EPA-650/2-73-053
DECEMBER 1973
Environmental Protection Technology  Series
                                             mmmwmmwwmmymm

                                                  I
                                                  1
                               UJ
                               a
                                                                      
                                                            fli:i

                                                   ^^

-------
                                                EPA-650/2-73-053



   PROCEEDINGS OF THIRD INTERNATIONAL

CONFERENCE ON FLUIDIZED-BED COMBUSTION
                           Sponsor:

                 The Environmental Protection Agency
                  Office of Research and Development
                National Environmental Research Center
                     Control Systems Laboratory
                      Program Element 1AB103

                    Project Officer: P.P. Turner
                  Chief, Advanced Processes Section
                   Clean Fuels and Energy Branch
                    Control Systems Laboratory
                         Prepared for
                  Office of Research and Development
                 U.S. Environmental Protection Agency
                      Washington, D.C. 20460

                         December 1973

-------
   These proceedings have been reviewed by the Environmental Protection Agency and approved
for publication. Except for minor editing for consistency of style, the contents of this report are as
received from the authors. Approval does not signify that the contents necessarily reflect the views
and policies of the Agency, nor does mention of trade names or commercial products constitute
endorsement or recommendation for use.

-------
                                    PREFACE

   The Third International Conference on Fluidized-Bed Combustion was held October  29-
November 1, 1972,  at Hueston Woods Lodge, RFD No.  1, College  Corner, Ohio, under  the
sponsorship  of the U.S. Environmental Protection Agency (EPA).

   The Conference,  under the general- and vice-chairmanship of Messrs. P.P. Turner and D.B.
Henschel, respectively, was open to early arrivers Sunday, October 29.

   The Conference proper, consisting of six sessions, got underway Monday morning with  the
official welcome extended by EPA's Robert P. Hangebrauck. Following Mr. Turner's introductory
remarks, Dr. A.M. Squires delivered the keynote address, "The Clean Fuel Technology Gap 
Opportunities for New Fluidization Procedures."

   The six-part Session I, chaired by Mr. A.A. Jonke, followed the theme, "Coal Combustion and
Additive Regeneration." Session II, held Monday evening,  consisted of five presentations on  the
general topic, "Non-Coal Fluidized-Bed Combustion Processes;" Mr. Alvin Skopp was chairman.

   Dr. Everett Gorin was chairman of Session III, titled "Gasification/Desulfurization," which
opened Tuesday's activities. The session consisted of six presentations.  "Conceptual Designs and
Economics" was the theme of Session IV. It was chaired by Dr D.H. Archer and consisted of seven
papers.

   Wednesday began with Session V which presented six papers on the topic, "Pilot Plant Design,
Construction, and Operation." The Session was chaired by Mr. H.B. Locke. Session VI, the last of
the Conference,  was a Wednesday afternoon discussion,  led by  a panel of six, summarizing
thoughts presented during the Conference and providing a final opportunity for comments from
the floor. Chaired by Mr. Hangebrauck, the panel consisted of Dr. Archer, Professor D.E. Elliott,
Mr. Henschel, Mr. H.B. Locke, and Dr. Squires.

   All papers presented during the Conference are included in these proceedings.
                                                                                     iii

-------
                        TABLE OF  CONTENTS



                                                                              Page

Preface	iii

R.P. Hangebrauck
Welcoming Remarks	0-1-1

P.P. Turner
Introductory Remarks 	0-2-1

A.M. Squires
Keynote  Address:  Clean Fuel  Technology  Gap:
Opportunities for New Fluidization Procedures    	0-3-1


                                   SESSION I

1.  G.J. Vogel, E.L. Carls, J. Ackerman, M. Haas, J. Riha, and A.A. Jonke
   Bench-Scale Development of Combustion and Additive Regeneration
   in Fluidized Beds  	1-1-1

2.  R.C. Hoke, H. Shaw, A. Skopp
   A Regenerative Limestone Process for Fluidized-Bed Coal
   Combustion and Desulfurization  	1-2-1

3.  R.L. Rice and N.H. Coates
   Combustion of Coals in Fluidized Beds of Limestone	1-3-1

4.  S.J. Wright
   The Reduction of Emissions of Sulphur Oxides and Nitrogen Oxides
   by Additions of Limestone or Dolomite During the Combustion
   of Coal in Fluidised Beds  	1-4-1

5.  A.A.Godel
   Selective Extraction of Clinker at the Bottom of a Deep Self-Agglomerating
   Fluidized Bed	1-5-1

6.  E.P. O'Neill, D.L. Keairns, W.F. Kittle
   Kinetic Studies Related to the Use of Limestone and Dolomite
   as Sulfur Removal Agents in Fuel Processing  	1-6-1

-------
                                                                                  Page


                                     SESSION II

 1. H.W.Schmidt
   Combustion in a Circulating Fluid Bed	II-l-l

 2. G.G. Copeland
   Disposal of Solid Wastes by Fluidized-Bed Combustion 	II-2-1

 3. G.L. Wade and D.A. Furlong
   Fluidized-Bed Combustion of Municipal Solid Waste in the
   CPU-400 Pilot Plant   	H-3-1

 4. B.J. Baxter, L.H. Brooks, A.E. Hutton, M.E. Spaeth, and R.D. Zimmerman
   Fluidized-Bed Combustors Used in HTGR Fuel Reprocessing	II-4-1

 5. W.E. Cole and R.J. Essenhigh
   Studies on the Combustion of Natural Gas in a Fluid Bed	II-5-1


                                    SESSION III

 1. J.T. Stewart and E.K. Diehl
   Fluidized-Bed Coal Gasification  Process and Equipment Development	III-l-l

 2. F.G. Schultz and P.S. Lewis
   Hot Sulfur Removal from Producer Gas 	III-2-1

 3. D.H. Archer, E.J. Vidt, D.L. Keairns, J.P. Morris, and J.L.  Chen
   Coal Gasification for Clean Power Generation	III-3-1

 4. J.W.T. Craig, G. Moss, H.H. Taylor, and D.E. Tisdall
   Sulphur Retention in Fluidised-Beds of Lime Under Reducing Conditions	III-4-1

 5. G.P. Curran and E. Gorin
   The COj Acceptor Gasification Process  A Status Report 
   Application to Bituminous Coal  	III-5-1

 6. G.P. Curran, B. Pasek, M. Pell, and E. Gorin
   Low-Sulfur Producer Gas Via an Improved Fluid-Bed Gasification Process	 III-6-1


                                    SESSION IV

 1. D.E. Elliott and MJ. Virr
   Small-Scale Applications of Fluidized-Bed Combustion and Heat
   Transfer (The Development of Small, Compact Fluidized-Bed Boilers)  	IV-1-1

vi

-------
                                                                                 Page
2. D.L. Keairns, W.C. Yang, J.R. Hamm, and D.H. Archer
   Fluidized-Bed Combustion Utility Power Plants  Effect of Operating
   and Design Parameters on Performance and Economics   	IV-2-1

3. B.R. Dickey and J.A. Buckham
   Application to Combined Cycle Power Production of Fluid-Bed
   Technology Used in Nuclear Fuel Reprocessing  	IV-3-1

4. A.N. Dravid, CJ. Kuhre, and J.A. Sykes, Jr.
   Power Generation Using the Shell Gasification Process	IV-4-1

5. R.A. Newby, D.L. Keairns, E.J. Vidt, D.H. Archer, and N.E. Weeks
   Fluidized-Bed Oil Gasification for Clean Power Generation 
   Atmospheric and Pressurized Operation  	IV-5-1

6. F.L. Robson
   Fuel Gasification and Advanced Power Cycles  A Route to Clean Power	IV-6-1

7. C.W.Matthews
   A Design Basis for Utility Gas from Coal	IV-7-1


                                    SESSION V

1. G. Moss and D.E. Tisdall
   The Design, Construction, and Operation of the Abingdon Fluidised Bed Gasifier	V-l-1

2. M.S. Nutkis and A. Skopp
   Design of Fluidized-Bed Miniplant	V-2-1

3. D.H. Archer, D.L. Keairns, and  E.J. Vidt
   A Pressurized Fluidized-Bed Boiler Development Plan	V-3-1

4. E.F. Sverdrup, J.R. Hamm, W.E. Young, and R.L. Strong
   Gas Turbines for Fluid-Bed Boiler Combined Cycle Power Plant	V-4-1

5. H.B. Locke, H.G. Lunn, and A.G. Roberts
   Combustion of Residual Fuel Oils in Fluidised Beds	V-5-1

6. H. Harboe
   Fluidized-Bed Air Heaters for Open and Open/Closed Gas Turbine Cycles	V-6-1

                                    SESSION VI

1. D.B. Henschel
   Minutes of Panel Discussion and Summary Session	VI-1-1

APPENDIX  Attendance List	

                                                                                    vii

-------
WELCOMING REMARKS:
  Mr. R.P. Hangebrauck, EPA

INTRODUCTORY REMARKS:
  Mr. P.P. Turner, EPA

 KEYNOTE ADDRESS:
  Dr. A.M. Squires, City College of the City University of New York
                             0-0-1

-------
Welcoming Remarks
                                         1.   THIRD  INTERNATIONAL
                                                        CONFERENCE   ON
                                      FLUIDIZED-BED  COMBUSTION

                             R. P. HANGEBRAUCK
                    Environmental  Protection  Agency
                        Control   Systems   Laboratory
   I feel honored this morning to welcome on
behalf of EPA all those attending the Third
International Conference on Fluidized-Bed
Combustion. A special welcome goes to those
of you here from other countries, old friends
and new alike.

   Each time at the end of our Conference we
have  asked  the attendees   and ourselves
whether it was worthwhile to assemble as we
have here in the hills of Ohio, and each time I
believe we have come up with a very positive
yes. Aside  from  the up-to-date information
exchange, it seems  that a  critical mass  of
expertise is achieved at this meeting causing
significant  reaction  to occur  during  the
meeting and long after.

   The aim in our work on fluidized-bed com-
bustion  and in our Conference is to develop
environmentally  and   economically  sound
systems for steam and power generation which
will enable  them  to  meet  new  source
performance  standards and  ambient   air
quality standards for sulfur oxides, nitrogen
oxides, and particulates, as, provided by the
Clean Air Act. Such systems must be  com-
patible  with these  and other present  and
future   environmental  considerations,
including  water,  land,  heat,   and noise
pollution.

   A variety of technologies  for  controlling
pollution from stationary combustion sources
 will be forthcoming  and should  fit together
 like pieces of a complex puzzle to solve the
 applications  problem in  a  cost and  time
 optimized  fashion.  The  technology  we are
 working on   here today is  most  directly
 applicable to the power industry,  but because
 this technology will allow the use of lower cost
 dirty fuels and problem fuels, it will free clean
 fuels for smaller residential, commercial, and
 industrial  fuel users constituting area-type
 combustion sources.

    The projected application of  technology
 and fuel resources will be such that, initially,
 part of the utility clean-fossil-fuel energy gap
 will be filled by increased use of low sulfur
 coal, physically cleaned  coal,  low sulfur and
 desulfurized oil, and  most critically, flue gas
 cleaning systems which will act as a counter
 balance to prevent the demand pressure for
, clean and/or cleaned fuels from driving fuel
 prices too high. From now into the 1980's, the
 increased demand for electrical energy  gen-
 erated  from  fossil  fuel, the shortage  of
 naturally occurring clean fuels in the users
 locations, and  improved  economics for flue
 gas cleaning should cause a great expansion in
 the application of these systems. High-Btu gas
 and liquids from  developing  coal  conversion
 systems will be limited to combustion sources
 considerably  smaller  than utilities,  whose
 consumer classification is such that it justifies
 the inherent,  much higher  price.  These
                                        0-1-1

-------
 processes should be commercially significant
 in the 1980's in filling the gas and petroleum
 gaps,  but they will  have their  own  set  of
 environmental problems.
   We realize that even though the needs for
 effluent cleaning will be satisfied  by such
 measures, the  ultimate  extent  of environ-
 mental control and lowest cost will  not have
 been  achieved.  Considering the  size  and
 growth of the power  industry and the multi-
 billion dollar annual control costs  involved,
 the public, government, and industry alike can
 no doubt  see the potential payout  for more
 effective, lower cost technology.

   Fluid-bed combustion systems hold great
 promise for reducing or eliminating the excess
 costs  created  by the technology gap.  The
 systems unfortunately are not here today, but
 are  under development and  nearing demon-
 stration, as will be evidenced by the nature of
 our   meeting  here  this  week.  We  feel
 reasonably confident the technology will be
 available on a commercial basis  as  we move
 into the end of this decade. Aside from the
 built-in  low-pollution nature of such systems
 for several pollutants, their cost effectiveness
 should cause a rapidly expanding share of the
 U.S. power-generating base to  be  filled by
 straight fluid-bed  combustion systems  and
 advanced  power  cycle  gasification systems
 incorporating  limestone/dolomite  fluid-bed
 technology. To  this end, EPA has invested
 considerable funds over the last five years, an
 investment which should  be  dwarfed,  if
 successful, by the payout to the public and
 industry for more cost-effective use of fuels in
 generating power.  I trust the progress we will
 hear reported this week will bear this out.

   EPA  program plans rely on  an increased
 shouldering of the cost of development and
 demonstration  by industry as the scale and
 cost of systems  increases in the final stages of
 development. Considerable  effort has  been
 made  in  the  EPA   sponsored  work   to
 concentrate   on  the   most   promising
 approaches, and this has been caused in no
 little way by the  small  amount of funding
 available. However,  it  is  hoped  that  each
 promising  avenue will be  explored  at  least
 somewhere in the world. It is also hoped that
 all information  available can and will be used
 by other groups in a way that will prevent large
 expenditures of funds on scale up of systems of
 questionable environmental merit.

   In any event, if we are to avoid getting into
 an expensive technology gap as new and more
 restrictive  emission   standards  are  set
 consistent with  the health and welfare of this
 Nation and the world community, we will have
to continue to look and move ahead with vigor
on  more productive  and  pollution-free
processes.

   Once  again  we  welcome  you  to  the
Conference and encourage the fullest partici-
pation  possible.

   Thank you.
0-1-2

-------
Introductory Remarks
                           2. EPA-CSL PROGRAM TO CONTROL
                 POLLUTION FROM STATIONARY SOURCES

                                 P. P. TURNER

                     Environmental Protection Agency-
                         Control Systems Laboratory
   I would like to take this opportunity to
repeat Mr.  Hangebrauck  and  once again
welcome  each of you  to  the  Third Inter-
national  Conference   on  Fluidized-Bed
Combustion.

   As was the case  with the two preceding
International Conferences, the purpose of this
meeting is to bring  together workers in the
field of fluidized-bed combustion, and related
areas, in  an  atmosphere  conducive to co-
operative information exchange. In addition,
we  have  a   number  of  organizations
represented  here  which,  although  not
currently directly involved in the research and
development effort,  will become involved in
manufacturing  or  operating fluidized-bed
boilers and  their auxiliaries, in designing
fluidized  boiler plants, or in performing some
other necessary function, when this promising
new combustion technique  is  commercialized
in the future. It is hoped that by getting all of
you  together,  discussing  your  individual
projects,  and  applying  your  individual
expertise  in this area, each of us can go home
better informed of the  overall  international
effort and with new ideas for direction of our
individual efforts.
   A  large  amount of  work has been
completed since the Second International
Conference was held over two years ago.  For
our part,  EPA has now  spent a total  of $7.6
million on work  related  to fluidized-bed
combustion through the end of fiscal  year
1972, of  which  about  $5.2  million  was
committed since the last Conference here at
Hueston Woods. Also during the  past two
years, design  has  been completed  and  con-
struction begun on  a  630-kW continuous
fluidized-bed combustion pilot plant, capable
of operating over the full range of conditions
of interest, including  pressure.

  This Miniplant is designed to continuously
regenerate the  partially-spent SO  2 control
sorbent generated in the fluidized combustor,
and  return the  regenerated  material to the
combustor for re-use. Studies on the Mini-
plant will enable EPA to obtain answers for a
number of important outstanding questions,
and will provide the continuous  combustion-
sorbent regeneration-data required  to design
the 20- to 30-MW pilot plant envisioned as the
next stage of the development effort.

  Also within the past two years conceptual
designs have been completed for a 30-MW
industrial coal-fired fluidized  boiler, and for
300-MW and  600-MW utility-scale  coal-fired
fluidized boilers  both at atmospheric pressure
and  at 10 atm pressure.

  EPA's emphasis has turned toward fluid-
dized boilers  operating at  elevated furnace
pressures,   although   atmospheric-pressure
systems  are  also  being  evaluated.  We are
taking a close look at continuous regeneration
of sorbent sulfated  in the combustor,  but
operation with a once-through sorbent  non-
                                       0-2-1

-------
regenerative system has  not  been ruled  out.
During the past two years, a fair amount of
evaluation has been conducted,  and data have
been generated, regarding pressurized oper-
ation and sorbent regeneration, in preparation
for,  and  complementing, the  forthcoming
Miniplant program.

    Development of the chemically active fluid -
ized-bed  (CAFB)  atmospheric-pressure  add-
on residual oil gasification system, currently in
the 750-kW stage, has advanced to the point
where utility  partners are being sought to
provide   a  site  for  a   82-MW  prototype
installation.

    Work  is also continuing  at an increased
scale  toward the  development  of a high-
temperature  fuel-gas  desulfurization  process
to produce clean low-Btu fuel  gas from caking
bituminous coals.

    All of this research and development effort
sponsored by EPA will be described during the
course of this meeting by the individual con-
tractors involved.  We also look  forward, of
course, to the  discussions  of work being
conducted by organizations other than EPA in
these and related areas.

    Expressing the hope that  the  activities of
this conference during the next three days will
lead to the healthy exchange of information
between all the participating members, and to
numerous individual contacts, I declare  this
Third International Conference on  Fluidized-
bed Combustion open.

   It is now my pleasure to introduce our key-
note speaker who will address the  conference
on "The Clean Fuel Technology Gap:  Oppor-
tunities for New Fluidization Procedures." He
will review the problems which one sees ahead
in the natural gas and petroleum markets; he
will characterize the substitute technologies
which  will be  needed; and he  will  review
advantages of new fluidization procedures for
treating coal and residual  oil.
   He is indeed well qualified to set the tone
for this conference.  Though his formal back-
ground  is  in   chemistry,   his  interest  in
engineering was aroused during World War II
through his  association with  Dr.  Manson
Benedict,  whom he assisted in the process
design  of  the Oak Ridge  gaseous diffusion
plant.
   After the war he was Director  of Process
Development at Hydrocarbon Research, Inc.
until 1959, when he resigned to  become  an
independent consultant. He joined the faculty
of the Department of Chemical Engineering of
The City College of The City University of New
York  in  September 1967,   and was named
Chairman of that Department  in the fall of
1970.

   He  has published  extensively  on  fluid-
ization,  oil  and   coal  gasification,  fuel
desulfurization, gas cleaning, and power gen-
eration; he  holds  15 U.S.  patents  in these
fields. He has conducted research at The City
College under  EPA grants  relating to  the
development  of new systems for  generating
clean power from fossil fuels. His team at The
City College began work last June on the first
18 months of a 5-year  effort under a grant
from  the  National  Science Foundation  to
support "studies toward improved techniques
for gasifying coal." The  objective  of these
studies is to convert coal into pipeline gas and
a light  aromatic liquid fuel  as well as low-Btu
gas for  power generation.  Gentlemen,  I
present to this conference  Dr.  Arthur  M.
Squires.
0-2-2

-------
Keynote Address
                             3.  CLEAN FUEL TECHNOLOGY  GAP:
                   OPPORTUNITIES  FOR NEW  FLUIDIZATION
                                                               PROCEDURES
                                 A. M. SQUIRES
           The City College of the  City University of New York
   Before we take up the specific problems of
 interest to this Conference, we may well first
 view these problems in the broader context of
 the markets for clean and convenient energy.
 To do so even briefly creates a sense of urgency
 and  a demand  for boldness.  The World's
 appetite for clean fuels is sharply rising. That
 the World's resources of cheap clean fuels are
 finite is  an inescapable fact, and economic
 consequences of this fact are beginning to be
 felt.

   Consider Figures 1  and 2,  which  depict
 broadly the supply and demand situation for
 oil and  natural  gas  in the United  States
 between  1955 and 1985. The range of uncer-
 tainty  in  demand  beyond the  present  is
 approximately the range among recent pro-
 jections for 1985.

   Figure 1 implies  that America's economic
 growth is in jeopardy if we cannot import from
 one -half to two-thirds of our oil  supplies in
 1985.  Imports  to  reach  the  upper  curve
 amount to substantially the entire present out-
 put from  the Middle East. Notice that oil from
 the North Slope has small relative effect. We
 would  need to discover a North  Slope each
 year between now and 1985 to reach the upper
 curve of  Figure 1 from domestic  supplies.

   Gas cannot be imported from  overseas as
 readily as oil, and Figure 2 implies a sharp
 flagging in the growth of the gas market.  We
 are using a quantity of gas yearly that is
 greater than the average annual  discoveries of
 gas made over the past 20 years and more than
twice the discoveries of last year. No doubt
discoveries  can be  increased by  removing
artificial restrictions on the price of gas at the
wellhead, but no prompt effect on gas supplies
could result. It takes several years to bring a
new field to production, and in the meanwhile
old fields  decline. Nuclear stimulation  is  a
doubtful proposition in light of concern over
spread of radionuclides. All substitute natural
gas from sources now in view will cost at least
about $1 per 106 Btu,  and often more.  This
includes gas from the Far North. It should be
remarked that not enough is  known yet of
North Slope  gas  reserves and production
problems to project with assurance its supply
to the United  States market, yet some gas
from the Arctic can probably be delivered by
about  1980  with  vigorous  development,
Canada  willing.

   Since  much of the historic growth  in
demand  for gas has been for fuel to  fire
boilers, a projection of the demand for clean
boiler fuel would reveal a much greater gap
between  supply and demand than  Figure  2
would suggest. The Nation wishes to eliminate
emissions from boilers fired  with  coal and
untreated residual oil.  Satisfactory  engineer-
ing solutions to the problem of ridding stack
gases of sulfur'dioxide are not yet in hand, and
the problems  of the large coal-fired power
stations projected for the Southwest illustrate
the dislocations of the forced shift from gas to
coal in new plant construction.

   An ironic illustration of the sharp change
in the gas market is furnished by the news that
                                        0-3-1

-------
 Coastal States Gas Producing Company  will
 build a plant to manufacture synthetic natural
 gas  from  petroleum feedstocks in Texas.

   It is hard to escape the impression that our
 energy markets will undergo price upheavals
 in the next decade. In view of this, it must
 seem astonishing to  a  layman  that our  fuel
 industries are so little prepared with substitute
 technologies. For example, to convert coal  into
 clean  gaseous  or  liquid  fuels,  they  can
 absolutely rely, for immediate construction,
 only upon  technology  introduced  nearly 40
 years  ago to fuel the German war  machine.
 Moreover, nothing better will be ready to have
 much effect in the time span of Figures 1  and
 2. Our research and development efforts, both
 private and governmental, have been far too
 inadequate.  The layman must  be further
 astonished  to  learn that private efforts  are
 being reduced.

   At least five major oil companies have shut
 down  laboratories  or  layed  off  personnel
 engaged  in synthetic  fuels  research.  One
 architect-engineering firm that caters to  the
 fuels industries has shut down an historically
 important laboratory, and  other  such firms
 appear  to   be  decreasing  their  research
 budgets. Yet even a Manhattan Project on new
 fossil fuel technologies could not be expected
 to make  a  major dent  in the clean fuel  gap
 before  1985. It  is  highly  improbable  that
 synthetic  liquid fuels made from coal  could
 play a significant role before  then.  Most
 substitute natural gas will be  made from
 petroleum feedstocks, which will themselves
 become short in supply. The extent to which
 we   have  been  minding  the ^ store,  in this
 particular respect, is  ironically illustrated by
 the fact that we will license the SNG processes
 for petroleum feedstocks from  Great Britain,
 Japan, and  Germany at fees  that will total
 more than $100 million.

   It will be important to understand how the
 clean fuel technology gap has arisen. To what
 degree has research and development failure
 been due  to governmental interference with
 "normal"  economic   processes?  To  what
degree is  it  due to size and maturity of  our
energy industries and  concomitant risks and
high costs of research  and development for
these industries? Is it due to a general loss of
appetite for risk-taking among the managers
of our technology?  Is it  due to  a  general
migration of creative talent, promotional and
managerial  as well  as technological,  into
glamorous activities such as the space effort?
Are there other factors?

   Understanding  these  questions  will  be
important, not only  so that our domestic
arrangements for fuels research and develop-
ment may be overhauled, but also so  that we
may better prepare ourselves for an even more
serious clean fuel technology gap which lies in
the not distant future.

   Figure 3  gives  a   gloomy but plausible
scenario  for the future course of  the world
petroleum market. If the attitude of the World
toward oil parallels that of the United States
toward  natural  gas,   something  like  the
scenario  of Figure 3  will  inevitably  unfold.
Substitutes for oil  will be developed too late;
production of oil will  reach the  limits of the
World's capability to yield oil before sufficient
experience with substitute  technologies has
been  acquired;   growth  of  technologies
dependent upon oil will be choked off; and
economic  disarray as well  as  insufficient
experience  will prevent   rapid  growth  of
substitutes.

   The historic excess  in  our  capacity  to
produce gas, seen  in Figure 2, was  important
to the  petrochemical  industry in the United
States. Disappearance of the excess capacity,
along with unwise import regulations  for light
hydrocarbon feedstocks, is creating serious
difficulties for this industry, which illustrate in
miniature the dampening effect that  the dis-
appearance of excess  oil production  capacity
will have upon invention and  development of
new technologies for better use of oil.  Existing
equipment and existing technologies will pre-
empt supplies,  and opportunities to divert oil
from uses of lower to higher value will be
missed.
0-3-2

-------
   Electricity from whatever source (nuclear
fission or fusion, solar, or geothermal) cannot
be readily substituted for clean, portable fuels.
Estimates of the U.S. fuel mix in the year 2000
postulate only  about 25  percent nuclear at
best. For the distant future, electricity  could
electrolyze water to yield hydrogen. This could
be used directly as  a fuel, stored either as
liquid  or  at  high  pressure  or  reversibly
adsorbed  upon  a  solid.  Alternatively,
hydrogen and carbon dioxide could  be con-
verted to hydrocarbons by  Fischer-Tropsch
synthesis, or hydrpgen and nitrogen could be
converted to ammonia. By what steps and over
what kind of time span might such a synthetic
fuel  economy  be introduced?  In  an  early
transitional stage,  electrolytic hydrogen might
add to the heating value of carbon-rich fuels of
natural origin,  converting  them to lighter
materials. Even earlier, if natural gas is  still
available in regions like  Venezuela and  the
Middle East, hydrogen  might be made from
the gas and added to carbon-rich  fuels.

   To what extent can the world  rely upon
synthetics based upon coal? It should be noted
that the Northern Hemisphere is much richer
in  coal  than  the  southern.  Even in  the
Northern   Hemisphere,  coal  deposits  are
concentrated  in  a  few  favored  countries
(notably the   United  States   and  Russia,
especially the latter). If SNG from Texas  is a
surprise, what  about  exports  of  synthetic
liquid hydrocarbons from Wyoming? Our coal
resources   seem  vast,  but  would  they if
measured against the projected world appetite
for oil in the 21st Century?

   What other substitute technologies based
upon electricity can replace liquid fuels? Can
nuclear energy  (or solar  or geothermal)  be
increased beyond  present projections before
the  year 2000?  On  the other  hand,  what
happens if our second try for a  liquid-metal-
fast-breeder-reactor is a flop, as was the first?

   Can solar-pond algae plausibly contribute
to the gap? Hydrocarbons based upon human
and  animal wastes? Cellulose?  What about
the "substitute" of making do  with less? If
there must  be a  belt-tightening anyway, it
would be better earlier than later.

   Can  a  combination  of  these  or  other
developments, carried out in a timely manner,
produce the more attractive scenario of Figure
4?
   Figures 3 and 4 carry no scale for  time.
Hubbert has drawn his celebrated  "dimple"
for two estimates of "recoverable"  world oil
reserves (meaning recoverable at costs not far
advanced from current costs): 1350 and 2100 x
109 barrels. At the lower figure, production
begins  to  depart   from  a   substantially
exponential rate of growth'at around 1980 and
reaches a peak at about 66 x 106 bbl/day
in about  1990. At Hubbert's  higher figure,
he  projects a "dimple" with   substantially
exponential growth until about 1990 and a
peak production at about 102 x 106 bbl/day
shortly after the year 2000. Current produc-
tion is nearing 50 x 106 bbl/day, and appears
to be ahead  of Hubbert's projections.

   At Hubbert's lower estimate of reserves, the
crisis in Figure 3 could come as early as about
1985. At the higher estimate, about 2000. This
is assuming at least modest cooperation from
both  oil-producing nations  and authorities
responsible for leasing off-shore drilling sites.

   The key to the scenario of Figure 4 is timely
research  and  development  and  timely
construction of full-scale installations in which
to practice the new technologies. It is already
too late to prepare for  the  world clean fuel
crisis if it should appear as early as  1985.
There is just barely time, perhaps, to get ready
for  problems  that  might  reach  critical
proportions  in 2000. Thirty years is not  too
long to develop and test a new technology and
to learn it sufficiently  well that it may be
expanded as rapidly as the scenario of Figure
4 will require.

   Energy decisions made in the United States
in the next  few years can be crucial in  the
choice of scenarios. That of Figure 4 may well
imply  the lower  curves of  projected  fuel
demand  in  Figures 1  and 2  as  well  as a
                                                                                     0-3-3

-------
 Manhattan Project to  acquire  good  tech-
 nologies for converting coal to gas and  oil.
 Simultaneously, we must not neglect vigorous
 attack on  the  substitute  technologies whose
 commercialization must begin in the 1980's if
 they are to be ready to forestall a world energy
 crisis.
 OPPORTUNITIES   FOR   AGGLOMER-
 ATING AND FAST FLUIDIZED BEDS

   Our first chapter has raised more questions
 than it has provided answers.

   My second chapter, turning to the  interests
 of this Conference must inevitably be a more
 particular response to the urgent call for  new
 fossil fuel technologies.

   My colleagues at The  City College and I
 believe that  the time  is  coming soon when
 economics  will turn against the practice of
 burning chemically-bound hydrogen for large-
 scale production of electricity.  Instead,  the
 bound hydrogen in coal or residual oil will be
 viewed  as too valuable to burn and  send as
 water vapor  up a stack.  The hydrogen  can
 become a part of some clean, convenient  fuel
 having  an economic value higher than coal's
 or residual oil's. The hydrogen-rich fuel would
 be  "creamed off  the coal or residual  oil
 leaving a carbon residue that would be burned
 to generate heat or  electricity.

   The idea of creaming off valuable products
 from coal or  oil is of course not original with
 us. The byproduct coke oven is 100 years  old;
 many attempts to  displace it with improved
 coking procedures have been recorded;  and
the oil industry has made steady advances in
technologies for reducing the yield of residue
and increasing yields of lighter products. Most
attendees at this Conference will  be  familiar
with Consolidation Coal Company's efforts, as
well  as the recent  achievements  of the FMC
Corporation.

   Any   good  idea, however,   can  stand
constant review  in  light of the appropriate

0-3-4
 technological context. We have tried to review
 the  foregoing idea in  light of the ongoing
 development to supply the military with better
 gas turbines, funded at about $300 million per
 year. Progress in engines for civil aircraft and
 stationary  power  has  historically  followed
 military achievements after a lag of only a few
 years. The 747 flies today with a temperature
 of  1300C at the  inlet  to  the turbines.
 Stationary machines larger than 100 MW are
 promised  for  this temperature before  1980.
   Existence  of such   machines  will allow
design  of electric  power   installations  of
sharply advanced  efficiency  and  reduced
capital cost. Gas turbines (inherently cheap in
cost) will supply about 50 to 60 percent of the
power,  and a steam system  (using  modest
steam conditions  and  a  cheap  boiler)  will
scavenge heat from gas turbine exhaust.
   The  prospect of these designs creates an
 imperative to develop  better techniques  for
gasifying coal and residual oil to provide a
 cheap "power gas" to run the gas turbines
 i.e., a low-Btu gas  made using air and a little
 steam as the gasification medium. (Power gas
was  Ludwig Mond's term  for  producer  gas,
which he  used to  generate  electricity in gas
engines.  He   founded  The   Power   Gas
 Company.)

   The  City College team believes that power
gas technologies will inevitably evolve to allow
some cream-skimming  that pulls out  fuel
products of value greater than coal or residual
oil. Nevertheless, power gas technologies will
probably arise in the first instance for treat-
ment of raw fuels. It is  doubtful that at first
any more than very slight consideration can be
given to the potential of these technologies for
profitable   evolution.  Let  us first consider
power gas - technologies for  raw fuels,  and
second,  how they might evolve.
PRODUCTION OF POWER GAS
   Among gas-to-solid contacting procedures,
fluidization will be the strongest candidate for
employment in better technologies for gasi-
fying coal or residual oil.  To a large degree,
this is so  simply because  of the scale  of the

-------
power stations to be built in the  1980's and
beyond. Sites for over 4000 MW are already in
service. Many more will be built. Typical sites
will process coal, for example, on the scale of
scores of thousands of tons per day; oil, at
hundreds of thousands of barrels  per day.
Only  fluidization   procedures  can  readily
provide equipment of capacities that will avoid
the necessity of processing  coal or  oil in an
unattractively large number of vessels oper-
ating in  parallel.
   Although the science of the gravitating bed
is far ahead of our knowledge of fluidization,
the former art is difficult  to build in large
capacities. The blast furnace, after more than
a century of development, gasifies up to about
4000 tons of coke per day, but coke of course is
a processed and closely-sized solid. More than
20 Lurgi gravitating-bed  pressure gasifiers of
the current design  would be needed for 1000
MW; scale-up of this approach may prove dif-
ficult and uncertain. Although the gravitating-
bed had  significant yield advantages, it did not
win  in  competition with  the  fluidized-bed
catalytic cracker because it could not easily
reach capacities appropriate to the scale of oil
processing after 1960.
   Gasification of either coal or oil in absence
of a  bed  of  solid  appears plagued  with a
carbon loss  problem which  may  inevitably
require  extra  equipment for extinction of
carbon.  Either Texaco or Shell "partial oxi-
dation" of residual oil must provide for carbon
recovery  and  recycle  to  achieve  complete
carbon  utilization.  Nothing  was  published
from  Texaco's large experiment with  a slag-
ging,  dilute-phase gasifier at Morgantown in
1957, but  experience  at  Bell and elsewhere
suggests  that carbon utilization  in such a
gasifier may be poor, especially for a coal with
a refractory ash.

   I  have  not  seen any  advantage in gasi-
fication of coal in a pool of iron, a procedure
of doubtful integrity at elevated pressure and
doubtful operability at  atmospheric  pressure.

  The City College view is that the  strongest
candidate for  gasifying coal to obtain power
gas will combine the ash-agglomerating fluid -
ized bed about which we will hear from Godel
and the circulating fluidized bed that Schmidt
will describe. The  combination would operate
at about 2000 F and about 10 ft/sec velocity.
A single vessel could easily handle coal for
1000 MW;  at  20  atmospheres, the diameter
would be less than 20 feet. The role of the ash-
agglomerating bed would be to gasify large
particles of coal, up to about 3/4-inch  as well
as to agglomerate and  separate ash  matter
from the carbon-rich bed. At The City College
we have re-dubbed  Lurgi's highly expanded
circulating  fluidized bed the  "fast fluidized
bed;" we have a two-dimensional fast  bed of
plexiglas in operation that is exciting to watch.
The fast bed in the gasifier combination would
gasify fines and would provide a zone of high
velocity  and  intense  circulation for  intro-
duction of a caking coal. It is reasonable to
hope that a caking coal could be successfully
introduced  into the fast bed, fine  particles
joining  the bed and a  large  particle being
coked sufficiently  on its surface to render it
harmless before it reaches the ash-agglomer-
ating bed below.
   Data by  Dent,  which I have discussed  in
"Role  of  Solid  Mixing  in   Fluidized-Bed
Reaction Kinetics" to appear  in  an  AIChE
Symposium Series volume, strongly suggest
that the kinetic performance of the proposed
combination will be excellent.  Because of the
high  temperature   and  good  kinetic
performance,  flow of steam  will be  small
relative  to  air. Table  1  compares  Lurgi
gravitating-bed power gas with gas from the
proposed   gasifier  calculated  with  the
assumption  that steam-carbon  equilibrium is
substantially achieved therein, an assumption
suggested by Dent's experience. Table 2 com-
pares  electricity-generating  efficiencies   of
combined-cycle  installations   (with   power
equipment   according  to  United Aircraft's
"second generation" design) using the Lurgi
gasifier  with wet gas cleaning, or the  ash-
agglomerating-fast-bed   gasifier  with  wet
cleaning,  or  the  latter  gasifier with  gas
cleaning by the high temperature procedures
that we have under study at The City College.

                                      0-3-5

-------
Table 1. COMPARISON OF CRUDE POWER GAS
FROM THE LURGI GASIFIER AND  A CANDI-
DATE FOR DEVELOPMENT
Composition, % by vol
Methane
Carbon monoxide
Hydrogen
Carbon dioxide
Water vapor
Nitrogen
Hydrogen sulfide
Heating value, Btu/ft 3
Lurgi
gravitating-bed
gasifier
4.4
10.7
15.7
10.7
27.8
30.2
0.5
100.0
129
Gasifier combining
ash-agglomerating
and fast
fluidlzed bads
0.5
31.8
15.6
0.5
0.5
50.4
0.7
100.0
157
   The high loss of latent heat in the stack
from a combined-cycle plant that depends
upon the  Lurgi is inherent for this gasifier.
The loss arises from two sources: air supplied
to   the   gasification  bed   is   necessarily
accompanied by a high flow of steam in order
to limit the temperature and to keep the ash
free-flowing; and gas leaving the bed must be
quickly reduced  in temperature by a  water
quench in order to prevent formation of heavy
tars that would lead to deposits of coke.

   A fluidized-bed  gasifier   operating  at
2000 F will not make tars  or  tar-forming
species, and its   power gas  need  not  be
quenched.


   One cannot be so confident in respect to a
gasifier working at 1700F. The Wirikler did
not  make tars, but secondary oxygen was
introduced into the Winkler above the fluid-
ized bed in order to raise the temperature (to
reduce methane  yield?  to  eliminate  tar-
formers?). Late-model  Winklers  were  pear-
shaped and had enormous freeboard regions.
For  operation  at  1700F, a  carbon-burnup
step must be provided, as Pell will describe,
and  there will probably be a price  also in
capacity and in loss of latent heat to the stack,
as well as a possible problem with tar-formers.
 Nevertheless, until development of a gasifier
 exploiting   Godel's   ash-agglomeration
 principle is far enough along, an approach at
 lower temperature is well worth carrying as an
 alternate.
   Experience  at  Hydrocarbon  Research,
 Inc.and elsewhere suggests that defluidization
 phenomena might plague an  attempt  to
 develop a gasifier to  operate between about
 1800 and 2000F, unless perhaps the develop-
 ment were to adopt a fluidization velocity so
 high as to approach  the fast-fluidized state.
 The  limits  would   of  course  vary  with
 properties of coal ash, and  may extend below
 1800F for some important coals.

   I have  not yet seen  advantage  for  an
 approach  using an initial step carbonizing
 coal at about 1700F  followed by gasification
 at 2000F, the flows of solid and gas between
 the two steps  being countercurrent.  During
 this Conference we will perhaps hear evidence
 supporting an advantage for this  approach.
 However,  there would appear to be a sub-
 stantial risk that tar-formers will appear in gas
 from   a  carbonization  step  at   1700F,
 requiring a rapid quench  of the gas to prevent
 coke  laydown  in transfer  lines.  To avoid
 unnecessary degradation of heat, a quench is
 of course best avoided  except in a case where a
 useful quantity of liquid-fuel byproduct can be
 recovered.
   I  have  also  not   seen  advantage  for a
separate  ash-agglomerating zone  such  as
Je'quier provided. This was a zone  lean  in
carbon and was both  more dilute and hotter
than Je'quier's primary fluid bed. It afforded
complete combustion of air  furnished thereto,
and it served to density ash agglomerates and
to reduce their carbon content. If reduction of
carbon content in ash agglomerates  should
become a  problem, Godel  has provided two
simpler approaches. His  grate emerges from
the  bed,   and  agglomerates   thereon  are
exposed to air in absence  of  coke particles
external to the agglomerates. Godel has also
demonstrated burnup  in  a  gravitating-bed of
agglomerates from an anthracite of high ash
0-3-6

-------
           Table 2. ILLUSTRATIVE ENERGY BALANCES FOR COMPLETE POWER-GENERATION
                                        FLOW SHEETS



Category of energy
loss, %
Electricity sent out
Heating value of sulfur
Loss of sensible heat
in stack gases
Loss of latent heat
(water vapor)
Loss of heat at steam
condenser and elsewhere
Loss of unburned fuel
and heat leakage
Mechanical losses and
power for auxiliaries

Efficiency, allowing
credit for heating
value of sulfur

Conventional
steam power
equipment
without
recovery of
sulfur
39.5
-

5.0

3.8

47.7

2.0

2.0
100.0


39.5
Combined-cycle power equipment
with recovery of sulfur3
Lurgi gasif ier.
gas cleaning at
low temperature
45.0


4.6

14.1

28.4

4.9

2.0
100.0


45.5
Gasif ier combining ash-agglomerating
and f ast f luidized beds
Gas cleaning at
low temperature
49.1
1.1

4.5

5.6

35.7

2.0

2.0
100.0


49.7
Gas cleaning at
high temperature
50.5
1.3

4.7

4.5

35.0

2.0

2.0
100.0


51.2
a United  Aircraft  "second  generation"  design parameters.

-------
 content, the bed being supplied with air at the
 bottom.  If  carbon  burnup  resists  these
 solutions, a separate burnup step exploiting
 Battelle's carbon-lean ash-agglomerating bed
 for complete combustion would be preferable
 to Je'quier's arrangement.

   In light of work at Esso Abingdon by Moss
 and his colleagues, gasification in a fluidized
 bed containing lime emerges  as a  strong
 candidate for production of power gas from
 oil. Presence of tar-formers in the Abingdon
 gas gives cause for concern, however. Periodic
 burnout of coke  deposits  would  not  be an
 attractive procedure  at  high pressure. In a
 commercial  embodiment  of the Abingdon
 ideas,  could the cyclone on the gasification
 bed be dispensed with?  My thought would be
 to introduce secondary  air directly above  the
 gasification vessel at  the elevation where  the
 fuel gas enters the larger cross-section of the
 boiler  being served. For production of power
 gas at high pressure, a comparable idea would
 be to introduce secondary air above the fluid-
 ized bed, raising the temperature of the power
 gas  (eliminating  tar-formers?  as   in   the
 Winkler?). If this does not work, we might be
 forced  to  supply  much  more air  than
 Abingdon uses to a gasification bed  working
 at high pressure, thereby producing  a leaner
 gas containing less hydrocarbon species. Heat-
 removal surface would  need to be provided.

 EVOLUTION TO THE FUELPLEX

   W.C. Schroeder published data in his U.S.
 Patent 3,030,297 (April 17, 1962) that seem to
 The City College team to provide a strong lead
 in respect to a candidate for the first  fuel-
 treating step  in  a  "Coalplex"  producing
 substitute natural gas, liquid  fuel, and elec-
 tricity.  Solihull and Bruceton long ago taught
that the treatment of  raw coal with hydrogen
 at elevated  pressure and  at temperatures
between about 1500 and 1800F can result in
 attractive  yields of methane at high  concen-
tration. In Solihull and Bruceton experiments,
the residence time of vapor product generally
 ran into the minutes. Schroeder's contribution
has  been to call  attention to the attractive
yields of benzene, toluene, xylene, and sub-
stantially  the  same  yield   of  methane,
accompanied  by very  little heavy tar  if the
vapor product residence time is kept  short,
preferably  around 5  seconds.  At the long
residence time  of the earlier  experiments,
aromatic products polymerized, condensed as
heavy tar upon the coke present in the reaction
zone, cracked to  form additional coke, and
vanished. By arranging for a rapid quench of
vapor  species  to about   700 F, Schroeder
preserved tar-forming aromatic  species and
thereby obtained an  attractive yield of a light
aromatic liquid.

   The  City College team sees Schroeder's
chemistry as opening up the possibility for a
Coalplex yielding  roughly  25  percent of the
coal's heating value  in form of methane and
perhaps  15  to 20  percent  as  BTX, with the
remainder being converted  to electricity at an
efficiency beyond  40  percent.  The  heat
degradation  that  results  from  Schroeder's
quench  is tolerable,  amply rewarded by ;the
yield of liquid  product  that  the  quench
preserves.

   For Western coals, the fast fluidized bed is
a candidate  device for conducting Schroeder's
chemistry. The  fine coke  product could be
circulated from the fast bed to a heating step
and   returned   to   sustain    the   slightly
endothermic Schroeder reactions.

   For  Schroeder's  chemistry with  Eastern
coals, the coke-agglomerating fluidized bed
that  I discussed here  during our Second Inter-
national  Conference  is a  candidate  device,
perhaps with a superposed fast fluidized bed
of a fine solid that  is circulated to  provide
heat.

   For either coal, the coke product could be
gasified in the aforementioned  combination of
Godel's ash-agglomerating  fluidized bed and
a superposed fast bed to deal with fines.

   The coke product  could  also be burned up
in a  fluidized bed boiler. Although The City
College team regards production of power gas
0-3-8

-------
to be the main chance, nevertheless we have a
healthy respect for the difficulties of gasifier
development,  and  the boiler  is  welcome
competition. Thanks to Godel's  35 Ignifluid
"gasifiers"  (to call them  "boilers" in this
context is misleading), the race at the moment
appears about even. 

   A counterpart of Schroeder's procedure as
applied to residual oil should yield even higher
quantities of methane and BTX and less coke.
Coke from oil treatment is practically certain
to contain negligible sulfur. In form of beads
roughly  1/12 to 1/2-inch in  size,  the  new
petroleum coke, after calcination, should find
lively markets for electrode manufacture and
metallurgy.

Whither Fossil Fuel Development?

   Our first chapter set down a gloomy picture
of fossil fuel research  and  development.  Past
efforts, both private and governmental, have
not been responsive to urgent needs, and the
former are  contracting.

   Our second chapter,  although it focussed
upon The City College's view of the future,
nicely  illustrates  the  necessity  for  sharply
expanded research budgets.

   In perhaps no engineering procedure other
than fluidization is art  so far ahead of science.
Neither   ash-agglomerating,   coke-agglom-
erating, fast fluidized beds,  nor other com-
peting procedures for  processing coal will be
developed  through  computer modelling or
"optimization" studies. Godel discovered his
ash-agglomeration phenomenon  in  1954
during  hands-on hardware  development of
processes  for making  activated carbon from
coal. Although  he had  no experience with
boilers, Godel promptly recognized a  new
capability  to  gasify  and  render  burnable
anthracite slacks from the  mine at  La Mure,
near his boyhood home at Vif,  Isere. With
help from Babcock-Atlantique, he had a small
boiler operating inside a year. Although the
fast  fluidized  bed arose from Lothar Reh's
work on fluidization in cones for his Dr.-Ing.
Degree, Lurgi had to carry the development
through frustrating difficulties as Reh's team
learned how to achieve the fast-fluidized state
in  large-scale, practicable  equipment.  The
origins of agglomerating beds that make dense
beads of agglomerated material are obscure to
me. Chance observation by Dorr-Oliver during
development of  a fluidized  roaster  for a
"sticky"   Katanga  ore  may   have  been
important.  Dorr-Oliver  had  a  commercial
process for calcining calcium carbonate slimes
before 1957, and a number of pilot operations
were  underway  at Dorr-Oliver,  Fuller,  and
Battelle before 1960.  Yet in 1972  we have
hardly begun to acquire a scientific knowledge
of these beds.
   A number of factors have joined to create
an  illusion, shared by public and  political
leaders, research managers, and alas far too
many  engineers, that  a  new development
requires only laboratory results and design of
appropriate equipment  based simply upon
scientific analysis. One factor has been the
dominance of the physics establishment in
selection of recent  national  R&D goals. A
second  factor  has been  emphasis  among
educators upon  engineering  science to  the
detriment of an appreciation for hands-on
development  of  new  engineering  hardware.
The latter has somehow not been  respectable
by  comparison  with   fancy  mathematical
analysis; it also  gets a  fellow dirty.  A third
factor, of special importance to the chemical
engineering profession, has been the success of
the  approach  exploited   so  brilliantly  by
Scientific Design Co., Inc.  immediately after
World War II;  viz.,  bench development of
fixed-bed catalytic processes followed directly
by commercial-scale equipment. A literature
has even grown  up  on the  theme, "the pilot
plant is obsolete." This situation may exist for
fixed-bed processes in which there is no risk of
minor unwanted  products undeteptable at the
bench yet troublesome in the  field,  but  this
special  case  should not  be  elevated to  a
philosophy.  Permit me a bit of personal
history. I got an altogether  false  idea of
                                                                                    0-3-9

-------
 engineering from my first job working on the
 design  and  startup  of the  first  gaseous
 diffusion  plant at Oak Ridge: this plant was
 built from bench data, was "scientifically de-
 signable," and was a justly celebrated success.
 1 learned the other side of engineering through
 my  participation  in  the  classic  flop  of
 fluidized-bed  process  development. The
 hydrocarbon  synthesis  reactors  failed  at
 Brownsville because the development lacked a
 timely pilot unit of adequate size. If  Dubie
 Eastman's 12-in.  pilot plant  at Montebello
 had operated 3 years sooner, we would have
 known. Our research budget had been too
 small.

   Budgets unresponsive to our  clean  fuel
 needs have led  to a  preference for "safer"
 experiments relying more heavily upon earlier
 experience  and  representing   a  "simpler"
 approach. We have learned that the "simple"
 is not always so: witness the misfortunes of the
 dry-limestone-injection  approach  and  the
 troubles of limestone scrubbing.  In the mean-
 time, for  lack of making ourselves  ready,
 unanticipated opportunities arise that cannot
 be   seized.   Studebaker-Worthington's
 Turbodyne and Southern  California Edison
 are showing how old steam turbines can be
 converted to a combined cycle by scrapping
 old boilers and adding gas turbines followed
 by new waste heat boilers. General Electric
 and   Westinghouse  report brisk  sales  of
 combined-cycle  equipment.  A  market  for
 power gas is developing right under our eyes,
 and  only Lurgi is ready for coal and, except
 Texaco-Shell, partial oxidation for oil.

   In relation to  the possibilities and  the
 urgency of our needs, progress in areas to be
 covered during this Conference must seem
 disappointing to any visitor at  Alexandria 5
years ago who admired  the remarkable hands-
on  hardware  development  already accom-
 plished  by Pope, Evans  & Robbins by that
date; or, to a visitor to Leatherhead more than
3 years ago who saw BCURA's feat for  burn-
 ing  coal   at  elevated  pressure  at  rates
 approaching 1000 pounds per hour. Fluidized-
 bed boiler development in the United States

 0-3-10
needed Esso's "miniplant" 3 years earlier and
10 times bigger, but our hosts at this Confer-
ence simply have not had available to them the
requisite funds. Nearly 3  years  ago at  the
Christmas AAAS meeting in Boston (and on
many occasions subsequently), I pointed  out
the potential value of Godel's ash-agglomera-
tion phenomenon in combination with the fast
bed in a maker  of power gas, and I said it was
a shame that we had no  commercial exper-
ience here with the phenomenon. It is  still a
shame. Our fossil fuel industries have not had
the simple  curiosity to buy a small Ignifluid
boiler, a commercial proposition on  which
nothing would be lost, to gain firsthand exper-
ience from  its operation. Our hosts here have
not had  available  funds to make good  this
omission.

   Let me close by quoting from F.C. Dent's
Melchett Lecture of 1965. This great develop-
ment engineer,  now enjoying life on his yacht
out of Malta, can reflect with pleasure on  the
scores of millions in royalties that  the  SNG
processes he developed at Solihull will bring
during the  next few  years  to the  British
economy. In terms of the quotation  to follow,
our own  efforts to develop SNG processes for
coal may be  said to be just beginning.

  It is significant. . .that we  usually had
  reason to regret  any protracted  period
  of exploratory laboratory investigation.
  Small-scale  experiments   have   often
  been time-wasting even when large-scale
  conditions  have  been  reproduced  as
  faithfully as possible. .  .Operation on  a
  reasonable scale at an early stage is most
  desirable to throw difficulties into their
  proper perspective. Laboratory work was
  of most  value after the problems had
  been  recognized  in this way.


   All technologies  addressed  toward closing
the clean fuel technology gap must include a
major materials-processing  step,  handling
solids, liquids, or gases on a scale almost with-
out precedent  in  chemical engineering  art.
Serious effort does not begin until this step is

-------
addressed by hands-on hardware development
on a practical scale. Budgets must  be big.
Programs must be bold.


APPENDIX

Energy   Studies  at  The  Department  of
Chemical Engineering of The City College of
The City University of New York

   Nine studies are in progress. All except III,
IV, and IX are supported by Grant GI-34286
from the RANN Program (Research Applied
to National Needs) of the National  Science
Foundation. Professors  Michael Gluckman,
Robert Graff, Robert Pfeffer, Reuel Shinnar,
and  Joseph  Yerushalmi;  Dr.  Norman
Holcombe; and Messrs. Samuel Dobner, Kun-
Chieh Lee,  Dennis Leppin, Jeffrey Silverstein,
Eugene Yu, Stanley Dobkewitch, and Nurettin
Cankurt are participating in the effort. In the
past, Drs.  Leon Paretsky,  Melvyn  Pell, and
Lawrence  Ruth,  and  Messrs.  Richard
Angiullo, Richard Earth, Ralph Levy, Basil
Lewris,  Michael  Somer, and Lauris  Sterns
made substantial contributions.  Messrs. John
Bodnaruk,  George   Dilorio,  Michael
Askenazy, and John Spencer have helped with
experimental arrangements.

   Research on Power Gas:

   I. Study of the Godel Ash-Agglomeration
Phenomenon.

   II.  Study of  Kinetics  of Carbon  Gasi-
fication in a Fluidized Bed.  Our objective is to
test F.J. Dent's  hypothesis that the superior
kinetics afforded by a fluidized bed for the
steam-carbon reaction are  a consequence of
solid  mixing  in the bed, bringing about
repeated exposure of carbon to fresh gasi-
fication medium.
   III. Study of Kinetics of Removal of Sulfur
Compounds from  Power Gas by Action  of
Calcined Dolomite.

   IV. Study of Removal of Fine Dust from
Power Gas by a Panel Bed Filter. The filter
can be built to clean gas  at temperatures
approaching 1800F. We have achieved clean-
ing efficiencies  beyond  99.99  percent  for
power station fly ash at a normal stack dust
loading in small-scale tests at  atmospheric
temperature.

   Research on the Coalplex:

   V.  Study of Reaction of Coal with Hydro-
gen at High Temperature and Pressure  and
Short Residence Time of Vapor Products. We
believe  this  reaction, disclosed by W.C.
Schroeder, can be  the first  coal-treating step
in a Coalplex shipping substitute  natural gas,
BTX,  and  electricity  at  costs  below  the
combined  cost  of making  each  product
separately from coal.

   VI. Study of Coke-Agglomerating Fluid-
ized Bed, a candidate device for conducting
Schroeder's chemistry  (as  in  V  above)  on
Illinois and other Eastern coals.

   VII. Study of the Fast  Fluidized  Bed, a
candidate device for conducting  Schroeder's
chemistry (as in V above) on Western coals;
also,  for  gasifying  fine particles of carbon
blown from an ash-agglomerating fluidized
bed.

   VIII. Flowsheet Studies  for the Coalplex.

   Research on the Oilplex:

   IX. Flowsheet  Studies for an Oilplex in
which  oil is  first  treated by  reaction with
hydrogen  at high temperature and pressure
and short residence time  of vapor products.
                                                                                  0-3-11

-------

 10
 &
    35
     30
     25
     20
     15
     10
            W|= IMPORTS
            ^"''  AND SYNTHETICS
            CONSUMPTION
   DOMESTIC
   PRODUCTION
,_ CAPABILITY
             "DOMESTIC
             PRODUCTION
                       FIRST OIL
                       FROM
                       NORTH
                       SLOPE
       1955
          1965        1975
               YEAR
 Figure 1. United States petroleum
 supply and demand.
1985
                                                   30
                                                   25
                                         S. 20
                                               3
                                               ce.
                                                f  15
                                                   10
                                                          I     DOMESTIC
                                                               PRODUCTION
                                                         ,-*  CAPABILITY^
                                                       CONSUMPTION
                                    FIRST GAS
                                    FROM NORTH
                                    SLOPE
                         I   DOMESTIC
                          N ppnnnrTii
                                                                    PRODUCTION
                                                                I = IMPORTS
                                                                  AND SYNTHETICS
                                                                      I
1955
         1965         1975
              YEAR
Figure 2.  United States natural gas
supply and demand.
                                                                                       1985
            l:: SYNTHETICS AND
                SUBSTITUTE ENERGY
                TECHNOLOGIES
       WORLD OIL
       PRODUCTION
       CAPABILITY
                        CONSUMPTION
                       WORLD OIL
                       PRODUCTION
                      TIME
       Figure 3.  World petroleum supply and
       demand:  a gloomy scenario.
                                                            CONSUMPTION
                                                            (INCLUDING SUBSTITUTES)
                                                WORLD OIL
                                                PRODUCTION
                                                CAPABILITY
                                                              WORLD OIL
                                                              PRODUCTION
                                                                TIME
                                               Figure 4.  World petroleum supply and
                                               demand:  another scenario.
0-3-12

-------
SESSION I:
  Coal Combustion and Additive Regeneration

SESSION CHAIRMAN:
  Mr. A.A. Jonke, Argonne National Laboratory
                               1-0-1

-------
                                1. BENCH-SCALE DEVELOPMENT
       OF COMBUSTION AND ADDITIVE  REGENERATION
                                                   IN FLUIDIZED BEDS

    G. J. VOGEL, E. L.  CARLS, J. ACKERMAN, M. HAAS, J. RIHA,
                              AND A. A. JONKE

                       Argonne  National Laboratory
 ABSTRACT
   This paper discusses information obtained since the last Hueston Woods Conference on the
 combustion of coal and oil with an excess of air and the combustion of coal in a deficiency of air.
 The paper is also concerned with the thermodynamics of several proposed regeneration processes
 and the regeneration of sulfur-containing additive by the two most promising processes  a one-
 step reductive decomposition of CaSO4 and a two-step (reduction-CO2/H2O) procedure.
 INTRODUCTION
   Research on the combustion of fossil fuels
 (particularly coal) in a fluidized bed of solids is
 currently under investigation  in the United
 States and other countries. Most of the U.S.
 effort is supported by the U.S. Environmental
 Protection Agency (EPA), Office of Research
 and Monitoring.

   In  applications of fluidized-bed combus-
 tion, fuel is burned in a fluidized bed of solids
 in which boiler tubes are immersed to take
 advantage  of the  high  heat-transfer char-
 acteristics  of fluidized beds. Additive, either
 crushed limestone or crushed dolomite, can be
 continuously fed to a fluidized-bed combustor
 to react with SO2 released during combustion
 and provide a means  of in situ control of the
 emissions of SO2.

   Two different combustion modes are possi-
 ble, one with complete and the other with par-
 tial combustion of the fuel in the fluidized
 bed. In the complete-combustion mode (also
 called  one-stage  or  oxygen-excess combus-
tion), oxygen in excess of the  stoichiometric
amount required to burn the fuel to CO2 and
H2O is added to the  fluidized bed. In the
second mode (called two-stage or oxygen-defi-
cient combustion), a stoichiometric deficiency
of air is added to the fluidized bed, and the
resulting H2, CO, and hydrocarbons are com-
busted  to  CO2 and H2O  by  providing
additional oxygen (air) either  in the region
above the bed or in a separate combustor.

  At Argonne  National Laboratory (ANL),
data have been  obtained  on combustion of
coal and oil  with an excess of oxygen and on
coal with a  deficiency  of oxygen. All experi-
ments have been  made  at  atmospheric
pressure. The objectives of these experiments
have been as follows:

  1. To determine  how sulfur retention  is
    affected by  independent  fluidized bed
    operating variables such as bed temper-
    ature, gas velocity, oxygen  concentration,
    bed height, calcium to  sulfur ratio, type of
    additive and coal,  and additive and coal
    particle  size.
                                      1-1-1

-------
   2. To determine the level of NO in the flue
     gas at different operating conditions.
   3. To obtain information on combustion effi-
     ciency,  combustion  products, limestone
     utilization, extent of calcination, decrepi-
     tation rates.
   4. To obtain information on the mechanism
     of the lime sulfation reaction.
    Efficient removal  of SO2  from the  gas
phase requires moderately large quantities of
limestone (compared to the quantity of coal
ash). When coal is burned, it will be desirable
to regenerate the partially sulfated lime. Ther-
modynamic  calculations  and  experimental
data are presented on the two most promising
reactions; i.e.,  high temperature  (~2100F)
reductive decomposition of CaSO4 and a two-
step process,  low  temperature   (~1600F)
reduction of CaSO4 followed by reaction of the
CaS with CO2/H2O.
1-1-2

-------
BENCH.SCALE   ATMOSPHERIC
COMBUSTION EXPERIMENTS

Materials, Bench-Scale Equipment, and
Procedure

   Figure 1 is a schematic diagram of the
bench-scale fluidized-bed combustor system.
The combustor is a 6-in. diameter stainless
steel  vessel.  The  fluidizing  air  enters  the
combustor through  a bubble-cap air distribu-
tor mounted on the bottom flange.  Feeding
and metering of coal and additive is done by
variable-drive   volumetric   screw  feeders
mounted on scales. The solids are fed  pneu-
matically (entrained in a transport air stream)
into the fluidized bed at a point just above the
gas distributor.

   Solids are removed from the off-gas by two
high-efficiency cyclone  separators in  series
and a glass fiber final filter. Downstream from
the cyclones, approximately 20 percent  of the
total  flue gas is diverted to  a gas-analysis
system  and its water content is  reduced to
3000 ppm (by condensation and refrigeration)
to  prevent moisture  interfering  with  gas
analysis. Continuous analyses of the dried gas
for NO, SO2, CO, CH4, and O2 are conducted
with infrared  analyzers and a paramagnetic
oxygen  analyzer.   Gas  chromatography
provides  intermittent  analyses  for  CO2-
Periodically during a run, the bed and  the
overhead   solids  are  sampled   to   permit
chemical  analysis  and  to  obtain  material
balances.  All  instrument signals, pneumatic
and electrical, are  routed to  a  data logger
which  produces  a paper  tape  record  for
further data processing and a typed output of
the signal values.

   In  a  startup,  the  fluidized  bed   of
particulate solids is preheated to ~1000F by
passing heated air through the bed and using
heaters  mounted on the reactor wall. Coal is
then introduced into the bed and  ignited,
increasing the bed temperature to the desired
operating  temperature (e.g.,  ~1600F). The
bed is maintained at a selected temperature by
passing air  or  an air-H^O mixture  through
annular  chambers on  the  exterior  of the
combustor wall.

   In one-stage combustion experiments (with
an  oxidizing  atmosphere in  the bed), all
combustion  air is  introduced at or near the
bottom of the fluidized bed.

   In two-stage combustion,  a stoichiometric
deficiency of air is introduced at the bottom of
the bed to partially burn the fuel. All or most
of the oxygen in the air fed to the first stage is
consumed, and reducing conditions prevail in
the bed. In some ANL experiments, additional
air  was introduced into the freeboard above
the bed through a tube located about 6 inches
above the fluidized bed; oxygen in this air feed
reacted with the CO, hydrocarbons,  and the
unburned carbon elutriated from the bed.

   The coals used in  the various series of
experiments were : (1) Illinois coal from Seam
6,  Peabody Coal  Co.  Mine  10,  Christian
County, Illinois (furnished by Commonwealth
Edison);  and (2) Pittsburgh Seam Coal from
the  Humphrey  Preparation  Plant,  Osage,
West Virginia. Sulfur contents of the coals (on
an as-received basis) were 3.7 and 2.4 weight
percent respectively. The as-received coal was
crushed to pass a  -14 mesh sieve whereupon
more than 80 percent of the coal was in the
-14, + 170 mesh fraction. Limited additional
size reduction occurred as the  coal passed
through the screw feeder.

   A  residual  crude oil was  obtained from
Esso Research and Engineering Co. Its sulfur
content  was  1.9  weight  percent,  viscosity
(Seconds Saybolt  Furol) was  162.5, and  the
flash  point was 178C.

   The natural.gas (obtained  from Northern
Illinois Gas  Company) had a heat content of
1035  Btu/ft3.

   The additive materials studied include: (1)
limestone No. 1359, Stephen  City, Virginia
(97.8  wt  %  CaCO3, 1.3 wt %  MgCO3); (2)
limestone No. 1360, Monmouth, Illinois (69.8
wt % CaCO3,19.2 wt % MgCO3); (3) dolomite
                                                                                   1-1-3

-------
 No.  1337,  Gibsonburg,  Ohio  (53.4  wt  %
 CaCO3,  46.5  wt  %  MgCOs);   and  (4)
 Tymochtee dolomite, Huntsville, Ohio (49.3 wt
 %  CaCOa, 36.6 wt % MgCO3).

 Results and Discussion

    1. Air   Excess,    Coal    Combustion
 Experiments

    The first experiments discussed here were
 made with sufficient fluidizing air introduced
 at the bottom of the reactor so that the  gas
 leaving the fluidized bed contained unreacted
 oxygen.  A  large-particle-size limestone  was
 fed, little of which was elutriated from the bed
 (which consisted of calcined, partially sulfated
 lime).

    a. Effect of  operating variables  on SOi
 retention  For this  mode of operation,  the
 operating variables having significant  effects
 on sulfur retention were the Ca/S mole ratio in
 the feed streams (the ratio of moles of calcium
 in the additive to moles of sulfur in the coal),
 the fluidized-bed temperature, and the  super-
 ficial gas velocity. (Sulfur retention is defined
 as the percentage of the sulfur associated with
 the coal feed that is not contained in the off-
 gas as SO2-) Less significant variables were  the
 types of coal and additive, the particle sizes of
 coal and additive, the height of the fluidized
 bed, and the amount of excess air fed  to  the
 fluidized  bed.  Three variables  having  no
 demonstrable effect on sulfur retention were:
 (1) premixing of coal and additive before they
 were fed to the combustor (instead of feeding
 separate streams of  coal and additive),  (2)
temperature of the gas in the  freeboard above
the fluidized bed, and (3) addition  of small
quantities of water to the fluidized-bed.

   Ca/S mole ratio. Figure 2  shows the effect
of Ca/S mole ratio  on  sulfur retention  at
1450F  for Pittsburgh coal and at 1550 and
1600F  for  Illinois   coal. Sulfur  retention
increased  as  additive  (Ca)  feed rate was
increased in relation to the coal (S) feed rate.
Relatively  good  removals were attained  at
 Ca/S ratios above 3.

   Fluidized-bed temperature.  Experimental
 results indicate that there is a temperature at
 which sulfur retention is at a maximum over
 the range of bed temperatures studied (Figure
 3). With Illinois  coal, a Ca/S mole ratio of
 ~2.5, and  limestone No. 1359 additive, the
 optimum bed  temperature apparently  was
 1500-1550F. With Pittsburgh  coal, a Ca/S
 mole ratio of~4.0, and the same limestone
 additive,  1450-1470 F appears to  be  the
 optimum temperature.  Those  results  were
 obtained in experiments using a gas velocity of
 3  ft/sec and 3 percent excess O2 in the  flue
 gas.
   The  difference in optimum  temperatures
 may be associated with different properties of
 the coal or alternatively, the temperature for
 optimum sulfur   retention  may  have  been
 influenced by Ca/S mole ratio. In any case, an
 operating temperature  of 1500F would  be
 near  optimum.

   Superficial gas  velocity.  Sulfur retention
.was observed  to  increase  with  decreased
 superficial gas velocity (in the range of 3.5 to
 7.4 ft/sec) at a coal combustion temperature of
 1550F   and  with  addition  of  limestone
 No. 1359 (> 1000 fxm average particle size) and
 Illinois coal at  a  Ca/S mole feed ratio of ~4
 (Figure 4). The relatively coarse additive  was
 selected to ensure that additive particles would
 be retained in  the fluidized bed at high gas
 velocities. At gas velocities of 3.5, 5.5, and 7.4
 ft/sec, the average SO2 concentrations in the
 flue gas  were  770,  1250,  and  1500  ppm,
 corresponding to  retentions  of 83, 73, and 66
 percent of the sulfur fed to the reactor. These
 data may be correlated with the equation
where:
          R = 101.79 e-
          R = SO 2 retention, %
          v = superficial gas velocity, ft/ sec
                                         (1)
   Results of British  experiments (1) using
Welbeck coal, 440-nm British limestone, Ca/S
mole ratios of 1 and 2, and a coal-ash fluidized
1-1-4

-------
bed show that sulfur retention is greater at a
gas velocity of 2 ft/sec than at 3 ft/sec (Figure
4). The slopes decrease as the Ca/S mole ratio
increases;  thus,  at sufficiently  high  Ca/S
ratios, sulfur retention may be essentially
independent of superficial gas velocity.

   Excess air. The oxygen level in the off-gas
was varied by adding pure oxygen at several
rates to the fluidizing air before it entered the
preheater.  At 1550F, ~3 ft/sec  gas velocity,
and Ca/S mole ratio of ~3, the sulfur  reten-
tions were 67, 71, and 75 percent, respectively,
at 0.7, 2.4, and 5.6 percent C*2 in the flue gas.
Apparently,  oxygen  concentration   affects
slightly the reaction of SO2  with limestone,
and sulfur retention can be  expected  to in-
crease when  oxygen concentration in the flue
gas is increased.

   Fluidized-bed  height.  Runs  were  per-
formed at  1550F, 3 ft/sec gas velocity, and
Ca/S ofM with three different bed heights 
14,24, and 46 inches (length to diameter (L/D)
ratios of  2.3,  4.0,  and  7.7).  The   sulfur
retentions  were  78,  80,  and   83  percent,
respectively,  indicating that bed height has a
small but real effect.

   Type  of coal. The effect of type of coal on
sulfur retention  could  not  be  evaluated
because  of  insufficient data in  ANL  ex-
periments.   However,  qualitatively,   sulfur
retentions  for  Illinois  and  Pittsburgh  coals
differed little. Work by the British and by the
U.S.  Bureau of Mines has  evaluated this
variable  in greater detail.

   Type of additive. Sulfur retentions for runs
performed at a Ca/S  mole ratio of 2.5 and
fluidized-bed  temperatures   of  1550   or
1600F with  several types of additives may be
compared in Figure 2.  Sulfur retention varied
about 10 percent indicating that differences in
these  additive types had only  small effects at
these  operating conditions.

   Size  of coal particles. To determine  the
effect of sulfur retention of the particle size of
the coal feed, two experiments were completed
with -12  +50 mesh and -50 mesh (a -12 +50
fraction ground to -50 mesh) Illinois coal at a
Ca/S mole feed ratio of 2.4 and a fluidized-
bed temperature of 1550F. Sulfur retentions
by No. 1359 limestone additive were 81 and 75
percent,  respectively, for the -12 +50 mesh
and  -50 mesh feeds. An experiment in which
Illinois  -14 mesh  coal  was  burned  under
operating conditions nearly  identical to those
used in the above experiments  yielded a  78
percent sulfur retention. Thus, similar sulfur
retentions were obtained by burning coal  of
three  particle  size distributions, but the
coarsest coal feed appears  to yield the best
results.

   Size of additive particles. Sulfur retention
was  calculated by  interpolation to  be ~87
percent for larger additive particles (1000 /xm
average) and  93 percent for smaller particles
(630 Mm average) in runs with a Ca/S mole
ratio of 4.0, a temperature of 1550F, and a
gas velocity of 3 ft/sec. This suggests that at
least in the region of high sulfur  retention,
additive particle size in this range has only a
moderate  effect on sulfur retention.

   b.  NO levels in the flue gas  Nitrogen
oxides,  principally  nitric  oxide  (NO), are
formed during the combustion of fossil fuels
and  are  an  important contributor  to  air
pollution.  Although  the  quantities  of NO
observed  in  the  flue  gas  from    high-
temperature conventional combustors may  be
accounted for  by  the  equilibrium  of the
nitrogen fixation reaction, this is not the case
for  low-temperature  fluidized-bed coal
combustion in which NO concentrations far in
excess of  those expected on the basis of the
equilibrium have been observed. At 1600 F, a
common  temperature for  fluidized-bed
combustion, the equilibrium concentration of
NO from the fixation of atmospheric nitrogen
ranges from 50 to 200 ppm,  depending on the
oxygen  concentration.  (The  oxygen  con-
centration, in turn,  depends on the  level  of
excess air employed.) However, in actual fluid-
bed coal combustion, nitric oxide levels of 400
to 800 ppm in the flue gas have been measured
with  oxygen concentrations  of ~3  volume
                                                                                      1-1-5

-------
 percent in the off-gas (15 to 20 percent excess
 air).

    In previous ANL work, the major source of
 nitric oxide during the combustion of coal was
 determined to be the nitrogenous content of
 the coal (1 to 1.5 weight percent in U.S. coals).
 In more recent  work, the effect  of moisture
 content of the coal on NO level was studied by
 adding water at several rates to the fluidizing
 air. Pittsburgh coal and limestone No. 1359
 were the feed materials.  The fluidized-bed
 temperature was 1450 F and  the Ca/S mole
 ratio  was  1. The  concentration  of  NO
 decreased from 530 ppm to 510 ppm  when 10
 cm3/min water  was added (equivalent to 26
 weight percent water in the coal), and to 380
 ppm when  the  rate of  water addition was
 further increased to 30 cmVmin (equivalent to
 51 weight percent water  in the  coal). These
 decreases in NO concentration may be due to
 chemical reduction of NO by hydrogen pro-
 duced by the water-gas-shift reaction, but  the
 effect  is  not great enough to warrant further
 attention.
   c.  Calcium utilization    The  relative
extent of conversion of CaO to CaSO4 for bed
and elutriated materials was  calculated  from
calcium and sulfur concentrations determined
by  wet chemical  analysis. In several experi-
ments with No. 1359 limestone of an average
particle  size  of  490 jum, at  temperatures
ranging  from  1400  to 1600F, and  gas
velocities from 2.5 to 2.8 ft/sec, the conversion
of  calcium  oxide to  calcium  sulfate was
highest (2/3 converted) for particles collected
on the final filter. These particles have a high
surface to volume ratio and would be expected
to react rapidly even though  their residence
time in the bed is relatively short. Next highest
conversion, ~2/5,  was obtained in bed par-
ticles  which have a relatively long residence
time in the bed. Lowest conversion, ~l/4, was
obtained  with  solids removed from cyclones.
These particles are larger than  final   filter
particles, smaller than the bed particles, and
are present in the bed for  a  relatively short
time.
    2.  Air Deficient,  Coal Combustion  Ex-
 periments
    The  concept  of  two-stage  combustion
 provides for  a substoichiometric quantity of
 air (that is, less air than is required to burn the
 coal completely to COa and H2O) introduced
 into the first stage of the fluidized bed to
 which coal is fed. Additional  air may be in-
 jected into the disengaging section  above the
 fluidized  bed (the  second stage)  to  burn
 gaseous hydrocarbons, Ha, and CO  in the gas
 stream from  the  first  stage.

    Two-stage combustion, experiments  of an
 exploratory   nature  were  conducted  to
 determine if this combustion mode might have
 benefits,  as   compared   with   single-stage
 fluidized-bed  combustion. To simulate  the
 conditions of combustion  in  the first  stage
 only, experiments were performed in which a
 substoichiometric quantity of air was intro-
 duced into the bottom of the fluidized bed, but
 no  secondary air was fed. In other experi-
 ments, secondary air was introduced  above the
 fluidized bed. The bed consisted of coarse lime
 particles in all of these experiments.

   The  experimental   results  include   in-
formation  on  the concentrations of SOa,  HaS,
NO, and CO in the off-gases at various air feed
rates and bed temperatures, as well as data on
the sulfur content of solid products when a
substoichiometric quantity of air  was fed to
the first stage only.
   a. Effect of decreasing air input on ratio of
 HaS to HaS + SO2  in flue gas  The concen-
 tration of Ha S in the off-gases was  measured
 to determine  which operating conditions  affect
 the  formation of this sulfur  compound.  In
 experiments in which air was introduced to the
 first stage only, the amounts of H2S and SOa
 in the off-gas were compared. The percentage
 of sulfur in the off-gas as HaS was sensitive to
 the amount of air introduced into the fluidized
 bed, increasing drastically when the air feed
 rate was reduced below a value corresponding
 to ~70 percent of the stoichiometric quantity
 of air necessary to react with the coal feed (see
 Figure 5). (Although the parameter, air feed
1-1-6

-------
rate as a percent of stoichiometric, was based
on feed rates of coal and air, it is recognized
that the quantity of  coal  actually oxidized
varies with  other parameters  (i.e.,  tempera-
ture, etc.). For certain correlations, it may well
be more suitable to use the parameter, stoichi-
ometric air feed rate based on the coal actually
oxidized.) At an air feed rate equivalent to
~50 percent of  the stoichiometric  quantity,
the concentration of H2S  (611  ppm)  was
nearly equivalent to the concentration of SO2
(660 ppm). At air inputs of 70 to 80 percent of
the  stoichiometric   quantity,  the  relative
amount of sulfur as H2S fell to about 2 percent
of the total sulfur in  the gas.
   In those experiments in which secondary
air was  introduced  above the fluidized bed
(Figure 5), the E^S level in the off-gas was low
 corresponding to less than 1 percent of the
total sulfur  in the gas. This suggests that any
H2S  in the gas leaving  the first stage is
oxidized to SOa by air fed to the second stage.
   No consistent  relationship  was  apparent
between H2S level and either the temperature
of the  fluidized  bed (1450-1650F)  or the
temperature of the off-gas in the freeboard
above the bed (1100-1800F).

   b. Sulfur retention  Sulfur retention is
defined  as  the  percentage  of  the  sulfur
associated with the coal feed but not contained
in the  off-gas  as SO2  or H2S. (Because  a
fraction of the carbon was not burned in these
experiments, the sulfur retention values given
are probably higher than would be realized if
all of the carbon  were burned.) Experiments
were performed  with  no  introduction  of
secondary air (Figure 6) to determine sulfur
retention as a function of the Ca/S mole ratio
in the feed at  1450, 1550, and 1650F. Also
shown in Figure 6 (to allow comparison) are
data for experiments carried out earlier under
one-stage oxygen-excess conditions  at 1450-
1470 F.

   The    data   presented   for   the   sub-
stoichiometric  air experiments show a  large
amount of scatter principally due to variation
in the quantity of air fed to the fluidized bed,
which was not the  same in all experiments.
Stoichiometric air added in each experiment
ranged  from  51 to  91 percent.  The  best
retentions  were  observed  at  stoichiometric
additions  of less than 60 percent.

   For experiments carried out at a Ca/S ratio
of about 2 and temperatures of 1450-1650 F
(Figure 7), no simple relationship between the
amount of air introduced  into the bed and
sulfur retention was evident; however, a line
has been fitted to the points as shown. At an
air feed rate of 100 percent of stoichiometric,
sulfur retention is about 65 percent.  As the air
feed rate  is decreased, sulfur  retention first
decreases to about 45 percent as the air rate is
decreased to  75 percent of  the  calculated
stoichiometric requirement and then increases
rapidly as the  air rate is decreased further.
This  suggests  that in an oxygen-deficient
region (75 to 95 percent of calculated stoichio-
metry), removal of sulfur by lime in the form
of SO2 is poor, but that at lower air flow rates
sulfur is in the form of H2S and is  efficiently
removed.  This would be  expected because
oxidizing  conditions  are  required for  the
retention of SO2 by lime (to convert a CaSC>3
intermediate  to  CaSC>4), whereas  reducing
conditions are required for the retention  of
H2S by lime.

   The introduction of secondary air above the
bed ^resulted in erratic but  generally  lower
sulfur retentions.  Decreases were  about 10
percent at 1450F, 20 percent at 1550F, and
40 percent at 1650F. The increased sulfur
content of the off-gas after secondary  air was
introduced was probably caused by burning of
entrained  coal particles in the second stage to
produce additional SCh.

   c. NO  levels in the flue  gas  When coal
was burned with a deficiency of air fed to the
first stage, concentrations of NO in the off-gas
from  the  first  stage  as a function  of  the
amount of air introduced into the bottom of
the fluidized bed were as shown in Figure 8.
To obtain these data, only the  first stage  was
operated.   The  NO   concentrations  were
                                                                                      1-1-7

-------
 generally < 250  ppm and  apparently  were
 affected by both the amount of air introduced
 and the temperature of the  fluidized bed. At
 the  lower air  feed rates, the NO levels  were
 generally lower. At a  given air feed rate, NO
 levels were higher at lower bed temperatures.
 Since  earlier  work   at  ANL  showed  that
 nitrogenous compounds in coal are oxidized to
 NO during fluidized-bed combustion, it can
 be postulated  that the lower levels  of NO
 observed at higher temperature are  due to
 more  rapid  decomposition of NO.  This
 decomposition may  be promoted  by  the
 presence of CO; another possibility  is that
 nitrogenous compounds other than NO may
 be formed in the highly reducing atmosphere
 of the bed.

   The data presented in  Figure  8 are not
 corrected to  an equivalent off-gas  volume
 basis. However, if this correction were made,
 the dependence of NO emissions on air feed
 rate  would  be  even  more   pronounced,
 assuming  that  feed   rates of coal  were
 equivalent.

   Upon the  introduction  of secondary  air
 above the fluidized bed, NO levels in the off-
 gas  varied  eratically   usually increasing.
 Possible explanations for this behavior are: (1)
 any reduction  of NO by CO in the zone above
 the bed would be suppressed by introducing
 secondary air  or  (2) if a nitrogen compound
 such as ammonia were present in the gas, it
 may be oxidized to NO by the secondary air.

   d. Sulfur species in the bed  Results show
 that the sulfide content  of bed  material
 decreased as air  flow was  increased.  At  air
 inputs   corresponding  to   50   percent  of
 stoichiometric, as much as 100 percent of the
 sulfur in the bed was sulfide. In  most  experi-
 ments  in which the  air input  exceeded  65
 percent of  stoichiometric,  sulfide  content
dropped off rapidly to  less than 1 percent. No
 relationship  was  found   between   sulfide
content and bed temperature.

   The  sulfite content of bed  samples was
 erratic, ranging between 6.2 and < 0.1 weight
percent. No correlation of sulfite content with
either bed temperature or air feed rate could
be found.
   e. Carbon balance  Carbon balances were
made for three experiments. For three other
experiments, all data for making the balances
except the COa  level in  the  flue gas  are
available (see Table 1). The small quantity of
carbon not accounted for is represented by
hydrocarbons (other than  CH4) for which no
analyses are made and by a small loss of fine
carbon  participate  from  the  combustion
system. The data show that as the volume of
air added to  the bed decreases (experiments
14-1A, -IB, -2) at the same temperature, the
CO content of the flue gas increases markedly,
the CH4 content increases slightly, and  the
quantity of carbon elutriated to the first and
second cyclone separators from the fluidized
bed increases.

   The carbon  content of the bed  under low
stoichiometric air additions (55 percent) was
as high as 31 percent (experiment 14-3B).
Under these conditions the amount of carbon
elutriated  was about  15 percent of that fed.

   f. Preliminary evaluation of the  concept 
Although  the work conducted on two-stage
combustion was  exploratory in  nature, a
preliminary evaluation of the concept can be
made. The principal advantages of the two-
stage combustion  concept  over  one-stage
combustion are:  (1) lower  NO emissions; (2)
retention of sulfur in the  form  of calcium
sulfide (rather than sulfate),  allowing  for
potentially easier regeneration of the additive;
and (3) production of a combustible gas that
could be  used  in conjunction  with  a  gas
turbine.

   The principal disadvantages are: (1) greater
elutriation  of carbon, (2)  possible compli-
cations in additive regeneration owing  to the
high carbon  content of  the  bed,  and  (3)
necessity  for removal of heat from the bed
under conditions  that might be corrosive to
immersed steam tubes.
1-1-8

-------
  Table 1. CARBON BALANCES AND CARBON CONTENTS OF SOLIDS IN
       SUBSTOICHIOMETRIC AIR-COMBUSTION EXPERIMENTS

Carbon in coal, g/hr
Carbon out, g/hr
First cyclone
Second cyclone
Flue gas
CH4
CO
CO 2
Total carbon out, g/hr
Carbon concentration
in solids streams, wt%
Bed
First cyclone
Second cyclone
Run conditions
Temperature, F
Coalfeed,lb/hr
Air, % of stoichiometric
14-1A
1428

71
22

>24
42
1230
>1470

<1
29
52

1450
5.0
90
14-1 B
1428

136
11

32
350
NDa
ND

<1
39
53

1450
5.0
86
14-2
1684

174
24

39
441
840
1632

6
46
55

1450
6.3
71
14-1C
1799

167
7

42
521
ND
ND

<1
39
56

1550
5.9
54
14-3A
1485

235
7

28
330
ND
ND

20
47
35

1600
5.2
64
14-3B
1713

233
16

32
455
800
1682

3.1
62
53

1600
6.0
55
No data available.
                                                                1-1-9

-------
   Sulfur retention  appears to  be roughly
 equal for the two concepts. It is notable, also,
 that no problems of coal  caking were en-
 countered even at high bed carbon contents.

   Further work might be warranted at lower
 air  addition  rates   and higher  bed  tem-
 peratures to avoid heat generation in the bed.
 Under  such  conditions the  process  would
 become gasification  rather than combustion.

   3. Air Excess, Oil Combustion Experiments

   To assess the  removal of SOi  from com-
 bustion gases when residual fuel oil is burned
 in a fluidized bed of sulfated lime with con-
 tinuous feeding of limestone additive, experi-
 ments were performed in the  6-in. diameter
 fluidized-bed combustor at a variety of oper-
 ating conditions. Residual fuel oil was burned
 in an excess of oxygen at bed temperatures
 ranging from 1450  to 1650F,  Ca/S mole
 ratios up to 11.9, a gas velocity of ~3 ft/sec
 (except  for one experiment at 5.5 ft/sec),  and
 with 3 volume percent oxygen  in the flue gas
 (except  in one experiment with  1 volume per-
 cent oxygen in the flue gas).

   The  following results were  obtained  in
 these experiments.

   1. The  effect  of temperature  on  sulfur
 retention is similar to that observed in coal
 combustion experiments, in  which there is a
 temperature  yielding  maximum  sulfur
 retention. In the oil-combustion  experiments,
 maximum  sulfur  retention  was  at  1500-
 1550F.

   2. The  shape of the curve   for  sulfur
 retention as a function of Ca/S_ mole ratio is
 similar to that obtained in  coal combustion
experiments.   Sulfur  retention  in  the  oil
combustion runs increases as Ca/S mole ratio
increases to about 5, then levels off at a  90
percent sulfur retention level as the Ca/S ratio
is increased further. The slope of the curve for
sulfur retention as a function of Ca/S ratio is
less steep than the slope for Illinois coal (3.7
weight  percent sulfur)  at similar operating
conditions.
   When oil was combusted, the NO levels in
the flue gas ranged from 110 to 150 ppm for
the experiments with 3 percent O2 in the flue
gas. This may be compared with the 400 to 800
ppm range observed when coal was burned.
However, the nitrogenous content  of residual
oil is also less than that of coal. No correlation
of NO level with  bed temperature or  Ca/S
mole  ratio was  observed.  Combustion  ef-
ficiency in  these experiments  is discussed  in
the section of combustion efficiencies below.

   4. Miscellaneous
   a. Additive decrepitation  rates during coal
combustion experiments  Decrepitation ^and
attrition  of several  additives  during coal
combustion experiments has been estimated
from the calcium content of elutriated fines.
(The fraction of additive carried over can only
be  estimated  because  the  particle  matter
elutriated during the combustion of coal in a
fluidized bed is a mixture of solids of different
origins and compositions.)

   In most experiments a gas velocity of ~2i6
ft/sec was used; at this velocity all of the flyash
and  additive  particles having diameters  of
<177 nm are.expected to elutriate from the
fluidized bed.

   The expected elutriation for each of several
series  of  experiments  calculated  in  this
manner is shown in Table 2.  The actual elutri-
ation was determined from  calcium material
balances (with an allowance  made for the
calcium content of the flyash,  which was also
expected to elutriate). The difference between
actual  and expected  elutriation  gave  an
estimate  for the decrepitation of larger ad-
ditive particles (Table 2).

   The indicated  decrepitation of BCR-13.59
limestone was  ~8  percent,  but no  decrepi-
tation  of a British limestone was  evident.
Decrepitation   of  limestone  BCR-1360  and
dolomite BCR-1337 was more  severe40 arid
85 percent, respectively. These results indicate
that decrepitation of BCR-1359 and British
limestone is low and  that limestones of this
 1-1-10

-------
              Table 2. ESTIMATED DECREPITATION OF ADDITIVE MATERIALS
                          FROM FLUIDIZED-BED COMBUSTOR


Experiments
Amer-1 ,-3,-4
BC-6,-7,-8
AR-1,-2,-4,-5,-6 ,
BC-9
BC-10
Brit-1,-2,-3,-3A
and Amer-Brit


Additive type
BCR-1359
BCR-1359
BCR-1359
BCR-1360
BCR-1337
British limestone

Distribution of total calcium
in combustor system, wt %
Expected
elutriation3
12
2
13
2
<2
37

Actual
elutriation b
19
11
21
42
~87
33

Estimated
decrepitation
7
9
8
40
85 c
	

  Calcium contained in particles < 177/4 m diameter in the additive feed to the system.
  These particles are expected to elutriate at a superficial velocity of 2.6 ft/sec.
  Fine particles in the starting fluidized bed are not included.

 bCalcium fed with the coal was deducted from the total calcium found in the elutriated material.

 c Derived from  both calcium  and magnesium material balances.
type are desirable materials for use in a full-
scale fluidized-bed  combustor  with  regen-
eration and recycle of additive. Higher decrep-
itation rates for BCR-1360 and BCR-1337 may
make these materials less promising for regen-
eration and recycle. Data are for one cycle of
use only.


   b. Cyclone  collection efficiencies during
coal combustion experiments  Data on the
particle removal efficiency of the ANL cyclone
separators have been compiled as  a basis  for
estimating the dust loading and filter area of a
cartridge filter  for a pressurized combustion
bench-scale plant now being designed. In the
present atmospheric  pressure system, flue gas
passes through two cyclones in series and a
final filter. (The diameter of the first in-line
cyclone is 6-5/8 inches;  the diameter of the
second is 4-1/2 inches.) It is planned to use the
same two  cyclones  with  the  pressurized
combustor.

   To determine  the  adequacy of the glass
fiber mat filters used in the atmospheric plant,
collection efficiencies (defined as  ratio of the
weight of particles removed to the weight of
particles entering the  cyclone) were compiled
for 26 earlier ANL experiments. In these one-
and  two-stage  runs,  the  flue  gas  flowrates
ranged from 8 to 14  ft3/min,  the coal feed
rates from 4 to 7.3 Ib/hr, the additive feed
rates from  1.1 to 2.3  Ib/hr, and  the dust
loadings at the -combustor exit from  0.16 to
1.78 g/ft3.  Combined efficiency  of the  two
cyclones was above 80 percent in 24 of the 26
experiments and above 90 percent in 21 of the
experiments. The dust loading in  the flue gas
leaving the second cyclone averaged 0.06 g/ft3
for the  26 runs;  the  maximum loading  was
0.22 g/ft3.
                                                                                    1-1-11

-------
   c. Combustion efficiencies for coal, oil and
 gas  with  excess air  The  combustion ef-
 ficiency  for  experiments  performed in  the
 combustor was determined as the ratio of
 carbon burned to carbon fed,  multiplied by
 100. The carbon loss is calculated by deter-
 mining unburnt carbon leaving the system by
 three routes:  (1) carbon  associated with the
 elutriated  solids;   (2)  incompletely  burned
 gases  (e.g.,  carbon  monoxide and  hydro-
 carbons);  and  (3)  carbon associated  with
 fluidized-bed material taken from the system.
 All  experiments  were  conducted  without
 recycle of fines.

   Combustion efficiencies in ten experiments
 with coal ranged from 93 to 96 percent. In all
 experiments, carbon losses in the bed material
 were negligible. Only about 10 to 20 percent of
 the carbon loss was due to the formation of
 carbon monoxide  and hydrocarbons.  The
 major  carbon loss (80 to 90 percent) occurred
 as a result of elutriation of fine particles in the
 exhaust  gases  before  they were  completely
 combusted.  Combustion  efficiency can be
 increased by  recycling the  elutriated  ash-
 carbon mixture to the  fluidized bed or to a
 carbon burnup cell. Oxygen concentration in
 the flue gas  in these  experiments  was ap-
 proximately 3 percent.

   Combustion efficiency in  oil combustion
 experiments was similar to that observed for
 coal combustion  experiments  under similar
 conditions, ranging from 94 to 96 percent for
 experiments with 3 percent O2 in the flue gas.
 However, sources of carbon losses for the two
 fuels differed. In coal combustion, most of the
 carbon  loss is represented by the  carbon
 content of solids elutriated to the cyclones; in
 oil  combustion,  inefficiency  results  from
 incomplete burning of the CO and hydro-
 carbons formed during combustion. Efficiency
 can probably be improved  by operating the
 combustor with a deeper bed or by increasing
the freeboard temperature. Lower combustion
 efficiencies were  observed  with less excess
oxygen  in the  flue gas and  at higher gas
velocities.
   During the combustion of natural gas, the
elutriation of carbon-bearing fine particles is
negligible, and the  major  loss of unburnt
carbon is in the carbon monoxide and total
hydrocarbons  in  the flue  gas.  Combustion
efficiencies,  calculated  from   analyses  of
samples of the flue gas, ranged from 94.1 to
99.2 percent at 1600F with 3 percent  excess
oxygen. At 1800F, combustion efficiencies of
94.1 to 98.8 percent were  observed.  These
results are similar to  data  reported by the
USSR on combustion of gas in a fluidized bed.
Although  the  USSR  data  indicate  that
combustion efficiency is principally affected
by bed temperature,  combustion efficiency is
also likely to be a function of bed depth, gas
velocity, and excess oxygen concentration. For
example,  in  experiment  NG-3  at 1800F,
combustion  efficiency  was decreased  to 91
percent  when the   combustion  was  in-
tentionally forced  toward  more  reducing
conditions; i.e., 1.5 volume percent O2 in the
flue gas rather than 1.8-4.5 volume  percent
02.

REGENERATION  OF  SULFUR  CON-
TAINING ADDITIVES
   When  coal is burned in a fluidized bed
containing limestone or  dolomite,   sulfur-
containing gases from the  combustion  of
sulfur-containing substances in the coal react
with the bed material and are retained  in the
bed. The reaction product is calcium sulfate if
combustion  is carried out  under oxidizing
conditions, or calcium sulfide if combustion is
carried out with a deficiency of air.
   Several regeneration processes are  under
consideration.  These are:
   1.  Reductive  decomposition  of calcium
      sulfate.
   2.  Roasting of calcium  sulfide in air  or
      oxygen.
   3.  Reaction of calcium sulfide with  water
      and carbon  dioxide.

   Processes 2 and 3 can be used to regenerate
not only calcium sulfide, but also material
containing calcium sulfate if the sulfate  is first
reduced to the sulfide.
1-1-12

-------
Thermodynamic Analyses

   Valuable information can  be gained  by
considering the   thermodynamics  of  the
process.  The  yields  of  gaseous   sulfur-
containing products, the composition of solid
phases, and the variations of these yields and
compositions with temperature, pressure, and
gas composition for a system  at equilibrium
can all be  obtained. Optimum reactant feed
ratios  and  gas compositions  can  also  be
calculated   easily   when   product  concen-
trations,  compositions,  and  pressures  are
specified.
   All of the following predictions and con-
clusions  are based on the  supposition that
chemical equilibrium is achieved among the
various phases. This implies that the rates of
all relevant chemical reactions are large on the
time  scales being used, which   scales are
determined by mass transport rates within the
system.  The maximum rates  at  which this
supposition is valid vary with temperature and
must be  determined in the laboratory and in
the pilot plant. It is further assumed that the
system is not stoichiometrically limited. There
must always be at least small amounts of the
appropriate  solid  phases present  for  the
results of these calculations to be valid. One
must not assume that all actual processes will
be operated with all these solid  phases present,
however.
   The assumption is also  made that  solid
solutions do not form to  any great extent.
Exploratory experiments to date support this
assumption.

   1. Reductive Decomposition of CaSO4 with
CO/CO2
   Before the relative amounts of the species
in an equilibrium mixture from the reduction
of calcium sulfate with  carbon  monoxide-
carbon dioxide mixtures can be calculated, the
solid phases present at the various conditions
of temperature and carbon monoxide/carbon
dioxide  ratio  must   be  determined.  The
possible  sulfur-containing  solids  are  con-
sidered to be calcium sulfate, calcium sulfite,
and calcium sulfide.
   a. Conditions for the presence of calcium
sulfate  and  calcium  sulfide  The solid
phases  present  at  equilibrium  with   a
PCO /Pco2  f 0.005-0.055 and  temperatures
of 1600 to 2400F are shown in Figure 9 (in
which  temperature  is   the  ordinate  and
Pco /PCO2 the abscissa). Examination of the
expression for Kp in reaction 2
    1/4 CaSO4 + CO - 1/4 CaS + CO2
                       C02
                       CO
(2)
shows that for any temperature, there is but
one  ratio of  carbon monoxide  to  carbon
dioxide at which calcium sulfate and calcium
sulfide can coexist at equilibrium. The co-
existence conditions appear as the line run-
ning from the lower left to the upper right part
of Figure 9 and represent  CO/CO2 ratios  at
which CaS and CaSO4 can both be present  at
equilibrium. In the area to the right of this
line, the gas mixture is so rich  in  carbon
monoxide that calcium  sulfate is  completely
reduced to calcium sulfide. To the left of the
line, the gas mixture is so rich  in  carbon
dioxide  that  calcium sulfide  is  completely
oxidized to calcium sulfate. This line is called
the coexistence line for  calcium sulfate and
calcium sulfide.
   b. Conditions  for presence  of calcium
sulfite  Calcium sulfite is not stable in the
presence of CO/CO 2 mixtures at any tem-
perature from  1500 to 2400F. This has been
established by plotting "coexistence" lines for
calcium sulfite with calcium sulfate and for
calcium sulfite with calcium sulfide. These are
analogous to the calcium sulfate-calcium sul-
fite coexistence line described above and are
determined in the same way from the equilib-
rium constants for reactions 3 and 4.

    CaSO4 + CO  - CaSO3 + CO2
                   C02
                                        (3)
                                                                                   1-1-13

-------
     1/3 CaS + CO 2-1/3 CaSO3 + CO
                                       (4)

   c.  Conditions for the presence of calcium
 carbonate  and  calcium  oxide    When
 calcium sulfate is reduced or calcium sulfide is
 oxidized by  a  mixture  of CO  and  CO2,
 calcium  oxide is formed.  However, in  the
 presence  of   carbon  dioxide   at  sufficient
 pressure,  calcium  oxide  is  converted  to
 calcium carbonate.

   A  coexistence line  for the carbonate and
 the oxide is determined  by the equilibrium
 dissociation pressure  of calcium  carbonate
 (reaction 5).
         CaC03 ^  CaO + CO2

                Kp  =PC02
(5)
It also appears as a nearly horizontal line at
about 1950F in Figure 9. This line represents
the temperature at which the partial pressure
of COa in the equilibrium mixture just equals
the equilibrium pressure of CO2 over calcium
carbonate.

   The  partial  pressure  of  CO 2  in  the
equilibrium mixture is obtained by assuming a
total pressure  of 10 atm and subtracting the
pressures of SO2 and CO. Clearly, if the total
pressure is lowered or if an inert gas is added,
the pressure of CO2 will  be  lower and the
horizontal line will be at a lower temperature.
It is also clear that the calcium carbonate does
not exist above the horizontal line and that
calcium oxide  does not exist below  it.

   d. Sulfur dioxide pressure  The pressure
of sulfur  dioxide in the equilibrium mixture
can be calculated from the CO/CO 2 ratio and
the equilibrium constant  of the  reaction
appropriate to  the part of Figure  9 under
consideration; however, in the areas labeled C
and D, one must  generate independent in-
formation about the CO2 pressure by making
assumptions exactly analogous to those made
       in the above discussion of calcium carbonate.

         In area A of Figure 9, SO2 is generated by
       reaction 6.
         CaS04 +CO ^ CaO + SO2  + CO2
                 Kp =-
SO2    CO2
   pco
                                              (6)
       The pressure of SO2 is shown as a family of
       isobars slanted down toward the right. In area
       B,  SO2  is  generated by  the oxidation  of
       calcium sulfide in accordance with reaction 7.

       1/3 CaS + CO2  ^ 1/3 CaO + CO + 1/3 SO2
                                                                          Lco
     C02
(7)
       The isobars of constant SO2 pressure in area
       B curve down to the left, meeting those of area
       A at the calcium sulfide-calcium sulfate co-
       existence line. At any temperature, the SO2
       pressure is at a maximum at this junction. For
       example, at 2000 F, the maximum attainable
       equilibrium pressure of SO2 is  0.46 atm at a
       CO/CO2  ratio of 0.020. This maximum in the
       SO2   pressure   may   be  understood   by
       examination of the appropriate equilibrium
       constants. For example, the expression for the
       equilibrium constant for reaction  6 predicts
       that the pressure of SO2  is directly propor-
       tional to the  CO/CO2  ratio.  Thus,  the
       pressure of SO2 must increase as the CO/CO2
       ratio increases as long as reaction  6 obtains.
       From the  expression  for  the equilibrium
       constant in reaction 7, it may be seen that the
       SO2  pressure is inversely proportional to the
       cube of the  CO/CO2  ratio.  Thus, the SO2
       pressure  increases  with decreasing CO/CO2
       ratio as long as reaction 7 obtains. Reactions 6
       and 7 occur simultaneously only along the co-
       existence line. Thus, as one moves away from
       the coexistence  line, the SO2 pressure must
       decrease.
1-1-14

-------
   In area C, reactions applies.

       CaSO4 + CO=^CaCO3 + SO2
                     P
              KP =
                      SO-
                      CO
(8)
In this area, the pressure of SO2 is dependent
only on Kp  and carbon monoxide pressure,
and the SO2  isobars are nearly vertical. The
main effect results from the variation of Kp
with the temperature.  In  area D, the  SO2
pressure is once again a strong function of the
CO/CO 2 ratio as may be seen from reaction 9.

   CaS + 4 CO2 ^CaCO3 +  SO2 + 3 CO
                  'SO.
                                       (9)
   The slope of the isobars in area D differs
only slightly from the slope  in area B  as  a
result  of  increased  dependence on  CO2
pressure in area D.

   e. Sulfur pressure  The pressure of sulfur
vapor  is  quite low  (<10~2  atm) over  the
temperature  range  1700  to 2300 F.  The
pressure of sulfur in area B was calculated
from reaction 10.
      CaS + CO2 -1/2 S2 + CO + CaO
                                      (10)

 Since  the formation  of sulfur  is  entirely
 analogous to  the  formation  of SO2, sulfur
 concentration  may be expected to exhibit the
 same sort of maximum at the calcium sulfide-
 calcium sulfate coexistence line.

   f. Carbonyl sulfide pressure  The car-
 bonyl  sulfide  pressure  was calculated  with
 reaction 11,
             CaS + CO 2- CaO + COS


                    K
                      COS
                     rCO,
                                             (11)
Since P^OS *s dependent on the pressure of
CO2, assumptions made in calculating COS
pressure  are  similar to those made in the
discussion of calcium carbonate.
(~10"3atm) in this system.

   g.  Solid-solid reaction of calcium sulfide
with calcium  sulfate  SO2 is generated by
the reaction of calcium sulfide with calcium
sulfate, as is  shown in reaction 12.

  1/3 CaS + CaSO4 ^4/3 SO2 + 4/3 CaO

                            4/3
                    KP = (PSO J
                                      (12)
       Since reaction 12 is exactly equivalent to the
       sum of reactions  6 and 7, the SO 2 pressure
       calculated from reaction 12 must be just that
       calculated from reaction 6 or reaction 7 using
       the CO/CO2  ratio at the coexistence line.
       Another way of saying this is that the presence
       of both calcium sulfide and calcium sulfate
       determines an oxidizing  potential for the
       atmosphere with which  it is in equilibrium;
       this  oxidizing  potential  determines  the
       CO/CO 2 ratio of the  atmosphere. If carbon
       monoxide and carbon dioxide are present in
       the gas phase over a mixture of calcium sulfide
       and calcium  sulfate,  they  serve as a  facile
       route to the production of SO2  so that rapid
       reaction rates  for mixtures of the two  solids
       are possible.

          h.  Reduction of  calcium  sulfate with
       H2/H2O  mixtures  The system calcium
       sulfate-calcium  sulfide-H2-H2O  is  exactly
       analogous to  the system  calcium  sulfate-
       calcium sulfide-CO-CO2. This means that all
       the features of the CO-CO 2 system are present
       in the H2-H2O system. The sulfur-containing
       solid  phases once again are calcium  sulfate
       and calcium sulfide, but calcium  oxide and
                                                                                  1-1-15

-------
 calcium   hydroxide   are   the   non- sulfur-
 containing solid phases. The pressure of SO 2
 at a given temperature has a maximum value
 where the calcium sulfate and calcium sulfide
 are in equilibrium with the H2-H2O mixture.
 The instability of calcium sulfite can be shown
 in the same way as in the CO-CO2  system. In
 fact, the  major  differences  between  the
 H2-H2O system and the CO-CO2  system are
 that carbonyl sulfide  is replaced  with H2S and
 that at any given temperature, the numerical
 value for the H2/H2O ratio differs from that
 of the CO/CO2 ratio.

   An  additional  difference  between  the
 systems is that in the  H2-H2O system, calcium
 hydroxide can form at lower temperatures and
 higher pressures of H2O (just as CaCOa can
 form in the CO-CO2  system). However, at 10-
 atm  H2O pressure,  Ca(OH)2  is  not  stable
 above 1200 F.

   For any temperature and SCh pressure, the
 H2/H2O ratio  can  be calculated from the
 equivalent CO/CO2  ratio via reaction 13.
         H2 + CO2 ^
                  P   / P
                  *CO  CO,
                   AJ2  "^          (13)

This is the familiar water-gas shift reaction.
The principle involved here is that equal SO 2
pressures are obtained in the Hj -H2 O system
and the CO-CO2 system when the oxidizing
potentials of the atmospheres are the same;
i.e., when the two  atmospheres are in equil-
ibrium with each other.

   An  important  conclusion is that  the
maximum pressure of SO2 from any system in
which calcium sulfate is reduced or calcium
sulfide is  oxidized  is the pressure  of SO 2
observed along either the H2 -H2O coexistence
line or the CO-CO2 coexistence line. The same
is true for 82  pressures. The basis for  these
rather far-reaching conclusions is that in any
process involving  a reduction  of  calcium
sulfate, the SO2 and 82 pressures will increase
with increasing reducing ability of the atmos-
phere until calcium sulfide is formed. At that
point, increasing the reducing ability of the
atmosphere no longer increases the amount of
SO2 or 82 formed, but rather causes  the
calcium sulfate to be transformed into calcium
sulfide.  Similarly,  in  a process  in which
calcium sulfide is oxidized, the SO2 and S2
pressures increase  with increasing oxidizing
ability of the atmosphere until calcium sulfate
is formed. This point is again a limit, and SO2
and 82 pressures cannot be increased further.

   2. Roasting of Calcium Sulfide
   In the roasting process, calcium sulfide is
oxidized with  oxygen  or air  according  to
reaction 14.

      CaS + 3/2 O2 - SO2 + CaO
                                                                 (
                           3/2
                                      (14)
As may be seen by examining the equilibrium
constant for reaction 14, the pressure of SO 2
at any temperature increases with increasing
pressure of oxygen. However, it follows from
the  arguments presented  above that above
some  definite oxygen  pressure (given  at any
temperature by Kp of reaction 15),

       1/2 CaS + O 2 -*l/2 CaSO4
                                                                                    (15)

                                              calcium sulfide is no longer stable,  but is
                                              converted to calcium sulfate. At this particular
                                              oxygen pressure,  the  SO2  pressure is  that
                                              observed along the coexistence line in the CO-
                                              CO2 system or in  the H2-H2O system. Thus,
                                              what appear to be two very different processes,
                                              the reductive decomposition of calcium sulfate
                                              and the roasting of calcium sulfide, are in fact
                                              very  similar.  Both processes  give  rise  to
                                              identical maximum SO 2 pressures at a given
                                              temperature.
1-1-16

-------
   3.  Pressure Effects in the Above Processes

   In  all  regeneration  processes  discussed
above, the pressure of SO2 is a function of the
temperature or of the oxidizing ability of the
atmosphere  (in the case of the roasting and
reductive decomposition processes). The pres-
sure of SO 2  as a function of the total system
pressure has not been discussed because SO2
pressure  is  independent  of the  total system
pressure in these processes. However, percent
of SO2 in the gas mixture is an inverse func-
tion of the  total system  pressure since  the
pressure of  SO2 is fixed at any temperature.
The pressure of SO2 is also independent of the
presence of inert gaseous diluents if sufficient
oxidizing or reducing gas is present. As stated
above, in the case of reductive decomposition
with CO-CO2 mixtures, the presence of inert
gases may affect the CO2  pressure enough to
change the reaction product from  calcium
carbonate to calcium oxide.

   4.  Acid-Base Reaction  of Calcium Sulfide
with H2O and CO2

   Reaction  16 has been  proposed  as a
regeneration reaction for  calcium sulfide.

     CaS + H2O + CO2  k CaCO3 + H2S
                       H2S
           KP ~  P
C02' PH2O
                                      (16)
Calcium sulfide is formed in the additive by:
(1) burning coal in a fluidized bed of limestone
or dolomite with a deficiency of air or by (2)
reducing the CaSO4 in the additive from a run
in which combustion was with an excess of air.
Unlike  all  other regeneration reactions dis-
cussed here, this reaction is pressure-sensitive.
The percentage and  the pressure of  H2S
increase with increasing total system pressure.
The pressure of H2S  is also sensitive to the
presence of inert-gas diluents, in contrast to
the  previously  mentioned  regeneration
schemes. The equilibrium constant for this
exothermic reaction becomes smaller  as the
                          temperature is increased (Figure 10). This is in
                          direct  contrast to  the  other (endothermic)
                          regeneration  schemes mentioned.

                             Maximum H2S yield is obtained when the
                          H2O/CO2  ratio in the feed gas is 1 to 1, as
                          may be seen by an examination of the equil-
                          ibrium constant expression. Figure 11 shows
                          the pressure  of H2S  as  a function of tem-
                          perature, assuming 10-atm total pressure and
                          an inlet gas stream composed only of H2O and
                          CO2 at various ratios. However, H2O may be
                          readily removed from the product gas stream
                          by condensation. Thus, higher values of H2S
                          concentration in a dried  gas  stream may be
                          obtained by operating with an excess of H2O
                          in the inlet  gas  stream. For an  inlet  gas
                          composition  of 50  percent  water  and  50
                          percent  CO2   and  temperatures  of  1000-
                          1400 F, Table 3 gives the percentage of H2 S
                          in the  gas  effluent from  the reactor and  the
                          percentage of H2S in the same effluent after it
                          has been dried.

                          TableS. H2S CONCENTRATION9 IN DRIED AND
                          UNDRIED PRODUCT GAS  STREAM AT EQUIL-
                          IBRIUM
Temperature,0 F
1000
1100
1200
1300
1400
% H2S in
undried gas
23.0
11.4
4.7
2.7
1.4
% H2S in
dried gas
37.4
20.5
9.9
5.1
2.8
                          Assumptions are 10-atm total pressure with  50
                           percent H20 and 50 percent C02 inlet gas.
                          B. Experimental Studies

                             4.  Reductive Decomposition of CaSO4

                             To test the  accuracy  of the  equilibrium
                          compositions calculated for the reduction of
                          CaSO4 with CO/CO2 mixtures,  experiments
                          have been performed in a static  system. The
                                                                                  1-1-17

-------
 apparatus consists of a horizontal tube reactor
 fabricated  from recrystallized alumina. The
 tube is 36 inches long and has an ID of 3
 inches with  1/4-in. walls. One end of the
 alumina reactor is closed and the opposite end
 is capped with a stainless steel  0-ring flange.
 The flanged  end is outside the furnace.

   The experiments are  performed in the
 following manner. A 3-gram sample of CaSO4
 (Drierite) is placed  in an alumina boat and
 loaded into the reactor. The system is closed
 and leak-checked.  The  CaSC>4 is  dried at
 500F under vacuum for 15 to 20 hours. While
 the system is still at 500F and  isolated from
 the vacuum pump, a predetermined pressure
 (
-------
   b. Reaction of CaS with CO2/H2O  The
product of each of the reduction experiments
was  carbonated  at  10  atm to  simulate  a
product that would be obtained in an actual
10  atm  combustion-reduction  experiment.
This material was then reacted batchwise with
an  equimolar   mixture  of CCh/HbO  at
temperatures ranging from 900 to 1100F, at a
gas velocity  of approximately 1  ft/sec and at
10  atm  pressure  in  the   2-in.  diameter
fluidized-bed reactor. The H^S  concentration
in the  outlet  gas  was  monitored  using  a
quadrupole  mass spectrometer.

   The  results to date have shown that:
   1. The reaction producing H2S is initially
rapid, but  the  rate decreases  after a  short
time. Typically, the reaction rate drops to near
zero after several minutes.
   2. The peak  concentration of  H2S in the
outlet  gas  is  high  and near  the  expected
equilibrium value.
   3. Typically, half or less of the CaS reacts.
In  continuing work, the effects of process
variables  are being studied  in an attempt to
increase the quantity of CaS that is reacted.

C. Sulfation-Regeneration   Cyclic
Experiments

   Since  it  will  be desirable  to  reuse  the
additive material several times in  commercial
application's, a  cyclic experiment has been
performed to obtain data on the pickup and
removal of sulfur from additive particles and
to determine decrepitation  and attrition  of
additive    particles   during  sulfation-
regeneration cycles. Six cycles  of simulated
combustion  and two-stage regeneration were
performed with a single bed of additive. The
starting material (1.2 kg) was obtained from a
coal combustion experiment in which dolo-
mite No. 1337 had been  used as additive. The
initial  sulfur content of  the bed was  15.4
weight  percent.  The experiment was  per-
formed  batchwise in the 2-in. diameter fluid-
ized-bed reactor:

   The sulfation portion of cycle  1 was omitted
since the  additive already contained sulfur.
For the remaining cycles, the constituents of
the sulfating gas were N2, CC2, H2O, Ch, CO,
and SO?. The sulfation reaction was allowed
to  proceed  until  the bed material  had
essentially ceased further pickup of SO2. After
the bed had been sulfated, the CaSCM was
converted   to  CaS,  using  H2  or  CO  as
reductant at 1550 to 1600F and 10 atm. The
bed was then  reacted with a CO2/H2O gas
mixture at  1000F and 10 atm to convert the
CaS to CaCOa. A sample of the bed material
was taken  after  each step  in the  cycle and
analyzed for sulfur and sulfide  content.

   The effluent gas stream was  analyzed for
H2S  concentration,  using  the quadrupole
mass  spectrometer. A plot of  H2S  concen-
tration versus reaction time in the six cycles is
presented in Figure 15.

   The results indicated that the conversion of
CaSO4  to  CaS  in the reduction step was
ineffective.  Only  in cycles  1 and 4 was the
conversion  to CaS greater than 50 percent. A
possible cause could  be the interaction  of
CaSO4 and CaS to form nonporous surfaces;
the formation of easily sinterable  cakes has
been  reported when  these materials  are
present.

   The  data  for  the  regeneration  step
(presented in Figure 15) showed  that the  peak
concentrations of H2S in  the   effluent gas
decreased from 13 volume percent (dry basis)
for cycle 1 to 0.5 volume percent (dry basis) for
cycles 5 and 6. The percent calcium  sulfide
converted to CaCOs decreased to a very low
indeterminate value after  several  cycles. It
appears from these  data  that a  layer  of
material of low permeability is built up on or
within  pores of  the  additive  particles,
inhibiting the removal of sulfur.  The high sul-
fur loading of the  bed  particles in  these
experiments may be  a factor  in the  poor
conversion.

   In  continuing  work,  it is  planned  to
investigate  the  use  of  a high  reduction
temperature (e.g., 1800F)  to promote more
complete reduction and to remove part of the
                                                                                   1-1-19

-------
 sulfur as SO 2. The remaining sulfur would
 then be removed as H2S by reaction with CO2 -
 H2O. The two gas streams could be combined
 for conversion to  elemental sulfur.
 PRESSURIZED  COMBUSTION  AND RE-
 GENERATIONPILOT PLANT DESCRIP-
 TION

   Equipment   has  been   installed  for
 combusting coal at pressures up to 10 atm and
 for continuously regenerating sulfated lime for
 reuse. A  simplified  equipment schematic  is
 shown in  Figure 16.  The regenerator and the
 fluidized-bed combustor have  a common off-
 gas  system (cyclones,  filters,  gas-sampling
 equipment, pressure  let-down  valve, and
 scrubber)   and   will   not   be    operated
 simultaneously.   Either  the  combustor  or
 regenerator will be disconnected from the off-
 gas line and flanged off when the other  unit is
 in operation.

    The combustion unit consists of a 6-in.
 schedule 40 pipe (Type 316 SS) approximately
 11 ft long, with an outer shell consisting of 12-
 in. schedule 10 pipe (Type 304  SS) over  nearly
 the entire length. A bellows expansion joint is
 incorporated  into   the  outer  shell   to
 accommodate  the  differential   thermal
 expansion of the inner and outer vessels.

   The unit is of a balanced pressure design;
 i.e.,  the  annular chamber between  the two
 pipes is maintained under pressure so  that a
 differential pressure does not exist across the
 hot inner pipe wall. The balancing  pressure
 for the shell is supplied by a bank of nitrogen
 cylinders.

   A  bubble-cap-type  gas  distributor   is
 flanged to the  bottom end of the inner vessel;
thermocouples, solids feed lines, and  solids
take-off   lines  extend  through   the  gas
distributor. The outer wall of the 6-in. pipe is
wrapped   alternately  with  sixteen  3000-W
tubular resistance heaters  and  3/8-in. OD
cooling  coils  that  are  spray-metal-bonded.
Internal cooling coils of 3/8-in. pipe extend
down into the  interior of the 6-in. vessel from
 the  flanged top  to provide additional heat
 transfer area. Water flow to the cooling coils is
 regulated using flow indicators and is adjusted
 on the basis of the temperatures of the fluid -
 ized bed  and  reactor wall.

   Both the annular pressure chamber and the
 reactor itself are  equipped with rupture disc
 assemblies and pressure relief valves vented to
 the room ventilation exhaust ducts.

   The regenerator has a 3-in. ID surrounded
 by 2-1/2 inches of Plibrico castable refractory
 and encased in an 8-in. schedule 40 pipe (316
 SS).  This entire  assembly is enclosed by a
 pressure  shell  made  of  12-in.  schedule 20
 carbon  steel   pipe.   Differential  thermal
 expansion between the inner and outer pipes is
 accommodated by the use of packing  glands
 on the lines entering the bottom flange of the
 unit.  The unit is  of a  balanced  pressure
 design; i.e., the annular chamber between the
 two pipes is maintained under pressure so that
 a  large differential pressure  does  not exist
 across the hot inner pipe wall. The balancing
 gas is nitrogen. Since the annular space is not
 gas tight with respect to the regenerator inner
 vessel, the pressure  in the annular space will
 be maintained  slightly  higher  than  the
 regenerator pressure to prevent process gases
 from entering  the annulus. A pressure alarm
 gauge will monitor the pressure in the annular
 space and will  be set to warn of both high and
 low pressure.
   A  bubble-cap-type  gas   distributor   is
connected to the bottom of the inner vessel via
a slip fit and  held  in place  with  retaining
screws. Thermocouples, solids feed lines, and
solids take-off lines  pass  through  the gas
distributor and then through packing glands
on the bottom flange of the  outer  pressure
vessel. The wall of the inner vessel is wrapped
alternately with 3000-W  tubular resistance
heaters and 3/8-in. OD tubing coils. Both the
annular chamber and the  regenerator  itself
are equipped with rupture disc assemblies and
pressure  relief valves vented to  the  room
ventilation exhaust ducts.
1-1-20

-------
   The primary filter cartridges are suitable     ASME code requirements . The design rating
for  temperatures  up  to   350F  (epoxy-     of the unit is 150 psig at 1500F. Air (or gas)
impregnated cellulose-base material with glass     passes through an annulus, reverses direction,
fiber  substrate).  The  secondary  filter     and passes through  a heated  section.
cartridges are Rigimesh (woven metal wire).'

   The gas preheater is of a balanced pressure       The feeders are of the rotary pocket type
design and was designed in accordance with     equipped with  hoppers.
                                                                                   1-1-21

-------
                                                                             TO
                                                                          GAS-ANALYSIS
                                                                             SYSTEM
 PREHEATER
                                                                     TO GLASS FIBER
                                                                     FINAL FILTER AND
                                                                     VENTILATION EXHAUST
                             Figure 1.  Bench-scale combustor system.
1-1-22

-------
    100
     90L_-
     70
     60
o
5
LU
ce
oc
                                      PITTSBURGH	
                                         COAL
      ILLINOIS
        COAL
SYMBOL
                                     A
                                     A
                                     
                                     O
                                     9
  RUN NO.
  HUWP-1A,
-ID, -2A, -2B
   AR-4, -5
   AR-1E
   BC4
   BC-9
   BC-10
             TEMP.,  ADDITIVE
COAL TYPE      T      NO.
                          PITTSBURGH   1450  1359
                          ILLINOIS
                          ILLINOIS
                          ILLINOIS
                          ILLINOIS
                          ILLINOIS
              1550   1359
              1600   1359
              1600   1359
              1600   1360
              1600   1377
                                  AVERAGE PARTICLE SIZE RANGE FOR ADDITIVE:  490-630 jjm
                                  GAS VELOCITY IN OOMBUSTOR: 2.6 TO 2.8 ft/sec
                                2            3           4            5
                                            Ca/S MOLE RATIO
                        Figure  2.  Effect of Ca/S mole ratio on.sulfur retention.
                                                                                         M-23

-------
    UJ
    LLJ
    Cd
         100
          90
          80
          70
          60
          50
          40
          30
          20
                          O ILLINOIS COAL, Ca/S~2.5

                           PITTSBURGH COAL, Ca/SM.O
1300
1400
                                                                 GAS VELOCITY, 3 ft/sec
                                                                 EXCESS OXYGEN, 3%
                                                                 LIMESTONE NO. 1359
                                                                                          J
1600
                                                    1500

                                        TEMPERATURE, F

                  Figure 3.  Effect of fluid! zed-bed temperature on sulfur retention .
                                                          1700
1-1-24

-------
100
 90
 80
 70
 60
 50

 40
Ca/S=l
 30
  20
 10
              CRE DATA (1470 T)

                                  WELBECK COAL, BRITISH LIMESTONE (440
              ANL DATA (1550 "F)
                                 ILLINOIS COAL, LIMESTONE NO. 1359 (>1000 pi)
                           A    ILLINOIS COAL, POINT TAKEN FROM CURVE (FIGURE 2)
                                  FOR LIMESTONES AND DOLOMITE (*630jim), Ca/S = 4
                      2                  4                  6

                               SUPERFICIAL GAS VELOCITY, ft/sec
            Figure 4.  Effect of superficial  gas velocity on sulfur retention.
                                                                                      1-1-25

-------
              70
              60
              50
CO
+
CM
^
           "3.
           in
              40
              30
              20
              10
               0
                               BED TEMPERATURE

                                   P1450F
                                   O1550F
                                      1600 F
                                      SECONDARY AIR INTRODUCED
                                           LIMESTONE ADDITION
                                           	NO LIMESTONE ADDITION     	
      H
                0

            Figure
                50
          60            70            80
AIR, % of stoichiometric (based on feed rates)
90
         5.  Effect of air feed rate on percent of sulfur in off-gas as H2S.
1-1-26

-------
o
I-
LU
a:
cc
                                                                 O 63
                                                                    O  58
                               O 53
                                                                              O 60
                                                              NUMBER NEAR POINT IS THE
                                                              PERCENT OF STOICHIOWIETRIC
                                                                  AIR ADDED
 SINGLE-STAGE OXIDIZING
     EXPERIMENTS
    (1450-1470 F)


BED TEMPERATURE, SUBSTOICHIOMETRIC
AIR EXPERIMENTS

   [] 1450F

   O1550F

   A1650F
                                           Ca/S MOLE RATIO

         Figure 6.  Sulfur retention in oxygen-excess and oxygen-deficient experiments.
                                                                                       1-1-27

-------
    100
     90
      80
      70
      60
      50
      40
      30
      20
      10
                      50
                                             o
                   BED TEMPERATURE
                      D  1450 F
                      O  1550F

                       A  1650F
60           70           80
         AIR, % of stoichiometric
90
100
110
                 Figure 7.  Effect on sulfur retention of air-feed rate to first stage.
1-1-28

-------
    300
     250
     200
     150
     100
      50
      40
                                                            1450 F
                                                                   1650F
                                                                               1550"F
        40
50
60             70             80

 AIR,  % of stoichiometric (based on feed rates)
90
100
Figure 8.  Effect of air feed rate and fluidized-bed temperature on NO concentration  in off-gas
from the first stage during combustion of coal.
                                                                                            1-1-29

-------
  2400
  2320
                                                                    6.0  5.0     4.0   3.40

                                                                7.0   I 5.5 |   4.5 j 3.70 J  3-0
  1600
                                                                                               0.001
    0.005    0.010     0.015    0.020    0.025    0.030      0.035    0.040     0.045    0.050     0.055

                                           PCO/PC02

    Figure 9.  Pressure of S02 in equilibrium with CO/C02 mixtures as a function of
    temperature  (10 atm total  pressure).
1-1-30

-------
                     FOR CaS +H20 +C02^H2S + CaC03
      1000
1100
 1200             1300
TEMPERATURE,F
Figure 10.  Equilibrium constant as function of temperature, CC>2 +1^0 + CaS-*CaCC>3 +H^S-
                                                                                   1-1-31

-------
                         2.0
                                        H20 +C02 + CaS;CaC03 +H2S
                                    1100       1200

                                     TEMPERATURE, F
                         Figure 11. Pressure of H2S in undried gas
                         stream as function of temperature at
                         10-atm Ptotal-
1-1-32

-------
   0.18f


   0.17



   0.16



   0.15


   0.14



    0.13


    0.12



    0.11
E
eg

  0.10



^  0.09
 ;
_i
    0.08


    0.07


    0.06


    0.05



    0.04


    0.03


    0.02


    0.01


    0.00
                                           TEMPERATURE:  1900 F
                                        TOTAL PRESSURE:  1 aim
                             A-14
                               1-3/4 hr
                  A-9,1 hr
                                                    A-13,2 hr
                 A-12,2 hr
                                     , A-14,20-3/4 hr
                              A-15
                               Ihr
                          C0/C02= 0.034'
                                A-8, 1 hr
                      A-7, 1.5 hr
   A-9,19-1/4 hr


CALCULATED EQUILIBRIUM PARTIAL
    PRESSURE OF S02 FOR
   CaS04-j.CO  CaO+C02 +S02
                                                       A-15,18-1/4 hr
                                                                               \      I      I
0.002
0.004
0.006
0.008
                                                                 0.010
                      0.012
                                                                                          0.014
               Figure 12.  Partial pressures of S02 over a range of PCO/PC02 ratios.
                                                                                              1-1-33

-------
                      CONDITIONS
 o
 UJ
 CQ
 UJ
      20
      15
      10
 o
 o
 g   5
             CATS-16  100%  H2, 1 atm, 1600 F
             CATS-18  100%  H2, 1 atm, 1450 F
             CATS-19  100%  H2, 1 atm, 1350 F
                100
      200

TIME, min
300
 Figure 13.  Sulfide content of bed during re-
 duction of partially sulfated dolomite with
 hydrogen at various temperatures.
                100


                 90


                 80
             ss

             g   70
             u_
             _j
             3   60


                50
                                                  LL.
                                                  O
                                                      40
                                 30
                              
                              LU
                                                   o
                                                   o
                                 20
                                                      10
                                                           CATS-17
                                    CATS-190-
                                   REDUCTANT
                                   o HYDROGEN
                                   CO
     1200  1300  1400  1500  1600  1700   1800  1900

                 TEMPERATURE, F
Figure 14.  Effect  of temperature on reduction
of CaS04 (dolomite), 4.5 hours of reduction
time.
1-1-34

-------
s
CO
=C
CO
UJ
                                                           QCYCLEl


                                                           D CYCLE 2


                                                            CYCLE 3
A CYCLE 4

 CYCLE 5


A CYCLE 6
to    7 
       0    2      4
                                            TIME, min
                            30
          Figure 15.  H2S concentration in effluent gas stream in regeneration step.
                                                                                     1-1-35

-------
                REHEATER
H. P. STEAM
                                                                                                    PRESSURE
                                                                                                    LET-DOWN
                                                                                                      VALVE
                                                                                    ROOM VENTILATION
                                                                                           AIR
                                                        H2   CO  C2H2
                 Figure 16.  Simplified schematic of combustion-regeneration equipment.

-------
                 2. A REGENERATIVE LIMESTONE PROCESS
                                        FOR FLUIDIZED-BED COAL
                       COMBUSTION  AND DESULFURIZATION

                  R. C.HOKE, H. SHAW, AND A. SKOPP

                Esso Research and Engineering Company
ABSTRACT
   The factors influencing NOx emissions from a fluidized limestone bed coal combustor were
 studied. NOx emissions were decreased by a decrease in temperature and a decrease in excess air.
 Sulfated lime depressed NO x emissions compared to an inert alundum bed. An apparent cause
 of these effects is reduction of NOX by CO. The NO/CO reaction was then studied further in fixed
 bed units. CaSO4 catalyzed the reaction slightly compared to alundum in a dry system. CaO
 promoted the reaction significantly, giving over 90 percent conversion in the absence of CO2- CO 2
 was found to inhibit the rate, possibly due to a kinetic limitation caused by the presence of the
 COa. Study of the reaction of SO 2 and NO  indicated that the reaction  is catalyzed by partially
 sulfated lime, but not by CaSO4 or alundum. Temperature was found to have a negative effect on
 the reaction, apparently due to  the  thermal instability of an intermediate,CaSO3. Two-stage
 combustion of coal was studied to promote the CO/NO reaction and reduce NO emissions further.
 NO  emissions were  reduced by two-stage  combustion, and the  reduction was  enhanced by
 operating the first stage at lower air levels.

   Regeneration of CaSO 4 to CaO and SO2 by CO and H2 was studied at pressures up to 9.5 atm.
 SO2 levels in the off-gas  as high  as 7.5 percent were measured at pressures up to 6 atm. The
 maximum concentration measured to date at 9.5 atm is 2.2 percent. The measured levels are 50-60
 percent of the levels calculated from equilibrium considerations.

 INTRODUCTION

   Esso  Research and Engineering Company
 is conducting an experimental program for the
 Environmental Protection Agency under con-
 tract CPA 70-19 to  develop a regenerative
 limestone process for  fluidized-bed coal com-
 bustion and desulfurization. This  is a  part of
 EPA's overall program to  examine fluidized-
 bed combustion as  a possible new power
 generation technique. The potential of fluid-
 ized-bed combustors for air pollution  control
 is good  because  the  intimate gas-solid  con-
tacting in a fluidized bed promotes high SO2
 removal  efficiency on suitable materials such
as limestone or dolomite.

   A schematic diagram of the process is
shown in Figure 1. In the combustor, the sul-
fur in the coal is burned to SO2 which then
reacts with the lime to form CaSO4. The
system being studied by Esso  involves trans-
ferring the partially sulfated  lime from the
combustor to a separate regeneration vessel
where the sulfated lime is regenerated accord-
ing to the reaction

      CaSO4+CO -> CaO+SO2+CO2
              H2             H20.    (1)
                                      1-2-1

-------
 The regenerated stone  (CaO) can  then be     regenerator has a high SO2 concentration and
 returned to the  combustor for further  use,     can be used as feed to a by-product sulfur or
 thereby substantially reducing the fresh lime-     sulfuric acid plant.
 stone  requirement.  The off-gas from  the
1-2-2

-------
Previous Studies

   Various laboratories  including  Esso
Research  have  studied fluidized-bed  coal
combustion over the  past  few  years.  The
results of the studies are summarized in Table
1 and  have shown  that coal can be  burned

    Table 1. PREVIOUS FBC FINDINGS

Coal combustion efficiency high
Over 90% removal of S02
NOX emissions reduced
Sulfated lime can be regenerated
    CaS04 + CO -* CaO + S02 +  C02
Activity maintenance of recycled lime satisfactory
    after 7 cycles
Pressurized  FBC system more attractive

efficiently with over 90 percent removal of SO 2
and with reduced NOx emissions.  Regener-
ation of sulfated limestone has been studied
using a number of regeneration methods. The
method studied at Esso Research consisting of
the one step reduction  of CaSO4 to  CaO and
SO2 gives 6-10 percent SO2 in the product gas
when carried out at 1 atm and about 2000F.
The recycled lime was also shown to maintain
a reasonably high level of activity after seven
combustion/regeneration cycles.

   Economic studies  were  carried out by
Westinghouse Research  Laboratories under
contract to EPA.1 These studies indicated that
operation of the combustor and regenerator at
higher  pressures,  approximately  10  atm,
would  be significantly  more  economical  than
atmospheric pressure operation. As a result,
the current studies  are being made  at higher
pressures.

Objectives
   Objectives  of  Esso  Research's  current
experimental program are  summarized  in
Table  2 and consist of (1)  investigating the
factors  influencing the  reduction  of  NOX
emissions in fluidized-bed combustion,  and (2)
studying the regeneration of sulfated lime at
pressures up to  10 atm. The latter objective
Table2. OBJECTIVES OF CURRENT EXPERI-
MENTALWORK	
1. Investigate  factors influencing  reduction  of
  NOX emissions
    Effect  of  temperature,  excess  air, bed
    materials
    Reaction of NO with CO and S02
2. Study regeneration of sulfated lime
    S02 levels attainable at higher pressures
    Kinetics of regeneration and stone activity
    maintenance
has required construction of higher  pressure
experimental equipment.

EXPERIMENTAL EQUIPMENT
    A number of experimental units were used
in the  current program. A flow diagram of the
atmospheric  pressure fluidized-bed  combus-
tion is shown in Figure 2. The reactor is a 3-
in.-ID Incoloy tube. Four continuous flue gas
analyzers are used including IR SO2 and CO
analyzers  and polarographic  NOx  and O2
analyzers. A  high pressure regeneration unit
capable of operating up to 10 atm was recently
built and  is  shown in Figure  3. The reactor
consists of a  3-in.-ID  alumina tube contained
in a 12-in. carbon steel vessel. The reactor is
15  feet long.  The interior of the steel vessel is
lined with 4-1/2 inches of castable refractory
insulation. The  regeneration  feed   gas  is
produced by  combustion  of propane. N2 and
CO2 can also be added to the burner to adjust
the composition of the regeneration gas. Most
of the input heat is provided by combustion of
propane, but  additional heat input is provided
through  electrical  heaters adjacent  to  the
alumina tube and an air preheater. The unit is
heated by burning the propane under  excess
air conditions. When  the operating tempera-
ture has  been reached, the air/fuel  ratio is
instantaneously changed to substoichiometric
conditions and  N2/CO2  flow  is  started,
thereby assuring a rapid change from heat-up
to operating  conditions.

   Two small fixed bed units were also used in
these studies. Three different reactors were
used: a 2-1/2-in. alumina tube operating at 1
atm, a 1-in.  stainless  steel tube  operating at
                                                                                    1-2-3

-------
pressures up  to 10  atm, and a  specially
constructed  regenerator.   The  regenerator
consisted of a 1-in. alumina tube contained in
a 3-in. pipe with fiber insulation between the
pipe  and tube. This  unit was capable  of
operating at pressures up  to 9.5 atm at tem-
peratures up to 2000 F. All fixed  bed units
were  electrically heated.

EXPERIMENTAL RESULTS

Factors Affecting NO x Emissions

   It  was  determined  previously that NOX
emissions measured  at the low temperatures
occurring  in  fluidized-bed  combustion  are
formed by oxidation of nitrogen compounds in
the coal. Oxidation of atmospheric N2 occurs
only at higher temperatures. In this study, the
effects of temperature,  excess air, and fluid-
ized  bed material  on  NO emissions  were
measured. The effect of temperature using a
bed of CaSO4 in the combustor is  shown in
Figure 4. As temperature decreased, NO emis-
sions  dropped rather sharply below  1500F.
The effect of excess air using a bed of CaSO4 is
shown  in  Figure  5.  Actual  NO  emissions
decreased  as  excess air  (percent   O2)  was
increased. However, when  the emissions were
normalized  to  a constant gas volume (at 3
percent 2), the NO emissions increased as
the excess  air increased. The NO formation
rate was thus increased by the higher average
oxygen concentration in the bed. The effect of
bed material is shown in Figure 6. CaSO4 gave
lower emissions than alundum. With a CaO
bed the emissions were high initially, but as
the bed sulfated the emission level approached
that of CaSO4.

   One consistent explanation for these results
is  the reaction of NO with  CO.  Carbon
monoxide emissions are higher at the lower
temperatures and at lower excess air condi-
tions.  The higher CO levels  then give lower NO
emissions. The effect of bed materials appears
to be  a catalytic effect.

Reactions of NO and CO

   The reaction of CO  and NO was studied

1-2-4
further in fixed-bed units. The effects of bed
material, temperature and feed gas composi-
tion were studied. In a  dry  system,  CaSO4
catalyzed the reaction slightly arid showed a
small effect of temperature, but alumina  and
an  empty  bed  gave essentially no reaction.
This  is  shown  in Table  3. However,  the
Table 3. NO-CO REACTIONS-EFFECTS OF
TEMPERATURE AND BED MATERIAL
Bed material
Bed temperature, "F
Inlet gas composition
NO, ppm
CO, ppm
NO conversion, %
CaSO4
1500

1990
4000
4
CaSO4
1700

2025
4000
6
Alumina
1500

2010
4000
1
None
1500

2035
4000
0.5
addition of water enhanced the reaction and
gave the same NO conversion regardless of the
presence of the bed material. This is shown in
Table 4. But when CaO was used  as the bed
Table  4. NO-CO  REACTIONS-EFFECT  OF
WATER VAPOR
Bed material
Inlet gas composition
NO, ppm
CO, ppm
H20, %
NO conversion, %
CaS04
1990
4000
0 7.2
4 12
Alumina
2010
4000
0 7.2
1 13
None
2035
4000
0 7.2
0.5 15
Temperature, 1500"F


material in a dry system, a very rapid reaction
occurred which gave over 90 percent conver-
sion of the limiting reactant as shown in Table
5. The reaction proceeded in 1:1 mole ratio of
CO and NO suggesting the reaction
      2 CO + 2 NO
2CO
(2)
   Carbon dioxide was then added to the feed
and reduced the conversion significantly over

-------
both  calcined  limestone  and  calcined
dolomite. This is shown in Table 6,

Table  5. NO-CO  REACTIONS-EFFECT OF
CaO
Bed source
Inlet gas composition
NO, ppm
CO, ppm
Outlet gas composition
NO, ppm
CO, ppm
Conversion, %
Lime #1359
1400
940
400
10
99
1800
1870
20
160
99
Dolomite #1337
1400
900
350
20
98
1990
2080
240
100
95
 Temperature, 1600"F

 Residence time, 0.3 sec

 Table  6. NO-CO  REACTIONS-EFFECT  OF
 CO 2
Bed source
Inlet gas composition
NO, ppm
CO, ppm
CO2, %
Conversion, %
Lime #1359

1400
940
0
99

860
990
17
26
Dolomite #1337

1400
900
0
98

840
980
16
19
 Temperature, 1600F

 Residence time, 0.3 sec
Table  7. NO-CO  REACTIONS-EFFECT  OF
PRESSURE AND RESIDENCE TIME
Bed source
Pressure, atm
Residence time, sec
Inlet gas composition
NO, ppm
CO, ppm
CO,, %
Conversion, %
Lime #1359
1
0.3
860
990
17
26
10
3
890
980
18
82
10
0.3
1150
1240
13
16
Dolomite #1337
1
0.3
840
980
16
19
10
3
840
980
16
88
                                              Temperature, 1600"F
TableS. NO-CO REACTIONS-EFFECT OF O-
Bed source
Inlet gas composition
NO, ppm
CO, ppm
CO-., %
02, %
Conversion, %
Lime #1359

1150
1240
13
0
84

970
1060
16
2.4
95
Dolomite #1337

840
980
16
0
88

980
1080
15
2.3
91
                                               Temperature, 1600F

                                               Pressure, 10 atm
were considered as possible explanations, but
were ruled out after closer examination.
   The effect of  pressure was  studied  and
 although increasing the pressure to 10 atm in
 the presence of CO 2 apparently increased the
 conversion, the increase was  probably due to
 increased  residence  time.  At  equivalent
 residence times, increasing pressure appeared
 to decrease  the conversion slightly. This  is
 shown in Table 7. Changing the background
 gas from N2 to argon appeared to increase the
 conversion very slightly, but the effect may not
 be significant. Oxygen was added to the  feed
 and appeared to increase conversion slightly.
 This  is shown in Table 8.
   The  most  likely explanation for  these
 effects is a kinetic  limitation caused by the
 presence of the CO2. Formation of CaCOa
 and inhibition caused by chemical reversibility
Reactions of NO and SO2

   Further studies of the reaction of NO and
SO 2 were made in a fixed-bed reactor. The
effects of bed material and temperature were
studied. The effect of bed material is shown in
Table 9. The results show that NO and SO2
did  not  react  in  the  vapor phase  or over
alundum or CaSO/^. However, a reaction  did
occur  over  partially  sulfated  lime  and
appeared   to   be  dependent  on  SO2
concentration. Further rate studies indicated a
0.5 order dependence  on the NO concentra-
tion. Temperature had a negative effect on the
rate, decreasing the rate with increasing tem-
perature, as shown in Figure 7. A proposed
mechanism for  the   reaction involves  the
reversible formation of CaSOs  intermediate
from CaO and SO 2- The sulfite then  reacts
                                                                                    1-2-5

-------
Table 9. NO-SO2  REACTIONS-EFFECT OF
BED MATERIAL AT 1600F
Bed material
Gas phase
Partially sulfated
limestone
Alundum
CaSO
NO concentration, ppm
Before SO 2
introduced
900
840
830
820
860
After SO 2
introduced
900
440
180
820
860
SOZ concentration, ppm
Inlet '
1290
780
1510
1000
670
Outlet
1290
300
480
1000
670
with NO to form N2 and CaSO4. However, it is
known that the sulfite becomes unstable in the
temperature  range  where  the  SCb/NO
reaction  rate drops; this instability is the
probable explanation for the negative temper-
ature effect.

Two-Stage Combustion

   The reactions of NO with CO suggest the
possible   lowering  of   NO  emissions  by
operating a staged combustion  system. Air
would be injected  at  two  points  in the
combustor giving an 02  lean section at the
bed inlet which should promote NO reduction
because of the relatively high CO levels. The
second step would then complete combustion.
The fluid-bed combustor was  modified to
operate in a staged fashion by injecting second
stage air 6 inches above the grid. The results of
two runs are shown in Figure 8. As the ratio of
the second stage air to the first stage air was
increased, the NO emissions dropped. Nitric
oxide emissions  were lowered to 200  ppm.
Although these conditions may not be feasible
in commercial  operation, the  principle of
stage combustion appears attractive.
Regeneration of Sulfated Limestone

   Regeneration  studies were carried  out  in
fixed and  fluidized  beds  using  CaSO4  at
pressures up to 9.5 atm. The results are shown
in Table 10.
   Concentrations of SO2 in the off-gas  as
high as 7.5 percent have been measured  at
pressures up to 6 atm. At 10 atm,  the highest
SO 2 concentration measured to date was a
little over 2 percent. Comparisons were also
made  with  SO 2  levels  estimated  from
equilibrium calculations  made by  Argonne
National Laboratory.2 The equilibrium  SO2
partial pressure is determined by the tempera-
ture and the CO/CO 2  ratio in  the  gas  in
equilibrium with the solids. However, at each
temperature, there is a CO/CO2 ratio which
gives the maximum  attainable SO2  partial
pressure for the  temperature in question. In
the fixed-bed  runs, the off-gases were not
analyzed for CO and CO 2, and the measured
SO2 concentrations had to be compared to the
maximum equilibrium SO2  concentration. In
the  fluidized-bed  runs,  comparisons  were
made at the actual CO/CO 2 ratio measured,
although these ratios were probably in error
due to  oxidation  of CO in the exit lines from
the reactor. The comparison of measured and
calculated SO2 levels is given in Table 10. The
SO2 levels  measured  in  the fixed-bed  unit
were less than 50 percent of the  maximum
attainable at the temperatures  of the runs.
The results  from the fluidized bed  were closer
to the  equilibrium  concentrations  calculated
for the  CO/CO2 ratio measured for each run.
In general,  the measured SO 2 concentrations
were 50-60  percent of the equilibrium  levels;
   Table  10. REGENERATION  OF  SULFATED LIMESTONE, CaSO4 AS BED MATERIAL
Unit
Rxed
Rxed
Fluidized
Fluidized
Fluidized
Fluidized
Pressure,
atm
3
9.5
3.2
6.2
6.0
6.0
Temperature,
F
2000
2000
1990
2100
1950
1870
SO2 concentration,%
MeasursH Calculated
7.2
2.2
5.2
7.5
2.0
1.8
15.0
4.8
9.8
11.3
3.7
1.8
SO 2 ratio
measured/calculated
0.48
046
0.53
0.66
0.54
1.0
1-2-6

-------
the last run met the  calculated equilibrium
concentrations. Further work is planned in the
fluidized-bed regeneration unit  to determine
the SO 2 levels attainable at pressures up to 10
atm as  a function of temperature,  regener-
ation gas composition and flow  rate, particle
size,  and  sulfated  lime  source.  A   new
pressurized  combustor  unit  is  being  built
which will be used with the regenerator to
measure cyclic activity maintenance of various
stones.
BIBLIOGRAPHY

1. Archer, D. H., D. L. Keairns, J. R. Hamm,
   R.  A.  Newby,  W.   C.  Yang,   L.  M.
   Handman,  and L.  Elikan. Evaluation of
   the  Fluidized-Bed Combustion  Process.
  Westinghouse Research Laboratories,
  Pittsburgh, Pa. Prepared for the Environ-
  mental  Protection Agency,  Research
  Triangle  Park,  N.  C.  under  Contract
  Number CPA 70-9. November 1971.

2. Jonke, A.A., G.J. Vogel, J. Ackerman, M.
  Haas, J. Riha, C.B. Schoffstoll, J. Hepperly,
  R.  Green, and E.L. Carls. Reduction  of
  Atmospheric  Pollution  by the Application
  of  Fluidized-Bed  Combustion.  Argonne
  National  Laboratory,  Argonne,  111.  Pre-
  pared  for the  Environmental  Protection
  Agency,  Research  Triangle  Park, N.C.
  under agreement EPA-IAG-0020. Monthly
  Progress  Report  Number 38,  December
  1971.
                                                                                   1-2-7

-------
                          FLUE GAS AND
                          COAL FLY ASH
FLUID BED
COMBUSTOR
        FRESH
        SORBENT
        &COAL
                                      SULFATED
                                      SORBENT
                                    REGENERATED
                                      SORBENT
                                                                      HIGH S02 GAS
                                                                      TO BY-PRODUCT
                                                                      PLANT
                                                                  FLUID BED
                                                                  REGENERATOR
                                                                       DISCARDED
                                                                       SORBENT
                          FLUIDIZING
                          AIR
                                                           REDUCING
                                                           GAS
                                                               -CaO+S02+C02
                                                             H2          H20

       Figure 1.  Esso proposed fluidized-bed combustion-lime regeneration system.
1-2-8

-------
           CYCLONE
             AND
           FILTER
            FEEDER
            SCALE
nnnnnik
                               WTM
                                                      REFRIGERATOR
                       VENT-
                           INSTRUMENT CALIBRATION BY-PASS
   COAL
  HOPPER
AND FEEDER
N2  CO  $02
  02  NO  AIR
              GAS ROTANIETERS
                                                    CONDENSATE
                     Figure 2. Esso fluidized-bed combustion unit.
                                              IRS02
                                              ANALYZER
                                                                        IR CO
                                                                       ANALYZER
                                                                              WTM
                                                                     NOX ANALYZER
                                                                       1
                                                                              -WTM
POLAROGRAPHIC
02 ANALYZER
                                                 T
                                                  I
                                                 J_
                                            INTERMITTENT
                                            GAS SANIPLER
                                                                               1-2-9

-------
                  VENT
                                                  ^ COOLER
                                                  Q-TT-!
                                                                            VENT
                                                                              DRY TEST
                                                                               METER
                                                                        ANALYZERS
                                                                  KNOCKOUT
                  PREHEATER
                        Figure 3. Fluidized-bed regeneration unit.
     800

     700


     600


     500


     400
I   300
     200

     100
                          U=6 ft/sec
                         Ho=6 in.
                         02=4%
                                                               -\
       1300
1400
                                                                     1700
1-2-10
                 1500             1600
                    BED TEMPERATURE, F
Figure 4. NO emissions as  a function of bed temperature.
1800

-------
900
                   U=6 ft/sec
                   T-1600F
        BED MATERIAL'CaSOi (-1000)
             (dp)COAb300jul:     "
 700
 600
 500
 400
                                                                NORMALIZED
                                                                TO 3% 02
                                                                      ACTUAL
1000
 800
                    2
                                   8
                       4                 6
                         02 IN FLUE GAS, %

Figure 5.  Effect of 02 in flue gas on NO emissions (CaSO4 bed).
                     CaO
                 10
                                                        ALUNDUM
                       CaS04
 600
 400
 200
                   T1600 F
                   U = 6 ft/sec
         02 (FLUEGAS)=4%
                  H=6 in.
                    0.5
1.0
                                       1.5
2.0
                                     RUN TIME, hr
                  Figure 6.  NO emissions using different bed materials.
2.5
                                                                                    1-2-11

-------
      60]
      50

      40
      35
      30
   ,20

oT   15

 """o
   z
 ^   10
 *    9
       8
       7

       6

       5

       4
            INLET GAS COMPOSITION
                N0 = 485ppm
               S02 =700ppm
                N2 = BAL.


            GAS RESIDENCE TIME=0.04 sec
                 BED MATERIAL-16.6% SULFATED LIME
                                                                    * RATE MEASURED 15
                                                                       MINUTES AFTER GAS
                                                                       FLOW STARTED TO
                                                                       REACTOR
         "I   I   I   I   I   I   I   I   I   I   I   I   I   I   I   I   I   I   I   I   I   I   I   I   I
         0.86       0.88         0.90        0.92          0.94         0.96         0.98

                                              1000/t, K

               Figure 7.  Temperature dependence of NO-S02-CaO reaction system.
                                                                                          1.0
     800


     700


     600


 E  500
 3-
 V)
 B  400
 CO
 CO
 
 o  300


     200


     100


       0
                 I     1     I     I
                                                                 BED MATERIAL = CaS04
                                                         TOTAL AIR =110% of STOICHIOWIETRIC
                                                                T=1600F
           TOTAL AIR =130% of STOICHIOMETRIC
                   T=1750 F
                                0.5
                                                        1.0
                                                                                   1.5
1-2-12
                                          SECOND STAGE AIR
                                          FIRST STAGE AIR
                                   Figure 8.  Staged FBC results.

-------
                                      3. COMBUSTION OF COALS
                                                IN FLUIDIZED BEDS
                                                      OF LIMESTONE
                      R. L. RICE AND N. H. COAXES

                  Morgantown Energy Research Center
                             Bureau of Mines
                      U.S. Department of the Interior
INTRODUCTION

  The experimental work reported here was
performed by the Bureau of Mines, Morgan-
town Energy Research Center, under contract
to the Control Systems Division, Office of
Research  and Monitoring,  Environmental
Protection  Agency.  The phase  of  work
assigned to the Bureau was to test various
coals as fuel in fluidized beds of limestone, to
compare sulfur retention, and to measure heat
transfer with tubes immersed in the bed. The
program involved testing five types of bitumi-
nous coal from high volatile A to low volatile,
which varied in ash content from 8 to 24 per-
cent and in sulfur content from 2 to 4 percent.
                                    1-3-1

-------
 EQUIPMENT AND PROCEDURE

    Figure 1 shows the 8-ft  high  combustor
 used in the tests. The bed was supported by a
 cone-shaped plate, perforated by 1/8-in. holes
 fitted with welded stainless steel 90 elbows to
 inject the fluidizing air axially and parallel to
 the cone surface. Water passing  through a
 heat exchanger made of 3/4-in. pipe extracted
 heat from  the  combustion bed. Tests were
 performed at superficial fluidizing velocities of
 3 ft/sec and 6  ft/sec. In the 6-ft/sec tests, an
 additional  heat  exchanger was installed to
 control the temperature of the gases leaving
 the combustor.

    Figure 2 is  a flow diagram of the system.
 Coal was metered by a screw conveyor, and
 then fed pneumatically near the base of the
 bed. Limestone was fed through the side of the
 combustor  just  above the  bed by  a screw
 conveyor. The  bed level was maintained by
 periodically removing material  from  the
 bottom   with   a   3-in.   screw   conveyor.
 Combustion products were  passed  through
 two centrifugal separators for removal of most
 of the entrained solids, and then to a bag filter
 for final cleaning. Solids from the first cyclone
 could be reinjected into the combustion zone.
 Combustion   gases   were   monitored
 continuously for O2, CO2, CO, and SOa  by
 .infrared  analyzers  except  for  O2  which
 utilized  a paramagnetic system. After each
 test the residue was removed from the bottom
 of the combustor by the screw conveyor.

   Startup  was accomplished in about  two
 hours by burning natural gas in the combustor
 and then  injecting coal mixed with limestone
 into the combustion chamber. A 2-ft bed was
 established  at about  1200F, after  which the
 gas was shut off, and coal and limestone were
 fed at a rate that is compatible with the super-
 ficial air velocity and the  designated run
 conditions.

 Combustion Tests

   Five types of coal were burned in beds of
limestone to determine its  effectiveness for
retaining sulfur. The  limestone was the type

1-3-2
 designated BCR 1359  (97  percent CaCC-3,
 Northern Virginia); various sizes of limestone
 were used in tests at 3 ft/sec but in tests at 6
 ft/sec the limestone was sized to 1/4- by 3/16-
 in. The coals were crushed by a hammer mill
 to the range of sizes shown in Table 1. Typical

 Table 1. TYPICAL SIZE  RANGE  OF COALS
 BURNED  IN   FLUID-BED  COMBUSTION
 TESTS
Screen size,
mesh(USS)
-1/4-inch + 20
-20 + 40
-40 + 100
100 + 150
-150 + 200
-200
Weight percent
41-53
16-24
15-20
2-6
2-5
4-10
 analyses of the various types of coal are given
 in Table 2. These analyses varied somewhat
 throughout the test series because batches of
 the same coals were purchased  at different
 times.

   Combustion tests, generally of over 70-hr
 duration, were made at fluidizing velocities of
 3 and 6 ft/sec. At 3  ft/sec, several tests were
 made with each coal; at 6 ft/sec, only one test
 was made with each  coal.  Results of the tests
 are given in Table 3.  Figure 3 shows the effect
 of Ca/S mole  ratio on sulfur  retention in the
 bed. It should be noted (from Table 3) that for
 most  of the tests at 3 ft/sec, material from the
 primary cyclone was recycled to the bed. At 6
 ft/sec, recycle  was possible only with one coal,
 hvbb, due to cooling  of the bed by reinjection
 of the large volume  of solids.

   The data of Figure 3 generally show that
 for the coals burned  at 3 ft/sec there appears
 to be a trend' in  which S  removal increases
 rapidly to 90 percent as Ca/S is increased to
 approximately 2.  Two of the tests.however,
 appear to deviate from this general pattern:
hvab (A-5-L), Ca/S  = 1.8, S removal = 73
percent; Ivb (G-5-L), Ca/S = 2.0, S removal =
68 percent. The test with hvab (A-5-L) was one
of the first tests made and was more cyclic in

-------
                           Table2. TYPICAL ANALYSES OF COALS
                        BURNED IN FLUID-BED COMBUSTION TESTS


Proximate analysis, wt %
Moisture
Volatile matter
Fixed carbon
Ash
Ultimate analysis, wt %
Carbon
Hydrogen
Nitrogen
Sulfur
Oxygen
Heating value, Btu/lb
hvcb
(III. #6)
2.4
33.8
55.6
8.2

72.4
5.1
1.2
2.6
8.I

13,045
hvbb
(Ind. #5)
6.0
35.5
50.6
7.9

69.6
5.2
1.2
2.9
7.2

12,530
hvab
(Ohio)
2.3
39.8
49.7
8.2

72.3
5.5
1.5
3.9
6.3

12,930
hvab
(W. Va.)
1.1
33.1
58.0
7.8

76!5
4.8
1.4
3.0
5.4

13,820

mvb
1.1
19.8
62.4
16.7

71.2
4.4
1.0
3.3
2.3

13,050

Ivb
1.1
17.3
69.9
11.7

76.0
4.0
1.2
2.3
3.7

13,120
nature than the later tests. In the test with Ivb
(G-5-L), the SOa meter functioned only a part
of the time, so the average SO 2 concentration
is suspect. Thus, there is reason to believe that
the average results reported for tests A-5-L
and G-5-L are not representative.

   The results from tests at 6 ft/sec show  a
slightly different pattern, but duplicate tests
would have to be made to confirm this. In only
one of these six tests, that with hvbb, material
from the primary cyclone was recycled to the
bed. Four of the remaining five tests without
recycle indicated the S retention increases as
Ca/S is increased, but not as rapidly as in the
3 ft/sec tests, and that Ca/S of 3  or more is
required to retain approximately 90 percent of
the sulfur in the bed. The test at 6 ft/sec with
Ivb does not fit the pattern, but no explanation
can be offered. In the one test at 6 ft/sec when
recycle was used, S removal was 89 percent at
Ca/S  = 2.1,  closely approximating results
from  tests  at  3  ft/sec  when recycle  was
employed. Therefore, based on the few tests at
6 ft/sec, it is difficult to determine whether the
reduced S removal was caused by the increase
in fluidizing velocity or the absence of recycle.

 Increasing the  fluidizing  velocity  has the
   following effects:
  1, Gas residence time in the bed is reduced.

  2. Bed density is decreased which lessens gas-
    solids contact.

  3. Higher rates of limestone  are  required
    resulting in a reduction of solids residence
    time.

  4. Larger gas bubbles are formed which per-
    mit more bypassing of the SO2.

All of the above effects of  increasing the
fluidizing velocity would tend to decrease S
retention. In addition, since the entrained
solids leaving the combustor likely contain
some unreacted limestone,  recycling of this
material would  be  expected  to improve  S
retention. Thus, the lower  S  retention at 6
ft/sec is probably caused by both the higher
velocity and the absence of recycle.
                                                                                     1-3-3

-------

Table 3.  RESULTS OF FLUID-BED COMBUSTION OF COALS IN BEDS OF LIMESTONE, BCR-1359

Duration, hr
Bed temperature, F
Superficial velocity,
ft/ sec
Coal rate, Ib/hr
Air/coal, scf/lb
Limestone rate.
Ib/hr
Sulfur in coal, wt %
Ca/S mole ratio
SO2 in POC, ppm
Sulfur removal, wt %
Recycle in use
Carbon utilization, %
Type of coal/run number
hvcb
E-4
84
1500
3.0
52.9
95.5
15.0

3.7
2.5
163
95.8
Yes
-
E-5
79
1510
3.0
51.0
98.8
12.7

3.7
2.2
416
90.8
Yes
89.8
E-6
72
1505
3.0
49.5
101.5
13.3

3.7
2.4
461
89.5
Yes
92.7
E-7
50
1520
5.9
72.7
140.2
17.4

4.1
1.8
1094
69.8
No .
75.6
hvbb
Enos Mine
C-4
87
1560
3.1
36.9
137.1
4.0

4.0
0.85
993
71.4
Yes
98.8
C-5
84
1510
3.0
33.8
148.0
5.0

4.0
1.2
520
83.8
Yes
97.1
Blackfoot Mine
C-6
73
1510
3.0
34.0
148.2
12.7

3.1
3.8
305
87.8
Yes
97,3
C-7
70
1520
3.0
31.7
161.6
13.4

2.9
4.4
189
91.7
Yes
92.5
C-8
84
1545
6.2
67.4
154.4
13.2

2.9
2.1
273
88.7
Yes
96.3
Ivb
G-5
84.5
1535
2.7
31.1
143.5
5.6

2.8
2.0
800
67.7
Yes
85.3
G-6
60
1470
3.0
36.3
138.8
5.7

2.3
1.8
192
92.5
Partial
88.6
G-8
46
1530
6.0
87.6
112.4
11.1

2.6
1.5
121
96.1
No
65.0

-------
  Table 3 (continued).   RESULTS OF  FLUID-BED COMBUSTION OF COALS IN BEDS OF LIMESTONE, BCR-1359


Duration, hr
Bed temperature, F
Superficial velocity,
ft/ sec
Coal rate, Ib/hr
Air/coal, scf/lb
Limestone rate.
Ib/hr
Sulfur in coal, wt %
Ca/S mole ratio
SO2 in POC, ppm
Sulfur removal, wt %
Recycle in use
Carbon utilization, %
Type of coal /run number
hvab (W. Va.)

Humphrey
Mine
A-5
84
1525
3.1
34.4
144.8
4.4

2.2
1.8
496
73.1
Yes
98.0
A-6
71
1525
3.0
31.8
157.4
3.9

2.2
1.7
63
96.3
Yes
95.1
Love-
ridge
Mine
A-7
84
1525
3.0
31.4
160.3
9.7

3.0
3.4
134
93.6
Yes
96.6
Ire-
land
Mine
1-1
84
1535
6.3
66.6
157.5
13.8

4.2
1.5
848
73.9
No
91.1

hvab (Ohio)
B-4
75
1580
3.0
34.3
144.1
4.3

3.9
1.0
1003
69.4
Yes
-
B-5
84
1520
3.0
34.8
149.0
5.9

3.9
1.4
194
94.1
Yes
93.2
B-6
75
1560
6.0
58.9
170.7
15.2

4.1
2.0
594
80.0
No
85.3

mvb
F-4
60.5
1505
2.8
30.1
155.2
4.6

5.2
0.9
2310
42.7
No
-
F-5a
84
1490
3.0
41.8
118.0
8.0

3.3
1.6
445
88.5
No
82.1
F-6
36
1435
2.6
52.2
86.3
10.0

5.2
1.2
2036
72.4
No
81 .4
F-7a
84
1495
3.0
36.1
140.0
9.3

3.3
2.2
196
93.9
Partial
89.9
F-8a
80
1575
6.0
65.3
154.7
15.2

2.5
2.9
293
86.3
No
74.1
     Coal was air-table cleaned.
en

-------
    Results in Table  3  also  show that when
 material  from  the primary  cyclone  is  not
 recycled to the bed, carbon burnup decreases.
 This is even  true  at the lower velocity of 3
 ft/sec. In a commercial boiler, if recycle was
 not used, the boiler would have to incorporate
 a method for increased carbon utilization such
 as  the "carbon burnup cell"  proposed  by
 Pope, Evans and Robbins.1  Results from tests
 at  3  ft/sec show carbon  utilizations  with
 recycle to range from about 90 to 99 percent.
 To  consistently achieve acceptable burnup,
 i.e., more  than 99 percent, recycle might not
 obviate the need for a burnup cell.

 Heat Transfer

    Heat transfer  in  a fluid-bed  boiler  is
 important  in  establishing  the commercial
 potential of this  combustion technique and
 would also be important in the design of fluid-
 bed  boilers.  In  the  combustor  previously
 described, which contained  a series of water-
 cooled  U-tubes immersed in the  bed, data
 were taken on one U-tube during the combus-
 tion tests.  Values of Ui were calculated from
 the data and  values for the water  coefficient
 (hi) were calculated using the Dittus-Boelter
 relationship.  Values   for  the  bed-to-tube
 coefficient (he) were then calculated by  the
 relationship
                             D0hB
                                       (1)
            =  overall   heat    transfer
               coefficient based on inside
               area of pipe
hj          =  inside  film  coefficient,
               steam
hg         =  outside film coefficient,
               fluid bed
Di,Dav,D0  =  inside, average, outside
               pipe diameters
x           =  pipe wall thickness
k           =  thermal conductivity of
               pipe.
   Results are given in Table 4 for 18 tests at 3
ft/sec and six tests at 6 ft/sec. The results from
tests at 3 ft/sec are generally consistent, except
for two tests (B-5-L, C-7-L). Neglecting those
two tests,  the  average   bed-to-tube  film
coefficient is 67.5 Btu/hr-ft2-F. At 6 ft/sec,
the average bed-to-tube coefficient  was 32.3
Btu/hr-ft2-F.

   Heat  transfer  also  was investigated  in
another 18-in. diameter combustor  that was
operated  to evaluate  the   performance  of
various alloy tubes. This combustor contained
a  steam-cooled  tube bundle  which passed
horizontally through the fluid  bed.  Figure 4
shows the layout of the tube bundle. During
three tests of approximately 500 hours each in
duration, one at 3 ft/s6c and .two at 6 ft/sec,
heat  transfer  to  the   various  tubes  was
measured.

   Heat transfer data from this steam-cooled
tube bundle were examined  via  graphical
interpretation of overall heat transfer coeffi-
cients.  The  overall heat transfer  coefficient
based on the inside area  of the pipe is given by
preceding equation.  Since conditions  in  the
fluid bed were  essentially constant in each
long-duration -test, h B  should  be essentially
constant. Neglecting thermal expansion of the
tubes, Di, Do, D av., and x are constants for the
tubes; k is constant for the various alloys over
the temperature range of the alloys (k is some-
what  higher  for  carbon   steel,  but  the
resistance term for the metal wall  is  so small
that it  is insignificant).  Hence, the resistance
terms for the metal wall and the outside film
can be combined into one constant  R ]
                                                                                       (2)
                                                  Thus,  U. =
                                                                    +  R,
                                                                                       (3)
1-3-6

-------
                         Table 4. HEAT TRANSFER RESULTS FROM
                                WATER-COOLED U-TUBE

Test

A5L
A6L
A7L
B4L
B5L
C4L
C5L
C6L
C7L
E4L
E5L
E6L
F4L
F5L
F6L
F7L
G5L
G6L


B6L
C8L
E7L
F8L
G8L
I1L

Limestone size
Uj
Btu/hr-ft2-F
Fluidizing Velocity = 3 ft/sec
1/4 in. x 10 mesh
3/8 in. x 3/16 in.
3/ 1 6 in. x 30 mesh
3/8 in. x 3/16 in.
8 x 28 mesh
1/4 in. x 10 mesh
3/8 in. x 3/16 in.
8 x 28 mesh
1/4in.x3/16in.
8x28 mesh
8x28 mesh
3/16 in. x 30 mesh
3/8 in. x 3/16 in.
3/8 in x 3/16 in.
3/8 in. x 3/16 in.
8 x 28 mesh
3/8 in. x 3/16 in.
8 x 28 mesh
71.2
68.9
55.4
64.0
36.I
61.3
60.8
62.4
31.2
69.6
69.1
66.0
68.5
79.1
67.7
65.1
70.8
68.8

Fluidizing Velocity = 6 ft/sec
1/4 in. x 3/16 in.
1/4 in. x 3/16 in.
1/4 in. x 3/16 in.
1/4 in. x 3/16 in.
1/4 in. x 3/16 in.
1/4 in. x 3/16 in.
36.9
39.9
37.6
39.9
34.4
34.9
hBed
Btu/hr-ft2-F

75.5
70.1
52.0
60.9
31.9
61.8
60.9
61.3
26.7
67.7
68.3
64.3
70.1
83.5
70.7
66.0
73.4
72.7
Avg. 67.5a

32.0
34.9
32.3
34.7
29.5
30.1
   a Neglecting two abnormally low values.
                                                                      Avg. 32.3
 or
                      -I-  R,
         nG-8
          D.0.2
                                       (4)
                                             where:
                                                                                    (5)
For a gas  or  vapor flowing inside smooth,
circular pipe  in  turbulent  flow,  the film
coefficient  hj  is given  by  a  number  of
relationships of the  type:
n   =  a constant  that  depends on
        the physical properties of the
        fluid
G  =  mass velocity of the fluid
D;  =  inside diameter of the pipe.
                                                                                  1-3-7

-------
    For steam over the temperatures in these
 tests, n can be considered constant, and D{ is
 constant for all the tubes.  Therefore,
and
where:
hj - CGU
1 1
Ti- CG"'
i
C is a constant.
                            +  R,
    Data from each tube were used to calculate
 Q, the rate of heat transfer. Values of Q were
 then substituted into the formula Q = UiAAt
 to obtain  values for Uj.  A  plot of 1/Ui (as
 ordinate) versus 1/G'8 gives a  straight line
 with a slope of 1/C. The vertical intercept of
 this line, b, represents Rj, the sum of the resis-
 tance of the metal wall and the outside (fluid
 bed) film.

    Figures  5  and  6  show the  graphical
 interpretations  of the overall coefficients  of
 heat transfer. The  linear correlation for the
 first test (Figure 5) has a vertical intercept of
 0.0127,  and  a calculated fluid-bed  film
 coefficient (hs) of  63.4 Btu/hr-ft2-F. The
 correlations for second and third  tests (Figure
 6) were combined since both tests were made
 at  6 ft/sec  and with  -1/4  in.   +  3/16 in.
 limestone. In  this  latter  case,  the vertical
 intercept is 0.0185, and a calculated fluid bed
 film coefficient (hB) of 43.3 Btu/hr-ft2-F.

   The  graphical  interpretations appear
reasonable. Results from the test at 3 ft/sec
with -8 + 30 mesh stone gave a bed coefficient
of 63; results from the two tests at 6 ft/sec with
beds of -1/4  in. + 3/16 in.  stone gave a bed
coefficient of 43. The differences between the
two  values  were caused  by differences in
particle  size  and  fluidizing velocity.  The
location  of data points from the second test
suggests  there is a difference in heat transfer
between the top and  bottom rows; data from
the first  and third tests  do not appear to
support this.
    At the same superficial fluidizing velocity,
 results  from  the  steam-cooled  and  water-
 cooled  exchangers' were expected   to  be
 comparable. At 3  ft/sec,  the  agreement was
 quite good,with bed coefficients of 63.4 for the
 steam-cooled tubes and  67.5  for the  water-
 cooled U-tube. At 6 ft/sec, the agreement was
 not as good: 43.3 for the steam-cooled tubes
 and  32.3 for the water-cooled U-tube.
 CONCLUSIONS

    At a fluidizing velocity of 3  ft/sec,  and
 when fines  from  the  primary  cyclone are
 recycled to the bed, S retention by a limestone
 bed increases rapidly to 90 percent as Ca/S is
 increased to approximately 2. At 6 ft/sec and
 without recycle, it appears that Ca/S must be
 at least 3 to retain 90 percent of the S  in the
 bed.
   Carbon burnup was too low for commercial
 boiler  operation   when   recycle  was  not
 employed, regardless of the fluidizing velocity.
 Even when  recycle was  used at  the  lower
 velocity of 3  ft/sec, carbon burnup might not
 be commercially acceptable so that a separate
 burnup cell  would be  required.  .
   At 3 ft/sec fluidizing  velocity, using beds
 ranging in size from 8 x 28 mesh to 3/8-  x
 3/16-in., heat transfer coefficient from the bed
 to a tube immersed in the bed was 60 to 70
 Btu/hr-ft2 -F. At a velocity of 6  ft/see and
 with  beds of 1/4- x 3/16-in. particles, the
 coefficients to steam-cooled and water-cooled
 tubes were 43 and 32,  respectively.
REFERENCE
1. Bishop,  J.W.,  E.B. Robinson, E.  Erlich,
   A.K.  Jain, and P.M.  Chen. Status of the
   Direct Contact Heat Transferring Fluid-
   ized  Bed  Boiler. (Presented  at  Winter
   Annual Meeting,  ASME, New York. Paper
   Number  68-WA/FU-4.  December   1-5,
   1968.)
1-3-8

-------
                              PRODUCTS OF
                              COMBUSTION
  CASTABLE REFRACTORY
CARBON STEEL
SHELL, X in.
 INSULATING
 FIREBRICK
CASTABLE
 REFRACTORY
  SIGHT GLASS
                                                         LIMESTONE
                                                        WATER
                                                        IGNITER PORT
             AIR

             NATURAL GAS
AIR DISTRIBUTOR
                                              COAL-AIR MIXTURE
                    Figure 1. Fluid-bed combustor.
                                                                           1-3-9

-------
               PRODUCTS OF COMBUSTION
TO ANALYZERS FOR  02 , NO, S02, CO, AND COj
                                                                             CYCLONE
                                                                            SEPARATOR
    LIMESTONE
       AIR
NATURAL GAS
RESIDUE
       FLUID-BED
      COWBUSTOR
                   Figure 2.  Flowsheet for fluid-bed combustion system.
1-3-10

-------
1UU

90


** 80
"5
_T
o
u5 70
cc
u_
_J
c/>
60



50

40
ra O A
*  191 D
D AA A
6



A
A ?
 
6
O
A | 6 OA5L
 ^ * A
w 6 13 G5L





O hvab (W.Va)

> hvab (Ohio)
	 A hvbb 	
A hvcb
L"] mvb (washed)

 E
| mvb (unwashed)
0 Ivb 
6 test at 6 ft/sec

others ,3 ft/sec
!
                              1                     2                    3
                                       CALCIUM /SULFUR, mole ratio

                    Figure 3. Effect of Ca/S  mole ratio on sulfur removal.
                - 5/8-in. TUBING, 0.065-in. WALL
1-1/4 in.
H/4i,
TUBE
A
B
C
D
E
F
G
MATERIAL
SEAMLESS CARBON STEEL
WELDED 340 STAINLESS STEEL
SEAMLESS 316 STAINLESS STEEL
SEAMLESS 410 STAINLESS STEEL
SEAMLESS 446 STAINLESS STEEL
SEAMLESS 304 STAINLESS STEEL
WELDED 316 STAINLESS STEEL
                      GAS FLOW
                    Figure 4.  Arrangement of steam-cooled tube bundle.
                                                                                       i-3-ii

-------
 
                                              BOTTOM TUBE ROW

                                             OMIDDLE TUBE ROW

                                              TOP TUBE ROW
                                                                            10
12
     Figure 5. Graphical analysis of overall  heat transfer coefficients: 3 ft/sec fluidizing
     velocity, -8+30 mesh limestone.
1-3-12

-------
                                                     BOTTOM TU BE ROW, 2nd TEST

                                                 O  MIDDLE TUBE ROW, 2nd TEST

                                                     TOP TUBE ROW, 2nd TEST

                                                 A  ALL TUBES, 3rd TEST
                                    104/GO-8

Figure 6. Graphical analysis of overall heat transfer coefficients: 6 ft/sec fluidizing
velocity, - 1/4 in. +3/16 in. limestone.
                                                                              1-3-13

-------
                            4.  THE REDUCTION OF EMISSIONS
           OF SULPHUR OXIDES AND NITROGEN OXIDES
           BY ADDITIONS OF LIMESTONE OR DOLOMITE
                                  DURING THE COMBUSTION OF
                                         COAL IN FLUIDISED BEDS
                                S. J. WRIGHT

                  National Coal Board., London, England
INTRODUCTION
  It has been known for a number of years, at
least since the First International Conference
on Fluid ised-Bed Combustion held in the
autumn of 1968, that additions of limestone or
dolomite to fluidised-bed  combustors could
materially reduce the proportion of the sul-
phur which was emitted in the flue gases as
sulphur dioxide.

  As research work progressed, both in the
U.S.A. and the U.K., it became apparent that
there were significant and unexplained differ-
ences between results  obtained  at  different
establishments under apparently similar con-
ditions. At that time the National Coal Board
(NCB) had  in operation the most comprehen-
sive range of fluidised-bed combustors avail-
able. In May 1970, therefore, the  National
Coal Board and the Environmental Protection
Agency agreed to jointly finance a consider-
able experimental programme designed (1) to
establish the causes of some of the anomalies
in the extant data, (2) to establish, within the
range of the rigs available, the effects of scale,
and (3) to  systematically investigate some of
the many variables effecting sulphur retention
in fluidised combustion beds.
  The programme was scheduled to cover a
period of 12 months and involved the follow-
ing  work  at the  NCB's  Leatherhead  and
Cheltenham laboratories.


 1. Experiments on a number of pilot-scale
   combustors to measure the effect on emis-
   sion  of sulphur,  nitrogen  oxides,  and
   particulates of a selected range of process
   conditions; e.g. coal type; the quantity,
   size and type of additive (limestone/dolo-
   mite) fed to the combustor to retain sul-
   phur;  combustion conditions as regards
   temperature, pressure,   and  fluidising
   velocity; plant scale; and design features
   such as bed depth and  the recycling of
   incompletely reacted  fuel .and additive.

 2. Experiments on selected  pilot-scale com-
   bustors to assess the extent to which the
   addition of limestone or  dolomite to coal
   in  a fluidised bed in a large test  rig
   influences the  corrosion, erosion,  and
   deposit   formation   on    specimens
   representative   of  typical   evaporator,
   superheater, and reheater tube metals.

 3. Laboratory scale experiments to charac-
   terise the coals and additives used.

 4. Development of a mathematical model to
   assist  in correlating  the factors which
   influence the pollution  control charac-
   teristics of a fluidised combustion system.
                                      1-4-1

-------
THE  SCOPE   OF  THE   RESEARCH
PROGRAMME
   The  main objectives  of  the  research
programme were:
 1. To assess the effectiveness of the fluidised-
   bed combustion process, with and without
   the  addition  of  limestone  or  dolomite,
   towards the reduction of SO2 emission; to
   show  those  operating  parameters that
   significantly affect the attainment  of the
   immediate target emission (300 ppm v/v);
   and to indicate how these data may affect
   plant design.

 2. To gather data, over the  same range of
   operating conditions, on the levels of NOX
   emission that occur during  the  combus-
   tion process  when  the  SO 2 is  partially
   absorbed by added limestone or dolomite.

 3. To measure  the particulates elutriated
   from the fluidised-bed combustor in order
   to provide data for the design of a particu-
   lates  removal system which will reduce
   atmospheric  emissions to an  acceptable
   level.

 4. To contribute towards an understanding
   of the  way in which the porous properties
   of  limestone   or dolomite  affect  SO 2
   retention under the conditions prevailing
   in a fluid-bed combustor, and to develop a
   simple method of classifying  limestones
   and dolomites according to their utility for
   SO?  retention  in the  fluidised-bed
   combustion process.

 5. To develop a mathematical model  of the
   retention of SOa  in a fluid-bed combus-
   tion system to allow the performance of
   new plant, with respect to SO 2 emission,
   to be predicted at the design stage from
   the design and other basic data.
 6. To study corrosion of typical steels used in
   boiler construction  when immersed in  a
   fluid-bed  burning coal,  both  with and
   without the addition  of limestone/dolo-
   mite.

   The research  programme  to meet these
objectives was organised into  eight tasks, at
various plants or locations as follows:

 1. To compare the performance of the 36-in.
    rig with that of the 6-in. rigs at C.R.E. and
    Argonne,  and to extend  the range  of
    operating  conditions  for  which experi-
    mental  data are  available  (36-in.
    combustor, CRE).

 2. To obtain, for operation under pressure,
    data on the emission of sulphur and nitro-
    gen oxides and on corrosion/deposition of
    boiler metal and turbine blade specimens
    (48-  x  24-in.  pressurised   combustor,
    BCURA).

 3. To carry out long-term tests to assess the
    effect of limestone addition on corrosion
    of evaporator, superheater, and reheater
    metals immersed in the fluid bed  (27-in.
    combustor, BCURA).

 4. To obtain data on corrosion of evaporator,
    superheater, and  reheater materials for
    lower fluidising velocities  (12-in. combus-
    tor, CRE).

 5. To obtain data on sulphur retention for  a
    range of coals and limestones, in particu-
    lar to allow comparison to be made with
    the 6-in. rig at Argonne (6-in. combustor,
    CRE).

 6. To  complete  the  development   of  a
    mathematical model of sulphur retention,
    to compare its predictions with the results
    of laboratory and  rig experiments, and to
    up-date  it as appropriate  (mathematical
    work, BCURA).

 7. To investigate the distribution of sulphur
    in a range of coals and  in the residue from
    the rigs (laboratory work, CRE/BCURA).

 8. To  investigate the  pore  structure and
    related factors that  affect sulphur  reten-
    tion by lime  (laboratory work, BCURA).

   The features of the combustors used, the
range of operating conditions explored, and
the   aggregate   number   of  test   hours
1-4-2

-------
accomplished are shown in Tables 1 and 2. It
will be noted that experiments were carried
out at combustion pressures up to 5 atmos-
pheres absolute, fluidising velocities up to 11
ft/sec, bed temperatures up to 1680F, using
four different coals, three limestones, and two
dolomites; the test running time totalled 5300
hours.
MATERIALS USED

   The experimental work in the programme
was carried out using four coals,  three lime-
stones, and two dolomites. Typical analyses of
the materials are given in Tables 3 and 4.
DISCUSSION OF THE RESULTS

   There were considerable differences, both
in size and geometry, between the various rigs
used in the programme. For instance:

1.  Rigs were either circular or rectangular in
   cross section and ranged in area from 0.2 to
   8.0ft2 .
2. The geometries of cooling surfaces within
   the bed ranged from deep banks of closely
   spaced 1-in. diameter tubes to relatively
   shallow banks of widely spaced 2.4-in.  di-
   ameter tubes.

3. The area of the cross section served by a
   single coal feed point varied from 0.2 to 4.5
   ft2.

4. Some  combustors had only internal fines
   recycle systems,  some had both internal
   and external recycle systems, and  others
   had only external recycle systems.  Results
   referred to as being without recycle  are
   from  rigs with only  external, and hence
   controllable, recycle systems.

5. In some combustors the walls of the bed
   and freeboard were  uncooled;  in  others
   they were cooled throughout.

   Despite those differences it was found that
geometry as such was not a variable; the whole
body of the results could be discussed in terms
of the process variables.
               Table 1.  MAIN FEATURES OF THE PILOT-SCALE COMBUSTORS
Feature
Designation
Bed cross
section
Bed depth, ft
Operating
pressure, atm abs.
Fluidising
velocity, ft/sec
Coal rate,
Ib/hr
Total
running hours
Location
BCURA, Leatherhead
27 in.
27 in. dia

1.5-2.0
1

6-11

200-300

2150

48 in. x 24 in.
(pressurised)
48 in. x 24 in.

3.5-4.0
up to 5

2

300-500

430

CRE, Cheltenham
36 in.
36 in. x 18 in.

2-7
1

2-8

75-300

1000

12 in.
(corrosion)
12 in. x 12 in.

2
1

3

20 - 25

1100

6 in.
6 in. dia

2-3
1

2-3

4-6

600

                                                                                   1-4-3

-------
                             Table 2. THE VARIABLES EXPLORED
Operating
variables
Coal
Ash content, %
Volatile matter, %
Sulphur, %
Chlorine, %
H20asfed,%
Ash fusion, F
Size
Bed depth, ft
Temperature, F
Fluidising velocity, ft/sec
Excess air, %
Recycle of cylone
fines
Additive
Ca/S mole ratio
Pilot-scale combustor(s)
Non-pressurised
U.S. Pittsburgh
U.S. Illinois
U.K.Welbeck
U.K. Park Hjll
12-18
37-47
1 .3 - 4.4
0.1 -0.6
1-10
1800-2600
-1/8in.and-1/16in.
1.5-7
1420-1680
2-11
-12 to +29
Zero, partial, full
U.S. Limestone 18
U.S. Limestone 1359
U.K. Limestone
U.S. Dolomite 1337
0-6
Pressurised
U.S. Pittsburgh
U.K.Welbeck
13-18
30-41
1.3-3.1
0.1-0.6
1-6
2100-2600
-1/16in.
3.5-4
1470
2
1 1 to 33
Partial
U.S. Limestone 18
U.S. Dolomite 1337
U.K. Dolomite
0-3
   The order in which the operating variables
are commented upon takes into account both
their relative importance and some of  the
interactions; e.g., through their effect on  gas
and  solids residence times.

   Pittsburgh coal  (3  percent sulphur)  was
used with  either Limestone  18  or Dolomite
1337 in the majority of experiments. While it
is believed that most of the comments in  the
following statement of the main findings apply
to other coals and limestones or dolomite, they
refer primarily to these materials unless other-
wise stated.

Ca/S  Mole Ratio:  The  SO2  is  reduced
asymptotically to zero as the feed rate of addi-
tive  to  the fluidised bed  is  increased. The
percentage SO 2 reduction obtained at a given
operating condition is a function of the mole
ratio of added calcium to sulphur in coal; it is
almost independent of the sulphur content of
the coal,  since the reaction is approximately
first order with respect to SO 2 concentration.
Clearly, in order to obtain a specified concen-
tration of SO2 in the off-gas when burning a
coal of high sulphur content, it is necessary to
achieve a higher percentage SO 2 reduction by
using a higher Ca/S mole ratio.

   For a given coal the lowest values of the
Ca/S mole ratio  were  required  at  (1)  low
fluid ising velocities (i.e. 2 to 3 ft/sec), (2) a bed
1-4-4

-------
     Tables. TYPICAL ANALYSES OF COALS USED

Analysis
Proximate analysis
Total moisture, % a.r.
Ash, % a.r.
Volatile matter, % a.r.
Ultimate analysis
Carbon, % d.b.
Hydrogen, % d.b.
Nitrogen, % d.b.
Sulphur, % d.b.
Oxygen + errors, % d.b.
Chlorine, % d.b.
Calorific value (d.a.f .), Btu/lb
Swelling number
Gray King coke type
Ash analysis
CaO, %
MgO, %
NaaO, %
K2O, %
S!O2,%
Size
(as received)
Coal
Illinois

9.8
11.8
46.6

67.8
4.5
1.3
4.4
8.5
0.2
14,300
4-1/2
D

10.1
1.0
1.7
1.8
40.8

-1/4 in.
Pittsburgh

1.6
13.5
41.1

71.7
4.5
1.4
2.8
4.4
0.1
15,100
8
G9

8.0
1.3
0.7
1.6
45.8

-1/4 in.
Park Hill

2.1
16.5
39.2

68.2
4.4
1.3
2.5
5.3
0.1
14,750
1
D

2.2
1.7
0.8
3.6
46.0

-1-1/2 in.
Welbeck

4.2
18.2
38.3

67.5
4.3
1.5
1.3
5.1
0.6
14,400
1
C

1.8
1.4
1.8
3.2
57.5

-1-1/2 in.
Table4. TYPICAL ANALYSES OF LIMESTONES AND DOLOMITES
Component
CaO
MgO
H20 + C02
Si02
Fe203
S03
Total
Dolomite
1337
28.9
22.9
47.4
0.5
0.2
-
99.9
U.K. Dolomite
29.3
21.5
46.3

-
0.1
97.2
Composition, %
Limestone
18
,45.7
1.4
36.6
13.6
0.3
-
97.6
Limestone
1359
55.7
0.3
43.6
0.5
0.1
-
100.2
U.K. Limestone
55.4
0.3
43.5
0.7
0.1
-
99.8
                                                          1-4-5

-------
temperature of around 1500F, and (3) when
most of the fines larger than about 10 Mm were
recycled. See Figure 1.

Bed Temperature:  The optimum bed  tem-
perature was  between 1400 F and  1600F.
The level of SC)2  emission and the change in
emission with change of temperature on each
side of the optimum  appeared to depend on
the type of additive used and to some extent on
the  Ca/S ratio employed.  The  increase in
emission  on  either side of  the  minimum
tended to be greater at low than at high Ca/S
ratios; i.e. under conditions where the fraction
of calcium  sulphated was  higher. The  opti-
mum  temperature  was  found to be 1500-
1550F for limestone additive and  1400F-
1500F for the dolomite.  The data  suggest
that maintaining the same level of sulphur
emission (e.g.,  85 percent sulphur retention)
at,  for example,  100 F above the optimum,
would involve increasing the Ca/S ratio by a
factor of about two.  The effect of changing
bed temperature  was not investigated on the
pressure combustor.

   The rapid increase in sulphur emission at
bed temperatures  above  about  1550F  is
unexpected from  laboratory measurements of
the reaction rate between CaO and SC2. Since
the effect of temperature appears to be rever-
sible (i.e., the SC>2 reduction reverts to a high
value  as  soon as  the  bed  temperature is
reduced) it cannot be accounted for by irrever-
sible factors such as sintering  or slag forma-
tion at the particle  surface.  One tentative
explanation  postulates  that  an  oxygen-
containing  species   (e.g.,   hydroxide  ions
derived from traces of water, which are known
to be difficult  to remove) is involved in the
conversion of CaSOs to CaSCM. It  is possible
that above  the  optimum  temperature  the
hydroxide ions  become more  mobile   and
hence less able to participate in the reaction.

  At low temperatures (i.e., below  1350F)
sulphur retention with dolomite  was higher
than with  limestone, because of  the lower
calcination temperature  of the MgCOs in
dolomite which leads to the development of

1-4-6
 pore structure  below  the  temperature  of
 calcination of the CaCOs. See Figure 2.

 Fluidising  Velocity:   Increase  in  fluidising
 velocity  resulted in  an increase in sulphur
 emission.   An   empirical  correlation  was
 derived and is reported later in the summary.
 Increase in velocity without any compensating
 action results  in reduction in both gas and
 solids residence times. To maintain the same
 sulphur retention (e.g., 85 percent, at 8 ft/sec
 as at 2 ft/sec fluidising velocity) the Ca/S mole
 ratio (at a bed temperature of  1500F and
 without recycle)  would have to be  increased
 from about 2 to about 4 (Figure 3).
         /
 Bed Height:  Increase in bed height usually
 resulted in a reduction in SCh emission. An
 empirical correlation  is reported later in the
 summary. In principle it should be possible to
 counteract  the  adverse  effect  of increasing
 velocity by a  proportionate increase in bed
 height. At atmospheric pressure the attendant
 increase in pressure loss for other than a small
 increase in bed height could be prohibitive,.  In
 addition,  because the  tube  bank  required
 would occupy only a part of the bed height, the
 effectiveness of increasing bed height may  be
 reduced by the formation of large gas bubbles.
 The effect  of bed height in super-charged
 boilers is potentially  of greater significance.
 Here the deep banks  of close packed tubes
 may assist  in  breaking up large gas bubbles
 and hence  may improve the contact between
 gas  and  solids. Further, the   increase  in
 pressure loss due to increasing bed height is
 less important under  pressure.  See Figure  4.

 Fines Recycle:   A high proportion of the addi-
 tive is elutriated  from the bed before being
 fully  utilised.   By efficient  recycle  of fines
 larger than 10 urn, to the bed, SOi reduction
 was increased significantly; e.g., from 73 to  99
 percent at a fluidising velocity of 2 ft/sec and a
 Ca/S mole ratio  of 1.6.

Operating Pressure:  The effect of operating
 pressure  on SO2 reduction was  negligible
when dolomite was used as an additive. This is
to be expected with a reaction which is first
order with  respect to the partial pressure  of

-------
SO2. With  limestone  as an additive,  the
reduction obtained at 5 atm  was appreciably
lower than with dolomite or with limestone at
atmospheric  pressure. This  was  also to  be
expected, since at 1470 F, calcination of lime-
stone to give a porous structure would  not
occur at operating  pressures above 2 atm.
Penetration  of the  particle  by  SO 2  would
therefore be  difficult and only a surface layer
of sulphate would form. It was found.however,
that the performance of the limestone was
better than  this reasoning would imply; it
suggests that the exposure of fresh surface by
attrition plays a significant role. Nevertheless,
from the point of view of both the Ca/S mole
ratio and the  total  quantity  of additive
required to attain a target level  of sulphur
retention, dolomite was superior to limestone.
To   retain  85  percent  of the sulphur,  for
example, the estimated  Ca/S  mole ratios for
dolomite  and limestone were 1.1 and  3.25
respectively;   the  estimated  quantities   of
additive were 7.6 Ib and 12 Ib per Ib of sulphur
removed, respectively.

Particle Size: For coarsely crushed limestone
the percentage SOi reduction increased when
the particle size of limestone was reduced; this
effect was probably due  to  the  consequent
increase in available reaction surface. On the
other hand, with dolomite there was no effect
of  particle size, suggesting  that  access  to
internal surface is not  a  limiting factor  for
dolomite.
   With additive ground to -125 pm or -150
urn, it was found that the fluidising velocity
had a  profound  effect  on  SO 2 reduction.
Whereas  at low velocity  (3  ft/sec)  the fine
additive improved SOa reduction; the reverse
was true at high velocity (8 ft/sec). The data of
Pope, Evans and Robbins suggest that, with a
Ca/S mole ratio of 2.6 in beds 10 in. deep
fluidised at 12 ft/sec, limestone  1359 gave 80
percent SO2 reduction when ground to 44 nm,
and 60 percent reduction when ground to -74
Mm. Evidently  the  SOj  reduction  is  very
sensitive to the  size of finely ground particles,
so the Pope, Evans and Robbins data are at
least qualitatively consistent  with  those for
-150 vm limestone from  the present study.
Limestone ground to -150 urn may have too
short a  residence time  at high velocity to
achieve a high degree of sulphation; superfine
material will become highly sulphated,  since
its residence time will not be markedly less
than  that of -150  Mm  material.  However,
superfine material may cause  serious gas-
cleaning problems.
Type  of Additive: The type and  source of
additive affects the reduction in SO 2 that can
be achieved.  At atmospheric pressure  Lime-
stone  18 was the most effective additive on
both molar and weight bases; the least effec-
tive on a mole basis was Limestone 1359,  and
on a weight basis Dolomite  1337. To achieve
the same level of retention  with the poorer
Limestone  1359 as with  the. Limestone  18
would require an increase of up to 100 percent
in  the  Ca/S  mole  ratio.  As  mentioned
previously, for operation under pressure both
the dolomites were superior to Limestone 18
on weight and molar bases.

   An  important finding  from the point of
view of simplifying prediction of suitability of
stones  was that measurements made at room
temperature  and  at combustor temperatures
showed the same accessibility of the structure
to gases of similar molecular  size to SO 2.
Temperature cycling (as may occur in some
plant  designs when elutriated  particles  are
recirculated) does not affect the pore structure
significantly  from the point of view of SOi
uptake.  An   empirical reactivity  test  was
considered to be the most economic method
for classifying stones.  For  limestones  the
results of laboratory experiments give  pessi-
mistic  predictions  of  plant  performance.
These  results are thought to be because the
tests do not take into account the beneficial
effect  of attrition in the combustor  which
results in removal of the sulphated surface
layer. The effect of attrition was particularly
important for limestones  in the pressurised
combustor and for Limestone 1359 at atmos-
pheric pressure. For dolomite,  access  to the
                                                                                    1-4-7

-------
 internal surface of the  particles  does  not
 appear to  be a limiting factor.

    Whereas thermal losses are incurred from
 the sensible heat requirement and the heat of
 calcination of an additive, the heat of sulpha-
 tion represents a thermal gain. Up to a Ca/S
 mole  ratio  of  about  2  using  limestone
 (sufficient  to retain 85 percent of the sulphur
 of a  3 percent  sulphur coal under  good
 operating conditions), it is estimated that the
 heat of sulphation  will  counterbalance  the
 sensible heat requirements and heat of calcin-
 ation.  With   dolomite,  however,   the  net
 thermal loss would be about 1-1/2 percent of
 the coal heat input. Under pressure calcina-
 tion of CaCOs  is inhibited,  and the thermal
 loss incurred  by using dolomite would be
 negligible  for Ca/S mole ratios up to 2.

 Type  of Coal:  The  most  important  coal
 property in this context is the sulphur content,
 which  determines not  only the  quantity of
 additive required for a given Ca/S mole ratio,
 but also the percentage SOa  reduction (and
 hence  Ca/S mole ratio needed) to  meet  set
 limits  of  SOi  emission.  Since the  SOi
 absorption reaction is first order with respect
 to  SO 2 concentration,  it could be expected
 that the same relationship between percentage
 reduction and Ca/S mole ratio would hold for
 all coals irrespective of the sulphur  content.
 However,  the  experimental  results showed
 that, for a given Ca/S mole ratio, similar SCh
 reductions  were obtained for three of  the
 coals, but the reductions were up to 15 percent
 higher with Welbeck coal. Differences in the
 rate of  sulphur release  have  been  found
 between  coals and might partly account  for
 differences in  performance. A more likely
 explanation of the higher SO 2 reduction with
 Welbeck coal is its low sulphur content, which
had the consequence that additive was fed at a
lower rate and hence had a longer residence
time. This  could  have resulted in the higher
degree of sulphation, particularly if particle
attrition was  an important effect.

 Plant Design:  As mentioned  earlier, it  was
 concluded that despite the difference of scale

1-4-8
and design over the range of combustors used
there was no significant difference between the
SO2 reductions obtained in different combus-
tors with  the  same  operating  conditions.
Nevertheless, direct application of the present
results  to  combustors  of  commercial size
requires some  caution. Some  factors which
might  alter the SO2 emission, such as  the
depth  of the tube  bank, have already been
mentioned. Further, observation of a radial
distribution of SO 2 concentration in the free-
board of one of the larger pilot plants suggests
that coal feed  spacing, if greater than that
used in the pilot-plants, may assume signifi-
cance in commercial boilers.
Mathematical Model and Correlation of Data:
The mathematical model has been developed
to give fairly satisfactory prediction of the con-
sequence  of  changing some  operating
conditions.  Additional development is needed:
(1) to take further account of attrition of addi-
tive  and (2)  to  extrapolate the  results  to
combustors that differ  significantly from the
present pilot plants. The model in its present
form has not been useful in correlating the
experimental  data. However a  number  of
empirical correlations have  been derived as
follows.
   There is  an  approximately  exponential
relationship between the SO 2 reduction and
the Ca/S mole ratio of the  form

           R = 100[l-exp(-MC)]        (1)
 where:  R  =  percentage SO 2 reduction
         C   =  Ca/S mole ratio
         M  =  empirical constant depending
               on  the  coal, limestone,  and
               operating conditions.

   The  effect  of fluidising  velocity on SO 2
reduction may be approximately correlated by
                A = X,/V               (2)
 where:  A =  absorption ratio, defined as
               R/UOO-R)
         V  =  fluidising velocity
         X] =  empirical constant depending
               on the Ca/S mole ratio and
               other  operating  conditions.

-------
   The effect of bed height on SO2 reduction
may be approximately correlated by
                 = X2H
                         (3)
 where: H
        X,
bed height
empirical constant depending
on the Ca/S mole ratio  and
other  operating  conditions.
Emission of Nitrogen Oxides

   Emission  of  NOX  from   the  pressure
combustor (50-200  ppm)  was  significantly
lower than from the non-pressurised combus-
tors   (300-600  ppm).  All  the  combustors
produced  less NOx pollution,  both with and
without the use of additives, than is common
with conventional plant.  The  reason for the
superior  performance  is   uncertain.   NOx
emission could not be correlated with SO2
emission,   although   on   some  occasions  a
decrease in the SO2 emission (due to feeding
limestone or dolomite) was accompanied by an
increase  in  the  NOx  emission.   It  was
concluded  that more information was needed
on the mechanism of NOx formation before a
contribution could be made towards reducing
emission.

Emission of Alkalis and Chlorine

   As expected, the low combustici tempera-
tures in fluid-bed combustors resulted in low
alkali emissions. The combustion gases from
the pressure  combustor  contained  about  2
ppm of Na; i.e., about one tenth of the lowest
concentration reported  for the  gases from
conventional  plant.  The  concentration of K
was less than  0.5 ppm. Higher emissions were
measured  when limestone was  added to the
pressurised combustor instead of dolomite (5
ppm of Naand 1.5 ppm of K) and from one of
the  non-pressurised   combustors  that  was
being operated at the higher bed temperature
of 1560F (6 ppm of Na and 3  ppm of K). As
expected most of the chlorine of the coal was
released into  the  combustion gases.
Emission of Particulates

   Particulate  matter  elutriated  from
fluidised-bed combustors comprises  5  to  15
percent of the carbon and 80 to 100 percent of
the ash and additive. By using primary and
secondary cyclones having collection  efficien-
cies  of 90 percent  at about 10  Mm it was
possible to  collect  95 - 98  percent of this
material to  give dust emission of 0.2 -  0.6
gr/scf. Within this  range the emission was
approximately proportional to the feed rate of
ash  plus  additive.  Increasing the fluid ising
velocity increased elutriation from the bed,
but because of more efficient cyclone opera-
tion  with higher gas flow rates there was little
effect on emission. Fines recycle in the  36-in.
combustor increased the dust emission  to 1.4
gr/scf.  The pressurised  combustor  had  an
internal recirculation  cyclone in addition  to
primary  and  secondary  cyclones;  dust
emissions in the range 0.05-0.1 gr/scf were
obtained.

   Based  on these results it is unlikely that
there  would be  any  problem  in  meeting
projected  statutory limitations on  particulate
emission.

Corrosion and Deposition

   The addition of limestone or dolomite had
no significant effect on corrosion or deposition
of tubes in the bed or in the gas space  under
the range of operating conditions likely to be
experienced in  a commercial plant.

   The amount of  material settling on the
turbine blade cascade at the outlet of the
pressure combustor was slight and was judged
to be unlikely to affect turbine performance.
There were  no signs  of sintered deposits  or
erosion.

CONCLUSIONS

   The  main conclusions reached  from the
work  are:

(1) With fluidised combustion and the addi-
    tion  of  limestone  (or  dolomite)  the
                                                                                   1-4-9

-------
    emission  of  sulphur oxides from coal
    burning power plant can readily be con-
    trolled to meet the very rigorous restric-
    tions (100 ppm  v/v SO2)  planned for
    certain  densely populated  areas in the
    U.S. For a power plant burning a 3 per-
    cent sulphur coal this would involve feed-
    ing sufficient additive to retain 95 percent
    of the sulphur. Under the best  combina-
    tions of operating  conditions about 1.8
    times the stoichiometric quantity of addi-
    tive would be required;  for a  100-MW
    plant this  would involve  supplying 160
    ton/day of limestone or 280 ton/day of
    dolomite. The less stringent restrictions
    that have been proposed for built-up areas
    (300  ppm),   and  for power  stations
    generally in the U.S. (700  ppm),  would
    require  sulphur retentions of 85 percent
    and 67 percent respectively, for  3 percent
    sulphur coals. These limits can be met
    under a wider range of operating con-
    ditions  and/or   at  less   expense for
    additives.

 (2) Emission  of  oxides  of  nitrogen  from
    fluidised  combustion systems  can  be
    expected to be at least 60 percent less than
    from conventional  combustion systems
    but additional measures would be needed
    over  and  above those used  for  SO2
    reduction to  meet the  very  stringent
    restrictions envisaged for the latter part of
    the century (i.e., 100-200 ppm).

 (3) The particulates emitted  from  fluidised-
    bed systems are unlikely to cause prob-
    lems in meeting current or possible future
    restrictions.

 (4) The use of limestone/dolomite additive to
    restrict  sulphur emission  is  unlikely  to
    affect adversely the  exemplary behaviour
    of the fluid-bed combustion  system from
    the  point of  view  of (a)   fouling and
    corrosion of tubes immersed in the bed
    and (b) deposition or erosion of turbine
    blade materials exposed  to the  com-
    bustion gases.
   In terms of sulphur retention the most
important variable is the Ca/S mole ratio. The
most stringent requirement for SOa emissibns
yet proposed can readily be  met if sufficient
calcium  is present in the  bed. If economic
factors   require    it,    the    usage   of
limestone/dolomite  can be  minimised,  by
reducing  the design fluidising  velocity. This
will  make the boiler bigger and hence more
expensive.

   In terms of boiler operation and control the
most important variable is bed temperature. It
has been  shown that under some  conditions
the  efficiency  of  sulphur  retention  is very
sensitive to bed temperature; for ease of boiler
start-up and flexibility during load following it
is useful to be able to let the bed temperature
vary through the maximum allowable range.
For  American  coals  this  range is probably
from about  1460 to 1800F.
ACKNOWLEDGMENTS

   The author wishes to thank the National
Coal  Board for permission to publish this
paper. Any views expressed are his own and
not necessarily those of the Board.

   This paper  is  no more  than  a brief
summary  of research work which was fully
written  up  in  a  Main  Report  and nine
Appendices involving the  efforts of more than
20 people.

The    author,    therefore,   gratefully
acknowledges the efforts of all the following:

   The NCB Contract Manager  was D.  H.
Broadbent,   assisted   by   S.  J.   Wright.
The research programme was directed by A.D.
Dainton (CRE) and H.R. Hoy (BCURA). The
project co-ordinator was  D.J. Loveridge. The
pilot  plant experimental work at CRE was
administered by J. McLaren. The  following
personnel were  involved  in the tasks into
which the programme was divided.
1-4-10

-------
Task I     The project leaders were D.  C.
           Davidson and D. F. Williams;  A.
           A.  Randell was responsible for
           operation of the plant; D. G. Cox
           and J.  Highley carried out  the
           data  processing and assessment
           of results.

Task  II    The  project  leader was  A.  G.
           Roberts;  D.  M.   Wilkins  was
           responsible for operation  of the
           plant; J. E. Stantan carried out
           the data  processing and  assess-
           ment of results.

Task  III   The  project  leader was  D.  J.
           Loveridge; M.  H.  Barker  was
           responsible for operation  of the
           plant; D.  M. Wilkins carried out
           the  final  data processing  and
           assessment of results.

Task  IV   The  project  leader was  M.  J.
           Cooke;  B.   J.   Bowles  was
           responsible for operation of the
           plant; E. A. Rogers carried out
           the corrosion studies.

 Task V   The  project  leader was  D.  C.
           Davidson; A. W.  Smale  was
           responsible for operation of the
           plant; D. G.  Cox,  J. Highley and
           J. Holder carried  out the data
           processing  and   assessment  of
           results.

Task  VI   The  project  leader was D.  W.
           Gill;  he was  assisted  by F.  V.
           Bethell  and B. B. Morgan.

Task  VII   The  project  leader was  D.  C.
           Davidson; the experimental work
           was  carried   out  by  R.   F.
           Littlejohn.

Task  VIII  The project leader was D. H.  T.
           Spencer; he was assisted by A.  A.
           Herod   and   B.   A.  Napier.
           Additional work to obtain data
           for the  mathematical model was
           carried  out by F. V. Bethell and
           G.  McDonald.
Work on emission of NO x was carried out by
J.  T. Shaw.

   The  monthly  and  quarterly  progress
reports were prepared by D. C. Davidson and
D. J.  Loveridge.  Assessment of the experi-
mental  results for the  Main Report  was
carried out by J. E. Stantan, J. Highley, and A.
G. Roberts.  The  Appendices to the Main
Report were edited by J. Highley and  W. K.
Joy.

   The OAP representative in the U. K. was E.
L.  Carls. His valuable contribution both in the
experimental work and in  the preparation of
the progress reports  and the final report is
acknowledged.
GLOSSARY OF TERMS

Absorption ratio; SOj  reduction  divided by
   100 minus SO2 reduction.
Ca/S mole ratio: moles of calcium in additive
   divided by moles of sulphur in coal.
Excess air: air input minus stoichiometric air
   tor coal input divided by stoichiometric air
   for coal input, times 100 percent.
Fluidising velocity: volume flow rate of gas at
   combustion  temperature  and  pressure
   divided by  cross  section of  combustor
   (neglecting tubes).
SO 2 reduction: SO2 emission without additive
   minus SO 2 emission with additive divided
   by SO2 emission without additive,  times
   100 percent.

Sulphur retention: sulphur in coal minus sul-
   phur in gas divided by sulphur in coal,
   times  100 percent.

Unburnt carbon loss:  unburnt  solid carbon
   divided by carbon in coal  input, times 100
   percent.

Utilisation of  additive: moles   of sulphur
   retained by additive divided by moles of
   calcium in additive, times 100 percent.
                                                                                  1-4-11

-------
     100
      60
 CM
O
o
o
      40
PITTSBURGH COAL
LIMESTONE 18 (-1680 urn)
BED DEPTH:  2ft
FLUIDISING VELOCITY: 3 ft/sec
BED TEMPERATURE:  1470 F
NO RECYCLE
O 6-in COMBUSTOR
 36-in COMBUSTOR
     20
                            1                    2                    3

                                          Ca/S MOLE RATIO

             Figure 1.  Comparison of S02  reduction in 6-in. and 36-in.  combustors.
1-4-12

-------
    100
v

O
     80
      60
UJ
O.
    PITTSBURGH COAL
    LIMESTONE 18 (-1680 urn)
    BED DEPTH: 2 ft
    FLUIDISING VELOCITY: 4 ft/sec
    NO RECYCLE
    36-in. COMBUSTOR
    Ca/S MOLE RATIO:  2.2
      40
                      PITTSBURGH COAL
                      DOLOIVIITE1337 (-1680 iim)
                      BED DEPTH: 2 ft
                      FLUIDISING VELOCITY:  4 ft/sec
                      NO RECYCLE
                      36-in. COMBUSTOR
                      Ca/S MOLE RATIO:  2.7
             1400          1500          1600

                      BED TEMPERATURE, F
                                    1400           1500         1600

                                               BED TEMPERATURE, F
     100
      80
                                                    1700
 I    60


 LU

 s
      40
PITTSBURGH COAL
LIMESTONE 18 (-3175 Jim)
BED DEPTH: 2 ft
FLUIDISING VELOCITY: 8 ft/sec
WITH RECYCLE
27-in. COMBUSTOR
Ca/S MOLE RATIO: 2.8
                                                          PITTSBURGH COAL
                                                          DOLOMITE 1337 (-1587 pi)
                                                          BED DEPTH: 2 ft
                                                          FLUIDISING  VELOCITY: 8 ft/sec
                                                          WITH RECYCLE
                                                          27-in. COMBUSTOR
             1400
         1500
1600
                                                      1400
                                                  1500
                                         1600
                   BED TEMPERATURE, F                           BED TEMPERATURE, F

                          Figure 2.  Effect of bed temperature on S02 reduction.
1700
                                                                                          1-4-13

-------
                                     VELOCITY, ft/sec

                                  5                 3
                           D/
                           /
                     r1
                                              A-
                                                    85
                                                        0*
                                                        a*
                                                                                      "I
                                                                                         E
                                                                                      60  
                                                                                      50
                    0.1
0.2          0.3           0.4


  1/VELOCITY, ft'Vsec
0.5
SYMBOL

o
A
V

D
COAL
PITTSBURGH
11
WELBECK
11
It
PITTSBURGH
ADDITIVE
18
11
U.K. L'STONE
11
11
18
SIZE,
jum
3175
1680
11
11
3175
11
BED DEPTH,
ft
2
2
2
2
2
2
BED TEMP.,
F
1560
1470
?!
11
1560
TT
Ca/S
MOLE RATIO
1.7
2.2
0.8
1.8
2.8
2.7
                            36-, 27-, AND 6- in. COMBUSTORS

           Figure 3.  Empirical relation between SO2 reduction and fluidising velocity.
1-4-14

-------
oe
I
1
SYMBOL

0
A
V
COAL
WELBECK
PITTSBURGH
ILLINOIS
1)
ADDITIVE
U.K. L 'STONE
1337
1359
ii
SIZE,
jim
3175
11
1680
Tf
VELOCITY,
ft/sec
8
TT
3
1)
BED TEMP.,
F
1560

1470
11
Ca/S
MOLE RATIO
2.8
5.0
1.1
2.2
                              36- AND 6-in. COMBUSTORS

         Figure 4. Empirical relation between 802 reduction and bed depth.
                                                                                     1-4-15

-------
                     5.  SELECTIVE EXTRACTION OF CLINKER
                                                 AT THE BOTTOM  OF A
            DEEP  SELF-AGGLOMERATING  FLUIDIZED  BED

                                  A. A.  GODEL

                            Societe Anonyme Activit
   In  our  former  Conferences,  very  few
communications have dealt with self-agglom-
erating fluidized beds. To my knowledge, the
only case of the subject which has been studied
   at  least  in  the  field  of  industrial
achievements  was the use of agglomerating
fluidized beds for coal combustion  via the
"Ignifluid Process" which has been the sub-
ject of several reports given here: in 1968, by
myself on behalf of the Activit Company, and
in 1970, by Mr. Svoboda, Manager of the
Babcock-Atlantique  Company,  and  by  Mr.
Demmy,  Vice-President  of  the  U.G.I.
Corporation.
   Professor  Squires has  honored  me  by
crediting me with having brought to light the
process in which  self-agglomeration  in  a
heavily turbulent fluidized bed results  from
preferential bonding of  slag particles when
they reach their sintering temperature (about
1100C for most coal ash); on  the contrary,
when slag  particleseven adherent ones
encounter coal particles, there is little chance
of their  agglomeration.  This  phenomenon
stems from laws governing the probability of
encounter of particles.

   The result of our study, quite fortunately,
is the possibility of forming slag agglomerates
in a  state of quasi-purity. Slag  agglomerates
fall to the bottom of the bed when they acquire
sufficient weight and must  be  eliminated
promptly in  order to avoid blocking fluidi-
zation. Extracting ash by slag agglomeration
is doubly interesting, for it also permits rein-
jecting the dust carried over into the fluidized
bed since it has  no excess of cinders.

   In most cases,  this leads us to prefer using
agglomerating fluidized beds which require no
temperature limit.

   As I stated here in 1968, the Ignifluid
combustion process achieves  extraction quite
simply by using  an inclined upward-moving
fluidization  grate.  This grate supports  the
fluidized bed with clinker deposited at  the
bottom; it passes through  the surface of the
bed, operating thereby like  a clinker extractor.

   Such grates are relatively narrow and may
be  easily placed  under rectangular boilers.
This position  has  enabled  their  successful
industrial development in equipping boilers of
various steam output for more than 16 years
by the Babcock-Atlantique Company (exclu-
sive license-holder for the process). A 60-MW
Ignifluid power plant comprising two boilers
has been in successful operation for the past
three years. Equipping larger  boilers raises no
problems; several projects have been designed
with such  equipment   and  exported  from
France since the  use of  coal is  constantly
decreasing in Europe.

  Having stated this, I should like to discuss
the essential reason for my communicationa
new process for extracting slag at the base of a
deep fluidized bed.  This new process is quite
different from the former in  both  the means
                                        1-5-1

-------
 used and the aims envisaged, although both
 the Ignifluid Process and the new one operate
 in self-agglomerating fluidized beds.

   Our goal in the new process is principally to
 achieve  certain chemical  reactions in  deep
 fluidized beds, sometimes  under  pressure;
 e.g., the  processing  of  mineral  ore,  the
 processing  of chalky  marl  for  producing
 cement clinker, etc.

   It should be noted that the new process
 may be adapted to the  use  of beds for  fuel
 gasification in conjunction with appropriate
 complementary   treatments,   such   as
 desulphurization and, in certain cases, cataly-
 tic reactions.

   In my invention, I have been guided by the
 obvious principle that it is easier to sort slag in
 a shallow  bed than  in  a deep one.  I  have
 therefore combined the  use  of two fluidized
 beds of different  depths, profiting from the
 communicating vessel principle or the "diving
 bell" principle, as you will see in Figure 1.
 (The figures are produced  simply to  give a
 theoretical explanation of the process.)

   Figure   1  shows   a  horizontal   grate
 supporting  a fluidized bed (A) of a certain
 depth and containing slag at its base.  I shall
 call this bed the "principal fluidized bed." On
 the right, communicating with the first one by
 opening (O), is a shallow fluidized bed (B),  also
 containing slag at its base, which I shall call
 the  "auxiliary fluidized bed."

   As you will notice, the  auxiliary fluidized
 bed is contained in a closed space (C), consti-
tuting what I shall call a "fluidization cell."

   Since gas escaping  from  the  auxiliary
 fluidized bed finds no outlet to the right, it
 escapes on the left through opening (O) which
 communicates with the  principal  fluidized
 bed.

   Under these conditions, the auxiliary fluid-
ized bed is in hydrostatic  equilibrium with the
principal fluidized bed; i.e., the pressure is the
same at corresponding levels.

1-5-2
   Theoretically,  the surface of the auxiliary
fluidized bed should thus be horizontal; but
the fluidization gas leaving this bed and flow-
ing laterally through its surface toward one
end causes  the surface to  assume  a concave
shape, dropping away considerably toward the
opposite end, as  shown.

   Figure 2  shows the same theoretical layout,
but an air intake  is provided at the  end of the
fluidization  cell by opening (d) in order to
increase the flow escaping through (O) toward
the principal fluidized bed (A). The  result is
that the surface of the  auxiliary fluidized bed
(B) drops away much mor,e quickly than shown
in Figure 1,  until  it lays bare the slag lying on
the bottom of the auxiliary fluidized bed.
Perfect  separation  is  thus achieved  between
slag particles and the  fluidized matters.

   To practically  implement the new process,
we use a  cylindrical  or  cylindro-conical
reactor, fitted at its base with a circular fluid-
zation grate which supports a principal fluid-
ized bed and a very shallow auxiliary fluidized
bed, the latter being  contained in  a small
fluidization  cell  immediately underlying the
principal fluidized  bed.

   The  grate and the fluidization cell are in
relative  motion to one  another;  the  fluid-
ization cell is wide open at the base to permit
communication  between  the  two  fluidized
beds, which are in hydrostatic equilibrium as I
mentioned previously.

   The direction of the motion of the grate in
relation to that of the cell is such that the slag
deposited on the grate penetrates into the cell,
where slag particles are decanted, extracted,
and sorted outside the principal fluidized bed
and finally poured into an ash pit in a state of
quasi-purity.

   These principles offer the possibility for a
variety  of embodiments,  of which  I  shall
mention only two.
   The  isometric  projection  in  Figure  3
represents  a cylindrical reactor  (a)  with a
horizontal grate (b) revolving around its ver-
tical axis (x-y). The fluidization cell (c) is fixed

-------
and provided with an outlet (d) toward an out-
side ash pit, not shown.

   The reacting and fluidizing gas is injected
by pipe (e); gas resulting from the reaction is
evacuated by pipe  (f)  after  having passed
through a cyclone (g), which assures collection
and reinjection of fine dust into the fluidized
bed. Pipe (h) supplies the reactor with  gran-
ulated solid material to be treated.

   The   grate   revolves  counter-clockwise,
which causes the slag particles deposited on
the grate to penetrate into the cell via opening
(O).  The size of this opening depends on the
size of the apparatus, and may reach 30  or 40
cm (1 ft) for a large apparatus in order to  avoid
blocking by slag.

   For a better understanding of the manner
in which  the apparatus shown  in Figure 3
operates,  I should like to ask  you to refer to
Figure 4 which shows an isometric projection
of the fluidization cell at a larger scale. This
cell is provided with vertical side-walls (a-b-c-
d) topped by an upper wall forming a sealed
cover.

   Inside the  cell,  the  grate supports the
auxiliary  fluidized bed  which communicates
with the principal fluidized bed  via opening
(O)  situated  at the  base  of wall  (a), as
mentioned previously.

   The shape of the various vertical cell walls
must be adapted to their specific functions, as
follows:

1. Wall  (a) is open at the  base at (O) to
   establish the communication  between the
   two fluidized beds and to provide a passage
   for the slag.

2. Wall (b) is identical to the reactor cylinder
   wall; it  contains an opening (g) giving
   access to an ash-pit placed   outside the
   reactor (not  shown).

3. Wall (c) is preferably laid out  in the shape
   of a logarithmic spiral, with its convex face
   turned toward the outside of the cell to
   force  slag particles  brought  in by the
   revolving  motion  of the  grate  toward
   opening (g). This arrangement profits from
   the very special fact that this  curve has, at
   all points, a constant angle in relation to
   the incident direction of slag brought in by
   the grate (for this purpose, an angle of 30 
   is very favorable). Another  essential func-
   tion of wall (c) is to prevent the principal
   fluidized bed from entering the cell on this
   side and flooding it.

4. Wall (d) serves as a simple connecting ele-
   ment between  walls (a)  and (c).

   This is only a schematic outline of the cell
for, in fact, the cover of the cell will be prefer-
ably in the  shape  of a peaked  roof,  following
the angle of repose or natural slope of products
being processed.

   The second example which I should like to
illustrate is shown  in Figure 5. Here you see a
cylindro-conical  fluidization  reactor (a)
containing in the center an ash pit surrounded
by a fixed annular horizontal  grate (b).

   The fluidization cell (c) revolves around its
vertical axis (x'-y') and  has   an evacuation
toward the inside ash pit situated  at (d), which
is also rotating.

   Reacting gas is injected into the reactor at
(e);  gas  resulting  from  the  reaction   is
evacuated  by pipe (f) after  having  passed
through a group of two cyclones  (g)  mounted
in series;  the latter collect and  reinject flue
dust into  the fluidized  bed.

   The  reactor  is  supplied with granulated
solid material in this example through pipe (h)
so that solid material circulates by counter-
current flow with relation to gases in interest
of heat recovery.

   In this second type of installation the  grate
is fixed; the relative motion between grate and
cell  results from   the  cell's  own clockwise
motion to assure the penetration of slag into
the cell by  opening (O').

   The arrangement of the fluidization cell (c)
is based  on the  same  principle as in  the
                                                                                       1-5-3

-------
preceding example (Figure 4). However, wall
(c) in  Figure 6 is curved  in  the  shape of a
logarithmic  spiral  with  the concave face
turned toward the inside in order to force slag
particles  toward the inside ash pit.

    You will note that wall  (a') is wide open at
the base  in (O')  to  permit inserting  the
fluidized  bed and the slag particles. Vertical
wall (V) is circular. Vertical wall (c') is laid out
in the shape of a  logarithmic  spiral with
concave  part turned  inwards,   as already
mentioned. As with  the preceding  cell, it is
closed by a  sealed cover.

    From  the point of view of the  mechanical
fabrication of the various parts, I feel that it is
unnecessary  to  enter into details concerning
them, except to mention that the cell must be
cooled by circulating water or by vaporizing
water.

    Circular  fluidizing  grates  are  of  the
ordinary  type and may consist of refractory
cast-iron  links supported  by  an appropriate
frame.

    Startup is realized  by  a  gas  or fuel  oil
burner (not shown) placed in the upper part of
the reactor, fired on top of the bed and is then
static. Fluidizing gas (in fact, air)  must  be
injected only when the temperature of the bed
has reached  a sufficiently high level.

    For example, if the fluidized bed consists of
coal, processing capacity may  be estimated at
about 2 to 3  tons/hr of coal gasified/sq meter
of fluidizing grate.

   The resulting processing capacity could  be
for an 8-meter diameter reactor, for example
150 to  200  tons  of coal  gasified/hr  at
atmospheric  pressure. This capacity could  be
considerably  increased to reach an equivalent
power  production of 1000 MW  if pressure
gasification is used.
   Naturally, reactors designed according  to
this new technique may be perfected in various
manners, particularly in  view of obtaining
high  temperatures which  are necessary for
producing cement, by heating the fluid bed
with coal fines or fuel oil and recovering heat
carried off by gases  with  care.  For  this
purpose,   one  might  use   a   series   of
superimposed  cyclones  in which  processed
solid material circulates against the current  of
reaction gases  (Figure 5).

   In addition, residual heat of reaction gases
may be recovered by diverting these gases to a
heat exchanger.

   Finally,  for  gasification   or   for   other
chemical reactions at very high temperatures
(500  to 1000 C),  reacting  gases may  be
injected  by nozzles (Figure  5,  "j") into the
truncated part of the reactor. This injection, if
sufficiently  great, may give rise to a dilute
fluidized  bed,  which  obviously  must  be
stabilized at the top  by  cyclones adapted for
this purpose. This latter use of dilute fluidized
bed   permits,  as it   is   well   known,   a
considerable increase in unit production.

   From all that  has  been stated, I feel that
one may conclude that the process offers the
advantage of an exceptionally simple achieve-
ment with a consequent low  cost, given  its
production capacity.

   This process has undergone successful cold
experiments on different types of  1:20 scale
models. The development is so recent that no
industrial nor semi-industrial testing has been
carried out but a pilot unit is now in  the course
of realisation for the gasification of 250 to 750
kWhr of coal. So  far, no major difficulties
have appeared.
1-5-4

-------
            PRINCIPAL FLUIDIZED-BED
                                                                        (C)
                                                                  FLUIDIZATION CELL
	'"	uiiiiBiiiiiiiiiiiiiiiiiii|iiiiiiin	iiiiiniiii.il	i	nun linn in	niiiuiniiinniiiiiiiiiiiii	i
  (0)
OPENING
                                                                         (B)
                                                                      AUXILIARY
                                                                     FLUIDIZED-BED
                               Figure 1.  Fluidized-beds of different depths.
                                                                                                       1-5-5

-------
                 .       .
         PRINCIPAL FLUIDIZED-BED  .
                           "
                           */
                                       (.0)     \
                                     .OPENING  (B)

                                           AUXILIARY

                                        FLUIDIZED-BED
                     Figure 2.  Fluid ized-beds of different depths with air intake.
                                                                                             (D)
                                                                                         AIR INTAKE
1-5-6

-------
                             (D) VERTICAL
                            '   SIDE-WALL
                                          VERTICAL SIDE-WALL
     (0) OPENING

(A) VERTICAL SIDE-WALL
                                                          (G) OPENING
         (B) VERTICAL SIDE
                                                 (F)
DIRECTION
OF FLOW
                                             (E) HORIZONTAL
                                                   GRATE
           Figure 4.  Detail of isometric projection of the fluidization cell.
                                                                               1-5-7

-------
                                                    (F) EVACUATION OF
                                                       REACTED GAS

                                                    G)  CYCLONE
                                                   (A) HORIZONTAL
                                                      GRATE
                                                   (H) SUPPLY OF
                                                      GRANULATED
                                                      SOLID MATERIAL

                                                   (B) VERTICAL AXIS
                                                  (C) FLUIDIZATION
                                                      CELL

                                                  (D) OUTLET
                                                  (E) INJECTION OF
                                                     REACTING AND
                                                     FLUIDIZING GAS
                                             -(0) OPENING

                                   Figures.  Cylindrical reactor.
1-5-8

-------
               (0)  OPENING

            (A) SIDE WALL
                  (D) SIDEWALL

                  (G)  OPENING
                                             (C) VERTICAL
                                                 WALL
(F) DIRECTION
    OF FLOW  (B) CIRCULAR
                VERTICAL
                  WALL
E)  HORIZONTAL
    GRATE
              Figure 6.  Detail of fluidization cell.
                                                                        1-5-9

-------
                           (G) CYCLONE
            (J)  INJECTION OF
               REACTING GASES
                                                 (F) EVACUTION OF REACTED GAS
                      I)  GRANULATED
                          SOLID
                         MATERIAL
                                                 (G)  CYCLONE
                                                    (A) FLUIDIZATION REACTOR
                 (B) HORIZONTAL GRATE


                -(D) ASH PIT


               "(C) FLUIDIZATION CELL


                    )  REACTOR



                  (I)



(0) OPENING


      Figures.  Cylindrical reactor.
1-5-10

-------
           6. KINETIC STUDIES RELATED TO THE USE OF
                                       LIMESTONE AND DOLOMITE
  AS SULFUR REMOVAL AGENTS IN FUEL PROCESSING

           E. P. O'NEILL, D. L. KEAIRNS, AND W. F. KITTLE
                    Westinghouse Research Laboratories
ABSTRACT

   A pressurized thermogravimetric analysis system adapted to handle corrosive gases has been
used to obtain sulfur removal and regeneration data for limestone and dolomite. The kinetic feasi-
bility of the desulfurization processes proposed for fluidized-bed gasification and combustion has
been demonstrated at ten atmospheres. Calcium carbonate can be regenerated for gasification
processes. Regeneration for the combustion system requires further study.
INTRODUCTION

   Power generation through  fluidized-bed
fossil fuel processing, which uses limestone or
dolomite in the bed to capture sulfur, has the
potential to meet SOa, NOx and  particulate
pollution abatement goals at reduced energy
cost. The application of fluidized beds to the
gasification and combustion of oil and coal at
elevated pressures, with  combined cycle power
generation, and to oil gasification at atmos-
pheric  pressure for retrofit on conventional
plants, is being developed.1  The  efficacy of
sulfur removal is  a major criterion in eval-
uating  these concepts.

   Four aspects  of the chemistry of sulfur
removal will directly affect the  usefulness of
limestones or dolomites as traps for sulfur in
fluidized-bed gasification and combustion of
fossil fuels. The rate of sulfur removal from
the bed gases  during gasification or combus-
tion and the capacity of the stones are primary
concerns for  ' all the  proposed  systems;
regeneration of spent stone in an active form
with concomitant sulfur recovery and disposal
of waste stones which contain sulfur are of
complementary importance.
   The cycle of reactions in Figure 1 encom-
passes the reactions of concern. Table 1 lists
the reactions  and  the  conditions  for  the
process options which are being assessed by
Westinghouse.1 Thermodynamic feasibility is
necessary but  insufficient for success of the
proposed processes; a knowledge of the ki-
netics of the essential reactions under the pro-
cess conditions can only be obtained by exper-
iment.  Despite atmospheric  pressure  data
obtained by Pell et al.4 on hydrogen sulfide
reactions with dolomite,  and various studies
on sulfur dioxide. "sorption" by limestones
and dolomites6'7'8 there is a need for primary
kinetic data with  which  the behavior of the
fluidized-bed desulfurization of fuels can be
predicted and explained.
   The objectives of the current program are
to:
   1. Establish which reactions occur,
                                       1-6,1

-------
     2. Determine the reaction kinetics,                7. Recommend optimal  operating condi-

                                                      tions.

     3. Determine stone utilization achievable,        In this paper we describe how a thermo-

                                                gravimetric analysis (TGA) system, designed
     4. Study the effect of regeneration on stone     to obtain the desired  data  is being  used  to


       reactivity,                                survey the chemistry of the processes, and how

     ..    ,      .       ,   .                    it is focusing attention on critical areas where
     5. Probe reaction mechanisms,                       ,, ,     f..       ,-            r  xl
                                                successful use of the reactions requires further

     6. Assess the influence of side-reactions,     study.
1-6-2

-------
Table 1. REACTIONS AND CONDITIONS OF WESTINGHOUSE PROCESS OPTIONS
Reaction
Sulfur removal
1 	 tiiy., 4- QD i 1/? n *r*^fri _L
' rorn - l'^U2--LabO4 +
CaLU3 ^^2
9 CaO + j_| 5_Qa2 + ^ Q + 
papr\ ^ ^ r*O
OdLrvy-i L* vj -i
J z
Stone regeneration
3 CaSOi + . ? ^-CaQ + ^O i ?V,
co 2 CO 2
4. CaS04 + 4H2 -*CaS + 4H20
4 CO 4C02
(followed by reaction 5, 6, 7 or 8)
5. CaS + H20 + CO2-CaC03 + H2S
6. CaS + H20 + C02-*CaC03 + H2S
7. CaS + 3C02-CaO + SCO + S02
8. CaS + 3/2 02-*CaO + S02
9. CaS04 + C + H20-*CaC03 + C02 + H2S
10. CaS + 2 O2-*CaS04
Operating conditions
1300F2000F
T<1600F
T<1500F
P>5atm
T~200F
atm
T>2000F
atm
T>1800F

atm
Applicable
fuel processing
option
Combustion
Gasification

Combustion
(low S02 concentration
at elevated pressure)
Combustion
(at elevated pressure)
Combustion/gasification
(at elevated pressure)
Combustion/gasification
(not recommended due to low
temperature & water purification)
Combustion/gasification
(not recommended due to low SO
concentration, ~2%)
Combustion/gasification
(low S02 concentration
at pressure)
Combustion
(advanced concept)
All gasification

-------
 EQUIPMENT

   The design conditions for both gasification
 and combustion (i.e., pressures of 10 to 30
 atm, temperatures  up to 1200C,  and  the
 corrosive gas compositions of hydrogen sulfide
 and sulfur dioxide) define an area in which
 little or no kinetic investigations of gas-solid
 reactions have been reported. For this study, a
 duPont  thermogravimetric  balance   was
 mounted inside  a pressure shell, so that it
 could record continuously the weight  changes
 of a solid suspended in a reacting gas stream
 of pre-selected  composition  at  temperatures
 up to 1200C. Figure 2 shows the apparatus;
 Figure  3 shows a closeup view of the balance;
 and Figure 4 is  a diagram of the system. The
 corrosion prevention system is based  on that
 described by Ruth.5 Despite this precaution,
 the  balance  has a limited lifetime owing to
 corrosion;  alternative designs are currently
 under investigation.

   The advantages of a TGA are the ability to
 isolate   chemical  reactions  for study,  the
 relative ease with which the desired conditions
 can  be attained and  controlled, and  the
 accuracy and rapidity  with  which reactions
 can  be  studied, from the point of view of the
 solid.

   The  chief disadvantage for our purposes
 lies in the fact that it is not a fluidized bed;
 translation of TGA results into likely fluid-bed
 behavior is difficult. A second disadvantage is
 the small size of the stone samples used (~ 10
 mg)* which makes it occasionally difficult to
 use chemical analysis to assess the importance
 of minor competing reactions; there  is little
 product for such physical characterizations as
 BET analysis.

   Operation of the pressurized  TGA  was
tested  by studying limestone  calcination.9
When the operating conditions were suitably
modified,  kinetic  results  in  reasonable
 agreement   with   studies   at   atmospheric
 pressure were obtained.

 MATERIALS

   Solids used in the experiments were lime-
 stone  1359, Tymochtee  dolomite,  dolomite
 1337, sieve fractions (420 to 590 Mm) except
 where stated. The  gases, N2, Oi, CO, CO2,
 H2S, and  SO 2 were taken  from commercial
 cylinders. Steam was generated in the mani-
 fold.

 RESULTS

   The strategy adopted was to examine each
 reaction in turn before proceeding  to  study
 cyclic behavior.

   A preliminary survey of the principal reac-
 tions of Table 1 has been carried out; Table 2
 summarizes the results.

 Sulfur Removal

   Combustion:

      CaO + S02 + l/2 O2 -*  CaSO4   (1)

   The  reaction between sulfur dioxide and
 calcined dolomite  1337  in  excess  oxygen
 followed closely the results obtained by Borg-
 wardt7 in a differential  bed  reactor; e.g., rate
 at 10 percent Ca utilization:

 TGA results (850C) 2.40 x  10 "5  [mole SO 3
             gram calcined dolomites-sec-1 ]
 Borgwardt's data    2.10 x 10'?  [mole SO 3
             gram calcined dolomites-sec"1 ].

   The  rate of reaction became  increasingly
 smaller  at  40  percent  Ca  utilization   at
 pressures of 1 atm, but at  10 atm rapid
 reaction  continued  beyond  70  percent
 reaction, as shown in Figure  5. The initial rate
 in the pressurized case is probably limited by
the supply of sulfur dioxide to the solid in our
 particular reaction  system.
*10 mg is kinetically optimal for particle sizes being studied;
balance can accept up to 1 gram samples.
   Gasification:
   CaO 
                                                                                       (2)
CaCO
             H2S
CaS + H20 +(C02)
1-6-4

-------
                         Table  2.  THERMOCRAVIMETRIC ANALYSIS DATA SUMMARY
Reaction
(1)
CaC03 + S02
+ 1/2 02
-> CaSOi, + CO2

CaO + S02
+ 1/2 02
- CaS04

(2)
CaC03 + H2S
-* CaS + H20 + CO2




(4)
CaSOft + 4 CO

-> CaS + 4 CO2

(5)
CaS + H2O * CO2

- CaC03 + H2S

(10)
CaS + 2 02

 CaSO,,



Original
substrate
500-^m dia

Tymochtee
Dolomite
Dolomite
1337
Tymochtee
Dolomite
Dolomite
1337

Tymochtee


Limestone
1359


Tymochtee
Dolomite,
Dolomite
1337

Tymochtee
Dolomite,
Dolomite
1337

Limestone
1359,
Tymochtee
Dolomite
(4000 ^m -
1000 ^m)
Pressure,
atm

1

10

1
10



1


10



1

10


1

10


1

1



Cas . a
composition

0.25% SO2
4%02


0.25% SO2
4% 02



15% H2S
1.5% H2S '" N2

0. 5% H?S in No
C02


25% CO

CO/'CO2 = 2/1


CO2, H2O + 25%




Air

Air



Temp.,
C

800-
850


800-
850



600-
845

760



750,
820
850


590-
700



400-
950
400-
800


Results

40% sulfation at 1 atm.

90% sulfation at 10 atm

Comparable to CoCO3
results



 90% sulfidation
Some suppression at
high H2S concen-
trations
normal calcination


-90% yield




Depends on source of
CaS



7% yield

Up to 90% yield



Conclusions



Makes once-through
systems competitive





Attractive utilization
and rate achievable


under CO.2 pressures
equilibrium

Gives poor substrate
for RX5



For gasification
promising; for
combuslion.
unproven

Appears impractical
as impervious
sulfate shell forms.
Suitable for disposal
al 800 C

Future
work

Higher temperature
composition
Investigate utilira-
lion .n low SO>
concentration




High pressure.
fuel gas effect





Investigate SO>
loss and effect
of initial sulfate
porosity






Seek higher
conversion

Chock loss
of SO 2

T*
VI
          The balance is nitrogen.

-------
   Sulfidation  of dolomite  at  atmospheric
 pressure gave results in broad agreement with
 Pell's work10 (Figure 6). Rapid rates and high
 yields  are  attained. Some  suppression  of
 reaction at  high  hydrogen sulfide concentra-
 tions and lower temperatures  (700C) was
 observed here, confirming Pell's finding.

   In limestone  sulfidation  at  pressure (10
 atm) complete reaction was achievable  when
 the  stone was  calcined;  a  shrinking  shell
 model  empirically  describes  the  kinetics.
 Sectioning of partially reacted stone reveals an
 outer  layer  of calcium sulfide  and  an
 unreacted core of  lime.  No sulfidation was
 apparent  before  the   normal   calcination
 temperature.

 REGENERATION

   The   two-stage  regeneration  of  calcium
 carbonate from  calcium  sulfate has  been
 studied. Reduction  of calcium  sulfate
 according to reaction

       CaSO4 + 4CO  -* CaS + 4CO2    (4)

 was essentially  complete at  temperatures
 between 750 and 850C,  and  10 atm  total
 pressure (Figure  7).

   Regeneration of carbonate from the sulfide
 produced in reaction (3)  above proved to be
 difficult.  AfteF 25 percent regeneration, the
 rate became extremely slow (Figure 8).  By
 contrast, yields greater than 70 percent were
 obtained  when  the  dolomite  was  directly
 sulfided with hydrogen sulfide (Figure 9). The
 latter regenerated carbonate was calcined, and
 its reactivity with sulfur dioxide was tested. It
 proved to react as rapidly as freshly  calcined
 dolomite.

   Unregenerated sulfide is concentrated at
 the  core  of the  particle, as evidenced  by
 sectioning the reacted stone.

 STONE  DISPOSAL

   Spent limestone or dolomite from gasifica-
tion  processes contain calcium sulfide.  This
 compound  liberates  hydrogen   sulfide  on
contact with carbonated water, preventing its
disposal in an untreated form. The problem is
important for once-through and  regenerative
processes.

   Conversion  of  the calcium  sulfide  to
calcium  sulfate  before disposal,  has been
proposed.

       CaS + 2O2  -*CaSO4           (10)

   This reaction has been studied  at  atmos-
pheric  pressure  using  sulfided  limestone,
sulfided dolomites, and air as reactants.

   Sulfided limestone 1359 cannot be oxidized
to sulfate, probably due to formation of the
same impervious  layer observed when the
stone is directly sulfated.11  Surface reaction is
observed but it rapidly decays. The reactivity
can be renewed by lightly crushing the stone
and repeating the reaction, which proceeds to
an additional degree of oxidation  about equal
to the  first  stage. At high  temperatures
(900 C),  about  1  percent reaction  occurs
extremely  rapidly,  followed   by  complete
cessation of reaction. The exothermic  nature
of  the  reaction   may  have  formed  an
impervious  dead-burned lime  layer  on the
solid  surface.

   Sulfided  dolomite  may be oxidized  to
calcium sulfate. At low temperatures (550 C)
an initially rapid rate of reaction falls off after
14 percent of the sulfur  has  been  oxidized
within eight minutes. By contrast, 92 percent
of the sulfide is oxidized within three minutes
at a nominal temperature of 800C. The stone
temperature may be higher (Figure 10).

MODELS

   The development of kinetic models which
encompass the data, with a view to making
predictions  of the course  of  reaction  in
fluidized beds, is  one of the goals  of our
investigation. Empirically, limestone sulfida-
tion  fits  a  contracting  sphere  model,  in
agreement with the physical form of sulfided
stones.  Reduction  of  sulfated  dolomite  by
carbon monoxide is apparently first-order in
sulfate. However,  sufficiently  detailed
1-6-6

-------
variation of the controlling parameters has not
yet been sufficiently studied to permit predic-
tions of reaction behavior over the wide range
of parameters  applicable  to the proposed
processes.

CONCLUSIONS

   The TGA has provided information on the
rates of reaction and the degree of utilization
which can be achieved by the proposed  sulfur
removal,  stone  regeneration,  and  solid
disposal processes.

   Removal of sulfur dioxide and hydrogen
       sulfide:  high stone  utilization  (>70
       percent) has been achieved within 30
       minutes.

   Regeneration  of calcium sulfate by the
       two-step,  low  temperature process:
       calcium sulfate can be reduced to cal-
       cium sulfide ( > 90 percent). Calcium
       carbonate has  not  been successfully
       regenerated from the sulfide  <30
       percent regeneration at practical tem-
       peratures.

   Regeneration  of calcium sulfide: calcium
       carbonate  can be regenerated  from
       calcium sulfide produced from H2S
        70 percent regeneration in 15 minutes.
       minutes.
   Stone disposal: calcium  sulfide has been
       oxided (90 percent) to calcium sulfate
       using dolomite.
These  results are  for  a limited range  of
operating conditions. Further work is required
to assess these reactions over the full range of
operating conditions  projected  for  the
processes.

   Future work is planned to:
    1.  Assess the  low-temperature  CaSO4
       regeneration process.

    2.  Study the one-step, high-temperature
       CaSO4 regeneration process and other
       alternative processes.
    3. Assess the activity of regenerated stone
       as a function of the number of sulfur
       removal/regeneration  cycles.

    4. Further study sulfation of stones for
       disposal.

    5. Translate the kinetic data to fluidized-
       bed  systems  with  the aid  of  experi-
       ments  on a 2-in. hot-model fluid bed
       and pilot plant experiments conducted
       by Westinghouse and others.

    6. Understand  reaction mechanisms.

ACKNOWLEDGEMENTS

   We thank Dr. D.  H. Archer for guidance
and  support.   We   also  acknowledge  the
technical assistance  of  Drs. F. P. Byrne  and
C. R. Wolfe of the  Westinghouse Analytical
Chemistry Department.
REFERENCES

  1. Archer,  D.H., et al.  Evaluation  of  the
    Fluidized   Bed   Combustion  Process.
    Summary Report. Westinghouse Research
    Laboratories, Pittsburgh, Pa. Prepared for
    the  Environmental  Protection Agency,
    Research Triangle Park, N.C.  under Con-
    tract Number CPA 70-9. November 1971.

  2. Keairns, D.L. Fluidized Bed Gasification
    and  Combustion for Power Generation.
    In: Proceedings Frontiers of Power Tech-
    nology, Oklahoma  State University, Still-
    water, Oklahoma,  October  1972.

  3. Archer, D.H. Fuel Processing  Tailored to
    Environmental   Needs.   (Presented   at
    American   Chemical  Society  National
    Meeting. September 1972.)

  4. Pell,  M, R.A. Graff, and A.M. Squires.
    (Presented at meeting of American Insti-
    tute  of Chemical  Engineers.  Chicago.
    December 1970.)

  5. Ruth, L.A., A.M. Squires, and  R.A. Graff.
    (Presented at American Chemical  Society
    Meeting. Los Angeles. March 1971.)
                                                                                   1-6-7

-------
  6. Coutant,  R.W.,  J.S.   McNulty,   R.E.
    Barrett,  G.G. Carson, R.  Fischer, and
    E.H.  Lougher. Summary  Report.  Pre-
    pared For National Air Pollution Control
    Administration, Cincinnati, Ohio, under
    contract Number  PH-86-67-115,  1968.

  7. Borgwardt, R.H.  Environ.  Sci.  Technol.
    4(1): 59, 1970.

  8. Borgwardt,  R.H.   and  R.S.   Harvey.
    Environ. Sci. Technol. 6(4): 350, 1972.
 9. O'Neill, E.P., W.F.  Kittle, C.R. Wolfe,
   and L.M.  Handman. Unpublished results.

10. Pell, M. Ph.D. Thesis, City University of
   New York. 1970.

11. Davidson,  D.C.  and  J.   Highley.   The
   National Coal Board, London, England,
   Final Report to the Environmental  Pro-
   tection Agency, Research  Triangle Park,
   N.C. Appendix ;8,  September 1971.
                 H2S
                                          S0202

                                             SULFUR REMOVAL
                                     	STONE REGENERATION
                Figure 1.  The calcium carbonate/sulfur cycle basic reactions.
1-6-8

-------
OS

GO
Figure 2.  Thermogravirnetric analyzer for high temperature and pressure reaction studies
on limestone and char.

-------
                   Figure 3.  The duPont 950 thermogravimetric balance.
1-6-10

-------
                                   *
COOLING
 COILS
SAMPLER-
       BALANCE
      MECHANISM
                     FURNANCE
                   V1\1\\\V\VV
=!  SAMPLE
                     iiiinm
                     COOLING COILS
 PRESSURE
" GAUGE
                                            BY-PASS
                                      PREHEATERS
   10 ATMOSPHERES
    REACTION
      GAS
    MANIFOLD
I     I    I    I
                                                              NITROGEN
                                                             '  PURGE
                   RECORDER
                    CONSOLE
                 Figure 4.  Diagram of the TG system.
                                                                         1-6-11

-------
                                                                     850C, 5000 pptn SO?
                                                                           402
                                                                     10 rug DOLOMITE
                                                  10       12       14      16       18       20
                      Figure 5. SC>2 reaction with calcined dolomite 1337.
1-6-12

-------
         ATMOSPHERIC PRESSURE
         -30+40 TYMOCHTEE
                DOLOMITE
         15% H2.S IN N2
         845C
                             EXPERIMENTAL
                             PELL'S FIRST
                             ORDER MODEL
              10      20      30      40     50
                  TIME, sec

 Figure 6.  Sulfidation of calcined dolomite.
                                                 ea
                                                 UJ
                                                 UJ
                                                o
                                                LU
                                                Q_
90

80

70

60

50

40

30

20

10

 0
            OTG 95 38% SULFATED ~
            *TG 96 26% SULFATED
        CO/C02 = 2/1
            CO=20%
             P=10atm
             T=820C
             DOLOMITE 1337
       CaS04+4CO-*CaS44C02  

I    I   I   I   I   I   I    I   I
                                                      0  10  20   30  40  50  60  70  80  90
                                                                    TIME, min

                                                 Figure 7. Reduction of sulfated dolomite.
Ld
(
LU
      30

      25

      20
      15

      10

       5

       0
                10      20       30      40      50       60       70

                                           TIME, minutes

                              Figure 8.  Regeneration of dolomite.
                                                                         DOLOMITE 1337
                                                                         (H20)10%
                          =10%, 23%
                       = 650C
                        10atm
                      TG94
                    CaS  MgO + C02 + H20  CaC03  MgO + H2S
                                                                         80
                             90
                               100
1-6-13

-------
LU
    0.8
o  0.7
QQ
ce
o
    0.6
3  0.5
  0-4
o
LU
S  0.3
~  0.2
o

i  0.1
o

^  0.0;
              DOLOMITE 1337

           -30+40
           86% Ca2+SULFIDED
           700  C, 10 ATM'        
           [H20]=[C02]=20%

"CaS- MgOf H20+C02 = CaC03- MgO+H2S"
       04     8    12    16   20    24    28

                      TIME, n\\n
  Figure  9.  Regeneration of carbonate from
  sulfided dolomite.
                                             O  TG 121 DOLOMITE 1337,55% Ca SULFIDED

                                             A  TG21  LIMESTONE 1359, 99% Ca SULFIDED
                                                      ATMOSPHERIC PRESSURE
                                                              AIR
                                                             70QC
                           CaS+202CaSOa
                                              100
                                            TIME, sec
                     Figure 10.  Oxidation of sulfided limestone and dolomite.
1-6-14

-------
SESSION II:
  Non-Coal Fluidized-Bed Combustion Processes

SESSION CHAIRMAN
  Mr. A. Skopp, Esso Research and Engineering, Linden, New Jersey
                              II-O-l

-------
                                                   1.  COMBUSTION m A
                                           CIRCULATING FLUID BED
                                   H. W. SCHMIDT

                  Lurgi Chemie und Huettentechnik GMBH
INTRODUCTION

  To reduce the high heat consumption as
compared with the rotary kilns normally still
used for calcination of alumina, Vereinigte
Aluminium Verke AG (VAW) has developed a
fluid-bed  process with direct  combustion in
cooperation with the Lurgi Companies.

  The main  characteristic of this  process,
which is generally suited to the application of
endothermic processes at low particle diam-
eter, is the circulating fluid bed. Compared
with classical fluid beds with constant  bed
height and a defined surface, a gas/solid mix-
ture of varying concentration in a circulating
fluid beds fills up the complete reactor room.
The solid discharged  at the furnace top to-
gether with the combustion gas is fed back to
the reactor after being collected from the gas
flow in a recycling cyclone. Only by permanent
recycling is the system  maintained in a  con-
stant condition.

   This type of reactor  has the advantage of
passing large gas volumes through relatively
small reactor sections.


   Since the combustion gas and the fluidiza-
tion gas are identical, the problem is to opti-
mally coordinate the fluid dynamic conditions
for the fluid bed, such  as solid recycling and
distribution of concentration, with the pyro-
technical requirements, such  as mixing and
combustion.
                                       II-l-l

-------
CIRCULATING   FLUID-BED  PILOT
PLANT

   To investigate these processes, a pilot plant
(24 ton/day AhO3) for fundamental studies
was erected prior to the construction  of the
first industrial plant (500 ton/day A12O3).
Figure 1 shows a process flow sheet with the
most  important process steps.1

   The essential part of the process consists of
the calcination circuit which  works on the
principle of a circulating fluid bed. Optimum
thermal  efficiency of the whole  process  is
reached by preheating the solid with the com-
bustion gas flow, and the combustion air with
the solid discharge flow.

   The feed hydrate is preheated by the waste
gas flow  and  partially dehydrated in two
Venturi dryer stages. The alumina, which is
preheated to about 400C and has a loss on
ignition of 7-8 percent, is subjected to final
calcination in the  fluid-bed furnace.

   The  energy required for the endothermic
process, water evaporation, dehydration heat,
radiation and discharge losses, is supplied by
direct combustion of heavy fuel oil in the fluid-
bed furnace.

   The combustion air is divided into primary
air and secondary air. The two air streams are
preheated in a multi-stage fluid-bed  cooler by
the discharged alumina. The secondary air
flow, which fluidizes the alumina in the cooler
chambers, is preheated directly. The primary
air flow, which is passed through tube bundles
immersed in the fluid-bed chambers, is pre-
heated indirectly.

   Figure 2 shows a sketch of the material
balance  of a  circulating fluid  bed.  The
alumina m A discharged at the furnace top
together with the combustion gas is composed
of the circulating flow niR and the throughput
flow mD.  The ratio 9!" circulating flow  to
throughput flow m R: m D  is decisive for the
mean retention time of the product. The mean
retention  times of  the  alumina  may be
adjusted between about 10 and 60 minutes

n-i-2
depending on the intensity of circulations of
the circulating  solid.

   This adjustment of the alumina retention
time needs  a defined distribution of  solid
concentration in  the  axial  direction  of the
fluid-bed  furnace,  corresponding  to  the
division of primary and secondary air streams.
Because of the  high gas velocities compared
with conventional fluid beds, the problem is to
burn the fuel completely by the time it reaches
the furnace top. This must be  achieved by
optimum mixture of fuel  and air. The  studies
mentioned in this paper therefore concentrate
on  the combustion  efficiency  at  varying
conditions for the mixing ratios of fuel and air.

   The  phase  condition of the circulating
fluid bed is illustrated in the fluid-bed phase
diagram (Figure 3) as a  function of the Fr^
number and the Re^ number.  The perimeter
is determined by parameters K and M. K and
M  are  nondimensional  parameters  which
result from a combination of the  Fr^ number
and the Rek number which derive from the
fluid-bed phase diagram.

   The range shown in the diagram covers the
adjusted test conditions by means of the K and
M lines. This range is located in the zone of
the aggregative fluidization, a transition phase
between classical fluid bed  characterized by
the particulate  fluidization of the  individual
particle, and of pneumatic transport which
starts with the limiting characteristic 3/4-Fri?
7G/(/K  y&)= 1-  The  cross-hatched  zone
between lines  K  =  3.52  and  K  =  2.1
characterizes at a mean particle diameter dkm
=  45 fim of the  alumina  used, the phase
condition of the lower furnace section whose
concentration is determined by the primary air
flow. Lines K = 626  and K =  0.264 result
from the  minimum  and maximum particle
diameters, dkmin = 8 ion and d|jmax = 120
urn.  Lines M = 0.035 and M  = 1.33 are
determined   by  the   lowest  and  highest
fluidization velocities.

   It can be inferred from this illustration in
the fluid-bed phase diagram that the fine
particles are  preferably discharged from the

-------
lower furnace section and subjected to  a
higher amount of circulation than the larger
particles, since the K lines of the smaller
particle diameters remain above the boundary
line for pneumatic transportation. However, it
can  be  demonstrated  by  tests  that   no
separation of the fine  particles occurs during
circulation.3  This   means   that  under
appropriate flow conditions all particles come
into the  range of pneumatic transport.  By
impact and frictional  forces of the particles
and  by   separation   of   particle  clusters,
conditions of fluidization may be changed.
Thus, the transport condition is delayed at
various times and in various places, resulting
in inconstant distribution of solid  concentra-
tion in the axial direction.

MEASURING  EQUIPMENT  AND TEST
PERFORMANCE

   The following parameters were varied  for
studying  the  combustion process.

1. Primary air velocity, Wp
   This is the velocity referring  to the free
   cross-section in the lower range  of the
   fluid-bed furnace below the secondary  air
   inlet.

2. Impulse of the secondary air flow, Is
   This is the force by  which the secondary air
   jets penetrate  horizontally  into the fluid
   bed produced by the  primary air.

3. Mean furnace  temperature, Tm

4. Air ratio of the combustion

   For the measurements, exhaust openings
were  arranged in six  measuring sections at
various levels of the fluid-bed furnace. It was
possible to take gaseous and solid samples by
two measuring directions  arranged perpen-
dicular to each other, with water-cooled bleed-
off lances. Figure 4   shows the  measuring
arrangement  for the tests.  The alumina
sucked off together with the gas is separated in
a cyclone. After cleaning and drying, the gas is
passed to continually  operated gas analyzers
for determining  the   concentration  of  the
components CO, CO2, and O2.
   The central part of the lance is composed of
a  thermocouple  for  ascertaining   the
temperature prevailing at the corresponding
measuring  point. By control of the suction
pump the exhausted gas rate can be adjusted
in accordance with  the temperature  at  the
measuring point in  the lance section so that
isokinetic suction conditions exist.

   With  this  measuring   equipment,   the
following  measuring  variables  can  be
determined  in a radial direction  of each
measuring section:
1. Gas concentrations of CO, CO2, and O2.
2. Temperature, Tm.
3. Concentration of solid, CM-

INFLUENCE OF SECONDARY IMPULSE
ON  THE EFFICIENCY OF COMBUSTION

   From the radial profiles  measured, mean
values of the measuring section  areas  are
formed by integration over the furnace cross-
section. From the integral mean values of the
various section areas, the axial distributions of
gas concentrations,  solid concentrations, and
temperatures  can  be  determined  over  the
height of the fluid-bed furnace.
   In Figure 5 the  axial temperature distri-
butions are stated for various mean furnace
temperatures (Tm).  For the greater part of the
fluid-bed furnace, the lines show a constancy
of temperature  which  does  not  occur  in
conventional  fluid  beds and can only be
explained by  the intensive solid  circulation.
The  decrease in temperature in the lower part
of the furnace is caused by the solid feed and
the  remaining  dehydration  heat  of  the
calcination process  still to be applied in the
range. Tests at  a higher ratio of circulating
solid to throughput solid mR:mD show that
the temperature drop in the lower furnace
section can hardly  be observed any longer.
   Distributions of solid concentration in axial
direction at varying velocities referred to in the
free cross-section are represented in Figure 6.
In the lower furnace section, the curves show
increasing  solid concentration  CM  (x)  at
decreasing primary air velocities (w). Starting
                                                                                   II-1-3

-------
 with coordinate X/D = 1.4, the solid concen-
 trations of all curves show a significant de-
 crease with  volume  flow  increase by the
 secondary air flow (Vs). Downstream  of co-
 ordinate X/D = 3.9 the axial distributions of
 the solid are almost constant up to the furnace
 top.

    To analyze the efficiency of combustion, it
 is necessary to determine the combustion rate
 of the fuel components.4 Except for the local
 gas analysis values CO, CO2,  and O2, it was
 possible to rely  upon  the unburnt  carbon
 percentage  which was  analyzed  from the
 sucked off alumina. These deposits of unburnt
 carbon on the alumina are mainly found in the
 lower measuring section areas, decreasing in
 accordance with the  course of combustion up
 to the furnace top.3

    The  active y-alumina  coming  from the
 preheating section at a temperature of about
 400 C is also fed to the zone of the fuel which
 is directly injected into the fluid bed. Due to
 the  catalytic  activity of this  y-alumina, the
 cracking process initiating the combustion of
 the  heavy  fuel oil is greatly  enhanced. The
 catalytic effect on the cracking process must
 be exclusively caused by the activity of the y-
 alumina since the remaining alumina, which
 has  a temperature of 1100C,  is  no  longer
 active as a  catalyst.

    On  the  basis   of  the  burnt  carbon
 percentage,  a molar  balance  of the  oxygen
 required for combustion,  and the  measured
 gas analysis values for CO, CO2,  and O2, it is
 possible to establish the necessary equations to
 determine the  development  of  combustion.

    From the radially measured values, such as
 shown for characteristic tests in Figure 7, the
. distribution   of  combustion rates in  axial
 direction of the fluid-bed furnace are obtained
 by radial and axial integration. Figures 8 and
 9  show the combustion curves from which the
 main factors of influence  on the degree of
 combustion  can be seen.
   Figure 8 shows efficiencies of combustion
with constant impulse of the secondary air jets
and varying primary air velocities (wp), which
are proportional to the velocity referring to the
free  cross-section (WG). At decreasing values
of the primary air  velocity (wp),  a steeper
course  of the  combustion  curves  can be
observed. This  result  is attributed to the
retention time of the combustion gases in the
furnace,  which  increases   at  decreasing
velocity.

   Figure 9 shows percentage of combustion
at a constant primary air velocity (wp), but
varying impulses (Is) of the  secondary air jet
penetrating into the fluid bed. Although the
high primary air  velocity (wp)  shortens the
retention time  of the   combustion  gases,
substantially steeper combustion degrees can
be reached at increasing secondary air impulse
(Is) than in the case of curves as  per Figure 8.
This result is due to the more intensive radial
mixing of the  fuel  components  with the
combustion air. The radial  distributions in
Figure 7 in the two different measuring section
areas clearly show this influence. At a lower
secondary air impulse an increase of the CO 2
concentration exists only in the zone near the
furnace wall. At a higher  impulse value a
homogeneous distribution of  concentration
over all the measuring  section areas exists.5

   The  results  show   that   by  suitable
distribution of  primary and secondary  air
streams, the  combustion  reaction can  be
substantially speeded  up. This will permit an
increase  of the  gas throughput and  a  higher
specific load  of the fluid-bed furnaces, along
with a reduction of the height of the fluid-bed
furnace.3

   In conclusion, it should be mentioned that
apart from the combustion process as studied
in a pilot plant, two industrial plants, each
having  a capacity of  500  metric ton/day
AhOs, are successfully operating according to
the described process; another  three plants,
each with a capacity of 650 metric ton per day
Al2O3, are under construction and will be
started up in the course of the next year.
 H-l-4

-------
SUMMARY

   In  a pilot plant for the calcination of
alumina according to the circulating fluid-bed
method, the  combustion process is  studied
experimentally. Since high gas velocities occur
in the circulating fluid bed, the retention times
of the combustion gases are short at direct
combustion in the fluid  bed. By a suitable
division  of  the   primary  and  secondary
combustion air flows, it is possible to increase
the radial mixing and to substantially speed
up the combustion of the fuel components. It
is thus  possible  to  increase  the  specific
throughput capacity by raising the gas rate.

   During  the tests,  deposits   of  unburnt
carbon on the alumina were observed in the
lower  furnace section.  Consequently,  the
endothermic  process can only be performed
according to the principle of circulating fluid
bed since the carbon is not completely burned
until  the fuel reaches the furnace top.

NOMENCLATURE
CM = Solid concentration, kg/Nm3
D = Diameter of reactor, m
dfc = Particle diameter, m
   = Froude number (-)
   = Acceleration of gravity, m/sec2
gy = Burned-out portion of fuel (-)
gu = Unburned portion of fuel (-)
Is = Impulse of secondary air flow, m kg/sec 2
g
 K  =
 M  =
                  Vc
              K-VG)
        ,3
mp
m R
Tm =
       Solid output on furnace top, kg/hr
       Solid throughput, kg/hr
       Circulating solid, kg/hr
       Reynolds number (-)
       Mean temperature,  C
                                             VP  = Primary air flow, NmVhr
                                             YS  = Secondary air flow, Nm3/hr
                                             V F  = Waste gas flow, Nm3 /hr
                                              WG . = Superficial gas velocity, m/sec
                                             wp  = Superficial velocity of primary air,
                                                    m/sec
                                             LM = Efficiency of combustion, gv/gu + gv)
                                             yG =  Specific gravity of gas, kg/m3
                                             y'K =  Specific gravity of solid, kg/m3
                                              v  =  Kinematic viscosity of gas, m2 /sec
BIBLIOGRAPHtY

1. Reh, L. Fluid Bed Processing. Chem. Eng.
   Progr. 67(2).  1971.
2. Ernst, J., L.  Reh, K.H.  Rosenthal, and
   H.W. Schmidt.  Experience with the Cal-
   cination  of Aluminum  Trihydrate in  a
   Circulating  Fluid  Bed. (Presented  at
   American.   Institute    of   Mechanical
   Engineers  Meeting.  New York.  Paper
   number A-71-4. February 1971.)
3. Schmidt, H.W.   Uber den Verbrennung-
   sverlauf in zirkulierender Wirbelschicht.
   Dissertation, Karlsruhe 1971.
4. Gunther,  R.  Ausbrand  von  Strahl-
   flammen.  Archiv f.d. Eisenhiittenwesen.
   39(7):515-519, 1968.
5. Martsevoi, E.P.  Spread of a Gas Jet in a
   Cross-Flowing Stream of Different Density.
   Gas-Institute, Acad.  of Sciences  USSR,
   translated  from  Teoreticheskie  Osnovy
   Khimicheskoi Technologic.  3(4):644-646,
   1969.
6. Reh, L. Das Wirbeln von kornigem Gut im
   schlanken  Diffusor  als  Grenzzustand
   zwischen   Wirbelschicht  und  pneu-
   matischer   Forderung.  Dissertation,
   Karlsruhe 1961.
                                                                                  n-i-s

-------
                              CYCLONES
        FAN

                   /!
ELECTROSTATIC
PRECIPITATOR
                         FLUID-BED
                          FURNACE
                                                        Al (OH)3
                                                         WET
                                                             FEEDING
                                                             SCREW
                                  \^
                                                  SECONDARY
                                                  AIR BLOWER
                                                                          FLUIDIZED
                                                                         BED COOLER
                                              PRIMARY
                                             AIR BLOWER
                         Figure 1.  Flowsheet of fluid-bed calcining plant.
n-i-6

-------
Figure 2.  Material balance of circulating fluidized bed.
                                                           IM-7

-------
IM-8
          IIP
                                   Figure 3.  Fluid-bed diagram.

-------
                                          TO RECYCLING CYCLONE
                                                  ANALYZER  ANALYZER  ANALYZER
                                     FILTER


                                FLOW METER I
                                          ALUMINA - FEEDBACK
                                           SECONDARY AIR
WORKING PLANE II

     FEEDBACK

    ALUMINA
ALUMINA INLET
    FROM
PREHEATING STAGE
             WORKING PLANE I
                                     SUCTION PUMP

                                   DRYER
                                                 	ALUMINA
                          Figure 4.  Measuring arrangement.
                                                                                II-1-9

-------
     1100
     1000
  o
  o
  0=





  fi!   900
      800
      700
     600
-O- T= HOOt


-O- T = 990 C


-- T = 900 'C



-- T = 810 T
                                                                                        O
                                              HEIGHT,-^



         Figure 5.  Axial temperature Tm distribution versus nondimensional height X/D

         of fluid-bed furnace.
IM-10

-------
    200

    180

    160

     140


j   12
5
S   100
60
40
20
                                                    Ow
                                                    A w
0.91 m/sec, WQ  1.83 m/sec
1.19 m/sec, WQ  1.98 m/sec
1.32 m/sec, WG.  2.05 m/sec
1.43 m/sec, WQ  2.36 m/sec
1.64 m/sec, WG  2.53 m/sec
2.03 m/sec, WG  3.05 m/sec
  01234567

                                NONDIMENSIONAL HEIGHT (X/D)

  Figure 6.  Axial solid concentration (C|y|) distribution versus nondimensional
  height (X/p) of fluid bed furnace.
                                                                                       Il-lrll

-------
                                                                    240 cm ABOVE GRATE
                                                                    wp = 0.91 m/sec
                                                                    ls=1.298m-kg/seez
140 cm ABOVE GRATE
WD = 0.91 m/sec   n
ls =1.298 m-kg/sec2

  o C02
   02

  * CO
                                                                       02
                                                                     OCO
                                                                    ls = 0.551 m-kg/sec2
                                                                       C02
                                                                      nco
      14

      12


      10
   CM
  O
  o
  O
                                     CM
                                    o
                                        14


                                        12


                                        10


                                         8
                                    o
                                    o
        490  350   210  70  0  70    210   350  490
8  6


    4


    2

    0
                                          490   350   210    70  0  70   210   350  490
                       RADIUS, mm
                                                         RADIUS, mm
                    Figure 7.  Radial distribution of gas concentration 140 cm
                    and 240  cm above grate at constant primary air velocity
                         and variable secondary  impulse (ls).
II-1-12

-------
CO

o
                                                                   ls=0.551 m-kg/sec2
                                                                   T =1100C
                                                                   Wp=L64m/sec

                                                                   Wp=2.03 m/sec
                                      3          4

                                          HEIGHT (X/D)

   Figure 8.  Efficiency of combustion (a/m) at constant impulse (l/s) and varying primary air
   velocities (w/p).
                                                                                      II-1-13

-------
                                       3         4          5

                                     NONDIMENSIONAL HEIGHT (X/D)

              Figure 9. Efficiency of combustion (am) at constant primary velocity (Wp)
              and variable secondary air impulse (ls).
II-1-14

-------
                              2.  DISPOSAL OF  SOLID WASTES BY
                                      FLUIDIZED-BED  COMBUSTION

                                G. G. COPELAND
                             Copeland Systems, Inc.
ABSTRACT

   Those of us who labor in the field of pollution control have grown to realize that solid waste
disposal is the largest pollution problem which society must face in the coming years. For too long,
we have left the problem to municipal fathers, who for one reason or another have either ignored
the problem or have installed equipment which has long since been made obsolete by modern
technology.

   We hear every day about the virtues of land filling garbage and using municipal sewage sludge
as soil conditioner, "a la night soil" techniques used in the Far East, because incineration is too
expensive, too dirty, makes smoke, and needs high chimneys. This thinking is obsolete in the face
of present  day  fluid-bed  technology which  is  substantially  less costly than  conventional
incinerators, is not dirty, and  cannot be made to smoke if operated properly. To our knowledge,
there is not  a fluid-bed system in operation anywhere in the world which is backed up with a smoke
stack.

   This paper covers the development of the largest fluid-bed solid waste incinerator in the world
 an installation which we feel is the forerunner of the next  generation of solid waste disposal
systems and the best means of solving the solid waste disposal problem.
 INTRODUCTION

   Having been one of the  midwives at the
birth of fluidized-bed technology in the 1940's,
it is a pleasure for me to be here today at this
Third International Conference on Fluidized-
Bed Combustion, and to be able to discuss the
technology  with  so many people who are
knowledgeable in its applications.

   As a matter of fact, I was one of the first
persons assigned the responsibility of selling
fluid-bed technology outside of the oil refining
industry.

   Looking back on those years of trying to
convince the technical world that we  weren't
crazy, and at the same time having to show
sales progress while doing research all day, all
night and weekends, it is a wonder that those
of us  who  were  engaged  in  this  exciting
development did not throw  up  our hands in
disgust and  go  to something more certain.
   In those, days when the new applications
were basically metallurgical, it seemed that all
fluid-bed installations would  always be in
out-of-the-way places like Red Lake, Ontario;
Arvida,  Quebec; or Berlin, New Hampshire.
For sure, however, all startups over a five-year
period were in the dead of winter, and could
only be  described as miserable.
                                                                               H-2-1

-------
   Today, our startups are in places like Flor-
 ida, Japan, South Africa, the Bahamas  as
 always, in the winter time; but the climate is
 certainly much better on the average, and we
 have the technical understanding and support
 of people who have  learned something about
 the  technique.

   The technology has come a long way in the
 intervening  25 years, and I believe it  is now
 coming into its own with  new  and exciting
 applications  popping up all over  the  world.
 Having been active in the business throughout
 these years,  I think my greatest satisfaction
 comes from  meeting young engineers, fresh
 out  of school, who have been taught at least
 the basic elements of the techniques and can
 go on to understand the many permutations
 and possibilities  of fluid-bed processing.

   The company, which bears my  name, has
 specialized in air and water pollution control
 since the early 1960's and uses fluid-bed com-
 bustion  technology  as  the  heart of many
 patented processes which the company sells on
 a worldwide basis.

   We  now  have fluid-bed waste disposal
 combustion  installations in virtually every
 major industry, many on a first time basis, and
 have successfully used the technique on both
 liquid and solid  waste  materials.

   We like to think that in developing certain
 fluid-bed processes  to  dispose of  pulp  and
 paper mill  wastes,  we  did  accomplish the
 impossible by burning organic  matter in a
fluid  bed in the presence of inorganic salts
having a low fusion point. We did,  in fact,
what  most of the early publications on fluid
beds  said was impossible.

   Our first such commercial installation was
in pulp and paper pollution control, where we
were  required to destroy  the organic matter
removed from wood pulp in an inorganic solu-
tion of sodium-sulfur salts. We knew that a
fluid-bed temperature in excess of 1300F was
necessary to get complete burn-out of organic
(polluting) matter; at the same time we knew
that the inorganic salts would fuse at  1350 to
1370F depending on the inorganic mix.

   By taking advantage of the partial fusion of
inorganics, we were able to force pelletization
of the inorganics to form the fluid-bed  me-
dium; our  systems actually  operate easily
between  1300 and  1350F. We have several
such  plants which  have  been in  successful
operation for over  10 years  providing a
solution  to  pollution problems which  were
considered impossible in  the 1950's.

   For the most part, our early development of
fluid-bed  technology for  pollution  control
dealt  with waste solutions. In  1971, however,
we built what we believe  is the largest  solid
waste fluid-bed  incinerator in the world at
Great Lakes Paper Company in Thunder Bay,
Ontario. We believe this  installation broke
some  new ground in solid waste disposal by
fluid-bed technology. This paper covers that
installation in some detail.
11-2-2

-------
FLUID-BED
GENERAL
INCINERATORS    IN
   The principle of fluidizing solid materials
at elevated temperatures in the presence of
and by means of a gas, was first commercially
developed by the oil refining industry in the
form  of fluid-bed catalytic  crackers. While
fluid-bed incinerators only vaguely resemble
"cat crackers" they do function because solid
particles are set in fluid motion (fluid ized) in
an enclosed space (fluid-bed zone) by passing
combustion air through the fluid-bed zone in
such a way as to set all particles in that zone in
a homogeneous  boiling motion.

   In this  state, the  particles are separated
from  each other  by an  envelope  of the
fluidizing gas (air for combustion) and present
an extended surface for a gas to solid reaction,
as for example, air to carbon-hydrogen. This
extended combustion surface makes  possible
the high  thermal efficiency found-in  most
fluid-bed reactors.

   Capacity is a  function of  reactor bed total
area,  but is  usually  expressed as fluid-bed
surface area.

   At combustion equilibrium, the fluid bed
resembles a boiling liquid and, in fact, obeys
most of the hydraulic laws. The dispersion of
fluidizing gas  throughout the fluid-bed  zone
by the specially designed orifice plate, assures
complete mixing; temperature variations from
any one spot in the fluid bed to any other will
not normally exceed 10 to  15F.

   The mass of fluid-bed medium is kept at
combustion temperature  by  oxidation of the
organic material in the feed by the oxygen
contained in the fluidizing air. There is little
or no flame, but rather a glowing condition.
Combustion is virtually instantaneous and the
fluid-bed  proper will contain no  unburnt
organic material. Complete  oxidation is the
key to the  control of air pollution.

   It is fundamental to any  incineration pro-
cess that there be no air pollution problem
resulting  from  incomplete  combustion  of
waste solids. In this respect,  fluid-bed units
are superior to any other type of combustion.
This is  again  due  to the  extended  surface
presented by the fluid-bed medium. Copeland
designed units normally will burn in excess of
300,000 Btu/ft2-hr.  This could be compared,
for example, with coal burning boilers which
do  well  to  consume 40,000 Btu/ft2 of grate
area/hr.

   A commercial unit burning sewage sludge
at 1400 F had the stack gas analyses shown in
Table 1.
    Table 1.  MASS SPECTR06RAPHIC
               ANALYSES^
Component
Volume, %b
C02
A
02
N2
Volume, ppmc
S02
COS
H2S
NOX
Hydrocarbon
Sample number
1

5.17
0.73
14.4
79.7
ND





2

8.33
0.07
9.2
81.8
ND





3

8.86
0.60
7.4
83.0
ND





                               Dry basis.
                               Samples  were taken at  11:30 a.m.,
                               October  30, 1968.
                              cNot detectable  by mass  spectro-
                               graphic  analysis.

                                Every installation built in the last ten years
                             is meeting the air pollution regulations of the
                             state in which it is located and complies with
                             the most stringent air pollution regulations  of
                             the country.  This applies to both gaseous and
                             particulate matter in exhaust gases.
                                Perhaps  the  most important  and  least
                             understood feature of Copeland incinerators is
                             the ability to pelletize  most  inorganic  ash
                             residues to form the fluid-bed medium itself.
                             Many fluid-bed incinerators use sized sand  as
                             the fluid-bed medium, but the presence of low
                                                                                  n-2-3

-------
melting point inorganics makes pelletization a
useful technique which in some cases permits
the  recovery of a  saleable product  uncon-
taminated by sand or other diluents. Pelletized
ash  is dust free and easy to handle. For this
reason,  where possible, we endeavor to  force
pellet growth in our incinerator systems.

   In every case of forced pelletization, we find
reduced dust  collection  problems  and  a
generally easier operation. Pellet growth is a
function of temperature  and surface area in
the  bed  itself;  its rate  is  controlled  by
controlling both temperature and unit  area in
the bed. By occasional screen analysis of the
bed product, we  can predict rate  of  growth
and adjust it to the needs of the system.

   Fluid-bed systems capacities are a function
of  superficial space velocity or the rate at
which the fluidizing gas is forced through the
fluid bed. A unit designed for a 2 ft/sec space
velocity  will  give  100   percent  additional
capacity if enough air is  forced through it to
raise the velocity to 4  ft/sec,  provided the
velocity is not  sufficient  to elutriate the bed
material out of the reactor. Space velocities
are  chosen to meet a   given  set of  feed
conditions and will generally be in the range of
1.0 to 5.0 ft/sec. Installations have been  built
for  other purposes where  the velocity has
exceeded 10 ft/sec.

   Of great importance in basic design is the
opportunity to  build into fluid-bed systems a
future capacity at minimal cost. By installing a
false brick lining in  the fluid-bed zone at time
of original construction, combustion area is
provided for future use. For example, a project
requiring a 10-ft diameter reactor can get 30
percent  future capacity built in by adding a 1-
ft thick  false lining, for less than a 5 percent
increase in present  cost.


DESIGN CONSIDERATIONS FOR FLUID-
BED INCINERATION PROCESSES

   Fundamental to the design of any fluid-bed
unit is  a clear understanding  of  the waste
material to be processed.  Fluid  beds will not
work on material too coarse in size to be
fluidized or having an ash content with a lower
melting point than the temperature necessary
for complete oxidation of the organic matter.
In this latter case, fluid beds do not differ
from the older type incinerators.

   Generally it can be said that fluid beds will
burn anything that can be fed into them and
fluidized. Solid wastes with a minimum of free
surface water are generally  blown  into the
reactor, whereas drier materials can be fed by
means of a sealing type screw conveyor. Semi-
plastic sludges such as sewage sludge are fed
by screw conveyor or more simply by a pro-
gressing   cavity   pump.   Thermoplastic
materials like grease are most readily fed by
first  being  melted  and  then  pumped  by
centrifugal pump.
   Wherever possible in our design, we try to
build  into   the  system  enough  freeboard
residence time to permit some heat exchange
between the  incoming  wet feed and  the out-
going combustion gases. Since these gases are
generally wasted, evaporation of water from
the incoming feed by direct heat exchange has
the effect of improving thermal efficiency.
   Before designing any fluid-bed system, we
pay  particular  attention  to the  chemical
composition of the material and look for trace
elements which might have a fluxing or fusion
point lowering effect on the ash content. We
have  found,  for example,  that  a fluid  bed
burning organic material, the ash content of
which is 100  percent sodium chloride, can be
operated in excess of 1300F without fusion
problems. Yet another product  of the same
type having 1.5 percent sodium chloride could
not be burned in excess of  1150F without
complete fusion. Obviously, this is a typical
eutectic problem,  but  critical in any incin-
erat^r design.


VIRTUES  OF  FLUID BEDS  USED  IN
SOLED WASTE INCINERATION

   The   higher   combustion  efficiency  of
fluidized beds is attributable to a number of
II-2-4

-------
characteristics found  only in part in other
combustion techniques. These  may  be
described  as follows:

Extended Surface

   Total surface area,  in or on which the com-
bustion process  takes  place,   is  a  very
important design consideration.

   We have a commercial installation burning
60 x 106 Btu/hr  (waste sulfite liquor) which
has surface area in the bed medium (90 tons)
equivalent to the surface of a super highway 70
miles long.  Bed medium in this case is pel-
letized  inorganic  salt  recovered  as  a  by-
product of combustion. Pellet size is basically
14 to 65 mesh Tyler  screen scale.

   Obviously,  the  more  surface  area,  the
better  the opportunity for reaction  between
oxygen in the fluidizing air and combustible
material.

Residence Time

   Combustion is a time-temperature reaction
which  is most efficiently carried out  under
conditions which  give instantaneous reaction.
Lack of time or  temperature will make  for
incomplete  reaction  and  produce  partial
products of combustion.
                            i
   In  systems  where temperature  must  be
controlled at lower limits because of other
thermal considerations,  residence time there-
fore becomes an  important factor.  We have
noted in systems  burning waste sulfite liquor
at 35 percent solids that combustion at 1300F
is instantaneous with no residual carbon left in
the bed. However,  at 1250F combustion is
slow and, if allowed to proceed for any length
of  time,  carbon  build-up  in  the  bed   is
noticeable.

TYPICAL APPLICATIONS OF FLUID-BED
INCINERATORS

   Fluid-bed incinerators are finding  appli-
cation  in the combustion of waste solids and
liquids because in  many cases  these waste
materials   could  not   economically   or
practically be incinerated by older more con-
ventional equipment. The very fact that many
waste solid materials have been used for land
fill, rather than completely destroyed by com-
bustion,  is  usually  an indication  of some
difficulty with conventional processes.

   Since disposal usually implies an outright
cost to the producer, the most efficient system
must  be found; fluid beds are  being  chosen
because of higher thermal efficiencies, better
control  of  odors  and  particulate  matter
emissions,  and  a  simpler  process   having
greater design latitude.

   Fluid-bed   incinerators   have   been
demonstrated by commercial practice  to be
readily applicable to the combustion  of the
following types  of solid  waste materials:


   1.  Domestic sewage sludge.
   2.  Municipal garbage.
   3.  Oil refinery wastes such as:
      API separator sludge,
      tank bottoms,
      waste caustic streams, and
      general refuse.
   4.  Petrochemical wastes such as:
      hydrocarbon compound sludges, and
      complexed waste inorganics.
   5.  Water treatment  plant carbonate
      sludges.
   6.  Packing house wastes.
   7.  Distillery slops.
   8.  Pharmaceutical plant wastes.
   9.  Clarifier effluents from most industries.
  10.  The destruction  of lethally poisonous
      materials.
  11.  Pulp and paper mill sludges and various
      solid wastes.
   The foregoing list is by no means complete;
it will be seen that some waste materials could
be destroyed by older techniques. However,
the  high  thermal efficiency of  fluid  beds
makes it possible to incinerate these materials
at much higher water contents without the use
of extraneous  fuel,  thus  giving  fluid  bed
incinerators the nickname  "water burners."
                                                                                   II-2-5

-------
 FLUIDIZED BEDS VERSUS POLLUTION
 CONTROL

 Water Pollution

   Combustion of waste material by fluid-bed
 techniques is  the ultimate of disposal tech-
 niques in that no liquid effluent need result.
 Ash produced by  fluid-bed combustion is
 completely burned out, impresses no BOD if
 used for land  fill, and many inorganic ashes
 can  be   reused  chemically.  Scrubbing  of
 exhaust  gases for air  pollution  control is a
 necessity; but in most cases we scrub with the
 waste  liquid  being combusted  and hence
 produce no new  effluent.

 Air Pollution

   Most fluid  beds used in industrial waste
 disposal use forced air fluidization which per-
 mits an  easier route  to high pressure drop
 treatment of exhaust gas.  We usually install
 dry cyclones after the reactor and take up to
 15 inches of water drop across them. These are
 generally followed by wet scrubbing where  up
 to 50 inches of water pressure drop are taken.

   We find that these systems, although costly
 in terms of power consumption, are exceeding
 by as much as 50 percent the most stringent
 air pollution regulations now in  effect.  Our
 experience  indicates  that,  if   sub-micron
 particulate  matter  is  to   be taken  out  of
 exhaust  gas,   high  pressure  drop  across
 scrubbing  equipment  is  necessary.  If  sub-
 micron particles escape the scrubber, a tail gas
 plume will  persist  in  the atmosphere.  Any
 persistent plume visible to the public is  an
 open invitation to investigation by pollution
 control  authorities.

   Regardless of  present  air  pollution
regulations,  our experience tells us that the
ultimate  regulation will demand  no visible
plume  whatsoever;  this may  include water
vapor plumes  as  well.

   We  believe  that  fluid-bed  combustion
systems, properly designed and incorporating
the newest  scrubbing  techniques, offer  the
best  answer  to  ultimate  air   pollution
regulations.

   A  recent trend in municipal sludge incin-
eration, promoted by new EPA regulations,
will require that exit gas from sludge incin-
eration be exposed to a temperature of 1600F
for 2 seconds to destroy malodorous gases and
chlorinated  hydrocarbon  compounds  which
are produced by incineration at lower temper-
atures. This will make it impossible to dispose
of municipal sludge by  older  incineration
methods without fuel burning after-burners.


WORLD'S  LARGEST   SYSTEM  FOR
BURNING BARK, DEBRIS, AND SLUDGE

   In  December of 1971,  we  brought on
stream a  solid waste,  fluid-bed incinerator,
which disposes of about 600 ton/day of pulp
and paper mill  wet  wood  waste at  a water
content of 65 to 70 percent without the use of
extraneous fuel. The composition of the feed is
given  in Table 2.
   The Great Lakes  Paper Company Limited
at Thunder Bay, Ontario, had two solid waste
disposal problems.

   The first problem was a huge pile of bark
(300,000 yd3) which had  been accumulating
for years, and  was  beginning to  make its
presence felt by spilling over  into the local
river.  The second problem arose from  the
installation of clarifiers on the effluents from
the groundwood mill, the sulfite mill, and the
Kraft mill, the sludge from all of which would
have to be disposed of.

   A conventional boiler was already in use at
the mill  for burning  bark fresh  from  the
barking drums.  But (apart altogether from
questions of existing capacity) disposal of the
clarifier  sludge  and  pile  bark  by  such
conventional means would require that they be
further dewatered. The old piled bark, how-
ever, contained all manner of junk, including
a generous proportion of stones, which repre-
sented a potential source of damage to existing
bark  presses. And the clarifier sludge, on
account of its slimy, fibrous nature, was very
II-2-6

-------
  Table 2. PROPERTIES AND DAILY QUANTITIES OF FEED TO COPELAND SYSTEM

Sludge feed
Sludge from groundwood clarifier
Sludge from kraft clarifier
Rejects from groundwood mill
Rejects from kraft mill
Average properties and total amount
Wood waste feed
Bark and wood debris
Surplus bark
Average properties and total amount
Dry solids
%

25
25
25
25
25

35
35
35
Ash, %

2.2
2.3
1.9
1.9
2.2

2.0
2.1
2.0
Btu/lb

7100
7155
7155
7155
7155

8500
8500
8500
Ton/day

40
5
5
5
55

50
20
70
 hard to dewater over 25-28 percent solids. At
 this  low   consistency,  it  would   cause
 combustion troubles in the conventional bark
 burning furnace.

   Consequently, Great  Lakes'  management
 turned to the Copeland fluidized-bed process,
 which offered as one of its characteristics the
 ability to burn woody materials with as little as
 30 percent solids without supplementary fuel.
 This would make  it possible to  burn, self-
 sufficiently, the  impressed  bark and/or  a
 mixture of the unpressed bark  and clarifier
 sludge.  A Copeland unit of 180 BD ton/day
 capacity was therefore decided on.

   The  next question  was whether this unit
 should be attached to a waste heat boiler for
 raising steam. It was recalled, however, that
 24,000 Ib/hr of steam was currently being used
 to heat  4000 gal/min of hot water for wood-
 room showers. Since this heat demand was of
 the same order as that expected to be available
 by recovery from the  Copeland unit,  it was
 decided to produce hot water directly in the
 unit's heat recovery-scrubbing system. In this
 way, the mill steam supply available for other
 uses was increased, without going  to the
 expense  of installing  another  waste heat
 boiler.

 Bark Feed
   The bark from the old pile and other wood
debris is picked up  in the mill yard by truck,
and combined with slasher sawdust, ground-
wood snipes,  and wood scraps in  a surge
hopper. From here the waste is mechanically
conveyed to a wood hog, which breaks it into
fragments which can be handled efficiently by
the subsequent pneumatic conveying system
(smaller  than 6x6x6  inches).  The  shredded
waste resulting is then transferred  by bucket
elevator to a storage silo. This silo holds 1
day's feed to the system, so that a man for
collecting bark and wood debris is needed  on
one  shift only. The silo  has a live-bottom
vibrating hopper, which discharges the waste
through a vibrating feeder into a pneumatic
conveyor which injects it  into  the reactor
immediately above the fluid bed.

   Stones up to 6-in. diameter in the bark and
wood debris, are conveyed along with wood
waste to the fluid-bed reactor by a pneumatic
feeder. Tramp metal is removed by an electro-
magnet.

Sludge Feed

   The   wet,   fibrous  clarifier  sludge  is
pneumatically conveyed to the disposal system
from  the  sludge  collecting  tanks,  after
dewatering by filtration. Though it could  be
injected directly into the reactor  above  the
fluid bed, the large amount of air required to
convey it would, under these conditions, enter
the reactor and critically lower its operating
temperature. Normally, therefore, the sludge
                                                                                   11-2-7

-------
 is separated from its conveying air in a cyclone
 from which it is mixed with the normal wood
 wastes.

 Combustion System
   The fluid-bed  reactor,  is  a carbon  steel
 vessel lined with  insulating  and refractory
 brick. A  five-stage, centrifugal, low-pressure
 blower supplies the fluidizing and combustion
 air.  The fluidizing air is distributed into the
 bed by an orifice plate separating the windbox
 from  the  bed section.  The  fluid-bed  zone
 contains the fluidizing medium, which is made
 up of sand removed  from  the waste wood.

   Above the bed zone, the reactor widens out
 to form the disengagement or  freeboard zone.
 The  increase  in  reactor  diameter here is
 sufficient  even in view  of  the increase in
 volume of gas phase due to the generation of
 water vapor  to reduce the upward velocity
 of the gases  so  that particulate  solids  will
 disengage and fall back into the fluidized bed.

   Overbed burners are provided for startups
 as necessary.  Auxiliary fuel is not  required
 when the system is fed waste materials at 75 to
 120 percent of design capacity.

 Sand Handling System

   It is necessary, from time to time, to with-
 draw excess sand and grit from the bed. This
 is done'by discharging it into  a sealing screw
 conveyor, and then to a storage silo.

 Gas Scrubber & Hot Water Generator

   The hot  combustion  gases leaving  the
 reactor at about 1600F pass into a two-stage
 scrubbing system.

   The first  stage  is an  adjustable wetted -
throat  venturi   scrubber,  in which   the
scrubbing water is introduced over  a weir and
atomized by the  energy  of the venturi. The
resulting  fine droplets  contact  the   ash
contained in the  combustion  gases  and  are
separated from the  gas  in  the separator
section. The secondary scrubber consists of
two  beds  of  fluidized packing,  on  which
scrubbing water is sprayed, thus trapping any
residual particulate matter while at the same
time picking up heat. Finally, the gases pass
through a demister to remove entrained fine
droplets before being vented to atmosphere.
Since  no particulate  matter,  objectionable
gases, or odor have been detected in these exit
gases, a high stack to disperse them has not
been found  necessary.

   The scrubbing  water from both first- and
second-stage scrubbers is collected in an 8500
gallon  reservoir below the separator section,
from which  it is recycled through  a pump to
the two stages of scrubber trays. The recycle
line is  provided with a purge line for main-
taining  the  desired  temperature  and  ash
concentration  in  the scrubber  liquid. From
this reservoir is  taken 4000 gal/min of hot
water at 155F used in the woodroom.
Operation & Control

   One man operates the plant from a central
control  panel.  During  startup   opera'ting
temperature  is reached  with  auxiliary fuel;
when the feed is ignited the auxiliary fuel is
shut  off.  Shut down  is  accomplished  by
shutting  off the waste  feed and the blower
furnishing  the  fluidizing air.  The  sand bed
normally loses less than 100F/day, so that the
unit can  be started up without auxiliary fuel
after being down for as long as 6 days.


   Cost  of the fluid-bed  installation  was
approximately $1 million.  Labor required is
1.3 man-days/day,  and maintenance costs are
expected  to be low. The operating credit of
24,000  Ib/hr  of   steam  is   reckoned  at
S100,000/yr. A further credit is the valuable
land that will become available when  the old
bark pile has been disposed of. Present feed to
the unit is  125 ton/day, but short runs have
shown that 180 ton/day is feasible. The Air
Management  group   of  the  Ontario
Department  of Energy  and Resources has
monitored  the  system thoroughly, and has
found the exit gases to contain  no particulate
matter or objectionable gases.
II-2-8

-------
SUMMARY


   In bringing this large fluid-bed solid waste
incinerator into operation, we have made some
interesting discoveries which lead us to believe
that the same technique can be used on such
other solid waste disposal problems, such as
garbage, etc.

   To our great surprise,  we have found the
reactor able to accept large numbers of stones
up to 6-in.  in diameter,  non-magnetic scrap
metal,  cans, and  a variety of  things  from
shredded truck tires to welding rods.

   We are fluid izing bed rock particles up to
1-in. cubes  and still maintaining good bed
stability.
   The unit has never produced the first wisp
of smoke; air pollution control is so good that
pressure drop  across  the  double scrubbing
system has been reduced from 50 in. H2O to
15 to 20 in. F^O, while still meeting the very
stringent  air  pollution  regulations  of   the
Province of Ontario.

   The ability of the unit to dispose of large
pieces  of  solid waste has encouraged Great
Lakes Paper to institute a change of feed flow
which will eliminate the hog completely. When
done, all  waste, including  waste pulp wood
sticks up to 4 ft. long will be fed directly into
the bed without any size breakdown.

   The  installation    argues    well  for
uncomminuted garbage incineration by fluid-
beds  a development  which is sorely needed
today.
                                                                                    II-2-9

-------
WATER
  VENTURI
  SCRUBBER
                         .HOT WATER
                           TO MILL
   ASH
DISCHARGE
                                                               PREHEATER   REACTOR
DIS. SCREW
BLOWER
FEEDER
                                      Figure 1. Copeland solid waste Incinerator system.

-------
         3. FLUIDIZED-BED COMBUSTION OF MUNICIPAL
              SOLID  WASTE  IN THE  CPU-400  PILOT PLANT
                     G.  L. WADE  AND D. A. FURLONG
                         Combustion Power Company
INTRODUCTION
   A series of experiments is currently being
carried out to develop a large, high pressure
fluidized-bed combustor for incineration of
municiple refuse. This publication is based on
work performed under Contracts PH 86-68-
198 and 68-03-0054 with the National Envi-
ronmental Research Center, Cincinnati, Ohio,
of the Environmental Protection Agency.

   The  solid  waste  management concept
known as the CPU-400 is sized to accept 400
ton/day of solid waste from municipal packer
trucks. Solid waste will be conveyed from
receiving pits directly into  shredders which
discharge  into  the  air classifier units. Com-
ponents  such as  metal and glass with  high
weight to aerodynamic drag ratios will be sep-
arated out and conveyed to ancillary disposal
or recycling processes. Lighter materials, pre-
dominantly papers and plastics, will be trans-
ported  to a storage conveyor  of sufficient
capacity to provide a continuous supply of
combustor fuel at a uniform rate.

   The shredded and classified solid waste is
fed into the fluidized-bed combustor through
high pressure  air-lock feeder valves and a
pneumatic transport line. In the fluidized bed,
inert sand-sized  particles  are  buoyed  and
mixed by an upward flow of air coming from
the compressor. Heat released by solid waste
combustion will maintain the fluidized bed
and exiting gas products between 1300 and
1700F.
   Several particulate removal  stages will
operate on the hot combustion gases prior to
their admission into  the gas turbine. Inert
granular  residue will be removed from the
fluidized  bed  and   particle  collectors   as
required.

   The economic basis for the CPU-400 lies
primarily in the recovery of the energy con-
tained in solid waste by virtue of its high con-
tent of paper, plastic, and wood products. The
recovery and sale of electric power from the
disposal of solid waste markedly reduces the
cost of the operation. Depending on the value
of electrical power and other local conditions,
estimated net operating costs range from 2 to
$5/ton compared with current incinerator
costs of 8 to $14/ton.  Where a stable market
exists for other reusable materials such  as
metals or glass, the CPU-400 will also permit
recovery of these resources; the process acts to
concentrate these  recyclable  materials   by
removing the  large  volume of combustible
materials. Revenues derived from  recycling
will serve to further reduce net operating costs.

  The CPU-400 is now in the early stage of its
development;  system  studies  and  subscale
experiments  have  been   completed,  and
development of the pilot plant is well  under
way. The  fluidized-bed combustor and partic-
ulate removal stages currently in development
testing are components of the  pilot plant.
                                       II-3-1

-------
ADVANTAGES   OF   FLUIDIZED-BED
COMBUSTION

   The fluidized-bed reactor is a relatively new
approach to the design  of high  heat release
combustors. The primary functions of the air-
fluid ized inert bed material are to  promote
dispersion  of  incoming  solid  fuel  particles,
heat them rapidly to ignition temperature, and
to promote sufficient residence time  for their
complete combustion   within  the  reactor.
Secondary functions  include  the  uniform
heating of excess  air and the generation of
favorable conditions for  residue removal.

   The fluidized-bed reactor greatly increases
the burning rate of the refuse  for three basic
reasons:

  1. The  rate of  pyrolysis of the  solid waste
    material is  increased by  direct contact
    with the hot inert bed material.

  2. The charred  surface of the burning solid
    material is continuously abraded by the
    bed  material,  enhancing the rate of new
    char  formation  and  the  rate  of char
    oxidation.

  3. Gases in the bed are continuously  mixed
    by the bed material, thus enhancing the
    flow of gases to and from the burning solid
    surface and  enhancing the completeness
    and   rate   of  gas   phase  combustion
    reaction.

   A  significant advantage of the fluidized-
bed reactor over conventional incinerators is
its ability to reduce noxious gas emission. Five
types  of noxious  gas are of potential  concern.
The anticipated  ability  of the fluidized-bed
reactor  to reduce each  of these  will be
separately discussed.

 1. Oxides of  Nitrogen.  The  relatively low,
    uniform temperature of the fluidized-bed
    reactor  (1300  to   1700F)  limits  the
    formation  of  oxides of nitrogen.  Most
    combustion chamber concepts require a
    hot,  primary combustion zone to  assure
    good combustion efficiency; it is in these
    hot zones that most oxides of nitrogen are
   formed.  Because of extensive  mixing of
   the  fluidized  bed,  excellent combustion
   efficiency is realized without a hot primary
   zone.

2.  Oxides  of  Sulfur.  The effectiveness  of
   limestone  for  control  of  SO 2 emission
   from coal  combustion chambers  is being
   demonstrated.  After  injection  into  the
   combustion chamber, the limestone is first
   calcinated to lime. The SO2 is oxidized to
   SO 3 on the lime surface and then reacted
   to calcium sulfate (CaSO^, which  remains
   with  the  ash. A disadvantage  of this
   process  for  coal   combustion  is  the
   relatively  large  quantity   of  limestone
   required. Test  data  show  that limestone
   may be only partially  reacted because of
   short residence time in the  furnace result-
   ing in calcium sulfate accumulation only
   on the  surface of the limestone. This
   problem is reduced with the fluidized bed
   since the increased residence time in the
   fluidized bed strongly favors the  capture
   of sulfur by limestone. In  addition^ solid
   wastes average less than 0.5 percent sulfur
   and the solid waste inerts already contain
   significant quantities of calcium and mag-
   nesium oxides. Thus, little  if any lime-
   stone additive  is required  when  burning
   solid waste in  a fluidized  bed.

3.  Hydrogen   Halides.   The   emission   of
   hydrogen halides, primarily HC1, can  be
   expected to be a significant problem for
   future incinerators; probably more signif-
   icant  than SO 2  emission.  Although
   limited experimental work  exists  on HC1
   suppression, chemical considerations
   indicate that reactants similar to  those
   previously described for SO 2 suppression
   may b.e effective for HC1  suppression.

4.  Carbon Monoxide, and

5.  Hydrocarbons.  The highly  mixed  oxygen-
   rich  environment of the  fluidized-bed
   reactor provides very favorable conditions
   for complete combustion, thus minimizing
   carbon  monoxide  and  hydrocarbon
   emission.
II-3-2

-------
   Finally, the  fluidized  bed is unique in its
ability to efficiently consume low quality fuels.
The   relatively  high  inerts  and   moisture
content of solid waste pose no serious problem
and require no associated  additional devices
for their removal.

THE CPU-400 PILOT PLANT

   The CPU-400 pilot plant development is
currently nearing completion in Menlo Park,
California. A systematic  evolution is planned
to ultimately include all major components on
a pilot plant level. The planned high pressure
configuration  (with  integrated gas turbine)
consists of four primary subsystems. Three of
these,   broken  down  into  their  major
constituent parts,  are illustrated in Figure 1.
These are the solid waste handling subsystem;
the solid waste combustor and gas preparation
subsystem; and the  turbo-electric subsystem.
A fourth area, the control system, is  both a
part of each of the other three subsystems and
also a separate  subsystem in itself, causing the
other subsystems to interact properly with one
another  and  respond  correctly  to external
commands.

   The  purpose of the solid waste  handling
subsystem is to prepare the solid  waste  for
combustion. This  includes separation of the
shredded  material  into  two   constituents
(materials dominated  by  combustible  and
inert  components, respectively),  storing the
combustibles  until  ready  for  use  in the
combustor, and metering the solid waste to the
combustor. Unprocessed municipal waste is
initially loaded onto  the shredder  conveyor by
a  skip loader. The  conveyor modulates the
feed to  the shredder  based upon  electrical
loading of the shredder  motor. After shred-
ding,  the material is fed to the air classifier
where the light,  combustible  materials are
pneumatically  lifted and transported  to the
storage bin, while the  heavy, inert materials
drop  out for subsequent separation  and
recovery. Metering of the prepared solid waste
fuel is accomplished through a variable-speed,
servo-controlled  outfeed  conveyor  in  the
storage bin along with variable speed transfer
and weighing conveyors to the combustor feed
points.

   The second subsystem of the high pressure
pilot  plant  consists  of  the  solid  waste
combustor, three particulate removal stages,
ash removal equipment,  and interconnecting
piping  and  valving.  The solid  waste   is
introduced into the  combustor  through air-
lock feeder valves. The material is burned  in
the solid waste combustor and  the resulting
hot gases are then cleaned of suspended solid
material in three stages of separation. Fly ash
material  is   removed   from   the  particle
separator  collection  hoppers  by pneumatic
transport to a baghouse  filter.

   The  third subsystem  is the turbo-electric
unit consisting of a gas turbine, generator,
switch gear, and load bank. The compressor
section of the turbine supplies the cold air for
the solid waste  combustor fluidization. The
resulting hot gases, after  being cleaned in the
separators, are used to power the compressor
turbine and the power turbine. The generator
is driven by the power turbine and generates
power which is subsequently controlled by the
switch gear. The electrical energy output  of
this pilot plant system  will be  dissipated in a
combination load and  light bank. In the full
scale  system,  the electrical  power will  be
delivered to  a  customer for subsequent use  in
the municipality.

   The control subsystem interacts with these
systems to control their respective outputs  in
response to  commanded  set  points.  This
system, which  features   analog  controllers
under the  supervisory control  of a digital
process computer, will also monitor numerous
signals  to provide data  acquisition, logging,
out-of-tolerance alarming, and status display
functions.

   In the low pressure configuration operated
to date, the gas turbine compressor is replaced
by  a  facility blower and exhaust gases are
cooled by water spray.  Consequently, there is
no further discussion of the turbo-electric sub-
system in this paper. Since incorporation  of
the process  control  computer into the pilot
                                                                                    II-3-3

-------
 plant  is  incomplete  at  this  writing,  few
 references  will  be  made  to  the  control
 subsystem.

 Solid Waste Handling Subsystem

   The  solid   waste   handling   subsystem
 processes municipal solid waste unloaded by
 packer trucks  at  the  pilot  plant  facility.
 Photographic  views  of the  subsystem  are
 shown in Figure 2. The flow of solid  waste is
 illustrated in Figure 3. A 16 ft3 skip loader is
 used for transfer of the raw solid waste to the
 shredder's  conveyor hopper. Through elec-
 trical  current   level controls, the  conveyor
 supplies material on demand to the shredder.
 Shredded material is ejected into the air class-
 ifier unit. In this unit heavy metallics, rocks,
 glass, etc., are  dropped out; the remainder is
 pneumatically  transported up through  the
 chamber  and   into the cyclone  above  the
 storage  tank  where  it is  disengaged   and
 deposited in the tank. The exhaust stack of the
 cyclone connects to a second blower on a dust
 filter cyclone. This unit pulls off vapors  and
 dust  from  the center of the  storage tank
 cyclone, depositing the dust in a  container at
 floor level  and venting excess air to  the
 atmosphere.

   Very little hand sorting  of the delivered
 solid waste  is conducted prior to shredding.
 On rare occasions  an  item too large for the
 shredder  inlet  (e.g.,  a  large  truck  tire) is
 encountered;  these are manually removed.
 Power limitations in the 75-hp shredder have
 dictated that massive metal items and fabric
 bundles (both infrequently found in municipal
 solid  waste) also be removed. Other highly
 visible  items such as automobile  tires  and
 mattresses,  though successfully shredded  on
 occasion,  are   normally  removed  in   the
 interests of maintaining high  through-put.
 Under these conditions a nominal rate of 1.5-
2.0 ton/hr has been established, and over  350
tons of solid waste have been shredded and air
classified through August  1972.  Concessions
to subscale  pilot plant  operation  essentially
disappear when 1000  hp shredders,  such as
planned for the CPU-400,  are employed.
   The shredded and air classified solid waste
fuel  form   is  a  mixture   whose   visual
appearance is homogeneous and dominated
by  identifiable  paper  products. A  typical
pound of the material consists of 0.30 Ib of
water, 0.52  Ib of ash-free combustibles, and
0.18 Ib of inerts (including ash). The latter two
fractions are typically sub-divided as follows:

         Description     % by weight
         Paper
         Thin metals,
           fine glass,
           and dirt
         Wood
         Dust
         Plastics
         Textiles
55
15
13
13
 2
 2
   As would be expected, all of the preceding
values are  subject to considerable variation.
As an example, moisture content in the final
fuel  form ranges from  10 to  40 percent by
weight depending upon the  origin  of raw
material  (e.g.,  residential or  commercial
sources), time of year, weather conditions, and
municipal collection policies.

   Air classified solid waste fuel is shown in
the storage tank in Figure 4. The accompany-
ing view of the empty tank interior shows three
of the four rotating bucket chains which sweep
the floor to move material into the  outfeed
conveyor located in the slot. Free trailing ends
of the bucket  chains allow edge contact with
the stored material pile regardless of pile size.
Variable speed hydraulic drives in the sweep
system and outfeed conveyors are controlled to
maintain material levels in a small hopper at
the outfeed conveyor discharge point.

   The hopper volume  lies above a  transfer
conveyor and  upstream of material  leveling
devices that produce a  fixed material height
on the transfer conveyor. Fuel is delivered by
the transfer conveyor to a  weighing conveyor
(equipped   with  load   cells  to   provide
continuous measurement of fuel flowrate) and
then   dropped  through   a  static  splitter
assembly -ifrto the two airlock feeder valves.
II-3-4

-------
Variable speed electric drive systems are used
on the transfer and weighing conveyors. Speed
command signals  are both slaved to  a  single
fuel  demand  signal;  this  implements  a
responsive,  continuous  volumetric  flow
element for a combustor temperature control
system.

   Two airlock feeder valves receive the solid
waste fuel from the weighing conveyor  and
deliver it to the air  transport lines, which in
turn pneumatically deliver the material to the
fluidized bed. Each valve (Figure 5) is powered
by a 25-hp electric motor drive and  associated
gear box. As the valves turn, empty pockets in
the valves receive material from the top, rotate
through 180 degrees past a sealing wall,  and
deliver the material into the transport lines at
the bottom. Material is continuously delivered
to the combustor through continuous  rotation
of the valve.
Solid Waste Combustor and  Gas Preparation
Subsystem

     The fluidized-bed combustor is contained
within  a  vertically  oriented  cylindrical
pressure shell with dished heads. The outside
diameter is 9.5 feet and overall height is 23.5
feet. Three layers of insulation protect  the 3/8-
in.  carbon steel  pressure  shell cylindrical
sections from high combustion zone tempera-
tures. A wear-resistant firebrick inner liner is
backed up by a liner of insulating brick. These
refractory layers are separated from the shell
by  a thin  layer  of packed  ceramic  fiber
insulation designed  to isolate the shell from
stresses  induced   by  differential   thermal
expansion. Insulation  in the top dome is
provided by a castable refractory held  in place
by  standard  hangers.

   The  fluidized bed is  supported by a  flat
carbon steel plate welded to the pressure shell
and covered  by two  layers of castable refrac-
tories that provide insulation and wear resis-
tance. Penetrating this assembly are 161 2-in.
pipes capped with wire mesh air diffusers.
Other penetrations  from  the   air  plenum
chamber  beneath  the  plate  permit   bed
temperature   and   pressure   measurements.
   The circular cross section fluidized bed has
an area of 40 ft2' and  is designed to operate
with a superficial velocity in the 5 to 7 ft/sec
range. A nominal 2-ft bed (unfluidized state) is
used  together   with   a   12-ft   freeboard
(unfluidized  bed  surface  to  exhaust  duct
centerline).

   Penetrations through the pressure shell and
refractory  insulation  into the  fluidized bed
provide for two solid  waste  feedpoints. Two
feedpipes  bolted to outer shell  bosses extend
into the bed. Solid waste is fed into the bed
along the  length of these pipes via a slanted
cut  on  the  bottom  side. The  design  and
positioning of these pipes is based  on earlier
tests   where    oxygen   concentration
measurements  established  the dimensional
characteristics  of  combustion  zones.   The
result is a configuration which, in low pressure
testing,  has  demonstrated  very  satisfactory
operation  with respect to geysering due to
feedpipe air flow, minimization of local fuel-
rich zones, and reduction of heat release above
the bed.

   Other bed penetrations provide for possible
removal of excess bed material and for six oil
guns  to permit  fluidization combustion of
auxiliary  diesel oil. This  normally  unused
auxiliary  fuel,  available  primarily  as  a
developmental  tool  for backup  service in
maintaining or establishing desired test condi-
tions, is mixed with air and  carried through
the inner  of two concentric  pipes. The outer
pipe of each gun carries cooling air.

   Initial bed heating is accomplished by hot
products of combustion  from an  oil burner
located in the top dome of the combustor. This
downward  firing  burner forces  hot  gases
through the bed in a  "back heating" mode
that is  capable  of heating the  bed  from
ambient conditions to  1100F in 90 minutes.
This bed temperature, being above the auto-
ignition temperature of either solid waste or
diesel oil,  is an appropriate initial condition
for successful fluidized combustion.
                                                                                    II-3-5

-------
   In  the  low  pressure  configuration,
fluidizing air is supplied to the combustor's air
plenum  by a  125-hp positive  displacement
blower which can deliver up to 7000 scfm at 3
psig. The blower also supplies fuel transport
air in  a parallel path to the fluidized bed. By
appropriate valving, the same blower is used
to drive the back heating mode.

   The top cylindrical section  is removable
and contains the  exhaust port, instrumenta-
tion, and observation ports. Exhaust into the
first particulate removal stage is carried by a
double-walled  pipe  with 26-in. carbon steel
outer  wall and 20-in. type 310 stainless steel
inner  liner. The annular space is packed with
ceramic fiber insulation.

   Two particulate removal stages have been
tested in low pressure operations by the time
of  this  writing.  The first, known  as  the
alumina/sand  separator,  treats a  problem
peculiar to fluidized-bed combustion of solid
waste  fuel;  the  removal  of inert particles of
elutriated bed material,  and the handling of
particles  generated   by  the   presence  of
aluminum in the fuel. While the majority of
aluminum is removed by  air  classification,
about 0.25 percent by weight of the processed
fuel is aluminum in the form of foils, beverage
can  pull-rings,  and  other thin or trapped
particles. The fluidized bed melts, fragments,
and  partially  oxidizes   this material  into
particles  having molten  internal  aluminum
and a frozen oxide (alumina) surface shell that
inhibits  further  rapid  oxidation.  Previous
testing has shown that elutriated particles of
this type tend to generate sizable deposits with
high  aluminum   content on  impingement
surfaces in the  exhaust  gas   stream.  The
possible participation of soft (at combustion
temperatures) bottle glass particles  or  other
binding agents in  this deposition mechanism
should not be discounted, even though present
evidence does not  seem to indicate that they
play a dominant role.

   A  promising solution to the problem of
handling  partially  molten  aluminum and
other sticky particulates is to provide a curved
surface in the first separator stage where the
turning exhaust gas can produce a controlled
deposit. Scrubbing action associated with the
continual impingement and turning of particle
laden exhaust gases acts to  promote  further
oxidation  at  the  deposit  surface.  It  also
promotes the necessary erosion of resultant
alumina-rich  particles  to  stabilize  deposit
shape  and size. Most of the inert alumina-
based particles thus generated as well as the
silica-based  elutriation  particles  are  then
collected  in the settling chamber formed by
the bottom of the vessel. In testing to date,
these separated particles have been allowed to
remain  in the  settling chamber.  For future
tests,  residue   particles collected  in   the
alumina/sand separator stage will be removed
on a continuous basis.

   Fly ash and fine bed material particles are
removed in two  stages of inertial  separation.
The  first  inertial  separator  stage has been
successfully tested by the time of this writing
and  the final stage (similar  design features)
will be  added in the  immediate  future.  An
inertial  separator  assembly  consists  of an
inertial  separator tube holder with a residue
collecting  hopper,  an insulated  cylindrical
housing with inlet and outlet flanges, a dished,
insulated  flanged  head,  and stainless steel
liners.  As  shown in Figure 6,  the hot gases
enter the  inlet  cavity to the  cyclone tubes
through an internal circumferential passage.
The entering gas then turns and flows through
the spirally fintied annular  section of each
tube. These fins  impart rotational motion into
the gas  which centrifuges the particles to the
outside wall. Particles spinning along the out-
side  wall of the  tube will decelerate and  fall
through the opening at the bottom of the tube
into the collecting hopper. To avoid plugging,
the gravitational flow  of particles from each
tube is  assisted by  a  secondary bleed flow.
Particulates are  removed  from the hopper
through the  opening  on the bottom  of  the
hopper  cone.  Gas  which enters the cyclone
tubes turns 180 degrees and exits through the
center  tube into the owtfet  manifold  of the
vessel.  In  the first inertia-1 separator, space is
II-3-6

-------
provided for 48 6-in. tubes; half of the spaces
are  plugged  for   low  pressure  operation,
however.  When  incorporated,  the  second
inertial  separator will have space  for 100 3-
1/2-in. tubes.

   Material from the inertial separator hopper
is pneumatically transported to the baghouse
filter through a finned line. The high tempera-
ture baghouse assembly includes a puffback
bag cleaning system, an  exhaust blower,  a
holding bin, and unloading valve.

   A photographic view of current subsystem
components (Figure 7) also  shows the gas
turbine  in the foreground. Present plans are to
install and integrate this turbine into the pilot
plant by the end of 1972.

SUMMARY  OF  COMBUSTION   TEST
RESULTS

   The  seventh combustion test conducted on
the  low pressure configuration  described in
preceding  paragraphs  was   completed  in
August  1972. The test featured 35 hours of
fluidized   combustion  and   produced  an
extensive  quantity  of data while  consuming
over 69,000 Ib of shredded, air classified solid
waste. During one particular 24-hr period of
special  performance  interest,  a  number of
solid waste and exhaust  gas  stack samples
were drawn for laboratory analysis. Most of
the discussion to follow will be based on data
obtained in this latter period.

   A laboratory facility has been established
at Combustion Power Company to perform
many of the required experiments on material
samples drawn from  the  pilot plant process
operation.  Included in the  installation are
equipment and procedures for the determina-
tion of:
 1. Solid waste moisture fraction,
 2. Solid waste inerts fraction,
 3. Solid waste heating value,
 4. Granular material size distributions,
 5. Particulate loading and size distributions
   from gas samples, and
 6. HC1  gas concentrations  in exhaust  gas
   samples using titration techniques.
Fuel Properties of Prepared Solid Waste

   A combination of 12 laboratory samples
and  adjustments  based on long term  pilot
plant mass balance measurements yielded the
following average  weight distributions for the
shredded and air-classified solid waste  used
during the 24-hr  period.
   Ash-free combustibles
   Moisture
   Inerts
0.516
0.301
0.183
   A   series   of   12   bomb   calorimeter
experiments in the laboratory on dried parallel
samples produced an average  higher heating
value of 6437 Btu/lb. Using an approximate
ultimate  analysis  of  C3QH48OJ9  for  the
combustibles fraction  and converting to  a
lower heating value based on the combustibles
only,  a corresponding average value of 8087
Btu/lb of combustibles is found. This result is
in  very   good  agreement   with  values
determined by applying heat balance relations
to observed pilot plant temperature and flow
measurements.  It  also  correlates well  with
expected values for a cellulose-like material
such  as  approximated  by the  C3QH48Oi9
formulation.

   Compared to more conventional  fuels, the
heating  value of solid  waste  is  in a sense
degraded  by  the  presence  of  inerts  and
moisture in greater-than-normal concentra-
tions. On  the other  hand, the greater mass
flows  of   combustibles  and   water vapor
required to vaporize and heat the water  are
exploited by the gas turbine cycle as  a natural
form of water injection. In addition, fluidized
beds have a demonstrated ability to consume
"low quality" fuels and hence to eliminate any
need for a  drier or other additional fuel pre-
processing.

   With the  preceding  fuel  properties,  the
steady  state   combustor  operating  point
described in Table 1 is typical. Note the  178
percent  excess  air  level associated  with this
operating  condition,  which is  one  of  the
reasons that very high combustion efficiencies
are realized.
                                                                                    II-3-7

-------
Table 1. TYPICAL OPERATING POINT OF THE
LOW PRESSURE CPU-400 PILOT PLANT CON-
FIGURATION
Solid waste fuel flowrate, Ib/min                  36.6
Combustor air flowrate (includes feed line), Ib/min      330
Excess combustor air, %                      178
Solid waste fuel inlet temperature, "F               70
Air inlet temperature, F                      130
Exhaust gas exit temperature, "F                1,500
Total combustor heat release, Btu/min            152,700
Heat loss to ambient (two vessels), Btu/min         5,000
Mole fraction of major exhaust gas contituents:
   Oxygen
   Carbon dioxide
   Water
   Inert mix of nitrogen and argon
0.122
0.063
0.099
0.716
   The  excess   oxygen   ratio  is  strongly
 influenced  by   fuel  moisture  fraction.  If
 moisture fraction  is  increased  to 0.50,  for
 example, and inerts  fraction is reduced  to
 0.131 (as would happen if the added moisture
 was obtained by direct addition of water to the
 previous fuel mixture), then the required solid
 waste flowrate jumps to 64.9 Ib/min, and the
 excess   oxygen   percentage  drops  to  119
 percent. Further increases in moisture fraction
 produce even sharper drops in excess oxygen
 ratio so that it is not possible to  use moisture
 fractions much over 0.60.  An interesting inter-
 pretation of the preceding numbers is that the
 combined  injection of 18.5 Ib/min of water
 and 30  percent  moist solid waste raises the
 consumption of the latter (and hence the total
 heat release) by  27 percent to 46.4 Ib/min.

 FIuidized-Bed Combustor Performance

   With a large, relatively shallow fluidized
 bed  served by one  or  two feed pipes, a small
 but significant portion of the total combustion
 process occurs in the lean phase  or freeboard
 above the dense bed. For a 40 ft2 bed area 2
 feet deep, for example, bed to exhaust temper-
 ature rises of 260 and 160F were observed
with one and two feed  pipes in  operation,
respectively. Since  these  levels of  gas  phase
combustion (afterburning) are quite stable for
well  controlled  solid  waste injection, it  has
been concluded that satisfactory operation can
be obtained in  either case at low pressure
conditions.
   Very high overall combustion  efficiencies
have   been   demonstrated   by   several
approaches. First, careful heat balance calcu-
lations  based on process measurements have
consistently produced apparent heating values
that,   when   compared   to   laboratory
calorimeter results,  produce efficiency values
slightly  in  excess  of  100 percent.  Flow
measurement  errors   and   occasional  high
heating value components in the solid waste
not  found  in  laboratory  samples  are  the
probable  explanations.  A  second  method
based  on a carbon  balance as indicated by
continuous  exhaust gas measurement of CO2,
CO, and hydrocarbons  plus post-test analysis
of free carbon in the separator residue stream,
generated an average efficiency value of 99.76
percent. Finally, there  is no visible exhaust
smoke,  odor, or other common evidence of an
inefficient process.

System Inert Material  Balance

   After installation of the combustor  was
completed it was loaded with a 2-ft "starter"
bed of commercial 16-mesh beach sand having
a bulk density of about 100 lb/ft3. After 18
hours of fluidized operation on solid waste fuel
in the  first six tests,  the  cooled  bed  was
observed to still be 24-in. deep and free from
excessively  large  particles  or  clinkers.  No
material had been removed except through the
elutriation  process. The size distribution was
somewhat larger (i.e., more fines  and more
coarse particles) and the bulk  density  had
dropped to  91 lb/ft3.

   The  bed generated  in this fashion  was
successfully used in  the following 35 hr test.
Again,  no  material  was  removed  except by
elutriation.  Post-test analysis and  inspection
then showed that the bed was 25-in. deep and
had  an average bulk density of 87.5 lb/ft3.
Thus  the  7,300-lb   bed   processed   and
elutriated more than 12,300 Ib of inerts in 35
hours while only growing about 10  Ib. This
natural replenishment of bed material by  the
inert content  of the fuel is  a  phenomenon
peculiar to  solid waste combustion.  Earlier
long duration tests (240 hours) on a 2.2 ft 2 bed
H-3-8

-------
coupled with computer simulation of transient
size distribution histories had shown that the
steady state bed  could  be expected to have
quite acceptable properties. Consequently, the
bed  maintenance  and  elutriation  control
problems appear to be  grossly simplified.

   In the  35-hour test, 72.0 percent of the
elutriated material (i.e., 8890 Ib) was collected
in the settling chamber of the  alumina/sand
separator vessel.  This inert granular  residue
has  considerable promise  as  a construction
material. Another 25.4 percent (3140  Ib) was
collected by the first stage inertial separator in
the form of fine  fly ash. The remaining 2.6
percent, 320 Ib of very fine fly ash, left  with
exhaust gases. Most of this  latter material will
be  collected  by  the second   stage  inertial
separator in subsequent tests.

   Gas samples were withdrawn from the first
stage  inertial  separator   outlet at  3-hour
intervals for laboratory analysis of particulate
loading. Early samples showed a total loading
of 0.057 gr/scf and a loading of 0.025 gr/scf
for particles greater than 5 f*m in size; both
values are indicative of satisfactory first stage
performance.  Later markedly  higher values
confirmed post-test findings that the  hopper
beneath the 24 active tubes has failed to drain
properly and therefore progressively plugged
the  ash discharge  sections of some tubes,
rendering  them  ineffective. Twin  vibrators
have been installed to solve this problem in
future tests.
Exhaust Gas Composition

   A set of instruments has been installed for
on-line concentration  measurements of six
specific constituents of the exhaust  gas. A
seventh gas chromatograph instrument for the
on-line measurement of HC1 is under develop-
ment and will be added  to replace  current
laboratory  titration  procedures.   The  gas
sampling,   conditioning,   and  distribution
systems are integrated with the instruments
and  analog  recorders  in  a   mobile,  rack
mounted,  complex. The six current  instru-
ments, measuring principles, and associated
gases are:
 1. Beckman 715, Polarography, Oxygen;

 2. Beckman  315B   IR,  Infrared,  Carbon
    dioxide;

 3. Beckman 400, Flame lonization Detec-
    tion, Hydrocarbons;

 4. Beckman  315B   IR,  Infrared,  Carbon
    monoxide;

 5. Theta  LS-800AS,   Electrochemical
    reaction, Sulfur dioxide;  and

 6. Theta  LS-800AN,   Electrochemical
    reaction, Nitrogen oxides.

   The  sampling probe leads into a  stainless
steel  sampling  train that includes  particle
removal elements*," gas  cooling  and drying,
controlled  reheating  to  100F,  a  common
manifold, and flow control elements  for each
instrument. Various  calibration  and  zero
adjustment  methods  have  also  been
incorporated.

   Measurements from the recent pilot plant
test are presented in Table 2 together with
pertinent emission standards. All instrument
records  were relatively steady and free from
apparent anomalies considering the potential
heterogenous composition of the fuel  form.
The 12 discrete  laboratory  samples  (2-hr
intervals)   and  subsequent   HC1   analyses
showed  more variance from a low value of 40.5
ppm to a high of 122.4 ppm.

   The first  two  entries of Table 2 correlate
well with   analytical  predictions such  as
contained in Table 1.  The very low values for
the second two, coupled with the average 45.8
ppm of  free carbon (weight basis), are indica-
tive of the high combustion efficiency. Sulfur
dioxide  presently appears not  to present  a
pollution control problem, probably owing to
the relatively low, sulfur content of municipal
solid  waste and the  apparent  capture  of
existing SO 2 by the bed material.

   Measured NOx levels are  much closer to
proposed  standards  and  somewhat higher
than originally expected. As confirmed  by
other investigators,  much of  the   nitrogen
                                                                                    II-3-9

-------
Table 2. EXHAUST GAS CONSTITUENTS AND
REFERENCE DATA
Constituent
Oxygen
Carbon dioxide
Hydrocarbons
Carbon monoxide
Sulfur dioxide
Nitrogen oxides
Hydrogen chloride
Proposed or expected
EPA emission standards
(maximum)
800 ppm
2000 ppm
300 ppm
NAa
NA
Mole fractions
measured during
24-hour CPU-400
pilot plant test
11.7%
6.8 %
14.4 ppm
4 1.7 ppm
20 ppm
162 ppm
90.3 ppm
a Not available.
 emitted in these compounds derives from the
 fuel rather than the combustion  air.  Other
 experience,   however,  indicates  that  the
 concentrations may  be expected to  drop as
 pressure is increased.

   Depending on emission standards yet to be
 established, the measured levels of HC1 pose a
 potential control  requirement. There appears
 to be a number of effective remedial measures
 that rely  upon fluidized-bed  characteristics,
 however, if forthcoming pressurized combus-
 tor tests should firmly establish a requirement.
 CONCLUSIONS

   Shredding and air classification operations
 on municipal solid waste produce a fuel form
 having very satisfactory physical and chemical
 properties   for   energy  recovery  through
 combustion in a fluidized-bed reactor. A solid
 waste  handling  subsystem  with   reliable
 components  has been  developed  and
 extensively operated.

   The fluidized-bed combustor has fulfilled
its promise as  a highly efficient,  easily fed,
readily controlled reactor of simple design and
capable of utilizing low quality  fuels.
   Grade 16 silica beach sand is an acceptable
 starter  bed material.  The inert  content  of
 typical  solid waste  provides  a natural bed
 make-up material that leads to a satisfactory
 steady state bed composition. As a result, the
 need for elaborate bed maintenance or anti-
 elutriation  devices  is  minimized  if not
 eliminated.

   Test results to date  show no problem with
 fluidized-bed residue buildup. Relatively large
 metal and inert particles entering the active
 bed  experience gradual oxidation or attrition
 to typical bed size particles and eventually are
 elutriated   to   be  collected   by  particle
 separators. No bed material agglomeration is
 experienced  in  the  1270  to  1450 F  bed
 temperature range.

   Gas phase combustion above the fluidized
 bed  has been reduced to acceptable  values
 without  resorting to  undesirable  remedies
 such as extensive  internal bed  structures,
 numerous fuel feed points, deeper bed,  multi-
 stage combustors, etc.

   Earlier combustor freeboard  and exhaust
 system deposit problems due to the aluminum
 content of solid waste appear to have been
 solved.

   Performance of the first stage inertial sepa-
rator has been very promising and is expected
to further improve at high pressure conditions.
Reliable  handling  and  transportation  of
removed hot, fine fly ash has posed develop-
ment problems.

   Exhaust gas sampling instruments indicate
 that the pollution control effort required for
 the  CPU-400 process can be expected to be
 minimal.

   No evidence of serious hot gas subsystem
 material corrosion or  erosion problems has
 been found.
II-3-10

-------
a
w
                                                                                       SOLID WASTE PROCESSING
               SHREDDED
                WASTE
               STORAGE
                                                                                ALUMINA/SAND
                                                                                SEPARATOR
                                                                                        FIRST
                                                                                       INERTIAL
                                                                                      SEPARATOR
                                                   GENERATOR


                                                     EXHAUST DUCTS
                 CONTROL ROOM
              BAG HOUSE
              FILTER
                                                                                                               SECOND
                                                                                                              INERTIAL
                                                                                                             SEPARATOR
AIR INLET
                                                Figure 1.  CPU-400 pilot plant pictorial.

-------
                                                           SHREDDER CONTROLS
                                                              wassail
                                                              CONVEYOR
                                           CONVEYOR
                                         H CONTROLS
         Figure 2.  Photographic views of the solid waste handling subsystem.
II-3-12

-------
H
oo
                                   PACKER TRUCK
                                                   DUST KOP
                            LOADER
           EIDAL SHREDDER WITH
                ANTI-BALLISTIC HOOD \
                    BLOWER
                 AIR CLASSIFIER
               SOLID WASTE SYSTEM CONTROL PANEL
                                                                    CYCLONE
                                                                             WEIGHING CONVEYOR.
                                                      SHREDDED SOLID WASTE
                                                         STORAGE TANK
                                                           AIRLOCK FEEDER
ATLAS CONVEYOR
INSTRUMENTATION PANEL
                                 Figure 3.  Pilot plant solid waste handling subsystem schematic.

-------
                 Sweep Bucket Chain
              Shredded,  Air Classified Solid  Waste Fuel
II-3-14
        Figure 4.  Photographic views of solid waste storage tank interior.

-------
Figure 5.  Airlock feeder valve.
                                                       II-3-15

-------
                           Figure 6.  Inertial separator schematic.
II-3-16

-------
                                 Inertia!
                                 Separator
Alumina/Sand
 Separator
                                                                       Gas Turbine
                                                                      (uninstailed)
    Figure 7. View of solid waste combustor and gas preparation subsystem.

-------
                                 4. FLUIDIZED-BED  COMBUSTORS
                           USED IN HTGR FUEL REPROCESSING

             B. J.  BAXTER, L. H. BROOKS, A. E. BUTTON,
                  M. E.  SPAETH, AND R. D. ZIMMERMAN

                       Gulf General Atomic Company
ABSTRACT

   High-temperature gas-cooled reactors (HTGR)  utilize graphite-base fuels.  Fluidized-bed
combustors are being employed successfully in the experimental reprocessing of these fuels. This
paper presents a general discussion of the reprocessing method and describes the two types of
fluidized beds being used.
INTRODUCTION

   The high-temperature gas-cooled reactor
(HTGR), as developed at Gulf General Atom-
ic, is a helium-cooled, graphite-moderated
reactor. The fuel in an HTGR consists of fis-
sile microsphere particles containing U-235,
recycle microsphere particles containing U-
233, and thorium fertile particles contained in
a hexagonal fuel element, shown in Figure 1.
The HTGR fuel recycle operation consists of
shipping spent fuel to recycle facility, repro-
cessing the  fuel to recover the U-233 and U-
235, refabricating the U-233 and U-235 into
recycle fuel, shipping the  refabricated fuel
from the recycle facility to  the  reactor, and
ultimately storing the radioactive fission pro-
duct  wastes.
  The fuel reprocessing sequence starts with
the head-end operation shown in Figure 2, in
which the fuel in the HTGR fuel element is
separated from the graphite body by crushing
and   fluidized-bed  burning.   Subsequent
head-end   operations  separate   particles
containing U-235 from ash containing U-233,
thorium, and  fission products. The  metal
oxide ash is dissolved to create a solution of
uranium, thorium, and fission products; the
silicon-carbide-coated  U-235  is the residue.
The U-235 is separated mechanically, and the
uranium  and  thorium  are   recovered
individually from the fission products. The
recovered U-233 and thorium are stored for
reuse  as  fuel. .The radioactive wastes are
disposed of in appropriate storage facilities.
                                       II-4-1

-------
 HEAD-END REPROCESSING

    Head-end reprocessing for HTGR fuel con-
 sists of a crush-burn-leach process.1  Fuel ele-
 ment size reduction is the first step in  head-
 end reprocessing. Two major criteria govern
 this step:  (1) the fuel must be crushed to  a
 suitable  size  for maintaining  fluidization
 quality in the  fluidized-bed burners, and (2)
 the crushing system must minimize fuel  parti-
 cle breakage to prevent undesirable crossover
 of fissile and fertile product uranium.2

    A three-stage  crushing  system  has  been
 adopted for the reprocessing plant, based on
 the experimental  testing  of commercially
 available equipment using full-sized fuel ele-
 ments. This crushing system is presently  being
 tested.

    Primary reduction is done in a large, over-
 head eccentric jaw crusher; secondary reduc-
 tion in  a  small, overhead  eccentric jaw
 crusher; and tertiary crushing in a double-roll
 crusher. The  tertiary crusher  product,
 nominally minus 3/16 in., is  pneumatically
 conveyed to the  fluidized-bed burner  feed
 hoppers.

    Crushed fuel is fed to the top or base of  a
 continuous, exothermic fluidized-bed burner,
 shown in Figure  3, by an auger feeder. (The
 term exothermic is used to describe the burner
 that generates sufficient  heat to  maintain
 operating temperature.) The feed rate is  auto-
 matically controlled  by the off-gas carbon-
 monoxide  concentration,  which  has  been
 shown to be proportional to the graphite sur-
 face area exposed in the bed.

   Both  the crushed graphite and the silicon-
 carbide-coated  fissile  particles serve as the
 fluidizing  media.  The  heat  generated  by
 burning  is removed by forced-air cooling in  a
clamshell jacket surrounding the burner and
an off-gas heat  exchanger. The  fluidizing gas
fed to the  burner is  oxygen with  a  small
amount of inert gas (i.e., CO2, N2),  and the
flow is automatically  controlled to  maintain
the  bed  temperature. The burner  product
removal rate is automatically controlled by the
 bed pressure drop, which is proportional to
 the bed weight.

   The burner off-gas with entrained fines is
 passed through a  cyclone separator  and a
 sintered metal  filter for fines removal, before
 being cooled and proceeding to off-gas treat-
 ment.  Off-gas  treatment removes the  fission
 products including noble gases before release
 to the  environment. Fines are presently being
 recycled to the burner in the  experimental
 program.

   If the exothermic burner is  operated with
 top feed and no fines recycle,  the elutriated
 fines from the burner (both TRISO/TRISO
 and TRISO/BISO flowsheets) are also  added
 to the  feed stream for the last burning step.
 This mixture now constitutes the feed material
 for the  endothermic (requiring heat  input
 from a furnace to maintain operating temper-
 ature)  fluidized-bed burner.

   The exothermic burner product is fed to a
 batch-operated  endothermic   fluidized-bed
 burner,  shown in  Figure  4,  where  the
 remaining graphite is burned and the thorium
 and uranium-oxide kernels are exposed. The
 silicon-carbide-coated fissile particles serve as
 the inert fluidizing media. The feed stream to
 the endothermic  burner  will  not  sustain
 exothermic burning to the low carbon level
 required in the subsequent processing steps.
 The  burning   in the  endothermic  burner,
 therefore, proceeds from  exothermic condi-
 tions, with the heat removed from a clamshell
 surrounding the burner, to endothermic con-
 ditions,  with  heat  supplied  by  resistance
 heaters located in the clamshell. The off-gas
 from the burner is treated in the same manner
 as that from the exothermic burner. The pro-
 duct from the endothermic burner is pneuma-
 tically  conveyed to the leaching system.

   The thorium and uranium oxides are dis-
 solved in acid thorex [13M HNO3 - 0.05M HF
 - 0.01M  Al  (NO3)3]  in  a  steam-jacketed
 cylindrical vessel with gas sparge mixing. This
leaching vessel is run  as  a refluxing,  batch
leacher.
11-4-2

-------
   The insoluble silicon-carbide-coated fissile
particles and the unburned carbon  must be
separated from the mother liquor before the
solution can be fed to the solvent extraction
system for uranium purification and thorium
recovery. A centrifugal separator receives the
entire slurry from the leacher. Solids retained
on the centrifuge screen are washed with fresh
leach solution, which becomes the leach solu-
tion for the  next batch of solids  from the
endothermic burner. The washed solids from
the leacher are then air-dried and transferred
to a screen classifier, where the fissile particles
are separated from the silicon-carbide hulls.
The waste solids are processed as wastes, and
the  fissile  particles  are  stored  for  later
processing, by a method similar to that for the
fertile  particles, to recover and  purify the
uranium.

   The  clarified leach solution is evaporated
and steam-stripped to an acid-deficient condi-
tion for use as  feed  to an  acid  thorex3'4
extraction process. The  acid thorex solvent
extraction   process   is   used   for   the
decontamination  and purification of the U-
233 and thorium and for the separation of the
U-233  and thorium from each other.

FLUIDIZED-BED BURNERS

Exothermic Fluidized-Bed Burners

   Figure 3  depicts the exothermic fluidized-
bed burner  presently   being  used  in the
experimental program.  Exothermic  burners
with both 4-in. and 8-in. diameters are being
used;  construction  of  a  larger burner  is
planned for early next year. The large burner
will become  the  full-sized commercial  plant
test unit.  Preliminary  nuclear  criticality
calculations have shown  that this burner can
be about 16 inches in diameter.

Operability

   The  exothermic  burners  have  been
operated  on  a routine basis for the last 18
months. Startup  is initiated  by heating  a
charge of coke (1200 g for the 8-in. and 400  g
for the 4-in.) to ignition temperature (700C)
with a carbon monoxide-oxygen gas mixture.
This gas mixture is introduced into the burner
with a standard cutting torch that is ignited by
two  spark  plugs. After the  coke is ignited,
fluidizing oxygen and the graphite-base feed
are introduced.

   The  steady-state bed  properties  of the
exothermic burners are listed in Table 1; the
bed contains 2 to 5 percent burnable carbon.

     Table 1. AVERAGE BED PROPERTIES
        FOR EXOTHERMIC BURNERS9
Size
-3/16 in.
-1/8 in.
-869 urn
-550nm
-420 jim
-375 jxm
-250 Mm
-125 /^m
-74|^m
-44jxm
-, 1/8 in.
+ 869 ^m
+ 550nm
+ 420 ^m
+ 375pm
+ 250 (^m
+ 125^01
+ 74 /urn
+ 44 ^m

Wt%of
total sample
0
0.2
67.9
4.9
0.2
1.0
5.5
4.3
2.8
13.2
 a Burnable Carbon: 3.0%
   Average Particle Size: 590 Mm

 Steady-state operation is achieved for experi-
 mental purposes in about 4 hours by adding
 the  estimated  steady-state  bed composition
 directly  to the burner  immediately - after
 startup. Product removed after 4 hours is also
 near steady-state and exhibits about the same
 properties  as  the bed. The present  feeding
 method is a variable-speed auger controlled
 automatically  by  the  carbon-monoxide
 concentration in the off-gas. The nominal off-
 gas  concentrations are 1 to 3 percent carbon
 monoxide, 0 to 6 percent oxygen, and 60 to 95
 percent carbon dioxide. This composition is a
 function of the inert fluidizing gas diluent (i.e.,
 air  versus CO2). Product removal  rate  is
 automatically  controlled using  the  pressure
 drop across the bed and regulating a variable-
 speed  drive motor on  the product removal
 auger.  Bed  temperature   is  automatically
 controlledrby regulating the fluidizing oxygen
                                                                                    II-4-3

-------
 supply to the bed. Temperature profiles, off-
 gas compositions, bed and filter pressure
 drops, and mass  flows of important streams
 are monitored continuously.
    Fluidization quality has been  difficult to
 define. Normal operation is with a well-mixed
 bed that occasionally slugs. Distributor plates
 are not presently being used but will be fully
 tested in the near future. At present, operation
 is with a cone base  and a ball check valve.
 Preliminary tests of perforated plates, bubble
 caps,  and  sintered  metal screens  were  all
 successful  to  some  extent,  and  the  beds
 appeared   to  maintain  good  fluidization
 quality.


 Feed

   The feed to the burners is presently defined
 as  minus  3/16-in. graphite-based material.
 This feed  size was established  by gradually
 increasing the  size from minus 1/16 inches to
 minus 1/4 inches. Poor fluidization occurred
 with minus 1/4-in. feed,  as witnessed by local
 "hot spots" in the bed. Returning to minus
 3/16-in. feed eliminated this problem.  Table 2
 shows the average properties of the exothermic
 burner feed.
      Table2. AVERAGE PROPERTIES OF
         EXOTHERMIC BURNER FEED3
Size
-3/16 in. + 869f*m
-869 J^m + 550 /urn
-550 ^m + 420 (xm
-420 |i*m + 375 urn
-375 /xm + 250 fxm
-250 j^m
Wt%of
total sample
64.6
25.0
1.5
1.5
1.5
5.9
% burnable
carbon in
fraction
100
21
78
97
98
98
aTop Density: 1.25g/cm3
 Bulk Density: 1.08g/cm3
 Angle of Repose: 35
 Average Burnable Carbon: 80%
 Average Particle Size: 854 jjtm
Heat transfer

   The heat transfer  problem encountered is
somewhat  different from  that  occurring in
most fluidized-bed work; because of nuclear
criticality considerations, the cooling medium
is limited  to  air. Figure  5 shows the  heat
balance  for   a   typical   8-in.  diameter
exothermic burner  run.  The over-all  heat
transfer coefficients  for  the off-gas  heat
exchanger  and  clamshell  cooler  are  also
shown.

Fines recycle

   One  of the  major problem  areas  in
exothermic burner operation is the burning of
fines that  have elutriated  from  the  burner.
Since burning efficiency and radioactive hot-
cell  constraints play  a major  role  in  the
process, work has focused on burning the fines
by recycle to the bed rather than in a separate
fines burnup cell.

   Fines carryover has been about 22  percent
of the burn  rate when  the furnace  has
operated at normal conditions; i.e., at a burn
rate   of approximately 200  g/min  and  a
superficial  fluidization velocity of 3 to 4 ft/sec.
Of these fines, about 98 percent is collected by
a cyclone and 2 percent by a  filter  chamber.
The  elutriated fines are described  in Figure 6.

   The present operating mode  is to recycle
the fines by blending them  with the graphite-
base feed stream.  This composite is fed to the
bottom of the fluized bed,  and the  fines
appear to burn successfully when steady-state
is  achieved.

   The nominal burn rate for the 4-in. and 8-
in.   exothermic   burners   is  about   33  g
carbon/hr-ft2, which  corresponds to  50 and
200 g carbon/min, respectively. A burn rate of
50 g carbon/hr-ft 2 (corresponding to  75 and
300 g carbon/min, respectively) is planned for
the 4-in. and 8-in. exothermic burners. Burn
rates of 125 and 475 g carbon/min for the 4-in.
and  8-in. burners (about 84 g carbon/hr-ft2)
have   been  achieved  for  short   periods.
Operation  at these high  burn  rates- is  not
II-4-4

-------
possible   for   long  periods   because  of
limitations with existing equipment.  Future
studies will be aimed at defining these values
over long-term run conditions.

Scale-up

   The  design considerations for the larger
exothermic fluidized-bed burner include both
a  theoretical approach and  scale-up factors
from the 4-in. and 8-in. burners. To date, the
problems encountered for scale-up have been
in defining both a suitable transport  disen-
gaging height and exact heat transfer values.
Although the  theoretical  and  experimental
predictions of heat transfer values in the area
of the bed-wall-clamshell are in agreement, it
is difficult to define the proper  heat transfer
coefficient in the transport disengaging height
in which  the overall heat transfer coefficient
rapidly decreases  with  reactor height. These
values will be determined experimentally in
future  experiments  in which   "sectioned"
clamshell coolers  will be utilized.       *

Endothermic Fluidized-Bed Burners

   Figure 4   depicts  the  4-in.  diameter
endothermic  fluidized-bed burner presently
being used in the expermental program. It is
planned  to  convert  the  8-in.  diameter
exothermic burner to an endothermic burner
by adding resistance heaters  to the clamshell
interior  and  moving the  filter  chamber to
directly above the burner for future scale-up
testing.

Operability

   The endothermic fluidized-bed burner has
been operated as a batch burner on a routine
basis for  the last  18 months  and is presently
being automated.  The automation consists of
the furnace temperature control loop and a
burner  control  system. The  burner  control
system is a repeating unit  of automatic tem-
perature control (by regulating the fluidizing
oxygen  flow)  and a  series  of  programmed
events. The programmed events control fluid-
izing gas flow,  batch  product  dump  valve,
batch  pneumatic  feeder, and  refluidizing
oxygen flow until the automatic temperature
control loop takes over and the sequence starts
again. This control cycle is repeated automati-
cally and  provides a batch-continuous oper-
ation. Although  the endothermic burner is
similar to the exothermic burner, the fines are
burned by containing them within the burner.
The low fluidizing velocities used during the
endothermic burn stage allow burning the bed
to less than 1 percent carbon.

Feed

   The feed to the endothermic burner is the
product from the  exothermic burner. The size
distribution of this feed  is highly variable,
depending  on the flowsheet being processed,
but at all  times  it is easily fluidized. The
average particle size of the feed varies from
200 to 400  jum  for  the various flowsheet
varieties.

Heat transfer

   The burn rates  achieved  in  the  47in.
endothermic burner are about one-half that of
the 4-in. exothermic burner, because the bulk
of the heat transfer  occurs  in the transport
disengaging height. Also, the operating period
in the endothermic stage of burning is a slow
burning  process. Burn  rates   of  25  to
35 g carbon/hr-ft2 are achieved during-the
exothermic   burn   period   and   5   to
10 g carbon/hr-ft2 in  the  endothermic burn
period. An average burn rate of about 20 to
25 g carbon/hr-ft 2 or 20 to 25 g carbon/min is
achieved during batch-continuous operation.

REFERENCES

1. Nicholson,  E. L.,  et  al.  Burn-Leach
   Processes  for Graphite-Base Reactor Fuels
   Containing Carbon-Coated or Oxide Parti-
   cles. Oak Ridge National Laboratory. U. S.
   Atomic  Energy Commission.  Oak Ridge,
   Tenn.  Report Number ORNL-TM-1096.
   1965.

2. Steward, H. B., et al. Utilization of the
   Thorium    Cycle   in   the   HTGR.
   In:  Proceedings, 4th Geneva Conference
   on Peaceful Uses of Atomic Energy. 1971.
                                                                                    II-4-5

-------
 3. Blanco, R. E.,  L. M. Ferris,  and D. E.     4. Blanco, R. E., L. M. Ferris, C. D. Watson,
   Ferguson. Aqueous Processing of Thorium       and R. H. Rainey.  Aqueous Processing of
   Fuels. Oak Ridge National Laboratory. U.       Thorium Fuels  -  Part II.  Oak Ridge
   S. Atomic Energy Commission. Oak Ridge,       National Laboratory. U. S. Atomic Energy
   Tenn. Report Number ORNL-3219. 1962.       Commission.  Oak  Ridge,  Tenn. Report
                                               Number ORNL-3418. 1963.
II-4-6

-------
                                                _E=\
31.22 in.
      196 lb=89 kg GRAPHITE
      235 Ib =116 kg TOTAL
                                                   CROSS SECTION
                                                                        BURNABLE
                                                                        POISON ROD
                                                                        COOLANT
                                                                        CHANNEL
FUEL ROD
STACK
                                                                        HELIUM
                                                                        FLOW
                       Figure 1. HTGR fuel element.
                                                                                II-4-7

-------
oo
                                                   OFF-GAS
OFF-GAS
     SECONDARY
     CRUSHER
     (JAW)
    TERTIARY
    CRUSHER
    (DOUBLE ROLL)
       SECONDARY
       BURNER
       (ENDO)
                               f    I  COOLANT OUT
                                                                                  SCREEN
                                                                      DISSOLVER   SEPARATOR
  ACID

TO LIQUID-
LIQUID
EXTRACTION
                                         FISSILE PARTICLES


                                         INSOLUBLE WASTE
                                    Figure 2. Head-end reprocessing simplified flow diagram.

-------
            CYCLONE
HEAT
EXCHANGER
PLATES
   COOLING
   AIR, IN
   COOLING
   AIR, OUT"

THERMOWELL
   BALL"
                            02 AND N2
                                                        H* OFF-GAS
                                                            16 FILTERS
                                             SOLIDS LEVEL INDICATORS
AUGER
      PRODUCT REMOVAL
        Figure 3. Exothermic fluidized-bed burner.
                                                                         II-4-9

-------
                          FILTERS
                   COOLING AIR
                   HEAT EXCHANGER
                     FLUIDIZED BED
                              02 AND
                                             =:	v OFF-GAS
                                                       FEED HOPPER
                                                                N2-PULSE
                                                                PNEUMATIC FEEDER
                                                          MOVABLE ELECTRIC
                                                          FURNACE
                                                           INTERMEDIATE SINGLE
                                                           BATCH HOPPER
                                                                    PNEUMATIC
                                                                    TRANSPORT
                              Figure 4. Endothermic fluid-bed burner.
II-4-10

-------
                            SENSIBLE HEAT
                            DIFFERENCE IN
                            PRODUCTS AND
                            REACTANTS AND INERTS
                            16 kcal/min  (1.0
HEAT OUT FROM
HEAT EXCHANGER SKIN
15 kcal/min
(0.9%)
     HEAT OUT FROM
     SPOOL PIECE SKIN
     125 kcal/min
     (7.7%)
      HEAT OUT FROM
      CLAMSHELL SKIN
      81 kcal/min
      5.0%
          RADIATION
          FROM CONE
          181 kcal/min
          (11.1%)
Ahr
 V
                   A H OF HEAT EXCHANGER
                    COOLING AIR
                    150 kcal/min
                   ' (9.2%)
                    U=lBtu/hr-ft2-F
, Ahr=1630 kcal/min
 (83% ACCOUNTED FOR)
                                                       AH OF CLAMSHELL
                                                       COOLING AIR
                                                       780 kcal/min
                    U=20Btu/hr-ft2-F
     Figure 5.  Heat balance and heat transfer coefficients for a typical 8-in.
     exothermic burner run.
                                                                               II-4-11

-------
                           44g/min
          EXO BURNER
          200 g/min
          BURN RATE
          -250 ji+125 )>
          125 ji -f74|i
          -74jj+44}i
          -44 ;i



1
CYCLONE

FILTER
CHAMBER
     43 g/min
     95% CARBON
 wt%
 0.25
 0.25
 4.4
95.1
    1 g/min
    96% CARBON
 0.0
 0.0
 4.0
96.0
                                 Figure 6.  Elutriated fines description.
II-4-12

-------
                                5. STUDIES  ON THE COMBUSTION
                               OF NATURAL GAS IN A FLUID  BED

                    W. E. COLE AND R. H. ESSENHIGH
                        Pennsylvania State University
ABSTRACT

   Natural gas has been burned with air in a fluid bed  of 1.4 ft2 cross-sectional area, using
expanded alumina of 14 to 16 ASTM mesh. The gas/air mixture  is supplied premixed.  Initial
problems of operation, now solved, concerned rapid ignition, and uniform distribution. Light-off
initially required two hours or more before combustion was uniform throughout the fluid bed. This
period is now consistently 5 to 10 minutes. With very uniform gas  distribution, obtained with a
distributor  of novel design, fluid  bed depths were reduced from 6 to 1 inch for complete
combustion, even  up to superficial (hot) flow velocities of 12 ft/sec. Combustion intensities at
2000F and 100 percent excess air were in the region of 106 Btu/ft3-hr based upon bed volume.
Experiments were  carried out with excess air ranging from 60 to about 150 percent, with extinction
at the higher value. Gas rates ranged from 6 to 11 ft3/min. Bed temperatures ranged from 1700 to
2400 F, rising with fuel/air ratio. Air rich extinction boundaries were mapped over a range of
fuel/air ratios. Bed temperatures, gas analysis, and pressure drop on a vertical axis through the
bed have also been measured.  Superadiabatic temperatures and a lowering of the  lean flam-
mability limit have been observed. These two observations  are explained qualitatively by the pre-
heat effect of the fluid-bed particles. Reaction rates are significantly faster than for free-burning
gas in a premixed flame at comparable temperatures and  gas concentrations. The  effect is
attributed to the bed particles. Data on the physical behavior of the bed, with good theoretical
agreements, are also given.
INTRODUCTION
   This  paper describes  experiments on the
combustion of natural gas in a fluid bed to
investigate problems of ignition, even distri-
bution,  and  combustion speed.  Combustion
applications   of  fluid-bed  technology  can
include  incineration, steam raising, possibly
heating of crucibles, billets, etc. However, use
of gas in such applications can present prob-
lems.

   Fluid-bed  combustion of natural gas is,
according to  conventional  belief, beset with
difficulties by comparison with combustion of
oil or coal. Reasons generally cited are:  (1)
excessive initial heat up time of the bed, up to
24 hours in -the case of very large units;  (2)
inefficient (incomplete) combustion in the bed,
leading  to (3) excessive freeboard (overbed)
temperatures from final burnup; (4) high bed
temperatures, said to be necessary compared
with liquid or solid fuels to allow an adequate
margin  of  safety  above   the   commonly
accepted ignition temperatures for gas; resul-
ting in (5) excessive thermal stress on heat
exchangers  included to improve the overall
thermal efficiency.
                                        II-5-1

-------
   The most significant problem  is combus-
tion efficiency, once light-up has  been com-
pleted. In large units it is customary to avoid
premixing the gas and air, because  of the
explosion  hazard; so  the  gas  is  supplied
directly to the bed through one or more supply
ports. A considerable fraction of the gas may
then bubble to the surface (2) and, in burning
overbed,  generate the  excessive  freeboard
temperatures  (3) that can overstress  heat
exchangers (5). Such bubbling behavior paral-
lels  remarkably the behavior of coal volatiles,
if a high volatile  coal is fed  too quickly into a
fluid bed at too  restricted  a point.
    As explanation of the incomplete combus-
 tion inside the bed, the bubbling effect sug-
 gests  immediately  that  it would be due to
 inadequate mixing before the gas breaks the
 surface. This view  is substantiated by esti-
 mates of transit time through the bed  which
 considerably  exceed the  expected reaction
 times. Nowhere, however, were we able to find
 any direct substantiation of this expectation in
 any published literature on gas  combustion in
 fluid beds. Indeed, information on this topic is
 conspicuous by its absence. There were a few
 references  available indicating that gas could
 be burned  in fluid beds, but not very satisfac-
 torily; and nowhere was there  any source of
 quantitative data, so far as we could establish,
 that could serve as any basis for engineering
 design involving commercial use of gas in fluid
 beds. Furthermore,  although our findings as
 reported here  have  now  substantiated  the
 expectation of very fast reaction in the bed
 once the gas and air have mixed, at the  outset
 of the  project there were some few indications
 that slow mixing might not be totally responsi-
 ble for the incomplete combustion in the bed.
The argument was based on the usual behav-
ior in the bed immediately after light-off. The
bed  acted as a flameholder with all reaction
above it and no reaction whatever inside  it
until the whole bed had heated up to a definite
temperature. Clearly, the cold particles were
providing thermal quenching and/or  chain
termination. Equally clearly,  once hot, the
particles could be expected to reverse  their
action and become sources for thermal and/or
chain initiation of reaction. However, this still
did not rule out the  possibility that the pres-
ence of very large surface areas,  even of hot
solids, in the middle of the  reaction  zone
might so significantly alter the reaction mech-
anism that  the  reaction rate  could  still be
appreciably   hindered  or  accelerated.  Our
findings do, in fact, suggest  that the reaction
rate  may be increased; there are also indica-
tions of some  other interesting  aspects of
behavior, such as wider combustion limits and
super adiabatic temperatures.

   Specifically,  however, our  starting  point
was the problem of even distribution and com-
bustion. In the process of investigating this, in
conjunction with developing an alternative to
the conventional distributor  plate, we found  a
means of reducing the light-off time. The unit
thus developed, having very even gas distribu-
tion and good fluidization, was then suitable
for more detailed  measurements  of gas  tem-
peratures and analyses in the  bed,  leading to
the results indicated above.

   What we have  to report  therefore  are po-
tentially valuable data for engineering design
and  data of more fundamental  significance
but, withall, not without relevance to design.
This is, we believe, the first public report on
quantitative behavior of gas burning in a fluid
bed.
n-s-2

-------
PRELIMINARY EXPERIMENTS
   The  first experiments carried out, sum-
marized here, led to a  new design concept
described below. The unit used for these initial
experiments utilized  a  square, 3/8-in.  steel
distribution plate, of 13-in.  side, carrying 81
perforated   studs  (a  9x9  array  on  1.5-in.
centers). The walls were uncooled, constructed
of  2.5-in.  series  superduty  (Rockspar) fire
brick. Air and gas were supplied separately to
mix in the bed. Air was supplied via an air box
under the distribution plate, flowing into the
bed through 64  of the distributor studs. Gas
was supplied through the remaining 17 studs
which were connected by a progressively bifur-
cating line  from a 15 psi supply  point. The
perforations in the distributor studs were 1/8-
in.  diameter, drilled horizontally  from the
outside  to  meet  hollow centers.  The  bed
material mostly  used was  Type  8-F  blown
alumina of 6 to  12-in.  depth in the initial
experiments. Instrumentation  included: gas
and  air meters,  wall  thermocouples,  wall
pressure taps,  suction  pyrometer  for gas
temperatures, and sheathed thermocouples
for bed temperatures.

    With the unit  as   described, the  bed
fluidized satisfactorily in cold flow; but light-
off, with combustion predominately  in the
bed,  was never initially achieved. The gas all
burned above the bed with the bed acting as a
flameholder. It was clear that mixing was the
problem because the flames  stabilizing on the
top of the  bed appeared in a set of rings of
flame whose pattern  was determined  by the
layout of the gas supply studs. Various means
of overcoming this were tried. Unsuccessful or
partially successful means included: other bed
materials including a fine sand that cascaded
through the stud  perforations  into the wind
box;  and   rotary  stirring  of  the  bed by
mechanical and  pneumatic means.  (For the
latter, a set of four supplementary air pipes
were  lowered into the bed to produce rotation
of  the  bed by horizontal jets  aimed at an
imaginary circle. This was partly  successful.)
   Success was finally achieved by resting the
fluid bed on  an  underlying bed  of limestone
rocks (Type  2B) 4.5 inches deep. This was
found to be an excellent mixer and distributor,
and  with this arrangement light-off became
possible with reaction ultimately drawn back
into the bed. (It was at this point that most of
the instrumentation was added.)

   Initially,  however,  light-off was still  the
excessively long process generally  claimed,
taking anything  up to 2 hours (for a 6 to 12-
in. bed). Furthermore, as the flame was drawn
into the bed it exploded in a random sequence
of strikebacks followed by blow-off (i.e., with
the flame oscillating between the top and bot-
tom of the bed). In some instances, the explo-
sions were violent enough to displace some of
the wall bricks even though they were mor-
tared into position and partly held there by
steel angle frames. Once hot, the bed burned
evenly without explosion.

   Light-off, however, was ultimately reduced
to 10 minutes or less (3 to 4 minutes is about
the shortest time so far). It was clear that the
flame would not  strike back into the bed until
the  whole  of  the   bed  had  reached   a
temperature that would  permit it. (The exact
temperature is not known.) The heatup took a
long time  because  the particles were  only
heated at the top of the bed (where the bed top
acted as a flameholder) and were cooled again
by the cold fluidizing air  and  gas as  they
mixed back into the bed. The solution was to
fluidize the bed in progressive stages and, with
correct gas  and air  settings, the process is
virtually  automatic.  The  gas   and  air  in
stoichiometric proportion are initially set  at
about  10  percent of the rate  required  for
incipient fludization in the cold; the gas is lit
to burn over the bed with the bed acting as a
flameholder.  The gas and air rates are then
promptly increased to about half of the cold
fludization requirements. The over-bed flame
heats the top (unfluidized) layers of the bed
and the rising gas and air. When the top bed
particles are  hot enough, reaction can  then
start just  inside the  top of the  bed.  The
consequent local rise in temperature of the
                                                                                    II-5-3

-------
 gases increases their velocity to above fluidiza-
 tion  velocity,  fluidizing  the top  layer  of
 particles. The same process  then operates to
 fluidize  the  next static  layer;  this is  then
 repeated, with smooth, non-explosive progres-
 sion of  the  fluidized  interface  down to the
 bottom of the bed. The gas  and air supplies
 are  then adjusted to the  required operating
 levels, and start-up is complete. There is no
 reason why the  same procedure could not be
 used to  cut  light-off time on a commercial-
 scale bed by an order of magnitude, or more.

   Solution of the even distribution and rapid
 light-off problems were the main tasks under-
 taken in these preliminary experiments. Work
 was then started on combustion behavior in
 the  bed  itself,  with initial  indications  that
 reaction was  substantially slower than it would
 be  in   free-burning  gas   at  the  same
 temperature. However,  problems   due  to
 leakage,  mainly  through  patched  cracks
 produced during the explosive  ignition tests,
 made  measurements erratic. An entirely new
 bed was therefore  constructed for further
 measurements, as described in the equipment
 section.

   Preliminary to the reconstruction, however,
 some cold model tests of air distribution and
 fluidization were carried  out to aid design.

   (1)  The first model used water  as  the
 fluidizing medium, fed into  the bottom of a
 12-in.   high,   3.5-in.  diameter,   Plexiglas
 cylinder, through a 1-in. pipe containing five
 4-in. sections of 3/8-in. copper tubing to serve
 as flow straighteners. An air  line to a hyper-
 dermic syringe  in  the  center  of the  flow
 straighteners served as  a fine bubbler  for
visual flow tracing in the mixing studies. The
fluid bed was simulated by glass beads on a
wire screen, about 1 inch above the 1-in. inlet
pipe. Beads of 3, 4, and 5mm diameter were
used. With water supply up to 6 gal./min., the
behavior of fixed, fluidized, and spouting beds
could be observed.

   (2) The second unit was an air model of a
3-in. diameter feed pipe feeding  into a square
clay pipe, of 13-in. side, or into 6-in.  diameter
Plexiglas  pipe,  again using  glass  beads to
represent the distributor with finer beads or
actual bed material above.

   These studies indicated the importance of
uniformity of the underlying distributor rock,
but provided the information to indicate that a
single  supply   point  of  relatively  narrow
diameter compared with the fluid bed  diam-
eter could  be used as long as the distributor
rocks (or coarse particles) were deep enough.


EQUIPMENT

   The present unit is illustrated in  Figure 1.
It consists  essentially of  a refractory lined
cylinder of castable refractory, cast inside two
oil drums, standing on a welded steel platform
30-in.  high.  This construction  provides  an
outside diameter of 22 in. and a height of 47
in. The shell is wrapped with  540 feet of 3/8-
in. copper tubing for cooling  and monitoring
the wall loss. Walls, floor, and roof are 3-in.
thick, of Hydrecon Tabcast, a  3600F erosion-
resistant  castable refractory.   The  air  is
supplied to the unit from two staged  24-oz.
blowers (360 scfm capacity) through a 3-in.
diameter  pipe  cast  in  the  floor,  and  the
exhaust gases leave through a metal cased flue
liner, 12-in.  diameter and 2 ft long, leading
into a 16-in. diameter stack.

   The bed is blown alumina,  of mesh size -14
+ 16 ASTM, resting on a 6-in.  deep bed of
crushed refractory, of mesh size -3 + 5 ASTM.
At the top of the 3-in.diameter air supply pipe,
the crushed refractory is supported on a 1/16-
in.   perforated   steel   plate.   The   crushed
refractory provides all necessary air  and gas
distribution across the full width of the unit.


   An ignition burner, at a height of 22 inches,
is provided for startup and safety. To observe
the bed, two inspection ports closed with 2-in.
Vycor discs are provided in the top; the top
also contains two holes for insertion of probes.
For access to the bed, the top can be removed
using a 1/2-ton differential chain hoist.
II-5-4

-------
   The air and gas are fed to the bed through
a mixer, after being metered by rotameters
reading  to  280  and  15  scfm  full  scale,
respectively. The mixer is a 2-1/2-in. diameter
(Pyronics) venturi unit, followed by a Tee to
carry the mixture into the bottom of the bed;
the other leg of the Tee provides a sump for
any material falling through the inlet hole. To
permit increased flow rate through the bed,  if
required, a bypass valve is provided  around
the air rotameter, with an "Annubar" flow
meter to measure these higher flow rates.

   Cast into the furnace at heights of 9 and 12
inches above the base are two rows of 11 holes
capped  with  1/4-in.  pipe   nipples  for
instrument access or solid feed. Twenty-three
pressure tap holes lined with 1/4-in. porcelain
are cast into the side at 1/2-in. intervals from
the bottom to a height of 8 inches, and every 1
inch thereafter to a height of 16 inches. Before
casting,  1.5-in. pieces  of stainless steel were
soldered to the oil drum shells at all pressure
tap and thermocouple  stations to protect the
porcelain from  breakage.  Shielded thermo-
couples flush with the inside wall are provided
at elevations of 1, 3, and every inch thereafter
up to 19 inches. Two sets of thermocouples are
also mounted at intervals  of 1/2-in. depths
into the wall  to monitor the wall temperature
profile.  All  thermocouples   are  Chromel-
Alumel,'with read out by a Leeds & Northrup
24-point recorder. Pressure is monitored by a
36-in. tube, well type manometer. A Chromel-
Alumel  thermocouple  is   used  for bed
temperature measurements with readout by a
single channel Honeywell recorder; a check on
bed temperature can also be made by optical
pyrometer.

   A draft gauge with range from + 0.005 to -
0.015-in. we  is connected  to the horizontal
flue-run for pressure monitoring. Two water
cooled probes 48-in. long  are used for gas
sampling, one in the stack and the other in the
bed. Gas is continuously monitored for CO2
and CO with infrared instruments, and for C2
with Thermox analyser using  a  fuel cell
as the sensor element. Incomplete combustion
of the gas is also determinable directly by an
MSA total combustibles analyser.

   Initial  startup,  when the unit was  first
completed,  was   accomplished  without
difficulty  using  the  technique  described
earlier. However, some difficulty was initially
experienced in maintaining fluidization for
any length of time. The bed would  start to
blind at  some point.  Fluidization would be
lost; the gas/air flow would be diverted from
those regions which would cool, thus tending
to prevent refluidization. Careful analysis of
the problem  suggested that it was probably
due to fines in the bed material.  At all events,
removal of fines by  screening on a 16-ASTM
sieve eliminated the problem.

   Progressive improvements to the  design
and method of operating the unit resulted in a
steady improvement in fluidization; as fluidi-
zation improved, the bed depth required for
complete combustion dropped steadily. In the
preliminary experiments on the first unit, bed
depths up to  12-in. were used. In the current
unit, beds were initially 4 to 6-in., but were
finally reduced to 2 inches and then to 1  inch
with  the  even  fluidization  achieved.  For
heating purposes (in the bed) or solid wastes
incineration greater depths are required. For
our immediate purposes here, however, since
no interesting combustion behavior occurred
above the  bottom 1 or 2 inches of the bed, the
top layers were "omitted"  to  allow easier
access by probes to  the regions of interest. It
should be emphasized, of course, that until the
present  unit  was completed, and the  very
uniform fluidization obtained,  the evidence
indicated  that combustion required  3 to 6
inches.


PHYSICAL PROPERTIES OF  THE FLUID
BED

   In  the   course  of  the   combustion
investigations,  the  porosity  and other
properties of the hot bed were measured, some
of which were  obtained  as  a   matter  of
necessity.  Because   of  their  potential
engineering  value   for   design   they   are
                                                                                   11-5-5

-------
summarized here. It is emphasized that these
are measurements made at high temperatures,
up to 2400F, with combustion following;
since the bed  depths were  small, the  data
should be valid for all bubble free conditions.
The  data of  principal  interest were  the
expanded  bed   heights,  porosities,  gas
velocities, and residence times, all of which are
needed for interpretation of the combustion
data.

   Values of pressure drop (AP) between the
top of the bed and a point below the top of the
bed were first measured by traversing the bed
with an open-ended, water-cooled probe.  (The
kinetic head contribution is too small to cause
any determinable error.) Measurements  were
taken for six  different  fluidizing velocities,
with  bed temperatures ranging from  1800 to
2350F; the normalized results of AP against h
are shown in Figure 2 where h is measured
from the bottom of the bed. The normalizing
parameters  used were   the  total  pressure
drop APtand the (expanded) bed depth L. The
top of the bed was identified by a break in the
slope of the AP  against h line. At first glance it
is evident that  the plot of Figure 2 is respec-
tably linear, in  agreement  with  theoretical
predictions quoted in standard tests; e.g., 1, 8.
With closer inspection, the slight curvature of
the line is self-evident; the curve is probably
due to the temperature variation  and hence
the velocity variation through the bed. How-
ever,  the departure  from linearity  is  small
enough that it  can be neglected for our  pur-
poses.
   The normalizing parameters, APt and L0,
were  also found to obey  simple  theory.
Equations given by Davidson and Harrisoni
have been used. Renormalizing their equation
(1.11) against the initial  bed depth  (LQ)  at
incipient fluid ization velocity (Uo) when the
initial porosity  is  o, we  obtain
where APot is the pressure drop across the bed
at incipient fluidization. However, (APt/AP0t)
is unity in the fluidization region, as Figure 3
illustrates. Note also that the value of AP0t is
below that required to balance the bed weight
This is in accordance with Trevedi and Rice's
experiments.5  A  further   simplication  of
equation (1)  is possible  by  using  equation
(1.22) of reference 1, again  renormalized to
yield
              = (/o)3(l-o)/(l-)
(2)
Figure  4  substantiates  this  equation,
illustrating the linearity obtained by plotting
3/(l-)  against U,  where U is determined
under the hot condition. Figure 4 also shows
the actual variation of  with U. Values of E
were calculated from the particle denisty  (a)
and the bulk bed density (p)  (calculated from
bed weight, depth, and areas)  using
        =l-p/a=l-0.464
(3)
The incipient  fluidization  porosity  (o)  is
0.536, which is somewhat above the value  of
0.476 for  cubic packing of uniform spheres.
Since the alumina particles  are quite good
spheres it is not unrealistic to assume this less
dense packing is due to particles bridging void
areas.
Substituting equation (2) in equation (1) yields
        L/Lo = (l-o)/(l-).
(4)
This relation is substantiated by Figure 5. The
slope,  equal to  (1 <- o),  has a  value 0.46
yielding Q= 0.54, which is good  agreement.
Figure 5 also shows  the variation of (L/LO)
with U, with the fitted curve calculated from
equation (2) adopting the experimental values
of U0 and Q. The gentle curve could be well
approximated by a straight line,  which is  a
consequence of e3 /(I-e) varying almost linearly
with !/(!-); it is approximately proportional
to U.
II-5-6

-------
   Absolute calculation of the incipient fluid -
ization velocity  is not quite  so satisfactory
although  the  experimental  value  can  be
bracketed.  Again, using an expression given
by Davidson and Harrison,1
U0 = 0.00081 (cgd2/n) (cm/sec)
(5)
where d is the particle diameter (=0.0510 in.)
and  (<  is the  dynamic  viscosity.  Uo  was
calculated  for  ambient  temperature   and
1800F,  knowing  the  average  weight of one
particle.  The predicted values were 3.28 ft/sec
and 1.21  ft/sec  (at the higher temperature).
This  bracketed the experimental value of 2.0
ft/sec. Since the numerical factor of 0.00081
was determined  for experiments on fluidizing
with  water,1  the  agreement  is  acceptable.

   From  the data given,  calculation of the
actual (as opposed to the superficial) velocity is
straight forward using a  porosity correction;
the same occurs for the residence time in the
bed Os). The results of the two calculations are
given in  Figures 6 and 7, respectively, with
bounding values for0 and  = 1 included for
comparison. The  residence  time  data,  in
particular, are needed for discussion of the
combustion behavior.

   The  general  agreement  with  theory
established here would support the use of the
equations tested for engineering design and
scale-up,  if used with care.

COMBUSTION BEHAVIOR

General

The  problems of light-off have been  fully
covered above. Once lit,  combustion  can be
maintained  indefinitely as long as no blinding
of the bed  by fine particles occurs.  In the
earlier experiments,  when fluidization and
mixing of the fuel and air were relatively poor,
the over-bed gases were  periodically  flecked
with yellow as bubbles of gas broke the surface
and burned in the  freeboard  space.  This
behavior  was  progressively  eliminated  by
improved fluidization and mixing. With the
present arrangement, utilizing premixing in a
venturi mixer, there can be no bubbles of fuel
rich gas. However, it was clear that the speed
of reaction was still strongly influenced by the
fluidization quality. When this was poor, 4  to
6 inches were apparently required for  com-
bustion.  As  fluidization was improved, the
space  required  was  progressively  reduced
until, as mentioned above, it could be accom-
plished well within a 1-in. bed. Under  these
conditions, all that can be seen through the
sight glass in the top is the red hot bed-top  in
continuous motion without  any spouts,  and
with particles welling up  and disappearing
again.

   In the experiments next  described there
were three objectives. With  the  expectation
that gas-fired fluid beds will be increasingly
used in commercial  practice, attention was
first directed at two aspects of safety: (1) if fuel
to a bed is cut off, at what minimum tempera-
ture will it relight? and (2) if a surge of air  or
neutral gases leans out the fuel-air mixture,  at
what  gas  percentage and  bed temperatures
will there be extinction? With information on
these first two, attention was then  given  to
behavior in the bed in an attempt to determine
whether combustion in a particle-filled volume
is affected in any way by the presence of the
particles.

Relight

   To investigate  relight  behavior,  the
procedure was to  set up the  bed in normal
operating condition and then to switch off the
fuel and the air, or to decrease the air. The bed
would  cool;   and periodically  at  recorded
temperatures  the  fuel   supply   would   be
restarted. The observation then  made  was
whether or not the  bed  temperatures would
start  to rise  again.  This was taken  as a
condition of relight.

   The experiments on relight were carried
out at a fairly early stage in the investigation,
with beds 4 to 6-in. deep, and mostly at fuel-
air ratios closer to stoichiometric than to the
lean limiir Under these conditions, relight was
                                                                                     ii-s-7

-------
 always successful down to bed temperatures of
 750 F (400C). Lower temperatures than this
 were not investigated for safety reasons. This
 is  substantially below  the values generally
 listed in standard data tables; e.g., 4, for auto
 (spontaneous)   ignition   temperatures.
 Reference 4, for  example, quotes:  1290F
 (700C) for the stoichiometric mixture; greater
 thai* 1200F (650C)  for  the most easily
 ignited mixture; and adds the comment  that
 under pressure, the  temperature is never less
 than 880F (470C). It seems fairly clear that
 the method of determination is too different
 from a  fluid bed for the results to be relevant.

 Extinction

    The procedure for determining extinction,
 at  the low limit, was to start with the bed in
 normal operating condition at some suitable
 fuel  and  air rate, and  then to lean  out  the
 mixture by stepwise reduction of the fuel rate.
 When the flame extinguishes the temperature
 falls rapidly. The extinction point can  only be
 judged  between two steps, always approached
 from the  flame side of the boundary. After
 extinction the bed was relit, the air rate reset,
 and the sequence repeated.

    A typical set of results is shown in Figure 8
 which  illustrates   plots of  temperature
 (maximum observed  by thermocouple) against
 the superficial velocity (calculated utilizing the
 maximum temperature) through the bed  for
 several run  sequences.  The extinction-
 temperature boundary is clearly marked as a
 heavy dashed line.  The lightly dashed lines
 represent   constant  fuel-air  ratio in  per-
 centages  by volume.  The  continuous  line
 marked 5.3 percent is the conventional  low
 limit. It  can be seen  that a considerable
 number of combustion points lie below  the low
 limit. This is more clearly seen on Figure  9,
 showing continued combustion down to 4 per-
 cent methane, more than 1 percent below the
 normal  low limit. Figure 9 also includes the
theoretical adiabatic flame temperature  line
with  some  temperatures  exceeding   the
 adiabatic   value.  The   source   of  these
 unexpected  peculiarities  lies   in  the  heat
 exchanger effect of the fluid-bed particles, as
 discussed below.

 Bed Profiles

   The same super-adiabatic  behavior also
 occurs in the bed itself. Figure 10 is a typical
 temperature  traverse down through the bed.
 Traverses were made both with sheathed and
 bare thermocouples,  and a displacement of
 the profiles was observed. By  using a special
 sheathed  couple  set  at  right angles to  the
 holder, it was established  that conduction
 down the sheath could  result in spuriously
 high temperatures at a given point. The data
 are,  in effect, translated  by  about 1/4-in.
 However, this may not entirely  account for  the
 differences.   Temperatures   exceeding   the
 theoretical adiabatic by 50 to 150F have been
 recorded in many of the temperature traverses
 made.

   Clearly, the rate  of  heat  removal  must
 exceed the rate of reaction at locations above
 the temperature peak (in the regions marked
 C and D) in order to allow the  temperature to
 decline. Furthermore, the reaction is probably
 totally   completed  at the location  of  the
 temperature peak. To check this, gas analyses
 were taken throughout the bed;  Figure  11
 illustrates one method of representing the fuel
 consumption  calculated from the gas analyses.
 The graph represents the unburned fuel, on a
 log-linear plot, declining with  distance  up
 through the bed. The fuel unburned was back
 calculated from the CC*2 analysis; the CO was
 never found to exceed 0.75 percent and was
 disregarded in the calculation. The two curves
 represent  two  different bed  temperatures,
 185_0F (1010C)  and 2350F  (1290C), with
 fuel burn-up  easily followed through the bed,
 although the times represented are only of the
order of milliseconds. The fuel concentration
 in each case decays more or less exponentially
through  the  bed.  Reaction is fast  with  90
 percent reaction in 16.0 msec and 2.2 msec  for
the  lower and higher  temperature  beds,
 respectively.  These  figures  are in agreement
with  the temperature profiles  and with  the
 prediction that combustion is mostly complete
II-5-8

-------
before the temperature peak is reached, and
all  (detectable)  combustion  is  completed
within the bed.
DISCUSSION

   The physical behavior of the bed is in good
agreement with expectation from established
theory and requires no further comment. The
combustion  behavior, on  the  other  hand,
shows a number  of unexpected  features.
   (1) The lean limit extension  and super-
adiabatic temperatures indicated in Figures 8,
9, and 10 were particularly unexpected.  They
are, however, simple to explain. They depend
on a heat recovery or heat exchanger effect
due to the bed particles moving up and down-
stream. To understand the behavior, consider
first the  effect of  a  heat  exchanger in the
exhaust  of  an  otherwise  adiabatic  flame
system.   With  the  heat   exchanger   not
connected, the gas exit temperature of the
flame system is the adiabatic flame tempera-
ture. If the heat exchanger is allowed to heat
the  incoming combustion  air for the flame
system, the gas  exit  temperature  must  be
boosted to adiabatic plus the preheat. Overall,
of course, there is no gain because  the  extra
heat in  the exit gases is removed by the heat
exchanger (exactly, in a no-loss system); the
gases now leave  the  heat exchanger at the
adiabatic flame  temperature.  This is,  of
course, no more than  the usual application of
the heat exchanger although the  potential of
heat exchangers  for  producing  super-
adiabatic  temperatures  is  not always
recognized for what it is.

   In the case of the  fluid bed, the net heat
exchanger effect is clearly evident, but it is not
particularly  efficient  in  this role since the
temperature excess  is only 50 to  150F. The
preheat  influence  is  believed to  be  most
marked in the early stages of the temperature
rise.  In  these  regions the  supply of bed
material must be predominantly  from above
(i.e., from hotter zones), whereas further Up in
the  bed  there can  be  as  much  material
supplied from below  as above,  thus contri-
buting  to  cooling  of the  upper  zones.
Nevertheless, the overall  consequence is to
accelerate the rate of heating, and therefore
the rate of combustion in the bed.

   (2) The same preheat effect is responsible
for the  extension of  combustion  below  the
usual low limit. Again consider the case of air
preheated to a very high temperature (say, by a
heat exchanger). If the temperature is  high
enough, any quantity of fuel, however small,
injected  into  that  air  stream  cannot  be
prevented from reacting completely. Between
this extreme condition and the  condition of
the normal lean limit there is a range of rising
temperatures  permitting   a   progressive
lowering of the lean limit to zero. Weinberg,7
for instance, recently quoted a system in which
stable combustion is  maintained at a  gas
concentration of methane in air of 1 percent;
combustion  contributed 250C  and preheat
contributed  1000C.

   Clearly, the heat exchanger effect can be at
least partly responsible for the widening of the
combustion  limit found in  the fluid  bed.
However,  the magnitude of the effect   a
drop of over 1 percent in the low limit  does
seem to  be rather large for the relatively small
temperature increase over adiabatic, of 50 to
150F. Some other effect, as discussed below,
may also be involved.  This view  is supported
by a few measurements using beds of double
the  depth.   The   maximum   temperature
increased by about 50 F, but the limit mixture
at extinction was unaffected. It may also be of
significance that the super-adiabatic tempera-
tures were only obtained at the higher  flow
rates, presumably because the heat exchanger
effect was stronger.

   (3) Indications that  factors,  other than
those already mentioned, could be influencing
the reaction  were obtained from  estimates of
reaction time. Table 1 lists some  estimates of
time to  complete 90 percent of  the reaction
(based on the Figure 11 plots). Included for
comparison  are  some  data   from  other
sources.2-3'6 The most directly comparable
                                                                                   II-5-9

-------
      Table 1.   ESTIMATES OF  REACTION TIME IN A FLUID BED AND IN CONVENTIONAL FLAME SYSTEMS
tyi

H*
e
Investigator
Present work













Dixon-Lewis
Levy
Van Tiggelen

Van Tiggelen

CH4. %
4.25








5.92




5.03
5.4
Stoichio-
metric
Stoichio-
metric
02, %
20.1








19.88




19.94
20.0
Stoichio-
metric
Stoichio-
metric
Inert, %
75.65 (N2)








74.2 (N2)




75.03 (N2)
74.6 (AR)
73 (AR)

65 (AR)

t
K
1290








(adiabatic)
1650
(observed)
1530

1528
1950
2110

2370

F
1865








(adiabatic)
2450
(observed)
2300

2300
3050
3340

3800

V
'ft/ sec
2.7





/



2.2



0.17
1
1.1

2.2

cm/sec
83









74



5.2
31
34

66

T
msec
22a

16a

11

6



2.2



17 + 3
4.9 + 1
0.060

0.015


Total bed
residence time
Temperature peak
90% reaction
Total bed residence
time using Tmgx
Temperature peak
90% reaction
using Tmax

Temperature peak
90% reaction
using T
max
90% reaction
90% reaction
Mean molecular

residence time

      Temperature does not show a significant rise until after 8 msec have elapsed.

-------
data are those given by Dixon-Lewis.2 For 90
percent reaction, 17  3 msec are required in a
flat-flame system, while only 2.2 0.5 msec are
required in the fluid bed at very  close to the
same temperature (2300F),  and only 4 1
msec for 99 percent reaction. This factor of 6
or 7 difference is clearly significant.

   The most probable explanation that comes
to mind is, of course, enhanced reaction due to
the particles. This could be a result of either
initiating more gas-phase reaction or catalytic
surface reactions.

   (4) Some choice between catalytic  surface
reactions  or  enhanced gas-phase  reactions
may be possible from further data developed
from the fuel consumption curves  of Figure
11. Assuming that the heat from consumption
goes exclusively  into  the  mixture and the
products at the  local level,  a temperature
profile through the bed was calculated. Figure
12   shows   how  this  compares  with  the
measured profile. The deviations at the top
end have already been explained as a result of
the  preheat effect. The discrepancy at the
lower  temperatures,  with   the  predicted
temperatures  substantially   in  excess  of
measurement, was totally unexpected.
              *
   In   accounting   for   the    observed
discrepancies a number of explanations were
considered; all but two were  discarded.  The
simplest explanation is that the gas analyses
may show spuriously high CO2 values because
of continued reaction in the sampling probe.
Against that, however, is the matter of the low
temperatures involved so that any substantial
cooling of the  gases would  freeze the compo-
sition. The other explanation is more involved
but is also considered more likely. It is based
on the assumption of significant temperature
difference between the  particles  and  gas,
which is quite possible considering  the rapid
translation of hot particles into the cooler bed
zones and  the short  times involved  for re-
equilibration.  If,  therefore, the hot particles
stimulate  surface catalytic reaction so that
most of the heat released goes directly into the
particles, the lead in the particle temperature
above the gas temperature will be maintained
until reaction decays. Either bare or sheathed
thermocouples  will then take up a tempera-
ture intermediate between the  gas and  the
particles, but  the  sheathed  couple  can be
expected to be more responsive to the particle
temperature  because of the enhanced  heat
transfer  coefficient between  particles and a
surface. This would then help to explain the
discrepancy noted above between the bare and
sheathed  couples as a factor  additional to
conduction as noted.

   (5) Finally, a brief examination of the rele-
vance  of  this information  to engineering
applications is in order. The outstanding point
is that reaction in unpremixed  systems  will
clearly be dominated by the mixing behavior;
the time for reaction can be virtually  ignored
unless temperatures  are very  low  indeed.
Clearly,  future experiments  should  include
measurements down to about 1000F or lower,
which may well be achieved on occasion  if a
very  wet slurry  or  sludge  is  incinerated.
However,  on  the  unit  used,  controlled
variation of the bed temperature has been  very
difficult. There could be some  advantage in
reducing the bed size to provide better  bed
temperature  control. The  other aspects of
possible engineering significance are the wider
combustion limits and enhanced  reaction  rate
(i.e., speed of ignition) due to the hot particles.
These could  also  increase the risk of serious
explosion of any  large bubbles  of premixed
fuel and air, if such bubbles are ever permitted
to form. The  extinction and relight conditions
are    also   important;    consequently,
development  of  analytical  models  for
experimental test  is now required to provide a
more reliable basis for  extrapolation. (Two
models have been developed, but they  are still
too limited in their assumptional basis to be of
much value yet.) Beyond that, what is mainly
needed  now  for  engineering  purposes is an
understanding of the behavior of jets and mix-
ing in fluid beds.

   In   conclusion,   therefore,   the  results
developed  iii this  paper  substantiate  the
                                                                                   II-5-11

-------
 reasonable expectation that reaction of gas in
 fluid  beds  is  fast,  and  that  problems  of
 incomplete bed reaction must be due to poor
 mixing  in  the  bed.   In  addition,  and
 unexpectedly, it was also found that reaction
 seems to be accelerated by the presence  of
 particles which also can widen the combustion
 limits and generate super-adiabatic tempera-
 tures. Other results include the development
 of a means  of  rapid   light-off  and  a
 demonstration that the  physical behavior  of
 the bed is in general accordance with expecta-
 tion  from  available  theory, in spite of the
 simultaneous  presence of combustion.

 ACKNOWLEDGMENTS

   We   have  pleasure  in  acknowledging
 support  for  this  work  from  Consolidated
 Natural Gas Service Corporation under Grant
 No.  C-70-29-2.  We  also gratefully acknow-
 ledge assistance in the construction of the
 equipment from Messrs. R. Frank, C. Martin,
 and D. Simpson. The first author also wishes
 to thank Mr.  M. Kuwata for assistance he  so
 ably  provided during this work.


 LIST OF SYMBOLS

 c jf  =  Initial fuel concentration, by volume

 h    =  Location  in  bed  measured  from
       bottom, in.

 L     = Bed thickness, in.

 AP  =  Pressure drop, in. EbO

 Tp  =  Flame temperature

 U    = Superficial velocity (calculated at bed
        maximum temperature), ft/sec

 V    = Flame velocity

   = Bed porosity

p    = Bed bulk density, lb/ft3

*     Completeness of combustion

 a    = Particle mass density, lb/ft3

 II-5-12
Ts   = Gas residence time in the bed, sec

Subscript

o    = Incipient   fluidization  valuesalso
       used  to  denote  fixed  bed  values
       where  applicable (AP, e, L).

REFERENCES

1.  Davidson, J.F. and D.  Harrison. Fluidised
   Particles.   Cambridge,   Cambridge
   University Press, 1963.

2.  Dixon-Lewis, G. and A.  Williams. Some
   Observations   on  the  Combustion  of
   Methane in Premixed Flames. (Presented
   at the llth Symposium (International) on
   Combustion. 1967.  pp 951-958.)

3.  Levy, A., J.W- Dredege, JJ. Tighe, and J.F.
   Foster. The Inhibition of Lean Methane
   Flames. (Presented  at  8th  Symposium
   (International)  on Combustion.  1961.  pp
   524-533.)

4.  Spiers, H.M. (ed.). Technical Data on Fuel.
   Edinburgh, British  National Committee
   World Power Conference,  1962. p. 260.

5.  Trivedi, R.C. and WJ. Rice.'Effect of Bed
   Depth, Air Velocity, and  Distributor on
   Pressure Drop  in an Air  Fluidized  Bed,
   Fluidized Bed  Technology,  Chemical
   Engineering  Progress  Symposium Series,
   American Institute of Chemical Engineers.
   62, 1966.

6.  Van Tiggelin, A. and J. Deckers.  Chain
   Branching  and  Flame  Propagation.
   (Presented    at    6th    Symposium
   (International)  on Combustion.  1957.  pp
   61-66.)

7.  Wemberg. Combustion Temperature:  The
   Future. Nature. 233:239, 1971.

8.  Zabrodsky, S.S. Hydrodynamics and Heat
   Transfer  in  Fluidized Beds.  Cambridge.
   MIT Press, 1966.

-------
ADDENDUM  ON RELIGHT BEHAVIOR

   Since this paper was written further data
have been obtained  concerning  the  relight
behavior of the bed. This further information
amplifies  the  results  quoted  in  the  section
titled "Combustion Behavior:  Relight."

   The method of experiment  was as follows.
In  the  tests  described  in the  section  on
"Relight" the bed was first fired up normally
till thermal  equilibrium was obtained. Then
the gas flow was decreased to a value below the
lean   flammability   limit   so   that   the
temperature declined slowly. Combustion was
still occurring in the bed, but the heat release
rate was insufficient to maintain equilibrium.
We also conclude that combustion in the bed
was not quite complete, because the reaction
was evidently continuing on surfaces in the
freeboard area, such  as  the  stainless-steel
sheathed temperature probe. It was observed
that the probe was glowing red when the bed
was black and, therefore, presumably cooler.
When the gas flow was increased (to the level
of stoichiometric gas-air mixture), relight was
always obtained down to 750F (400C), as
already reported. The glowing thermocouple
sheath  did   not  originally appear  to  be
important since  the  temperatures reported
were presumably those of  the thermocouple
(and sheath); the reported temperatures, being
greater  than those in the  bed,  provided  a
conservative margin.
   These tests were recently repeated, but with
the  gas  flow  turned completely  off. The
temperature declined much more rapidly since
there was no combustion occurring in the bed
to retard the temperature decline rate. This
rapid  drop  made  temperature   estimation
difficult which was the reason for trying to
control the  temperature rate  of  fall  in the
original  experiments.  There  was also  no
combustion in the over-bed region to provide
an over-bed ignition source as in the first case.
Relight under these new conditions was not
obtained  even at  1400F (750C).

   These results indicate that in the first case
the hot  areas  in  the  over-bed  region were
providing the ignition. Thus, for  engineering
applications,   an  over-bed  ignition  source
should be present for safety reasons.

   The relight  temperatures obtained  under
the second set of conditions are evidently at or
above the values  generally quoted for auto-
ignition.  The experiments  thus underline the
very   significant  distinction  between  auto
ignition and reactor ignition. They show very
clearly the  safety of the system  and ease of
relight down to very low temperatures, even if
the gas  concentration is  very substantially
below the low limit. Risk  occurs  only  in the
event of total failure of the gas supply, which is
easily  guarded against by  standard  safety
precautions.
                                                                                  II-5-13

-------
     IGNITION
     BURNER
     HP GAS
    WALL
    THERMO-
    COUPLES;
    18 TOTAL
    TYPEK
                                            WATER COOLED
                                            GAS SAMPLING PROBE

                                                TEMPERATURE PROBE,
                                                TYPE K THERMOCOUPLE
                                                            DRAFT GAUGE
                                                           FLUE GAS EXIT
                                                                                 12 in.
                                                             WATER COOLED WALL
                                                             ALUMINA BED (1 in.)
ROCK DISTRIBUTOR
NET (6 in.)
                                                             CASTABLE
                                                             REFRACTORY
                                                               AIR-GAS PREMIXER
                                                                  AIR INLET
                                                             GAS
                                                            INLET
                Figure 1.  Details of experimental apparatus used for these tests.
II-5-14

-------
    0.0
!
Q.
LU
V)


o
o
     0.2
                                                                                       1.0
                                  DIMENSIONLESS BED DEPTH, h/L

      Figure 2.  Variation of dimensionless pressure drop  Ap/Apt with bed depth h/L.
      Data taken with water cooled probe in a thin fluidized bed of 1-in. fixed-bed
      thickness.  Combustion in the bed produced a temperature variation of 1800-2350F,
      and velocities of 3 to 5 times incipient fluidization velocity (2 ft/sec).
                                                                                    II-5-15

-------
  o
  CM
      1.0	
       0.8
  CO

  
  V)
  O
  ce
  u
  
-------
o
ce.
o
Q-
                                                                                   0.3
                              HOT SUPERFICIAL VELOCITY, ft/sec

                                                          3
     Figure 4.  Variation of porosity  (see equation 3) and  / (1 -Q (see equation 2)

     with hot superficial velocity.  The second group shows correlation between theory

     and experiment.
                                                                                     II-5-17

-------
                               HOT SUPERFICIAL VELOCITY, ft/sec

        0.0        2.0        4.0        6.0         8.0         10.0       12.0
     1.0
                                       VALUES OF
                                                  (1-6)
           Figure 5.  Variation of the dimension I ess bed thickness L/Lo with both
           the superficial velocity U (curve calculated from equation 2 using experimental
           values of U  and L  ) and with 1/(1-G) as substantiation of equation 4.
II-5-18

-------
o
o
    2.0
                                                                            10.0
                                HOT SUPERFICIAL VELOCITY, ft/sec
           Figure 6,  Variation of actual velocity with hot superficial velocity u.
           The dotted lines indicate what the velocity would be if the bed had a
           constant porosity equivalent to fixed bed (=0.536) or if no bed were
           present (=i).
                                                                                    II-5-19

-------
  o>
  o
  LLJ
  a
  LlJ
  CO
 INCIPIENT
FLUIDIZATION
      10.0
      5.0
                                                                                      10.0
                                      HOT SUPERFICIAL VELOCITY, ft/sec
              Figure 7.  Variation of bed residence time Ts with superficial velocity.
II-5-20

-------
UJ
ct

ffi
0-
    2300
    2200
    2100
    2000
    1900
    1800
                 EXTINCTION LINE
            6.0
8.0
10.0
                                     HOT SUPERFICIAL VELOCITY, ft/sec
                Figure 8.  Variation in temperature with hot superficial velocity.
                The initial fuel concentration is also indicated. Extinction occurs
                at the low temperature and low  initial fuel concentration.
                                                                                       II-5-21

-------
            COLD SUPERFICIAL VELOCITY
    2400)	   O  1.27 ft/sec

               A 1.49 ft/sec


               V 1.75 ft/sec

    2300       D 2.02 ft/sec
    2200
    2100
    2000
    1900
    1800
                                                          LOW
                                                           LIMIT
                                                        I
                                INITIAL FUEL CONCENTRATION, Cif: %

          Figure 9.  Temperature as function of initial fuel concentration for differing cold
          superficial velocities.  Also indicated are the adiabatic line and the generally
          accepted lower flammability limit.
H-5-22

-------
                                    TIMESCALE,T: msec
                                               16
2000
ADIABATIC
TEMPERATURE
                        0.4        0.6        0.8        1.0        1.2

                              HEIGHT ABOVE BOTTOM OF BED, h: in.

    Figure 10.  Temperature as a function of height  in the bed with  inlet and adiabatic
    temperatures indicated. An auxiliary axis shows the^elapsed time in milliseconds.
                                                                                 II-5-23

-------
  o
  o
  o
  ce

  ca
                                          0.4               0.6



                                      HEIGHT ABOVE BOTTOM OF BED, h : in.
0.8
1.0
             Figure 11.  Unhurried fuel concentration (calculated from CC>2 analysis)

             as a function of height in the bed for two initial fuel concentrations.
II-5-24

-------
2000
           ADIABATIC
           TEMPERATURE
                                             TEMPERATURE
                                             FROM C02
                                             ANALYSIS
                                                                               1.4
                             HEIGHT ABOVE BOTTOM OF BED, h : in.
          Figure 12.  Actual temperature measurements and temperature calculated
          from the completeness of combustion as a function of height in the bed.
                                                                              II-5-25

-------
SESSION III:
  Gasification/Desulfurization

SESSION CHAIRMAN:
  Dr. E. Gorin, Consolidation Coal
                              III-O-l

-------
               1. FLUIDIZED-BED GASIFICATION-PROCESS
                               AND EQUIPMENT DEVELOPMENT

                     J. T. STEWART AND  E. K. DIEHL

                      Bituminous  Coal Research, Inc.
INTRODUCTION

   As part of its broad gas generator research
and development program sponsored by the
Office of Coal Research," U.S. Department of
the Interior, Bituminous Coal Research, Inc.,
(BCR), is developing a multiple fluidized-bed
coal gasification process for the production of
low-Btu  fuel gas. The goal of the multiple
fluidized-bed system is the gasification of both
caking and non-caking  coal,  with  fuel  gas
being the only product.

   Several government and industry co-spon-
sored coal gasification programs  are at the
pilot plant stage. These include the IGT Hygas
process,  the BCR BI-GAS process, and the
Consolidation  Coal Company's COa acceptor
process. Such processes are designed to gener-
ate high-Btu gas, i.e., gas having a heating
value in excess of 900 Btu/scf. Gas of  this
quality can be used as a direct substitute for
natural gas. Steam boiler and  gas turbine
applications for electrical power generation do
not, however,  need this high-Btu  gas. More-
over, the optimum fuel gas  heating value
required for combined cycle applications can
be of the order of 150 Btu/scf. The primary
purpose of the  multiple fluidized-bed coal
gasification system is, then, the production of
low-Btu fuel gas for the generation of electri-
cal energy by means of  the combined cycle.

   This paper describes  the BCR  fluidized-
bed gasifier concept and summarizes the work
done by BCR in its development program. The
program began with laboratory scale kinetic
experiments and has progressed through the
semi-continuous  operation  of a small fluid-
ized-bed batch reactor.

   With the aid of an engineering subcontrac-
tor,  a  process and equipment development
unit  (PEDU)  has  now  been designed. The
PEDU, to be located at the BCR Research
Center  at Monroeville, Pa., consists of three
fluidized-bed reactors, a gas quenching and
scrubbing system, facilities to preheat reactor
inlet gases, and solids handling equipment.
Designed to gasify 100 Ib coal/hr, 'the PEDU
will operate at 250 psia and at temperatures to
2100F.
                                      ra-i-i

-------
THREE   STAGE   FLUIDIZED   BED
CONCEPT

   The goal of  the  multiple  fluidized-bed
system is the gasification of both caking and
non-caking coal, with fuel gas being the only
product. End use of the gasification product
dictates the operating conditions as well as the
gasifying medium. Thus,  gasifying with  air
and steam  will yield a  low-Btu  fuel  gas;
gasification with oxygen and steam can yield a
higher Btu gas containing a greater proportion
of combustible  components; and gasification
with  carbon dioxide  can  yield  a carbon
monoxide-rich gas that could serve  as a fuel
for power generation  by MHD.
   A  3-stage system was chosen as one with
the probable minimum number  of stages
necessary to meet the requirement of starting
with any rank of coal and producing no tars or
oils as a waste or by-product. Figure 1 is the
design  material  balance  for  the  proposed
fluidized bed PEDU. Stage 1 receives raw coal
and functions as the pretreatment step. The
devolatilized coal flows  by gravity to Stage 2
arid then to  Stage 3, which operates as the
final   carbon  burn-up  reactor.   Several
pretreatment mediums have been investigated
by others and have been  shown to be effective.
These include air alone,  and steam or carbon
dioxide diluted with nitrogen and containing
small  amounts of air. In this scheme, Stage 3
flue gas is used as the fluidizing medium for
Stage 1.


   Stage 2 is the major gasification stage. The
devolatilized  coal is gasified  with air and
either steam or carbon dioxide to generate the
desired product gas. In addition, Stage 1 flue
gas is fed to Stage 2 where the entrained tars
and oils are gasified. Stage 3 operates at the
highest temperature  and serves to  maximize
carbon utilization. The  ash discharged from
Stage  3 will contain a minimum  amount of
carbon. Hot flue gas from Stage 3 flows to
Stage  1 and completes the cycle.
LABORATORY INVESTIGATIONS

   The literature  abounds with information
regarding the  kinetics of carbon/steam and
carbon/carbon dioxide reactions, but there is
often  little agreement  among the  results  of
different  investigators. For example, reported
activation energies for  the  carbon/carbon
dioxide   reaction   range  from  25  to  90
kcal/mole.  This   wide  variance  may  be
attributed to the different types and ranks of
carbon used, the  different temperature and
pressure  ranges investigated, and the various
simplifications  and  interpretations  of the
observed   data.  The   different  results  are
indicative  of   the   different  reaction
mechanisms and  rate controlling steps that
occur   under -  different   experimental
conditions. No correlations were found in the
literature that could adequately describe the
kinetics of the gasification reactions to the
devolatilized coal or "char"  that  would be
produced  in the  first  step of the multiple
fluidized-bed gasification scheme. Therefore,
laboratory scale kinetic studies were begun.

   Six widely different chars were  chosen as
the basis  for the laboratory studies.  The  chars
ranged in volatile content from 3 to 12 percent
and were produced from all ranks of coal,
from  lignite to a highly caking Pittsburgh
seam  coal.  Table  1   shows  the  chemical
analyses of the selected chars.
EXPERIMENTAL  EQUIPMENT  AND
PROCEDURE

   A thermogravimetric balance was used to
obtain the kinetic data. A schematic view of
the TGA is  shown  in Figure 2.

   Char is  placed  in the sample holder,  a
crucible or flat pan, which is connected to and
suspended beneath  the transducer coil  and a
precision spring.  This  entire  assembly  is
mounted  inside   a  quartz  and   Pyrex
housing. During  operation,  the  sample  is
located inside the well of the furnace where
temperatures are continuously monitored by a
chromel-alumel thermocouple.
m-i-2

-------
Table  1.  ANALYSIS OF CHARS USED  IN REACTIVITY TESTS
Char sample
number
2455
2469
2927
1963
2280
2655
Dry basis, %
Proximate
Volatile
matter
2.85
11.6
2.1
11.3
5.9
4.2
Fixed
carbon
88.9
58.4
74.6
71.4
82.5
83.5
Ash
7.65
30.0
22.5
17.3
11.6
12.3
Ultimate
Carbon
88.7
56.1
73.5
77.1
82.0
83.4
Hydrogen
0.8
1.92
1.35
1.03
1.8
1 .42
Nitrogen
1.3
1.16
0.62
0.49
-
0.9
Sulfur
0.6
9.39
2.89
0.90
0.37
0.76
Ash
7.7
30.0
22.7
17.3
11.6
12.3
Oxygen
0.9
1.43
0.0
4.18
-
1.22

-------
   The char is brought to the chosen reaction
temperature in an inert atmosphere. The inert
gas stream is turned off; simultaneously the
reaction gas stream is  turned on. Sample
weight loss is then recorded as a function of
time.   Data  precision   is   checked   and
maintained   within   0.1  percent  weight
loss/unit time by duplicating all experimental
runs.

   As  the test progresses, changes in sample
weight cause an extension or contraction of
the spring which changes the positional rela-
tionship of the armature and transducer coil.
A resulting electrical signal  proportional to
the change in sample  weight  is  developed,
amplified, and fed to the vertical (Y) axis of
the recorder. The input to the horizontal (X)
axis of the recorder is proportional to time or
temperature.
 MATHEMATICAL    MODEL    AND
 REGRESSION ANALYSIS

   The purpose  of  this kinetic  study  is to
 establish  the  rate-controlling step  and thus
 rate  equations   for the  char/steam   and
 char/carbon dioxide reactions. The  study will
 also determine  the  effects  of temperature,
 reacting gas concentration, and particle size.

   The conversion of solids in a heterogeneous
 gas-solids reaction  can  follow  one  of two
 extremes. At  one extreme  the  diffusion of
 gaseous reactant into the particle is rapid
 enough, compared to the chemical reaction
 rate, that the reaction takes place  at the same
 time and at  the same rate everywhere. This is
 called  the  continuous  reaction  model.  If
                    1
 diffusion into the particle is slow, the reaction
 is restricted  to a thin shell which moves  from
 the outside of the particle inward. This is the
 unreacted  core  model  with   diffusion
controlling.

   The appropriate model may be chosen by
determining the  time  needed for  complete
conversion of solids of different  sizes,  as
summarized below:
   The parameter tc,  defined  as the  time
 needed for complete  conversion is:

   1. Independent of particle diameter for
      the continuous reaction model.

   2. Directly  proportional to particle  dia-
      meter  for the  unreacted  core model,
      with chemical reaction at the reaction
      front as the rate controlling step.

   3. Directly proportional to the square of
      the particle diameter for the unreacted
      core model,  with diffusion through the
      ash layer as the rate  controlling step.

   Experiments were carried out with particles
 ranging  in size from 70 mesh  (2\Q\nm) to
 minus  325  mesh,   (44^m).  Within  the
 temperature ranges investigated, the reaction
 rate  was  independent of particle  diameter.
 Therefore, the continuous reaction model  was
 chosen to develop a rate equation, as follows:
 the  rate  of  carbon  is proportional  to  the
 concentration of reacting gas around particles
 times the amount of carbon left  unreacted.
 In terms of the  fraction of carbon reacted, X,
 this  becomes
            dt
                                        (1)
where: Cn is the concentration of reacting gas
to some power, n.

Rearranging equation (1) gives:
Integrating equation (2) yields:
                  i = -kCnt
                  or
(3)
The proportionality or rate constant, k, may
be assumed to vary  with  temperature in  an
Arrhenian fashion:
                       -  E
             k = ae
                        RT
                                        (4)
Combining equations (3) and (4) gives:

                       "
ffl-1-4

-------
Letting X now represent the amount of char
reacted, the complete rate equation becomes:
                             .E        (5)
                              RT
(1 - X) = Ash + (1 - Ash) e ' ae     nt
where:
X   = fraction of char reacted
T   = temperature, K
G   = concentration of reacting gas
t    = time, minutes
,                   .  .       Ib C reacted
k   = apparent reactivity,  ^ c ^^ min
Ash = weight percent of ash in unreacted char
                        cal
     = activation energy,
                        mole
                                 cal
R   = the gas constant = 1.987  mole OK

n   = the order of reaction with respect to
       reacting gas
a   = the  Arrhenius   constant   (frequen-
       cy factor), mhr1

   The experimental data are in the form of a
plot of (1 - X) as a function oft. The constants
a, E, and n may then be found from equation
(5) when (1 - X), Ash, T, C, and t are known
variables. The regression analysis proceeds in
two steps. For each value of Ash, C, and T, (1 -
X) versus t  is solved for the  constant M  as
follows.
                     -W/WI
                             ,n
                    -E/RT
    Let      M = ae       CJ
                               Mt
then   (1 - X) = Ash + (1 - Ash) e

and          (1-X-Ash)   ,.,
          ln   (1-Ash)  =Mt

The square of the error in this equality is:
                                        (6)

                                        (7)
 where:  c-  is  the  error  and  the  subscript  i
 denotes individual data points.  is minimized
 by differentiating and setting it equal to zero.
              = S  tjln   1-Ash
 M =
      I tjlnd-X-AshMnd-Ash) Ztj
                                                                                      (9)
   Once M is found from the above equation
for each C and T, a multiple linear regression
approach may be used to find the constants
-E/R, a, and n by solving
          M = ae
                 -E/RT
                                                                     (Cn)
                                      (10)
                                              or    In M = ln(a - E/RT - n)ln C

                                              Again, the  partial  derivatives  of the error
                                              function with respect to M, T, and C are set
                                              equal to zero; the three resulting equations are
                                              solved simultaneously for a, -E/R, and n.
 EXPERIMENTAL RESULTS

   A series of photomicrographs gave the first
 physical clues leading to the selection of a
 reaction model. Figures 3  through 6  show
 unreacted char, 50 percent  reacted char,  80
 percent  reacted  char,  and  ash.   It  is
 immediately   apparent  that  the  average
 particle  diameter  does  not  change  with
 increasing  carbon  burn-off.  The  particles
 become increasingly porous,  but even the ash
 residue retains a skeletal structure similar in
 overall  dimensions to the  unreacted  char.
 Along  with these  physical observations, the
 experimental   data showed   that  the  time
 needed for complete reaction  was independent
of particle diameter. Rate equations were thus
developed for the six chars. For example, the
rate equation for the reaction  of char No. 2455
with steam is:
                                                               (l-Ash)e-k(CH
The significance of  this  rate  equation is
demonstrated  in  Figures  7  and  8.  The
activation energies  are of the order of 40
kcal/mole. Both the char-steam and the char-
carbon  dioxide were  found to  be approxi-
mately half order with respect to reacting gas
concentration.
                                                                                   ra-i-5

-------
   The char-oxygen reaction was also studied.
Within the temperature range of importance
to  this   investigation,  the  reaction  was
controlled by mass transfer from the gas phase
to the surface  of  the particle. Again the
particles  do  not  shrink  as  the  reaction
proceeds. Diffusion of reactants and products
into and out of these small, porous chars is so
fast  compared to the chemical reaction step
that the reaction may be thought of as taking
place  continuously  throughout  the  particle.

   The next step in  the laboratory studies was
the  construction and operation of a  small
fluidized-bed  batch reactor. Figure  9 is  a
schematic diagram of this system, and Figure
10 shows the actual equipment. The reactor
was made from 1-1/2 in. schedule 40, type 310
stainless  steel pipe.  External  heating  was
provided for operation at a maximum reactor
temperature   of  2200F.  The  system   is
designed for operation at pressures to 10 atm.

Table 2. MATERIAL BALANCE FOR FLUIDIZED-
BED BATCH REACTOR AIR-BLOWN GASIFICA-
TION TEST NUMBER 4a
Component
Oxygen
Nitrogen
Carbon dioxide
Carbon monoxide
.Hydrogen
TOTAL
Feed
Mole.%
12.4
46.4
41.2
0
0
100.0
g moles/ min
0.00846
0.03160
0.02824
0
0
0.06830
Product
Mole.%
0.12
30.72
21.45
47.71
0

g moles/min
0.0001
0.0316
0.0221
0.0491
0
0.1029
Feed gas rate: 1 530 sml/min
Product gas rate: 2300 sml/min

Total g moles Carbon
Total g moles Oxygen
Total g moles Nitrogen
In
0.02824
0.03770
0.03160
Out
0.0712
0.0467
0.03160
       a Carbon gasification rate,  0.0429 g moles/min
         Carbon Dioxide utilization, 21.7 percent
         Reactor Pressure, 68 psia
         Reactor Temperature, 940C
         Initial Charge to Reactor, 20 g char, BCR Lot 2455
   The  batch reactor tests  had  a threefold
purpose:
   1. To verify the proposed rate equations.
   2. To determine the degree  of steam or
carbon dioxide decomposition that could be
achieved in a reactor of reasonable size.

in-i-6
    3. To  provide  physical  data   such  as
 minimum fluidizing  velocity, attrition  rate,
 elutriation losses, etc.
    Table 2 is a material balance from a typical
 batch reactor test. The gasifying medium was
 air  and  carbon dioxide; At the  reaction
 temperature of 940C, the gasification rate of
 0.0429 g moles/min was consistent  with the
 predicted value of 0.0519 g moles/min.  The
 product  gas had a gross heating value  of
 approximately 153 Btu/scf.
    The success of the laboratory studies led to
 the  design  of the 100 Ib/hr  process  and
 equipment development unit.

 PROCESS AND  EQUIPMENT DEVELOP-
 MENT UNIT (PEHU)
    The purpose of the PEDU is to provide the
 necessary parameters  for the  design   and
 operation of a pilot scale  or  larger  unit to
 demonstrate the process and  the economic
 feasibility of fluidized-bed gasification for the
 commercial  production  of  low-Btu fuel  gas.
 The PEDU was designed to conform with the
 desire of the  Office of Coal Research to have a
 flexible  system  designed   with   a  nominal
 capacity of 100 Ib coal/hr. Feed to the unit can
 be either coal or char with  air or oxygen, or a
 mixture,  as  the  oxidant and  steam  and/or
 carbon dioxide as the moderator.

    Figure 11 is a schematic diagram showing
 the major process equipment. Coal is metered
 from the pressurized lock  hopper through  a
 rotary air-lock feeder and flows by gravity into
 the first reactor. Stage  1, the smallest  reactor,
 has a reaction zone inside diameter of 10
 inches and a disengaging zone of 16 inches.
 Stage 2 is the largest reactor with a reaction
 zone inside diameter of 16 inches and  a 24-in.
 disengaging zone. The reaction zone diameter
 of  Stage 3  is  12  inches  with  a  16-inch
 disengaging  zone.  All  three  reactors   are
 approximately 11 feet  high.

   Refractory-lined cyclones are provided  for
 stages 2 and 3 to recycle entrained solids to  the
bed. Solids are scrubbed from the product  gas
stream in a venturi scrubber, and the gas flows
through iron oxide boxes for hydrogen sulfide

-------
removal and thence to a thermal oxidizer for    what extent and in what manner this project
disposal.                                     will  continue.

   The next step in the development program    ACKNOWLEDGMENT
is  the construction  and  operation  of the
PEDU. Detail engineering, procurement, and       This paper is based on work carried out at
erection will take approximately 12 months. A    Bituminous Coal Research, Inc., with support
definitive  cost  estimate  is currently being    from the  Office of  Coal  Research,  U.S.
prepared for this phase of the program.  The    Department of the Interior, under Contract
Office of Coal Research will then decide to    No.  14-32-0001-1207.

-------
Stream number


Coal (ash free)
Ash
H2
CO
C02
HjO
N2
HzS
Stage 1 oil gas
Total .
Average mole wt
Temperature,0?
Pressure, psia
sclm
aclm
1
Coal leed
Ib
93.8
6.2







100.0
wttf
93.8
S.2







100.0

77



2
Char from
stage 1
Ib
72.E
6.2







78.8
wt*
92.1
7.9







100.0

1200



3
Stage 1
flue gas
moles


0.95
1.53
0.37
0.55
2.48

1.06
6.94
VOl*


13.7
22.0
5.3
7.9
35.7

15.4
100.0
23.3
1200
250
43.8
8.2
4
Stage 3
Hue gas
moles


0.95
1.53
0.37
0.55
2.48


5.88
vol 3


16.2
26.0
6.3
9.4
42.1


100.0
23.9
2100
250
37.1
10.7
5
Product gas
moles


4.7
5.16
1.04
2.00
9.14
0.05

22.1
voIX


21.3
23.4
4.7
9.0
41.4
0.2

100.0
22.3
2000
250
140.0
38.9
6
Air to
stage 2
moles









8.43
29
1000
300
53.0
7.3
7
Char Irom
stage 2
Ib
26.4
6.2







32.6
wt!?
81
19







100

2000
250


8
Char
staj
Ib
3.6
6.2







9.S
from
e3
wt*
37
63







100

2100



9
Air to
stage:
moles









3.14
29
1000
300
19.8
2.7
10
Steam to
stage 3
moles









1.50
18
1000
300
9.5
1.31
11
Steam to
stage 2
molts









2.5
18
000
300
15.8
2.17
                                PRODUCT GAS
                         ^                   51
        BASIS COAL FEED
                                              /K/h
                                              TAM I
                                                                    STAGES


                                                                     2100 F
                                                                          4-10-
     STEAM


9 AIR
                                                  AIR
               Figure 1.  Material balance for gasification with air and steam.
IIM-8

-------
                 SPRING
            TRANSDUCER
              ARMATURE
             TRANSDUCER
                COIL
     SWEEP GAS INLET
     REACTION GAS INLET

INERT GAS INLET
           CRUCIBLE
           FURNACE
                              t
DEMODULATOR  J
                   RECORDER

                  Y  AXIS

                    X - AXIS
(U.
  t
                              5
                                           GAS
                                          OUTLET
                       TIME
                   TEMPERATURE
                       SWITCH
                      I
                  TEMPERATURE
                    CONTROL
                                                      SAMPLE
                                                    THERMOCOUPLE
      Figure 2. Schematic view of thermbgravimetric analysis equipment.
                                                                           ra-i-9

-------
                   Figure 3.  Photomicrograph of unreacted char.
Ill-1-10

-------
           SHREDDED, AIR CLASSIFIED SOLID WASTE FUEL
                                                .
Figure 4.  Photographic views of the solid waste storage tank interior.
                                                                          III-l-ll

-------
                               Figure 5.  Airlock feeder valve.
III-1-12

-------
                                     I    I    I     I    I     I    I    I     I
                                    	1.0 cm	
Figure 6.  Photomicrograph of char ash  (100% burn-off).
                                                                     III-l-13

-------
i
LU
 0

10

20

30

40

50

60

70
     90
    100
                                                                            REACTING GAS: 100% STEAM
                                                                            TEMPERATURE: 1000 "C
                                                                              SOLID LINE: EXPERIMENTAL DATA
                                                                            BROKEN LINE : PLOT OF PROPOSED RATE EQUATION
                                                                           12
                                                                                       15
18
21
                                                                TIME, min
                                             Figure 7. Typical correlation of reactivity data.

-------
30
40
50
 60
                                    REACTING GAS: 100% STEAM
                                    TEMPERATURE: 900 *C
                                       SOLID LINE:  EXPERIMENTAL DATA
                                     BROKEN LINE: PLOT OF PROPOSED RATE EQUATION
                                                    TIME, min
                                   Figure 8.  Typical correlation of char reactivity data.

-------
                                             FLUIDIZED-BED
                                             BATCH REACTOR
PRESSURE REGULATOR
        o
BACK PRESSURE
    REGULATOR
                                                                                                    TO IN-LINE
                                                                                                  CHROMATOGRAPH
                         Figure 9.  Flow scheme for fluidlzed-bed batch reactor

-------
Figure 10.  Fluidized-bed batch reactor system.
                                                            HI-1-17

-------
     TO SULFUR REMOVAL
     AND THERMAL OXIDIZER
oo
                                                                                            PRODUCT GAS
           i
         COOLER'
         SCRUBBER
          A
             COAL
             FEED
             HOPPER
STAGE 1 OFF GAS


PRODUCT
GAS
COOLER
.1
1
                                                                       CYCLONE
                                                                      SEPARATOR
                                                                        STAGE 3 FLUE GAS

1
i-J
",r
STAGE 3
r
.^^^
                                                                         2100F

                                                                         \
                                                             ASH
                                                             LOCK
                                                             HOPPER
                                                                       ASH
                                    F
                                                                                       CYCLONE
                                                                                      SEPARATOR
                                AIR AND STEAM
                                                                                                    BAG
                                                                                                    FILTER
                                                           CHAR
                                                           FINES
                                                           HOLD
                                                           BIN
                                                        \y
                              Figure 11.  Fluidized-bed gasification PEDU process flow diagram.

-------
                                 2.  HOT  SULFUR REMOVAL FROM
                                                           PRODUCER GAS
                      F. G.  SHULTZ AND P. S. LEWIS
                   Morgantown Energy  Research  Center
                             U.S.  Bureau  of Mines
                       U. S. Department  of the Interior
ABSTRACT
   Sulfur-free gas for power generation or catalytic conversion to pipeline gas is needed to meet
near term energy and antipollution requirements. Gasification of coal with air or oxygen and
steam at elevated pressures  supplies the gas, but  cleaning is required  to remove sulfur and
participate matter. A stirred-bed pressurized producer is described, and results are discussed for
caking coals. Progress is reported in developing a process using a regenerable solid sorbent for
removing hydrogen sulfide from hot producer gas with recovery of elemental sulfur formed during
regeneration.
INTRODUCTION

   Gasification and gas cleanup must be con-
sidered  jointly, in  view of  today's  clean
environment regulations, because their com-
bined action is required if clean gas is to be
obtained from coal. Much of the coal sulfur
appears in the gas; in addition, solid and tar
particulates are present in concentrations that
vary with  the gasification process and coal
composition.  Gasification concepts  under-
going development include gas cleanup in the
overall processing  scheme.  Innovations  are
introduced mainly in the gasification step, and
in some cases desulfurization is incorporated
at this point.  In other cases, gas purification
could take  place after gasification, and
existing commercial systems may be satisfac-
tory. However, new technology may be needed
to meet more stringent demands.

   Probably the least complicated system for
converting coal into either low-Btu fuel gas or
high-Btu pipeline gas  is the one described
herein. It bears the suggested name MORGAS
(Morgantown  gas).  It  incorporates  pressure
gasification in a stirred bed of mine-run coal,
which may have any free-swelling index  from
low to high. Hydrogen  sulfide is removed by
contacting the hot gas with a bed of solid sor-
bent containing iron oxide; elemental sulfur is
recovered during regeneration of the  sorbent.
These two basic elements can be combined
with other unit operations as required by the
end use of the gas.
                                       in-2-i

-------
EQUIPMENT

   The  Bureau's  stirred-bed  producer
resembles  the   conventional   fixed-bed
producer except it uses mechanical deep-bed
agitation. Principal dimensions and layout are
shown in Figure 1. The fuel is supported on a
revolving grate  with an area of 9.6 ft2.  The
bed depth  varies between 6 and 7 feet; the
depth is maintained by frequently adding coal
in batches  weighing 200 to 250 lb.; cinder  is
removed as needed. The water-cooled stirrer is
balanced by a counter weight and supported
in the  pressure vessel by  a  thrust  bushing
sealed by a packing gland. Compound motion
is imparted by combined horizontal rotation
and vertical reciprocation, which can be con-
trolled with respect to speed of rotation  and
vertical movement. Figure 2 shows the stirrer
in greater detail. The two lower arms are water
cooled, but the top arm is not cooled  as  it
normally remains in a reduced temperature
zone.  Steady  gasification conditions  usually
have been obtained by rotating the stirrer at
one-half revolution per  minute and limiting
the vertical  travel through a vertical distance
of 2 feet. In practice, the stirrer passes through
the  bed in  15 minutes, but this rate can be
slower or faster, as optimum rate varies with
coal properties. The lowest point reached by
the stirrer is usually set 2 feet above the top of
the grate, but the limit of travel is within 1 foot
of the grate.

   Nuclear density gauges are used as shown
in Figure 3 to indicate  conditions  within the
pressure vessel. Ash zone, bed level, and voids
in the bed are detected.  Control of operating
conditions are simplified by the use of these
instruments;  centralized,  fully   automated
controls seem to be feasible for multiple units.

   Continuous stirring of the bed maintained
a  dense fuel  bed, giving  good quality gas
having  constant  composition.  Vertical
movement was as important as rotation for
operating the experimental producer,  but
vertical movement may not be  necessary for
full-size  units. Stirring was needed to break
large clinkers  that formed in the combustion
zone, as well as coke in the gasification zone.
As shown in Figure 4, the torque applied to
rotate the stirrer varies directly with the bed
depth covering the stirrer. At maximum depth
the  normal torque was 1,300 foot-pounds,
reaching momentary  peaks  of  1,700 foot-
pounds. Measurements were  obtained  on
Upper Freeport coal, which gives a very hard
coke. No significant difference in the  torque
load was found for double-screened 1/4- x 1-
1/2-in. Upper Freeport and run-of-mine 0- x
1-1/2 in. Gasifying mine-run coal in a  stirred
bed was a  significant  advancement because
the size  limitation,  heretofore believed
necessary, can be  eliminated.  More  of the
market supply will be available for gasifica-
tion,  and preparation  will be  less costly.  A
screen analysis of run-of-mine Upper Freeport
coal is given in Table 1. Twenty-five percent of
the sample  passed through a  1/16-in. sieve
and 5 percent through  100 mesh. Some fine
coal particles were entrained  in the gas, but
most were  removed by a cyclone separator.
Gas vented to the atmosphere and burned had
a dust loading of about 0.5 to 0.7  lb/1000 ft3.

Table 1. SCREEN ANALYSIS, UPPER FREEPORT
                  COALa
Screen size
2-1/2x2 in.
2x1 in.
1 x 1/2 in.
1/2x1/4 in.
1/4x 1/16 in.
1/16 in. x 50 mes
50 x 100 mesh
100 x 200 mesh
200 mesh x 0 in.
Analysis, %
Direct
2.5
12.1
12.2
17.6'
30.2
h 16.9
3.5
2.1
2.9
Cumulative .
2.5
14.6
26.8
44.4
74.6
91.5
95.0
97.1
100.0
      Free swelling index No. 8-1/2


RESULTS

   Experimentally determined gas yields for
moderately  caking Illinois  No.  6 coal are
shown in Figure 5. Gas production was limited
when the gas flow reached a velocity at which
loss of fuel by entrainment becomes excessive.
m-2-2

-------
This plot shows that the  quantity of  air
limiting gas  yield increases with increased
pressure.
Table 2.  SORPTION  OF H2S  FROM  DRY
PRODUCER GAS BY SINTERED IRON OXIDE-
FLY  ASH8
   A mixture of iron oxide (hematite
and fly ash was the best sorbent found among
more than twenty materials tested.  Primary
requirements were that the sorbent be readily
available  and  relatively  inexpensive,  have
reasonable sorption capacity and useful  life,
be easily regenerated for repeated use, and be
resistent to fusion or disintegration over the
useful temperature range. Fly ash as received
could  be formed  into  a  durable   and
regenerably sorbent, but its sorption capacity
was improved by adding iron oxide, increasing
the  concentration  to  36  from 15 percent
originally  present. Other oxides present  and
inactive included silica 35 percent, alumina 18
percent, and small percentages  of oxides of
calcium, magnesium, sodium, potassium, and
titanium.  Iron oxide concentrations  greater
than 40 percent were unsatisfactory because
the bed fusion temperature was lowered  and
fusion took place during  normal operations.

   Pilot quantities  of the fly  ash-iron oxide
sorbent were made by two  catalyst manu-
facturers   by mulling  and   extruding  the
mixture to form 1/4-in.  diameter cylinders
with  1/4 to 1/2-in.  lengths, which were then
sintered  to  develop hardness.  Mercury
porosimeter  measurements  showed  pore
volume of new sorbent was 0.36 cm3/g, but
this  decreased to 0.13 cm3/g and remained
constant after 30 regenerations, as shown in
Figure 6.  Surface area measured by nitrogen
absorption ranged from 4.2 to 6.5 m2/g. Sorp-
tion of hydrogen sulfide from dry simulated
producer gas is given in Table  2 for materials
of essentially the  same composition but made
by three laboratories.

   The sorbent made  by MERC was  tested
through  174  regeneration  cycles  using
simulated producer gas and bed temperatures
of 1100,  1250, and 1500 F.  Producer gas
contains about  5 to  10 percent  steam by
volume, as excess  steam is used  to reduce
temperature in the combustion zone, and the

Commercial
laboratory 1
Commercial
laboratory 2
Mechanically
formed by
MERCC
Surface
area,b
m2/gram
6.5
4.2
4.2
Bed
temperature,
F
1000
1250
1500
1000
1250
1500
1000
1500
g S removed/100 g sorbent
From sorp-
tion data
12.5
14.7
22.2
7.5
11.5
22.5
10.5
27.6
From regen-
eration data
12.4
13.9
22.0
6.7
11.0
17.4
10.9
25.6
  aAII at 3 psig sorption pressure.
  bBET nitrogen sorption method.
  cMorgantown Energy Research Center, Morgantown, W. Va.
gas leaves the generator at a  temperature
around 1200F. Steam amounting to 7 percent
by volume was added to the gas for many of
the above tests to closely simulate producer
gas.  Results obtained with  gas containing
steam, Figure 7, indicate  a  reduction  in
capacity when compared with capacities for
dry gas  as  shown  in  Table  2.  This was
attributed to  the lowering of the hydrogen
sulfide concentration at the gas-solid interface
by the  added steam. Improving  the mass
transfer   coefficient  by  raising  the  bed
temperature  was  effective in increasing  the
capacity from 6  g  sulfur/100  g sorbent at
1100F to 10 g at 1500F.

   Iron  oxide catalyzes the water gas shift
reaction, HaO + CO = fib + CO2, and steam
in producer gas affected the composition of
the producer gas  in  passing  through  the
sorption   bed.  The  composition  change
resulting  from  the  shift  reaction  was
determined at 300 psig and temperatures of
1100,  1300, and 1400 F by passing producer
gas containing 18 mole-percent steam through
a bed of iron oxide fly ash sorbent using 1000
space velocity. Heating value was decreased by
dilution from the carbon dioxide added to the
gas; increased hydrogen and decreased carbon
monoxide  concentrations resulted in virtually
no net change in  heating value because they
have  nearly the  same value,  319 and 316
Btu/scf,  respectively.  The  shift would  be
                                                                                  in-2-3

-------
 beneficial if pipeline gas  is  the end  use,
 because additional shifting would be needed
 to  bring  the hydrogen to carbon monoxide
 ratio close  to 3:1. Increasing temperature
 favors higher carbon monoxide concentration
 at equilibrium. Results are shown in Table 3.

Table3. CHANGE IN COMPOSITION AT 300 psig
AND 1000 SPACE VELOCITY

Feed gas
Effluent gas:
1100F
(dry)
1100F
lwet)a
1300F
(wet)a
1400F
(wet)8
Hydrogen,
%
17.0
16.4
22.9
21.3
19.9
Carbon
Dioxide.
%
6.7
7.9
13.2
11.0
9.7
Nitrogen,
%
50.9
51.0
51.1
51.0
51.0
Carbon
Monoxide,
%
25.4
24.7
12.8
16.7
19.4
Healing
Value.
Btu/scf
134
130
113
121
125
 aSteam content 18 volume-percent. Composition and heating
  value on dry basis.
   Two  sorption-regeneration  cycles  were
completed, and cleaning gas was generated in
the pressurized gas producer  using Upper
Freeport coal; the results are shown in Figure
8. Gas from the producer was transferred to
the sorbent bed at system pressure of 120 psig
via a heated pipeline. Bed temperatures were
controlled to give 1100 and 1200F, and flow
rates  were  adjusted to give hourly  space
velocities of 710 and 940, respectively. Hydro-
gen sulfide concentration averaged 380 gr/100
ft3; the gas  contained  approximately 0.516
dust,  1 Ib tar, and 5  Ib  steam/1000 ft3
Hydrogen  sulfide   in the  gas leaving the
sorbent bed had its concentration reduced to
10 and 20 gr/100 ft3 and did not increase until
after 6 hours  on steam. Removal was 95 and
97 percent effective with respect to hydrogen
sulfide. Tar was not removed by the sorbent.
    Reaction  mechanism  is  chemisorption,
 whereby hydrogen sulfide difusses throughout
 the sorbent and reacts with Fe2O3 forming
 FeS and  FeS2. Analyzing the spent sorbent
 indicated  the  empirical  composition  was
 FeS1-3. Iron oxide,  Fe2O3, was regenerated
 and the sulfur released as SO 2'by passing air
 or  oxygen over the  hot bed. With oxygen
 regeneration, the; effluent gas was pure  SO2
 until some oxygen passed through unreacted
 after regeneration was 90 percent complete.
 Rather than recovering  the SO.2 as sulfuric
 acid or ammonium sulfate, it appears possible
 to reduce SO^2 to elemental1 sulfur. This  may
 be done by regenerating two beds of saturated
 sorbent at the same time. Oxygen or air is
 supplied  to  one ;;bed where the  sulfur  is
 oxidized,


 and the SO2-rich effluent, free of oxygen, is
 supplied to the second bed where oxidation-
 reduction at 1500F gives elemental sulfur,

       3 FeS + 2 SO2 = Fe3 O4 + 5 S.     (2)

 Before returning the second bed to  sorption
 duty, magnetite is oxidized to  hematite, as
 follows:

 CONCLUSION

   Results indicate that hydrogen sulfide can
 be removed from producer gas by chemisorp-
 tion using sintered pellets of iron oxide-fly ash.
 Long life is indicated for the sorbent used in a
 fixed  bed.  Fluidized-   or  expanded-bed
 operation may be  possible if the pellets are
.reduced in size and shaped as spheres. The gas
 is generated  and  cleaned at pressure  and
 temperature,  thus  conserving   space   and
 energy expended  in gas compression.  The
 initial results  indicate that elemental  sulfur
 may be recovered.
ffl-2-4

-------
                                        AGITATOR DRIVE
    BED LEVEL
GRATE DRIVE
                                                            27 ft - 9 in.
        STEAM
        RUPTURE DISK
            Figure 1.  Schematic drawing of gas producer.
                                                                         HI-2-5

-------
                                     WATER OUT-*
ROTATION, time per revolution

   7 min, 13 sec SLOWEST
   1 min, 30 sec FASTEST

VERTICAL TRAVEL, ft/hr
        1.9 SLOWEST
        9.4 FASTEST
        -1/4-1..| 
        c pipe     //
           M/4-
           CS PIPE
           TUBING
           3-in. OD
           1-in. ID
                                                                 WATER IN
                                                                   PACKING GLAND
                                                                         WATER IN
                                                                           PRODUCER SHELL
                                                cm
                                     TUBING -
                                  6-1/4 -in. OD
                                  3-1/2-in. ID
                                                                          BEARING ASSEMBLY

                                                                              HIGHEST
                                                                                    OSITION
                                                                              BOTTOM SURFACE
                                                                                  6 ft-4 in.
                     SECTION A-A
1
1

l_
01
14
J 	 1-
nr
23 in.
ilr i
1 71 
in.
.J LOWEST POSITION i
12 in.

                                                                    t
                                                                            TOP OF GRATE
                                Figure 2.  Stirrer arrangement.
III-2-6

-------
                                         STIRRER
COAL FEED
                                                            COAL FEED
                                                                 1
                Bfc.^M^^        ...  .1-. n. i....i       ii..i. i  





    Figure 3.  Nuclear density gauges applied to gas producer.
                                                                            III-2-7

-------
   i'
   o
                       10
                     20           30



                               TIME, min



                  Figure 4.  Torque applied to stirrer.
                                                                             50
                              60
      160
 
 o
 S   140
     120
 ILLINOIS NO. 6 COAL

(H.V. B BITUMINOUS)
                                             3025
                      2350 LIMITING AIR FLOW, Ib/hr
                         20
                            40
60
80
                                               PRESSURE, psig


                             Figure 5. Pressure raises limiting air flow.
100
I1I-2-8

-------
 o
 111
 Q
 s.
     12


     10
s
m
o
CO
O
OC

I    4
                                                                                 60   90
                                       NUMBER OF REGENERATIONS
                 Figure 6. Pore volume reaches constant value after 30 regenerations.
                                              A.
              C.
                      BED TEMPERATURE

                         A. 1500 F

                         B. 1250 F

                         C.1100F
                           50
                                               100
150
200
                                     NUMBER OF REGENERATIONS
                  Figure 7.  Sulfur sorpfion  increases with bed temperature.
                                                                                      m-2-9

-------
       400
       300
   "S:
   exi
(A) INLET CONCENTRATION

(B) OUTLET CONCENTRATION
   SPACE VELOCITY, 940
   BED TEMPERATURE, 1200  F

(C) OUTLET CONCENTRATION
   SPACE VELOCITY, 710
   BED TEMPERATURE, 1100 F
   LU
   O


   I
      200
      100
                                                TIME, hr

                             Figure 8.  Removing H2$ from Producer gas.
HI-2-10

-------
                              3. COAL GASIFICATION FOR  CLEAN
                                                   POWER GENERATION
               D. H. ARCHER,  E. J. VIDT, D.  L. KEAIRNS,
                       J. P. MORRIS AND  J.  L.  CHEN
                   Westinghouse Research Laboratories
ABSTRACT
   The growing demand for electrical energy in the U.S. requires the construction of new coal-fired
power plants. Coal gasification, coupled with combined  gas  and steam turbine  generation,
provides a basis for  a low cost, high efficiency, non-polluting plant.  A fluidized-bed coal
gasification process adapted to power generation has been devised.  It uses air and steam for
gasification and  limestone or dolomite  sorbent for  desulfurization.  A development effort is
underway which includes the construction of a 1200 Ib/hr coal gasifier and the performance of a
supporting laboratory program.

INTRODUCTION
 Need for Power Generation

   In the  next  20  years  the  quantity  of
 electrical  energy  generated  in  the  U.S.  is
 expected to increase by a factor of almost 4, as
 shown in Figure I.1 Efforts to reduce the rate
 of growth in demand for electrical energy have
 been proposed. On the other hand it has been
 suggested   that   additional  quantities   of
 electrical energy will be required as programs
 are carried out to improve the  environment
 and  to  maximize the efficiency of energy
 utilization. It  seems prudent,  therefore,  to
 determine how the projected demands can
 best  be  met.

   Both  nuclear  and  fossil fuelscoal, oil,
 and  gaswill be  needed  to  supply this
 demand. Projections of fuel usage are made in
 Figure 2.1  Nuclear fuel usage is limited by the
 number  of nuclear power  plants which can
 conceivably be constructed in coming decades.
Natural  gas shortages in the U.S. have led  to
the prediction that its use by utilities will be
severely  curtailed in the coming decade. Coal
and oil, therefore, must provide the difference
between  total fuel demand  for electrical
generation  and  that  portion  supplied  by
nuclear fuel. An upper limit may be placed on
imported oil to avoid problems resulting from
dependence  on  foreign  nations  and  from
unbalance of payments in foreign trade. If so,
the use of coal  in power generation  must
increase by  a factor of 3 in the next two
decades.  But  if coal  is  not  available in
sufficient quantities to meet the demand, the
use of oil in power generation will of necessity
continue  to increase.

   Additional  power   plants   must   be
constructed to meet the demands for electrical
energy,  as  illustrated  in  Figure 3.1  The
generating capacity must increase from 325
GW in 1970 to 1400 GW in 1990.  Fossil fuels
will be used primarily  in  intermediate and
peak load plants.  Intermediate  plants vary
their output during each day to match the
varying demand; their overall electrical energy
                                       in-3-i

-------
 output is  about 40  to 60 percent  of  the    seasonally or in emergencies to  match the
 maximum (if the plant operated continuously    varying demands; their energy output is 2 to
 at rated capacity). Peaking plants operate only    20 percent of the maximum.
m-3-2

-------
Criteria for Power Plant Concepts

   In order  to compete  successfully with  a
conventional coal burning steam power plant
with a stack gas scrubber for SO 2 removal, an
improved  power generation system  should
have lower  capital costs,  higher operating
efficiencies,  and pollutant  emissions which
meet established requirements. Targets have,
therefore,  been  established for the  overall
economics and performance of a new power
plant concept:2

 1. Capital costs for 250 to 600 MW plants
    operating in 1975, less than $330/kW

 2. Overall operating efficiencies greater than
    39 percent

 3. Sulfur dioxide emission less than 1.2, NOx
    emission  less than  0.7, and particulate
    emission less than 0.1 lb/106 Btu heat of
    combustion   of   the   fuel;  thermal
    discharges prevented by the use of cooling
    towers.
Proposed Power Plant Concept

   An efficient, economic power plant burning
coal and providing low cost electrical energy
for intermediate or base loads can be provided
by  coupling a  coal gasification  and  gas
cleaning system with a combined gas- and
steam-turbine generation plant as shown in
Figure 4.  A  two-stage fluidized-bed process
gasifies coal using air and steam at tempera-
tures of 1400 to 2100F and pressures of 10 to
20 atm. The process desulfurizes the fuel gases
at high temperature, 1400 to 1800F, imng a
limestone  or  dolomite sorbent. The resulting
CaS is treated for disposal; or the sorbent is
regenerated for  return  to  the process, and
sulfur is recovered. Particles are removed from
the hot fuel gases  by cyclones, pebble bed
filters, or  porous ceramic filters. Most of the
fuel gases, 80 to 90 percent, flow to gas turbine
combustors where they burn with excess air to
provide hot  gases  for  expansion in a  gas
turbine. The remaining 10 to 20 percent of the
fuel gases  flow to a heat recovery boiler which
provides steam at 1200 psig and 950 F to the
steam turbine. About half the electrical energy
output from the plant is produced by the gas
turbine generator;  and  half  by the  steam
turbine generator. The gas turbine also drives
the main compressor  for air flowing to both
the combustor and  the  gasifier. A booster
compressor for the air flowing to the gasifier is
used  to overcome pressure  losses in  the
gasification and gas cleaning systems.

   Similar combinations  of coal gasification,
gas cleaning,  and combined  gas  and  steam
turbine generation have  been proposed and
explored.3-4 These differ in the type of gasifier
proposed, in  the design  of the gas cleaning
system, and in the configuration of the  power
generation system. For example, STEAG5 has
built   a plant  employing  fixed-bed  coal
gasifiers, low temperature aqueous scrubbers
(for particulate removal  desulfurization is
not included in the plant), pressurized boilers
(for combustion  of the fuel gases), and a gas
turbine expander.


GASIFICATION FOR POWER PRODUC-
TION

General Background

   Coal gasification  for  power  production
produces a clean fuel gas  with minimal
sulfur and ash  which can be utilized  either
at atmospheric pressure in a conventional gas-
fired boiler or at elevated pressure  in  a gas
turbine combustor. Gas turbine generator  or
combined gas and  steam turbine generator
plants  are  preferred  for new  installations,
because estimated costs for such plants are
appreciably lower than those for conventional
steam power plants. Gasification may also be
useful in the future preparation of a clean fuel
for  an MHD6 or fuel cell power plant.7

   Tn gasification, air (or oxygen) is supplied
to fuel in a quantity insufficient to complete
the conversion of its carbon and hydrogen to
COa  and  HaO.  A   number  of  possible
sequential processes become important. Some
                                                                                  in-3-3

-------
 of the oxygen added to the fuel reacts to form
 COz and H2O:
                O,
CO,
                                         (1)
            H2+1/2O2  -> H2O
                  (oxidation)
 These  reactions release  large  quantities  of
 heat. But unburned carbon from the fuel re-
 mains, and it reacts with CO2 and  H2O to
 form CO and H2:
               C02              CO
                 (gasification)            '2'
 These  reactions absorb  large  quantities of
 heat. Hydrogen can also react  with  carbon
 from the fuel to form  methane:
             C + 2H.
   CH,
             (hydrogasification)           (3)
 This  reaction  is  moderately  exothermic.
 Finally  the fuel,  when  heated,  can  also
 undergo
      Fuel + heat  *  C + CH4 + HC
              (devolatilization)
 where HC indicates higher hydrocarbons and
 tars. This  reaction may also yield heat.

   In a gasification process all of these pro-
 cesses can  occur simultaneously throughout a
 reactor, or each reaction may be localized in a
 region of a reactor or in  a separate vessel.
 Most  gasification  processes,  however, are
 carried out so that the heat released by oxi-
 dation, hydrogasification, and devolatilization
 balances the heat required by gasification and
 the  sensible heat of  the overall  reaction
 products. This overall  heat balance can be
 achieved by controlling the amount of air (or
 oxygen), the amount of steam, or the amount
 of an inert gas added to the gasifier. If the
 reactions are carried out in separate regions  or
 reactors, some  mechanism is  required   to
transfer heat between these regions.

   The gas composition produced by a gasifi-
cation  process  depends primarily  on  the
nature of the fuel  and on the temperature,
pressure, and gas composition in the regions
 where  gasification,  hydrogasification,  and
 devolatilization  occur. These  quantities
 determine the kinetic rates  and  the ther-
 modynamic limits of the various processes 
 oxidation, gasification, hydrogasification, and
 devolatilization.

   When H2O, CO, H2, and  CO2 coexist  at
 high temperatures, they can also undergo the
 shift  reaction:
              C0
                                            - H20
                                                                              C0.
                                     (5)
This reaction has a negligible heat effect, but
its equilibrium does affect the gas composition
according to the  relative  quantities of the
gases involved in the shift.

   In gasification,   sulfur  in   the  fuel  is
converted    primarily   to    H2S.    A
limestone/dolomite  sorbent can be utilized in
a fluidized bed to remove this pollutant  from
the fuel gases:
CaCO3
 CaO
                                    H2S
                                   CO,
                     CaS + H7O+  _2
                           Over the past ten years much effort has
                        been expended in the United States to develop
                        processes to produce pipeline gas, consisting
                        primarily  of methane.  Recently, interest in
                        gasification processes to produce a fuel for
                        power  plants  has  increased.  There  are
                        important  distinctions  between the gasified
                        coal properties required for pipeline gas and
                        those  required  for  power  plant  fuel.  These
                        distinctions are both technical and economic;
                        they have an important effect on the nature of
                        the optimum gasification process for each of
                        these  applications.

                           In  general, pipeline gas processes employ
                        either pure O2  or H2 together, with H2O at
                        pressures well in excess of 20 atm to produce a
                        produ^ high  in CH4.  Fuel  gas for  power
                        processes uses air and H2O at 20 atm or below
                        to  produce a lower-cost product.  Table  1
                        summarizes   the  differences   in   the
                        characteristic properties of  a pipeline gas and
                        a power plant fuel produced by a gasification
                        process.
in-3-4

-------
     Table 1. FUEL GAS PROPERTIES REQUIRED FOR PIPELINE GAS AND FOR COMBINED
                               CYCLE POWER PLANT FUEL

Heat content, Btu/Sft3
Pressure, atm
Temperature,F
Composition
Cleanliness
Sulfur
Particulates
Fuel cost target, $/106 Btu
Pipeline gas
~1000
>60
~70
Primarily CH4
<1ppmb
 0.01 lb/106 Btu b
0.50-1.00
Power plant fuel
> 90
10-20
70-1800
CO, H2,N2,C02, H20,CH4
1.2 Ib SC-2/106 Btu (H,550 ppm)
0.1 lb/106 Btu
0.25-0.50
             A high temperature is advantageous and may be necessary if the heating value of
              the gas is low.
           b Limits established by process requirements.
   Fuel is currently processed in three reactor
types   fixed  bed,  suspended  bed,  and
fluidized bed. In a fixed-bed reactor,  gases
pass through  a  bed of solids at a velocity
sufficiently low that the solid particles are not
blown from the bed and are not  supported by
the flowing gases. The  weight of the particles
rests primarily on other particles which  make
up the bed. A  boiler with a chain grate stoker
is one  type of fixed-bed reactor.

   In a suspended-bed  reactor, gases flow at a
sufficiently high velocity that solid  particles
are carried along with  the gases; their weight
is supported  by drag  forces exerted by the
gases. Contact between particles is limited  to
occasional  collisions. A pulverized fuel boiler
is one  type of suspended-bed reactor.

   In a fluidized-bed reactor, the gases flow
through a bed  of particles at a sufficiently high
velocity to  support their weight  but  not high
enough to carry them  out  of  the  bed.
Fluidized-bed  gasification  reactors have not
yet been applied commercially to utility power
generation. But at  least  five  fluidized-bed
gasification reactors  are  currently  under
development to produce pipeline gas and/or
liquid  fuels from  coal.  Other  fluidized-bed
gasification reactors  are  currently  under
development to produce pipeline gas from oil;
fluidized-bed  reactors   are now used  com-
mercially in  the catalytic cracking of oil,
roasting of sulfide ores, incineration of oily
wastes and  sludges,  production of organic
chemical  monomers,  making  of  cement,
conversion  of  nuclear  materials  for  fuel
elements, etc.

   Fluidized-bed  reactors  provide  the
following features in processing solids  and
gases:

 1. Ease and versatility in solids flowing and
   handling. Solid materials can  readily be
   added to or removed  from fluidized-beds.
   Gas velocities can be chosen to  promote
   particles mixing in the  bed or to cause
   separation between particles of different
   size and density.

 2. Rapid heat transfer. The free movement of
   particles in a fluidized bed promotes rapid
   heat  transfer both within  the bed  and
   between  beds.   Bed  temperatures  are
   therefore uniform and  easy to  control.
   Heat can be transferred  between beds by
   the circulation of solids.

 3. Effective gas-solid contact.  Because the
   relative velocity between  gas and  solids is
   high, exchange of mass and heat  is rapid.
   A fluidized  bed  also provides  a  large
   amount  of solid surface in contact with
   flowing gas in a relatively small volume.

   Table 2 summarizes various reactor types
for coal gasification  their  applications,
advantages,  and problem areas.
                                                                                    in-3-s

-------
                  Table   2.   REACTOR  TYPES   FOR  COAL GASIFICATION
Reactor type
     Application to
      gasification
     Advantages
         Problems
Fixed bed
Suspended bed
Ruidized bed
Lurgi, McDowell Wellman
gasifiers.
Texaco  partial  oxidation,
 BCR two-stage gasifier.
Consol  gasoline  and  ac-
ceptor,  IGT  hydrogasifica-
tion,  FMC  gasification,
Bureau of Mines synthane
gasifiers.
Developed  technology;
countercurrent  flow of gas
and solids possible.
High temperatures do not
lead to excessive agglom-
eration of coal or ash.
Versatility  and  ease  of
solids  handling,  uniform
temperature; effective gas-
solids contact.
 Maintaining  uniform  gas
 and  solids flow,  adding
 coal,  removing ash, tem-
 perature control, transfer of
 heat.
 Separating  ash solids from
 gases, temperature control;
 co-current flow.
Pretreatment  to   prevent
coal agglomeration; multi-
stage  beds  required  to
achieve    counter-current
flow.
    The process of gasifying  coal  involves a
  number of process steps including:

     Drying - The water content of the coal is
     reduced so that the particles  are  free-
     flowing and more readily transported and
     introduced   into   the   gasification
     equipment.

     Pretreatment - The coal is oxidized and/or
     devolatilized  superficially  in  order  to
     prevent  sticking  and agglomeration of
     particles.

     Desulfurization - The sulfur released from
     the coal as H^S during the  gasification
     process is  sorbed  by limestone/dolomite
     particles of the bed.

     Hydrogasification  and devolatilization -
     Volatile products are driven off the coal in
     an  atmosphere   containing   hydrogen,
     which  reacts  with  the  coal  and  char
     forming  methane   and  higher  hydro-
     carbons and releasing heat.

    Gasification - Steam reacts with the char
     (or devolatilized coal) absorbing heat while
     forming fuel gases  Ha  and CO.
                                     Combustion - Air reacts with the carbon of
                                     the  char forming  combustion products
                                     and releasing heat.

                                     The  rate  and  extent of  each of  these
                                  processes depends on temperature, pressure,
                                  atmospheric   composition,   and   time.   It
                                  appears advantageous and probably necessary
                                  to perform  groups of  these  processes in
                                  individual reactors or  in individual reaction
                                  steps  within  a single  reactor;  in  this case
                                  provision  must be  made for  the  flow of
                                  reactants  and  heat  between  reactors  or
                                  regions.

                                  Proposed   Coal  Gasification  Process  for
                                  Electric Power Generation

                                     A   proposed   improved   multi-stage
                                  fluidized-bed  coal  gasification for  power
                                  production process concept is  illustrated in
                                  Figure 5. It comprises three process units  a
                                  dryer,  a recirculating bed devolatilizer-de-
                                  sulfurizer,  and  a  fluidized   bed  gasifier-
                                  combustor.

                                     Crushed coal is dried in a fluidized bed and
                                  transported  to the  devolatilizer-desulfurizer
                                  unit.  Here  the  devolatilization,  desulfuri-
                                  zation,  and  partial  hydrogasific'ation
 m-s-6

-------
functions are combined  in a  single  recir-
culating fluidized-bed  reactor  operating  at
1300 to 1700F. Dried coal is fed into a central
draft tube of this reactor. In this tube, the coal
feed and large quantities of recycled solids 
char  and/or  lime sorbent    are carried
upward by gases from the total gasifler flowing
at velocities greater than 15 ft/sec. The recycle
solids  needed to dilute the  coal feed and  to
temper the hot inlet gases flow downward in a
downcomer  a fluidized bed surrounding the
draft tube. These solids, flowing at rates up to
100 times the coal feed rate, effectively prevent
or control agglomeration of the coal feed as it
devolatilizes  and passes through  the plastic
and sticky phase. A dense dry char is collected
in the fluidized bed at the top of the  draft
tube. Lime sorbent  is added to this bed  in
order to remove sulfur which is present as FhS
in  the fuel  gases.  (An alternative  concept
would   employ  a  separate  desulfurization
process. The fuel  gases could be cleaned  at
high temperature  in a fluidized bed of lime
sorbent;  or  they could  be  cleaned,  after
cooling, at low  temperature in a scrubber.)
Spent (sulfided) sorbent is withdrawn from the
reactor after stripping out the char either  in
the transfer line or in a separator of special
design. Char is withdrawn from the top section
of the  bed. Heat is primarily supplied to this
unit from the  high  temperature  fuel gas
produced in the total gasifier. Additional heat
can be  transported to the devolatilizer by
solids  carry-over in the gases from  the total
gasifier or by solids exchange between the two
process  units. Alternatively, additional  heat
can be generated  in  the  devolatilizer by
supplying air to the downcomers surrounding
the draft tube.

   The final gasification of the low sulfur char
is conducted in a fluidized bed with a lower leg
which  serves  as a combustor. In this section,
char   obtained  from  the  devolatilizer-
desulfurizer is burned with air at ~2100F to
provide  the  gasification  heat.   Heat   is
transported from the combustor to the gasifier
both by combustion gases and by solid? which
flow up and down between the combustor and
gasifier. The ash at this temperature agglome-
rates and segregates in the lower bed leg for
removal.  Gasification occurs in  the  upper
section of the bed at 1800-2000 F with  the
sensible heats of both gas and char being used
through solids exchange, to provide the heat
requirements for the devolatilizer-desul-
furizer.

   This   concept  has  the   potential   for
overcoming the limitations of other gasifica-
tion  processes and  providing  a lower cost
gasification system.

   Utilization of  wide variation in fuels 
   Caking coals and high ash coals can  be
   used  without  costly  and  inefficient
   pretreatment. This feature is achieved by
   employing the recirculating bed to prevent
   agglomeration  of coal particles.

   Utilization of a  wide variation in coal  size -
   The sizing of the coal to the system is not
   critical. Coal with a size range of 1/8- to
   1/4-inch x 0 can be used in the fluidized-
   bed  system.

   High   thermal  efficiency  -  Good  heat
   economy is realized through the counter-
   current movement of gases  and  solids
   between stages. The multi-stage arrange-
   ment  provides the long  residence  time
   required for high carbon conversions, with
   good control over the temperature in both
   stages of gasification. Fluidized-bed  gasi-
   fication systems also provide a means  for
   minimizing high carbon ash leaving the
   bed.  This  is achieved  in the proposed
   design by using  the  agglomerating bed.
   Fluidized-bed agglomeration of coal ash
   with low carbon loss 1 to 2 percent) has
   been demonstrated on both small and
   large scale equipment.

   Reduced heat losses - Clean fuel gas can be
   produced  without  the  heat  loss  as
   occasioned in cooling the fuel gases. The
   fluidized-bed concept permits fuel desul-
   furization by  limestone or  dolomite  at
   elevated temperature.
                                                                                   in-3-7

-------
    Although this advanced gasifier concept is
 unique, it is composed of sub-systems which
 have been successfully operated by others. For
 instance,  the  recirculating  bed  has  been
 developed by the Gas Council in England 8"13
 and utilized by others.14 The desulfurization
 step employing  fluidized char and dolomite
 has been investigated by Consol Coal,15 who
 embodied  this  idea  predominantly  for
 producing a low sulfur char; and by Esso
 (UK),16  who uses fluidized beds of lime  for
 gasifying  and  desulfurizing  oil.  Similarly,
 FMC  employs multiple  fluidized  stages  to
 produce  char22   Fuller, Chicago  Bridge,
 Battelle,19 and others  1?,18,20,21  have used
 agglomerating  fluidized-bed  combustors  in
 their processes. The use of multiple fluidized
 stages, with countercurrent flow of product
 gas, to achieve total gasification with desul-
 furization is a  logical but novel method  for
 utilizing all  the  inherent advantages of all
 these systems.

 DEVELOPMENT PROGRAM

    To  realize this gasification concept and to
 achieve  its  potential  benefits  in  power
 generation  a  development  program  is
 currently underway. It involves three parallel
 efforts:

  1. The design, construction, and operation of
    a process development plant for gasifying
    1200 Ib  coal/hr  and for  cleaning  the
    resulting gases.

  2. The planning and pursuit of laboratory
    studies in fluidization; H2S sorption and
    lime regeneration;  coal  devolatilization,
    char gasification, and ash aggolmeration.

  3. The conduct of  systems studies on overall
    power plant performance and economics.

Process Development Plant Design

   The purpose  of  the process development
plant is to provide  a means for investigating
the gasification system  the devolatilizer-de-
sulfurizer and  the  gasifier-combustor,  both
individually and in combination. The investi-
gation is to:

 1. Establish the operability of the  equipment
   over a suitably wide range of conditions 
   flow rates,  pressures, temperatures,  and
   types of coal and lime sorbent;

 2. Verify  the  suitability  of  the fuel  gas
   produced  for power  production in  a
   combined   gas  and   steam  turbine
   generator plant;

 3. Produce the data required for engineering
   scale up and economic evaluation of the
   gasification equipment for a power plant.


Preliminary plans for the plant  have been
completed. These plans include flow diagrams
(Figure 6)  material and  heat balances,  and
dimensional sketches for the  devolatilizer-de-
sulfurizer and the gasifier combustor. Special
propane burners are included with the devola-
tilizer-desulfurizer to  supply  hot reducing
gases  so  that  this  unit can be operated
independently of the gasifier-combustor.  If a
supply of  char  is  provided,  the  gasifier-
combustor can also be operated independently
of the devolatilizer-desulfurizer.

   In addition to the devolatilizer-desulfurizer
and   the  gasifier-combustor   reactors,  the
process development plant includes:

 1. Coal,   char,  and, limestone  sorbent
   receiving, storage, and  feed systems;

 2. Ash and spent sorbent removal, treating,
   and discharge systems;

 3. Primary cyclones  for removing particles
   from gases  leaving the  reactors; and

 4. Gas scrubbing and quench systems  for
   cleaning fuel  gases prior  to incineration.

Provisions  have  also  been  made  for  the
addition of other features to the development
plant including:
ffl-3-8

-------
 1. A secondary cyclone and filter to clean fuel
   gases leaving the primary cyclone of the
   devolatilizer-desulfurizer.   These   gases
   must be sufficiently clean to pass through
   a  gas  turbine with  minimal  corrosion,
   erosion, or deposition.

 2. A gas turbine combustor to burn the clean
   fuel gases efficiently with  a minimum
   production of NO.

 3. A turbine blade test  unit to demonstrate
   that  the  combustion  gases  have  been
   effectively cleaned.

 4. A lime sorbent regenerator to convert the
   spent (or sulfided) sorbent back to a form
   (a carbonate or oxide)  which will absorb
   additional  sulfur.

   Operating conditions  have been selected
for the plant. Initially a coal feed rate of 300
Ib/hr was selected as  a  basis for sizing the
plant. Heat losses from  the plant, however,
were appreciable at this scale  amounting to
20 percent or more of the enthalpy exchanged
between hot gases  and  coal in the devola-
tilizer-desulfurizer.   Electrical  heating  was
used  as a  means for  minimizing  this loss.
Some of the critical internal dimensions of the
reactors were also small   2 inches or less. It
was   decided,  therefore,  to   increase  the
capacity of the process development plant to
1200  Ib coal/hr.  Electrical  heating is not
required; minimum clearances of 3 to 4 inches
are achieved. Finally, no  increased cost of the
plant is predicted. The  reactor designs are
simplified; outside dimensions of the pressure
vessels remain unchanged. The solids feed and
discharge   systems   are  adequate   for  the
increased capacity without modification.

   The operating pressure for the gasification
system is that  required to  supply fuel to the
gas turbine combustor  a minimum of 10 to
16 atm. Current  large industrial gas turbines
use a combustor  pressure of  about 9  atm
(gauge); advanced models in the next decade
will probably use increased pressures, around
16 atm  (gauge).
   Operating  temperatures  have  been
estimated for the two reactors. The tempera-
ture for the devolatilizer has been chosen as
1600F  high enough to crack higher hydro-
carbons and thus to minimize carryover of
tars.23 This temperature also is close to that
required to maximize the effectiveness of lime
sorbents in sulfur removal.16 A somewhat
lower temperature  1400 F   may still be
effective  for   sulfur  removal   and  would
decrease the heat requirements for processing
the coal in this reactor. The temperature for
the  combustor has  been  chosen  as 2000-
2100F as required  to agglomerate  the  coal
ash.19 The temperature in the  gasifier will
depend on the effectiveness  of the solids in
transferring heat between the combustor and
the gasifier; the more effective the transfer,
the smaller the temperature difference.

   Gas velocities  in the fluidized bed  of the
devolatilizer  and  gasifier  (and  thus  bed
diameters) have been selected  on  the basis of
assumed  particle  size   distributions   and
densities to achieve  high capacities without
excessive carry over of solids. Gas  velocities in
the draft tube and downcomer of the devola-
tilizer are chosen to  achieve a ratio of about
80  to  1  of recycle solids  to coal  feed.9"13
Finally, in the combustor  gas velocities  are
chosen to minimize the quantity of char in the
agglomerating  section;  it  is expected  that
txcessive char will inhibit agglomeration.

   The depths  of various bed sections have
been chosen on the basis of various criteria.
The combustor is deep enough to complete
combustion of the char24 and to capture 80 to
90 percent of the ash.19  The gasifier is deep
enough to  react about  half of the  steam
fed.25'26 The  devolatilizer is dimensioned to
circulate the solids  at  the  desired  rate,  to
devolatilize the coal, and to remove 95 percent
of the sulfur from the fuel  gases.

   The  compositions   of   gas  streams
throughout the system have been estimated or
selected  on the  basis  of  prior  results  or
practice in coal gasification. Water gas equili-
brium has> been  assumed  to relate concen-
                                                                                   ra-3-9

-------
 trations of H2O, CO2, H2 and CO. It seems
 possible,  however,  that in  this  particular
 system greater amounts of the H2O will react
 with the char due to the fluidization of this
 material27 and greater quantities of CH4 will
 be produced due to rapid heating of the coal to
 1600F in the presence of H2.

    Material  and  heat  balances  have  been
 carried out which indicate that heat input to
 the devolatilizer-desulfurizer in addition  to
 that provided by the hot  gases  from  the
 gasifier may be  required. This heat can  be
 provided by exchange of solids between the
 gasifier and  desulfurizer with additional heat
 generated directly within the devolatilizer  by
 supplying air to the downcomer.

    Detailed design of the process development
 plant  is  now about 50 percent  complete.
 Figure 7 shows a model  of the plant. Specifi-
 cations for process vessels will shortly be ready
 to send to suppliers. The  current schedule
 calls for mechanical completion of the plant in
 October 1973 assuming that complete funding
 is immediately available. The estimated cost of
 the plant is  $4.2  million.

 Laboratory Studies of Coal  Gasification for
 Power  Generation

    To  support the  design  of the  process
 development  plant  and the  evaluations  of
 commercial  performance  and  economics,
 laboratory studies are now underway in two
 areas:

 1. Cold model studies of fluidized beds  to
    study  bed circulation,  jet penetration,
    solids  separation  and  elutriation,  and
    solids attrition.

 2. Limestone  sort/eat behavior studies  to
    study sulfur  capture, sorbent  regenera-
    tion, and  sorbent disposal.

Work will be initiated shortly  in a third area:

    Coal behavior  studies to  study  devola-
    tilization,   char  gasification,   and  ash
    pgglomeration.
   Cold model studies have been carried out in
a two dimensional fluidized bed (Figure 8) to
study the transport of particles upward  in  a
draft tube and  downward in a downcomer.
The  penetration  of  the  jet  of  particles
emerging from a draft tube into the fluidized
bed at its upper end has also been studied.
The transfer of particles from the downcomer
to the draft tube at the base of the devolatilizer
is now being observed. Correlations relating
gas  flow  rates;  particle flow rates, pressure
drops,  and  bed  dimensions  have  been
produced. These correlations  are  useful  in
designing the devolatilizer and in determining
recirculation rates within  it  from  process
measurements. Observations  are now being
carried out on bed slugging, solids separation
and elutriation, solids  attrition, and  solids
movement in  fluidized beds  simulating the
gasifier.

   Limestone and dolomite behavior has been
studied  in  a  thermogravimetric analyzer29
(Figures  9 and  10), which  measures weight
changes of small samples  of the solid as  it is
exposed to various atmospheres.  The sorption
of H2S and the regeneration of the resulting
CaS by  H2O and CO 2 have been  studied at
pressures, temperatures, and gas compositions
projected  for the process  development plant.
Both the degree of sorbent utilization and the
rate of  sorption  and  regeneration  appear
sufficiently high for the process to be practical
(Figures  11 and 12). The  oxidation of the
sulfided CaS sorbent in air to CaSO4 has  also
been studied as a means for rendering spent
sorbent inert for disposal. A problem has been
encountered  in completing  this conversion
(Figure  13). Tests  of sorbent behavior  are
continuing to  determine optimum  operating
conditions for sulfur removal and  regenera-
tion, to estimate  the number of sorption-
regeneration cycles the sorbent can  effectively
undergo, and to develop an improved sorbent
disposal process.

Systems and Economic Studies

   Preliminary studies have  been carried out
to compare the cost of the  gasification' process
HI-3-10

-------
proposed  here  with  alternative  fixed,
suspended, and  fluidized  bed processes. It
appears that cost reductions  of 20 to 40
percent can be achieved by the development of
the process. Cost calculations have also been
made for overall plants comparing  a coal
gasification, combined gas and  steam turbine
cycle plant with a conventional steam plant
using a stack gas scrubber for sulfur removal.
An electric  power plant  using the  coal
gasification process is estimated to cost 20 to
30 percent less  than the conventional plant.

CONCLUSION

   A comprehensive program is underway to
develop a coal gasification  system  for a
combined gas and steam  turbine generator
plant. Such a power plant is expected to be
lower in  capital  costs,  lower in pollutant
emissions, but equal in overall efficiency to a
conventional  power plant. The goal  of the
overall program is to  complete the  demon-
stration of such  a plant before the end of this
decade. Further improvements in  gas turbine
technology are  expected which will improve
the efficiency and reduce the costs for com-
mercial power plants in the  1980's.

BIBLIOGRAPHY

 l.Hauser,  L.G., W.H. Comtois, and  R.R.
    Boyle. The Effect of Fuel Availability on
    Future R  &  D  Programs   in  Power
    Generation.  (Presented at the American
    Power Conference. Chicago. April 18-20,
    1972.)

 2. Evaluation   of  the   Fluidized-Bed
    Combustion  Process. Summary  Report,
    Vol.   1.   Westinghouse   Research
    Laboratories. Pittsburgh, Pa. Prepared for
    Office of Air Programs, Environmental
    Protection  Agency,  Research  Triangle
    Park, N.C. under Contract  Number CPA
    70-9. October 1972.

 3. Rudolph, P.F.H. New Fossil-Fueled Power
    Plant  Process Based on  Lurgi Pressure
    Gasification of Coal. (Presented at Joint
    CIC-ACS Conference, American Chemical
   Society, Toronto. May 24-29,1970. pp. 13-
   38.)

 4. Robson, F.L., A.J. Giramonti, G.P. Lewis,
   and   G.  Gruber.   Technological  and
   Economic Feasibility of Advanced Power
   Cycles and Methods of Producing Non-
   polluting Fuels for Utility Power Stations.
   Final Report, United Aircraft Research
   Laboratories.  East  Hartford,  Conn.
   Prepared for the National Air Pollution
   Control Administration,  Durham, N.C.
   under Contract Number CPA 22-69-114.
   December 1970.

 5. Bund, K., K.A.  Henney, and K.H. Krieb.
   Combined  Gas/Steam-Turbine
   Generating Plant with Bituminous-Coal
   High-Pressure Gasification Plant in the
   Kellermann  Power  Station  at  Lunen.
   (Presented   at   8th  World   Energy
   Conference.  Bucharest. June  28-July  2,
   1971.)

 6. Way,  S.   New   Directions   in   Power
   Generation-MHD.   In:  Proceedings   of
   North  American   Fuel  Technology
   Conference, Ottawa, Canada, June 1970.

 7. Keairns,  D.L.  and  D.H. Archer. New
   Directions in Power Generation  Fuel
   Cells. In: Proceedings of North American
   Fuel  Technology  Conference,   Ottawa,
   Canada, June 1970.

 8. Dent, F.J. Methane from Coal. BCURA
   Quarterly Gazette, 42:1-14, 1960.

 9. Dent, F.J., R.F.  Edge,  D. Hebden, F.C.
   Wood, and T.A.  Yarwood. Experiments
   on the Hydrogenation of Oils to Gaseous
   Hydrocarbons.   Midlands  Research
   Stations,  The  Gas  Council,  Research
   Communication  GC37,  Inst.   of Gas
   Engineers Trans., pp. 594-643.

10. Murthy, P.S. and R.F. Edge. The Hydro-
   genation    of    Oils    to    Gaseous
   Hydrocarbons.   Midlands  Research
   Station, The  Gas Council,  Gas  Council
   Research  Communication  GC88,  IGE
   Journal, August  1963, pp. 459-476.
                                                                               IH-3.11

-------
 11. Thompson, B.H.,  B.B. Majumdar, and
    H.L. Conway. The Hydrogenation of Oils
    to  Gaseous  Hydrocarbons.  Midlands
    Research Station,  The Gas  Council Gas
    Council Research Communication GC122.
    IGE Journal, pp. 415-428, June  1966.

 12. Horsier,  A.G.  and   B.H.   Thompson.
    Fluidization in the Development of Gas
    Making  Processes.  Midlands  Research
    Station.  The Gas Council  pp. 51-59,
    March 1967.

 13. Horsier,  A.G.,  J.A.  Lacey, and  B.H.
    Thompson. High Pressure Fluidized Beds.
    Chemical    Engineering    Progress.
    65(10): 59-64,  1969.

 14. Pilot-Scale  Development  of  the  CSF
    Process.  R & D  Reports  Number 39,
    Volume  IV,  Book  3.  July 1,   1968  -
    December 31, 1970.  Consolidation  Coal
    Company, Research Division, Library, Pa.
    Prepared for Office  of  Coal Research,
    Department of the Interior,  Washington,
    D.C. under Contract Number 14-01-0001-
    310 (1).

 15. Theodore, F.W. Low  Sulfur Boiler  Fuel
    Using the Consol CO2 Acceptor  Process.
    Report Number 2.  Consol Coal. Prepared
    for   Office  of Coal  Research,  U.S.
    Department of the Interior,  Washington,
    D.C. under Contract Number 14-01-0001-
    415,  November 1967.

 16. Craig, J.W., G.L. Johnes,  G. Moss,  J.H.
    Taylor,  and  D.E.  Tisdall.  Study  of
    Chemically Active Fluid Bed Gasifier for
    Reduction of Sulphur Emissions. Final
    Report. Esso Research Center, Abingdon,
    Berkshire, England. Prepared for Office
    of Air Programs, Environmental Protec-
    tion  Agency,  Research Triangle Park,
    N.C. under Contract Number 70-46. June
    22, 1970 to March 1972.

17. Godel, A.A. Ten Years of Experience  in
    the  Technique of Burning  Coal in  a
    Fluidized Bed.  Revue  Generale  de
    Thermique. 5:348-358, 1966.
18. Godel, A.A. and P. Cosar. The Scale-Up of
   a  Fluidized Bed  Combustion System  to
   Utility  Boilers.  American  Institute  of
   Chemical Engineers Symposium  Series.
   67(116):210-218,  1971.

19. Goldberger, M.W. Collection of Fly Ash in
   a  Self-Agglomerating Fluidized-Bed Coal
   Burner.  (Presented at  Winter  Annual
   American  Society   of   Mechanical
   Engineers Meeting.  Pittsburgh.  ASME
   paper 67-WA/FU-3. November  1967.)

20. Jequier, L., L. Longchambon, and G. Van
   de Putte. The Gasification  of Coal Fines.
   J.  Inst.  Fuel. 33:584-591, 1960.

21. Jequier, L. et  al. Apparatus for  Dense-
   Phase  Fluidization. U.S.  Patent 2,906,
   608, September 29, 1959.

22. Jones,  J.F.,  F.H.  Schoemann, J.A.
   Hamshar, and R.T. Eddinger. Char Oil
   Energy Development. Chemical  Research
   and Development Center, FMC  Corpora-
   tion.   Prepared   for   Office   of  Coal
   Research, U.S.  Department  of the
   Interior,  Washington,  D.C.  under
   Contract  Number   14-01-0001-498,
   October 1966 - June 1971.

23. Consol Coal, Private communication from
   E. Gorin and G.  Curran.

24. Hoy, H.R.  and A.G. Roberts. Fluidized
   Combustion of Coal at High  Pressure.
   (Presented  at  Annual  AIChE  Meeting,
   San Francisco. November  1971.)

25. von  Fredersdorff,  C.G.  The  Reaction
   Rates  Between Carbon-Carbon  Dioxide
   and Carbon-Steam at 2000 F and 1 atm.
   I.G.T.  Research  Bulletin No. 19, 1955.

26. Blackwood,  J.D.  and  F. McGrory. The
   Reaction Rates Between the  Gas and the
   Carbon at 1600F and High Pressure (1 to
   50 atm). Australian J.  Chemistry. 11:16-
   33, 1958.

27. Blackwood, J.D.  and A.J.  Ingerme. The
   Reaction Rates Between the Gas and the
ra-3-12

-------
   Carbon at 1600 F and High Pressure (1 to
   50 atm). Australian J. Chemistry. 13:194-
   209, 1960.

28. Squire, A.M., M.J. Gluckman, R.A. Graff,
   R.  Shinnar, and J.  Yerushalmi. Studies
   Toward Improved Techniques for Gasi-
   fying  Coal. Part II:  Technical Presen-
   tation. Submitted to the National Science
   Foundation (RANN) by The  City College
   and Research Foundation of The City
   University of  New York, June 1972.

29. O'Neill, E.P.,  D.L.   Keairns,  and  W.F.
   Kittle. Kinetic Studies Related to the Use
   of Limestone  and  Dolomite  as Sulfur
   Removal  Agents  in  Fuel   Processing.
   (Presented   at   3rd    International
    Conference on Fluidized Bed  Combustion.
   Hueston Woods. October 29 - November
    1,   1972.)  (Session  I,   Paper  6,  this
   document.)
APPENDIX A

Heat and Material Balances for 1200 Ib/hr
Process Development Plant

   The heat and material balances pertain to a
mode of operation in which air is fed to the
downcomer of the devolatilizer-desulfurizer to
generate part of the heat requirements of the
vessel. Exchange of hot solids between the two
reactors  is another way to increase the heat
input to  the  devolatilizer-desulfurizer; this
mode is not compatible  with the  planned
independent  operation   of  the  reactors,
however.

   Air fed to the  downcomer reacts mainly
with carbon since fuel gas is kept out of the
zone by the flow of. fluidizing gas. The air for
combustion is  introduced  as  part of  the
fluidizing gas. Rapid, countercurrent flow of
solids and gas in the downcomer  prevents
excessive temperature rise.

   The material balance is given in Table Al
and the individual flow streams are identified
in Figure Al. Separate heat balances for the
two reactors are presented in Table A2 and
A3.

   The heating value of the fuel gas product is
dependent on the assumptions made and, in
the present example, is 123 Btu/scf. Assump-
tions regarding  stream  temperatures,  fines
elutriation rates, and transport and fluidizing
gas  requirements are incorporated  in  Table
Al.

Other  assumptions  follow:

  1. Ultimate analysis of coal:
c
H
N
S
O
Ash
74.0 wt%
5.0 wt'%
1,5 wt %
3.5 wt %
6.0 wt%
10.0 wt%
            100.0 wt %

 2. The products of devolatilization, including
decomposition of tar and oil, were estimated
as follows:

    The volatile matter includes all  of the
oxygen and hydrogen and half of the sulfur in
the coal.

    Oxygen  divides equally between  qarbon
and hydrogen in the volatile matter, forming
CO and H2O.

    Sulfur is evolved as H2S.

    The remaining hydrogen forms Ha and
CH4 in the molecular ratio of 2 to 1.*

    Carbon in the volatile matter in excess of
that producing CO and CH4 reverts to  solid
carbon.

 3. The heat of carbonization of the  coal is
small and can be neglected.

 4. Twenty-three percent of the  total  char
carbon derived from the coal is burned in the
*Bituminous Coal Research, Inc., Gas Generator Research
and Development. Survey and Evaluation, Phase Qne, Vol
Two. BCR Rept. L-156, pp. 223-5 (1965).
                                                                                 m-3-13

-------
Table A1. MATERIAL BALANCE FOR 1200 Ib/hr PROCESS DEVELOPMENT PLANT
Stream no.
Temperature, F
Solids
Lb/hr
Composition, wt %
Fixed carbon
Volatile matter
Ash
MgO-CaCOa
MgO-CaS
Gas
Lb/hr
Comoosition, mole %
N2
CO
C02
H2
H20
CH4
H2S
Air
1
77
Coal
1200

55.0
35.0
10.0


Transport
480

53.7
19.0
9.2
14.3
1.0
2.8


2
1000
Calcined
dolomite
696




100

Transport
47

53.7
19.0
9.2
14.3
1.0
2.8


3'
1600
Char
440

82.3

16.7













4
1600
Fines
560

66.7

33.3


Fuel gas
7354

50.0
17.7
8.6
13.3
7.9
2.6


5
1000
Ash
120



100













6
1600
Spent
dolomite
659




Tin
22.3
Stripping
94

53.7
19.0
9.2
14.3
1.0
2.8


7
2000
Fines
280

50.0

50.0


Gasifier
product
4680

48.6
19.5
8.8
11.4
11.2
0.1
0.4

8
1000
Char
440

83.3

16.7


Transport
88

53.7
19.0
9.2
14.3
1.0
2.8


9
1000
Fines
560

66.7

33.3


Transport
112

53.7
19.0
9.2
14.3
1.0
2.8


10
1000








Air
3113








100
11
400








Steam
741








100
12
665








Fluid-
izing
1546

10.5






89.5

-------
Table   A2.  COMBUSTOR-GASIFIER   HEAT
                BALANCE
                 (Btu/hr)
Input
Air and steam preheat
Transport gas preheat
In char and fines feed
Reaction heat: S -* H2S
C -* CO
C -* C02
H20-> H2

Utilization
Heat losses
Out in ash
Out in fines
Out in product gas


827,400
55,900
369,600
5,700
1,633,300
2,649,500
-2,138,000
3,403,400

321,900
25,800
167,100
2,888,600
3,403,400
downcomer and produces CO and COa in the
molecular ratio of 2 to 1.

 5. Recycled fuel gas, after drying to 1 percent
moisture, is used to transport solids  and to
Table   A3.  DEVOLATILIZER-OESULFURIZER
             HEAT BALANCE
                 (Btu/hr)
Input
In combustor/gasifier product gas
In fines from combustor/gasifier
In dolomite feed
In transport gas
In fluidizing gas
Combustion of carbon in downcomer
Water gas shift reaction

Utilization
Heat losses
Out in char and fines
Out in spent dolomite
Out in product gas
Desulfurization reactions


2,888,600
167,100
165,600
38,400
127,100
1,337,700
16,700
4,741,200

267,000
493,700
298,300
3,626,300
55,900
4,741,200
strip char from the spent dolomite.
 6. The concentration of CO2, CO, H2O, and
H2 in the product streams of the two reactors
are related in accordance with the water gas
shift equilibrium.
                                                                             m-3-15

-------
   cs

   >^
   o
   cr
   o
   LLl
   LU
       2.0
       1.0
          1960
1965
                                                  YEAR
                       Figure  1.  Annual  electric energy  generation in the U.S.
III-3-16

-------
3.0
    I960
1970
                                            YEAR
1980
1990
                 Figure 2.  Forecast of power generation by fuel in the U.S.
                                                                                 IU-3-17

-------
            Figure 3.  Generation additions by, fuel and type in the U.S., 1970-1990.
IH-3-18

-------
             FLUIDIZED BED
             DEVOLATILIZER/
              DESULFURIZER
    COAL-


LIMESTONE
(CaC03)   '
 10-15
  aim
 1600 F
                             HOT FUEL GAS
        CHAR
E
 120-160 Btu/scf)

  SPENT
-STONE
  (CaS)
                      FUEL GAS
     ASH
                10-15
                 atm
                1900 F
          FLUIDIZED BED
           COMBUSTOR/
            GASIFIER
               STEAM
                                                                                         HEAT RECOVERY
                                                                                             BOILER
                                                                                            STACK GAS

                                                                                                *
                            PARTICIPATE
                               REMOVAL
                                             HIGH TEMPERATURE
                                              GAS COMBUSTOR
                                                                 COMPRESSOR
                                                                     TURBINE
                                            BOOSTER COMPRESSOR
'  GENERATOR
                                                                                              AIR
    0.5 Et
                                                                                     0.5 Et
                                                                                                  ,.
                                                                                                         TURBINE
                                         Figure 4. Coal gasification - combined cycle plant.

-------
                                        j*
                                                        CLEAN FUEL GAS
                                                    -- TO GAS TURBINE











CRUSHED _
COAL

HOT
GASES 	





LIME
SORBENT ^^
CaO

DRY COAL


i
lp
:-v.;?;

J




'.*&
88
ii
":'


^,
1

/-ii
I
\
  .



\
i
^H4R
  _



> SPENT
SUKBtNl UBS



 *





V HOT FUEL GAS
T
I

RECIRCULATING BED
DEVOLATILIZER/DESULFURIZER

(

i*.

f










)
;.'.-''.'-'../ IUIML
$#:::'. GASIFIER
:.''::>'.':
i";-

J}
1
1
>| AGGLOMERATING
i COIWBUSTOR
I 	 AIR
 _ fTFftM


ASH




         COAL DRYER
           Figure 5. Westinghouse multistage fluidized-bed gasification process.
III-3-20

-------
       CONTACT    COOLING AND
       COOLER.    FLUIDIZING
                     GAS
COOLING GAS    QUENCH      CYCLONE
     COOLER  SCRUBBER
 GASIFIER
ASH AGGLOMERATION
                                                                                              CHAR
                                                                                          LOCK HOPPERS
(N) NITROGEN PURGE
    PRESSURING HEADER

(A) COOLING GAS
    RETURN HEADER

 DEPRESSURING
    HEADER
                                                                                                                CHAR FEED
                                                                                                                CONVEYOR
                                                                                                               PROCESS
                                                                                                              PREHEATER
                                TCOOLEITFUEL

         u,-rctf   /  nrowr., r      600M Btii/hr
         WATER     RECYCLE  WATER    INCINERATOR  CIRCULATING
     CIRCULATING COMPRESSOR J,RCULATNG
             COOLER
           3800M Btu/hr
                                                           PROPANE C02 STEAM
                                                     PROCESS AIR               CHAR
                                                     COMPRESSOR            STORAGE
w
to
                            Figure 6.  Westinghouse coal gasificatjon process development plant flow sheet.

-------
CO
            DOLOMITE
            CONVEYOR
                DOLOMITE FEED
                  LOCK HOPPERS
                                 DEVOLATILIZER
                                 -DESULFURIZER
                                                                                    CYCLONE
                              QUENCH
                              SCRUBBERS
                                                                                     SECONDARY

                                                                                     SPE?ARATOR
                     COMBUSTION  (FUTURE)
                       UNIT
                 ^-v (FUTURE)
                               COAL FEED
                                CONVEYOR
                 COAL FEED
                 LOCK HOPPERS
                                                                             CHAR
                                                                          DRAWOFF - POT
                           SPENT
                          DOLOMITE
                           OXIDIZER
                                                            CHAR
                                                          LOCK HOPPERS
                                                                                  FINES
                                                                             LOCK HOPPERS
DISENGAGING
  POT
                                                                                  .  STORAGE
            DOLOMITE    COAL
            STORAGE   STORAGE
        OXIDIZED DOLOMITE
          COOLER
                       DOLOMIE
                        REGENERATION
                           (FUTURE)
 CHAR
COOLER   CHAR CONVEYOR
QUENCH WATER
 CIRCULATING
                             Figure 6.  Westinghouse coal gasification process development plant flow sheet.

-------
ALUMINA/SAND
 SEPARATOR
                                                                             GAS TURBINE
                                                                             (UNINSTALLED)
      Figure 7.  View of solid waste combustor and gas preparation subsystem.

-------
                 Figure 8. Two-dimensional cold model recirculating bed.
IH-3-24

-------
Figure 9.  Thermogravimetric analyzer for high temperature and pressure reaction studies on
          3nd ch3r.

-------
                                                                                          N
                                                                                          >.
                                                                                          I CO

                                                                                          Is
                                                                                           Q>
                                                                                           1
                                                                                           CO


                                                                                           O
                                                                                           o
                                                                                           I
                                                                                           O)

                                                                                           L
IH-3-26

-------
    0.9
                     T=760C
             (    p=i*
    0.7
S   0.5
u.
    0.3
    0.1
10
15
20
25
                                                           30
35       '40
                                       TIME, min

          Figure 11. Sulfidation of calcined limestone at H2S + CaO--CaS
                                                                         IH-3-27

-------
  o
  LU
  (3
  UJ
  CE
     0.80
     0.70
     0.60
     0.50
  C  0.4
  o
  fe
     0.30
     0.20
     0.10
      0.0
T=700C   P=10atm
 H20 = C0 = 20%

-35+40 DOLOMITE
(86% Ca FRACTION SULFIDED)
         0   24   6    8  10   12   14   16  18   20   22  24  26  28  30   32   34  36  38  40

                                                TIME, min
             Figure 12.  Regeneration of carbonate from sulfided dolomite CaS+ h    +
             CaCOs + H2S.
IH-3-28

-------
N
O
X
o
o
o
                                                 ATMOSPHERIC PRESSURE
                                                     AIR
                                                     700 C
                                                          O DOLOMITE 1337
                                                               55% Ca SULFIDED
                                                            LIMESTONE 1359
                                                              99% Ca SULFIDED
                                           100

                                         TIME, sec

              Figure 13.  Oxidation of sulfided limestone CaS-f-202--CaS04.
                                                                                    HI-3-29

-------
                                                    FUEL GAS+FINES
                                                        1600F
                                DEVOLATILIZER-
                                DESULFURIZER
                        100-1000 F
             FLUIDIZING GAS 
|    
                       GAS+FINES 2000 F
                                                                  CYCLONE
                                                FUEL
                                                GAS
                                                          CHAR
                                                           1600 F
                                                SPENT
                                                DOLOMITE
                                                     1600F
         STRIPPING
           GAS
         1600F
                                 COMBUSTOR-
                                   GASIFIER
                                                                              FINES
                                                 T.G.  | COOLERS

                                                     SPENT
                                                    DOLOMITE
                                                                CHAR
                    vv
1
  FINES
v
                                                      1000F
                                                                   -*T.G.
                                                     1000F
               STEAM
       1000 F
                                                         T.G. TRANSPORT OR STRIPPING GAS
              Figure A1.  Flow diagram for 1200 Ib/hr process development plant.
III-3-30

-------
              4.   SULFUR RETENTION  IN FLUIDISED BEDS
                   OF  LIME UNDER REDUCING CONDITIONS

    J. W. T. CRAIG, G. MOSS, J. H. TAYLOR, AND  D. S. TISDALL

                Esso Ltd., Abingdon, Berkshire,  England

ABSTRACT
  Work done at the Esso Research Centre at Abingdon under EPA, OAP Contract CPA 70-46 has
amply confirmed the early promise of the oil-fired fluidised-bed desulphurising gasifier. Since the
Second International Conference on Fluidised-Bed Combustion an extensive test programme has
been carried out on a batch basis. In addition, the 7 x 106 Btu/hr continuously regenerating
gasifier, which was previously described, has been  constructed and successfully operated. The
results show that the gasifier will operate satisfactorily at lower stoichiometric ratios and at higher
temperatures than were previously established.
 INTRODUCTION

   Papers presented at the First and Second
 International Conferences on Fluidised-Bed
 Combustion dealt with the mechanisms of sul-
 phur absorption in fluidised beds of lime,
 under  oxidising and  reducing conditions.
 information was also given for the gasifying
 case, concerning the effect on desulphurising
 efficiency of variations in stoichiometric ratio,
the mean particle size of the bed material, and
bed replacement  rate in cyclic  operation.
Work  carried out  under OAP contract CPA
70-46 has enabled this information to be sup-
plemented considerably.  In this more recent
work an American limestone, BCR 1691, was
used in conjunction with a western hemisphere
fuel oil.
                                     in-4-i

-------
 Batch Test Equipment

   Two  batch  gasifiers  were  specially
 constructed  for  the programme  but  were
 essentially similar to those used previously.
 Figure 1  shows a drawing of the new reactor
 vessel with its various connections. The reactor
 is made  of mild  steel  lined with a  castable
 refractory. The lower section containing the
 fluid bed is 7 inches in diameter and 33 inches
 high.  Fuel oil enters the  reactor through a
 single 1/4-in. diameter nozzle which protrudes
 1 inch from the reactor wall at a point 5 1/2
 inches  above  the  bottom.  The upper, disen-
 gaging,  section of the reactor  contains  two
 cyclones which can be drained externally. The
 distributor is of radial form cast in refractory
 cement, with  16 horizontal holes distributed
 around  its  circumference. The  units  are
 brought into operation by underfiring  with gas
 and  kerosine.  When gasifying, product  gas
 leaves through a cyclone outlet and is flared
 outside the laboratory. A portion of the gas is
 burned in a sample flare located  above the
 reactor; the combustion products are analysed
 for SO2, O2, CO, and COa. Because of the
 wide range  of  sulphur  compounds which
 might be present in the product gas itself, this
 is the only practical method for measuring
 desulphurising efficiency.  During regenera-
 tion  similar   analyses  are  made  of  the
 undiluted gas.

 Batch Test Methods

   Two types  of tests were made,  fresh bed
 tests and cyclic tests.  A  new batch of calcined
 lime was  used  in each of the fresh bed gasi-
 fication tests. These tests were used to rapidly
 screen the effects of the following variables:
 bed  depth, gas velocity,  particle  size,  and
 air/fuel ratio.  The cyclic tests are the nearest
 simulation to  continuous  gasifier  operation
that can be obtained in batch units. The same
charge of lime is subjected to repeated cycles
of sulphur adsorption and regeneration. After
each regeneration a portion of lime is removed
and is replaced by an equivalent amount of
fresh   limestone.   The  added limestone  is
calcined to lime during the early part of the
 next gasification cycle. Without  this replace-
 ment the activity of the bed gradually declines.
 With replacement, the activity falls initially,
 but after a few cycles it reaches an equilibrium
 level determined by the replacement rate.

 Batch Test Results

   Because the fuel  handling capacity of a
 gasifier of given dimensions increases as the
 air/fuel  ratio  decreases, efforts  have  been
 made to  operate at the lowest possible air/fuel
 ratios and to  define the limitations  that may
 exist. Under  adiabatic conditions there is a
 relationship   between  air/fuel   ratio  and
 operating temperature. In  the batch  units,
 however, the operating temperature was in any
 case lower than the adiabatic level because of
 heat losses through the walls and  could  be
 lowered  still  further by  means  of the bed
 cooling heat exchanger.

   The results plotted in  Figure  2 show how
 changes   in   operating  temperature   affect
 desulphurising  efficiency  over  a  range  of
 air/fuel  ratios. All of these tests were  made
 using  different batches  of the  same bed
 material and in each case there was 5 percent
 by weight of sulphur in the bed material when
 the plotted result was obtained. The figures in
 parentheses by each point indicate the carbon
 content of the bed material. It can be  seen that
 there  is  a  tendency  for  desulphurising
 efficiency  to  fall  as  the  air/fuel  ratio  is
 lowered,  and that this tendency  is increased
 when the  operating temperature is  lowered
 from between  840 and 870C to 800C. It can
 also be  seen that  there is,  as would  be
 expected, a tendency for the carbon content of
the stone  to  rise  as the  air/fuel  ratio  is
 lowered;  this tendency, too, is  increased when
the operating temperature is lowered. It would
 not  seem  unreasonable  to infer  that the
 presence of a surface layer of  carbon reduces
the reactivity  of the stone.  Further  evidence
 supporting this view is shown  in Figure 3.  In
 this case desulphurising efficiency is plotted
 against percent  of calcium utilisation at four
 different air/fuel ratios. The figures in paren-
theses again relate to the carbon content  of the
ffl-4-2

-------
stone at the adjacent point. If the point at 0.38
percent by weight of carbon is compared with
the point at 4.68 percent by weight of carbon,
it will be seen that  although the temperature
was about the same in the two cases, 850 as
against 845C, the desulphurising efficiency at
the higher carbon  content was much lower.
The results obtained  at  an air/fuel ratio of
14.8 percent of stoichiometric and a tempera-
ture of 780C might also have been affected by
some degree of recarbonation of the stone; this
might  account for  the relatively poor desul-
phurising efficiency at very low calcium utili-
sations when  the carbon content  was also
probably quite low.

   At  this point  it  might be appropriate to
discuss the factors  controlling  stone carbon
content. It has been found that much  more
carbon than  hydrogen  is  oxidised  in the
gasifier. The injected oil cracks on the surface
of  the stone  laying down carbon which is
oxidised when the stone reaches the vicinity of
the distributor. It follows that the  amount of
carbon on the stone  is a function of the rates of
deposition and removal; the rate of deposition
reflects the rate at which fuel is injected, and
the rate of removal  reflects the  availability of
oxygen and the  relative proportions of CO 2
and CO which are  produced.

   It is evident that CO2 is the predominant
combustion product which is produced in the
vicinity of the distributor where the incoming
air meets  the carbon coated bed material.  As
the combustion products pass up through the
bed, however, there  is a tendency for the CO 2
to be reduced to CO at a rate dependent upon
the temperature and the availability of carbon.
This point is aptly illustrated  by the data
plotted in Figure  4.  These data  relating
CO/CO.2   ratio  with  temperature  were
obtained  during  the regeneration of a bed
which  initially contained  7 percent by weight
of carbon. Readings of temperature and gas
composition  were  taken  at  one   minute
intervals, and no oxygen appeared during the
period  under consideration. It  may be seen
that during the first two intervals there was a
steep   increase in  CO/CO2  ratio  as  the
temperature rose. Subsequently, however, as
the carbon content of the bed fell so did the
CO/CO2  ratio  despite   a  further  rise in
temperature.  As a  matter of interest, the
thermodynamic  equilibrium gives a CO/CO 2
ratio  of 40  at 900C.

   An  example of  the  working  of  this
mechanism during gasification is given by the
data plotted in Figure 5. In this figure, percent
weight of carbon on lime  is plotted against
duration of exposure to gasifying conditions.
The bottom curve  lines out  at  a very low
carbon content within the first 60 minutes of
operation, whereas the top two curves show a
progressive increase in carbon content over the
first 120 minutes; though both of these curves
show  a tendency for the carbon content to
reach an equilibrium level.  The  figures in
parentheses   are   the   desulphurising
effeciencies  at 120 minutes gasification time;
the 865C run gave a much better result than
the 845C run,  although the stoichiometric
ratios as well  as the temperatures were very
close  to each other, being  24.8  and 24.1
percent. It seems reasonable to attribute this
difference in performance to the difference in
carbon  content,  these  being  0.2  and  6.0
percent by  weight,  respectively.  It  may  be
deduced  from  these  results  that  it  is
advantageous to  run the gasifier at as high a
temperature as possible at low stoichiometric
ratios. Unfortunately, due to the heat "loss
through the walls of the small batch reactors it
has not been possible to approach adiabatic
conditions;   but  experience  with  the
continuously operating gasifier has confirmed
that good results are obtainable  at  air/fuel
ratios in the region of 18 percent of stoichio-
metric and  temperatures in  the  region of
870C. This is an area which will be explored
in more detail.

   The tests which gave the results which have
so far been discussed were made in beds 15-in.
deep and  at superficial gas velocities  in the
region of 4 ft/sec. In this work, the  practice
has been to discuss gas velocity in terms of the
superficial air rate, i.e., the velocity which the
air supplied  to the reactor would reach at the
                                                                                   m-4-3

-------
 operating temperature when flowing through
 the empty vessel. The actual  superficial gas
 velocity is  higher,  due to  the  presence of
 cracked  oil  products;  the  deviation  will
 increase  as the air/fuel ratio falls.

    The curve plotted  in  Figure  6 shows the
 basic  effect of variations in superficial gas
 velocity. In these tests the bed depth was 15
 inches, the operating temperature was 870C,
 the air/fuel ratio was  25 percent of stoichio-
 metric, and there was 5 percent  by weight of
 sulphur in the bed material. These data were
 obtained using a U.S. fuel and fresh beds of
 U.S. stone, BCR 1691. It can be  seen that
 satisfactory   results were  obtained  at
 superficial gas velocities up to 6 ft/sec, but at 8
 ft/sec there was a marked deterioration in
 performance. The next step was  to check the
 effect of varying the bed depth. Figure 7 shows
 the results which  were  obtained  at  a gas
 velocity of  6  ft/sec, with  all variables other
 than bed depth held at the levels used in the
 previous set of tests. In this case satisfactory
 results were obtained  at bed depths greater
 than 15  inches but a marked fall  in  desul-
 phurising efficiency occurred when  a  10 in.
 bed was used. At a gas velocity of 8  ft/sec the
 desulphurising efficiency was only 40 percent
 with a 10-in. bed, but reached nearly 100 with
 a 20-in. bed.
    The size of the stone which  was used to
 obtain these results was in the range 300 to
 3000 /xm. This size range  is obtained by de-
 dusting  the 1/8-in. diameter tailings from  a
 normal limestone quarry screening operation.
 As a matter of interest some comparative tests
 were made with narrower size range fractions
 sieved from this stone. The results of these
 tests are shown in Figure 8. It can be seen that
 the narrower  sized fractions gave  markedly
 poorer results than the material from  which
 they were obtained. The results listed in Table
 1, however, show that when the smaller of the
 two narrow cut fractions  was substituted for
 the  make-up  feed  of  a full size range bed
 under cyclic conditions, contrary to expecta-
 tions,  the desulphurising performance was
 improved. There is no clear-cut explanation
 Table 1. EFFECT OF LIMESTONE PARTICLE SIZE
 ON DESULPHURIZING EFFICIENCY IN CYCLIC
                   TEST

Size range
of makeup, /urn
Replacement rate
CaO/wts
Lined out SRE, %
Sulfur Input
Test number
T-3
(Cycles 1-14)
300 to 3175
2.38
61
3.0
T-3
(Cycles 24-31)
600 to 1400
2.37
68
3.1
for this effect, but it is possible that the single
cycle results reflected  quality of fluidisation as
well as particle size.
   In the case of fresh bed tests, the extent to
which  the stone is reacted is  not  important;
but when cyclic tests are made it is necessary
to choose  realistic levels. It  is possible  to
deduce on thermodynamic grounds that under
continuously operating adiabatic  conditions
the reaction in the sulphur content of the stone
per pass through the regenerator will be in the
region of 1.0 to  1.5 percent by  weight.  It
follows that it is not possible to operate a cyclic
test, in the absence  of progressive  coking,
which  adds  more fuel, unless  a sulphur
content higher than about 2 percent by weight
is obtained at the end of each absorption cycle.
This  somewhat  higher  sulphur  content  is
required  because  of the  additional heat
demand which occurs  in a batch  regeneration.
In continuous  operation, more  or less adia-
batic conditions obtain; in a batch operation
the  refractory  lining  must  be  raised  in
temperature during regeneration as well as the
bed, and the heat losses through the wall  of
the reactor are appreciable. Because a fair
amount of-oxidation  occurs at  relatively low
temperatures there is a  tendency for more
calcium sulphate to be formed  during batch
regenerations than would be expected under
continuous conditions. This sets the  level  of
the  sulphur  content of  the  bed  at the
beginning of the subsequent test. This level is
generally  found  to be about 2 percent  by
weight.
111-4-4

-------
   The fuel which was used for the cyclic tests
had  a sulphur  content  of 2.3  percent by
weight; for convenience,  a  standardised run
duration  of 45 minutes was  generally used.
When account is taken of  the range of gas
velocities and stoichiometric ratios which were
covered this gave sulphur inputs ranging from
1.5-3.7 percent by weight on lime  per run.

   The first series of cyclic tests were aimed at
obtaining  a  direct   comparison   of the
reactivities of BCR 1691  and the U.K. stone
which had previously been  tested.  The U.K.
stone is  about 98 percent  CaCOa,  whereas
BCR 1691 is of inferior quality, containing
only 88 percent CaCOa. Fresh bed tests had
given results indicating that the two stones
were equally effective when reacted to the
same degree. Cyclic  tests, however,  revealed
that BCR 1691 is so inferior in performance
under these conditions that it requires about
three times the stone replacement  rate for
an equal desulphurising performance in a bed
15.5-in. deep. Fortunately, it was also found
that a modest increase in the depth  of the bed
gave an  improvement in performance  which
was  quite disproportionate, allowing the stone
replacement rate to be substantially reduced.
An indication of the nature of the relationship
between  bed depth,  stone replacement rate,
and  desulphurising efficiency is given by the
data shown in Table 2. All of these tests were
made at 870C with  25  percent of stoichio-
metric air and a superficial  gas velocity of 6
ft/sec.
Table2. EFFECT OF BED DEPTH AND STONE
REPLACEMENT IN RATE DESULPHURIZING
       EFFICIENCY IN CYCLIC TESTS
Bed depth.
inches
15.5
20.0
20.0
22.0
Moles CaO
per moles S
1.39
1.43
2.9
1.4
Lined out
S.R.E.,%
61
90
98
97
   It will be seen that increasing the bed depth
from  15.5 to 20 inches improved  the desul-
phurisation efficiency from 61 to 90 percent at
roughly the same  stone replacement rate of
approximately 1.4 mole CaO/mole sulphur. In
order to  improve  the  desulphurisation
efficiency to 98 percent with a bed depth of 20
inches, it was necessary to increase the stone
replacement  rate  to  2.9  mole  CaO/mole
sulphur.  When, however, the bed  depth was
increased to  22   inches then  a  1.4  mole
CaO/mole  sulphur stone replacement  rate
gave  a  desulphurisation efficiency  of  97
percent.  It  remains to  be seen whether BCR
1691  will give a similar performance in the
continuous gasifier at a superficial gas velocity
of 6 ft/sec.
Results Obtained Operating the Continuous
Gasifier

   The  continuously  operating gasifier has
been described  previously; a  detailed
discussion   of   its   construction   and
commissioning is not within the scope of this
paper. A general view of the installation,
however, is  shown in Figure  9. During the
commissioning period three runs were made
giving a total operating time under gasifying
conditions of about 460 hours. U.K. stone was
used  in  these tests,  but with  U.S.  fuel
containing 2.5 percent by weight  of sulphur
and 350 ppm of vanadium. The prime purpose
of  these  runs  was . to  demonstrate  fuel
gasification  with  sulphur  removal  on  a
continuous basis; the study also took a quick
look at the effects of some of the controllable
variables. The study showed that the gasified
fuel  ignites readily and  burns with a bright
luminous stable flame. Smoke  free operation
was obtained with about  1.5 percent oxygen in
the flue gas over long periods. This is a better
performance   than  that  given   by  the
conventional  oil burner which the gasifier
replaced. When tested prior to conversion it
was found that the package boiler required
about 3 percent oxygen in the flue gas for
smokeless operation.

                                    m-4-5

-------
   The operating conditions covered  during
the  test runs are indicated in Table  3.  In

Table 3. CONTINUOUS GASIFIER OPERATING
              CONDITIONS
                               Table 4. GASIFIER PERFORMANCE
Programme item
Number of test
Test duration
Limestone used
Oil used
Gasifier temperature
Regenerator temperature
Bed depth
Superficial gas velocity
Lime particle size range
Lime replacement rate
Air fuel ratio
Oil feed rate
Pilot plant operation
3
91-202 hu-
ll.K., Denbighshire
Venezuelan 2.5% S
820-920 C
1050-1100  C
13-23 in.
2.8-4.3 ft/sec
300/3200-800-320% m
0.54-4.8 mole CaO/moleS
15-31 % of stoichiometric
61-82lb/hr-ft2
 general the superficial gas velocity was about 4
 ft/sec;  when  the fuel  rate  was  varied  the
 operating  temperature  was  controlled  by
 recycling  flue gas.  During a  considerable
 proportion of  the  operating  time  desul-
 phurisation was virtually complete. Although
 this result was  highly gratifying, it  did  not
 yield much information concerning the effects
 of the independent variables. In one period of
 19 hours duration in which virtually  no SO2
 was detected in the flue gas,  the  running
 conditions were  as  shown in Table  4. The
 gasifier  temperature was  about 900C,  the
 pressure drop through the bed averaged 14.5
 inches water gauge the air/fuel ratio was 23
 percent  of  stoichiometric,  and  the  stone
 replacement  rate   was  about  1.4  mole
 CaO/mole sulphur  entering the bed.  The
 superficial gas velocity averaged 3.7 ft/sec. In
 another period of 25 hours duration at the end
 of the test, the  operating temperature was
 about 880C, the pressure drop through the
Duration of
experiment, hr
Gasifier temper-
ature, C
Air/fuel ratio,
% stoichiometric
Bed pressure drop,
in. H2O
Gas velocity, ft/sec
Stone replacement
rate, moles CaO/moles S
Sulphur removal
efficiency, %
19

900
23

14.5
3.7
1.4
100
25

880
22

19.5
3.9
0.85
95
                         bed was about 19.5 inches water gauge the
                         air/fuel ratio was 22 percent of stoichiometric,
                         and the desulphurisation efficiency averaged
                         95 percent. During the first 16 hours of this
                         period the stone replacement rate was about 1
                         mole  CaO/mole   sulphur,  during  the
                         remaining 9 hours the Ca/S ratio was (X6 of
                         stoichiometric giving an average figure for the
                         25 hours of 0.85 mole CaO/mole sulphur. In
                         the last  hour the  air/fuel ratio fell  to  18
                         percent of  stoichiometric, but  the  desul-
                         phurising  efficiency  never  fell below  91
                         percent.

                            The  operational problems  which  were
                         encountered during these runs, which totalled
                         460 hours under gasifying conditions, were of
                         a minor nature and remedial action has since
                         been taken. In view of the results which were
                         obtained there is  no doubt  at  all that the
                         process  is a feasible proposition.
m-4-6

-------
WATER COOLANT
CONTROL VALVE

GAS SAMPLING POINT.


MANOMETER TAPPING
THERMOCOUPLE

FILL POINT
SAMPLING FLAME BURNER
                FLARE
BED COOLING LOOP
      NITROGEN
         IMIM^
              1

          r**
BED SAMPLE POINTS


BED DRAIN

DISTRIBUTOR PLATE



MANOMETER/THERMOCOUPLE
                                                   TWIN GAS VENT PIPES

                                         TWIN CYCLONE
                PILOT BURNER
                                         REFRACTORY LINING
MANOMETER TAPPING

MANOMETER TAPPING
THERMOCOUPLE
FUEL INJECTOR
THERMOCOUPLE
MANOMETER TAPPING
IGNITER AND GAS INLET
AIR INLET
                                                                   SECTION ON A-A
                           Figure 1. CAFB batch unit reactor.
                                                                               IH-4-7

-------
  on
  ce.
                                         5 % BY WT SULPHUR IN BED

                                                D 840 to 870 C
        80
                                                                                             35
                                           STOICHIOMETRIC AIR, %
                  Figure 2.  Interaction between air/fuel ratio and bed temperature.
III-4-8

-------
    100
     80
Q
Q
E    *
      40
                                                                          0.38)
                                                                             (4.68)
                                       , (11.88)
                             10
20
30
40
                                         CALCIUM UTILISATION, %
 Figure 3.  Result at 6 ft/sec in number 4 units ( sulphur removal curves at different air/fuel
 ratios).
                                                                                        III-4-9

-------
                                        TEMPERATURE,^
               Figure 4." CO/CO? profile during regeneration, fresh bed test no. 7.
IH-4-10

-------
o
00
      0
                                               (83.6)
                                               (85.3)
                                        T,C
                                        810
                                                            845
                                                               865
                                       % STOIC.
                                         20.5
                                                      24.1
                                                       24.8
        0
40
80
200
240
                                    120          160
                                   GASIFICATION TIME, min
Figure 5.  Change in carbon content of lime during batch gasification cycle 1-A unit.
280
                                                                                          ffl-4-11

-------
      100
                           5 % BY WT SULPHUR IN BED
       40
         4.0
5.0               6.0               7.0                8.0
            SUPERFICIAL GAS VELOCITY, ft/sec
   Figure 6.  Basic effect of superficial gas velocity.
9.0
III-4-12

-------
a
OC
    80
                         5 % BY WT SULPHUR IN BED
      70
      60
        10.0
                15.0
         BED DEPTH (STATIC), in.
Figure 7.  Basic effect of bed depth.
20.0
                                                                                        IH-4-13

-------
   O
   cc
   Q-
        40
        20
D  300 to 3175 Ji

  600 to 1400 JJ

O 1200 to 3175 ji
                                       10
            20
                                              Ca UTILISED, %
                              Figure 8.  Basic effect on particle size range.
30
m-4-14

-------
       5. THE  CO2 ACCEPTOR  GASIFICATION PROCESS-
                             A STATUS REPORT-APPLICATION
                                               TO BITUMINOUS COAL
                       G. P. CURRAN AND E. GORIN
                        Consolidation Coal Company
ABSTRACT

   This paper discusses experience gained and problems encountered during startup operations of
the Rapid City pilot plant. A project schedule is given for completion of that phase of the work
aimed at production of pipeline gas from low-rank western coals.

   Process revisions that must be made in application of the CO2 acceptor system to high-sulfur
bituminous coals are discussed. The major revisions are  installation of pretreatment facilities to
handle caking coals and an increase in gasification temperature to accommodate the poorer
reaction kinetics.

   Experimental work on the pretreatment of bituminous coals to render them  suitable for
pressurized gasification by preoxidation. Highly fluid coals such as Pittsburgh seam do not.
Promising results are reported via a "Seeded Coal" type process.

   A revised flow sheet and heat and material balance is given for application of the CO2 acceptor
process to the processing of bituminous coals. Recycle-of CO 2 is a key feature in this operation.
INTRODUCTION

   The CO 2 acceptor process has been under
development for a number of years. The major
goal of this work has been the production of
pipeline gas from low-rank western coals. The
process has  been extensively described in
numerous  previous  publications  and  no
description  is  deemed  necessary here. A
relatively complete description of the technical
basis of the process and its economic potential
is available in reports to the Office of Coal
Research.
   A pilot plant to test the process has been
constructed at Rapid City, South Dakota. The
project is financed jointly by the Office of Coal
Research and the American Gas Association.

   The purpose of this paper is to give a brief
status report on the Rapid City operations and
of the contemplated development schedule. It
also  discusses  problems  and opportunities
involved in the application of the process to
treatment of bituminous coal. The use of bitu-
minous oals in the process  is not only of
                                     in-5-i

-------
  interest for production of pipeline  gas, but
  more broadly for the production of low-sulfur
  boiler fuel.
    The production of low-sulfur boiler  fuel
  from bituminous coals by an adaptation of the
  process to produce a low-Btu gas without2 and
  with low-sulfur char3 as co-product has been
  studied in the  course of a research  contract
  between Consolidation Coal  and the EPA.
This work is discussed in more detail in a suc-
ceeding  paper at  this  conference.  It  also
formed the subject of a paper presented at the
Second International Conference on  Fluid-
ized-Bed Combustion. This paper gives a brief
resume of some work now being conducted for
the EPA in the pretreatment of caking bitumi-
nous coals to establish  operability  in pres-
surized fluid-bed gasifiers.
ra-s-2

-------
STATUS  OF PILOT PLANT  DEVELOP-
MENT

   The construction of the pilot  plant  was
completed with formal acceptance of the plant
on December 28, 1971. Mechanical testing of
the various plant components occupied the
next period through the end of March 1972.
During  this test period,  a number of  unit
operations   were  successfully  carried  out.
These   included  operation  of   the  gas
purification, char grinding, and lockhopper
systems. The fired heaters were put on stream
and the  process vessels  were successfully
heated to 1400-1600F by hot gas circulation.

   Pilot plant operations since April 1972 were
aimed at initiation of an actual gasification
run using  lignite char as a feedstock.  The
initial  run  was  chosen  to  demonstrate  a
simplified two vessel version of the process to
be conducted at 150  psig. The system to be
demonstrated is illustrated  by two vessels
shown  in   Figure  1    the gasifier  and
regenerator.   The  details  shown  on  the
remainder   of the  flow  sheet  should  be
disregarded  at  this  time since  they  refer
specifically  to future  operations  with
bituminous coal which will be discussed later.
It should also be noted that in the pilot plant
runs  the  temperature in the  gasifier  was
programmed  for lower  temperatures than
indicated  in Figure 1, i.e., at 1520F.

   One of  the  unique  features,  from the
engineering point of view, in the COa acceptor
process is the dual fluo-solids handling system
wherein acceptor, which originally  is  either
dolomite or limestone, is  fed to fluidized beds
of char. The acceptor particles are bigger and
heavier than the char particles and shower
down through the char bed. They collect as a
separate and segregated fluid bed of acceptor
in the boot at the bottom of the gasifier from
which  the  acceptor  is  recirculated to the
regenerator.

   The  ability  to maintain  a  segregated
fluidized acceptor bed reasonably free of char
is one of the key elements in achieving a
successful demonstration of the process.
   This feature has been well demonstrated in
the prepilot scale work at Consolidation Coal
Company's Research Laboratories, but one of
the purposes of the Rapid City pilot plant is to
demonstrate that  this  operation  can  be
successfully scaled up.

   Acceptor circulation tests were carried out
in April,  May, and June 1972 between the
gasifier  and   regenerator   preparatory  to
initiation of the gasification  tests. Operations
in June were hampered by the June  9 Rapid
City flood and its aftermath. Difficulty was
encountered  in these tests  due to  chronic
plugging  of the  pressure  probes  used to
control the operation. The plugging was due in
part to an inadequate purge gas system.

   The situation  was rectified by increasing
the diameter of the pressure probes from 1/4
in.  to  1/2 in., installing a rod out system to
break  plugs, installing duplicate probes at
critical measuring points, and improving the
purge  gas supply system.

   Successful hot continuous circulation of
dolomite-based acceptor  was  demonstrated
for a period of 14 hours during the end of this
period. The MgCO3 portion of the stone was
calcined to MgO  and the stone circulated in
the "half calcined"  condition.  The  run was
terminated  involuntarily due to  a  pressure
upset  and  corresponding loss  of  pressure
balance between  the two vessels.

   One of the  problems encountered during
this period was a  high rate of attrition of the
acceptor in its soft, half-calcined condition.
This led to excessive generation of fines which
caused plugging difficulties at the entrance to
the quench  towers.

   The softness of the stone  is of a transitory
nature since it is known from  our prepilot
experience that it hardens rapidly when it is
cycled  through the actual process operations
which   were not attempted during  these
circulation tests.

   During July  a successful  hot  acceptor
circulation  test  was  performed,  and
preparations were made  again to  start an
                                                                                   in-s-3

-------
 actual gasification run. Beds of half-calcined
 acceptor and  char were  established in the
 gasifier and regenerator, respectively. Circu-
 lation of acceptor through the char  bed was
 initiated, but  difficulty was experienced in
 obtaining a distinct char-acceptor interface in
 the gasifier boot.

    The acceptor-char mixture was transferred
 via the lift leg to the regenerator. The  acceptor
 was, however,  rich in  char, and since the lift
 gas was nearly pure air combustion of the char
 in the lift line  was  initiated. This caused
 development of a hot spot in the line which
 resulted in its  rupture and termination of the
 operation.

    The above incident is not typical of normal
 operation since the lift gas usually is recycle
 regenerator  offgas   substantially   free  of
 oxygen. In this case air was present because
 circulation   was  started   before   char
 combustion in the regenerator was initiated.

    A  fourth startup  was  initiated after the
 ruptured lift  line was repaired  in  August.
 Difficulty occurred at all times, however, with
 blockages in the  acceptor and char  transfer
 lines.   It  was   obvious   from  both   the
 temperature and pressure  profiles that  fluid-
 ized beds of acceptor in the gasifier boot and
 char in the gasifier was not being maintained.
 In  spite of  this, sufficient  acceptor  was
 transferred through the lift line  to establish
 the regenerator bed, and sufficient char was
 fed to the regenerator to establish combustion
 therein. The  regenerator  was increased  to
 1700F. The unit was shut down on August 20
 for inspection when the char feed line to the
 gasifier plugged. This  was  done to determine
 the  cause for  the failure to establish  the
 desired fluid beds  in the gasifier.

   Inspection of the gasifier revealed that the
fluidization difficulty was due to  a failure of
the refractory in the gasifier boot,  particularly
at the seam between the head and the gasifier
shell, which allowed gas to bypass the solids
bed through massive holes and cracks in the
refractory.
   The gasifier was taken down and a new
refractory configuration installed. Specifically,
the outer layer of soft insulating refractory was
removed from the gasifier boot to be replaced
with harder castable refractory. The unit, at
the time of this writing, was scheduled to be
put back in operation again early in October
1972.

   The initial operating difficulties have not
been  related to fundamental deficiencies  in
the process itself. It is expected that successful
demonstration of the  COa  acceptor process
will be achieved in the Rapid City pilot plant
in the near future.

   The original schedule has been set back
about  6 months,   first  by some delays  in
completion of construction and second by the
startup difficulties  outlined  above.  A  new
operating schedule has been  drawn up and
submitted to the Office of Coal Research for
their  approval.  The new schedule calls for
completion of the original pilot plant program
in July of 1974. The contemplated program,
however,  does  not  encompass testing of
Eastern bituminous coals.

   The processing of bituminous coals is  of
interest not only for production of pipeline gas
but also for production of low-sulfur boiler
fuel. A series of new problems are introduced
when the process is adapted to use of caking
bituminous coals. Extension of the pilot plant
operating period as well as some modification
of the  equipment would be required to study
bituminous  coal processing.

PROCESSING OF  BITUMINOUS  COALS
VIA CO2 ACCEPTOR PROCESS

Pretreatment via Preoxidation

   The use of fluidized-bed technology for the
gasification of caking coals requires  that the
feed coal be pretreated to render it non-caking
in order to  sustain an  operable bed.  The
problem becomes more severe as the operating
pressure is  increased and  may also  be a
function  of hydrogen  partial pressure.  The
specific role of hydrogen partial pressure  as
m-s-4

-------
distinct  from  total  system  pressure  in
intensifying the pretreatment problem has not
been fully defined.

   The effect of increasing total pressure is
illustrated   by  the  two   experimental
observations outlined below. A highly caking
Pittsburgh  seam  coal  was  successfully
processed in 1949 in an atmospheric pressure,
1 ton/hr fluid-bed gasification unit without
any pretreatment.5 The above admittedly was
accomplished at a relatively low coal through-
put rate of 25 lb/hr-ft2, but the  effect of
higher rates was not explored.

   The  other  observation  was,  that  in
processing non-caking sub-bituminous  coals
at 20  atmospheres  pressure  in  the hydro-
devolatilizer of the  CO 2  acceptor  process,
agglomeration  of the bed solids  occurred
unless the coal feed  was pretreated  by mild
preoxidation.10

   The work  on  the development  of the
synthane  process  at the  USBM6  again
illustrates the fact that successful operation of
a  pressurized fluidized-bed gasification
process with bituminous coal requires that the
feed be pre-treated by preoxidation.

   A  study  of the  degree  of pretreatment
required for pressurized fluid-bed gasification
of two types of bituminous  coals has  been
carried out for the EPA  One coal was from
the Ireland  Mine in northern West  Virginia
and is typical of the  highly fluid, high sulfur
Pittsburgh  No. 8 seam. The other coal was
from the Hillsboro Mine in central Illinois and
is"  typical of the more weakly caking,  high
sulfur  Illinois No. 6  coals. The experimental
investigation was carried out in the same 4-in.
ID reactor previously used in the development
of the CO2  acceptor process.1 c  The experi-
mental method is described in detail in the
Annual Report  to  the EPA. Only a  brief
summary of results will be presented here.

   The gasification  conditions  chosen for
testing operability of the pretreated coals are
those outlined in Table 1  and correspond to
conditions selected for adaptation of the CO
acceptor  process  to  produce  low-Btu  gas.
Under normal conditions in the CO2 acceptor
process, i.e., for production of pipeline gas, a
higher partial pressure of hydrogen prevails
and even more severe pretreatment may be
required.

   The severity of  preoxidation  conditions
required to establish  operability for the two
coals in the gasifier operated at the conditions
cited in Table 1 are given for the case of a 28 x
100 mesh feedstock in Table 2.

Table 1.  TYPICAL CONDITIONS  USED FOR
TESTING   OPERABILITY  OF   PRETREATED
            COALS IN GASIFIER
Temperature, F
Pressure, psig
Feed rate,lb/hr
Feedstock


Fluidizing velocity, ft/sec
Input, scfh
Stearn
C02
N2
Air
Percent carbon burnoff
Mean particle density,
Ib/ft3
1700
206
4.83
28 x 100 mesh,
Pretreated,
Hillsboro Coal
0.34

77
38
103
105
55
43.2

1700
206
5.50
20x1 00 mesh
Pretreated,
Ireland Coal
0.33

78
35
111
105



Table 2. MINIMUM CONDITIONS OF SEVERITY
OF  PREOXIDATION TO  PROVIDE OPERABLE
           GASIFIER FEEDSTOCK
Coal
Temperature of preoxidation, F
Size consist of coal feed
Stages of preoxidation
Oxygen consumed, wt %
(referred to raw coal)
Stage 1
3 Stage 2
Total
Ireland Mine
750
28 x 100 mesh
2

18.6
9.3
27.9
Hillsboro
810
28x100 mesh
1

8.7
-
8.7
   Two criteria are  used to  evaluate  the
impact  on economics of  the  results of the
preoxidation tests. The first is the percent pre-
oxidation  required as  compared with  the
"adiabaticr" quantity, i.e., with the amount of
                                                                                  in-s-5

-------
 oxidation  by heat  balance  to  sustain  the
 reaction at the  desired temperature.

   This relationship is illustrated in Figure 2.
 It is readily seen that the  amount of pre-
 oxidation required for Ireland  Mine  coal is
 four times  the  adiabatic quantity at  750F.
 This is a strong  economic debit since, in order
 to carry out this process in practice, large
 amounts of heat must be removed from the
 preoxidizer. Somewhat larger amounts of pre-
 oxidation  are  permitted  by   adiabatic
 operation if the temperature in the preoxidizer
 is allowed  to rise. However, in the case of
 Ireland  Mine   coal,  substantially  higher
 preoxidation temperatures are precluded since
 the  preoxidizer  itself becomes inoperable.

   The demonstrated  preoxidation  severity
 required for the  Illinois No. 6 coal, however, is
 only slightly above the adiabatic  level. As  a
 matter of fact, lower extents of preoxidation
 may in fact  be permissible in this case since an
 investigation of the lower limit of preoxidation
 was not carried out.

   The other desired property of the  pre-
 oxidized  coals  relates  to  the  fluidization
 behavior. In order to operate the gasifier at a
 practical throughput and for the preoxidized
 coals to have a relatively high particle density,
 it is necessary to use a relatively coarse feed.
 These  properties permit  operation  of  the
 gasifier at reasonable gas velocities without
 excessive entrainment,  and they maintain  a
 reasonably high bed  inventory to satisfy  the
 demands of the gasification kinetics. It is seen
 from the data  in Table 3  that  significant
 particle swelling occurred in the preoxidation
treatment of both coals.

   Reduction in the amount of preoxidation
required  for  Ireland  Mine  coal can  be
accomplished by use of finer  coal. For  the
reasons cited above,  however, the  use  of fine
coal  is economically undesirable.

   The conclusion from  these  preoxidation
studies was that  Illinois  coals  may  be
pretreated   successfully  with   pressure
gasification  by   use  of  "adiabatic"
  Table 3. PROPERTIES OF COAL AND PRE-
              OXIDIZED COALS
Coal
Treatment
%Sulfur
Ash
Mean diameter, inch
Mean density, tb/ft3
Pittsburgh seam
Ireland Mine coal
Raw
4.52
11.36
0.0165
81.0
28.3%
Preoxidation
at750F
3.75
13.76
0.0156
56.4
Illinois No. 6
Raw
4.93
21.86
0.0176
80.0
8.7%
Preoxidation
at810F
3.40
15.39 a
0.0175
51.8
 aAsh is low because of segregation and selective removal of
  mineral matter in the preoxidizer.
preoxidation.  Highly fluid  Pittsburgh seam
coals, however, require economically excessive
amounts  of  preoxidation,   unless  an
impractically small size consist feed coal is
used.

Pretreatment via "Seeded Coal Process"

   The principle of the preoxidation method
of pretreatment is to convert the coal to a more
rigid  structure via oxidation, such that the
fluidity  is  severely reduced when  the  coal
undergoes pyrolysis.

   The "Seeded Coal Process" would operate
on  just  the  reverse  principle  and  actually
utilize the natural fluidity of the coal. In the
process visualized, char would be circulated at
a high rate by  means of a lift gas through a
draft tube immersed in  a  normal  fluidized
bed. Coal and fine size seed char would be fed
into the  draft  tube. The external  fluid bed
would be maintained at 1000 to 1400 F either
by injection of air or hot fluidizing gas from a
gasification step,  as shown  in Figure 1.

   The  coal  melts,  smears  out  over  the
surfaces  of the seed  char and  external bed
material, and then solidifies on completion of
pyrolysis.

   The demonstration of such a device7 was
successfully   carried   out   in   the   low-
temperature carbonization section of the CSF
Coal Liquefaction Pilot Plant at Cresap, West
Virginia. The feed material, in this instance,
was  somewhat  different and constituted the
ra-s-6

-------
underflow from  the  hydrocyclone  separation
of the extraction effluent. Coal extract in this
case was used instead of the fluid coal; the
extraction residue was used instead  of the seed
char. Other differences were that the mixture
was sprayed into the draft tube as a slurry and
operating temperatures and pressures  were
lower, i.e., 825 to 925F and approximately 4
psig, respectively.

   In this particular installation, a 36-in.  ID
carbonizer was employed in which there was
installed a 6-in. ID  x 11-ft high draft  tube.
Solids were circulated through the draft tube
by injection of about 3500 to 4500 scfh of lift
gas into the bottom of the tube. The feed was
sprayed into the circulated char stream within
the draft tube  by means  of a nozzle 3 feet
above the lift gas injection point.

   Solids circulation  rates of  the order of
100,000  Ib/hr were  achieved in this device,
while complete  operability and  product size
consist control  was maintained with extract
feed rates up  to 200  Ib/hr.  The  ratio of
extraction residue solids to extract was in the
range  of about 1.5:1  to 3:1. The  above
throughput rates do not necessarily represent
the capacity of the system since high extract
feed rates were not available and consequently
were not tested.

   The  above results led to an attempt to
apply the same system to  coal  even though
coal is a less fluid material than extract and
the  operating  conditions,  particularly the
pressure, required are higher.

   Tests  with an inert bed  of 48 x 100  mesh
char at 1500 F  and 15 atm system pressure
showed that the external baffle was effective.
The solids circulation rate upward through the
tube was measured by substituting a known
amount of air for some of the N2 entering the
solids   feed   line.1C From   the   measured
temperature  rise,  the solids flow rate was
calculated  as 900 Ib/hr  by  heat  balance.
Calculations involving the pneumatic transfer
line   model, devised in the course  of develop-
ment of the CO2 acceptor project, showed that
without the external baffle about  270 of the
340 scfh of N2 fed to the bottom of the exter-
nal bed had entered the draft tube;  with the
baffle, the flow was reduced to about 60 scfh.

   Seven tests were made with the modified
draft tube, using an external bed of 48 x 100
mesh  char  at   15  atm  system  pressure.
Common conditions for the runs are listed
below:

Ireland mine coal, sized to 100 x 200 mesh
Coal feed rate:                     2.0 Ib/hr
Duration of feeding:              3.3 minutes
Air to coal feed line equivalent  to 100% of
  adiabatic   preoxidation level  at  the
  temperature used.
Gas flows, scfh
       Air + N2 to coal feed line          65
       N2 to accelerating line             85
       N2 to bottom of external bed      340

   Tests were made at temperatures from 900
to 1500F, in 100 degree increments. Temper-
ature limits of operability were established as
follows: (1) at  900dF little or no smearing
occurs as was  shown by  presence  of  coal-
derived material in the form of hollow spheres
in the bed after the run, and (2) at 1500F
caking  occurred  in  the   draft  tube.
Unfortunately, we were severely handicapped
by the small scale of the equipment available,
since the draft  tube  principle  had to  be
adapted to the existing 4-in. diameter gasifier
vessel.

   The potential advantages of the process are
that it will supply a feedstock that is assuredly
operable with  respect to  agglomeration at
gasifier conditions;  and it can  produce  a
dense,  closely sized feedstock substantially
free of fines. This will permit a higher gasifier
throughput than otherwise.

   A series of exploratory tests were carried
out  with  the  configurations A,  B, and  C
indicated  in  Figure  3. Best  results  were
obtained with configuration C, but even here
two basic deficiencies were noted. From the
appearance of the agglomerates obtained,  it
was apparent  that  insufficient  mixing
occurred in the draft tube between the injected
                                                                                   ra-s-7

-------
 coal  and the circulating char. Part of the
 difficulty is associated with the small scale of
 the equipment,  since  calculations show that
 the Reynolds number in the  draft tube  is
 barely above the Stokes Law range. Also,  it
 was apparent that most of the'fluidizing gas
 bypassed the main bed in favor of the draft
 tube. The result was that a fluidized bed was
 not maintained external to  the draft tube.

    To  overcome  these  limitations,  the
 configuration C of Figure 3 was modified as
 follows:

    To allow installation of an external baffle
 which  would maintain fluid ization  of  the
 external  bed, the  draft tube  was raised  2
 inches  and the inlet  lines  were lengthened
 accordingly. An elliptical baffle 3-5/8 x 1-3/4
 x 1/16-in. thick was welded to the accelerating
 gas line below the mouth of the tube at a slope
 of 60 from the horizontal.  To help  promote
 mixing a conical baffle was installed inside the
 tube with the apex of the cone positioned 1/2-
 in. above the end of the coal inlet tube.

    The  products from the  runs at  1000  to
 1400F all  showed more uniform smearing
 than in any of the previous  runs without  the
 internal baffle. At the end of each run,  the
 system  was depressured  and  the bed  was
 drained by removing the coal inlet line. The
 hot bed material was quenched  rapidly by
 contact with dry ice in the catchpot. The entire
 bed material then was  screened at 28 and 48
 mesh. All the run products contained some
 +48 mesh agglomerates which were external
 bed particles cemented together by a thin film
 of  coal-derived-material.  No  agglomerates
 larger than 28 mesh were found. The fewest
 agglomerates occurred at 1300F, indicating
that this may be the optimum temperature
with respect to  uniformity of smearing.  The
amounts of +48 mesh agglomerates which
formed are listed below:
Temperature,
F
1000
1100
1200
1300
1400
+ 48 Mesh Agglomerates, wt%
of Bed Inventory
18.0
16.0
15.5
8.0
10.7
The particle density, measured in mercury, for
the +48 mesh agglomerates formed at 1300F
had a high value of 85 lb/ft3.

   An attempt  was  made  to  run for  a
prolonged  period  at 1300F  and  15   arm
system pressure  to determine the size distri-
bution of  the   "equilibrium"  product.  To
simulate  the seed char  in  the commercial
embodiment (fines from the internal  cyclones
in the gasifier) an initial external bed of -100
mesh precarbonized  char was established.
Then, 100 x 200 mesh Ireland Mine coal  and
additional  -100  mesh char were fed to  the
draft  tube  at   rates  of 2  and  4 Ib/hr,
respectively.  The  fine   char  contained  a
considerable amount of -325 mesh material
which was  elutriated from the  reactor.  The
outlet piping system of the present equipment
was not designed to handle large amounts of
solids. The run had to be terminated after 35
minutes of  feeding coal  because the outlet
system began to plug. Thus,  an equilibrium
bed was not established. However, analysis of
the bed showed  that it contained  50 weight
percent of  +100 mesh agglomerates, with  a
top size of 24 mesh.

   The  high particle  density  achieved  is
favorable, in that smearing of liquid coal over
the  seed   particles  apparently  occurs   as
desired.

   The small size of the  existing equipment
precludes any further meaningful  studies of
the seeded coal process. The radial clearance
between the inlet line and the wall of the draft
tube is only 0.15 inch. The mouth of the tube
eventually would become choked by the larger
agglomera-vs which  inevitably  would  be
formed.
ra-5-8

-------
   The  results  of the  exploratory  studies
strongly indicate that future studies should be
made.

   Several  essential factors are  required to
achieve success in such an operation. Intensive
mixing in the draft tube is required to achieve
smearing of the "liquid" coal over both  the
seed  and  recirculating char.  A  sufficient
residence time in one pass through the unit of
the recirculating burden is needed to complete
the "drying out" or carbonization of the coal.
Finally, the draft tube must be large enough to
handle the largest size particles made in  the
process  without  choking.  All these factors
point to a need for a larger unit in which  the
draft tube would be at  least 2  inches in
diameter as opposed to the present 0.680 inch.
Such a unit, of course,  would have  a much
higher capacity for coal feed which would lie
approximately in the range of 30 to 300 Ib/hr.

Pretreatment via Pre-Extraction

   This  method  would be   a more direct
application of the  draft tube pyrolysis method
already demonstrated at the CSF pilot plant at
Cresap,  West   Virginia. Two  principal
differences would be required here.  First of
all, the extraction slurry would be injected into
the  draft tube  unit operated at  15-20  atm
pressure instead of at substantial atmospheric
pressure, and  secondly, the ratio of extract to
extraction  residue normally would  be greater
since little or no extract need  to be recovered
as such. This technique would possibly prove
to be more operable.

Reaction Kinetics

   The other limiting factor is the gasification
of bituminous coals  via the  CO2 acceptor
process is the poor reaction kinetics relative to
sub-bituminous   coals   and   lignites.   The
treatment  of  the gasification  kinetics is
described  in  more detail in a  companion
papers  presented at this  conference which
deals with  gasification of bituminous  coals to
produce low-Btu gas. Suffice to say, that data
available to us indicate that gasification rates
with bituminous coal chars are about l/15th
of lignite chars. This necessitates  increasing
the gasification temperature to  1650F  to
achieve adequate rates. There is little incentive
from the kinetic point of view to increase the
temperature  with  the  available  size  and
density of the char treated, i.e., 28 x 100 mesh
and 45 lb/ft3 particle density. The limiting
factor in throughput in this instance becomes
the fluid dynamics of the char particles rather
than  kinetics of  gasification.  The  use  of
coarser feedstocks, of course, would remove
this limitation and would require higher gasi-
fication  temperatures  to   achieve   higher
outputs.

   The acceptor process becomes deficient in
heat supplied to  the gasifier at  this high
temperature (1650F) unless one or both of the
following  expedients  is  employed.   More
sensible heat as opposed to chemical heat may
be  supplied  by  increasing the  acceptor
circulation rate; or heat may be supplied  by
recycle of carbon  dioxide.

   The  amount  of  supplementary  heat
required by either of the two above expedients
also may be lowered by increase in operating
pressure.
Process Description

   An outline of the proposed process is given
in Figure  1  previously mentioned.  This in-
corporates the system of COi recycle to supply
the heat deficiency in the gasifier,  and  the
"countercurrent" contacting of the feed coal
with  the  gasifier  offgas  in  the draft tube
pretreater. This latter step not only "decakes"
the coal feed but  also significantly increases
the Btu of the product gas.

   The heat and material balance relationship
is given in Table 4 for the processing  of
bituminous coal via the flow scheme of Figure
1. The heat and  material balances were
derived  by  adaptation   of the  computer
program as previously devised  for the OCR
project  on the development of the CO2  ac-
ceptor process. The process assumptions used
are generally quite similar to those outlined in
the companion paper.8
                                                                                    m-5-9

-------
                Table 4.  HEAT AND MATERIAL BALANCE IN TREATMENT

Product
F
moles
Ib
Composition,
mole %
CH/i
H2
CO
CO2
H20
H2S
N2
MgO-CaS
MgO-CaCOs
MgO-CaO
Hydrogen
Carbon
Stream number
1
Steam
1200
2.759















2
CO
1200
1.100















3
Recycle
gas
400
0.157



Same
as
10









4
Pre-
treater
char
1250
4.222
60.9












5.01
94.99
5
Calcined
acceptor
1 873
5.750










27.2
-
72.8


6
Casifier
except
recycle
1650
5.064



3.0T
46.43
21.30
7.52
21.74
-
-





7
Spent
acceptor
1650
0.230










27.2
25.5
47.3


8
Fuel
char
1650
2.124
37.3












2.61
97.39
 aDry,  H2S-free basis.  HHV = 424 Btu/ft3.
 r-
 Total sulfur content = 100 ppm.

 "Water content only.   Dust,  tar, and phenols content not known.
   System  pressure:   14.84  atm  (204 psig) .
   Basis:  100-lb dry Ireland Mine coal
              wt % (dry basis)     mole
      H
      C
      N
      O
      S
      Ash
 4.8
69.8
 1.2
 7.6
 4.3
12.3
2.381
5.812

0.475
0.1341
   6 percent moisture, as fed;  50 percent of coal sulfur removed in preheater;
   95 percent carbon burnout in regenerator; cold efficiency, 79.9 percent;
   total carbon gasified, 64.4 percent; and fixed carbon gasified, 56 percent.
ffl-5-10

-------
OF BITUMINOUS COAL VIA CO2 ACCEPTOR PROCESS (FIGURE 1)
                           Stream number
9
Pretreater
gas except
recycle
1250
7.120

17.15
40.41
19.92
9.89
11.70
0.93

10
Water
gas

100
6.279

19.63a
46.25
22.80
11.32
-
-

11
Air


550
9.051





0.44


12
Regenerator
gas includ-
ing lift
1873
13.548


-
1.98
27.27
0.78
b
69.97
13
Absorber
gas

-
4.538

Same
as
12




14
Makeup
acceptor

60
0.230
48.1







15
Dirty
liquor

200
0.841





100C


16
Lift
gas

1200
2.500


-
2.7
2.5
0.4
-
94.4
17
Excess
gas

-
0.938


as
16




                                                            in-5-ii

-------
    The  gasification temperature 1650F is
 compatible with an overall gasification rate of:
 RX = 39 x 1(H Ib fixed carbon gasified/lb
 carbon  in bed/minute,  and  the other  con-
 ditions cited with respect to carbon and steam
 conversions, etc.

    Recent  experiments   under  the  EPA
 contract with the seeded  coal process show
 that the pretreated coal from the draft tube
 (actually, char, since the temperature is in the
 range of 1.200 to 1300F) will have  a high
 particle density of  about  85  Ib/ft3.  The
 particle density of the gasifier bed then would
 be about 45  Ib/ft3 after gasification of 56
 percent of the fixed carbon (see Table 4).

    Calculations  involving the gas flow rates,
 the above gasification rate, and our fluidized-
 bed density correlation10  showed that one
 train with the gasifier bed dimensions shown
 below will be capable of processing 272,000
 Ib/hr of MF coal.
    Fluidized bed height:
    Gasifier ID:
    Bed density:
    Fluidizing velocity:
          50ft
        25.4ft
14.7 Ib/ft3 char
    0.81 ft/sec
   The estimated  Btu  content  of the  dry
 product gas is 424 Btu/ft3  and the overall
 thermal efficiency on a cold gas basis is 79.7
 percent. Thus, a single train is capable of
 producing about  1.5 x 108 ft3 or 6.5 x 1010
 Btu/day of raw gas. This compares favorably
 with  the  3.5  x   1010   Btu/day (22.5-ft-ID
 gasifier) raw gas output projected for a single
 train in the CO2  acceptor process when used
 for lignite gasification.

   There  are a few features incorporated in
 the process which should be mentioned.

   A small amount of recycle gas (Stream 3) is
 added to the gasifier boot to prevent oxidation
 of CaS in the recirculating acceptor  stream to
 CaSC>4  by the incoming steam and CO 2-

   The fate of the coal sulfur is not known at
present. It was assumed for the purposes of
the balances that are presented that half of the
sulfur is eliminated as H2S in the pretreater; it
will  have to be  removed  by scrubbing  the
product gas.

   The regenerator is operated sufficiently on
the reducing side, such that the remaining
sulfur is largely retained and discarded in  the
spent acceptor as CaS. It also will be recovered
as H2S  via  the  Claus-Chance  reaction  as
proposed previously.1 d

   The sulfur content in the regenerator off-
gases is low enough, such that it may be flared
without  scrubbing; but prior  incineration of
residual  reduced  sulfur  forms  to SC2   is
required.

   The system as shown here  is certainly  not
optimum  and  some improvements   are
potentially possible  as listed below:

  1. Increase  operating pressures from 18-20
    atm. This will reduce the quantity of CC2
    recycle needed for heat balance and also
    raise Btu content of the product  gas.  As
    offsetting features,  higher  regenerator
    temperatures and a  system  to  recycle
    product  gas  to  the  gasifier would  be
    required.

  2. The CO 2 recycle  requirement could  be
    derived from the product gas. Sufficient
    CC>2 would be removed from the product
    gas,  after water-gas shift and prior to
    methanation to provide the CO 2 recycle
    requirements. As a matter of fact, the hot
    pot  absorbent could  be  regenerated in
    such a way as to generate  directly  the
    steam-CO2 mixture required for gasifica-
    tion  under  full  system  pressure, thus
    eliminating  the   CO2  compressor.   A
    difficulty here is the presence of H2S in
    the product gas. Processes are available,
    however, which  afford  at least  partial
    selectivity in H2S versus CO2 removal.

  3. Recycle tar oil and tar to pretreater rather
    than the gasifier. It would be used as a
    vehicle to pressurize  and transport  the
    coal to the pretreater.  By preheating  the
    slurry to obtain  partial extraction,  im-
    proved operability may be achieved.
ra-s-12

-------
 4. By  use  of  a  catalytic  afterburner  to
   combust the CO in the regenerator offgas,
   it may  be possible to generate surplus
   power through the expansion turbine. An
   afterburner will be required in any case to
   convert traces of H2S, COS, and S2 in this
   gas to SO 2-

ACKNOWLEDGMENT

   Appreciation for financial support of the
work described in this paper  and for per-
mission to publish the results is expressed to:

 1. Office of Coal  Research, Department of
   the  Interior,  and the  American  Gas
   Association for the portion describing the
   Rapid City pilot  plant.

 2. Environmental  Protection Agency for the
   portion  describing the  bench-scale
   processing of bituminous coals.

BIBLIOGRAPHY

1. Colsolidation   Coal  Co.,   Research  &
   Development Report No. 16 to the Office of
   Coal Research, U.S. Dept. of the Interior,
   Washington, D.C. under Contract Number
   14-01-0001-415.
   a. Interim  Report No.  1.  Pipeline Gas
      from Lignite Gasification  A Feasi-
      bility Study. U.S. Dept. of Commerce,
      National     Technical    Information
      Service,-PB-166817 (feasibility  study),
      PB-166818 (appendix). February 1965.
   b. Interim  Report No.  2. Low-Sulfur
      Boiler  Fuel  Using the  Consol  CO
      Acceptor Process. U.S. Dept. of Com-
      merce, National Technical Information
      Service PB-176910.  November  1967.
   c. Interim  Report No.  3.  Phase  II 
      Bench-Scale Research on CSG Process.
      January 1970.

        Book  I,  Studies on  Mechanics  of
          Fluo-Solids Systems. U.S. Govern-
          ment   Printing  Office  Catalog
          Number 163.10:16/INT3/Book 1.
        Book 2, Laboratory Physioco-Chemi-
          cal Studies. U.S.  Government
          Printing  Office Catalog Number
          163.10:16/INT3/Book 2.
        Book  3,  Operation  of the Bench-
          Scale  Continuous   Gasification
          Unit. U.S. Government Printing
          Office      Catalog     Number
          163.10:16/INT3/Book 3.
   d.  Interim  Report No. 4.  Pipeline Gas
      from Lignite  Gasification  Current
      Commercial Economics.  U.S. Govern-
      ment Printing Office Catalog Number
      164.10:16/INT4.
2. Curran, G.P., J.T. Clancey,  C.E. Fink, B.
   Pasek, M.  Pell,  and  E.  Gorin.  Annual
   Report to Office of Air  Programs,  En-
   vironmental  Protection   Agency   under
   Contract Number EHSD-71-15. September
   1,  1970 to November 1, 1971.
3. Curan, G.P., W.E. Clark, and E.  Gorin.
   Low-Sulfur Char as a Co-Product in Coal
   Gasification.  Environmental  Protection
   Agency,  Research  Triangle Park,  N.C.
   EPA-R2-76-060,  October  1972.
4. Curran, G.P., C.E. Fink, and E.  Gorin.
   Proceedings of Second  International
   Conference on Fluidized-Bed Combustion,
   1970.  Office of  Air Programs,  Environ-
   mental Protection Agency, Research
   Triangle Park, N.C. Publication Number
   AP-109. pp.  III-l-l to III-l-ll.
5. Unpublished work carried out jointly  by
   Consolidation Coal and Esso Research.
6. Forney, A.J., SJ.  Gasior, R.F. Kenny, and
   W.P. Haynes. Proceedings of the Second
   International Conference of Fluidized-Bed
   Combustion, October 1970.  Office of Air
   Programs,  Environmental Protection
   Agency, Research  Triangle Park,  N.C.
   Publication Number AP-109. pp. III-3-1 to
   III-3-21.
7. Pilot  scale  Development of  the  CSF
   Process. Period July 1,1968 - December 31,
   1970, R & D Report No.  39, Volume IV,
   Book 3. Consolidation  Coal  Co. Prepared
   for Office of Coal Research, U.S. Dept. of
   the  Interior,  Washington,  D.C.  under
   Contract Number 14-01-0001-310.
8. Curran, G.P., B.  Pasek, M. Pell, and  E.
   Gorin. Low-Sulfur Producer Gas via  an
                                                                               m-s-13

-------
B
                             BFW
           QUENCH
           TOWER
 TOH2S

 REMOVAL
                                         SPENT ACCEPTOR
                                          TO GLAUS-CHANGE
                                                                                                             STEAM
"0
2
1
                          Figure 1. Processing of'bituminous coal by the C02 acceptor process.

-------
    22


     21


    20


     19


     18


     17


     16


     15


G   14
o
i-
Q
X
O
IU
as
a.
     12


     11


     10
            Si PREOXIDATION -100 (Ib 02 CONSUMED/lb DRY COAL)
                                                      BASIS

                                                      COAL IN AT 60 F
                                                      6% MOISTURE IN COAL
                                                      AIR IN AT 398 F
                                                      ALL OXYGEN CONSUMED
                                                      HEAT OF REACTION:
                                                      200,000 Btu/lb mole 02 CONSUMED
        400   500    600    700     800    900   1000    1100   1200   1300    1400    1500   1600

                                        TEMPERATURE, F


         Figure 2. Percent preoxidation versus temperature for adiabatic constraint.
                                                                                       III-5-15

-------
                  RECYCLE
           4 in.
                   8 in.
                   24 in.
                    100 x 28 x
                    200m 100m
                    COAL CHAR
                                        RECYCLE
                                                 RECYCLE
                                   4 in.
                                 AIR
                                           8 in.
                                           24 in.
                                            100 x   28 x    48 x
                                            200 m   100 m   100 m
                                            COAL   CHAR   CHAR
                                              4 in.
                                 N2
                                                              NZ
                                                                                                            6 in.
                                                                                                            34 in.
                                                                                                             ex
                                                                                                             o
                                                                                                             cj
                                                                                                             a:
                                                          100 x   48 x
                                                          200 m   100 m
                                                          COAL   CHAR
                                                                                                                               AIR
                                                                                                                                N2

CONFIGURATION A
CONFIGURATION B
CONFIGURATION C
DRAFT TUBE
COAL FEED LINE
TIP POSITION
ACCELERATING
  GAS LINE
0.500 in. OD x 0.444 in. ID
0.250 in. OD x 0.180 in. ID
HALFWAY INTO SKIRT
0.75 in. OD x 0.680 in. ID
0.250 in OD x 0.180 in. ID
1 in. ABOVE  TUBE BOTTOM
0.750 in. OD x 0.680 in. ID
0.250 in. OD x 0.180 in. ID
5 in. ABOVE TUBE BOTTOM
0.375 in. OD x 0.305 in. ID
(TIP POSITION 1 in. ABOVE TUBE BOTTOM)
                                 Figure 3.  Configuration of draft tubes used in seeded coal tests.

-------
                     6.  LOW-SULFUR  PRODUCER GAS VIA AN
          IMPROVED FLUID-BED GASIFICATION PROCESS
           G. P.  CURRAN, B. PASEK, M. PELL AND E. GORIN
                        Consolidation  Coal Company
ABSTRACT

   This paper describes the evolution of the process concepts for generation of clean low-Btu gas
from bituminous coals via fluid-bed gasification. The improved process now under development
for the EPA does not involve the CO2 acceptor principle. Hot sulfur recovery from the gas is
achieved by the use of dolomite. The residual char from the gasifier is utilized in a carbon burn-up
cell. The heat sink utilized in this case is the sensible heat of the air and steam feed to the gasifier.

   Dolomites show activity for hot sulfur cleanup via the reaction,

                         CaCO3 + H2S = CaS + H2O + CO2.

A single limestone was tested and was substantially inactive.

   Various dolomites have been assessed and best results are obtained with pure crystalline type
stones.

   Experimental background data around other key process steps are also briefly presented.
INTRODUCTION

   The production of low-sulfur producer gas
via an adaptation of the CO 2 acceptor process
was described in a paper1 given at the Second
International Conference on  Fluidized-Bed
Combustion. The system was  described in
some detail along with some supporting back-
ground experimental data.

   A detailed process design and feasibility
study of the sytem as well as an experimental
evaluation was  carried out under a  contract
with the  EPA.  The  results are reported in
detail  in the Annual Report.2 The economics
will be briefly summarized later in this report.

   The experimental evaluation of the system
indicated  feasibility of all steps in the process
with one exception. It was found that the sul-
fur recovery from  the  acceptor would be
incomplete from the regenerator. This neces-
sitated addition of another step in the process
in which sulfur is rejected by the reaction first
proposed by Squires,3
CaS + H2O + CO2 = CaCO3 + H2S.
(1)
With this added complication introduced, fur-
ther thought was given to refining and simpli-
fying the overall process.

   It became apparent that there is no real
advantage in using the CO 2 acceptor reaction
simultaneously with the sulfur acceptor reac-
tion in the gasifier when low-Btu fuel gas is the
                                      III-6-1

-------
 desired  product. Disposition of the residual    thermic  calcining reaction in the  acceptor
 char from the gasifier can be accomplished by    regenerator.
 use of a carbon burn-up cell which preheats all
 the steam and  air required for the gasifier.
 The sensible heat duty involved in preheating       Since the acceptor no longer needs to be in
 serves as the "heat sink" .which is necessary to    the gasifier,  the sulfur acceptance reaction
 prevent ash slagging during combustion of the    now can be carried out in a separate, external
 residual gasifier char. In  the  CO2  acceptor    reactor containing a dense-phase fluidized bed
 process the heat sink  is provided by the endo-    of dolomite in the form of CaCO3 MgO.
ra-6-2

-------
IMPROVED   FLUID-BED  PROCESS
DESCRIPTION

   A schematic diagram of the revised process
is  shown  in Figure  1.  The  simplest  con-
figuration occurs when the burn-up  cell is
integral with the gasifier. In this instance, it
would  be analogous to the  CO2  acceptor
gasifier in that the burn-up cell would be in
the form of a "boot" which would contain a
fluid bed of coarse inert solids such as "dead
burned lime." Combustion of the char .residue
from the gasifier would take place in the boot.
In figure 1, the burn-up cell is shown  as a
separate reactor. This is a costlier configu-
ration  but permits more selective rejection of
ash. The hot fuel gas is desulfurized in the
H2S sorption bed by the reaction,
+ CaCO=CaS
                                        (2)
 The bed temperature is held at a level below
 which the acceptor can calcine by the reaction,
           CaCO3 =CaO + CO2.
                               (3)
 The low-sulfur hot gas is cooled to 1300F by
 heat  exchange  with  the water  needed  to
 generate the  gasifier steam,  and  then is
 cleaned of particulates  and alkali  by high
 pressure drop cyclones.

   The sulfided acceptor is conveyed to the
 regenerator  by  continuously recirculating a
 stream of CO2  and steam. In the  regenerator
 the "Squires" reaction takes place at about
 1300F,

    CaS + CO2 + H2O = CaCO3  + H2 S.  (4)

 The  regenerated  acceptor  is  returned  by
 gravity to the sorption reactor.

   A  computer program  has been devised to
 evaluate the  heat  and material  balance
 relationships and overall thermal efficiency of
 the scheme shown in Figure 1.

   The program  evaluates the  interaction
 between the various components of the system
 consistent with the  thermodynamic,  fluo-
 solids mechanic and kinetic restraints on the
 system.
   The entire system was represented by 27
simultaneous linear and non-linear equations
which represent the five basic process  steps
given below:

 1. Carbon burn-up cell,
 2. Gasifier,
 3. Sulfur reactor,
 4. Steam-product gas exchanger,
 5. Squires reaction - impact of temperature
    only,

which also are interrelated by the following
quantities:

 1. H, C,  and O balance,
 2. Heat  balances  around   components  1
    through 4, above,
 3. Water-gas-shift  equilibrium in  compo-
    nents  2 and 3,  above,
 4. Methane yield correlation,
 5. Equilibrium  in the reaction,
    CaCO+HS =
                                                                             (5)
   The  above equations  were solved by an
iterative procedure for the moles of air and
steam fed to the burn-up  cell as a function of
the following variables:

 1. Burn-up cell temperature.
 2. Gasifier temperature.
 3. S/l Ca, mole ratio.
 4. "Squires" reactor temperature.

   Once having  computed the  input  and
output flows and compositions by the method
outlined above, it was necessary to determine
the gasification reactor sizes. The vessel sizing
is determined by the interaction of the fluid i-
zation mechanics of the char particles and the
gasification kinetics.

   The basis' used here was to provide for a
single train to process  120,800 Ib/hr of coal.
The fluidized-bed height was fixed at 50 feet.
The fluidized-bed density was then calculated
using the  correlation  developed  during  the
work on the CO2 acceptor process.  It was
assumed here that a high-density, closely sized
char particle  would  be  generated  by  the
"seeded coal process" from Ireland Mine coal
(the mean particle diameter was taken as 0.04
                                                                                  m-6-3

-------
 inch and the initial  particle density  of 85
 lb/ft3). The reduction in particle density as
 affected by carbon burnoff was based on the
 relationships developed during the work  on
 the CO 2 acceptor process.6

    It is now necessary to compute the vessel
 cross  section and  fluidizing velocities  which
 are compatible with the estimated bed  inven-
 tories by use of reaction kinetic data.

    Extensive differential rate  data7"9   were
 obtained some time ago on the gasification of
 bituminous coal chars as a function of tem-
 perature,  pressure and  mole  fraction  of
 hydrogen in hydrogen-steam mixtures. Subse-
 quently, extensive kinetic data were obtained
 on  the   gasification   kinetics   of  lignite
 char.6'10 In this  work it was  found that the
 reaction  rate was strongly inhibited by the
 presence of CO as well as hydrogen. Thus, the
 prior data on bituminous coal chars could not
 be used since the inhibiting effect of CO was
 not taken into  account.  However,   under
 comparable conditions (in the absence of CO)
 it  was  found  that  on the  average  the
 bituminous coal  chars had about l/15th the
 reactivity of the  lignite  chars.

    Therefore, in developing the kinetic calcu-
 lations the equations  developed  for  lignite
 char10 were  used  with introduction  of  a
 correction factor of l/15th to  account for the
 lower reactivity of the  bituminous coal chars.
 The differential kinetics were translated into
 integral kinetics; i.e., they were averaged over
 the  whole bed by the method given in Ap-
 pendix D of the Annual Report.2

   Having calculated the integral rate Rf, the
 required bed inventory is calculated from the
 equation,      Jj
    R  _  lb fixed carbon gasified  x JQ -4
     T    min/lb fixed carbon  in bed

The fluidization  calculations  outlined above
 are then used to determine the required vessel
cross section and fluidizing velocity.

   A total of 14 cases were computed covering
gasification temperatures between 1650 and
 Table 1. RANGE OF  VARIABLES STUDIES IN
             SYSTEM ANALYSIS
Range of independent
variables studied
Gasifier temperature, "F
Burn-up cell temperature, "F
S/I Ca,mole ratio
Squires reactor temperature, "F
Range of calculated quantities
Cold gas efficiency, %
HHV dry product gas, Btu/ft3
Sulfur removed, %
Steam conversion, %
Carbon to burn-up cell, wt % of C
in coal feed
Gasification lb C (gasified) x 104
rate lb C in bed, min
Gasifier, ID, ft
Char particle density
Gas fludizing velocity, ft/sec
Gasifier cross section index3
Constant parameters
System pressure, atm
Gas outlet temperature, "F

Base Case
1650-1750
1750-1950
0.1-0.4
1200-1300

79.1-81.2
143-149
93.1-97.5
49.4-55.7

14.3-16.9
45-78
24.9-25.3
28.7-35.8
1.46-1.52
392-415
1725
1883
0.2
1300

80.6
147
96.7
53.6

15.4
68
25.1
30.2
1.49
398

- 15 -
- 1300 -
   JFt2/109 Btu-hr (HHV of product gas).


 1750F. Because of the kinetic and  thermo-
 dynamic  limits  the  system  is  highly  con-
 strained,  and the response of the system is
 quite limited. This is  illustrated by the ranges
 given in Table 1.
   The conditions for the base case, which is
felt to be close to a practical "optimum" for
the system, are also given in Table 1.  A more
complete heat and  material  balance around
the base case is also given in Table 2.
   The process concept given here has several
potential advantages over the original process
which utilized the CO2  acceptor  process, as
outlined below:

Operability  May Be  Improved in the  New
Process

1. The O2 partial pressure to the burn-up cell
   is lower than to  the previous regenerator.
   Steam and air N2  serve as the heat sink.
   There  is  less  chance of  ash slagging,
   especially since  the burn-up  cell  can be
   operated at much lower temperatures  than
   are needed to regenerate  CaCO3.
m-6-4

-------
                               Table 2.  KEY STREAM  FLOWS  AND ANALYSES   CASE  6

Identification
F
moles
Ib
Composition, mole %
CH4
H2
CO
C02
H20
H2S
N2
02
NH3
MgO-CaS
MgO-CaCO3
Hydrogen (as Hj!
Carbon
Stream Nnmhpr
1
Air
398
11 .41






0.40








2
Steam
870
3.504















3
Fuel
char
1725
0.936
23.14












4.25
95.76
4
Burn-up
cell gas
1883
14.68


X
X
X
6.10
24.42
X
61.31
8.17





5
Raw
product
gas
1725
20.37


1.72
16.31
18.96
7.85
10.06
0.66
44.18
X
0.26




6
Spent
acceptor
1624
0.648











20
80


7
"Squires"
offgas
1300
3.522


X
X
X
48.16
48.16
3.68
X
X





8
Regenerated
acceptor
1300
0.648











0
100


9
Sulfur
320
0.1296
4.15







/






10
CO 2
200
0.1296















11
Clean
product
gas
1300
20.50


1.71
16.56
18.50
8.78
10.29
0.02
43.90
X
0.25




      Basis:   Ireland Mine Coal (100 Ib dry coal)

H
C
N
O
S
Ash
wt %, dry basis
4.8
69.8
1 .2
7.6
4.3
12.3
moles
2.381
5.812
-
0.475
0.1341
-
            moisture as fed;  system pressure:  15 atm (206 psig)
C/i

-------
 2.  The Oa  partial pressure to the gasifier is
    lower, since the air is diluted with all of the
    input steam and all of the products of com-
    bustion  of the burn-up  cell. Thus,  one
    possibly  can  raise  the  temperature  to
    1750F   to   improve  kinetics  without
    increasing the danger of ash slagging.

 3.  More positive contact of dirty gas with the
    dense-phase  bed  of acceptor  in sulfur
    reactor is effected. Also, in certain circum-
    stances it may be possible to use the sulfur
    reactor as a fluid-bed  filter to remove
    "residual" particulate matter.

 The  Cold Gas  Efficiency is Definitely  Im-
 proved

 1.  Lower duty to  calcine make-up acceptor.
    Circulation rate is about 10 percent that of
    the original concept.
 2.  100 percent burn up of carbon (versus 98
    percent).
 3.  Improved gasification kinetics require less
    steam. Thus, less latent heat in  product
    gas.
 4.  Less air required. Thus, less sensible heat is
    lost with  N2.


 EXPERIMENTAL BASIS

 Pretreatment

    This was discussed in the preceding paper4
 and only the conclusions need be reiterated
 here. Pretreatment by preoxidation is a viable
 procedure for the  more weakly caking Illinois
 No. 6 coals,  but is not a desirable procedure
 for use with the more highly fluid Pittsburgh
 Seam coals. For the latter coals, pretreatment
 by  the seeded coal process appears promising
 but further  development  to  prove  out  the
 method is required.

 Gasification Operability and Kinetics

   Studies were carried out to demonstrate
operability of the gasifier with respect to both
caking and ash fusion using pretreated Illinois
No. 6 coals. The conditions studied were those
that correspond to the original adaptation of
the  CO2 acceptor  process  to low-Btu  gas
production.  Complete operability in the 4-in.
diameter gasifier was achieved on both points.

   It is felt that the conditions in the present
system  are less stringent, as was pointed out
above, such that operability problems due to
ash  fusion are less likely to occur.

   The  background  data  on   differential
kinetics which were used to calculate integral
gasification rates were outlined  in the previous
section.  Integral gasification rates  were  ob-
tained also in operation of the continuous unit
with both Disco char and Illinois No. 6 coal.4
The results were given in the Annual Report.2
The results  are in approximate  agreement
with the basis  used for reactor design given
above, although the rates obtained with Disco
char tend to be somewhat lower  than those
obtained  with  the  pretreated Illinois coal.
Further data on the kinetics  of gasification of
bituminous coals  are required to  provide  a
firmer basis  for reactor  design.


Carbon Burn-up Cell

   No data on the operability of this unit are
available at this time although it is planned to
obtain such data in the course of the present
EPA  contract.  The  operation  comprises
combustion at full system pressure of residue
char in the presence of a  fluid bed of inert
solids,   such  as  "dead  burned  lime."  An
analogous operation is the regeneration of the
CO 2 acceptor process  where residue char is
used  as  fuel   for acceptor  regeneration.
Operability   of  this   process   has  been
demonstrated in prepilot scale work on  the
process. A full-scale pilot  test, of course, is
scheduled at the Rapid City pilot plant.

Desulfurization and Sulfur  Recovery  Steps

   General

   An  experimental program  to  test these
steps is now under  way in  our  library bench-
scale unit as specified  in  our present EPA
contract.
m-6-6

-------
   Preliminary  results are now  available for
desulfurization  of simulated producer gas at
about  1600F  by  means of half  calcined
dolomite and for regeneration at  1300F by
means of the "Squires" reaction.  These will be
discussed  below.

   The offgas from the Squires reaction at the
specified  operating  conditions  (1300F)  is
relatively  low in  H2S content due to equili-
brium limitations.   Special  techniques  are
required  to  recover sulfur  from this  gas
economically without condensation of steam
or removal of carbon dioxide. For this  pur-
pose,  a  liquid-phase  Claus reaction  was
proposed  using hot water under pressure as
the  reaction  medium. This  is  the so-called
"Wackenroder"  reaction,  and the system is
described in  detail  in the Annual Report.2
   Laboratory equipment to test this process is
now being assembled but no results are as yet
available  to report.
   Description of Experimental System

   A flow diagram of the new experimental
 unit is shown in Figure 2. Acceptor in the form
 of CaCOs'MgO is fed continuously at a known
 rate to the top of the H2S-sorption reactor via
 a pneumatic lift line.  The  carrier  gas  is
 recycled product  gas and it  does not  pass
 through the  fluidized  bed in  the sorption
 reactor. Hot, H2S-laden, producer gas is fed to
 the bottom of the bed.  Steam, N2, CO2, H2,
 and H2S are added to a stream of recycle gas
 to simulated the  partial pressures  of the
 various components of the product gas from
 the gasifier shown in the process flow diagram
 in Figure 1. The reactor, previously used as the
 CO2 acceptor regenerator, is  3-in. ID with  a
 bed height of 18  inches.

   The sulfided acceptor is fed  by  gravity to
 the top of the Squires  reactor and is  regen-
 erated while being fluidized  in a stream of
 steam-CO2. The reactor, previously used as
 the gasifier vessel, has been necked down from
 4-in. to 2-in.  ID and has a bed height of 48
 inches.
   Continuity of  acceptor recirculation  is
maintained   by  withdrawal   and   feeding
through  parallel lockhoppers  as shown  in
Figure 2. In both reactors the  acceptor is fed
to the top of the bed and is withdrawn from
the bottom. The height of each fluidized bed is
held at the desired level by means of a  AP cell
placed across the upper part of the bed, which
actuates a solids control valve  located below
the acceptor stand leg.

   The product gas from either reactor can be
monitored continuously for H2S content by
means of a dualprange infrared analyzer. The
continuously recirculating inventory  of ac-
ceptor is sampled  periodically  from  both
reactors and analyzed for CaS and
   Note that the physical arrangement of the
two reactors is reversed from that shown in
Figure 1, for the process reactors. From an
experimental  standpoint,  it  is  immaterial
which reactor is the upper vessel. It is possible
that regeneration  of  the  acceptor  by the
Squires reaction  will  require a  greater ac-
ceptor retention time than that in the H2S-
sorption reactor.  The  existing pressure shell
and electrical furnace for the lower reactor is
considerably larger than that for the upper
reactor. It was chosen to house  the  Squires
reactor because the bed volume can be easily
increased by a factor  of 3  over  that  of the
present design if initial operations show the
need.

   Preliminary Experimental Results

   A series of experiments were carried out, all
with conditions   similar to  those  listed in
Tables 3 and 4 for the gas desulfurizer and
regenerator, . respectively.

   The first series of runs were conducted with
Tymochtee dolomite which  was used  in the
previous work on the CO2 acceptor process.6
A fundamental difficulty found  here is the
very high attrition rate, which ran as  high as
18 percent of the acceptor fed per pass.

   Run A4 (Tables 3 and 4) was made with air
injection into the regenerator in an attempt to
harden the stone  by  partial oxidation to
                                                                                    in-6-7

-------
        TableS. GAS DESULFURIZER
         Table 4. REGENERATOR


Acceptor


Feed rate, Ib/hr (half calcined basis)
Solids residence time, min
Input, scfh
Recycle to bed
H2S
CO2
H2
H20
N2
Purges (CO2lto bed
Purges (N2) above bed
Recycle, acceptor lift gas, above b
Output, scfh in cycle
Exit gas rate, scfh (dry basis)
Composition, mole %
H2
CO
C02
Nl2 (by difference)
H2S
Outlet gas, top of bed
composition, mole %
H20
2
CO
C02
N2
H2S
Flow rate at top of bed , scfh
Fl-uidizing velocity, ft/sec
Attrition, % of feed rate
Duration of circulation with H2S feed.hr
Removal of feed sulfur, %
% H2S in outlet/equilibrium %H2S
Conversion of acceptor/pass, mole %
Run number
A4
35x48 mesh
Tymochtee
dolomite
5.9
32

130
3.5
35
60

65
5
15
ed 92
1-3
148

18.3
17.3
12.1
52.1
0.09


10.9
16.3
15.4
10.8
46.5
0.08
275
0.40
5.6
7.1
97
2.3
23
Temperature = 1600"F, Pressure = 206psig
A7
28x35 mesh
Canaan
dolomite
6.5
33

175
3.5
54
73

96
5
15
71
1-2
215

17
18
12
53
0.05


9.7
15.4
15.9
11.2
47.7
0.04
416
0.60
0.7
25.2
97
1.4
19

CaSC>4 in situ. Prior work in the CO2 acceptor
process  development showed  that hardening
of the stone occurs when this is done at higher
temperatures  due  to  the formation  of a
transient liquid  in the CaS-CaSO4 system.6
The attrition rate apparently was  somewhat
reduced over the comparable  run  where air
injection was  not used but was  still unac-
ceptably high. The operating temperature
(1300F) was apparently too  low to achieve
hardening by the transient liquid mechanism.

   A Tymochtee  dolomite which had been
hardened by cycling through the CO 2 acceptor
process was also tested. This material showed
the expected good resistance to attrition. Only
preliminary  results  are  available  at  this
writing,  but the activity of the stone appears to


Solids residence time, min
Input, scfh
H2O
COa
Air
H2
Purges (N2) to bed
Purges (N2) above bed
Purges (C02) above bed
Output, scfh in cycle
Exit gas rate, scfh (dry basis)
Composition, mole %
C02
N2
H2S l
S2
Outlet gas, top of bed
Composition, mole %
H20
C02
N2
H2S
S2
Flow rate, top of bed^scfh
Fluidizing velocity, ft/sec
Regeneration of acceptor, mole %
Run number
A4
38

89
82
7.5

10
10
5
1-3
110

76.1
23.0
0.8
0.05


43.5
47.3
8.6
0.5
0.03
186
0.53
6.7
A7
37

110
110


8
10
5
1-2
133

85.7
13.5
0.7



48.3
47.8
3.5
0.4

228
0.64
5.1
Temperature = 1300F,  Pressure = 206 psig
be less  than that of the  fresh  Tymochtee
dolomite.
   A Nebraska limestone was  also tested; it
shoi:"*H  very low absorption of H2S as com-
pared with half calcined dolomites. This is in
accord with prior laboratory studies by Ruth,
et al.    Other limestones such as BCR-1692
used in  the Esso, Ltd. work will  be tested to
determine whether the low  activity is generic
to all limestones.
IH-6-8

-------
   A  pure  dolomite  from  Canaan,   Con-
necticut, was also tested under the conditions
indicated  in  Table 3 and 4.  This material
showed excellent  attrition resistance with an
average weight loss of only 0.8 weight percent
per pass  through the system.

   Both the Tymochtee and Canaan dolomites
showed excellent fresh activity for removal of
HaS in the  gas  desulfurizer.  Ninety-seven
percent removal of H2S was achieved in both
cases. The removal of CaS in the regenerator
was highly incomplete in  both cases. Thus, it
would appear that the kinetics of the Squires
reaction is a limiting factor in this  system and
this factor requires more study.

   The  definition  of results  in  terms  of
number of cycles is difficult because of the
semi-continuous nature of the acceptor circu-
lation, loss of material due to  attrition,  and
sampling  of material for  analyses. An ap-
proximate,  and  "conservative" method for
calculating the number  of cycles  was used.
Each pass of a charge through  the unit is
calculated as a fractional cycle  An using the
relationship: An equals the actual amount fed
divided by  the internal  inventory plus the
actual external inventory.

   The  number  of  cycles that  could  be
achieved  with the Tymochtee  dolomite were
limited due to the high attrition losses. In the
case of the Canaan dolomite, attrition was not
the  determining  factor;  the  run  was  ter-
minated when breakthrough of H2S occurred
in the gas  desulfurizer.

   The composition of the offgases from the
gas desulfurizer and regenerator as a function
of the number of cycles for runs A4 and A7 are
given in Figures 3 and 4, respectively. In both
cases the H2S content of the regenerator off-
gases increases with the  number of cycles.
This, as we will show below, is due to accumu-
lation of CaS on  the  acceptor and points to
poor kinetics in the Squires reactor. The E^S
content of the desulfurizer offgas  remains
relatively steady until the acceptor becomes
heavily loaded with sulfur towards the end of
the run.
   The sulfur content of the stones for the two
Runs A4 and A7 as a function of cycle number
are given in Figures 5 and 6, respectively. It is
immediately clear that in both cases there is
rapid buildup of CaS  on the acceptor due to
incomplete  regeneration  of the  CaS.  The
efficiency  of sulfur removal and recovery is
further  illustrated by  the  tabular  data
presented for Run A7 with Canaan dolomite
in Table 5.

   At 19 percent conversion per  pass to CaS
(Table 3, Run A7), the  known percent of sulfur
rejected (Table 5)  makes possible a rough
estimate of the Ca/S  ratio required if fresh
dolomite feed were added continuously to the
system operated under the above conditions.
The Ca/S ratio would be approximately 0.45.
                 Tables. RUN A7 H2S CONTENT OF EXIT GASES (DRY BASIS)
                                     CANAAN DOLOMITE
Gas desulfurizer
Cycle No.
0.3
1.4
5.5
7.5
8.6
12.6
H2S,
mole%
0.046
0.049
0.048
0.058
0.064
0.672
H2S
Removal, %
97
97
97
96
96
59
Regenerator
Cycle No.
0.5
1.5
5.7
7.8
9.2
12.5
H2S,
mole%
0.671
0.927
1.39
1.69
1.40
1.90
Recovery of
H2SFeed,%
25
35
53
64a
53a
72 a
            The condensate contained about 3% of the feed ulfur as elemental sulfur.
                                                                                   III-6-9

-------
   The acceptor at the  end of Run A7 was
nearly completely converted to CaS, hence the
break-through observed in the desulfurizer. At
the  end  of the  run,  2  hours   additional
residence time was given to the batch of ac-
ceptor remaining in the regenerator. The CaS
content of the acceptor was reduced from 85 to
76 mole percent. This indicates again that the
poor kinetics in the Squires  reactor is  con-
trolling.  Better results should  be  achieved if
longer residence times in the  regenerator are
used. It  does  appear  that  a considerable
amount of inactive CaS  is inevitably formed;
at the present time, however, we are unable to
clearly  distinguish  between  CaS of  low
reactivity and  the  "dead" material.

   It thus appears  that high purity crystalline
dolomites have acceptable physical strength
and  activity   for  use  in   the  process.
Economically acceptable make-up rates can
be  achieved at the proper  operating  con-
ditions.  The  geographical distribution of
dolomites with acceptable strength and ac-
tivity is now being studied. A variable study is
also  planned   with  selected  dolomites to
determine optimum conditions for their use.


ECONOMICS OF PROCESS

   No economic figures  are available for the
improved process.  The potential economics of
the CO2 acceptor based  process were given in
the Annual Report.2 Figure 7 is  reproduced
from that report. It gives the cost of low-Btu
gas as a function  of coal cost delivered to  a
1200-MW boiler from  a  large-scale  gasifi-
cation plant. The  figures are  based on  1976
operation, 15 percent capital charges and 7.5
percent/yr escalation on materials and labor,
and 7.5  percent/yr  interest  during  con-
struction and an operating factor of 70 per-
cent.

   The  economics  of the new  process are
expected to be somewhat better.

ACKNOWLEDGMENT

   Appreciation is expressed to the Environ-
mental   Protection  Agency   for financial
support of the work presented in this paper
and  for  permission to publish the results
given.

BIBLIOGRAPHY
 1. Curran, G.P., C.E. Fink, and E. Gorin.
   Proceedings of Second International Con-
   ference on  Fluidized-Bed  Combustion,
   1970. Office of Air Programs,  Environ-
   mental   Protection   Agency,  Research
   Triangle Park, N.C. Publication Number
   AP-109. pp.  III-l-l to HI-1-11.

 2. Curran G.P., J.T. Clancey, C.E.  Fink, B.
   Pasek, M. Pell, and E.  Gorin. Annual
   Report to Control Systems Division, Office
   of Air Programs, Environmental Protec-
   tion Agency, Research  Triangle  Park,
   N.C., under Contract Number EHSD-71-
   15. September 1, 1970-November 1, 1971.

 3. Squires, A.M. Fuel Gasification. Advances
   in  Chemistry Series  No. 69. American
   Chemical   Society.   Washington,  D.C.,
   1967. pp.  205-229.

 4. Curran, G.P.  and Everett Gorin. The CO2
   Acceptor  Process  A  Status Report.
   (Presented  at 3rd International Confer-
   ence  on   Fluidized-Bed  Combustion.
   Hueston Woods. October 29-November 1,
   1972.) See Session 3, Paper 5 this volume.

 5. Book  1, Studies on  Mechanics of Fluo-
   Solids Systems.  Consolidation Coal Co.,
   Research & Development Report No. 16 to
   the Office of Coal Research, U.S. Dept. of
   the Interior,  Washington,  D.C., under
   Contract    Number     14-01-0001-415.
   Government  Printing   Office  Catalog
   Number 163.10:16/INT3/Book 1.

 6. Book  3,  Operation  of the  Bench-Scale
   Continuous Gasification Unit. Consolida-
   tion Coal  Co., Research  & Development
   Report  No.  16  to  the  Office  of  Coal
   Research,  U.S.  Dept.  of the  Interior,
   Washington,  D.C., under Contract Num-
   ber 14-01-0001-415. Government Printing
   Office Catalog Number 163.10:16/INT3/-
   Book  3.
ra-6-io

-------
7. Goring, G.E., G.P. Curran, R.P. Tarbox,    10. Book  2,  Laboratory  Physico-Chemical
  and E. Gorin. Ind. Eng. Chem. 44:1051,       Studies. Consolidation Coal Co., Research
  1057, 1952.                                  & Development Report No.  16 to the
                                              Office of Coal Research, U.S. Dept. of the
8. Goring, G.E., G.P. Curran, C.W. Zielke,       Interior,  Washington, D.C. Government
  and E. Gorin, Ind. Eng. Chem. 45:2586,       Printing    Office    Catalog    Number
  1953.                                       163.10:16/INT3/Book 2.

                                           11. Ruth, L., A.M.  Squires, and R.A. Graft.
9. Zielke,  C.W. and  E.  Gorin.  Ind.  Eng.       Environmental  Science  and Technology.
  Chem. 47:820, 1955, and 49:396,  1957.         November 1972.
                                                                            m-6-ii

-------
O3
        COAL
PRODUCT GAS
                                         WITH H2S
                                          GASIFIER
                                        CHAR
                                                      STEAW

                                                 "SQUIRES"
                                                  REACTOR
                                                                           SULFUR '
                                                                          REACTOR
                                                                            H2S
                                                                          SORPTION
                              1   CaS- MgO
                                                                                        H2S
                                                                                       STEAM
                                                                                         C02
                                                                      SULFUR
                                                                     RECOVERY
                                                                         I
                                                                       SULFUR
                                                                                                                  CLEAN
                                                                                                                 PRODUCT
                                                                                                                  GAS
                                                                                                                  1300 F
                                                                                                          . H20
                                  Figure 1. Two-stage fluidized-bed partial  combustion process.

-------
                                                                                                         DPCV
                                                                                                                      VENT
                  WATER


                 UDEMINER
                 11 ALIZER
                                                              CONDENSER
                                                               RECEIVER
                                                                                                              DRY GAS
                                                                                                             'METER
                                               H2S SORPTION

                                                REACTOR
                            ACCEPTOR
                            STAND LEG
                                        TOGAS
                                       ANALYZER VV
                                                 MAKE-UP
                                                ACCEPTOR
                                                                                                RECYCLE
                                                                                               COMPRESSOR
                                                 ACCEPTORr*lF"-TERS
                                                 CHARGE POT
            CONDENSER
             RECEIVER  COOLER
                            cu
                  ,  BALANCE
    H,  No  CO,   H,S   GAS
HSOLIDS
 CONTROL
  .VALVE
VENT A,TFRTnpVENT  EMERGENCY
     ACCEPTOR  i     SOLENOID
                                    ACCEPTOR
                                 VFNTSAMPLE
                                 VEN.T  TUBE  ^
DEMINER-
 ALIZER
                                           DRY GAS
                                           METER
                                     ACCEPTOR
                                      STAND LEG
ACCEPTOR
  FEED
 HOPPERS
          WATER FEED
             PUMP
                    BALANCE
                       GAS
           FEEDER   SYSTEM
                                                                "SQUIRES"
                                                                REACTOR
                                                                                                     COIVTPRESS^R
                                ACCEPTOR WITHDRAWAL
                                      HOPPERS
                                      Figure 2.  Flow diagram of new experimental unit.

-------
*   1-4
    1.2
    1.0
    o4
    0.6
                               REGENERATOR
                  1         2         3
                      CYCLE NUMBER
                                                   0.12
                                                   0.11
                                                   0.10
                                                   0.09
                                                  0.0
                                                   0.07
                                                C/9
                                                 CS1
                                                   0.06
                                                   0.05
                                                   0.04
                                                                     GAS OESULFURIZER
                                            4        0


                              Figure 3. Run A4-tymochtee dolomite.
1        2         3
   CYCLE NUMBER
III-6-14

-------
    0.07


    0.06


*  0.05
o>


I"  0.04
_c
6*
  0.03
E
IS)
*  0.02


    0.01


      0
                                                                                              14
                                    /  TO 0.67 at 12.6 cycles
                                      GAS DESULFURIZER
                                  4             6
                                         CYCLE NUMBER
                            8             10


Figure 4. Run A7-canaan dolomite.
12    14
                                                                                         III-6-15

-------
     90







     80








     70








     60

E



5


1    50
  JD
  O



  

  <3

  +
  00
     40
       30
       20
       10
                    GAS DESULFURIZER
                   1          2

                      CYCLE NUMBER
                                                    90
                                                   80
                                                    70
                                                    60	
                                                S  50
8   40
s
                                                TO
                                                o
    30
    20
   10
           REGENERATOR,
                1          2

                   CYCLE NUMBER
                        Figure 5. Run A4-molar sulfur content of exit solids.
III-6-16

-------
                4     6     8     10

                  CYCLE NUMBER
12
                      4     6     8     10

                       CYCLE NUMBER
                         Figure 6.  Run A7-CaS content of acceptor.
     80
S
o
cx>
<
CJ
     70
     60
      28       30       32       34       36       38       40

                                 COAL COST, 
-------
SESSION IV:
  Conceptual Designs and Economics

SESSION CHAIRMAN:
  Dr. D.H. Archer, Westinghouse
                            IV-0-1

-------
    1. SMALL-SCALE APPLICATIONS  OF FLUIDIZED-BED
                           COMBUSTION AND HEAT  TRANSFER

                      D. E. ELLIOTT AND M. J. VIRR
                University of Aston, Birmingham, England
  Research on fluidized-bed combustion has
primarily been aimed at improved economics
and anti-pollution measures for coal- and oil-
fired power stations or for relatively large
packaged boilers. At the last Hueston  Woods
Conference, however,  the  Keynote Address
hinted that advantages might be gained from
applying fluidized-bed combustion and heat
transfer techniques even   on  very  small
systems.

  In the two years since then,  considerable
progress has been made in research to identify
and solve problems associated with small-scale
application; a  small development  company
has been started to exploit areas of likely com-
mercial interest.

  As some of the work involved may  give
feed-back  to  large boiler  technology,  this
paper reviews the state of the art and gives
data on combustion and heat transfer.
GAS-FIRED
COMBUSTION
FLUIDIZED-BED
   Up to now, research  at the University of
Aston Mechanical Engineering Department
has mainly concentrated on gas firing. That
gas can be burned successfully in deep fluid-
ized beds is well-known. Reference 1 describes
Russian  work.  Reference  2  cites  French
research. The Coal Research Establishment
(NCB) used gas combustion in the early days
of fluid-bed coal burning work to investigate
aspects of volatile burning. Most of this work
was done in  beds 1-ft  deep or more;  it
required the use of  relatively high pressure
blowers, therefore necessitating either high-
speed motors or some form of multi-stage
rotors. In a practical plant the former solution
would result in noisy appliances, while  the
latter incurs high manufacturing  costs.  A
further disadvantage  with deep fluidized beds
when applied to small heat inputs is that the
total surface area from which heat can be lost
is high in relation to the throughput,  so .that
beds with an L/D ratio of more than about 0.5
have to be surrounded with a high  degree of
lagging to  prevent undue heat losses  during
startup.

   Research was, therefore, initiated  to find
the minimum  bed depth which would give
stable and  efficient  combustion. Because of
the very poor lateral gas mixing in fluidized
beds, the idea of introducing separate gas jets
into an air fluidized bed was discarded;  the
bed  depth needed for complete  mixing and
good combustion would be too great unless an
extremely large (and costly) number of gas jets
were used. The experiments  were conducted
with pre-mixed  gas/air mixtures fed through a
porous ceramic distributor into a bed of silica
sand (Figure 1).  Provided that  the  gas/air
mixture  is initially within the  flammability
limits (2.2  to 9.5 percent by volume for  the
propane normally used) and  the initial fluid-
izing velocity is less than the flame propaga-
tion speed, then such a bed can be ignited by
simply lighting the gas/air mixture above the
surface of the bed. Whether or not the system
heats up to obtain  controlled fluidized-bed
                                       IV-1-1

-------
 combustion, however,  depends upon having
 suitably sized  solids particles, choosing  an
 appropriate heat input rate  for the system,
 and preventing undue heat loss at startup.


    Various stages of startup on a bed with cor-
 rectly chosen  conditions (1-in. deep,  6-in.
 diameter, 18,000  Btu/hr input,  10 percent
 excess air) are as follows:

 1.  Air  is blown through the bed at a rate
    which is  around the incipient fluidizing
    velocity.   The   bed  surface   is  hardly
    disturbed. A near  stoichiometric ratio of
    gas  is then admitted.

 2.  On ignition, the gas/air mixture burns at
    the top of the bed with a blueish flame,
    which dances around in an irregular pat-
    tern as it is not attached to any particular
    stabilizing system.

 3.  Particles which had been thrown up from
    the top surface of the bed into the flame
    and heated, now return to the bed and
    carry the heat down.

 4.  The  gas/air  mixture  now  becomes
    preheated as it passes into the bed and
    burns more readily at the  surface with a
    distinct popping, roaring noise. The pre-
    heating causes the fluidizing velocity auto-
    matically to increase; it also causes more
    particles  to  be thrown up,  which in turn
    maintains the rate of temperature rise in
    the bed. The flame structure modifies, and
    red-hot particles begin to tinge the colour
    of the flame.

5. When the bed temperature reaches 600C,
   combustion  begins  which glows dull-red
   behind the reddish-blue flame. Noise level
   increases. By now the fluidizing velocity is
   some three times greater than the initial
   cold fluidizing velocity and more bubbling
   takes place.  Combustion in the bubbles
   becomes  more"  violent;  mild  explo-
   sions/detonations occur  which  tend  to
   throw up far more particles from the sur-
   face of the bed.
6. When  the  bed temperature  has reached
   700 C, much of the combustion occurs in
   the bed; by 800C only a small proportion
   of the gas burns at the top. The noise level
   is now at a peak,  probably indicating that
   the combustion  is  occurring after  the
   gas/air mixture has formed itself into bub-
   bles within the fluidized bed. The surface
   splashing of particles has decreased,  which
   indicates that although it is suspected that
   combustion occurs in the bubbles, these
   bubbles are further below the surface when
   combustion occurs than in the case  of the
   600C  bed.

7. By 900C the noise  level is reducing;  at
   1000 C the level has dropped several deci-
   bels below its maximum level. The ultimate
   temperature level  reached depends mainly
   on the gas and air input to the bed. Tem-
   peratures up to  1200C  can be readily
   maintained with silica sand, but particles
   with higher fusion temperatures are needed
   to  go  much  beyond  this   temperature
   because of sintering. The noise level  is
   reduced in  deeper beds over  800 C and
   disappear above 880 C.

   Figure 2 shows a typical heating up rate for
a shallow bed  combustor; Figure 3 shows the
noise level spectrum  emitted from the bed at
various temperatures. It is believed that lower
noise levels occur at  the higher temperatures
because combustion is extremely rapid in the
first few millimeters of the bed before the
gases have time to form into bubbles.  Thus,
the likelihood of detonations  occurring  is
reduced because the  combustion is quenched
by the large number of particles present.

STABILITY

   Once a temperature above 800C has been
reached, the system  can be operated over a
wide range of  gas/air mixture strengths well
outside the normally accepted  limits  set by
flame propagation phenomena. The fluidized
bed operates as  a very effective pre-heater,
bringing the incoming gas up to bed tempera-
ture within the first few millimeters of the bed.
rv-i-2

-------
Provided the external heat losses from the bed
are small, achieved by lagging and by placing
reflecting surfaces above the bed to radiate the
heat back down to the bed, extremely weak
gas/air mixtures can be  burned.

   Combustion can be maintained with ex-
tremely shallow beds (below 0.5-in.), but  it  is
not yet clear whether or not complete temper-
ature equilibrium between the exit gases  and
the solids is achieved. With very shallow beds,
some combustion probably takes place after
the gases leave the bed. For beds 0.5-in. deep
and above the exit gases appear to be more or
less in temperature equilibrium with the bed,
and combustion  efficiency is  excellent with
CO/CO2 ratios dropping below 0.002 (a factor
of ten better than the standards insisted upon
by the UK Gas Council  for domestic  appli-
ances).

   Because of the quenched low temperature
combustion it was expected that NOX produc-
tion would be low. This has been borne out by
samples  drawn   through   Dager   tubes
measuring NO + NO2 which indicated  less
than 5 ppm in the exhaust gases. These figures
represent a considerable  reduction compared
with  emission from normal flames.

   It  is interesting to note that the combustion
intensity of these beds is  in the region of one
million Btu/ft3-hr based  on the full depth  of
the bed. As  it is likely that most  of the com-
bustion takes place in the bottom half of the
bed, the actual combustion intensity must be
at least twice this rate. A similar bed, but  with
heat  transfer by direct contact  between the
particles and cooling surfaces, can be operated
at more than twice the above heat release rate;
figures of 3 x 106 Btu/ft3 -hr have already been
achieved.

   The ease of operation of this form of gas
"burner" in the temperature range of 800  to
1200C,  combined with  the  excellent  heat
transfer which occurs when small objects are
placed in the bed makes possible the use  of
such  furnaces for laboratory, workshop,  and
factory metallurgical  processes,  e.g.,   har-
dening, annealing, heating small  billets prior
to forging etc. Figure 4 shows a typical heating
curve for a 3-in. long x 1/4-in. diameter alloy
steel bolt immersed in a fluid-bed combustor
operating at 960C (the extra heat  absorbed
during  the  transformation zone  is clearly
seen). With these possibilities in mind a small
company, Fluidfire Development Limited, has
designed and built a range of furnaces:
1. 6-in. diameter x 1-in. deep self-contained
   units  for  laboratory investigations  and
   demonstration work.

2. 6-in. diameter x 6-in. deep units  again for
   laboratory work which contain a  two-stage
   blower or can be  used from a  shop air
   supply.

3. An 8-in. diameter  x 8-in. deep  metallur-
   gical furnace with  automatic temperature
   regulation for use in small quantity har-
   dening and annealing work.

4, A  larger 12-in.  diameter  x  12-in.  deep
   metallurgical unit, again fully temperature
   controlled, shown in Figure 8.

   The  metallurgical  furnaces  which  can
operate over a  temperature range  of  from
about 700 to 1200C can replace traditional
salt  bath  furnaces  for hardening  and  tem-
pering a wide range of materials. The fluidized
furnace has the following advantages:

1. Higher operating efficiency allows  lower
   fuel costs per pound of metal processed.

2. The cost of special salts and the  difficulty
   of handling and disposing of the  spent salt
   are eliminated.

3. The furnaces have a short startup time and
   therefore can be switched off overnight.

4. The atmosphere in the heating zone can be
   adjusted to suit the  requirements of treat-
   ment. It is usually made to be  reducing.

5. The furnaces can operate over a wide range
   of temperatures and  with slight modifi-
   cations  can  be changed to carburising
   duties in which case they eliminate the use
   of cyanide with all its attendant safety and
   disposal .problems.
                                                                                   IV-1-3

-------
  6. All types  of steels may be heat treated.

   The  performance  of  these  units  with
 respect to operating cost,  heating rates, and
 oxidising rates at various temperatures and
 the  hardness achieved under various condi-
 tions is  described in reference 3.

   A special unit has  also been designed to
 operate in line with automatic continuous pro-
 duction of hardened steel components.  This
 unit uses a continuous  wire  mesh belt  to
 support  and  convey  articles through  two
 separately controlled fiuidized- bed combus-
 tors: the first higher  temperature combustor is
 the  heating zone;  the  second combustor,
 running at a precisely controlled temperature,
 allows a short soaking period and ensures that
 the  articles are fed into the quenching system
 at the correct  temperature.

 RADIANT HEATERS

   Shallow-bed combustion systems of this
 type are  interesting in their own right since
 they are a new way  of making more effective
 radiant  gas heaters. Normally,  radiant  gas
 heaters rely on convective  heating of ceramic
 plaques by very hot gases and the re-radiation
 of   the   heat  from  the   plaques   to   the
 surroundings. The exit gases are at a substan-
 tially higher temperature than the  radiation
 surfaces; the overall effectiveness is  generally
 such that only 25 to 35 percent of the input
 heat is radiated.

   The immense surface area exposed to the
 combustion gases in  a fiuidized bed allows the
 temperature differential between the gases
 and  the  solids  to be negligible.  Thus, a bed
 operating stoichiometrically at  1000C  will
 radiate approximately  50 percent of its  heat
 away from the  bed.  The operating tempera-
 ture/radiation efficiency of an ideal  bed  (i.e.,
 perfect  burning of  the  air  and gas  before
 leaving the  bed  and thermal   equilibrium
between  the gases  and  the solids)  can be
readily calculated. The radiation  efficiency is
given by the expression
          n = radiation = 1 -
                           ili
                            H
where: ht is the enthalpy of the exhaust gases
at the bed operating temperature, and H is the
calorific value of the fuel burned.

   There are only two sources of heat loss from
the bed, the enthalpy of the exhaust gases and
the  radiation  from the bed. The  external
convective heat losses  are negligible  if the
height of the containment wall  is low. Any
heat directed  downwards to the distribution
plate is returned to the combustion bed as pre-
heat in the gases.

   Figure  5 plots  the  radiation efficiency
versus bed temperature  for various gases. It
will be noted that the radiation efficiency for
stoichiometric mixtures  is not very different
for the various gases and is not a function of
the bed emissivhy which only affects the rating
of the bed per unit area. Figure 6 shows how
the heat input and  the radiant output of-a-6-r
in. diameter shallow-bed combustor varies as
the temperature of the bed changes for stoi-
chiometric gas/air  mixtures. The full curves
correspond to a bed of unit emissivity and the
dotted curves to a  bed of emissivity 0.7.

   It will be noted  that  as the bed operating
temperature is lowered by reducing the gas/air
volume fed to the bed, the radiation efficiency
increases. Above about 800C the small fiuid-
ized beds which have been produced perform
very nearly as shown provided that  due ac-
count is taken of the effective emissivity from
the bed surface. It was originally expected that
the emissivity of the granulated surface of a
gently bubbling fiuidized bed would act very
nearly as a black body for most solids. But this
does not appear to  be the case; the emissivity
for silica sand is approxmately equal to that of
silica sand itself.

   Below 800 C the radiation effectiveness is
not  as  good  as predicted;  either  thermal
equilibrium is not established or some gas is
leaving the surface unreacted.

   Although the emissivity of a  gently bub-
bling bed falls into line with the individual
particle emissivity, the overall effective radia-
tion from  a fluidized-bed combustor using
IV-1-4

-------
very fine particles may be significantly higher
than this. As might be expected, the influence
of a particle  cloud above  a  fluidized  bed
materially affects its radiation characteristics.
This effect exists because the particles during
their stay in the  gas  space above  the  bed
radiate their heat very rapidly, cool down to a
temperature below that of the  off-gases,  and
thereafter tend to act as a second-stage cooling
medium for  the gases leaving the bed. Thus, it
should be possible to  operate a fluidized-bed
combustor at a temperature which is signifi-
cantly more  than the temperature of the gases
leaving  the  system. In this case, the cloud of
particles above the bed shows a duller colour
than one would normally expect  from a bed
operating at the same temperature. Research
into this phenomena  is underway at Aston.

   It will be seen from Figure 6 that the heat
output  of a 6-in. diameter  bed operating at
1000C is approximately the same as that of a
traditional British radiant gas fire. This  has
led to the concept of using  shallow fluidized-
bed combustors as room heating devices.  The
constantly varying pattern  of the fluidized
bed, coupled with the similarity  to the open
coal fire, was thought to offer an  attractive
alternative to the traditional fire. Thus, self-
contained  units  incorporating brushless
electric   motors, fans,  controls,  and safety
devices  are  now being developed  by Fluidfire
Development Limited  in order to assess  the
potential of such appliances. Although much
more work is still  needed to satisfy stringent
safety requirements, many problems inherent
in such  a radically  new system have been over-
come,  and  there  appears a fair chance of
success.

NOISE

   One of the requirements for room heaters is
that they should not generate a great deal of
noise, and the UK Gas Council recently pro-
posed a standard  of  acceptable  noise levels
shown by the dotted line on Figure 3. It will be
seen that the noise level from an open radiant
bed is slightly higher than the permitted noise
levels. Thus, to quiet  the  fire somewhat  and
also to lessen the danger  of having  a com-
pletely open fire, a glass screen could be incor-
porated in front of the fire. Tests with ceramic
glass show that this reduces the noise level to
well below the acceptance levels and  ensures
that clothing cannot be ignited by direct con-
tact with the fire.

BOILER APPLICATIONS

   It is obvious that if a radiant fluidized-bed
combustor  operating at a temperature level of
800 to 900C is surrounded by a water jacket,
then 55-60 percent  of  the heat input will be
radiated to the water walls, even if the fluid-
bed particles do not come into contact with the
walls. However,  the rating of such a boiler
would be relatively lowa 6-in. diameter bed
having an  output of something  like  15,000
Btu/hr. Hence, the unit would not be particu-
larly compact and could not be considered as a
viable  commercial alternative to the highly
rated gas-fired boilers which are now being
produced.

   If an attempt were made  to place heat
transfer tubes in the fluidized-bed combustor
in a similar manner to  the way in which large
fluidized boilers have been designed, then the
simple  startup procedure  described  earlier
would not be effective,  and compartmentation
of the  bed for  startup  purposes would  be
necessary.  For small-scale  appliances, this
would  be prohibitively costly. An alternative
was  sought; the idea of locating heat transfer
surfaces just above  and around the settled
fluidized bed was formulated and has been
successfully developed. This solution relies on
the principle  that the expansion of a  shallow
fluidized bed (as a percentage) is very high
compared with a deep  fluidized bed. Thus on
startup with cold air the bed expansion is very
small,  and the particles  do not contact the
heat transfer surfaces.  If a  gas/air mixture is
then lit above the bed, the  bed  heats up in a
similar fashion to the  radiant bed described
above; but  by the time a temperature of 700 to
800 C  has  been  reached, the  fluidizing
velocity is some three to four times the initial
velocity and the bed has expanded to contact
                                                                                   IV-1-5

-------
 the  heat transfer  surface. Further  contact
 takes place by virtue of splashing of particles
 from the bed upwards and sideways.

   For beds of  small  output  (up  to  100,000
 Btu/hr it is sufficient to surround the bed with
 a water-cooled wall which is  insulated from
 the  settled bed.

   The direct contact between the particles
 and the  cooling surfaces allow the heat input
 to be two to three times that  which could be
 sustained in a bed which was only cooled by
 radiation. The principle could  still  hold good
 for  larger outputs, but  in  this  case some
 additional heat transfer surface would have to
 be placed in such a position that it contacted
 the  expanded bed and received splashing heat
 transfer.

   Because the  fluid-bed combustor will not
 operate  very satisfactorily below 800 C,  and
 even if we could exploit father the principle of
 particle cloud radiation to cool the off-gases to
 below bed temperature, the overall efficiency
 of the combustion system in transferring heat
 to water would be too low for domestic central
 heating  systems. Some form  of second-stage
 heat recovery  is therefore   necessary.  The
 incorporation  of  convective   heat transfer
 surfaces would leave the system with many of
 the  disadvantages  of normal systems,  e.g.,
 large heat transfer volumes or the use of high
 extended  surfaces  with  the  possibility  of
 condensation  and  corrosion  troubles.  A
 second-stage shallow fluidized bed was there-
 fore incorporated above the combustion bed of
 a trial 40,000 Btu/hr laboratory unit which
 was  supplied by air from an external source.
 With a 1-in. bed depth and an 8-in. diameter,
 the heat transfer area of about 1/6 square foot
 around the periphery of the bed allowed the
 existing  gases to  reduce in temperature to
 400C, giving an overall efficiency of about 80
 percent.

   The addition of a number  of thick fins to
 the walls increased  the surface area  to  1/2
 square foot and enabled the gas temperature
 to be reduced to less than 250C, giving an
 overall efficiency of about 87.5 percent.
   An important point is that these high effi-
ciencies are achieved without operating with
metal temperatures below about 110C. This
is  possible  because of the  very high  heat
transfer  coefficients  between  the  fluidized
solids and the metal surfaces.  These coeffi-
cients result in the significant advantage that
condensation of the exhaust gases on the fins
does not  occur and corrosion is  eliminated.
This is not the case if highly extended surfaces
are used in normal convective heat transfer.

   Figure 9 shows a prototype domestic fluid-
bed  boiler  incorporating  blowers,  controls,
safety cut-outs, etc.  This latter unit is now
carrying out endurance trials to determine the
rate of loss of particle mass during extended
running.

   With  regard to particle life, experiments
have already been conducted at room temper-
ature using sand in a fluidized bed operated at
2 ft/sec; no measurable attrition loss occurred
in 500 hours. Whether or  not the continual
heating and cooling of particles in a combus-
tor with the consequent thermal shocking will
produce   more  severe  attrition  is  as  yet
unknown.

   Apart from the very effective heat transfer
in the second-stage bed, a  further significant
advantage is that it produces an extremely
effective  silencer for the combustion system.
Provided, therefore, that  particles  to  resist
degradation can be  found, it  appears  likely
that this type of approach can provide central
heating units with the following characteristics:

1.  Very low NO x emission  5 ppm).

2.  Very low CO/CO2 ratios.

3.  Lower aldehyde formation than for normal
   flames due to the low temperature combus-
   tion.

4.  Efficiencies as high as 90 percent without
   significant extra cost and with no fear of
   condensation or corrosion.

5.  Very compact plant with overall ratings of
   100,000 Btu/ft3 obtained (including fans,
   motors and gas controls).
IV-1-6

-------
6. Use of well-established cast iron techniques
  which are cheap and produce long life, low
  maintenance units, to produce boilers.

7. Easily adjustable units for  all gases  and
  Wobble  irrelevant  number  and  flame
  speeds.

8. Good turn down ratio of  units without
   losing efficiency.

9. Units readily developed  to  burn oil  and
   with a little  development could possibly
   burn solid fuel, (but could still be switched
   back to gas easily).

10. Application   of techniques   to   visible
   fire/back boiler systems providing a central
   focus in the living area as well as full house
   heating.

SHALLOW BED HEAT TRANSFER

   Early studies  by the  Central Electricity
Generating Board showed that uneconomical
high pressure drops would be  incurred if an
attempt were made to use  plain tubes in a
fluidized  bed  as  a  straightforward heat
recovery system (i.e., one not using a combus-
tion reaction).  Early work  by Petri et al.4
showed that  if a tube were  provided with an
extended surface finning system with an area
15 times that of the base tube, then the overall
heat  transfer would be increased approx-
imately six-fold, i.e., an effectiveness of 40
percent. As the overall tube diameter would
not be more than doubled by  employing the
fins, the net effect would  be  a far  greater heat
flux  per  unit volume  of  bed.  Thus, the
pressure drop penalties when using extended
surfaces would be significantly reduced.

   Following up this idea,  a  new  form of
extended surface heat exchanger was built,
and some preliminary results were presented
to the Second  International Conference  on
Fluid-Bed  Combustion.  These  preliminary
results were obtained with the  extended  sur-
face systems in a comparatively deep bed with
fairly large  particles. Further  work  showed
that if this particular arrangement of vertical
fins was placed low down in a fluidized bed, its
performance was unexpectedly higher. Figure
7 shows the performance of vertical-finned
extended surface tubing operated in a  very
shallow fluidized bed compared with that for
1-in. tubing and with the results of Petri et al.
The graph plots the bed-to-metal heat transfer
coefficients  against  particle  size.  It  will be
noted that the shallow bed results lie above the
generally  accepted heat  transfer  coefficient
line. These  results do not necessarily imply
that the vertical surfaces have an effectiveness
of over 100 percent, although they  may in
some instances, but do show that shallow-bed
performance is superior to that of deep beds, to
which most  of the world's data on heat trans-
fer  relate.
  The reason for the superiority of the partic-
ular configuration  of  vertical  surfaces in
shallow fluidized beds is believed to be the
absence of large  bubbles in the system.  The
absence is partly because the  vertical surfaces
prevent lateral mixing of the gases  which
again  restricts   bubble  formation.   It  is
interesting  to   observe  that when  these
extended surface tube bundles are placed  in a
shallow bed they do not appear to disturb the
bubbling pattern.

  A  further possible  explanation  for  the
improved heat transfer is that the viscosity of
shallow fluidized beds varies almost directly as
the  bed depth. Shear stress/shear strain data
derived from a Stormer-type viscometer with a
hollow cylindrical rotor is given in reference 5.
Because of the better fluidization  in  shallow
beds, the resistance to shear  of the fluidized
solids is much less. As we know that fluidized-
bed heat transfer depends upon  the  rate of
exchange  of particles  at the heat  transfer
surfaces, it would be logical to expect that heat
transfer would  improve  if particle mobility
improves.  Thus,  shallow beds   would  be
expected to  be superior to deep beds  from a
heat transfer aspect.

  Coupling extended surfaces with extremely
shallow beds  has been  patented  with  the
concept  that  we  need  no  longer  regard
fluidized beds  as isothermal devices.   This
enables the  overall thermal effectiveness  of a
                                                                                    IV-1-7

-------
 fluidized bed to be improved by the correct
 design of the heat transfer bundle/distributor
 unit.

    As pointed out at the last conference here,
 in  contrast  to   conventional  convective
 extended-surface heat transfer systems where
 the improvement  in  heat  fluxes  per  unit
 volume is accompanied by a higher pressure
 drop,  the  use of fluid-bed extended surfaces
 increases the heat flux per unit volume and
 decreases  the  pressure  drop.  The  pressure
 drops in some of the shallow-bed units we have
 investigated are so low  that the system can
 now compete favourably with normal convec-
 tive heat transfer, even for gas turbine waste
 heat recovery where low pressure drops are of
 paramount  importance.  It  is  contemplated
 that a two or three stage fluid bed can be
 operated with an overall pressure drop of less
 than 12 inches  water  gauge.

    Theoretical and experimental studies of the
 mechanism of heat transfer between the fluid-
 ized solids and the vertical fins and along the
 fins themselves  are underway; a vast amount
 of data has been obtained for various proprie-
 tary extended surface tubing as well as for
 specially   designed  fin/tube arrangements.
 These data, which are at the moment being
 written up for presentation in the near future,
 are sufficiently complete to allow an economic
 appraisal of various forms of heat exchangers
 to be made. Further advances in performance
 are expected when a better appreciation of the
 various phenomena is  acquired.

   It is suggested that for fluidized-bed boilers
 or  for  combined  gas  turbine  steam  cycles
 where steam is generated in the high tempera-
 ture exhaust from the gas turbine, there  is
 already a  case  for  investigating the use of
 shallow    fluidized-bed    extended-surface
 systems instead of  normal  convective  heat
 transfer.

   Preliminary  trials of  a single stage unit
picking up  waste  heat from  a diesel engine
have been very encouraging. Heat transfer
coefficients of the bed  which was completely
covered by vertical fins were just as good as for
the  original laboratory  unit.  The  fluidized
solids and the fins became coated with carbon,
suggesting that there may be some possibility
of enhancing this effect to reduce pollution
from the sub-micron carbon  in the  exhaust
gases.  The unit  ulso acted as  a very effective
silencer.


FUTURE  RESEARCH AND
DEVELOPMENT

   Initial experiments into  burning distilate
oil in fluidized beds have been successful; it
seems likely that a dual fuel gas/oil system can
be developed quite quickly. The unit would
have no higher pressure'drop than the existing
gas-fired  unit  and,  therefore, would  have
advantages over normal pressure  jet burner
furnaces with regard to fan pressure and noise
levels.  Its pollution control level would be  far
superior.  It  is  expected  that the  unburnt
hydrocarbons will be much reduced compared
to normal  oil flames.

   Research  on small-scale  solid fuel-fired
units has also started; there appears to be no
insuperable problems in producing a shallow
open-hearth  solid  fuel fire giving radiation
outputs of over 60 percent. A high-efficiency
domestic fluid-bed solid fuel-fired  boiler also
appears to be a practical proposition.

   The extension of these ideas into the field
of packaged boilers is underway; it is expected
that  economically  viable   units  can  be
developed.

   In addition to the low pressure, hot water
boiler developments, studies of high tempera-
ture, high pressure steam  systems  indicate
that such units can be designed to startup and
operate satisfactorily and economically. These
units would  employ high temperature  alloy
tubing  capable  of  operating dry  during
startup,  thus avoiding heat  losses  to the
cooling system during startup.A unit capable
of producing steam for a 150-hp engine would
have a diameter of approximately 2 feet with a
combustion bed 6-in. deep followed by a
further 6-in.  deep  economiser. With shallow
rv-i-8

-------
fluidized-bed extended  surface  systems  the
bed depths would be even smaller.

CONCLUSIONS

   Domestic  fluidized-bed combustion/heat
transfer systems which are no more expensive,
are just as compact, and have far less pollution
than  their conventional  counterparts have
been developed.

   Endurance and reliability trials are in pro-
gress so that by the end of 1972 we should be
able to assess the full potential  of gas-fired
fluidized-bed combustion as applied to small-
scale boilers. The results so far  suggest that
additional work should be undertaken on oil-
and coal-fired systems.

   Fluidized-bed   combustion   and  heat
transfer lends itself well to metallurgical heat
treatment processes in which the antipollution
aspects combined with rapid processing have
been  shown  to  lead  to  environmental and
economic benefits. Batch processing furnaces
are now  available,  and  in-line continuous
furnaces show distinct promise.

   The development of shallow bed, extended
surface heat transfer systems promises to open
up a completely new field in heat  recovery and
could help to reduce costs and space require-
ments in many types of  plant.
   It  appears,  therefore,  that  fluidized-bed
combustion and heat transfer techniques can
be  usefully   employed   for   antipollution
measures  over the whole range  for units
having an output of a bunsen burner up  to
extremely large power station sizes; ,in many
cases  it will  be accompanied  by economic
benefits.

REFERENCES

1. Ukilov, V. M., G. K. Rubtsov, and A.  P.
   Baskakov. Gas Combustion in a Packing
   under   a   Fluidized   Bed.   Gazovaja
   Promyshlennost.  5, 1969.

2. Tamalet, A. J. Application of Fluid  Bed
   Heat Transfer to Metallurgical Processes.
   Inst. Chemical  Engineering  Symposium
   Series.  In:  Proceedings  Symposium  on
   Chemical  Engineering  in Iron and Steel
   Industry, pp. 105-114,  1968.

3. Virr, M. J.  and R. Reynoldson. Heat Treat-
   ment in Fluidized Beds. Industrial Process
   Heating.

4. Petrie,  J.  C, W. A.  Freeby,  and J. H.
   Buckham.  Bed Heat Exchangers. Chemical
   Engineering Progress,  64(7),  1968.

5. Botterrill, J. S. M., M. Van der Kolk, D. E.
   Elliott, and  S. McGuigan.  The  Flow  of
   Fluidized  Solids.  Powder  Technology,
   6:343-351, 1972.
                                                                                  IV-1-9

-------
  POROUS CERAMIC
   DISTRIBUTOR
                    GAS/AIR
0   40   80  120  160   200  240  280  320
               TIME, sec
 Figure 1. Radiant cooled shallow fluidized bed.      Figure 2. Start-up temperature/time
                                                     sequence.
  OQ
       40
                                           FREQUENCY, Hz
         Figure 3.  Noise levels 2 ft away from radiant  bed (including unsilenced fan noise).
IV-1-10

-------
   1000
o
o
LU
a.
                                   20
40
                            30

                        TIME, sec

Figure 4.  Heating of 1/4-in.-diameter alloy steel bolt.
50
60
                                                                                         IV-1-11

-------
  ,
  UJ

  o
  o
  =t
  cc
                                          -	PROPANE     STOICHIOMETRIC


                                                        METHANE     STOICHIOMETRIC
                                                        TOWN'S GAS ( )

                                                        STOICHIOMETRIC MIXTURE STRENGTH
       50
       45
       40
         1000
1100
1200
1300
1400
                                           TEMPERATURE.'K

               Figure 5.  Percent radiation from shallow fluidized bed combustors.
IV-1-12

-------
24,000
20,000
 16,000
 12,000
                            800
        900
TEMPERATURE,0 C
1,000
   Figure 6.  Input and radiant output from a 6-in.- diameter shallow bed combustor (based
   on gross calorific value).
                                                                                    IV-1-13

-------
      600
      500
      300
      200
      100
                          ASTON DATA ON EXTENDED SURFACE/SHALLOW BEOS
                                                  1-irt. x 2-in. dia. TUBE DATA (DEEP BEDS)
                                       PETRIEETAL
                                       EXTENDED SURFACE
                                       (DEEP BED)
                                  550                        1000

                                        PARTICLE DIAMETER, fin

                      Figure 7. Comparison of shallow/deep bed heat transfer.
1500
IV-M4

-------
      2. FLUIDIZED-BED COMBUSTION UTILITY POWER
         PLANTS-EFFECT OF  OPERATING AND  DESIGN
   PARAMETERS  ON PERFORMANCE  AND  ECONOMICS

    D. L. KEAIRNS, W.  C. YANG, J. R. HAMM, AND D. H. ARCHER
                  Westinghouse Research Laboratories
ABSTRACT

   Pressurized fluidized-bed  boiler power  plants have the potential to  meet SO2, NO, and
participate emission standards at energy costs 10 percent below conventional plants with wet
scrubbing. This paper analyzes the sensitivity of the operating and design parameters selected for
the plant design on plant performance and economics. Results show that the plant costs and per-
formance are essentially invariant with projected changes in operating and design parameters2.5
percent change in energy cost. The concept has the potential for achieving plant efficiencies of
~45 percent.
INTRODUCTION
   A pressurized  fluidized-bed boiler power
plant has been designed using state-of-the-art
power generation equipment.' -2 Performance,
costs, and pollution abatement were projected
for the system. The results show the concept
has the potential to meet SO2, NO, and par-
ticulate emission standards and may reduce
energy costs 10 percent below a conventional
plant with stack gas scrubbing.
   Operating conditions  and  design param-
eters for the pressurized  boiler were selected
based on an  evaluation of available  data,
power cycles, and alternative boiler concepts.
It is important to know how sensitive the oper-
ating and design parameters selected for the
base design are to the plant  economics. An
understanding of the effect of changes in the
proposed design  on plant cost will provide a
basis for evaluating current pressurized fluid-
bed combustion  pilot plant data, planning
experimental programs, designing the devel-
opment  plant,   and  understanding  the
economic margin  for solving technological
problems.
   The sensitivity analysis evaluates the effect
of the following variables on plant design,
cost, and performance:

Fluidized bed boiler operating conditions
    Bed temperature
    Fluidizing velocity
    Excess air
    Pressure

   Fluidized bed boiler design
    Heat  transfer  surfaceconfiguration,
    heat transfer coefficient, and materials
    Module capacity

   Particulate carry-over from the boiler
    Loading    *
    Size distribution

   Power   plant    equipment    operating
    conditions
    Gas turbine inlet temperature
    Steam temperature and pressure
The evaluation is  performed by considering
each variable separately; it is general in order
to permit the  coupling  of different effects to
assess alternative designs. It is performed to
indicate relative effects of variable  changes in
plant costs relative to each other and the total
plant cost.
                                      IV-2-1

-------
 BASIS FOR SENSITIVITY ANALYSIS

   The basis for the sensitivity analysis is the
 boiler  and   plant   design   developed  by
 Westinghouse under contract to EPA. * '2 The
 power plant cycle is shown schematically in
 Figure 1. The plant subsystems included in
 this sensitivity analysis are enclosed within the
 broken  lines.  The  pressurized boiler  was
 designed  by  Westinghouse   and  Foster
 Wheeler and is shown schematically in Figure
 2. The preliminary  boiler design was for a
 nominal 300-MW plant. The  boiler design
 consists of four  modules;  the  modularized
 design provides for a maximum of shop fabri-
 cation  and  turndown  requirements.  Each
 module  includes  four primary  fluidized-bed
 combustors, each containing a separate boiler
 functionone bed for the pre-evaporator, two
 beds for the superheater, and one bed for the
 reheater. Evaporation takes place in the water
 walls. All of the boiler heat transfer surface is
 immersed in the beds, except for baffle tubes
 above the bed to minimize particle carry-over.
 Each module contains a separate fluidized bed
 or carbon burn-up cell to complete the  com-
 bustion of carbon elutriated from the primary
 beds. The philosophy used to design the boiler
 was to maximize shop fabrication. Thus, the
 300-MW plant utilizes boiler modules which
 can be  completely  shop  fabricated. From
 roughly  300  to 600  MW,  the boiler  can  be
 partially shop fabricatedthe pressure shell
 being  too  large  for rail  transport.  Larger
 plants,  utilizing  the  four-module concept,
 would be field erected.

   The operating conditions and design para-
 meters for the boiler and the power cycle are
 summarized in Table 1. The power plant per-
 formance and economics were based on these
 specifications. The cost  breakdown for the
 fluidized-bed steam generator is summarized
 in Table 2. The fluidized-bed  boiler design
was scaled to 600-MW capacity and the  costs
estimated.  These  costs  are  summarized  in
 Figure 3. A breakdown  of the  power plant
equipment  costs  for  a  635-MW  plant  is
presented in Table 3. Limestone or dolomite
regeneration is not included in this analysis.

IV-2-2
Thus the  costs presented  are for a  once-
through system. The energy costs used for this
analysis are also presented  in Table 3. The
costs of a conventional plant with wet scrub-
bing on the same basis are also indicated.

   The following assumptions are made for
the sensitivity analysis:

1. The plant  concept  maintains  the four
   module with two modules per gas turbine
   concept.

2. Coal feed rate  is maintained constant for
   each  variable  analysis.   Thus  the  coal
   feeding and handling system is assumed to
   remain unchanged.  This  may  not  be
   completely true if the bed area is changed
   significantly and the number of feed points
   increased or decreased. The cost of the coal
   feed system is considered if the bed design
   is altered.

3. Structural and erection costs are constant.
   The structural steel and concrete costs for
   the boiler plant equipment are~$2/kW.
   The maximum change in these costs for the
   cases considered is~$0.20/kW and will be
   significantly less in general. The  cost was
   thus assumed constant for this  analysis.
   Erection cost changes  are  negligible.

4. Coal and   stone  feed  size are  assumed
   constant1/4 in. x 0. This parameter is
   important when considering particle carry-
   over,   but  insufficient  information   is
   available to permit a quantitative analysis.

5. The ash and dust handling system cost is
   assumed  constant.  This   cost  will  be
   affected by the particulate carry-over, but
   it is considered a second order effect for the
   cases considered.

6. Stack and foundation, instruments  and
   controls, and other costs are constant.

7. All  variables  not  being  evaluated are
   assumed constant unless  stated otherwise.

-------
   Table 1. PRESSURIZED FLUIDIZED BED BOILER POWER PLANT
              OPE RATING AND DESIGN CONDITIONS
Cycle

   Steam system
   Gas turbine expander
     Pressure ratio
     Inlet temperature
     Air cooling

    Coal feed rate
    Number of boiler modules

    Boiler modules/gas turbine

    Fuel/air ratio
2400 psia, 1000F superheat, 1000F reheat

10:1
1600F
5%

53,910 Ib/hr/module for nominal 300-MW  plant
design

4

2

0.0919
 Boiler design
    Bed area

    Heat transfer surface
      Walls

     Bed
    Gas side heat transfer
     coefficient

    Tube materials
    Bed depth (expanded)

    Gas temperature drop from
     primary beds to gas
     turbine expander
35 ft2 (5x7 ft) -for ~80-M W module
2-in. OD tubes on 3-1/2 in. welded wall spacing

1-1/2-in. OD tubes in pre-evaporator and  super-
heater; 2-in. OD in reheater (details in text)

50Btu/hr-ft2-F
SA-210-A1 pre-evaporator
SA-213-T2lower superheater
SA-213-T22water  walls;  upper  superheater
(lower loops); reheater
SA-213-TP304H upper   superheater   (upper
loops)

11to14ft

150F
 Boiler operating conditions

    Bed temperature (100% load)  1750F
    Fluidizing velocity

    Excess air

    Particle carry-over
     carbon from primary beds
 6 to 9 ft/sec

 17.5%

~7 gr/scf
 6% of carbon feed
 Auxiliaries

    Coal feed system

    Primary particulate
     removal

    Secondary particulate
     removal


    Stack gas coolers
Petrocarb feed system

4 size 355 VM 8/0/150 Duclone per module
nominal  300-MW design

2 model 18000 Type  S collectors per module-
nominal 300-MW design (quoted by Aerodyne Dev.
Corp.)

Conventional heat exchanger design.
                                                                                 IV-2-3

-------
 Table 2. COST OF A 318-MW PRESSURIZED
             FLUID-BED BOILER
Pressure parts3
Shell
Subcontracted and contracted
equipment
Drafting
Home office
Sub-total
Erection
TOTAL
$1,777,000
935,000
435,000
185,000
685,000
$4,017,000
500,000
$4,517,000
 a Pressure  parts include tubing  cost, headers,
 downcomers, risers, tube bending, tube welding,
 and  water wall fabrication.

 b Field erected ($3,856,000 shop assembled).

 BOILER PLANT EQUIPMENT

 Operating Conditions

   Bed Temperature

   The full load design temperature is 1750F.
 Lowering  the  design bed  temperature
 increases the total amount of heat transferred
 in the  bed  and thus  increases  the steam
 turbine power generation. At the  same time,
 lowering bed temperature decreases the gas
 turbine  inlet   temperatureassuming  no
 burning above the  bedand thus decreases
 the total gas turbine  power generation. The
 decrease in gas turbine power is larger  than
the increase in steam power, resulting in an
overall decrease in plant power  (Figure 4).
Lowering the bed temperature  increases the
total amount of heat  transferred  in the bed
and thus requires more heat transfer surface.
Assuming the cross  sectional area  of the fluid
bed (5 x 7 ft for a 300-MW nominal plant size)
and  tube size/tube pitch are  constant, the
expanded bed depth for each functional bed
increases with decrease in bed temperature as
Table   3.  PRESSURIZED   FLUIDIZED-BED
     BOILER POWER  PLANT COSTS

             Equipment Costs
                                               Component
                                Cost,$/kW
 Boiler plant equipment

  Boiler
  Particulate removal
  Piping/ducts
  Stack and foundation
  Coal handling and feeding equipment
  Ash and dust handling system
  Instruments and controls
  Miscellaneous equipment

 Steam turbinegenerator equipment

 Gas turbinegenerator equipment

 Other: land, structures, electric plant
      equipment, miscellaneous plant
      equipment, undistributed costs
Subtotal
                                                                                    14.49
                                                                                    12.76
                                                                                     4.43
                                                                                     0.47
                                                                                    14.94
                                                                                     1.55
                                                                                     3.10
                                                                                     0.94

                                                                                    44.14

                                                                                    14.80
     70.38
    182.00
Total capital cost (inc. escalation, IDC etc.)  265.00

($340/kW for conventional plant with wet scrub-
bing on same basis)
Energy costs
  Fixed charges
  Fuel
  Dolomite
  Operating and maintenance
mills/kWhr
     6.44
     4.04
     0.52
     0.71
     11.71
(13.45 for conventional plant)

shown in Figure 5. The bed depth of the two
superheater beds is assumed  to be the same
for convenience. This will not affect the total
heat transfer  surface requirement  and the
resultant module height shown in Figure 6.
The  bed depth and module height can  be
reduced by enlarging the bed area and module
diameter. However, since the module diameter
of 12 feet is considered to be the largest ship-
pable  railroad  size,  increase  in  module
diameter  to  accommodate  additional  heat
transfer surface may not be  economic for a
300-MW plant.
IV-2-4

-------
   The effect of changing design bed tempera-
ture on the steam generator cost is calculated
based on the boiler cost estimation shown in
Table 2  and on the assumptions that  the
module diameter is constant at  12 feet and
that the total number of modules is four based
on turndown  consideration.  The cost  (not
including  erection) of the 4-module steam
generator  as  a function of bed temperature
with constant module diameter (12ft) is shown
as curves la  and  Ib in Figure 7. Curve la
assumes the maximum allowable bed depth to
be 20 feet. That means any bed with expanded
bed depth larger than 20 feet will have to be
split into two beds  with  their separate air
plenums  and  freeboards.  Curve  Ib assumes
that there is no restriction on maximum bed
depth. The choice of 20 feet as the maximum
allowable bed depth is arbitrary, just to show
the importance of this variable on the cost of a
steam generator.  It is doubtful that the bed
depth of each fluid bed can be  unrestricted
without   creating   undesirable   bubble
formation and  slugging, poor bedtube  heat
transfer  coefficient,  and  temperature
gradients in the bed at some bed depth. The
maximum  allowable bed  depth  at  specific
operating conditions will  have to be experi-
mentally determined in a large unit. Without
the required experimental evidence, the steam
generator  cost  (not  including  erection) is
plotted against the maximum allowable bed
depth in Figure 8. The bed temperatures were
calculated by assuming the gas turbine  tem-
peratures of 1600, 1500,1400, and  1300F and
by assuming a linear temperature loss between
the boiler and the gas turbine inlet. The cost of
the  steam generator designed  for  1636F
increases ~ 20 percent ( ~  $2.8/kW) over that
designed at 1750F if the maximum allowable
bed  depth is  15 feet.  The  cost  increase is
primarily due to the splitting of beds with bed
depth higher than 15 feet. Bed splitting can be
avoided  by  either  decreasing  boiler  tube
diameter and spacing in the bed or increasing
the  module  diameter.  Decreasing   tube
diameter and spacing will change the bed-tube
heat transfer coefficient,  tube bending and
fabrication,  and  tube  wall   thickness.
Increasing the module diameter will not only
change the cost of the pressure shell but also
affect   constructioncomplete   shop-
assemblage versus partial field erection. All
these factors have to be taken into account in
designing an optimal boiler. These factors are
discussed in separate  sections.

   In the present cost estimation, the possi-
bility of using thinner  wall tubes  for  the
designs  at  lower bed  temperatures was also
taken  into  account  by   calculating   the
minimum tube wall thickness requirement. No
allowance for corrosion is provided.

   Change in the  operating bed temperature
will also change the gas temperature to both
the primary and secondary cyclones and  thus
change  the actual volumetric gas flow rate.
This will change  gas  inlet velocity to  the
cyclones which,  in turn,  affects  cyclone
collection efficiency. This  effect was estimated
to be small compared to the effect of the
change in pressure drop across the bed due to
a  change in the  design bed  temperature.
Decreasing  the  design   bed  temperature
increases the heat transfer  surface require-
ment in  the bed, which requires an increase in
pressure drop due to an increase in bed depth
if bed area and boiler tube configuration are
constant. The decrease  in  net  plant power
output  as  a  function  of the  design  bed
temperatures is presented  in curve 1, Figure 9.
The effect is smalla decrease of only  ~ 0.3
percent   if the design bed temperature is
reduced to 1407F.

   Bed  temperature is one of the primary
variables  used for load  turndown  in  the
present design. A 4:1 turndown can be met if
the design bed temperature is higher than
1600F.  The   primary   limitation  on   the
operating  bed temperature is  the  sulfur
removal efficiency of the sorbents in the bed.
At bed  temperatures higher than 1750F or
lower  than  1350F,  the  sulfur  removal
efficiency in the bed is too low.  Thus  it is
concluded from the above bed temperature
analysis  that  the  design  bed  temperature
should be the highest temperature required
                                                                                  IV-2-5

-------
for desirable  degree of load turndown and
sulfur removal in the bed.

   Fluidizing Velocity

   At constant fuel feed rate and excess air,
increasing the fluidizing velocity requires a
decrease in the bed area and in the module
diameter (Figures 10 and 11). For a constant
overall heat transfer coefficient and a specific
design bed temperature the total heat transfer
surface in the bed is constant; a decrease  in
the bed area will require an increase in the bed
depth at constant tube size and tube spacing
and thus an increase in the module height. An
economic design will depend on the balance of
thse  factors.

   The bed area and bed depth requirements
with respect to change in fluidizing velocity at
different design bed temperatures are calcu-
lated. The corresponding module height and
module diameter are presented  in Figure 11.
The  cost  of the pressure shell  at  different
inside diameters is  estimated based on the
data  from Foster  Wheeler Corporation' and
on an independent estimation  by Westing-
house (Figure 12).  The  discontinuity at  a
module inside diameter of 12  feet is due to the
cost difference between  the  shop assembled
and the field  erected shell.

   One additional cost which has to be taken
into consideration is the  fabrication cost.  In
addition to the shell cost, the change of fabri-
cation cost  of water walls and tube bending
cost  are  not   to  be ignored.  Taking  into
consideration the factors involved, the steam
generator cost is plotted against the superficial
fluidizing velocity in the pre-evaporator ash
shown in Fugure 13 for 318-MW and 635-MW
plant. The  superficial  fluidizing velocity  in
the pre-evaporator is used here since it is the
largest velocity  in all beds  inside  a single
module. The superficial fluidizing velocity  in
the superheaters and reheater can be calcu-
lated  accordingly, the  results  show that
increasing  the fluidizing velocity  tends  to
increase rather than decrease the total steam
generator   cost  at a  318-MW  plant size.
Decreasing the fluidizing velocity in the pre-
evaporator below  ~ 8 ft/sec requires a shift
from the shop-assemblage to the field erection
and escalates suddenly the steam generator
cost. A minimum cost does exist for a 635-
MW plant. Figure 13 is for unrestricted max-
imum allowable bed depth. If the maximum
allowable bed depth is limited to say 10 or 20
feet, the disadvantage of increasing the fluid-
izing velocity would be even larger at 318-MW
size. At 635-MW size, the minimum would
shift to lower velocity.
   The important thing here is to understand
why increasing the fluidizing velocity increases
the steam generator cost at 318-MW size and
produces a minimum at 635-MW. To better
illustrate  the  point,  cost  reduction  due  to
decrease  in   module diameter  and  cost
escalation due to increase in module height for
a four-module design  are shown in Figure 14
for the design bed temperature at 1750F.
Increasing the fluidizing velocity escalates the
steam generator cost almost linearly from the
basic design point due to increase in module
height (curve 2). At the same time,  the cost
decreases due. to decrease in module diameter;
however, the  decrease  is much more gradual
and levels off at higher  fluidizing  velocity
(curve 1).  This is because the  bed area alone
occupies less than 40 percent of the total cross-
sectional area of a pressurized module.  The
remaining area  is  required for  piping and
headers; this space is relatively unchanged at a
specific plant size even though the bed area is
reduced to increase the fluidizing velocity. At
635-MW size, however, the cost reduction due
to  decrease  in  module   diameter is larger
during initial deviation from the basic design
point and thus creates a minimum (Figure 14).

   Another approach  for analyzing the effect
of fluidizing velocity would be to change the
number  of modules  as well  as  the  module
diameter. It is preferred  to have a 5-module
design  based  0:1  turndown   consideration;
however, if a  3-module design shows  a sub-
stantial saving with negligible effect on turn-
down capability, it would be a better  choice.
IV-2-6

-------
Change in fluid izing velocity will change the
total bed area required  for  each functional
bed, but the total bed volume for each func-
tional bed will remain constant once the tube
size and spacing are fixed. Thus a design with
smaller number of  modules  will require a
larger  module diameter  at  the same  design
fluidizing velocity.  An estimation can usually
be performed to  evaluate the relative economy
between these  two  designs.   For example,
consider a 4-module and a 3-module design at
the same design fluidizing velocity and with
the same bed depths. Since the bed volume of
each functional bed is constant at fixed tube
size and spacing for both cases, we have
             D
3_
4
                                        (1)
 if bed height is  assumed  constant  for both
 cases  and the  bed  area  occupies a  fixed
 percentage of the total cross-sectional area of
 a module.  04  and  DS  are the  respective
 module diameters for the 4-module and  3-
 module  designs. Since  the  shell  cost  is
 dependent on the vessel diameter (Figure 12),
 the relative advantage of these two designs will
 depend on  the  plant size in  question. For
 example,  for D4=12.5 feet, D3 can be calcu-
 lated from equation (1) to be 14.4 feet. The
 shell cost can be found from Figure 12 to  be
 $0.94 x 106 for the 4-module design and $0.79
 x 106  for the 3-module design. If fabrication
 cost of the module internals is similar in both
 cases,  the 3-module design will have a  slight
 economic advantage. However, this advantage
 becomes  progressively  smaller  because  of
 rapidly increasing shell cost  at large module
 diameter and rapidly decreasing shop-fabrica-
 table portion in the design. It is estimated that
 the  largest  module  diameter which is still
 economic for the 3-module design is ~ 17 feet.
 This conclusion is based  on  the  assumption
 that boiler turndown is not a problem. If the
 module diameter is in  the shop-fabricatable
 and railroad -transportable range,  i.e.,  <  12
 feet, designing for the maximum shippable
 module will have definite advantages provided
 that turndown is not a problem.
   It is concluded from this analysis that for a
plant size around 300 MW,  the  module
diameter should be the largest within shipping
limitations (12 feet for railroad transporta-
tion); at a 600-MW plant size,  an optimum
fluidizing velocity exists,  and  it  should  be
found for each capacity.  However, the cost
deviation from that of the optimum design is
less than  Sl.OO/kW (Figure  13). Of course,
decreasing the bed  area  may  reduce  the
number of feed points but the saving is only ~
$0.10/kW.  Although   the   effect  of  the
fluidizing velocity on the steam generator cost
is  primarily  based  on  the  basic  design
conditions,  i.e.,   bed-tube  heat transfer
coefficient  =  50 Btu/ft2-hr-F, the  trends
would be the same for bed-tube heat transfer
coefficient at 35 or 75 Btu/ft2-hr-F. Change
in tube size and spacing may change the slope
of the curves or alter the minimum in Figure
13; the  conclusions will remain the same.

   The effect of changing fluidizing velocity
on  the  bed-tube  heat transfer coefficient,
combustion efficiency, and  total particulate
carry-over is not  taken into  account  in this
analysis due to lack in  accurate quantitative
data. The effect of dust loading and particle
size distribution on the cost of the particulate
removal system is evaluated separately.

   Excess Air

   Change in design excess air will affect the
cycle  efficiency and the cost  of the  boiler
module, the steam and gas turbine equipment,
and the particulate removal system. In order
to quantify  the effect  of excess air on total
boiler cost, the air/fuel ratio is allowed to vary
with the total fuel input  kept constant.  To
simplify the analysis,  other parametersbed
temperature, tube size and tube spacing, bed-
tube  heat  transfer  coefficient,  fluidizing
velocity, and number  of boiler  modulesare
held constant  at the basic design values.
   Thus, when excess  air is increased beyond
the design value, bed area has to be increased
if  the  fluidizing  velocity is kept constant.
When the bed area is increased, the module
diameter has to be increased as well; however,
                                                                                   IV-2-7

-------
 the  module height is decreased due to  a
 decrease in bed  depth.  Increasing air input
 into the bed will  also increase the amount of
 heat carried out  from the bed  by air and
 reduce the total heat transferred in the bed. At
 constant bed-tube heat transfer coefficient,
 the total heat transfer surface requirement is
 reduced.  At  constant tube size and tube
 spacing, the bed depth is also reduced  as well
 as the total module  height. At  100 percent
 excess air, a reduction of > 40 percent in heat
 transfer surface  and  a reduction  of  >30
 percent in module height is possible. The bed
 depths for  different  functional beds  are
 reduced to  ~4 feet which decreases pressure
 drop through  the beds and increases cycle
 efficiency.  The module diameter,  in turn,
 increases  from the original  12 feet  inside
 diameter to more than 16 feet for an 80-MW
 module. Transferring all these  changes into
 economics,  an increase in excess  air can
 reduce the  boiler cost  up  to ~$0.60/kW  as
 shown in Figure 15. Cost reduction due  to heat
 transfer surface  increases  continuously with
respect to excess air because of the decrease in
the total  amount of heat transferred in the
bed. Cost reduction due to pressure shell first
increases  because  of reduction  in  module
height and then decreases because of increase
in module diameter.
   Increasing excess air will increase the over-
all plant efficiency as shown in Table 4. Larger
gas turbines or additional gas turbines  are
needed  to handle the increased mass flow of
gas. If additional units are used, the increase
in gas turbine equipment cost  is  shown as a
cost adder in Figure 16. This does not account
for cost reductions which can be realized by
going to larger turbine capacities.  This cost
increase in gas turbine equipment is partially
offset by a decrease in steam turbine equip-
ment cost also shown in Figure  16. The major
equipment items  taken into consideration in
this analysis include gas turbines with external
manifolds, steam turbine system, circulating
water and  condensing  systems,  feedwater
          Table  4.   EFFICIENCY CALCULATIONS AT  VARIABLE AIR FLOW  RATES
Excess
air,
%
17.5
50.0
90.0
Fuel/air
ratio
0.0861
0.0674
0.0532
Plant
size,
MW
644.1
648.0
649.7
Heat
rate,
Btu/kWhr
9026
8972
8948
Gas
turbine,
MW
127.9
156.9
193.5
Steam,
MW
532.2
506.3
470.4
Gas
turbine/
steam
ratio
0.240
0.309
0.411
No. of
gas
turbine
2.000
2.450
3.025
IV-2-8

-------
system including station piping, and stack gas
coolers.

   Instead  of keeping the fluidizing velocity
constant, the bed area and module  diameter
can be kept constant and allow the fluidizing
velocity to increase  with excess  air. In  this
case,  the  cost  reduction  in  heat  transfer
surface will be similar, but the cost reduction
in the  shell will continuously increase with
increasing  excess air and not g6 through  a
maximum  (Figure 15). The cost reduction in
heat transfer surface and pressure shell at 100
percent excess air in this case (with 10 to 15
ft/sec fluidizing velocity) is estimated  to be
 ~$1.00/kW. However, increasing  the  fluid-
izing velocity to larger than 15 ft/sec may be
impractical in this design approach.

   Increasing the excess air will decrease the
total heat transfer surface required in the fluid
bed  until  no  boiler  tube surface  will be
required at an excess air of approximately 300
percent. In this case the power system would
become a combined cycle plant with the gas to
the turbine expanders supplied from a coal-
tired, adiabatic combustor.  The heat recovery
boiler would probably be unfired. This system
concept has  several  significant differences
from a  pressurized fluidized-bed boiler power
plant   concept:   for  example,  the  boiler
becomes an adiabatic combustor, particulate
removal equipment costs increase significantly
due to the  increased gas flow, gas piping costs
increase and the gas turbine power contribu-
tion increases from  ~ 20 percent up to  ~ 70
percent. An economic analysis of this system
has not been made as part of this evaluation.
The  heat  rate for  the  adiabatic combustor
plant is projected to be 100 to 500 Btu/kWhr
(depending on the gas turbine inlet tempera-
ture) greater than for the  pressurized  boiler
plant. Further evaluation of this high excess
air case is  required to perform a comprehen-
sive assessment.

   Increasing the excess air will provide more
flexibility in turndown. At 100 percent excess
air, an additional ^ 10 percent load reduction
is  possible as compared to operation at  10
percent  excess  air.  This  means  a  boiler
designed at 1600F and 100 percent excess air
will be able to meet a 4:1  turndown require-
ment. An adiabatic combustor system should
extend the turndown capabilities.

   More discussion on excess air and gas flow
rate  appears  in  the section  on particulate
removal.

   Operating Pressure

   The  full load design pressure is 10  atm.
When the design pressure level is changed and
the  other   operating   parameters  remain
constant, the  gas  density will change  in
proportion to the pressure,  and the volumetric
flow will vary accordingly. Therefore, the bed
cross-sectional area will have to be changed to
maintain constant fluidizing velocity and the
bed depth changed to maintain constant bed
volume.  The  heat transfer coefficient  may
change  because of changes in the  quality  of
fluidization.

   The changes in volumetric flow and in gas
density will affect the design of the particulate
removal  equipment.  Gas  turbine cycle
efficiency is also dependent on the  operating
pressure.  However,  analysis   of  the   high
pressure fluidized-bed coiler system indicates
a  reverse direction  in  pressure  level  effect.
Auxiliary equipment such as the coal and dol-
omite feeding systems are pressure dependent
as well. The relative importance of  these fac-
tors with respect to operating pressure is  ana-
lyzed in the  following paragraphs.

   First, consider the boiler module alone.  At
constant  fuel  feed  rate   and  excess  air,
increasing the operating pressure will decrease
the gas volumetric flow rate.  There are two
approaches in designing the boiler modules:
(a) keep the bed  area  constant and let the
fluidizing velocity change with the operating
pressure, and (b) keep the fluidizing velocity
constant and change the bed area according to
the  operating  pressure.  The  incremental
module cost for the constant bed area case is
                                                                                     IV-2-9

-------
the cost of reinforcing  the pressure shells.
Since  for  a constant overall  heat  transfer
coefficient and a specific design bed tempera-
ture, the total heat transfer surface in the bed
is constant. This amounts to $0.20/kW and
$0.30/kW for operating pressures of 15 and 20
atm, respectively, at 300-MW nominal plant
size. At 600-MW plant size, the respective cost
increments are $0.10/kW and $0.20/kW. The
credit  of decreasing particulate carry-over  by
operating pressure will require a decrease in
bed area. Since the total heat transfer surface
is constant, a decrease in bed area will require
an increase in bed depth at constant tube size
and tube spacing, and thus an increase"nr the
module height. In this case,  the  effect  of
operating pressure on boiler design and boiler
cost while  keeping  the fluidizing  velocity
constant is the same as the effect of changing
fluidizing velocity with the operating pressure
constant. Thus the Figures 10, 13, and 14 can
also be applied  for  this  approach if  the
coordinate  for the   fluidizing  velocity  is
changed to (fluidizing velocity at basic design)
x  (new  operating  pressure/basic  design
pressure). For  a design  bed temperature  of
1750F, the cost  increments are  $0.80/kW
and $1.90/kW for  operating pressures of 15
and 20  atm, respectively, at 300-MW plant
size. At 600-MW  plant  size, the cost  incre-
vents  are $0.20/kW  and $1.20/kW respec-
tively.
   Comparing these two design approaches,
the constant area case is the less costly one.
Moreover, if  basic  design  bed  area  and
fluidizing velocity are maintained,  increasing
operating pressure will mean a higher capacity
shop-fabricatable  module  (i.e.,  module
diameter   <12  ft).  At  15  atm  operating
pressure,  a module  of   ^120-MW capacity
can be shop-fabricated;  at  20 atm  pressure,
the  maximum shop-fabricable module  is
~ 160 MW.  However, the module height will
be considerably increased because of increase
in total heat transfer surface required. The
total increase in module height will depend  on
 the heat transfer surface arrangement in the
 bed.

   Increasing the operating pressure reduces
the size of the particulate removal equipment
because  of the  decrease in volumetric flow
rate.  It  reduces  the  particulate  removal
efficiency as well because of  changes in gas
density and  viscosity. Increase in operating
pressure will also require reinforcement of the
containment  vessel  and  the  piping  and
ducting.   The savings  in  the  particulate
removal equipment by operating at 15 and 20
atm   are   estimated  to  be  $4.0/kW  and
$6.0/kW, respectively, for a  300-MW  plant
size.  The  major  saving  comes  from  the
secondary  collectors where maximum  single
unit capacity  is assumed restricted to 30,000
acfm. An increase in operating pressure  will
result  in  fewer  units.  Cost  reduction   by
operating  at  higher  pressure .will  be even
greater  for larger  plant  sizes or  at higher
design excess air.
   Cycle  optimization  calculation  was
 performed to evaluate the effect of operating
 pressure. The parameters studied  are: inter-
 cooled and non-intercooled compression; gas
 turbine compressor pressure ratios from 10 to
 30; and cycle gas side pressure drops of 3 to 8
 percent. The results are summarized in Figure
 17 which gives plant heat rate  for the inter-
 cooled and non-intercooled cases. For the non-
 intercooled case, the best efficiency is obtained
 at a pressure ratio of 10; for the intercooled
 case, the optimum pressure ratio is 15 but with
 a higher heat rate and a more complex gas
 turbine.  Thus   an  increase  in  operating
 pressure  higher  than 10 atm decreases the
 overall plant efficiency. This decrease is small
 however:  ~ 0.2  percent  at 15 atm.


   Weighing the  above discussions, the only
 distinct advantage for operating at pressures
 higher than 10 atm is the capability of shop-
 fabricating a large capacity plant, especially at
 higher design excess  air.
IV-2-10

-------
Boiler Design


Heat Transfer Surface

  Configuration

   In the 300-MW design, the heat transfer
surface is provided by serpentine tubes having
the horizontal  sections spaced  as  shown  in
Figure 18. Tubes of 2-in. OD  are used  at
waterwalls and are space 3-1/2 inches  apart.
Tubes for pre-evaporator and superheaters are
1-1/2-in. OD and tubes for the reheater are 2-
in. OD. The tubes can usually be arranged in
staggered or rotated diamond arrays, or they
can be arranged in a square or rectangular
pitch  (Figure 19).
   The effect of tube pitch/diameter ratio on
the cost of the steam generator (not including
erection) for constant 12-ft module diameter
was evaluated  for three different tube sizes
and tube spacings with respect to change  in
the design bed temperatures. The results are
plotted  in Figure 7. Curve 2 represents the
.estimated cost for a staggered arrangement of
1-in. OD tubes, where H = 4 inches and V = 2
inches (see Figure 19 for definition) in all beds.
Curves 3a and 3b are for staggered arrange-
ment of 2-in.  OD tubes where H = 8 inches
and V = 4 inches in all beds. Some interesting
trends are present when these results are com-
pared to the steam generator cost for the basic
design. Ignore curves Ib and 3b for the time
being because the assumption of unrestricted
maximum allowable bed^ depth  is not con-
sidered reasonable. Then:


  1. Decreasing tube size and  tube  spacing
    increases  the steam  generator  cost   at
    design bed temperatures above  ~1520F
    (curve   2).  Further  decrease  in  bed
    temperature  necessitates  splitting  the
    reheater bed  in the  basic desig^ into two
    beds  and  substantially increasing the
    steam generator cost of the basic design.
  2. Increasing tube size and tube spacing also
    increases the steam generator  cost (curve
    3a).
The reasons  for these results are as follows:

 1. When tube  size  and  tube spacing are
    decreased, more heat transfer surface can
    be immersed in a unit  bed volume which
    results in lower bed height and module
    height; thinner wall tubes can be used
    which results in lower tubing cost. These
    are positive advantages.

 2. However,  smaller  tube  size  and  tube
    spacing   increase  the  amount of  tube
    bending and tube welding required. This
    is because more tubes of smaller diameter
    are  required   to  carry  the   same
    water/steam load at a constant flow rate in
    the tube.  Fabrication cost as a function of
    tube  wall thickness is  not  taken into
    account because of not enough informa-
    tion available. Pumping costs, a part of
    total operating cost, are not included here.

The balance  of these  two factors determines
the total steam generator cost.
   Since the tubing cost constitutes only about
20 percent of the cost of the pressure parts, the
increase  in  fabrication  cost  is  a  more
important cost. This can best be illustrated in
Figure  20 where  the component  costs are
plotted  against the  available  heat transfer
surface  per unit volume which relates to the
tube pitch/diameter ratio. Shell cost increases
almost linearly with decreasing  heat transfer
surface  per unit  bed volume.  The  cost  of
tubing,  headers,  downcomers, and  risers  is
almost constant at heat transfer surface larger
than 6 ft2/ft3 bed volume; it increases rapidly
at heat transfer surface lower than  ~ 6 ftVft3
bed volume, which corresponds to the use of a
tube  diameter of  1-1/2-in.  OD  or larger.
When larger tubes are used, the minimum
wall thickness  increases rapidly and  so  does
the tubing cost which contributes most of the
cost escalation  at lower heat transfer surface
                                                                                  IV-2-11

-------
 per unit bed volume. However, the cost of tube
 bending, tube welding, and water walls fabri-
 cation increases steadily with increase in heat
 transfer  surface   per   unit   bed  volume.
 However,  the  cost of tube  bending,  tube
 welding, and water walls  fabrication increases
 steadily with increase in heat transfer surface
 per unit bed volume. The balance of all these
 factors creates  a minimum  in total  steam
 generator cost at about 6.5 ft2 heat transfer
 surface per ft3 of bed volume. This can be
 achieved by arranging 1-in. OD tubes  where
 H = 4 inches and V = 3 inches. Fortunately,
 the  tube  size and  tube spacing used in the
 basic  design is  very  close  to this actual
 minimum.  Clearly,  there  are  different
 minimums  at   different   design  bed
 temperatures.  An  optimum  design for  a
 specific  operating  condition  requires   a
 separate evaluation. However, this optimum
 design  point is not as  critical as  may  be
 generally conceived. For example, at a design
 bed temperature of 1750F  (Figure 20) the.
 difference in the steam generator cost between
 the  optimum design and other designs  is
 within  $1.00/kW for available heat transfer
 surface of 3 to 11 ftVft3 bed volume, which
 covers the three tube sizes  and tube spacings
 in our current  evaluation.

   The maximum  allowable  bed depth is a
 more important variable. The steam generator
 costs are  also plotted against the maximum
 allowable bed depth at constant design bed
 temperature  and constant bed area in Figures
 21 and 22. Although the steam generator of 1-
 in.  OD tubes (where H = 4 inches and V = 2
 inches) is  not as economical compared to the
 basic design, it becomes  progressively  more
 attractive  at a lower maximum allowable bed
 depth. At maximum bed depths lower than 14
 feet (Figure 21), a steam generator using 1-in.
 OD tubes is  actually cheaper than the  basic
 design by up to 
-------
Heat Transfer Coefficient

   In the  basic design,  the overall  heat
transfer coefficients assumed are  47 Btu/ft2-
F-hr for the pre-evaporator, 45 Btu/ft2-F-hr
for the superheater, and 43 Btu/ft2-F-hr for
the  reheater. The  bed-tube  heat transfer
coefficient is assumed to be 50 Btu/ft2-F-hr
for all beds. When the bed-tube heat transfer
coefficient is changed, the overall heat transfer
coefficient will be changed as well. This leads
to a change in the total heat transfer surface
requirement and the bed depth.

   Change in the heat transfer coefficient will
also change the tube metal temperature.  A
change of tube material may be necessary for
some cases.  The design metal temperature is
assumed to be the maximum outside tube wall
temperature based  on  the  minimum
permissable wall thickness.

   Taking  into  consideration the aforemen-
tioned factors, the total steam generator costs
at bed-tube heat transfer coefficients of 35 and
75 Btu/ft2-hr-F were projected for different
tube pitch/diameter ratios at different design
bed temperatures.  At a  low bed-tube heat
transfer coefficient (35 Btu/ft2-hr-F) where a
large amount of in-bed heat transfer surface is
required, a  boiler with smaller tube size and
tube spacing is  much more economical than
the one with larger tube size and tube spacing.
Saving up to 20 percent of the  total steam
generator cost  is  feasible if bed area  is
constant and maximum allowable bed depth is
20 feet. If the maximum allowable bed depth
is less than 20 feet, the saving will be even
larger. At a bed-tube heat transfer coefficient
of 75 Btu/ft2-hr-F where the heat transfer
surface requirement is substantially reduced,
smaller boiler tubes and spacings  do not have
a clear  advantage.  If  the  bed-tube  heat
transfer coefficient is 35 Btu/ft2-hr-F rather
than 50 Btu/ft2-hr- F as assumed in the basic
design, the steam generator cost will increase
by $2.7/kW at 1750F design bed temperature
and by $5.4/kW at 1600F. If 1-in. OD tubes
are  used where H  =  4 inches and  V = 2
inches, the cost escalation would be $2.0/kW
at 1750F and $3.4/kW at 1600F. If the heat
transfer coefficient is increased to 75 Btu/ft2-
hr-F, the reduction in steam generator cost
from that of the basic design is only marginal,
~ $1.0/kW at 1750F.


   Figures 25 and 26 show the effect of the
bed-tube heat transfer coefficient on the steam
generator cost at constant bed temperature.
An increase of the bed-tube heat transfer co-
efficient from 50 to 75 Btu/ft2-hr-F (a  50
percent   increase)   decreases   the  steam
generator cost by ~ 10 percent ( ~ $1.40/kW).
A decrease  of bed-tube  heat transfer  co-
efficient from 50 to 35 Btu/ft2-hr-F (a  30
percent decrease) increases the cost by  ~20
percent (~ $2.80/kW) (Figure 25). The'curves
start  to  level off at higher bed-tube  heat
transfer coefficients. Thus, further increase in
bed-tube heat transfer  coefficient larger than
about 75 Btu/ft2-hr-F does not affect the cost
substantially. However, further  decrease  in
bed-tube heat transfer coefficient lower than
50   Btu/ft2-hr-F   increases   the  steam
generator cost  rapidly, especially for large
tube sizes and tube spacings and  at lower
design bed temperatures (Figures 25 and 26).
A 40 percent increase in cost occurs when the
bed-tube  heat transfer coefficient decreases
from  50 to  35 Btu/ft2-hr-F  at design bed
temperatures of 1636F for 2-in. OD tubes
where H = 8 inches and V = 4 inches (curve 3,
Figure 26). It is recommended to design the
steam generator at a  bed-tube heat transfer
coefficient about 75 Btu/ft2-hr-F if it is at all
possible, and to  avoid designing the steam
generator at a bed-tube heat transfer  co-
efficient   lower  than   50   Btu/ft2-hr-F,
especially if large tubes and spacings are used
and if lower bed temperatures are employed.


   To  complete  the  evaluation,  the  cost
information was also prepared for other cases
at different  design and operating variables to
show  the   interacting  effect  of  tube
pitch/diameter ratio,  bed-tube heat transfer
coefficient,  maximum  allowable bed  depth,
and the design bed temperature.

                                   IV-2-13

-------
    Again, the maximum allowable bed depth
 turns out  to  be the limitation  of the steam
 generator design and cost, especially at  low
 bed-tube heat transfer coefficient where large
 amount of heat transfer surface  is required in
 the bed. In this case, a smaller  tube size  and
 tube spacing and a higher design bed temper-
 ature are preferred. At a maximum allowable
 bed depth of 10 feet, the boiler designed at 35
 Btu/ft2-hr-F costs $3.5/kW  more than that
 designed at 50 Btu/ft2-hr-F and  S8.6/kW
 more than that designed at 75 Btu/ft2-hr-F at
 1750F design bed temperature. At design bed
 temperature  of  1407F,   the  figures   are
 $6.7/kW and  $12.5/kW, respectively. With 1-
 in. OD where H = 4 inches and V =  2 inches,
 the figures are $3.3/kW and  $4.4/kW  at
 1750F; and $2.8/kW and $8.1/kW at 1407F,
 respectively.
    Cost savings become smaller when the bed-
 tube  heat  transfer coefficient  is  further
 increased over 75 Btu/ft2-hr-F. If the bed-
 tube heat transfer coefficient  is decreased to
 lower  than  50 Btu/ft2-hr-F, the steam
 generator cost  increases rapidly.  Thus it is
 recommended that the bed-tube heat transfer
 coefficient be kept  at higher than 50 Btu/ft2-
 hr-F and preferable around 75 Btu/ft2-hr-F
 at  the current design conditions.
 Tube Materials
    The tube materials used in the basic design
 are conventional boiler tube material with SA-
 210-Al for tubes in pre-evaporator; SA-213-
 T2 for tubes in lower superheater; SA-213-T22
 for tubes in water walls, upper superheater
 (lower loops),  and  reheater;  and  SA-213-
 TP304H for tubes in upper superheater (upper
 loops). Changes in design  bed  temperature,
 bed-tube heat transfer coefficient  or steam
 temperature may require  higher grade tube
 materials; however,  these changes do not sub-
 stantially affect  the steam generator cost
because the tubing cost alone constitiutes only
~ 10 percent of the total steam generator cost
( ~ $1.40/kW).  Higher fabrication  cost  for
higher alloy material may  increase this cost
slightly. Nevertheless, the  total boiler cost is
not expected to  increase significantly due to
change of tube material unless the operating
bed temperature and heat transfer coefficient
are drastically changed.

   Module Capacity

   Boiler modules can  be  shop fabricated,
partially  shop fabricated  or field  erected
depending on the size. Modules up to 12-ft
diameter can be shop fabricated. Modules up
to  17-ft  diameter  can  be  partially  shop
fabricatedthe  boiler  internals  being shop
fabricated,  the  pressure shell  being  field
erected. The plant concept for a given capacity
can be either multiples of shop fabricated
modules, partially shop fabricated modules, or
field erected modules.

   The boiler plant equipment  cost is different
for each case. The steam generator cost will
depend on  operating conditions  and design
variables  such as  design bed temperature,
tube   size  and  tube   spacing,  maximum
allowable bed depth, etc. Auxiliary equipment
will also be affected: coal feeding, limestone or
dolomite  feed  and withdrawal, particulate
removal,  steam  piping,  boiler  feed  water
system, etc.

   The evaluation of  these  approaches is
based  on Figure  3, which presents the cost
variation  of  the  pressure   parts, . shell,
subcontrated  and  contracted   equipment,
drafting and home office, and erection with
respect to the plant size. Figure 27 presents
the resulting  costs  (including erection) for
design bed temperature at 1750F;  Figure 28
presents those for  design bed  temperature at
1636F at a  maximum allowable bed depth of
20 feet. The results show that at  a  plant size
larger  than   ~ 340  MW,  partially  shop
fabricated 4-module  plants with maximum
shop  fabrication  of the pressure  parts are
more economic than collective  multiples of
largest shop-fabricated  modules of the same
plant capacity. Changing the tube size  and
tube spacing does not affect this conclusion
(Figures  27  and 28). Change in heat transfer
coefficient should  not  produce  a different
conclusion because the cost escalation' due to
IV-2-14

-------
addition of a single module is more expansive
than  simple   enlargement  of the  existing
modules. However, this is no longer true when
the shell size is larger than  ~ 17 feet, because
the degree of shop fabrication of the pressure
parts decreases and the steam generator cost
again increases at a much higher rate.

   When the maximum allowable bed depth is
decreased from 20 feet, the cost saving of the
4-module partially shop-fabricated plant is
expected to increase. The splitting of the beds
in the shop   fabricated  module means  an
increase  in module height; however, in  the
partially  shop-fabricated module the  splitting
of beds can be avoided by simply enlarging the
module  diameter. From  the discussion  on
Figure 14,  the latter means a more economic
alternative except when the module diameter
is increased beyond  ~ 17 feet.

   The optimum design  variables (fluidizing
velocity, excess air, and pressure) in relation to
module  capacity  are  discussed  in their
respective  sections.

Participate Removal System  Economics

   The primary variables taken into consider-
ation in analyzing particulate removal system
 economics  are  dust  loading,  particle  size
 distribution entering and leaving the system,
 and  gas flow  rate. The effects  of boiler
 operating variables  bed temperature, free-
 board height,  and  superficial velocity  and
 design variables  tube  pitch/diameter ratio
 and bed-tube heat transfer coefficient  were
 also evaluated.

 Range of Dust Loadings and Particle Size Dis-
 tributions Considered

   The cases evaluated are outlined in Table
 5.
 Gas Turbine  Specification

   A  review  of  operating  experience  and
 assessment of erosion in gas turbines  was
 prepared by Westinghouse under contract to
 the Office of Air Programs.1 Specifications for
 the fluidized-bed combustion system based on
 that study are:

   Dust loading less than 0.15 gr/scf.

   Concentration of particles greater than 2
        less than 0.01 gr/scf.
   These design requirements will be updated
 as additional laboratory test data and opera-
, ting experience become available.
                   TableB. CASES EVALUATED FOR DETERMINING EFFECT
                   OF DUST LOADING AND PARTICLE SIZE DISTRIBUTION
Cases
evaluated
Group
Casel
Case 2
CaseS
Case 4
CaseS
Group 2
Casel
Case 2
CaseS
Dust loading
leaving FBC
Basic design
(6.7 gr/scf)
Double design value
Triple design value
Triple design value
Triple design value
Basic design
(6.7 gr/scf)
Double design value
Triple design value
Particle size
distribution
(Refer to Figure 29)
Curve 1 for particles
elutriated from FBC
Curve 2 for particles
elutriated fromCBC
Curve 2 for particles
elutriated from FBC
Curve 3 for particles
elutriated from CBC
Cyclone system
design
Figure 30
Figure 30
Figure 30
Figure 31
Figure 32
Figure 30
Figure 30
Figure 30
                                                                                 IV-2-15

-------
 Effect of Dust Loading and Particle Size Dis-
 tribution on Particulate Removal Equipment
 Cost

    The effect of dust loading and particle size
 distribution leaving  the boiler on the par-
 ticulates  going to the  gas  turbine for the
 particle removal systems selected are shown in
 Table  6.  The  gas turbine  specification  is

 Table  6.  DUST  LOADINGS  LEAVING THE
 SECONDARY  CYCLONE   FOR  DIFFERENT
             CASES EVALUATED
Cases
Evaluated
Group 1
Case 1
Case 2
CaseS
Case 4a
Case 5
Group 2
Casel
Case 2
Case 3
Total dust
loading leaving
secondary cyclone,
gr/scf
0.14
0.27
0.40
0.16
0.40.
0.27
0.46
0.67
Dust loading
for particles
> 2 /im, gr/scf
0.007
0.014
0.020
0.0006
0.019
0.013
0.022
0.032
Assuming same fractional collection efficiency for
the second secondary cyclone as that for the first
one. This assumption is too optimistic.

exceeded in several cases. In order to meet the
specification,  alternative particulate removal
systems could be considered  granular bed
filters,  electrostatic  precipitators,  ceramic
filters, etc.; additional mechanical collectors
could be used in series; or boiler operation
could  be  altered.  The  additional  use  of
mechanical  collectors is the approach used to
evaluate additional costs. This approach was
selected since the equipment is available and
thus provides the best cost data, and because
the effect of boiler  operating  conditions on
particulate  emission  is  difficult to  project.
Table 7 presents a summary of the economic
implications.

   If the gas turbine dust loading requirement
is defined as  < 0.01 gr/scf without reference
 to  particle  size,  no  centrifugal  separator
 presently available can meet this requirement
 within reasonable cost in all cases discussed
 above. If this were the case, high temperature
 ceramic filters  may have to  be  used.  The
 possibility  of  this   application   is  being
 evaluated. In this respect, the particle size dis-
 tribution is a far more important parameter
 than  the dust loading, since the  collection
 efficiency for the particle size smaller than 2
 l*m decreases rapidly. Thus, the particle size
 distribution  curves assumed for the present
 evaluation (as shown in Figure 29) are con-
 servative because a large amount  of fines  is
 assumed to be present.

 Effect of Total Gas Flow Rate on Particulate
 Removal Equipment Cost

   The effect of increasing gas flow rate was
 evaluated for four cases which correspond to
 the basic design flow rate, 30, 50, and  100
 percent  air.  The  selection  of  first  stage
 cyclones is based on the criterion of maximum
 efficiency at minimum cost with a minimum
 cyclone efficiency of  85  percent.  The cost
 increment for higher gas flow rate is shown in
 Figure 33.

   The  first  stage cyclone cost includes  not
 only the  cost of  the  first  stage  separators
 supplied by Ducon but also the cost of the
 separator pressure  vessel  and  all of the  gas
 piping from the steam generator outlet to the
 secondary  separator inlets.   The  costs  for
 pressure vessel and gas piping were estimated
 for two cases. In one case, the gas piping from
 the steam generator to the first stage separator
 is lined with hard refractory,  but the pressure
vessel and the gas piping from it are lined with
stainless steel. In the other case, hard refrac-
tory without an alloy liner was used through-
out. At 100  percent excess air, the increase is
$1.70/kW and $2.40/kW for hard refractory
liner and stainless steel liner respectively. The
major cost increase is  from enlargement of
pressure vessels and gas piping due to higher
gas flow rate. The increase in  separator cost
alone constitutes only about 18 percent of the
total cost increment.
IV-2-16

-------
                Table 7.  PARTICULATE  REMOVAL  SYSTEM COST
Case
Group 1
Casel
Case 2
Case 3
Case 4
CaseS
Group 2
Casel
Case 2
CaseS
Need for further clean-up to
achieve gas turbine specification

No
May not be required3
Yes, a second secondary cyclone in series
or granular bed filter
No
Yes, a second secondary cyclone in series
or granular bed filter
May not be required3
Yes, a second secondary cyclone in series
or granular bed filter
Yes, a second secondary cyclone in series
or granular bed filter
'articulate removal system
cost increase, $/kW

-

6.30
8.50
-
6.30
8.50
-
6.30
8.50
6.30
8.50
            3 Dust  loading is  larger  than  specification  but particle size  is  close to
             specification.
   The cost increase for the  second stage  is
presented in Figure 34. The incremental cost
for the second stage is more than 2 times that
of the first  stage ($5.00/kW versus $2.40/kW
at 100 percent excess air).  This is because
model  18,000  is  the largest  cyclone now
supplied  by Aerodyne. The capacity of model
18,000 with dirty gas  as the secondary gas  is
30,000 ftVmin. Any gas flow rate higher than
that will require  multiple units with their
individual  pressure vessel.  This  tends  to
increase  the  incremental cost of the second
stage; however, the rate of cost increase slows
down at excess air larger than 100 percent.. At
excess air larger than  100 percent, the rate of
cost increase for the first stage speeds up. This
is because increases in pressure vessel size and
gas piping  diameter increase the incremental
cost rapidly at very high gas flow rate.

   Taking into account the cost of gas piping
from the second stage cyclones to gas turbine,
the total incremental cost at different gas flow
rate is shown in  Figure 35.
   Combining cost figures from Figures  15,
16, and 35, an increase in total boiler cost of
$9.6 to S10.2/kW is required to operate  the
boiler at 100 percent excess air. This, however,
does not  take into consideration that  the
combustion efficiency in the primary beds will
approach 100 percent at 100 percent excess air
and that  the carbon  burn-up  cell can  be
eliminated.   In   addition,  the   particulate
removal system would  be much simpler and
the high  excess  air may also  provide  the
necessary flexibility to achieve plant turndown
if the  boiler  is  designed  at  lower  bed
temperatures. An additional   ~ 10 percent
turndown  capability is  obtained by operating
at  100 percent  excess air. The  lower bed
depths would also permit combining the two
superheater  beds into  one bed which  would
also reduce module height and cost. Thus it is
concluded that if the carbon burn-up cell can
be eliminated, operating at 100 percent excess
air may result in a lower energy cost with
increased cycle efficiency and flexibility. This
conclusion is true at 300-MW nominal plant
                                                                                  IV-2-17

-------
 size, but is  not  necessarily true at 600-MW
 plant size. This  is because at 600-MW plant
 size, the module has a 17-ft diameter at the
 basic design. Further increase  in  excess air
 requires either an increase in module diameter
 or in the total number of modules. At  100
 percent excess air,  four  modules with 23-ft
 diameter are required. Since fabrication cost
 of the internals  increases rapidly at module
 diameters larger than  17 feet because  of
 rapidly decreasing shop-fabricable portion, an
 increase of  $5.0/kW  in  boiler  module cost
 alone is conceivable. Adding the cost increase
 in boiler cost is about $15.0/kW. In this case,
 operating  at  higher  pressures   may  be
 beneficial.

    Instead of increasing the module diameter
 from  17 to 23 feet, the number of modules
 could  be  increased  keeping  the  module
 diameter constant. At 100 percent excess air,
 seven modules are required. In this case,  in
 addition to  module cost individual  particle
 removal equipment has to be  provided  for
 each module; ducting and piping manifolds
 have to be  increased;  coal feeding  systems
 become more complicated;  and  above all,
 instrumentation and control have to be more
 sophisticated. The total increase in boiler cost
 is estimated  to be $20.0 to $25.0/kW in this
 case. Consequently, even if the carbon burn-
 up cell could be eliminated, the increase in
 cost and complication in control do not clearly
 favor the operation at high excess air for large
 plant  size.  A  careful evaluation of overall
 design and control philosophy should be done
 if operation  at  high excess  air  is to  be
 attemped for large plant  size.

 POWER GENERATION EQUIPMENT

   Alternative  boiler design and  operating
 conditions will have three primary affects on
 the power generation equipment:

 1. Capacity   of   gas   and  steam   turbine
    equipment.

 2. Gas turbine inlet  temperature.

 3. Ability   to   achieve  higher  steam
    temperatures  and pressures.

IV-2-18
 Since this analysis assumes a constant fuel
 rate,  the capacity changes are  considered
 small except for the high excess air case. The
 effect  of capacity  was considered with the
 excess air analysis.

 Gas Turbine Inlet  Temperatures

   The design value for the gas turbine inlet
 temperature was 1600F for the base design of
 the pressurized fluid-bed boiler, which is well
 below  the current  state-of-the-art tempera-
 tures of  1800-1900F  for utility intermediate
 load  applications.  The 1600F  level  was
 established by assuming the flue gas leaves the
 bed at the 1750 F bed temperature, and that
 the temperature difference  between the  bed
 and  the  gas turbine  inlet  would  be 150F.
 Several factors may alter the gas turbine inlet
 temperature. These include:

 1. Temperature  drop  between  boiler  and
    turbine  expander150F was assumed
    which is probably excessive. A drop as low
    as  50 to 75 F may be  achieved.

 2. Bed  temperatureif the boiler  design
    temperature is changed, the turbine inlet
    temperature  will change. Sulfur removal
    considerations and ash agglomeration will
    determine the maximum temperature.

 3. Combustion above the  bedcombustion
    has been observed above the bed  of  a
    fluidized-bed  boiler which  increases the
    gas temperature 200-300F. Any combus-
    tion above the bed would increase the gas
    turbine inlet temperature. No combustion
    was assumed  in the base design.

 4. Modification of the cycle to provide for
    reheat of the product gas from the boiler
    prior to  the gas turbine. One  concept for
    doing this is  shown in Figure  36. Carbon
    carried out of the primary beds would be
    gasified to produce a low-Btu gas. The gas
   would  be used  in  the  second  stage
    combustor.

   Performance  calculations were  made to
determine the effect of a change in the gas
turbine   inlet   temperature   on   plant

-------
performance. The results are summarized in
Table 8. These results  are for 17.5 percent
excess  air  and  a boiler  efficiency of 88.6
percent. A  200 F change in the turbine inlet
temperature will change the  plant  heat rate
~1 percent using current technology.

Table 8.  PERFORMANCE  OF  PRESSURIZED
FLUID-BED  BOILER  POWER PLANT AS A
FUNCTION    OF   GAS   TURBINE   INLET
              TEMPERATURE

Gas turbine
inlet temperature.
F
1400
1500
1600
1700
1800
1900
2000

Plant
output,3
MW
625.6
634.9
644.1
644.8
645.3
641.8
636.2

Plant
heat rate,3
Btu/kWhr
9293
9157
9026
8921
8820
8773
8753
Fuel burned
after primary
beds,b
%


-
2.5
5.0
7.6
10.1
 3The decrease in performance at high gas turbine
 inlet temperatures is the result of increased bleed
 air required for turbine cooling and the increase in
 gas turbine waste heat which reduces steam cycle
 extraction for regenerative heating. The heat rate
 at 2000F would be reduced ~ 270 Btu/kWhr if
 turbine blade cooling was not required.

 b Assumes any increase above the design value of
 1600F .has to  result  from  burning  fuel either
 above  the bed or separately.
 Steam Temperature

    Plant  performance can  be increased by
 increasing steam temperature  and pressure.
 The effect of higher steam temperatures on
 the performance of the plant is shown in Table
 9.
Table   9.   EFFECT   OF   HIGHER   STEAM
TEMPERATURES ON PLANT PERFORMANCE
Steam
temperature,
F
1000/1000
1100/1100
1200/1200
Steam
pressure,
psi
2400
3300
4500
Gas turbine
inlet temperature,
F
1600
1600
1600
Power.
MW
644.1
673.9
690.8
Heat rate,
Btu/kWhr
9026
8627
8417
An increase of 100F in both superheat  and
reheat temperatures will give a reduction of
about 400 Btu/kWhr in plant heat rate.  The
increased performance and  inherently  less
severe  boiler tube corrosion in fluidized-bed
boilers make high steam temperatures attrac-
tive.

ASSESSMENT

   A summary of the sensitivity analysis is
presented in Table  10.  Each parameter is
indicated with  a  projection of what change
might  be required  as the result  of experi-
mental   data.   For   example,   the    bed
temperature was set at 1750F for 100 percent
load. This temperature may prove to be too
high for economic sulfur removal and have to
be reduced to 1600 or 1650F which may be
more favorable. The summary table indicates
how such a change would affect plant cost  and
performance assuming  no  other variable
restrictions.  In this case, the plant cost would
increase  < $3/kW, plant performance would
decrease  < 0.5 percent, and  plant turndown
to 25 percent load could still be achieved. This
would  result in an energy cost reduction of
< 0.3 mills/kWhr. Since the projected advan-
tage over a conventional plant with stack gas
scrubbing is greater than 1.5 mills/kWhr,1  the
penalty for  lowering the temperature is  not
significant. This conclusion holds for all of the
variables  considered.   The   conclusion  is
generally valid for plant capacities of 200 to
700 MW. The results of the sensitivity analysis
indicate that the base plant design is relatively
insensitive to changes in operating  conditions
or design parameters.

   Operating   conditions   and  design
parameter  changes can  occur which would
result  in a  significant cost increase for the
system. This would most likely occur as the
result  of additive  problems.  For  example,
suppose the  bed  temperature  had  to  be
decreased  to  1600F,  the  heat  transfer
coefficient was only 35 Btu/hr-ft2-F, and the
dust loading from the boiler was 3 times the
design value. The plant cost could increase by
7 to 8 percent and the efficiency decrease by
                                                                                  IV-2-19

-------
Table 10. SENSITIVITY ANALYSIS SUMMARY
Parameter
Boiler operating conditions
Bed temperature




Fluidizing velocity

Excess air



Pressure



Paniculate carry-over
Loading



Panicle size



Boiler design
Heat transfer surface
Configuration





Change

Reduction from
1750Fto16rjOF



Decrease to 5 ft/sec
Increase to 15ft/sec
Increase to 100%



Increase to 15 atm




Increase loading
to 3 times the
design value
(
-------
                                   Table 10 (continued). SENSUIVrTY ANALYSIS SUMMARY
        Parameter
     Heat transfer
     coefficient
     Materials


    Bed depth i
  Power generation equipmentk
    Gas turbine inlet
    temperature
    Steam temperature
     Change
Decrease from
5.5ft2/ft3to
3 ft2/ft3
Increase from
SO to 75
Btu/hr-ft2.F
Decrease from
50 to 35
Btu/hr-ft2-F
Assume tubing
cost 50% greater
than base design
Reduced to 10 ft
200Ffrom
1600F
100F increase
in superheat
and reheat
       Boiler
                                                    oj$1/kW increase9
$l/kW decrease
                                                    ""vSS/kW increase
<$2/kW increase
~$3/kW increase
<$2/kW increase
PJantcost

Auxiliaries
Effect

      Power generation
         equipment
                     Negligible
Performance
                              1% in efficiency

                              2% increase in
                            efficiency
aIncrease will depend on bed depth restrictions; $3/kW would correspond to a maximum allowable height of ~15 ft.
"The fluidizing velocity will affect the particulate removal equipment  see participate emission parameter for costs.
Considerable savings may be realized for large capacity (>600 MW) plants since higher velocities avoid the need for field erection.
"Additional savings would be realized for large capacity  (>600 MW) plants in order to avoid field erection. The $1/kW does not include a cost reduction
 which may result from elimination of the carbon burn up bed due to increased efficiency.
e$4/kW assumes the larger capacity gas turbine has the same unit cost as the base machine. Actual cost $1/kW of a larger machine would be lower.
'Based on projected gas turbine requirement of <0.01 gr/scf of particles <2 (im.            ,
9Assumes maximum allowable bed depth of 20 ft.
"Includes effect on fabrication.
'Assumes constant freeboard. (Change in freeboard requirement would have similar effect.)
'If the temperature is the result of equipment modification, the capital cost would be altered. Gas piping will be affected in any case.
kThe plant cost projections are based on the base case plant capacity and performance. An increase in efficiency will reduce the specific plant cost. Water
 cooling has not been included. A higher efficiency will result in lower cooling equipment cost.

-------
 ~ 0.5  percent.  This  would  result  in  an
 increase  in  the  energy  cost  of   ^ 0.7
 mills/kWhr. This is a significant increase but
 still  within  the  economic  margin. Caution
 must be exercised  in  interpreting multiple
 changes in the variables. In the case above, the
 cost effects were added. However, decreasing
 the bed temperature and increasing the heat
 transfer coefficient both increase the  heat
 transfer surface and the bed depth if the bed
 area  is maintained constant. Any restrictions
 on bed depth would also have to be considered
 in the evaluation. Parametric curves have been
 prepared to  enable this type of evaluation to
 be made.
   The effect of boiler plant pressure drop and
 the steam turbine condenser pressure on plant
 performance has been presented.2 The boiler
 plant pressure drop has a small effect on plant
 capacity and heat rate: 1 percent increase in
 pressure drop results in a 0.1 percent increase
 in plant heat rate.  An  increase in  the steam
 turbine condenser pressure from 1-1/2-in. Hg
 to 3-in. Hg for a cooling tower results in a 2.5
 percent increase  in heat rate.

   The potential performance of the plant was
 evaluated by assessing the effect of higher gas
 turbine inlet temperatures and higher steam
 temperatures and   pressures.   The  results
 indicate that plant efficiencies of ~ 45 percent
 can be achieved with gas turbine inlet temper-
 atures  greater   than  2000F,  with   high
 temperature   blade   material  to  minimize
 cooling requirements,  and   with  steam
 temperatures of 1200F.

   Advances  in boiler plant subsystem  con-
 cepts  have  not  been  considered, in  this
 analysis. Cost reductions may be achieved by
 using  alternative concepts. The  following
components have been  studied for potential
savings:

   Particulate  RemovalThe  projected
   system  utilizes four secondary  collectors
   for final  particle removal before the gas
   turbine.  Alternative systems  are  being
   considered which may reduce the number
   and  size  of  the units.  Cost  estimates
    indicate that reductions of $1 to $5/kW
    for the total particulate  removal system
    may be possible.

    High-Temperature Gas PipingThe high-
    temperature  gas piping  cost  would  be
    reduced if the particulate removal system
    were simplified  by  using  fewer units per
    module.   Additional  savings might  be
    realized if refractory lined pipe could be
    used between the secondary collectors and
    the gas turbine. The present design uses a
    high-alloy steel to assure protection of the
    gas turbine from additional  particulates.
    Coal Feeding  SystemThe coal feeding
    system design  is 'based  on systems which
    have been built and operated. The design
    provides a separate coal feeding system for
    each fluidized  bed in  order to  assure
    control of the coal feed rate to each bed. It
    may be possible, however, to reduce the
    number of coal feed systems from 16 to 4 if
    independent control of solids flow to each
    bed in  a module can be achieved from a
    single pressurized injector. The potential
    cost  reduction  is  estimated to be  >
    $2/kW.

    Stack Gas Cooler DesignCost estimates
    were obtained  for the stack  gas  coolers,
    but no  attempt was made  to optimize the
    design  or consider nonconventional
    designs,  such  as those using fluidized
    beds. Preliminary conceptual evaluation
    indicates that the cost might be reduced
    $4.40 to $3.40/kW.
CONCLUSIONS

   1. Base plant design is near optimal.

   2. Pressurized fluidized-bed boiler  power
plant maintains greater than 5 to 10 percent
energy  .--ost advantage over conventional plant
with stack gas  scrubbing.

   Effect  of  a  potential  change  in  bed
   temperature, fluidizing velocity, heat
   transfer  surface  configuration, gas  side
   heat  transfer  coefficient,   boiler  tube
IV-2-22

-------
  materials,  bed  depth  limitations or
  pressure will  result  in:  no  significant
  change   in   plant   operability   or
  performance, < 2 percent increase in plant
  cost, and  < 0.2  mills/kWhr increase in
  energy cost. Effect of increasing the  dust
  loading to three times the design value will
  increase the plant cost ~ 4 percent, energy
  cost < 0.3  mills/kWhr.

   Effect of increasing excess air will result
  in: increased turndown capability and per-
  formance, 3 to 6 percent increase in plant
  cost, and no significant change in energy
  cost.

  3. Plant efficiencies of ~ 45 percent may be
achieved for gas turbine  inlet temperatures
above 2000F and  steam  temperatures of
1200F.

ACKNOWLEDGMENT

  This work was performed under contract to
the  Environmental Protection Agency, Office
of Research and Monitoring. P. P. Turner
served as  contract  officer.  N. E.  Weeks,
Westinghouse  Power   Generation   Systems
Division,   performed  power  system  cycle
analyses.
REFERENCES

1. Evaluation  of  the  Fluidized   Bed
   Combustion  Process. Volumes  I-III.
   Westinghouse  Research  Laboratories,
   Pittsburgh, Pa. Submitted to the Office of
   Air Programs, Environmental  Protection
   Agency, Research Triangle Park, N.  C.
   November 1971.

2. Keairns, D. L.,  J. R. Hamm,  and D.  H.
   Archer.  (Presented  at Annual  AIChE
   Meeting.  San  Francisco. November  1971.
   AIChE Symposium Series, Volume "Air",
   1972.

3. Highley, J., D. Chandrasekeva, and  D.  F.
   Williams.  Fluidized  Combustion  Section,
   Coal  Research  Establishment,  National
   Coal  Board,  London,  England. Report
   Number 20, April 1969.

4. McLaren,  J. and  D. F. Williams.  J.  of
   Institute of Fuel. 303, August 1969.

5. Smith, S.  Private communication, Boiler
   Tube Division, The Babcock & Wilcox Co.,
   Barberton,  Ohio.

6. Genetti, J. W. and W. E. Bartel. (Presented
   at  American  Institute   of   Chemical
   Engineers   72nd  National Meeting.  St.
   Louis. May 22-24, 1972.)
                                                                               IV-2-23

-------
to
u
COMPRESSOR POWER
1 COMPRESS
"R-fL>
	 GAS FLOW 1
== SOLIDS FLOW j
	 STEAM CYCLE '
1
i SULFUR I
[ RICH GAS 1
SPENT ,
' DESULFURIZING

OR TURBINE TURBINE

i
PRIMARY
REMOVAL
1
tf
\J \l
^ SECONDARY
REMOVAL


DESULFURIZING AGENT FLUJ!?!PD CARBON
AGENT REGENERATED rnwiRiKTnR BURN"UP
REGENERATOR STONE  ngSjl^gfe, CELL

1 ft
1
1
IHttSTF ! MAKE-UP
SToJll STONE
1 COAL
1 (OIL)
i
1
1
1.
~TT
S\ 1 .XI ^

-------
         REHEATED STEAM
         REHEATER BED
         SUPERHEATED STEAM


         SUPERHEATER BED
         SUPERHEATER BED
        PRE-EVAPORATOR BED
        FEED WATER
PLANT
SIZE
320 MW
635 MW
VESSEL
DIAMETER
12 ft
17ft
                   GRADE ELEVATION
                                   ELEVATION
Figure 2.  Pressurized fluidized-bed steam generator for combined cycle plant (four required).
                                                                                 IV-2-25

-------
  2.0
  1.5
                              ^SHELL
81.0
        HOME OFFICE (ENGINEERING,
          CONTRACT RESERVE, ETC.
        ERECTION
  0.5
      300
                 CONTRACTED AND
                 SUBCONTRACTED
                   EQUIPMENT
                ' DRAFTING
                                               3.5
                                               3.0
                                             "&
                                             ii
                                             x
                                               ? 5
                                               > **
                                               2.0
                                               1.5
                                                                  ' PRESSURE PARTS
                    500
600
400
500
600
                      700     800      300
                           PLANT SIZE, MW
Figure 3.  Steam generator cost breakdown for the W-FW basic design.
700    800
IV-2-26

-------
    300
           PLANT POWER


           STEAM TURBINE POWER
     200
                                                        ASSUME BED DEPTHS ARE
                                                        CONSTANT AT BASIC DESIGN VALUES
I
o
     100
                                   GAS TURBINE POWER
        1300
1400
1500
                                       1600
1700
                                        BED TEMPERATURE, F
                   Figure 4.  Effect of bed temperature on power generation.
1800
     25
 Q_
 a
 LLJ
 oa
 o
 UJ
     20
      15
      10
        1300
1400
                                 1700
                       1500             1600
                    BED TEMPERATURE, F
Figure 5-  Bed depth requirement at different bed temperatures.
              1800
                                                                                   IV-2-27

-------
     150
     100
 LU
 31

 LU
 O
 O
      so
         1300            1400             1500             1600             1700

                                          BED TEMPERATURE, F
                     Figure 6. Effect of bed temperature on total module height.
1800
IV-2-28

-------
I
ce
UJ
9
C3

i
LU
     0
                       lb
                                                           BASIC DESIGN POINT
                     BED-TUBE HEAT TRANSFER COEFFICIENT: 50 Btu/ft2 -hr-F
              CURVE    MAX. ALLOWABLE BED DEPTH
               la



               lb
               2

               3a

               3b
            20ft



       UNRESTRICTED
            20ft

            20ft

        UNRESTRICTED
      TUBE ARRANGEMENT

I'/Hn. OD AT H=7 in. AND V=3 in. IN
PRE-EVAPORATOR AND SUPERHEATER
2-in. OD AT H=7 in. AND V=4 in. IN
REHEATER
              SAME
Hn. OD AT H=4 in. AND V=2 in. IN
ALL BEDS
2-in. OD AT H=8 in. AND V=4 in. IN
ALL BEDS
               SAME
       1300
1400
                   1700
                                 1500             1600
                                  BED TEMPERATURE, F
Figure 7.  Effect of bed temperature on the steam generator cost of 318-MW plant
(not including erection).
1800
                                                                                   IV-2-29

-------
       10
  O

  X
  GO
  O
  O
  (XL
  O
  DC
                                                                             1470 "F
                       BED-TUBE HEAT TRANSFER COEFFICIENT: 50 Btu/ft2-hr-  F
                       TUBE PITCH/DIAMETER RATIO: BASIC DESIGN
         0
10
20
30
                                  MAXIM ALLOWABLE BED DEPTH, ft
         Figure 8.  Dependence of the steam generator (318-MW) cost on the maximum allowable
         bed depth ( not including erection).
IV-2-30

-------
   100.5
UJ
V)
01
D-

O
o
Q.
   100.0
    99.5
    99.0
            TUBE PITCH/DIAMETER RATIO
            BASIC DESIGN
            J-in. OD AT H=4 in., V=2 in.
            2-in. OD AT H=8 in., V=4 in.
THIS IS AN ADDER APPLIED TO FIGURE 4
BED-TUBE HEAT TRANSFER COEFFICIENT.
50 Btu/ftz-hr-F
50 Btu/ft2-hr-F
50Btu/ft2-hr-F
         1300
           1400
                    1700
1800
                                       1500              1600
                                     BED TEMPERATURE, "F
Figure 9. Change in net plant power output due to change in pressure drop across the bed.
                                                                                        IV-2-31

-------
       20
        15
  O

  3
  UJ
  O
  LLl
  D.
        10
                       BED TEMPERATURE: 1750 F

                       TUBE PITCH/DIAMETER RATIO:  BASIC DESIGN

                       BED-TUBE HEAT TRANSFER COEFFICIENT: 50 Btu/ft2 - hr -f
                 SUPERFICIAL

                  VELOCITY
                     SUPERHEATER


                        PRE-EVAPORATOR
SUPERHEATER


    REHEATER


      PRE-EVAPORATOR
                                                                                   60
                                          UJ
                                          a

                                          o
                                          UJ
                                          00
                                                                                   40
                                                                                   20
                                         REHEATER
                                                    ?
                                                    I
                           10
                 30
                             20

                         BED AREA, ft2

Figure 10.  Dependence of bed depth on the fluidizing velocity.
40
IV-2-32

-------
    300
     250
     200
C9
LU
     150
     100
      50
            BED TEMPERATURE
                      1407F
                      1522 F
MAXIMUM ALLOWABLE BED DEPTH: UNLIMITED
TUBE PITCH/DIAMETER RATIO: BASIC DESIGN
BED-TUBE HEAT TRANSFER COEFFICIENT:
        50 Btu/ftZ-hr-F
                                                                                  16
                                         15
                                         14
                                         13
                                         12
                                         11
                          10
                      30
40
                                     20
                                  BED AREA, ft2
Figure 11.  Dependence of module diameter and module height on the bed area.
                                                                                  10
                                                                                    IV-2-33

-------
  ,
   iI
   x
   
       3.0
2.5
       2.0
       1.0
                              COST ESTIMATION BY FW
                     ...-._  INDEPENDENT ESTIMATION BY WESTINGHOUSE (WITHOUT       /
                               INTERNAL SUPPORT ALLOWANCE)                       '
                     	INDEPENDENT ESTIMATION BY WESTINGHOUSE (WITH      S
                               INTERNAL SUPPORT ALLOWANCE)                  /
                 !       4       6       8      10     12      14     16
                                       MODULE DIAMETER, ft
                  Figure  12.  Shell cost (four modules) at basic design height.
                                                                    18
20    22
IV-2-34

-------
                                                                                 '1750F
                        635 MW
O
CJ
                         COST INCREASES K160.000 FOR BASIC
                         DESIGN CASE ) CAUSED BY SHIFT FROM
                         SHOP-ASSEMBLED TO FIELD-ERECTED
                         VESSEL AT VELOCITIES <8 ft/sec
                                       318 MW
      1407 F


      1522 F

      1636 F

      1750 F
                  CURVE    TUBE PITCH/DIAMETER RATIO
                           BASIC DESIGN

                           1-in. 00 TUBES AT H= 4 in. AND V=2 in.
             BED-TUBE HEAT TRANSFER COEFFICIENT:  50 Btu/ft2.hr- F

             MAXIMUM ALLOWABLE BED DEPTH: UNLIMITED
                                                     10
15
                              FLUIDIZING VELOCITY IN PRE-EVAPORATOR, ft/sec
          Figure 13.  Dependence of the steam generator cost on the fluidizing velocity.
                                                                                    IV-2-35

-------
       2.5
               BED TEMPERATURE: 1750 F
                                             

               CURVE    TUBE PITCH /DIAMETER RATIO
                         BASIC DESIGN
       2.0
	lin.-OD TUBES AT H = 4 in.
          AND V= 2 in.
CURVE 1: COST REDUCTION DUE TO DECREASE
         IN MODULE DIAMETER.
CURVE 2: COST ESCALATION DUE TO
         INCREASE IN MODULE HEIGHT
   o
       1.5
       1.0
                           635 MW
                           318 MW
       0.5
                                5                      10                     15
                               FLUIDIZING VELOCITY IN PRE-EVAPORATOR, It/sec
                   Figure 14.  Dependence of the shell cost on the fluidizing velocity.
IV-2-36

-------
s
Q
o
    0.7


    0.6



     0.5



     0.4
     0.2
     0.1
      0
       	 CONSTANT FLUIDIZING
           VELOCITY

        CONSTANT BED
           AREA
              10
              20
80
90
100    110
                     30      40      50      60      70
                        EXCESS AIR IN PRIMARY BEDS, %
Figure 15. Cost reduction in heat  transfer surface and pressure shell at different air
flow rates.
o
                                             THESE WOULD NOT BE CONTINUOUS
                                             CURVES IN PRACTICE, BECAUSE TURBINES
                                             ARE ONLY MANUFACTURED IN SELECTED
                                             SIZE RANGES
                                                                                        5
                                                                                        a1
                                                                                a
                                                                                UJ
                                                                                on

                                                                                g
                                                                                o
            IB     20    10    40     so     0     n     so    so   100   no   120
                           OVERALL IXCESS AIR (INCLUDING AIR FROM CBC), %
                18,  0t <l@n in gas turbine equipment and cost reduction in steam
         turbifli equipminl at different air flow rates.
                                                                                     IV-2-3?

-------
     9500
                                                               BOILER PRESSURE   =2400psi
                                                               BOILER EFFICIENCY =88.
     8900
         10
25
30
                                    OVERALL PRESSURE RATIO
      9400
     9300
     9200
 2  9100
     8900
                             15                   20                   25
                                       OVERALL PRESSURE RATIO
                    Figure 17.  Plant heat rate versus compressor pressure ratio.
                    30
IV-2-38

-------
4 in.
-*
X
3'/2in.
=> p.
1
^r
$

f"^








^
r
 -
"*
^
,
T
O

/^







O
Q



s
2-in. OD WATER
^ WALL TUBES ~\
3 in.
^>- 2-in. OD TUBES



<
<

y/2 in.
c
'
P
6


o
^
C








V.

* 

I
)
s
O


O






o
Q
tf

o
^
2-in. OD WATER
^ WALL TUBES
>-l'/2-in.OD TUBES




           REHEATER                        PRE-EVAPORATOR AND SUPERHEATERS

                     Figure 18.  Tube arrangement in basic design.
 r)	
     O   ^
^()
                               O

                               O
                   ^
      O         O        O
            O         O
      O         O        O
            O         O
      STAGGERED OR ROTATED DIAMOND
               ARRANGEMENT
 	 
1 ' ^
0 -6
o o
o o
o

o
o
o
                                            SQUARE OR RECTANGULAR
                                                 ARRANGEMENT
                   Figure 19.  Definition for different tube arrangements.
                                                                               IV-2-39

-------
  o
                    BED TEMPERATURE: 1750 F, MAXIMUM ALLOWABLE BED DEPTH: 20ft,
                    BED TUBE HEAT TRANSFER COEFFICIENT: 50 Btu/ft2 . hr. -p
            TUBE BENDING, WELDING,
           ' WATER WALLS
            FABRICATION
                            DRAFTING,
                        HOME OFFICE, ETC.
                               SHELL
                          TUBING, HEADERS,
                              COMERS, RISERS
                                         <>
                                         tI
                                         X
                                         (s>
                                         H-"

                                         O
                                                                       I       I      I
  TOTAL STEAM
GENERATOR COST
                    6
                       12    14
                   8      10     12    14      2     4      6     8     10

                 AVAILABLE HEAT TRANSFER SURFACE/UNIT BED VOLUME, ft2/ft3
Figure 20.  Change of the steam generator cost with heat transfer surface per unit bed
volume (not including erection).
IV-2-40

-------
   10
x
,_   e
fr   Q
uj   4
UJ
o
                TUBE PITCH/DIAMETER RATIO

                1-1/2 in. OD AT H7 in., V-3 in.
                IN PRE-EVAPORATOR AND SUPER-
                HEATER                   _
                2-in. OD AT H=7 in., V=4 in.
                IN REHEATER
                1-in. OD AT H4 in., V*2 In.
                IN ALL BEDS
                2-in. OD AT H8 In., V4 in.
                IN ALL BEDS               
          BED TEMPERATURE:  1750 F

          BED-TUBE HEAT TRANSFER
          COEFFICIENT: 50 Btu/ft2-hr-F
                                                10
                                               o
                                               g
                                               5
                                               LU

                                                      BED TEMPERATURE: 1636 F

                                                      BED-TUBE HEAT TRANSFER
                                                      COEFFICIENT: 50 Btu/ft2-hr-F

                                                      SEE FIGURE 21 FOR CURVE DESIGNATION
                 10           20
           MAXIMUM ALLOWABLE BED DEPTH, ft
                                          30
        Figure 21.  Dependence of the steam
        generator (318-MW) cost on the maxi-
        mum allowable bed depth (not includ-
        ing erection),
       10           20
MAXIMUM ALLOWABLE BED DEPTH, ft.
30
                                                     Figure 22.  Dependence of the steam
                                                     generator (318-MW) cost on the maxi-
                                                     mum allowable bed depth  (not includ-
                                                     ing erection),
                                                                                   IV-2-41

-------
       50
       40
   
   E  30
   o
   o
   CO
   o
   UJ
   ca
       20
       10
               60mm
               TUBES
  A
  O
          35 mm
          TUBES
                          V
                          A
          NO. OF
          ROWS
          1
           2
           3
           3
           4
           PITCH/DIAMETER
               RATIO
                2
                4
                4
                6
                                       PITCH/DIAMETER RATIO
          Figure 23.  Change of bed-tube heat transfer coefficient with pitch/diameter ratio.3
       50
    ".  40
   UJ
   o
   o
       30
       20
   UJ
   LU
   m
   a
   UJ
   ca
       10
60mm
TUBES
  A
  O
35mm
TUBES
                          V
                          A
NO. OF
ROWS
 2
 2
 3
 3
PITCH/DIAMETER
     RATIO
       2
       4
       4
       6
        8
                      24             6             8             10
                                   NARROWEST GAP BETWEEN TUBES, in.
                 Figure 24.  Change of heat transfer coefficient with tube spacing.3-4
                                                                           12
IV-2-42

-------
 te.
 is
 1
                                                   CURVE    TUBE PITCH/DIAMETER RATIO
                                                     1       BASIC DESIGN
                                                     2       1-in. OD AT H=4 in., V=2 in.
                                                     3       2-in. OD AT H=8 in., V=4 in.
                                                   BED TEMPERATURE: 1750F
                                                   MAXIMUM ALLOWABLE BED DEPTH: 20 ft
                                                   BED AREA: 35 f(2
              10
20
30
40
50
60
70
80
90
100
110    120
                        BED-TUBE HEAT TRANSFER COEFFICIENT, Btu/ft -hr- F
         Figure 25.  Effect of bed-tube heat transfer coefficient on the steam generator cost
         (not including erection).
IV-2-43

-------
8
o
                                        TUBE PITCH/DIAMETER RATIO
                                        BASIC DESIGN
                                        1-in. OD AT H.4 in., V=2 in.
                                        2-in. OD AT H=8 in., V=2 in.
                               BED TEMPERATURE:  1636 F
                               MAXIMUM ALLOWABLE BED DEPTH: 20 ft
                               BED AREA: 35 ft2-
            10
20
30
40
50
60
70
80
90
100
110  120
                   BED-TUBE HEAT TRANSFER COEFFICIENT, Btu/ft -hr- F
   Figure 26.  Effect of bed-tube heat transfer coefficient on the steam generator cost
   (not including erection).
                                                                                    IV-2-44

-------
   10
8
O
LU
tu
O

1
te
 CURVE TUBE PITCH/DIAMETER RATIO
   1   BASIC DESIGN
_ 2   1-in. OD TUBES AT H=4 in., V=2 in.

 BED-TUBE HEAT TRANSFER
  COEFFICIENT: 50 Btu/ft2-hr-F
 BED TEMPERATURE: 1636 "F
 MAXIMUM ALLOWABLE
BED DEPTH: 20ft
                    FIELD ERECTED
                    WITH MAXIMUM    I
                    SHOP FABRICATION
                    (FOUR MODULES)
            // MUTIPLES OF LARGEST
           ''   SHOP FABRICATED
                MODULES
                                                10
                                              i-H

                                              X
S
k
       0   100   200    300   400   500   600 700

                    PLANT SIZE, MW
Figure 27. Dependence of the steam generator
cost on the plant size.
  CURVE TUBE PITCH/DIAMETER RATIO

     1    BASIC DESIGN
  _ 2    1-in. OD TUBES AT H-4 in., V-2 in.

   BED-TUBE HEAT TRANSFER
     COEFFICIENT: 50 Btu/ft2-hr-F
   BED TEMPERATURE: 1750F
   MAXIMUM ALLOWABLE
 r~ BED DEPTH: 20 ft
                                                  0
                       FIELD ERECTED
                       WITH MAXIMUM   -
                      SHOP FABRICATION
                       (FOUR MODULES
                                                   -   y
        ,
       /i   MUTIPLE OF LARGEST
           SHOP FABRICATED
            MODULES
 0    100    200   300   400   500   600 7 0
               PLANT SIZE, MW
Figure 28.  Dependence of the steam
generator cost on the plant size.
                                                                                     IV-2-45

-------
                                           TYPE SCREEN NO.
       0.1
       1.0

       5.0

      15.0
      25.0
      35.0
      45.0
      55.0
      65.0
                                               400     200     100   60  40
                                                                 20
                                                                10
    75.0
  
  g
  oE  85.0


      93.0


      96.0


      98.0


      99.0
                              URVES 1,2,&3 REPRESENT PROJECTED PARTICLE
                                  SIZE DISTRIBUTION LEAVING FLUID BED
                                  BOILER SYSTEM (SEE TABLE 5)
JJJ
I
                             CURVE 4 REPRESENTS PROJECTED PARTICLE SIZE
                                  DISTRIBUTION LEAVING PRIMARY COLLECTORS1*
                                  (GROUP 1)
CURVE 5 REPRESENTS PROJECTED PARTICLE SIZE
    DISTRIBUTION LEAVING SECONDARY
    COLLECTORS(GROUP 1)
 I    I   I  I MI 11      !
I    !
               1                    10                   100                  1000
                                       PARTICLE SIZE, urn
               Figure 29.  Particle size distribution for different gas streams.
                                                                            0.5
                                                                            3.0
                                                                            10.0
                                                                            20.0
                                                                            30.0
                                                                            40.0
                                                                            50.0
                                                                            60.0
                                                                            70.0

                                                                            80.0
                                                                  90.0
                                                                    .0

                                                                  97.0
                                                                            98.5
IV-2-46

-------
                                                                 s.c.
   FLUIDIZED
    BED
   COWI8USTOR
    (FBC)
         CARBON
         BURNUP
           CELL
          (CBC)
                                f
                                                     P. C.= PRIMARY CYCLONE
                                                     S. C.= SECONDARY CYCLONE
             Figure 30.  Flow diagram for particulate removal system.
                       P.C.
FLUIDIZED
  BED
COMBUSTOR
  (FBC)
  T
 CARBON
BURNUP
  CELL
 (CBC)
P.C.= PRIMARY CYCLONE
S.C.= SECONDARY CYCLONES OF SAME
     FRACTIONAL COLLECTION EFFICIENCY
                       Figure 31.  Flow chart for .group 1, case 4.
                                                                            IV-2-47

-------
        FLUIDIZED
          BED
        COMBUSTOR
         (FBC)
            \
                               CARBON
                               BURNUP
                                 CELL
                                (CBC)
                                                        PRIMARY CYCLONES OF SAME
                                                        FRACTIONAL COLLECTION EFFICIENCY
                                                  S. C,=SECONDARY CYCLONE
                          Figure 32.  Flow chart for Group 1, Case 5.
  <
 o-
  cc
  Ul
  a
  a


  o
  Of
  01
  UJ
  cc
                                               cc

                                               a


                                               o

                                               LU
                                               O
                                               
                                               cc.
                                               1
^   20  30   40   50  60  70   80   90  100  110
cc
2:                EXCESS AIR IN BED, %
  Figure 33.  Increase in first stage cyclone
  cost due to increase in gas flow rate (for
  318-MW plant).
                                                    20  30  40
                                                a:
                                                
                  50   60  70   80   90   100 110

                  EXCESS AIR IN BED, %
Figure 34.  Increase  in second stage cyclone
cost due to increase  in gas flow rate (for 318-
MW plant).
IV-2-48

-------
o
a
O
O
O
UJ
oc
3
o
      0
                             1: FIRST STAGE PRESSURE VESSELS
                                AND GAS PIPING ARE LINED WITH
                                STAINLESS STEEL
                             2: ALL LINED WITH HARD REFRACTORY
              10
20
30
70
80
90
100
                                40      50      60
                                EXCESS AIR IN BED, %
Figure 35.  Cost increments for participate removal system at different gas flow rates.
110
                                                                                     IV-2-49

-------
s
             SORBENT OUT
	AIR
	COMB. PROD.
	FUEL GAS
 STEAM AND WATER
      SOLIDS
               SORBENTIN

             COAL
                                                                                                            TO STACK
                            Figure 36.  Pressurized fluid-bed combustion power plant with secondary combustor.

-------
           3.  APPLICATION TO COMBINED CYCLE POWER
       PRODUCTION  OF FLUID-BED TECHNOLOGY USED
                              IN NUCLEAR FUEL REPROCESSING

                    B. R. DICKEY AND J. A. BUCKHAM

                            Allied Chemical Corp.


 ABSTRACT
   Fluid-bed processing is used extensively at the Idaho Chemical Processing Plant (ICPP) in the
 recovery of uranium from spent nuclear fuel elements.  Fluid-bed denitration of uranyl nitrate
 solutions and fluid-bed solidification of radioactive waste solutions are  used routinely  in plant
 operations. In addition, current pilot-plant development of processes for recovering uranium from
 graphite-based fuels depends largely on fluidized-bed combustion.
    During the course of fluidized-bed operations and development at ICPP, advantageous appli-
 cation of fluid-bed technology in areas other than nuclear fuel reprocessing has become apparent.
 Current and near-term fluidized-bed technology appears directly applicable to the conversion of
 the energy in wood  wastes and municipal refuse to electric power. Fluidized-bed operations and
 process development at the ICPP and a concept for combined-cycle power production based
 largely on fluidized-bed combustion are discussed herein.
INTRODUCTION

   Allied  Chemical  Corporation's  Idaho
Chemical    Programs-Operations    Office
operates the Idaho Chemical Processing plant
for the Atomic Energy Commission. The pri-
mary  mission of the facility is to recover
uranium from spent nuclear fuel elements in
an economic and safe manner. Metallic clad
nuclear fuels are dissolved in inorganic acids,
and the uranium and fission product solutions
are separated by solvent extraction. Uranium
rich solutions are denitrated and solidified to
uranium oxide; fission product solutions are
calcined to a mixture of metallic and fission
product oxides.

  Fluidized-bed  processing  is  used in  the
solidification of fission product solutions and
in denitration  of uranium  rich  solutions.
Fluid-bed solidification of radioactive waste
solutions has been used since December 1963.
Denitration by thermal  decomposition in a
fluidized bed was started in 1971. Both waste
calcination  and denitration  are endothermic
processes.  In-bed combustion of  kerosene
supplies the heat for waste calcination; heat
for fluid-bed denitration is obtained  from
wall-mounted electrical heaters.
   The capacity to recover enriched uranium
from graphite-based fuels will be required in
the near  future. Pilot-plant  development of
processes  for separation and recovery  of
uranium from these fuels is well underway.
The heart of these processes  is the fluidized-
bed combustion  of the graphite matrix.
                                      IV-3-1

-------
    In  the  course  of  plant  operation  and    wastes and municipal refuse through the con-
  process  development  using fluid-bed  tech-    version  to electrical  power.  The  proposed
  nology  at the ICPP, concepts based on  the    process is based on a combined gas turbine-
  application  of fluid-bed technology in non-    steam turbine  cycle;  the key  to .producing
  nuclear areas have evolved. A proposed appli-    power at high thermal efficiencies in the pro-
  cation of current interest is the recycle of wood    posed concept is a fluidized-bed combustor.
IV-3-2

-------
OPERATIONAL FLUID-BED PROCESSES
AT ICPP

Waste  Calcining Facility

   Radioactive wastes resulting from nuclear
fuel reprocessing can be solidified by various
methods that have been under investigation
over the past 20 years; however, only one pro-
cess has been demonstrated on a production
basis. The Waste Calcining Facility (WCF) at
the Idaho  Chemical Processing Plant  is the
first production  facility  in  the  world  for
converting  aqueous  radioactive  wastes  to
solids  using the fluidized-bed  calcination
process.

   A  schematic  flow sheet for  the WCF is
shown in Figure  1. The heart of the process is
the 4-ft  diameter calciner vessel where the
radioactive  aqueous wastes are continuously
atomized into the fluid ized bed of oxide par-
ticles by an external-mix pneumatic atomizing
nozzle.  In  the fluidized bed, which is  main-
tained  at  500 C,  water  evaporates,  the
metallic  salts   are  converted   to   their
corresponding oxides or fluorides which  are
deposited layerwise on the  spherical bed par-
ticles.  The  solid  particles  are  withdrawn
continuously from the calciner vessel to main-
tain a constant bed height and are transported
pneumatically to stainless  steel  solid storage
bins adjacent to the calcining facility.

   The  off-gas  leaving  the calciner  vessel
passes  through   a dry  cyclone where the
majority of elutriated fines are removed and
pneumatically  transported  to  the  solids
storage bins. The off-gas then passes through
a wet scrubbing system consisting of a quench
tank, venturi scrubber, cyclone separator, and
a demister. Here, the off-gas is contacted with
a nitric acid scrubbing solution which removes
the majority of  the  remaining  solids. The
scrubbing   solution    is   recirculated
continuously,  and  any  accumulation   is
recycled to  the waste feed tank.

   Heat  for  the  endothermic  calcination
reactions is supplied by in-bed combustion.
In-bed  combustion consists  of atomizing a
hydrocarbon fuel (kerosene) with pure oxygen
directly in the  fluidized  bed.  Startup of the
process is  achieved by heating the fluidized
bed to temperatures in the range of 360 to
400C  using   preheated  fluidizing  air.  A
nitrate-containing waste  is  then  injected
through  a separate waste  atomizing nozzle,
followed immediately by the  injection of the
fuel-oxygen mixture  through  the  fuel
atomizing nozzle. Ignition of the  fuel-oxygen
mixture is spontaneous at temperatures above
335C in the presence of nitrates. After the
startup,   the   wastes  are  calcined  at
temperatures in the range of 400 to 500_C
during routine  operation. The advantage of
this method of heating is that no heat transfer
surfaces  are involved that  can foul  or  limit
capacity;  the  heat  flow  paths   of in-bed
combustion heating  and  an in-bed  heat
exchanger  are compared in Figure 2.

Denitration facility

   The  denitration  process is based on the
thermal decomposition of uranyl nitrate  solu-
tion  to  uranium trioxide. In fluidized-bed
denitration, solution is continuously sprayed
through  an air atomizing nozzle into heated
fluidized bed of UO3. The bed temperature is
300C;  the pressure immediately  above the
support plate is approximately atmospheric.
Granular product is  continually withdrawn
into a product  collection vessel. The process
off-gas~consisting mainly  of  fluidizing air,
water vapor, and oxides of  nitrogen-flows
through a filtering section consisting of  three
sintered metal filters which remove over 99.9
percent of the entrained  and  elutriated  UC*3
dust particles.  The  filters are blown  back
intermittently, and the fine particles serve as
seed particles for particle growth. A schematic
flow-sheet of the fluid-bed denitration process
is  shown in Figure 3.

FLUID-BED BURNING PROCESSES

   Pilot-plant development of a fluidized-bed
combustion process for separating uranium
from spent graphite-matrix nuclear fuels has
been in progress since January 1966 at ICPP.
This unique combustion process  is required
                                                                                   IV-3-3

-------
because graphite, unlike the metallic cladding
of the more conventional nuclear fuels, is not
readily dissolved in common  inorganic acids.
As the graphite matrix is removed by combus-
tion,  the  uranium  and  other  metals are
converted to their oxides. Dissolution of the
resulting uranium oxide, U3O8, completes the
burn-leach head-end process. The uranium is
separated  from fission products  and  other
impurities  by conventional solvent extraction.

   The  fuel consists  of uranium  dicarbide
microspheres coated with pyrolytic  carbon
dispersed in a graphite matrix. A  protective
coating of niobium carbide is  present on some
of the surfaces of the fuel elements. Although
aluminum or stainless steel orifices are present
in  some   elements,   there  are  only   three
constituents of consequence to the combustion
process:  uranium, niobium, and carbon.

   The   fluidized-bed  combustion  process
developed  at  ICPP for nuclear rocket  fuels
involves the following concepts: (1) charging of
whole fuel elements to a fluidized bed of inert
alumina   particles,  (2)  combustion  of
essentially all carbon, (3) oxidation of uranium
and niobium carbides, and (4) elutriation of
uranium and niobium oxides from the burner.

   Requirements  for  the  fluidized-bed
burning  and elutriation  processes are:  (1)
combustion of at least 95 percent of the matrix
graphite and pyrolytic carbon,  (2) conversion
of the uranium dicarbide  microspheres and
niobium  carbide to  elutriable particles  by
oxidation and attrition, (3) negligible attrition
and elutriation of the inert bed material (a-
alumina), and  (4) adequate heat dissipation
for control of bed temperatures. If "steady-
state" conditions  can be achieved,  all of the
uranium  and niobium charged to the burner
will be elutriated in the burner off-gas. The
amount of alumina  and unburned carbon
carried overhead with the U3O8 product must
be minimized.

   The fluidized-bed burner  segment of the
Graphite Fuels Pilot Plant (GFPP)  (Figure 4)
was  constructed  to  permit  studies  of the
combustion process. The burner and leaching
equipment, located on two adjacent modules,
can  be operated independently or simultan-
eously.

   Major components of the  burner  module
are a fluidized bed for burning the fuel and a
dry product collection system for filtering and
retaining particulates from the burner off-gas.
The dry product collection system is bypassed
when  direct  introduction  of the   burner
product   into  the   leaching   equipment  is
desired.
   Several burner-designs were used  in  the
course of pilot-plant development.  The final
burner  design, the  concentric fluidized-bed
burner, is shown schematically in Figure 5. A
4-in.  diameter,  4-1/2-ft  long  first-stage
burner is located concentrically inside a 6-in.
diameter vessel which extends 8-1/2 feet above
the top of the first  stage. Fluidized beds  are
contained in the 4-in. diameter first stage,  the
surrounding annular space, and in the 6-in.
diameter  section above the first stage. The
wall of the upper  2  feet of the first-stage
burner  is  slotted to  allow particle  mixing
between  the annular and inner beds.
   The  concentric fluidized-bed  design  was
proposed  on the  basis of potential  increased
heat transfer rates from the  wall of the inner
first-stage burner. The outer wall of the first-
stage burner in an earlier  two-stage fluidized-
bed  burner was cooled by forced-air  convec-
tion; this proved to be adequate under normal
conditions but  inadequate in  the event of a
temperature excursion.  Substitution  of  the
annular bed for the forced convection system
was  expected  to  increase the heat transfer
coefficients at the first-stage wall by an order
of magnitude.  The  annular   fluidized  bed
provides the added  advantage of secondary
containment should melt-through of the inner
vessel occur.

   Tests conducted with a 6-in. diameter glass
column containing the actual 4-in. diameter
inner bed further showed  that increased heat
transfer could also be  expected from  particle
IV-3-4

-------
mixing between the inner and annular beds.
The intensity of slugging in the inner bed was
greatly reduced in the upper 2 feet due to the
transfer of gas and particles between the inner
and annular beds.

   The  anticipated  improvement  in   heat
transfer using the concentric  fluidized-bed
burner has been realized. The concentric-bed
design has proved superior to an original two-
stage burner with respect to heat transfer and
temperature   control.   No temperature
excursions  have   occurred  in  any  of  the
experiments. Temperature differentials within
the first stage have ranged  from a normal
spread of 25 to 45 F to a maximum spread of
75F.
   Combustion efficiencies equal to or greater
than the required 95 weight percent have been
obtained in the two-stage concentric fluidized
bed  over  the  following ranges of operating
variables:

  1. Nominal bed temperature1400 to 1500F.
  2. Fluidizing-gas composition80   to  100
   percent oxygen to both stages.
  3. Mean fluidizing velocities (average of inner
   and  annular fluidizing velocities)   1.05
   to 1.50 ft/sec.
  4. Fuel   charging  ratesup   to   33   kg
   graphite/hr-ft2.


APPLICATION     OF     FLUID-BED
COMBUSTION  TO  COMBINED-CYCLE
POWER  PRODUCTION  FROM  WOOD
WASTES AND MUNICIPAL REFUSE

Problem Definition

   At present, approximately 250  x  10 _  tons
(190 x  10'6 tons  of which are collected) of
residential,  commercial,   and   institutional
wastes are produced in the United States each
year. The per capita generation of such wastes
is rapidly  increasing;  combined  with  an
expanding population, the magnitude of the
problem may double by  the  turn  of the
century. The disposal of refuse is -becoming a
critical  problem for communities, especially
those in the heavily populated parts of the
country.

   The   lumber  industry,  particularly  the
small-to-medium size  mills,  has a  similar
waste disposal problem. Approximately  5 x
106 tons of wood wastes are produced in the
United  States per year. The volume of waste
makes landfill disposal  impractical; the wastes
are  usually  burned  in  inefficient  teepee
burners.  While  economical,  such  systems
cause  localized  air pollution by emitting
smoke,   particulate  matter,   and partially-
oxidized  chemicals.  Failure  to  meet  the
stringent air pollution  standards  (now being
introduced) will prevent future operation of
teepee burners.

Recycle of Solid Waste By Energy Production

   Through regional planning and  cooper-
ation, wood waste materials  and  municipal
refuse  would  be transported to  a   central
location  and  burned  in  a highly efficient
pressurized fluid-bed burner. Heat released in
the bed would  be  used to generate steam
within  an in-bed heat exchanger; both  the
steam and high temperature off-gases (1400 to
1500F) would generate electrical power using
steam and  gas  turbine  cycles, respectively.
Removal of heat by generating steam in an in-
bed heat exchanger would also minimize  the
amount of excess air  normally  required  to
control the bed temperature; thus, the volume
of off-gas requiring cleanup would be reduced.
Technology  required for plant-scale  process
demonstration is either already established or
in the latter stages  of  development.

 Advantages of the proposed system are:

 1. Significant reduction in present air and
   solid waste pollution (to meet present and
   future standards).

 2. Decreased  land requirements for disposal.

 3. Conservation of natural resources (fossil
   and nuclear fuels) by recycling wastes to
   produce electrical power.

 4. Decreased  cost of waste disposal.
                                                                                  IV-3-5

-------
   While the concept is  one in which  both
 wood  and  municipal wastes  are  available,
 successful development  and  plant-scale
 demonstration of the proposed concept could
 lead to use in areas where either type of fuel
 predominates.  The concept should find  wide
 application throughout the Pacific Northwest,
 the North,  and the Southeast.

   A schematic  flowsheet of the conceptual
 plant is shown in Figure 6. Although the plant
 would be designed to process wood  waste and
 municipal refuse as  the  principal  fuels, the
 basic concept is  compatible with other waste
 (e.g., industrial wastes, sewage sludge, etc.) as
 fuel. In the future, the plant could be modified
 to accept feed in the form  of low sulfur coal as
 fuel in the event that advancing technology
 results  in better use  of the waste materials.

   After  the  required  preparation   (e.g.,
 screening, shredding, and storage), the waste
 is partially dried and charged  to a fluid-bed
 burner operating  at  150 psia and  in  the
 temperature  range of 1400 to 1500F.
 Combustion occurs in an inert bed of sand;
 the efficient solids and gas contact in  the bed
 results in rapid and  complete combustion. Ash
 and  noncombustibles are  continually  with-
 drawn from the  burner;  some particulate is
 carried overhead in the off-gas.

   Steam   is   generated   from  condensate
 passing through  the tubes of an in-bed  heat
 exchanger.  The  superheated  high pressure
 steam then flows  to a steam turbine-generator
 to produce  electrical  power.  Flue gases  from
 the fluid bed are cleaned of particulate matter
 using a combination of  cyclones  and  high
 efficiency filters (ceramic  or sintered metal).
 The clean flue gas then flows to a gas turbine
 to generate additional electric  power.

   The  concept of combining gas and steam
turbine cycles in a system for generation of
electric power is not new, and the high thermal
efficiencies  possible in such cycles  are being
demonstrated in  utility power stations.  The
 San Angelo Station of West Texas Utilities,
for example, has achieved  41 percent thermal

IV-3-6
efficiency while burning natural gas in a gas
turbine exhausting to a conventional boiler-
steam turbine  system.

   The concept of generating steam in tubes
immersed in fluidized beds of solids also is not
new.  Work has been  under way  for  several
years  to  develop  this system,  both  in  the
United States  and  abroad.  The  British, in
1969,  speculated  that  the  advantages  of
pressurized fluid-bed boilers deserved further
study including a mixed cycle incorporating a
gas turbine.4 In the  United States, pilot-scale
work has been  under way for some four years
to develop a  system  for pressurized fluid-bed
combustion of solid wastes using hot exit gases
to produce electric  power in a gas  turbine
generator.5
   The concept  proposed  herein  includes
elements from all of the aforementioned work;
however, the cycle proposed is unique and  has
distinct advantages  over  other  proposed  or
existing processes for  processing wood waste
and  municipal  refuse.  The proposed cycle is
shown in  Figure 7;  wood wastes would be
dried to less than 10  percent moisture. This is
important  to the  overall thermal  efficiency
because low-level turbine exhaust heat, much
of which is wasted in  other cycles, would be
used  to dry the  wood wastes which  may
contain up to 50 percent moisture.

   Thermodynamically,  the   cycle  is  very
attractive, particularly for the combustion of
waste  materials  with  high  water  content.
Performance data for  one set of conditions,
not necessarily  optimum,  are  summarized in
Table  1.

   The plant can be divided into four major
sections: feed preparation, fluid-bed burning,
off-gas cleanup,   and   power   generation
facilities.  In  feed preparation,  the removal
efficiency   of  noncombustibles  from   the
municipal refuse has a significant impact on
the operation of the fluid bed.  Operation of
the fluid-bed burner  (including feed introduc-
tion, ash removal, and in-bed heat transfer)
and   off-gas cleanup  are   critical  to  the
successful  operation of the proposed plant.

-------
      Table 1. PERFORMANCE DATA FOR
             PROPOSED CYCLE
Basic parameters
Waste moisture content
Gross heating value, dry
Air pressure to fluid bed
Gas pressure to gas turbine
Air temperature to compressor
Fluid bed and exit gas temperature
Excess air
Turbine exhaust pressure
Turbine exhaust temperature
Steam pressure
Steam temperature
Feed water temperature
Condensing pressure
Fluid-bed dryer and exit
gas temperature
Calculated performance/1 00 Ib dry
fuel
kWhr, steam turbo-generator
kWhr, gas turbine turbo-generator
Energy, kWhr/100 Ib dry fuel
Thermal efficiency, %

50%
8075 Btu/lb
150psia
145psia
60F
1440F
15%
17 psia
720 F
1500 psia
1000F
415F
1.5in. Hg
150F



77.0
11.5
88.5
37.5
The power generating facilities will be conven-
tional and do not require detailed discussion.
A process  flowsheet of the demonstration
plant is shown in Figure 8.

   Feed Preparation and Storage

   Feed to  the fluid-bed burner  consists  of
wood  waste  (e.g.,  sawdust,  chips,  and
shredded  material) and municipal  refuse.
Wood wastes are relatively homogeneous and
require only sizing and  drying before feeding
to the bed. The heterogeneity of municipal
refuse requires separation of glass and  metals
and  sizing  before  being introduced  to the
tluidized bed. Based on a minimum amount of
engineering  development,  the  A-E  would
select  a   feed   preparation  scheme   for
facilitating  materials  handling,   reducing
environmental pollution, and providing safe
storage.

   Wood  Waste   Feed  Preparation  and
StorageWood waste would be properly sized
at the mill site for feed to the burner. Raw feed
which  is   sufficiently   dry   (<10  percent
moisture)  would be transported  directly to
 feed storage. All other feed  would be dried
 approximately  10 percent moisture before
 storage.
    Municipal Refuse  Feed Preparation  and
 Storage  With the exception of moisture
 content, a  typical municipal refuse content
 and composition is shown in Table 2. The
 municipal refuse consist of glass, dirt, metal,
 and  various  combustible  materials.  Size
 distribution  of the refuse varies  from large
 pieces of material to  dust particles; moisture
 content is a nominal 25 percent by weight. For
 rapid  combustion, good  quality fluidization,
 and  satisfactory  materials handling, refuse
 must be sized to less than 1-in. pieces. The
 water content is usually lowered to less than 10
 percent during normal feed preparation (e.g.,
 shredding  and classification);  therefore,
 drying  of the refuse is not anticipated.

    Table 2. TYPICAL MUNICIPAL REFUSE
              COMPOSITION       	
              (Yard-free basis)
       Material
    Paper, > 1/4 inch
    Paper, wood, fabric fines
    Wood
    Fabrics
    Plastics
    Inerts (glass and metallics)
    Dust
    Heating value, Btu/lb (dry basis)
 Wt%
48.9
11.7
10.2
 1.0
 2.0
14.0
11.7
6000
   A conceptual flowsheet for feed prepara-
tion of the municipal refuse is shown in Figure
9. The  raw refuse is  dumped  into a  feed-
conveyor hopper directly from a truck; no raw
feed storage is provided. After charging to the
shredder, the material is transported  to a
classifier where the large and/or dense parti-
cles  are  separated in the  underflow.  As
dictated by economics, ferrous metals may be
separated from the underflow for recycle.

   Feed preparation equipment will be  over-
sized to allow for routine maintenance without
limiting  the  plant  capacity  and  would
normally be operated only one shift per day. If
economics  require, refuse feed  preparation
                                                                                   IV-3-7

-------
 could be done at the source; this is especially
 true in  the  case  of a  large  city such  as
 Spokane.

    Fluid-Bed Combustion

    The  fluid-bed burner  (Figure 10) is the
 heart of the proposed system. The maximum
 bed temperature is limited by the gas turbine
 blade  materials and the minimum  fusion
 temperature of the noncombustibles and ash
 present in the bed. Fluid-bed burners have a
 high heat release rate per unit volume of bed;
 because of the variability in the heat content of
 refuse, control  of bed temperature  could  be
 difficult  when  charging municipal  refuse
 alone. However, temperature control problems
 in the proposed concept should be minimal,
 since the wood waste (relatively homogeneous
 and of constant heat value) will comprise one-
 half to two-thirds of the feed to  the burner.

    Condensate  within tubes of an in-bed heat
 exchanger is converted to steam by transfer of
 heat  from  the fluid  bed.  Depending  on
 economics  and technical considerations, bare
 or  finned  tubes will be  used. Based on  a
 practical carbon steel tube  bundle  configur-
 ation and overall heat transfer coefficients in
 the range of 300 to  600 Btu/hr-ft2-F, heat
' transfer rates in the range of 4 x 105 to 8 x 105
 Btu/hr-ft3  of bed  are  possible.  For a 125
 ton/hr plant, approximately 1 x 106 Ib/hr of
 steam (1200 to  1500  psia and 900 to 1000F)
 would be  generated. Based on  a nominal
 steam cycle efficiency of 35  percent, approx-
 imately 100 MW would be  produced  by the
 steam turbine.

    Flue gases from the fluid-bed burner pass
 through  a high  efficiency  off-gas cleanup
 system (series of cyclones followed by ceramic
 or sintered  metal filters)  for removal of essen-
 tially  all particles > 5 Mm in diameter. The
 cleaned off-gas  is then passed to a gas turbine
 where, based on a  125  ton/hr plant and  a
 turbine cycle efficiency of 20 percent, 14 MW
 are produced.

    Though most of the ferrous and nonferrous
 metals will be removed by magnetic separation
and   air  classification,   some  of   these
components will  remain in the feed to the
burner.  In addition,  glass,  dirt,  ceramic
material,   and   possible   agglomerates  of
material (e.g.,  plastics, etc.) must be  removed
from the burner.  If the material  is allowed to
accumulate on the air distributor, fluidization
quality  would  deteriorate as a result of gas
channeling. A reliable system  for removal of
ash  agglomerate and noncombustibles  is
required.

   Off-Gas Cleanup System

   Removal of  entrained  particulates  and
corrosive gases is a  major problem in the
design  of  conventional incinerator  off-gas
cleanup systems; expansion of the off-gas in a
turbine requires even a higher degree of off-
gas cleanup. Particulates, if not removed, will
erode and foul turbine blades and pollute the
environment. Corrosive gases will also corrode
the off-gas cleanup system and turbine blades.
Fewer  corrosive gases are  generated during
processing of wood waste; therefore, the wood
waste could result in off-gas concentrations of
corrosive  gases  of  acceptable   levels.
Although maximum particulate loading speci-
fications for turbine inlet gases vary,  2 x 10"3
gr/ft3 of off-gas, with no more than 10 percent
of the particles greater than 10 urn in  size, is a
typical requirement. Because of the control of
bed temperature by transfer of heat to an in-
bed steam  generator, the excess  air flow and
hence  volume  of gas requiring cleanup is
minimized  using  the proposed concept.

   Conventional  Particulate  Removal  
Particulates are normally removed from off-
gas  streams   by   combinations   of  cyclone
separators, scrubbers, and electrostatic
precipitators; each type of cleanup device has
distinct disadvantages. Cyclones  are normally
effective in removing particulates at high gas
rates as long as the particle  size is larger than
20  nm.  Scrubbers  require  low  operating
temperatures,  and electrostatic  precipitators
are often ineffective. Sand or granular filters,
though  sometimes employed,  are bulky and
regeneration is difficult.
 IV-3-8

-------
Combination Multi-Stage  Cyclone  and
"Candle" Filters  for Participate Removal 
The concept of filtering off-gases at high tem-
peratures is not new; sintered metal filters are
commonly  employed  up to  1500F.  Such
filters are  available  for removing  submicron
particles at pressure differentials less than  1-
in.  water.  Ceramic "candle" filters, used  in
conjunction with  multi-stage cyclones, appear
to be most attractive for removing particulates
to the levels recommended by turbine manu-
facturers.  Well  designed 3-stage  cyclones
could almost  satisfy the requirements  alone,
and  candle   filters  could  further  reduce
particulate loading  to levels below process
requirements.   The  cyclones  would  be
constructed of a material resistant  to chloride
corrosion at temperatures as  high  as 1500F.

    In the proposed concept,  off-gases at 145
psia and 1400 to  1500F  would  leave the
burner and pass through multi-stage cyclones
followed by candle filters for  final  particulate
removal, as shown in  Figure 11. The  multi-
stage cyclones would  remove  essentially all
particulates  greater  than 20 pirn. Off-gas
passing to the turbine would contain a max-
imum  loading  of  1  x 10"3  gr/ft3,  and
essentially  all particulates  greater than  5
microns would be removed.  The multi-stage
cyclone and candle filter systems are commer-
cially available.

 CONCLUSIONS

    Disposal of municipal refuse  and  wood
 waste while minimizing harmful effects on the
 environment is an existing problem. The per
 capita  increase  of refuse  generation  rates
 coupled with the  expected population increase
 will almost double the annual refuse generated
 by 1980. Wood waste generation rates are also
 expected to increase, though not as rapidly as
 municipal refuse.

    Existing methods  for municipal  refuse
 disposal (landfill and conventional inciner-
 ation) frequently  result in air and water pollu-
 tion and unsightly facilities. Disposal of wood
 waste by  incineration in teepee  burners  is
unacceptable from the standpoint of satisfying
air pollution standards. Development of more
advanced systems for disposal of municipal
refuse is based largely on the concept of pro-
duct recycle. The recycled products have ques-
tionable market value when compared with
virgin materials.

   Existing   and   near-term   fluid-bed
technology can result in a fluid-bed process for
power generation having a thermal efficiency
greater than 35 percent while reducing the off-
gas mass flow rates per MW-hour  of electrical
energy by a  factor of eight when compared
with a conventional gas turbine cycle. In such
a process, steam would be generated within an
in-bed  heat  exchanger and used   as  the
working fluid in  a  steam-turbine cycle. The
latest technology in finned-tube heat transfer
within fluid beds would be used.  The burner
off-gases would be cleaned of particulate  by
passage  through  a  high  efficiency  cleanup
system  consisting of  staged cyclones and
sintered metal (stainless steel and Hastelloy C)
or  ceramic   "candle"  filters.   Particulate
removal at the temperatures (1400 to 1500F)
proposed for refuse and wood waste combus-
tion has  been tested in development and
operations facilities in the atomic energy field.
Use of existing technology and commercially
available  equipment where  applicable will
significantly reduce the development costs and
time  required  to place  a  full-scale  waste
disposal-power generation facility on stream.
Power production costs are estimated to be in
the range of 5 to 6  mills/kWhr.

BIBLIOGRAPHY

1. Bendixsen, C.L. et al. The Third Processing
   Campaign in the Waste Calcining Facility.
   U.S.  Atomic  Energy  Commission, Oak
   ridge, Tenn. Publication Number IN-1474.
   May 1971.

2. Kilian, D.C. et al. Description  of the Pilot
   Plant for the Headend  Reprocessing of
   Unirradiated   Rover Fuels. U.S.  Atomic
   Energy  Commission, Oak Ridge,  Tenn.
   Publication Number IN-1181. May 1968.
                                                                                    IV-3-9

-------
 3. Cox, A.R., et al. Operation of San Angelo        Combustion. The Engineer, July 24, 1969.

   Power Station  Combined  Steam and Gas     c  ,,   .    .        T, . Ann  /-..*
   ,.,._     ,.     .    .        5. Combustion Power Umt-400. Combustion
   Turbine Cycles. In: Proceedings American             _   n ,  A1,   .-, ,.       , ~
         _  J_      A7 ,     VVTV    AM        Power Co., Palo Alto, Calif. Prepared for
   Power Conference, Volume XXIX, p. 401-        ,.          ~ c ,., l,T  .  .         ,
   41-             '               v             the Bureau of Solid Waste Management,

                                                Environmental Protection Agency, Wash-

                                                ington, D.C. under Contract Number Ph

 4. Big Economics  from Pressurized Fluid-Bed        86-67-259.
IV-3-10

-------
                 WASTE SOLUTION

               _/           RECYCLE
            ATOMIZING
              AIR
            ATOMIZING
              OXYGEN


            KEROSENE
                          SCRUBBER SEPARATOR.
                       VENTURISCRUBBER
                                                                                                                         ICPP
                                                                                                                        STACK
                  BLOWER
 SOLIDS TRANSPORT AIR
                                                                             SOLIDS STORAGE BINS
                                                                                                             BLOWER
                                                   Figure 1.  Waste calcinating facility.
CO

-------
                                  CIRCULATING
                                HEAT TRANSFER
                                  MEDIUM
               I
                                                   HEAT
                                                   SINK
           BURNING
            FUEL
METAL
WALL
LIQUID
METAL
METAL
WALL
FLUIDIZED
PARTICLES
ENDOTHERM1C
  REACTION
                                    IN-BED HEAT EXCHANGER
                         BURNING
                          FUEL
                 FLUIDIZED
                 PARTICLES

                IN-BED  COMBUSTION
                  ENDOTHERMIC
                     REACTION
                           Figure 2.  Comparison of heat flow paths.
            BLOWBACK
                AIR
                                                       TO VESSEL
                                                        OFF GAS'
                                        GLOVE BOX
                                                          BAG
                                                        Z30UT
                                                          PORT
                                                 CANNER
                           Figure 3. Product denitration facility.
IV-3-12

-------
                              TO EXHAUST STACK
            CHARGING GLOVE
                  BOX
 ALUMINA
CHARGING
  POT
COOLING
AIR

FLUIDIZING
GAS
                                                   A  TO PLANT OFF-GAS
                                                           SYSTEM
                                                      TO DISSOLUTION
                                                          STEPS
    GRAPHITE
    BURNER
                                                 DRY SOLIDS
                                                 COLLECTION
                                                   VESSEL
  Figure 4.  Graphite fuels pilot plant-combustion process flowsheet.
                                                                        IV-3-13

-------
       P  TW
    COOLING AIR
     INLET
   GAS INLET TO
    INNER BED
                                    OFF GAS
                                    OUTLET
                           ALUMINA CHARGING PORT
                            FUEL CHARGING TUBE
                             AUXILIARY GAS INLET
6-in. SCHEDULE 40 PIPE
TYPE 316 STAINLESS STEEL
                                 BED SAMPLING LINE
                                    COOLING AIR OUTLET
                                                     MlO
                                HEATING ELEMENTS
          PNEUMATIC
          OPERATOR

          BALL VALVE
          SAMPLING LINE
Hn. SCHEDULE 40 PIPE
TYPE 304-L STAINLESS STEEL
                               STAINLESS STEEL SHROUD
                             /-HEATING ELEMENT
                                      FLUIDIZING GAS
                                      DISTRIBUTOR
                                      PLATE
BED SUPPORT PLATE AND
GAS DISTRIBUTION
       GAS INLET TO
       ANNULAR BED
                                                     QUO
                                                             DEFLECTION BAFFLES
                                FUEL CHARGING TUBE
                                1-M-in. SCHEDULE 40 PIPE
                                HASTELLOY ALLOY-C
                                                             PERFORATED PLATE BAFFLES
                                                               FUEL CHARGING TUBE
                                  3/8-in.x3-in. PERFORATIONS
                                  STAGGERED ON %-ln. CENTER
                                     LEGEND

                              P-PRESSURE TOP
                              THNTERIOR TEMPERATURE
                                THERMOWELL
                              TW-VESSEL WALL TEMPER-
                                ATURE THERMOWELL
                                   INNER BED DRAIN
                                   ANNULAR BED DRAIN
                       Figure 5. Concentric bed burner.
IV-3-14

-------
                                                           INERT MATERIALS
                                                           (ASH, GLASS, AND
                                                           ROCKS) TO A
                                                           LANDFILL
RAW
MUNICIPAL
REFUSE

WOOD
WASTES
1
1
1
MUNICIPAL
REFUSE
PREPARATION
RECYCLE OF !
GLASS I
L
r
i
WOOD
STORAGE
               PLANT BOUNDARY
                                          FLUIDIZED
                                            BED
                                         COMBUSTOR
                                        ELECTRICAL
                                          POWER
                                          (STEAMS
 ELECTRICAL
   POWER
GAS TURBINE
                                      SALABLE ELECTRICAL POWER
Figure 6.  Schematic of waste disposal-power generation facility.
                                                                   IV-3-15

-------
CO
h-1
O5
                                                                                         TO ATMOSPHERE
                                                                                    SCRUBBER
     SUPERHEATED
        STEAM
                                                           PREPARED
                                                           MUNICIPAL
                                                           REFUSE
                                                                                             L
 FEED WATER
   PUMP
                                                                                          WET WOOD
                                                                                            WASTE
                                                                               HOT TURBINE
                                                                               EXHAUST GAS
Figure 7.  Proposed cycle for producing electric power from wood waste and municipal refuse.

-------
                                                                        FLUIDIZED-BED
                                                                        INCINERATOR
                     WOOD FEED
                     STORAGE
                      MUNICIPAL
                        WASTE
                                                                                 FLUIDIZED-BED
                                                                                 INCINERATOR
                                                   GENERATOR            I  COMPRESSOR

                                                              GAS TURBINE
                                                                                                         ASH
<3
CO
                                           Figure 8-  Flowsheet of proposed waste recycle-power plant.

-------
   RAW MUNICIPAL
     REFUSE
PREPARED FEED
     STORAGE
                                                                                TO STACK
                                                GAS
                                                                   TO FLUID-BED BURNER
                                                        NONCOMBUST1BLES
                     Figure 9.  Municipal refuse feed preparation facility.
IV-3-18

-------
   SUPERHEATED
      STEAM
       CONDENSATE-
                                                    -OFF-GAS
PEPARED
 FEED
       PNEUMATIC
       CONVEYOR

li'Vi-'''i ''j,j' ^' i''!!z,;.'fjiji''2,
/A v/a Vfa //t  ra r/a VIA v// m
                                                     -CYCLONES
                                                      SUPER HEATER
                                                        SECTION
                           t* -*STEAM GENERATOR
                                 EVAPORATIVE
                                   SECTION
                                                      GRID SUPPORT AND
                                                       AIR DISTRIBUTOR
                         HIGH-PRESSURE COMBUSTION
                           AND FLUIDIZING AIR

                      Figure 10.  Fluid-bed burner.
                                                                            IV-3-19

-------
                          BANK OF
                        CANDLE FILTERS
      MULTI-STAGE
       CYCLONES
     FEED
                      Figure 11. Off-gas cleanup system for proposed plant.
IV-3-20

-------
                  4. POWER GENERATION USING THE  SHELL
                                               GASIFICATION PROCESS

           A. N. DRAVID, C. J. KUHRE, AND J. A.  SYKES, JR.
                         Shell Development  Company

ABSTRACT
   Growing concern about sulfur and nitrogen oxide emissions has given rise to a search for means
of converting conventional fuels into clean,  non-polluting fuels for electric  power generation.
Through its ability for converting liquid fuels into partially oxidized gaseous fuels and recovering
the heat of partial oxidation in the form of high pressure steam, the Shell Gasification Process
(SGP), aided by the Shell Sulfinol or ADIP Process, offers an attractive means of converting sulfur-
laden heavy hydrocarbon feedstocks of high metals  content into  non-polluting fuel gas and
saleable elemental sulfur.
   Conceptual design and economics of an SGP-based power plant utilizing the Combined Gas
and Steam (COGAS) cycle are presented. Besides offering the simplicity, flexibility, and reliability
associated with the SGP, such a power plant can generate electric power at unit costs competitive
with those of future conventional power plants.
 INTRODUCTION

   Thp Shell Gasification Process (SGP)* is a
 process for the partial combustion of hydro-
 carbons,  and is particularly  suitable for the
 partial combustion of heavy, sulfur containing
 residual fuels and heavy crude oils to produce
 a mixture of hydrogen and carbon monoxide.
 From this mixture the hydrogen sulfide pro-
 duced during partial oxidation can be readily
 removed. A non-polluting fuel gas is thus pro-
 duced  which  can  be   used   for  power
 generation. This type  of fuel should be of
 particular  interest  for  power  generation
 because of the following factors:

 1. Natural gas, for many years a sulfur-free
   fuel, has slid into a declining reserve posi-
   tion in the face of an increasing demand at

 "Licensed by Shell Development Company,  Houston, Texas
 77001  and Shell Internationale Research, Maatschappij, N.V.,
 The Hague.
    present regulated prices.
 2. The cost and  difficulty of desulfurizing
    heavy  fuel oils, particularly  those  with
    high metals content, is very high.
 3. Eastern and mid western coals have  high
    sulfur  contents which  to  an  increasing
    extent  make them  unsuitable  for use for
    generation of power in conventional steam
    power  plants.
 4. There is a relatively long lead time for the
    development of low sulfur western coals,
    and also high  transportation  cost asso-
    ciated  with the use of these coals.
 5. There is a long lead time required for the
    installation of nuclear power  plants.

   As an alternative method  of  power  gen-
eration, the Shell Gasification Process, with a
                                        IV-4-1

-------
 moderate  investment  and  a high  thermal
 efficiency (>85 percent), converts fuels with
 high levels of sulfur, nitrogen and/or metals
 into attractive power generating fuels for use
 in  the COGAS  cycle  (Combined  Gas  and
 Steam cycle).

   An SGP-based power plant consists of a
 Shell Gasification Process unit, for converting
 residual fuels or low-value crudes into low-Btu
 fuel gas and recovered steam, and a gas tur-
 bine-steam turbine unit for converting these
 products into electrical power. An SGP-based
 power  plant differs  from  a  conventional
 thermal  power  plant:  in  a  conventional
 thermal power  plant  raw  fuel  is  burned
 directly, while in an  SGP-based plant the
 power plant fuel is the product of partial oxi-
 dation of the raw fuel.
   The power generation  unit recommended
for the  SGP-based power plant  uses the
COGAS cycle. The COGAS cycleJ-2'3 is ther-
modynamically superior to either the steam
cycle or the gas cycle. It is particularly suited
for an SGP-based power station, since in the
SGP the net exothermic heat of partial oxi-
dation is  recovered  as high  pressure steam
which can be integrated with the steam section
of the COGAS cycle.

   The following technical  and economic case
study shows that an SGP-based power station
is  not  only a  feasible  means  of generating
power  without  contributing to  atmospheric
pollution, but it is also  economically competi-
tive with conventional power  plants of the
future.
IV-4-2

-------
CHEMISTRY OF PARTIAL OXIDATION

   Partial oxidation describes the net effect of
a number of component reactions that occur
in a flame, supplied with less than  stoichio-
metric  oxygen.  This  net  effect   can  be
approximated:
and is actually a combination of several reac-
tions that occur within the reactor.

Heating-op and cracking phase

   In the fuel injection region of the reactor,
hydrocarbons leaving  the atomizer at about
preheat  temperature  are  intimately mixed
with air. Prior to combustion they are heated
and vaporized  by back radiation from the
flame and the reactor walls. Some cracking of
the  hydrocarbons to  carbon, methane,  and
hydrocarbon radicals  may take place during
this brief interval.

Reaction phase

   As soon as  the  ignition  temperature is
reached, part of the hydrocarbons will react
with  oxygen   according   to  the  highly
exothermic reaction:

CnHm+(n + )O2 -* n CO +H O.  (2)

As the  equilibrium  is far  to  the  right,
practically all  the available  oxygen is  con-
sumed in this phase. The remaining hydrocar-
bons which have not been oxidized react with
steam  and the combustion  products from
reaction  (2) according to  the  endothermic
reactions:
CnHm + nC02
and
                     2nCO+f-H2
 CnHm + n H20- n CO + (21 + n) H2 .  (4)
   In order to prevent excessive local tempera-
tures, it is essential that all reactants of equa-
tions (2) to (4) are intimately mixed so that the
endothermic  reactions  tend to  balance the
exothermic reactions. In this way the complex
                                             of reactions   is  brought  to  a  thermal
                                             equilibrium resulting in a measured tempera-
                                             ture of about 2350 to 2550F.

                                             Soaking phase

                                                Soaking takes place in the rest  of the
                                             reactor where the gas is at a high temperature.
                                             The gas composition changes only slightly due
                                             to secondary reactions of methane and carbon
                                             and the water gas shift reaction.

                                                Methane  produced  by  cracking will de-
                                             crease according to:
                                                    CH4 + H2O   ^   CO+3H2
                                                                                   (5)
                                                 and
                                                    CH4 + CO2
                                                                    2CO + 2H2.    (6)
                                             As the reaction rate is  relatively low, the
                                             methane content is  higher  than would  be
                                             expected from equilibrium.

                                                During the soaking phase a portion of the
                                             carbon also disappears according to the reac-
                                             tions:
                                                    C + CO
                                                                 2CO
                                                 and
                                                    C + HO  -  CO + H2.
(7)

(8)
                                              However, some carbon is always present in the
                                              product  gas from  the reactor in  a  quantity
                                              equivalent to about 3 weight percent of the oil
                                              feed.

                                                The composition of the fuel  gas is deter-
                                              mined by  the water-gas  shift equilibrium
                                              which appears to freeze after the gas enters the
                                              waste heat boiler at an equilibrium tempera-
                                              ture about  2200 to 2400F.
                                                   CO + HO ^ CO2 + H2
(9)
                                             DESCRIPTION  OF  SHELL  GASIFICA-
                                             TION PROCESS

                                                A simplified SGP flow sheet is given in
                                             Figure 1.  The hydrocarbon  charge  and the
                                             oxidant are preheated and fed to the reactor.
                                                                                 IV-4-3

-------
 The hot reactor-effluent gas containing about
 3 percent of the feed as soot is  passed  to  a
 waste heat boiler, producing  high pressure
 saturated steam. High heat transfer rates are
 achieved, with the result that the  temperature
 of the gas leaving the waste heat boiler closely
 approaches that of the steam produced in the
 boiler. The design and construction  of the
 waste heat boiler are such that  the surface
 remains clean for an indefinite period (without
 using any external cleaning devices); it may be
 noted that the waste heat boiler  of the Shell
 prototype unit has  been in  operation since
 1956 and never has been cleaned on the gas
 side. The waste heat boiler can be  designed for
 steam pressures  up to  about 1500 psig.

   The crude gas leaving the waste heat boiler
 at temperatures around 350F is  then  passed
 to the carbon removal system, consisting of  a
 bulk  removal  of the carbon  by  a special
 method of contacting the gas with water, and a
 final water wash. The product gas is virtually
 free of carbon (< 5 ppm).

   The carbon produced in the gasification is
 recovered  as a  soot-in-water  slurry (carbon
 content 1 to 2 weight percent). In most cases, it
 will  not be possible to dispose of this carbon
 slurry as such. Therefore, a special technique
 has  been developed for removing the carbon
 from the slurry, resulting in carbon-free water
 for re-use. Depending upon the metals content
 of the feedstock and  the  economics  and
 maintenance policy of the process operator, up
 to 100 percent of the soot can be recycled to
 extinction with the fresh feed.

   Sulfur  in  the  feedstock  is  converted
 primarily to H2S and  traces of  COS.  The
 carbon-free product gas is treated in a Shell
 Sulfinol or ADIP process unit where the sulfur
 compounds and  most  of  the   CC>2   are
 absorbed.  The  desulfurized  gas   typically
contains less than 5 ppm of sulfur. The acid
gas effluent from the Sulfinol unit is fed to a
 Claus process unit which recovers elemental,
salable  sulfur.

   Depending  on  the desired  LHV  (Lower
Heating  Value)  in the product  gas,  either

IV-4-4
 oxygen or air (enriched or unenriched) may be
 used as the oxidant.  Nitrogen  present in the
 air  acts  as a  moderator for  temperature
 control in the reactor. When either oxygen or
 air  enriched  with oxygen is   used as  the
 oxidant, a certain quantity of steam must be
 injected into the  reactor for  temperature
 moderation.  Air oxidation produces a  low
 heating value (120 Btu/scf) fuel gas due to the
 presence  of  nitrogen,  while   oxygen  feed
 produces a medium  heating value  gas  (300
 Btu/scf). Typical product gas compositions for
 air and oxygen gasification are shown in Table
 1.

Tabte 1. TYPICAL PRODUCT GAS COMPOSITION

Hydrogen
Carbon monoxide
Methane
Nitrogen
Argon
Sulfur
Total
% vol, dry basis
Oxygen
oxidation
48.0
51.0
0.6
0.2
0.2
5 ppm
100.0
Air
oxidation
12.0
21.0
0.6
66.0
0.4
5 ppm
100.0
 COGAS CYCLE THERMODYNAMICS

   Although  the idea  of  combining  a gas
 turbine and steam turbine  is old, its applica-
 tion to power generation has been studied only
 lately. Wood3 has  presented  an  excellent
 summary of the development of the COGAS
 cycle. In a COGAS cycle  (Figure  2),  air  is
 compressed and heated by  burning fuel in  it.
 The hot gases are then expanded in  a gas
 turbine coupled to the air  compressor  and a
 generator. The gas turbine exhaust, still at a
 high  temperature,  is  used  to  raise  and
 superheat high pressure steam; it is also used
 as a heat source for deaeration and boiler feed
 water preheating. The steam generated  by the
 gas turbine exhaust is expanded in a  steam
 turbine to produce additional electric  power.
 Heat rejection occurs in the stack exhaust and
 in the condenser of the steam cycle. It  is well
 known that the greater the difference between

-------
the heat source and heat sink temperatures of
any heat engine, the higher its thermodynamic
efficiency. In the COGAS cycle, the heat sink
of the gas cycle becomes the heat source of the
steam  cycle, increasing the  overall  spread
between source and sink temperatures  for the
combined cycle. As a result, the COGAS cycle
has a higher thermodynamic  efficiency than
either the  simple gas cycle  or the steam cycle.
In the case presented in this paper, a COGAS
cycle efficiency (based on the net useful energy
input to the power plant) of 44 percent was ob-
tained. Simple gas cycle and  steam cycle
efficiencies are of the order of 25 percent and
37 percent,  respectively.

   In application of the  Shell Gasification
Process to the COGAS cycle,  the addition of
the steam generated by the heat of  partial
oxidation  to the  steam  generated  in  the
COGAS cycle largely compensates for the loss
of heating value of the oil caused by gasifica-
tion.

POWER  PLANT FLOW  SCHEME

   A  flow diagram of the  power plant using
desulfurized fuel gas as fuel is shown in Figure
3. Air is compressed to 14 atm and split into
two parts. One part is cooled by heat exchange
with 100 F product gas from the SGP section
and compressed again in a booster compressor
to 18 atm  before entering the SGP unit as the
oxidant. The other part of  the compressed air
is combusted to a temperature  of about
2200F in  the  gas  turbine  combustor  by
burning the sulfur-free  fuel gas supplied  by
the SGP/Sulfinol units,  and expanded to 1.5
atm  (absolute) pressure in the gas turbine
which  is   coupled  to  an   electric  power
generator. The gas turbine exhaust is cooled to
about    350 F   in   a   waste   heat
boiler/superheater  and boiler feed  water
deaerator  before being vented  to atmosphere.
Steam (1250 psig) generated in the SGP waste
heat   boiler  is combined  with the  steam
generated  in the power plant waste heat boiler,
and the combined steam  is superheated  to
1000F in the superheater section of the latter.
The superheated steam is expanded  typically
to about 4-in. Hg vacuum in a steam turbine
coupled to a second electric power generator.
From the point of view of startup and control,
it is  advisable to use separate  generators for
the gas and steam turbines. Approximately 56
percent  of the  total  power  generation  is
contributed by the gas turbine.

SGP-BASED  POWER  PLANT    CASE
STUDY

   Using  the   foregoing  technology,   an
economic study has been made  of an SGP-
based  power  plant of  200-MW  nominal
generating capacity. Air was used  as  the
oxidant  in the partial oxidation  step.  The
composition of the typical heavy residue used
as feedstock is stated in Table 2. The material
balance of various gas  streams in the power
plant appears in the inset of Figure  3.

Table 2. TYPICAL RESIDUAL FEED PROPERTIES

       Gravity, API, 60F     12.0
       Specific gravity, 60/60    0.986
Composition, wt %
Carbon
Hydrogen
Sulfur
Nitrogen
Oxygen
Ash
Total
Viscosity, centistokes
470F
212F
100F
Ash analysis:
Nickel
Vanadium
Sodium
Iron
Others
Total
86.00
10.73
2.65
0.30
0.30
0.02
100.00
4
55
800
ppm in feed
30
100
1
4
65
200
% of ash
15.0
50.0
0.5
2.0
32.5
100.0
   Turbine manufacturers have indicated  to
us that industrial gas turbines are currently
designed for a  compression ratio of 12 and
                                                                                 IV-4-5

-------
turbine inlet temperature of 1800F. However,
it is predicted that by 1975, these conditions
are likely to be upgraded  to 14 and 2200 F,
respectively.  In  this  evaluation  we  have
assumed 1975 technology. Other simplifying
assumptions made were:

  1. The gas  turbine can  accommodate  an
    increase  of  approximately   25 percent
    (mole)  in  the  gas  flow   across  the
    combustor as low-Btu gas is  injected into
    the combustor. (This has been shown to be
    feasible.)

  2. Polytropic  efficiencies  of  the various
    components are : compressor 90 percent,
    gas turbine 90 percent, steam turbine 75
    percent2.

  3. Approximately 2 percent of  boiler feed
    water is evaporated in the deaerator, and 5
    percent is rejected in the boiler feed water
    blowdown. Thus, fresh boiler feed water
    requirement is 7 percent of the total steam
    generation.

  4. Friction and generator losses, which are
    normally small, have been neglected.

RESULTS AND  DISCUSSION

   A power flow diagram for this case (Figures
4A and 4B) shows the factors contributing to
an overall station  efficiency of 38  percent.
Figure 4A refers to the SGP unit and 4B shows
the power flow through the power generation
unit. The downstream end of Figure 4A thus
matches with the  upstream end of Figure 4B.
About 87 percent  of the energy  input  to the
SGP unit (LHV basis) is available to the power
plant as input energy.

   Table 3 is a summary  of the capital and
manufacturing costs. The major uncertainties
in the  unit power cost in this estimate, the
power plant capital cost and the oil price, are
shown parametrically (Figure 5) based on 1972
U.S. East Coast costs.

   Estimated unit power costs are about 1 to 2
mills/kWhr  higher  than  current   rates,
depending on the location. However, the latter
costs do not include the cost of stack gas
scrubbing  or  other  alternatives  of treating
sulfur and NOX emissions. With the growing
shortage of natural  gas and  restrictions on
sulfur, NOX, and  particulate  emissions, the
power cost is certain to rise rapidly in the next
few  years. At such  time, SGP-based  power
plants will offer an  attractive  alternative to
conventional power plants.

   Some of  the  advantages  of the   Shell
Gasification   Process   relative   to   coal
gasification processes are:

 1. The SGP  unit flexibly  accepts  a  wide
   variety of  fuels  ranging  from  heavy
   residues (e.g., flasher pitch) to natural gas.
   Thus, a consistent and continuous quality
   of fuel  gas can be  generated despite
   variations in the  quality of fuel supply.
 2. The SGP-based  power  station  handles
   fluids,   avoiding  the   complex   solids
   handling andash disposal steps involved in
   coal gasification  units.
 3. Both SGP  and  coal gasification   units
   require  comparable installation times.
   The lag in coal availability may  lead to a
   longer  project  realization  time for coal
   gasification-power units.
   Relative to other means of clean  power
generation, such as  the conventional  power
plants with stack gas  clean-up units, the SGP-
based power  plant offers  the   following
advantages:
 1. The SGP unit,  followed by a Sulfinol and
    Claus unit, produces the most marketable
    quality of recovered  sulfur.
 2. The NOX concentration in SGP-based fuel
    gas is low because the C-N bond in the fuel
    is broken mostly into CO and N2. With
    careful  gas turbine  combustor  design
    there is little breaking of the N-N bond, to
    produce NOX.

   The unit evaluated in this paper is intended
for intermediate to base load applications.
IV-4-6

-------
                  TableS. POWER GENERATION COST
 Power produced (gross), MW'
 Power consumed, MW
 Net power output, MW
 Overall efficiency, %

 Capital costs, $ x 106
Fuel processing unit
Power generation unit
Total capital cost, $ x 106 b
Operating cost
Variable costs
Oil @ $x/bbl
Sulfur credit @ $10/ton
Catalysts and chemicals
Cooling and boiler feed water
$x106/yr
2.61x
(0.11)
0.10
0.69
Total 0.68 + 2.61 x
Fixed costs
Operating labor @ $83,500/job
(4 operators)
Maintenance @ 3% of capital
Local overhead @ 100% labor
plus 25% maintenance
Taxes and insurance @ 1 %
of capital
Total
Net operating cost 3
Capital charges @ 14%
0.34
1.45
0.70
0.50
2.99
67 + 2.61 x
6.94
Total power cost 10.61 + 2.61x
18.2
31.4
49.6
Mills/kWhr
1.52x
(0.06)
0.06
0.40
0.40+ 1.52x
0.20
0.85
0.41
0.29
1.75
2.15+ 1.52x
4.06
6.21 + 1.52x
200.0
  4.7
195.3
 38.0
a Yearly average value. Actual capacity is 11% higher to compensate for
b  90% stream factor.
C1972, U.S. East Coast dollars.
 C.W.at1.5
   ...
 Cond. at 7% of circulation @ $0.50/10 3 gal.
                  Basis: 200 MW Nominal Generating Capacity
                            90% Stream Factor
                                                                          IV-4-7

-------
CONCLUSIONS

   The   Shell  Gasification   Process   in
conjunction with a COGAS power generation
unit offers an  attractive  combination  for
generating  electric  power from  high-sulfur,
heavy residual fuels. At the same time, sulfur
and NOX  emissions are  almost  completely
eliminated, and the sulfur is recovered  as a
salable   byproduct.  The  process  flexibly
accepts wide and frequent variations in the
feed quality  and  composition ranging from
natural gas to flasher  pitch. Effluents  are
minimized. Efficient  heat recovery in  the
gasification   unit   combined   with    the
advantages of the COGAS cycle in the power
generation  unit  ensures very little  loss  of
energy despite the additional fuel  processing.
High stream factors  render  this  process
 suitable for base load applications. In view of
 these merits, the SGP-based power station has
 promise for use in the electric utility industry.


REFERENCES

 1.  Business Week, March 11,  1972,  p. 44C.

 2.  Robson, R. L.,  et  al. Technological and
    Economic  Feasibility of Advanced Power
    Cycles  and Methods of Producing Non-
    polluting Fuels for Utility Power Stations.
    United Aircraft Research  Laboratories,
    East   Hartford,   Connecticut.   Report
    Number J-970855-13, p. 258.

 3.  Wood,  B.  Combined  Cycles:  A  General
    Review of Achievements.  Combustion.
    April 1972, p. 12.
IV-4-8

-------
     STEAM
   PREHEATERS
HIGH PRESSURE STEAM
OXYGEN
 OR AIR
               REACTOR
                          WASTE
                          HEAT
                         BOILER
                   CARBON
                   SLURRY
                  SEPARATOR
                                                            FUEL GAS TO
                                                            SULFINOL UNIT
                                                   FRESH
                                                   WATER
                                                                     CARBON-FREE
                                                                     CIRCULATION
                         BOILER
                         FEED
                         WATER
                 BOILER FEED STOCK
                                                            WASTE
                                                            WATER
                       Figure 1.  SGP for fuel gas manufacture.
                                                                                IV-4-9

-------
                           *  FUEL
                         COOLING WATER
                         Figure 2.  COGAS cycle for power generation.
IV-4-10

-------
     414 F
                                                                                                                 1250 psig
                                                                                                                 STEAM
                                                                                                               FROM WHB
                                                                                                               OF SGP UNIT
                                                                           BAROMETRIC
                                                                           CONDENSER
                                               WASTE
                                               HEAT
                                               BOILER
                                               (UNFIRED)
                                                                                             CONDENSATE
                                                                                               PUMP
                                                                                  1250 psig, 212 F
                                                                                                      350 F
STEAM
TURBINE
DRIVE
-*-  AIR TO REACTOR
     470 F, 18 aim
                                                                                                    PREHEATED
                                                                                                    BFW TO WHB
                                                                                                    OF SGP  UNIT
COMPONENT
HYDROGEN
CARBON MONOXIDE
CARBON DIOXIDE
METHANE
NITROGEN
OXYGEN
WATER
MOL
WT
2.0
28.0
44.0
16.0
28.0
32.0
18.0
B.P.







SP.
GR.







TOTAL
PRESSURE AND TEMPERATURE,0 F
scfm_x 103
ALL NUMBERS IN Ib-mole hr
CASE

4,158
6,490
397
85
17,230


28,540
116.
<3>




17,485
4,647

22,150





64,412
17,112

81,534
14atm,780F
179
140
516
j Btu/scf
<4>


6,972

81,642
11,668
4,308
104,560
2,200F
661
                            Figure 3 .   Power plant flow diagram.

-------
        AIR FROM
        POWER UNIT
                                     OIL
                                    ENTHALPY
                                     PLUS
                                    HEATING
                                     VALUE
                             CONDENSATE
                          OIL
                          PRE-
                          HEAT

-^:
M
I
GAS ENTHALPY
10 MW

H
-s;
525 MW
1786x106
Btu/hr
yr
SGP REACTOR
GAS
ENTHALPY
575 MW
\^^-^
y*
3 WASTE HEAT
g BOILER
o
fe
UJ
3=
fe
s -^
(

^
GAS
ENTHALPY
446 MW
^^ ^*
SOOT AND SULFl
RECOVERY
GAS
ENTHALPY
366 MW
^-^ ^^
;=
IR


?*
ft 15

\JWAT
ENT
WATE


MW


ER
HALPY
IR ENTH


OIL
PREHEATER
f.
LT>
ftLPY



CD
r-
"xT

cc
UJ "=*"
g =
i
o
o
0
l_, ^.J

*
Q.

-------
GAS ENTHALPY
   25 m
AIR TO SGP
  AIR
    mm
FUEL FROM SGP
WATER ENTHALPY
    114 m
STEAM FROM SGP
COWBUSTOR

GAS
ENTHALPY
IRQ MW

                                   CONDENSER
                                                                       GAS ENTHALPY
                                                                         50 MW
                                                                       STACK GAS






WASTE HEAT
BOILER

WATER
ENTHALPY
292 MW





T|
I* WATER ENTH
J1 15MW
GAS ENTHALPYj
65 MW n
                                                              DEAREATOR
                                                  CONDENSATE
                                    COOLING
                                     WATER
                                      LOAD
                                     205 MW
                                COOLING WATER

                   Figure 4B,  Power flow through power plant.
                                                                               IV-4-13

-------
       20
             PARAMETER: TOTAL CAPITAL INVESTMENT, $
                                                            5
2            3             4
                     OIL PRICE, $/bbl
  Figure 5.  Economics of power generation using SGP.
IV-4-14

-------
      5. FLUIDIZED-BED OIL  GASIFICATION FOR  CLEAN
                POWER  GENERATION-ATMOSPHERIC AND
                                         PRESSURIZED OPERATION
                R. A. NEWBY, D. L. KEAIRNS, E. J. VIDT,
                    D. H. ARCHER, AND N. E. WEEKS
                   Westinghouse Research Laboratories
ABSTRACT

   This paper evaluates high sulfur residual oil gasification for the purpose of clean power
 generation. It also considers both atmospheric pressure operation with conventional boilers and
pressurized operation utilizing combines operation with conventional boilers and pressurized
operation utilizing combined steam and gas turbine cycles.

   The outlook and status of atmospheric pressure fluidized-bed oil gasification is reviewed. The
process has been studied by Esso (England) on a 1-MW pilot plant scale and has been shown to be
an  effective pollution  control  device.  Preliminary cost  estimates  for  retrofit  systems  on
conventional plants  indicate a potential energy cost reduction of 30 to 50 percent  over wet
scrubbing or low sulfur fuel alternatives. A 30 to 100-MW demonstration plant is scheduled.

   A pressurized fluid-bed oil gasification process has been proposed, and preliminary assessment
is being carried out in which plant performance, capital costs, and energy costs are examined. The
projected ability of the fluid-bed process to meet both pollution regulations and gas turbine
requirements is based on Esso (England) atmospheric pressure data. Capital and energy costs for
the pressurized fluid-bed combined-cycle (PACE) plant process are compared with a conventional
oil-fired plant with wet scrubbing, pressurized fluid-bed combustion of oil, PACE plant operation
with No. 2 distillate fuel oil, and conventional pressurized oil gasification processes developed by
Shell and Texaco. The fluid-bed process has the potential to reduce energy costs > 20 percent
below a conventional plant with wet scrubbing  and  > 15 percent below a PACE plant using
distillate fuel oil.

INTRODUCTION

   Electric utility demand for residual oil is    electric power industry as the availability and
projected to increase by around 250 percent by   cost of natural gas and clean fuel oils de-
1980. Low sulfur residual to meet Federal and    creases  and  because the   technology for
local regulations will not meet the demand   removing sulfur dioxide from stack gases is
without an intense effort on desulfurization.    proving to be expensive. A gasification system
                                          must  provide a clean gas,  reliability, util-
   Gas producers for making  low Btu gas   ization of a wide range of  fuel, competitive
from oil are becoming more attractive to the   energy cost, and high efficiency.

                                     IV-5-1

-------
    Fluidized-bed  oil   gasification  can   be
  applied  to  power  generation  to  produce  a
  clean, low-Btu fuel gas200 to 500 Btu/scf.
  In a  fluidized-bed  gasifier, oil is added to a
  fluidized-bed  of limestone  or  dolomite  with
  sufficient air15  to  25 percent of stoich-
  iometricto maintain the bed at ~ 1600F
  and  react with the  oil to produce a fuel gas.
  The  limestone  or  dolomite removes sulfur
  from the fuel  gas during the  gasification
  process.

    Gasification  of oil  can be carried out  at
  either atmospheric  or elevated presure. Oper-
  ating at atmospheric pressure, a fluidized-bed
  gasifier  provides clean  fuel to a  combined
cycle gas and steam turbine power plant. The
potential for high efficiency and  low capital
cost  of  such  a  plant  makes  this  system
attractive.

   It has been demonstrated that oil can be
gasified and sulfur removed from the resulting
fuel gases in a fluidized bed. The design and
evaluation of a fluidized-bed oil gasifier oper-
ating at atmospheric pressure have been com-
pleted. A 30 to 100-MW demonstration plant
is  scheduled. A  pressurized fluidized-bed oil
gasification  system has  been designed. The
economics and performance  of a  pressurized
oil  gasification combined  cycle power plant
have been evaluated.
IV-5-2

-------
ATMOSPHERIC  PRESSURE  FLUID-BED
OIL GASIFICATION

   Westinghouse    is    evaluating     the
gasification/desulfurization of residual oil at
atmospheric pressure  under contract  to  the
Office of Research and Monitoring* (ORM) of
the  Environmental  Protection  Agency. The
concept is being studied for  the  purpose of
producing,  on-site,  a  low-sulfur fuel  gas
suitable for power plant utilization in  a con-
ventional boiler. Esso Petroleum  Company has
provided experimental information on  the oil
gasification/desulfurization process based on
small  scale batch  fluidized units  and  a 750-
kW  continuous  unit.   Westinghouse  has
carried out preliminary conceptual  design
studies to evaluate  the  commercial process,
and is presently attempting to locate a utility
partner with  which to carry out a demon-
stration plant operation.

Gasification/Desulfurization Concepts
   Two  possible  modes  for  gasification/
desulfurization operation are the regenerative
mode  and the once-through mode. Figure  1
illustrates  the  major  process  streams and
identifies the  basic  elements  of the  two
operational modes.
   The elements of the regenerative operation
are  the  gasifier vessel and the regenerator
vessel. The  gasifier is an air-fluidized  bed of
lime operated  at  1600F  with  sub-stoichio-
metric air, ~ 20 percent of stoichiometric. The
oil is injected into  the  gasifier vessel where it
cracks and  is partially combusted to form  a
hot, low-sulfur fuel  gas. Hydrogen sulfide is
produced during the gasification which reacts
with the lime to produce a sulfided lime.
     H2S + CaO  - CaS + H20
(1)
The fuel gas is transported to  the  boiler
burners where combustion is completed and
the sulfided lime is  sent to the regenerator.
The regenerator is  an  air-fluidized  vessel
operated with a slight excess of air at 1900F.
*Previously under the Office of Air Programs.
 Regeneration takes  place  by  reaction  of
 oxygen with the utilized lime to give an SO 2
 rich stream (of about 10 mole percent SO2)
 and  a regenerated  lime  having  a slightly
 decreased  activity compared to that of fresh
 lime.
            3

 The  SO 2 stream  is  transported to a  sulfur
 recovery system and  the regenerated lime is
 returned to the gasifier along with a stoichio-
 metric amount of fresh make-up  limestone.

   The  elements   of   the  once-through
 operation  shown  in  Figure  1  consist of a
 gasifier vessel and a sulfate generator, or pre-
 disposal vessel. The operation of the gasifier
 for once-through  limestone operation is  the
 same  as for the regenerative operation. The
 sulfate generator  operates similarly to  the
 regenerator,  but  at  a  lower  temperature
 ( ~ 1500F) so that the sulfided lime from the
 gasifier is converted to calcium sulfate rather
 than  calcium oxide. The dry  calcium sulfate
 may be disposed of while the gas stream from
 the sulfate generator  is sent to the gasifier. A
 limestone addition rate of 3 to  4 times that
 used in the regenerative operation is necessary
 in  the once-through operation  to  achieve
 similar sulfur removal of 90 to 95 percent.

Experimental Work

   Under contract to  the  Office of Research
and  Monitoring, Esso (England) is carrying
out laboratory tests on  atmospheric-pressure
batch fluidized-bed equipment to investigate
lime sulfur absorption and lime regeneration
operating variables.1'2 The results of the Esso
Petroleum  Company  experimental  program
have   identified   the  critical   phenomena
associated  with  atmospheric-pressure  gasi-
fication/desulfurization, and have established
the   probable operating   conditions   and
behavior of  a  commercial  system.  This
evaluation has been based on the results of the
small scale batch work. The  batch work has
been carried out in conjunction with the con-
struction and operation of a 750-kW con-
tinuous pilot unit.
                                                                                   IV-5-3

-------
 Design

   Energy and material balances provide the
 basic information with which the feasibility of
 applying the gasification/desulfurization con-
 cept as an add-on to an existing boiler, or as a
 new plant design feature, has been examined.
 The feasibility of the retrofit concept has been
 examined  in terms of the availability of space
 in an existing power plant, the modifications
 necessary to retrofit an existing boiler, and the
 performance of a modified boiler. The specifi-
 cations assumed for the conceptual design are
 presented  in  Table 1.

   The feasibility  of  converting  an existing
 boiler to one which utilizes the fuel  from a
 gasification/desulfurization  system  depends
 on a number of factors, many of which will
 differ from one boiler to the next. The  space
 available  for  a  gasification/desulfurization
 system in an existing power plant, the modifi-
 cation necessary to retrofit an existing boiler,
 and the performance  of a retrofit boiler will
 depend on the specific gasification/desulfur-
 ization system design and choice of operating
 conditions,  the   location   of  the   gasifi-
 cation/desulfurization system in the plant, the
type of boiler (coal-, oil-, or gas-fired), size of
boiler, the turn-down needed, the boiler load
factor, and the specific design features of the
boiler.

   The gasification system may  be  placed
internal  to the boiler (directly beneath)  or
external to the boiler. Preliminary considera-
tions seem to favor the external retrofit design
over the  internal  retrofit design,  and  the
feasibility  of  retrofitting  a coal-fired boiler
over the feasibility of retrofitting an oil- or gas-
fired boiler.  The internal  design   concept
appears to be limited  by the available space
beneath the boiler and the cost for modifying
the boiler. The external design concept should
reduce modification costs and should allow
greater uniformity and  flexibility  in system
design. Coal-fired boilers have more potential
space near the  boiler than oil-  or gas-fired
boilers and may provide some of the needed
solids  handling  and  particulate  clean-up
equipment. Boiler  derate  should be of little
concern with  coal-fired  boilers because they-
are designed to operate with slag on the boiler
heat transfer surface. Derate may be a concern
with oil-   and  gas-fired  boilers  due to  a
reduction  in  the  flame temperature and a
                  Tabte 1. SPECIFICATIONS FOR CONCEPTUAL DESIGN
                  Process specifications
              Design variables
            Sulfur removal: 90-95%
            Fuel oil: 3 wt % sulfur; LHV = 17,700 Btu/lb
            Limestone: 98.6% CaO yield
            Boiler size: 600 MW
            Load factor: 40%, 80%
            Turn-down: 4/1
Operating variables

Gasifier temperature
Regenerator (sulfate
generator) temperature
Stone make-up rate
Air/fuel ratio
Limestone utilization
Fluidization velocity
Minimum fluidization
velocity
Particle sizes (avg)
Gasifier bed depth
(static)
Regenerative system
1600F

1900F
1 mole CaO/mole sulfur
20% of stoichiometric
~ 5 wt % sulfur in bed
8 ft/sec

3 ft/sec
~2000 Mm

2.5-3.5 ft
Once-through system
1600F

1500F
3 moles CaO/mole sulfur
20% of stoichiometric
~ 19 wt% sulfur in bed
8 ft/sec

1 ft/sec
~ 1000 pirn

3.5-4.0 ft
IV-5-4

-------
redistribution of the heat  release within the
boiler which may  occur with  the fuel.

   In contrast to boiler retrofit considerations,
the  feasibility of  incorporating  a  gasifica-
tion/desulfurization system into a new boiler
design will  be limited only  by  the overall
economics of the  system and  the  market
potential for new  oil-fired boilers. The  total
space occupied by the gasification/desulfuri-
zation system will be a small percentage of the
total plant   volume.  Also, because  of the
flexibility in boiler design, boiler performance
will not be affected by the presence of the gasi-
fication/desulfurization system.

   Figure 2  shows a plant layout for a  600-
MW  once-through    gasification/desulfur-
ization system. The plant consists of two  gasi-
fication modules, each utilizing air from one
of the two   power plant  forced draft  fans
present in the existing boiler.
 Evaluation

   The  design  study,  coupled  with  the
 experimental   work   of  Esso  Petroleum
 Company,   points  out  the  technical  and
 economic feasibility of oil  gasification/desul-
 furization as a retrofit  SO 2 control system for
 utility boilers, or as an SC>2 control system to
 be incorporated into a new boiler  design. A
 market for  retrofit and  new oil-fired  boilers
 exists.3

   Preliminary  investigations  indicate that,
 overall, the once-through  operation may be
 somewhat more attractive to a utility customer
 than   the  regenerative  operation.  Capital
 investment is reduced  with the once-through
 operation and,  although the limestone feed
 rate is expected to be three  times the rate with
 regenerative operation, the  operating costs for
 once-through operation may be less than those
for the regenerative operation.  A complete
cost breakdown for once-through and regen-
erative operation has been  presented.4 Once-
through operation has fewer technical prob-
lems at this time, and is an overall  simpler
process than regenerative operation.
   A comparison is made in Table 2 between
atmospheric pressure oil  gasification/desul-
furization and the alternative schemes of low-
sulfur oil and stack gas cleaning. Capital costs
and operating costs are compared for new and
retrofit systems. Oil gasification/desulfuriza-
tion compares favorably with low-sulfur  oil
and stack gas cleaning, based on prelimimary
cost estimates.  A reduction of about 40 per-
cent in the capital costs involved in stack gas
cleaning is estimated for  new and  retrofit
gasification/desulfurization systems.   Oper-
ating costs appear to be about the same for
once-through stack gas  cleaning  and regen-
erative gasification/desulfurization with sulfur
recovery. Once-through gasification/desulfur-
ization may reduce the operating costs 30 to 50
percent as compared to stack gas  cleaning.
The cost estimate indicates that the operating
cost with low-sulfur fuel oil will be about 30 to
50 percent greater than the operating cost for
gasification/desulfurization.  These  conclu-
sions are based on the desulfurization of high-
sulfur  residual oil (3 weight  percent  sulfur)
and may be altered when a lower sulfur oil is
considered (1 to 1.5  weight percent sulfur).
Environmental factors are also compared in
Table 2. The low-sulfur oil  is advantageous in
that capital costs are limited to possible boiler
modifications necessary when  changing from
gas or coal to low-sulfur  oil.  On the other
hand, operating costs are higher than those for
stack gas cleaning or oil gasification/desulfur-
ization. Capital costs are extremely high with
stack  gas  cleaning, especially on the  retrofit
case, while operating costs are very near those
estimated  for oil gasification/desulfurization.

   Advantages  of atmospheric-pressure oil
gasification over  stack  gas  wet  scrubbers
include:

    Corrosion and fouling problems mini-
    mized   in SO 2 removal process  and
    boiler  (minimum SOx  and V),

    No flue gas reheat required,
    Uses crushed limestone  no lime-
    stone pulverizing system needed,
                                                                                    IV-5-5

-------
       Table 2. ASSESSMENT OF FLUIDIZED BED OIL GASIFICATION/DESULFURIZATIOIM

Cost
Capital, $/kW
New
Retrofit
Fuel adder, 
-------
Phase II.  Detail design  and construction of
          demonstration oil gasification pro-
          cess.

Phase III. Developmental operation of the gas-
          ification  process  and  integrated
          power plant.

   A utility (or utilities) is  required  as a
 cooperating party to carry out this  program.
 In Phase I the utility would supply technical
 information  on its existing  power  plant,
 provide price and availability information on
 oil fuels which might be utilized in  the plant
 currently  and/or   in   the  future,   supply
 information concerning the load requirements
 placed on the plant, cooperate in the selection
 of  an   engineering  firm  to  prepare  the
 preliminary design, assist in selecting various
 options  in  the  design  of the  system  and
 evaluate, in  cooperation with  Westinghouse,
 Esso (England) and EPA, the effectiveness and
 economy of the oil gasifier/desulfurizer in
 power generation and   pollution  abatement.

   An engineering firm will carry out  designs
 in sufficient detail that fixed price bids can be
 solicited for detailed design and construction
 of the system. The design effort will be based
 on experimental data from Esso (England) on
 their 1-MW continuous pilot  plant and  the
 conceptual   design   and   assessment  by
 Westinghouse.  EPA will  provide  general
 guidance and funding for Phase  I.

   If the preliminary   design  confirms  the
 effectiveness and economics of the  system, a
 proposal would be prepared for Phases II and
 III.  In Phase II the utility would work closely
 with the engineering firm in the design and in-
 stallation of the gasifier/desulfurizer system
 and share in the  costs of the design  and
 installation.

   In Phase  HI the utility would  operate  the
 plant, collect basic data on the operation of
 the  gasifier/desulfurizer system,  aid  in  its
 interpretation and  analysis, and cooperate in
 the  evaluation of the effectiveness,  technical
 and  economic,  of  the process  in   power
 generation and  pollution abatement.
   Presentations  have  been  made  to  13
utilities  who  currently  operate or  plan to
operate oil-fired power plants. Presentations
began in  March 1972  and  have included
utilities  from the  East  Coast, West Coast,
Florida and the South Central U.S. A West
Coast utility is actively interested in the pro-
gram. Seven utilities are currently evaluating
the proposal, and five utilities  have indicated
they are not interested in participating in the
program at this time.

   Problem areas  for  which  utilities  have
expressed concern are:

   1.  Hot fuel gas piping and  valves.

   2.  Control and emergency conditions.

   3.  Space requirements.

   4.  Solid waste disposal.

   5.  Modifications  and  time  required  for
     modifications to boiler.

   6. Availability.

These and similar technical  concerns will be
evaluated in detail  during  the preliminary
design phase. In addition,  the prospect of
obtaining funds from EEI  (Edison  Electric
Institute) has been pointed out.

Conclusions

Performance

The  concept has been  technically  demon-
strated  with  a  750-kW  development gasifi-
cation plant.

Sulfur removal  up  to  95  percent  can be
achieved.

Nitrogen oxide emissions of  150 ppm appear
possible.

Particulate  emissions will  be higher  than
conventional oil- or gas-fired systems but can
easily  be  removed  to  achieve   proposed
standards.

Further development effort is required in key
arease.g., calcium sulfate generation for
                                                                                     IV-5-7

-------
 once-through operation, temperature control,
 sulfur recovery.

 Economics for  comparable  pollution abate-
 ment
 Capital cost of a retrofit, once-through gasifi-
 cation system may be  50  to 70 percent less
 than a retrofit  wet scrubbing system.

 Fuel adder  cost for  a  retrofit,  once-through
 gasification  system may be  30 to 50 percent
 less than  wet  scrubbing,  low-sulfur oil, or
 desulfurized oil.

 Market

 Initial market is expected  to be small boilers
 ( < 600  MW)  on  the  East Coast, in  the
 Southwest where gas may be  limited, or on the
 West Coast. Once-through operation may be
 favored over operation with sulfur recovery for
 these plants.
 PRESSURIZED
 GASIFICATION
FLUID-BED
OIL
    Two different processes are  being con-
 sidered for the pressurized gasification of
 residual oil: (1) a pressurized  version  of the
 fluidized  bed oil gasification/desulfurization
 process, being developed by Esso (England),
 which has been explored only at atmospheric
 pressure, and (2) the pressurized oil gasifica-
 tion  process  of the type  which  has been
 operated for gas manufacture  by Shell5 and
 Texaco.6 The Shell and Texaco processes are
 not identical, but they are very similar in con-
 cept, performance and cost. The process con-
 cept is the important factor  for this study so
 the Shell and Texaco processes are not dis-
 cussed  individually. The Shell and  Texaco
 processes generate  a low temperature (100-
 250F), clean fuel gas having a low  heating
value of about 120 Btu/scf. Steam generated
by waste heat boilers in the  gasification pro-
cess is also  provided for the combined cycle
plant. The fluidized-bed process generates  a
hot (~1600F),  clean fuel gas  having a low
heating value of 200 to 500 Btu/scf (hot). Pre-
liminary cost and performance  estimates for
these  two   process   concepts   have  been
developed and are compared with alternative
oil-fueled power generation techniques having
pollution control.

Process  Concepts and Options

   Flow  diagrams for the Shell and Texaco
processes  and  the  pressurized   fluid-bed
process  are  shown  in  Figures  3  and  4,
respectively. The Shell and  Texaco  processes
consist of an  air-blown oil gasification  vessel
(partial  oxidation   reactor)  operated  at  a
temperature of about 2500F with an air/fuel
ratio of  about 45 percent of stoichiometric.
The hot  gas is cooled in waste heat boilers,
producing saturated steam,  prior to purifica-
tion of the gas. The gas purification process
consists of a carbon (soot) recovery step  and a
sulfur removal section. Recovery  soot  ( ~ 3
weight percent  of the fuel oil  feed rate) is
recycled  to  the  gasifier vessel,  and   H2S
produced in the sulfur removal section is sent
to a sulfur recovery plant to recover elemental
sulfur.

   The pressurized fluid-bed oil gasification
process  shown  in  Figure  4 consists  of  a
fluidized bed gasifier/desulfurizer  vessel,  a
limestone/dolomite  regeneration section, and
a  sulfur recovery section.  The  phenomena
taking place  in the pressurized fluid bed
gasifier are essentially the same  as have been
described for  the atmospheric pressure fluid
bed  case. The gasifier is operated  at  about
1600F with an air/fuel ratio of about  14-25
percent of stoichiometric. Both limestone and
dolomite are considered as sulfur absorbents
in the pressurized fluid-bed case.  The  major
process options which have been  examined
with respect to cost and performance are:

   1. The  gasifier   air/fuel  ratio.  Air/fuel
ratios  of  14 percent  and  25  percent  of
stoichiometric  giving fuel gas  low heating
values (hot) of about 500 and  280 Btu/scf,
respectively, have been considered. An air/fuel
ratio of  14  percent of  stoichiometric  is
assumed to be the minimum air/fuel ratio at
which a gasifier temperature of 1600F can be
maintained, while 25 percent of stoichiometric
 IV-5-8

-------
is assumed to be a conservatively high air/fuel
ratio according to the experience gained from
the  Esso  (England)  atmospheric  pressure
operations.   The   physical   feasibility   of
operating  at  an air/fuel  ratio  as  low as  14
percent  of stoichiometric without  excessive
carbon deposition  in the  gasifier  must  be
demonstrated.

   2. The  limestone/dolomite  regeneration
method. The regeneration of calcium  suflide
and the production of a high  sulfur gas  for
 sulfur recovery  can be achieved by either of
two processes:
H2O
to
    Regeneration  with   CO2  and
 produce H2S. With this process the reaction
                     -  CaCO+  H2S    (3)
 is utilized to produce calcium carbonate and
 H2S-rich gas. The reaction is favored by high
 pressure  and  would be carried  out  at  a
 temperature  of  about  1100F,  with H2S
 concentrations of about 9 percent by volume
 being projected. Two sources of CO2 for  the
 regeneration scheme  have been considered--
 CO2 provided by scrubbing flue gas  from the
 combined cycle plant, and  CO2 provided by
 scrubbing  a   gas   stream   produced   by
 combustion of  a portion of the fuel gas.

    Regeneration with air to produce SO2 . The
 reaction of oxygen  with calcium sulfide,
        CaS + 3/2O2  -  CaO + SO;
(4)
 is  favored   by  low   pressure   and   high
 temperature (~2100F). Although this regen-
 eration   scheme  would   yield   an   SO 2
 concentration of only  about  2  percent  by
 volume  at  pressures  of  10-15  atm,  it  is
 considered because of its apparent simplicity.

    If a high stone make-up rate is required for
 the regenerative processes and if high stone
 utilization can  be achieved in the gasifier, a
 once-through system may be attractive. Once-
 through operation would require conversion of
 calcium  sulfide  to  calcium  sulfate  before
disposal of the stone. The reaction

       CaS + 2O2 ^ CaSO4             (5)

could be applied  for this purpose, and would
be carried out at  1400-1700F. Limestone
utilization  in  a once-through   process  is
expected to be 35 percent or higher.

   Other process  options examined are  the
pressure drop required across the gas turbine
combustor and the option of cooling the  gas
produced  by  fluidized-bed  oil  gasification
before it is combusted in the combined cycle
plant.

Process Specification and Design Basis

   Table 3 lists the factors specified for the
conceptual design of the pressurized fluid-bed
oil gasification process. The specifications for
the  plant  capacity, capacity factors,  and the
turndown ratio are also assumed for the Shell
and Texaco  processes.  The  combined cycle
power  generating plant is a  Westinghouse
PACE Plant consisting of two Westinghouse-
501B gas turbines and a single 28.5-in. steam
turbine. The pressure of the fuel gas to the gas
turbine is 215 psia. Designs and energy costs
 are based on operation using limestone for de-
 sulfurization. Sulfur removal of 90 to 95  per-
 cent from a 3 weight percent sulfur residual oil

Table  3.  SPECIFICATIONS  FOR  FLUID-BED
                 OPERATION
        Plant electrical capacity
        Capacity factor
        Turndown ratio
        Number of gasifier modules
        Modes of operation
        Gasifier pressure
        Pressure of gas turbine
        Sulfur removal
        Residual oil
        Gasifier temperature
        Regenerator temperature
        Suifate generator terriperature
        Air/fuel ratios
        Lime particle diameters
        Regenerative lime utilization
        Once-through time utilization
        Limestone make-up rate
        Gasifier temperature control

        Regenerator temperature control
        Sulfate generator temperature
        control
        Plant heat rate
                                   250 MW (PACE Plant)
                                   70%
                                   4/1
                                   2
                                   Regenerative or once-through
                                   ~15 atm
                                   11 atm 1165 psia to turbine)
                                   90-95%
                                   3 wt % sulfur
                                   1600F
                                   1100Fa
                                   1700F with once-through option
                                   14 (minimum} Er 25% of stoichiometric
                                   500-2000 (im average diameter
                                   10%
                                   35%
                                   1 mole CaO/mole sulfur fed
                                   Stack gas recycle, steam or water
                                   injection
                                   Lime circulation rate; water injection
                                   Excess air circulation {-v 200%
                                   excess)
                                   9000 Btu/kWhr {assumed for purposes of
                                   material balances)
        With Co2/H2O regeneration; 2100F .with air regeneration.
                                                                                         IV-5-9

-------
 is specified for the fluid-bed  process, with
 specifications  for  vessel temperatures,  lime-
 stone utilizations, and limestone make-up rate
 based both on thermodynamic  information
 and the atmospheric pressure data of Esso
 (England). Control of the vessel temperatures
 is assumed to be easily carried out by any of
 the methods suggested in the table, based on
 atmospheric information. A plant heat rate of
 9000 Btu/kWhr was assumed for the purpose
 of finding the approximate fuel  consumption
 of a PACE Plant.

   Shell and Texaco both supplied energy and
 material   balance  information  for  their
 processes  along  with  capital   investment
 estimates. The Shell and Texaco processes had
 not necessarily been modified to provide  the
 optimum power generation performance, but
 were based  mainly  on gas  manufacturing
 experience.

   Material and energy  balance information
 for the pressurized  fluid-bed  oil gasification
 process  was   based  on   Esso   (England)
 atmospheric pressure data. Performance and
 vessel sizes  were  modified  for  effects   of
 pressure.  Regeneration system designs were
 based largely  on  thermodynamic behavior.
 The general behavior of the fluidized-bed gas-
 ifier for sulfur removal, vanadium removal,
 and carbon deposition  is assumed to  be
 independent of pressure once the gasifier bed
 diameter and  bed  depth has been scaled  for
 the affect of pressure.

 Material and  Energy Balances

   Figures 5 and  6  show  simple  block flow
 diagrams  for the Shell and Texaco processes
 and the pressurized  fluid-bed  oil gasification
 process, respectively. Block diagrams of this
type have been  utilized  to analyze the per-
formance of the gasification processes  and the
performance of the  complete power plant.

   Inputs  to the Shell and  Texaco processes,
Figure 5,  are  shown to  be residual oil, air,
water,  booster  compressor   power,   and
auxiliary power for pumping oil, boiler water,
cooling water,  scrubber recycle, etc. Output is

IV-5-10
cold fuel gas, steam, sulfur, and energy losses.
Energy  losses  for the  Shell  and  Texaco
processes arise from heat losses, cooling water
losses, and sensible heat of flue gas from gas
purification and sulfur recovery sections. The
thermal  efficiency  of the Shell and  Texaco
processes as presently conceived is less than 80
percent;  though  improvements  in this per-
formance factor  may be  made  by simple
process   alterations.   Shell  has  recently
indicated  87 percent  thermal  efficiency for
their  process.
   Inputs to the fluidized-bed process (Figure
6) are residual oil, air, water and steam, lime-
stone  (or dolomite), booster compressor power,
and auxiliary power for pumping oil, water,
solids circulation, and gas compression in the
regeneration section of the gasification system.
Output is hot fuel gas for the combined cycle
plant, sulfur, and energy losses in the  form of
sensible  heat of the  spent  limestone, heat
losses, carbon deposition losses, and sensible
heat of flue gas  from the regeneration and
sulfur  recovery    sections.   The  thermal
efficiency of the fluidized-bed  gasification
system will  depend slightly  on the  air/fuel
ratio  and the regeneration method used, but
will be  about 90 to  95  percent for  all the
options  considered.
   Fuel  compositions  and  heat  values  are
shown in Table 4 for  the two oil gasification
processes. Shell  provided expected  product

    Table 4. GASIFICATION PRODUCT
              COMPOSITIONS
s
N2
H2
CO
C02
H20
CH4
C2H4
H2S
Low heating
value (hot),
Btu/scf
hell process,
Vol%
60.79
14.57
22.92
1.37
0.00
0.35
0.00
0.00


117
Fluidized bed process,
Vol%
14A/F
50.74
0.82
13.40
6.70
0.00
9.43
18.86
0.05


~500
25A/F
47.74
2.64
7.97
7.97
20.56
6.55
6.55
0.02


~280

-------
compositions for a dry gas, while the product
gas compositions for the fluidized-bed process
at air/fuel ratios of 14 and  25 percent of
stoichiometric   have  been  estimated  from
atmospheric pressure data. Projected  plant
performance has been based on these product
compositions.

Capital   Investment  Evaluation   for   the
Pressurized Fluid-Bed Process

   Figure 7  is a  flow  diagram   for  the
pressurized fluid-bed oil gasification process
with limestone regeneration be reaction with
CO2 and H2O to form CaCO3  and H2S^ The
source of CO2 is flue gas from the combined
cycle plant.  The alternative CO2 source,
combustion of a portion of fuel gas, has been
eliminated  based   on  the   comparative
economics of the two options. The equipment
shown in Figure 7 consists of fuel oil and lime-
stone  handling   equipment,   air   booster
compressor, gasifier vessels  (2 modules) with
multi-stage   particle   collection,    and    a
CO2/H2O regeneration system.

   The process has been separated  into four
component systemsthe gasification system,
the  CO2/H2O  regeneration  system,   the
booster compressor, and  the  Claus plant.
These four  component costs  are  shown in
Figure 8, in units of S/106  Btu-hr of oil feed
(HHV),  as a function of the pressure at  the
gasifier product gas  outlet,  with the air/fuel
ratio as a parameter. Two cases of  pressure
drops required in the gas turbine combustor
and the  fuel  distribution equipment  are
considered in the figure8 psi  AP and 53 psi
 A P. The pressure drop across the gasifier  and
particle control equipment is assumed to be 9
percent of the gasifier  pressure.

   Figure 8 leads to the following conclusions:

 1. Reducing the air/fuel ratio from 25 to 14
    percent  of  stoichiometric  reduces  the
    capital investment by about 10 percent in
    units of $/106  Btu-hr.

 2. For  the  case  of  a  10  atm gas  turbine,
    increasing  the combustor  pressure drop
    from 8 psi to 53  psi  reduces  the process
    cost by about  10 percent because of the
    highly pressure  sensitive  nature  of the
    CO2/H2O regeneration  scheme.

 3. A once-through scheme would reduce the
    cost  of  the process by more  than  40
    percent, neglecting the slight cost increase
    due to a sulfate  generating vessel.

 4. The cost of the fluidized-bed process with
    an air-blown regenerator would  be some-
    where between the cost for the CO2/H2O
    regeneration system and the once-through
    gasification system.

   The true capital investment for pressurized
fluid-bed oil gasification will depend  on the
plant performance as a function of the design
options-air/fuel ratio,  combustor  pressure
drop, and regeneration  method.

Performance

   Cycle studies based on estimated material
and energy  balances and  approximate  gas
compositions (Table 4) have led to estimates of
the plant heat rate, the plant power,  capital
investment,    and    energy   costs,   using
pressurized  oil  gasification with combined
cycle  power  generation (PACE Plant).  The
following    conclusions   concerning    the
combustor pressure drop have been  deduced:

 1. With  the  pressurized   fluid-bed   oil
    gasification process an  increase  in  the
    required  combustor pressure  drop of 10
    psi increases the plant heat rate by only 10
    Btu/kWhr with operation at  an air/fuel
    ratio of 14 percent of stoichiometric,  and
    30 Btu/kWhr percent.-

 2. With the pressurized fluid-bed  oil gasifi-
    cation process an increase in the required
    combustor pressure drop of 10 psi reduces
    the  plant power  by  about 0.5 MW  with
    operation at 14 percent of stoichiometric,
    and about 1.0 MW at an air/fuel ratio of
    25 percent of stoichiometric.

 3. With the  Shell  and  Texaco processes,
    increasing   the   required   combustor
    pressure drop  10 psi increases  the plant

                                   IV-5-11

-------
     heat rate  by about  70 Btu/kWhr,  and
     decreases  the  plant  power  by about 3
     MW.

    From these points it is concluded that the
 combustor pressure drop  is a fairly insensitive
 parameter with respect to plant  performance
 and should be determined by the requirements
 for fuel distribution and control.

    Table  5  summarizes the performance,
 capital cost,  and energy cost of the Shell and
 Texaco  processes  and   the  fluidized-bed
 process with a combustor pressure drop of 50
 psi for the two air/fuel ratios, and for regener-
 ative  operation  by  CO2/H2O,   and  for
 once-through operation. The figures shown in
 the table refer to the case  in which the 1600F
 fuel  gas from the  fluid-bed  process  is  not
 cooled before combustion.  Table 5 indicates
 that the plant heat rate with the  fluid-bed
 process  increases  with  increasing  air/fuel
 ratio, while the heat rate is comparable for the
regenerative  and  once-through operations.
The plant heat rate  is  less  than  10,000
Btu/kWhr even for the high air/fuel ratio case
of 25 percent of stoichiometric. The Shell and
Texaco processes yield  a plant heat rate  of
about  13,000 because  of the  low thermal
efficiency, the high booster compressor power
requirements, and the relatively high ratio  of
power  produced by the  steam turbine to the
power  generated by the gas turbine. A plant
heat rate of 11,000 Btu/kWhr is estimated for
87 percent thermal  efficiency. Plant power is
comparable for all of the cases shown, with the
Shell  and Texaco  processes yielding the
highest power. Capital investment is based on
PACE  Plant  cost information and  the cost
estimates shown in Figure 8. The cost basis is
listed in Table  5. Capital cost is reduced for
the  fluid-bed  process  with  once-through
operation due to the great expense involved  in
limestone  regeneration  by  the  CO2/H2O
method. The  change in plant capital cost  in
going  from 14 percent air/fuel ratio  to an
                               Tables. PROCESS PERFORMANCE

Plant heat rate
(HHV), Btu/kWhr
Plant power, MW
Plant capital cost,
$/kW
Total energy cost,
mills/kWhr
Break-even distillate
cost, tf/106 Btu
Fluid bed process
CO2^H2O limestone regeneration
14% A/F 25% A/F

9,007
266.0

179.5

9.43

65.2

9,828
276.0

182.1

9.93

71.1
Fluid- bed process
once-through operation
14% A/F 25% A/F

8,906
269.0

158.8

8.96

59.7

9,716
279.0

159.5

9.42

65.1
Shell and Texaco
thermal efficiency
75% 87%

13,000
282.4

246.7

13.34

111.1

11,000
282.4

246.7

12.44

98.7
Conditions Assumed:
1. 50 psi combustor pressure drop.
2. 1600F fuel gas temperature from fluid-bed process with no cooling.
3. Fixed charges at 15%.
4. Residual oil at 45i|;/106 Btu.
5. Limestone at S6/ton.
6. 5% contingency; 4% escalation; 8% interest during construction; 2 years construction time; 2%
   A&E; 70% capacity factor.
7. No credit for sulfur recovered.
IV-5-12

-------
air/fuel ratio of 25 percent of stoichiometric is
slight.  The plant capital investment for the
Shell and Texaco processes  are based on the
estimated gasification system capital  cost of
about $100/kW provided by both Shell and
Texaco.

   Energy cost assumptions are listed in Table
5. The 14 percent air/fuel ratio operation of
the fluid-bed process is more economical than
the 25 percent air/fuel ratio, while the energy
cost for  once-through operation  is also less
than regenerative operation. The  relationship
between energy cost of regenerative and  once-
through operation is dependent on the cost of
limestone disposal,  but even if the limestone
cost should  double to  $12/ton  the   once-
through and regenerative system energy costs
will be about  equal.

   Break-even costs for No. 2 distillate fuel oil
is given in $/10&  Btu and represent the price
at which No. 2 distillate must be available for
the PACE Plant to operate at the same energy
cost as with residual oil gasification. With No.
2 distillate  prices ranging  from  85-95
-------
     capital about $10/kW, not including the
     cost of a waste heat boiler system and soot
     removal equipment.  The  second  option
     would result in inefficient operation with
     high energy losses, as with cooling of the
     gas by water injection. Both options  seem
     unattractive for power generation,  though
     further analysis  is  required  to develop
     quantitative conclusions.

 Comparisons   With   Alternative   Power
 Generation Systems

    Tables 6 and 7 compare  capital  invest-
 ments and energy costs of pressurized fluid-
 bed oil gasification with alternative oil-fueled,
 pollution   controlled,   power   generation
 systems. Capital and energy costs for a con-
 ventional  oil-fuel power plant utilizing lime-
 stone-wet  scrubbing  for  pollution  control,  a
 pressurized fluid-bed combustion plant fueled
 by oil, and a PACE Plant fueled with No.  2
 distillate,  are compared with  capital  and
 energy  costs  for  a  PACE Plant with  pres-
 surized fluid-bed oil gasification. The fluid-
 bed process is  carried out at  a 25 percent
 air/fuel  ratio with limestone regeneration by
the  CO2/H2O  method,   representing  the
highest capital cost and lowest efficiency of
the fluid-bed cases considered. Due to the low
cost of the PACE Plant package,  the  PACE
Plant with fluid-bed oil gasification is  about
140$/kW  cheaper than a conventional power
plant and 60$/kW cheaper than a regenerative
pressurized fluid bed combustion plant fueled
with oil. Energy cost of the PACE Plant with
fluid-bed  oil gasification,  assuming 4520
percent are projected. This conclusion  holds
over the range of factors explored  air/fuel
ratios, combiistor pressure drops,  and  lime-
stone regeneration methods.
                        Table 6. CAPITAL INVESTMENT COMPARISONS

Total capital cost,
$/kWc
Assumptions
Construction time,
years
Sulfur removal equip-
ment, $/kW
Plant capacity, MW
Conventional oil-
fired power plant
with scrubber
323.28

4.5
50.0
635
Pressurized
fluid bed oil
composition3
243.67

3.5
12.68
635
PACE plant
with no. 2
distillate fuel
137.80

2.0

269
PACE plant
with residual
oil gasification6
182.1

2.0
44.3
269
       aOperated with limestone regeneration.

       bC02/H20 regeneration of limestone and air/fuel ratio of 25% of stoichiometric-no
        cooling.

       C5% contingency; 4% escalation per year; 8% interest during construction; 2% AErE;
        70% capacity factor.
IV-5-14

-------
                          Table?. ENERGY COST COMPARISONS
                                      (mills/kWhr)

Fixed charges
Fuel
Limestone
Operating and main-
tenance
Total
Conventional oil-
fired power plant
with scrubber
7.91
4.11
0.12-
0.91
13.05
Pressurized
fluid bed oil
combustion
5.96
4.35
0.13
0.82
11.26
PACE plant
with no. 2
distillate fuel
3.37
7.80

0.52
11.69
PACE plant
with residual
oil gasification
4.45
4.42
0.13
0.93
9.93
     Assumptions
      1. Fixed charges at 15%; 70% capacity factor.
      2. 3 wt%  sulfur residual oil at 45^/106 Btu.
      3. Limestone at $6/ton.
      4. No.2 distillate at gO^/IO6 Btu.
      5. No credit for sulfur recovered.
 2. Cooling of the 1600 F fuel gas  produced
   by pressurized  fluid-bed oil  gasification
   with no recovery of the energy or by water
   injection   is   unattractive   from   the
   standpoint  of capital  and energy  cost.
   Cooling to 400 F with 25 percent recovery
   of the  energy may be uneconomical  and
   further analysis is required.


3. An  experimental  study  of pressurized
   fluid-bed oil gasification should be carried
   out to determine the process behavior in
   critical areas  and  to  obtain  a  better
   understanding     of   the    economic
   advantages to be gained from the process.
4. The process concept utilized by Shell and
  Texaco,   though  not  as  efficient   or
  economical   as  the  fluid-bed  process
  concept,  should   be   considered   for
  combined cycle power generation because
  it depends largely on existing technology.
  Improvement of the process performance
  may be possible.
ACKNOWLEDGMENTS

   This work was performed under contract to
the  Office  of  Research and  Monitoring,
Environmental   Protection  Agency.  P.  P.
Turner served   as  Project  Officer.  W.  L.
Wright,  Westinghouse  Power  Generation
Systems Division, arranged  and was  a con-
tributor to utility presentations on the atmos-
pheric pressure demonstration  plant.  Esso
(England) also contributed to the utility pre-
sentations.

REFERENCES

1.  Moss, G. The Desulfurizing Combustion of
   Fuel  Oil  in  Fluidized  Beds   of Lime
   Particles. (Presented at First International
   Conference  on  Fluid-Bed  Combustion.
   Hueston Woods. November 1968.)

2.  Craig, J. W. T., G. L. Johnes, G.  Moss, and
   J. H.  Taylor. Study of Chemically Active
   Fluid Bed Gasifier for Reduction of Sulfur
   Oxide Emissions.  Interim Report.  Esso
   Research  Centre,  Abingdon,  Berkshire,
   England. Prepared  for the Air Pollution
                                                                                IV-5-15

-------
    Control Office, Environmental Protection
    Agency,  Research Triangle Park, N.  C.
    under Contract  Number  CPA  70-46.
    August 1970.

  3. Hauser, J.  G. Optimum Uses of Energy
    Sources.  (Presented at Spring Conference,
    Southeastern  Electric  Exchange.   New
    Orleans.  April 1971.)

  4. Newby, R.  A., D. L. Keairns, and D.  H.
   Archer. Assessment of Fluidized Bed Oil
   Gasification   for   Power   Generation.
   (Presented at the 65th Annual Meeting of
   the Air  Pollution  Control  Association.
   Miami Beach. June 1972.)

5.  Information provided  during 'discussions
   with Shell,  May 3, 1972.

6.  Information provided  during discussion
   with Texaco, May 24, 1972.
IV-5-16

-------
                                              T0
                                            STACK
COMBUSTION AIR
LIMESTONE
MAKE-UP^
HH^

~l
I
GASIFIER
CLEAN
BOILER
FUEL GAS k
                             r^*^   TEMPERATURE
                                     *^ CONTROL
      AIR
                         REGENERATED
                             LIME
                         S02RICH
                         STREAM
                                           STREAM
                                           SULFUR
                                           RECOVERY
                     SULFATED
                       LIME
                     DISPOSAL
            REGENERATIVE MODE
COMBUSTION AIR
                    CLEAN FUEL GAS
                    ,	^\
                                            A   T0
                                            I STACK
                                            I	
                                      BOILER
                              LIMESTONE FEED
                 GASIFIER
  FUEL
|   OIL

I            A
I.	,
                      I
                SULFATE
               GENERATOR
                                           TEMPERATURE
                                              CONTROL
                                              STREAM
                                          	GAS
                                          	 LIQUID
                                                   SOLID FEED
                                                   SOLID CIRCULATION
                      SULFATED
                        LIME
                      DISPOSAL

            ONCE-THROUGH MODE
                 Figure 1.  Modes of operation.
                                                                        IV-5-17

-------
                                            40ft
              1. GASIFIER-DESULFURIZER
              2. SULFATE GENERATOR
              3. SULFATED-LIME BUNKER
              4. LIMESTONE BUNKER, FEEDER
              5. HIGH-EFFICIENCY CYCLONE
              6. FORCED DRAFT FAN
              7. GASIFIERFAN
              8. SULFATE GENERATOR FAN
 9. OIL FEED LINE
10. FUEL GAS LINE
11. LIME CIRCULATION LINE
12. SULFATED-LIME DISPOSAL LINE
13. BURNERS

	SOLIDS TRANSPORT
IHHIHIHH GAS TRANSPORT
             Figure 2.  Retrofit of 600-MW boiler, external once-through design.
IV-5-18

-------
                                                OIL
         ELECTRIC
         POWER

         COOLING
         WATER

         BOILER FEED
           WATER
        RESIDUAL
         FUEL
         OIL
   TO
MOTORS, etc.

TO C. W.
 USERS
    -*  TOW. H.
        BOILERS
    FUEL OIL
STORAGE, PUMPING,
 AND PREHEATING
       OIL TO/FROM 
       C RECOVERY -*-
                                                AIR
          FUEL
          GAS
                                      AIR
             AIR
                                 1
                                 STEAM
STEAM
WASTE HEAT
  STEAM
                             GASIFICATION
                                 AND
                             WASTE HEAT
                                 STEAM
                             GENERATION
      AIR
   COMPRESSOR
   AND MOTOR
                                              ASH
                                            SLURRY
           TO DISPOSAL
               POND
                                            TO
                             RAW
                             FUEL
                             GAS


UR
)SAL
RBON
SULFUR
RECOVERY
i
H2S
REMOV
SYSTE

HS
AL
M
STACK
GAS
I
FUEL
                           CLEAN
                           FUEL
                           GAS
                           TO
                           WESTINGHOUSE
                           SYSTEM
                                            FREE
                                            FUEL
                                            GAS
                                       CARBON
                                      RECOVERY
                                                           CARBON/OIL
                              SLURRY TO
                              FUEL OIL
                              STORAGE

                                FUEL OIL
           TO
           OIL
         PREHEAT
       MAKEUP
       NAPHTHA
                                       FROM STORAGE
                                    Figure 3.  Pressurized oil gasification for gas turbine fuel.
en

-------
                                         FUEL GAS
                          COIBBUSTOR
                         f
   H2$ OR SOj

   STREAM TO
   SULFUR
   RECOVERY
   LIME/DOLOMITE
   DISPOSAL
     REGENERATOR
                           SULFIDED
                           UME/
                           DOLOMITE
                                                    GENERATOR
                                                                               BOOSTER
                                                                              COMPRESSOR
            H20, C02

               OR
              AIR
STEAM
  OR
 WATER
INJECTION
              BOILER
                     ^	JI
                                                    SUPERHEATED
                                                      STEAM
                                                                             CONDENSER
                 Figure 4.  Regenerative pressurized oil gasification process.
IV-5-20

-------
SATURATED
STEAM TO
COMBINED CYCL
PLANT
WATER
RESIDUAL
FUEL
OIL
I. i COLD (100 TO 250 F)
FUEL GAS TO
E COMBINED CYCLE PLANT
SHELL/TEXACO
OIL GASIFICATION
SYSTEM
ENERGY LOSSES
	 	 . ^. (HEAT LOSSES SENSIBLE HEAT
FLUE GAS, COOLING WATER)

i i *
T A | AUXILIARY POWER (BOILER FEED WATER PUMPING
1 / \ | PUMPING, GAS CLEANING SYSTEM PUMPING)
OF
,OIL
  BOOSTER   I
 COMPRESSOR
   POWER
                    BOOSTER COMPRESSOR
  AIR FROM GT COMPRESSOR
(AIR/FUEL RATIO*45% OF STOICHIOMETRIC)
Figure 5.  Energy balance for the Shell/Texaco oil gasification process.
                                                                     IV-5-21

-------
                                  HOT (1600 F)
                                   FUEL GAS TO
LIMESTONE/
DOLOMITE
80F
STEAM _
WATER _

UUINDINC.U UTL,

FLUIDIZED BED
Oil G/KIFlPATinN
SYSTEM

LC ru\n I
** ENERGY LOSSFS (HFAT 1 nF<:,
CARBON DEPOSITION, SENSIBLE
HEAT OF FLUE GAS, COOLING WATER)
^e|i| Clip

fl A AUXILIARY POWER (COOLING WATER PUMP-
X 1 ING, OIL PUMPING, SOLID CIRCULATION,
- /\ 1 COMPRESSION FOR REGENERATION)
                     BOOSTER
                    COMPRESSOR I
                     POWER
                                   BOOSTER COMPRESSOR
AIR FROM GT COMPRESSOR
(AIR/FUEL RATIO=14 TO 25% OFSTOICHIOMETRIC)
              Figure 6.  Energy balance for the fluidized-bed oil gasification process.
IV-5-22

-------
BASIS: 250-MW COMBINED CYCLE PLANT
TWO 125-MW GASIFIER MODULES
15-atm GAS TURBINE PRESSURE
       50 ton
                                     FUEL GAS
                                    TO COMBINED
                                    CYCLE PLANT
       COVERED
       HOPPER
        CARS
No. 2 FUEL
OIL FOR
STARTUP    TO
                                               I  FLUE GAS FROM
                                               y COMBINED CYCLE PLANT
                                                __                 __
                                                                                TO STACK
                                                                            CONDENSER
               LIMESTONE
               RECEIVING
                HOPPER
                   ton
  C02
ABSORBER,
 STRIPPER
                                                         WATER -*-
                                                         (ACIDIC)
                                          MULTI-
                                            STAGE
                                          CYCLONES
  STARTUP
    OIL
  STORAGE
                                                                   COMPRESSOR

                                                                   V
                                                      REGENERATED
                                                    _ LIMESTONE
                            INJECTION .
                                                       TRANS-
                                                       PORT
                                                       AIR
                OIL PUMPS
                150 gal/min -
                SOOpsiAP
 PETROCARB
 INJECTION
TO 2nd MODULE
 TO 2nd
MODULE
                   8,800 bbl/day
                                         T__L_'k
                                   BOOSTER       I
                                    AIR
                                 COMPRESSOR  TEMPERATURE
                                              CONTROL
                                               WATER
                                                             STONE
                                                           DISPOSAL
                                                                           WATER FOR
                                                                          TEMPERATURE
                                                                       CONTROL INJECTION
                                                                            PARTICLE
                                                                            COLLECTOR
                                                                     REGENERATOR GAS
                                                                     TO CLAUS PLANT
 SOLID  	
 LIQUID	
 GAS    	
                 Figure 7-  Pressurized oil gasification plant flow diagram.
                                                                                   IV-5-23

-------
      3200
      2800
      2400
     2000
     1600
  o  2200
  2   800
                           BASIS:

                                MID-1972 COSTS
                               AP GASIFIER/P GASIFIER OUTLET=0.09
                                3 wt % SULFUR IN FUEL OIL
                                95% SULFUR REMOVAL

                           AIR/FUEL RATIO=14% OF STOICHIOMETRIC
                           AIR/FUEL RATIO =25% OF STOICHIOMETRIC
                           INDEPENDENT OF MR/FUEL RATIO
                     53 psi AP IN COMBUSTOR AND FUEL CONTROL
                 I Vps\ A P IN COMB. AND FUEL CONTROL I         ICLAUS PLANT
                                                                           .  I   BOOSTER
                                                                             ?COMPRESSOR-
                                                                           J
          120    140
160
    180     200      220     240      260

PRESSURE AT GASIFIER PRODUCT GAS OUTLET, psia
280
300
32
                     Figure 8.  Fluidlzed-bed oil gasification capital investment.
IV-5-24

-------
         6.  FUEL GASIFICATION AND ADVANCED  POWER
                         CYCLES-A  ROUTE  TO CLEAN  POWER
                                F. L. ROBSON
                  United Aircraft Research Laboratories
ABSTRACT

   The United States is currently faced with a growing gap between th,e demand for electrical
energy and the supply of economic fuels for generating this energy with minimum environmental
impact. The use of advanced power cycles utilizing technological spinoffs from the aerospace
industry in conjunction with fuel gasification/desulfurization offers a solution which could prove
to be not only technically feasible but economically attractive. A review of one such  system, the
Combined Gas And Steam (COGAS).is presented and the technical and economic advantages are
enumerated. There are, however, several problem areas, particularly in the interface between the
power system and the fuel system which must be resolved before the overall concept becomes a
commercially viable one. These problem  areas are  presented with the intent  of provoking
thoughtful discussion  and perhaps of opening new areas of research  among the conference
attendees.
INTRODUCTION

   The majority of the people attending this
conference were aware of the current energy
crisis facing this country well before there was
in fact a crisis. This is not the time to talk of
the reasons for the current situation but rather
to discuss methods of alleviating it by  using
the Nation's vast supply of coal, a fuel source
now held in low esteem by a large segment of
the air  pollution regulatory agencies.  It is
apparent that unless extensive effort is directly
applied towards this goal, or unless equally
extensive institutional  changes are brought
about in this country, we will be faced  with a
utility system based upon foreign sources for
one portion of our fuel and what can only be
termed an adolescent nuclear industry for the
remainder.

   The use of fossil fuels in the utility industry
will hopefully not parallel that currently being
followed in the transportation industry where
each  new  reduction   in  emissions   is
accompanied by a corresponding increase in
fuel consumption. To assure this, any power
system utilizing the advantages promised by
fluidized-bed  combustion/gasification  must
have the  potential  of  achieving operating
efficiencies significantly higher than currently
attainable in conventional boilers.

   However,  the various  economic forces,
natural and imposed, which now and in the
future affect the energy scene are such that a
delay in the  introduction of these advanced
power systems could result in electrical power
becoming more of a  luxury than  a necessity.
Thus, these advanced power systems must be
based upon technology which is now well in
hand, but which  will continue to grow, thus
affording better performance and economics
as each  new generation of power system is
achieved.
                                      rv-6-i

-------
    The  system  described  briefly  in   the    using evolutionary changes of technology cur-
  following paragraphs does indeed use the cur-    rently being  demonstrated  will offer  power
  rent technology available in the aircraft  gas    systems which could make use of the pollution
  turbine industry applied to the industrial seg-    reductions  offered   by  fluidized-bed  com-
  ment to realize efficiency and economic gains.    bustors and still attain efficiencies well beyond
  Second and third generation power systems    those of conventional steam-electric plants.
IV-6-2

-------
THE POWER SYSTEM

   Current   steam   power   plants   have
efficiencies approaching 40 percent at 1000 F,
a   value   which   is   limited   not  by
thermodynamics  but  by  economics.  There
have  been  steam  systems   designed  and
operated at temperatures  of 1200 F, but the
initial boiler and turbine costs were high and
the maintenance associated with operation at
this temperature eventually caused derating to
1000 F.1 What,  then,  are  the  alternatives
available to increase the  efficiency of power
systems? The answer  comes  of course from
Carnot's law which defines the efficiency limit
of  any heat engine operating between two
temperature limits. Referring to Figure 1,  it
can be seen that while the theoretical Carnot
efficiency is well above that obtained by a real
power cycle, several of the advanced  power
cycles  demonstrate efficiencies  which  are
nearly 70 percent of the theoretical limit. The
three  fossil fuel-fired sytems  having the
highest efficiency are based on the potassium
topping cycle, the COGAS cycle, and an MHD
topping cycle. Each of these cycles uses a high-
temperature cycle to top  a more or  less
conventional  steam   cycle.   Of the   three
mentioned, however, only  the COGAS  system
has been  demonstrated  in commercial size
and, in fact, nearly 2500-MW of these systems
are currently on  order by the utilities.2

    The advantage the COGAS system has over
the potassium and MHD topping systems  is
that of development. The COGAS  system  is
evolutionary in nature, adapting the advances
demonstrated in military and commercial air-
craft to  the large industrial turbomachines.
Thus,  the  technology  of  the JT9D engines,
used  on the 747,  which allows  operation at
2000 F and above will  be  adapted to the next
generation of industrial machines.

    In order to estimate the performance of the
COGAS  system,   the relation  of   system
elements must be established. While there are
many ways of combining the  gas turbine and
steam  system  equipment,   essentially  two
generic types  emerge: (1)  the waste-heat
recovery type  in which the turbine exhaust
raises steam with or without additional firing,
and (2) the  pressurized-boiler type  in which
steam  is  raised and superheated using heat
from   a   gas  turbine  combustor.  These
variations are shown  in  Figures 2 and  3,
respectively. The effect of configuration  of
performance can be obtained by interpreting
Figure 4, which shows station efficiency versus
gas turbine  participation  for the waste-heat
recovery type configuration.

   The parameter gas turbine participation is
really a measure of oxygen utilization. For a
given  turbine  inlet temperature,  a  fixed
amount of oxygen is required for combustion;
the remainder, plus a large  amount in the
dilution air, is exhausted through the turbine.
In the waste-heat recovery system, if there is
no  additional firing,  the efficiency  is the
highest value, i.e., the right end of the lines of
efficiency in Figure 4. If there is firing using
up the additional oxygen and generating more
steam, the gas turbine  participation declines,
and the efficiency becomes lower until the left
end  of the  line, the point of which all the
oxygen is consumed, is reached. The foregoing
applies  to   combined-cycle   systems   with
turbines   operating at  2000F  and  above.
Below this  turbine inlet  temperature, the
combined-cycle  efficiency could,  in   fact,
increase with high steam fractions since it is
possible to have steam-cycle efficiencies signi-
ficantly better than the efficiency of the lower
temperature gas turbine.

   In the  supercharged cycle, fuel and air are
burned essentially at stoichiometric conditions
and  the combustor exhaust is cooled to tur-
bine inlet temperature by raising steam rather
than be dilution air. This  raising of steam  by
combustion has the same effect as supple-
mentary  firing, i.e. the system  efficiency is
reduced. Actually these systems are, obeying
Carnot's law,  which in essence says that the
system which uses the entire heat input at the
combined-cycle efficiency will be more effi-
cient than the system which utilizes  only part
of the heat input at combined-cycle  efficiency
and  part  at the  lower  steam-cycle efficiency.
                                                                                   IV-6-3

-------
   There are also considerations outside of the
 power cycle which  influence the ratio of gas
 turbine  power  to  steam power which  will
 appear again when the integrated  gasifica-
 tion/power system  is discussed.

 THE GASIFICATION SYSTEM

   There is  a  wide  variety of methods for
 converting solid or liquid  fuels to  gaseous
 form.  Rather  than discuss the  operational
 characteristics of any one of these processes, it
 would prove  more fruitful  to discuss those
 characteristics which are important from the
 viewpoint of utilizing the fuel in an advanced-
 cycle power system.

   Perhaps the most important aspect of fuel
 supply is cleanliness. While the turbine can
 handle a wide variety of fuels, each of these
 fuels must meet rather rigid contamination
 criteria. For gaseous fuels, the current specifi-
 cations3  will  allow  no more  than   0.08
 lb/106ft3 of total solids. No  particulate size is
 specified, but filters  are located in  the  fuel
 lines capable  of removing  particles  of 30
 jum  and above. To minimize blade erosion,
 however, particulates should be small enough
 to follow the air stream through the blading
 without  impingment. The  exact particulate
 size has not been firmly defined but measure-
 ments of smoke particles seem to indicate few
 larger than 20 to 25 jum.

   Current specifications  limit  total  sulfur
 content in  the fuel  to 162 lb/106ft3 of which
 HzS can be  no  more than  0.18 percent by
 volume.  This amount,  assuming it was all
 converted by combustion to SC>2, would be the
 equivalent of about 3 lb/106 Btu of methane-
 type fuel gas or  about 2 lb/106 Btu  of low-
 heating-value  gas from coal.  Both of these
 values   are  above  EPA   regulations  (no
 allowable SO2 for gaseous fuel, 1.2 lb/106 Btu
 for coal)  so that any sulfur removal methods
 which  will  meet EPA regulations could be
 suitable from the turbine viewpoint.
   A gas turbine can handle a wide latitude of
fuel heating values, ranging from  blast  fur-
nace gas ( ~ 100 Btu/ft3) to propane ( ~ 2000

IV-6-4
 Btu/ft3).  Thus, the chemical heating value,
 per se, is  not a problem. There is, however, a
 heating value dependent problem which must
 be considered. This  is the problem of fuel
 delivery  pressure.  The  first  aspect  of this
 problem is essentially one of hardware. The
 sizing of  fuel  manifolding, injection nozzles,
 etc., is a function of fuel heating value, i.e., a
 given Btu/min must be supplied to the engine
 and a fuel with a heating value of 150 Btu/ft3
 requires a higher volume throughput than one
 with  1000 Btu/ft3. Thus, to  reduce the fuel
 handling  equipment to sizes compatible with
 high release gas turbine combustors, the low-
 heating-value  fuels should be supplied  at a
 pressure higher than the pressure ratio of the
 engine.  The relationship  between the  fuel
 delivery pressure and the Btu/ft3 requirement
 differs among engine types but is of the form:
 P = (f)  LHV/(Specific  gravity  x delivery
 temperature)1/2.  Unfortunately   the  func-
 tional form is inverse; thus,  as the  heating
 value decreases or temperature increases, the
 fuel delivery pressure increases.  This has a
 significant  effect  on  overall  system  per-
 formance and will  be discussed later in the
 paragraphs dealing with the integrated station
 performance.

   The  most  important characteristic  of a
 gasifier designated for use with a power system
 is its efficiency in converting coal Btu's to fuel
 gas Btu's. There are two gasifier efficiencies
 that need to be considered. The first is the cold
 gas efficiency (chemical heating value of fuel
 gas/chemical heating value of coal) and the
 second is  the  hot gas  efficiency (chemical &
 sensible heat  in fuel gas/chemical  heating
 value of coal). The manner in which these two
 efficiencies  affect  the   integrated  system
 performance is  complex, but  if one thinks
 again in  terms  of Carnot's  law,  all of the
 chemical  heating value  (represented  by the
 cold gas efficiency)  is  used at  combined-cycle
 efficiency, while the sensible heat may or may
 not be so utilized. If the hot gas can be used in
 the engine, there would be no degradation in
the performance. However, if the gas must be
 cooled by raising steam, then the sensible heat

-------
could be used only at steam cycle efficiency.
There is of course, a "however",  attached to
the foregoing. If the sensible heat were to be
used in regenerating the gasification system, it
would be used at  essentially combined  cycle
efficiency. This use of the fuel gas sensible
heat requires relatively  expensive  heat ex-
change  equipment.

   Another gasifier attribute necessary for use
with power  plants is operational flexibility.
While power systems designed for  base-load
operation  do  not   require   fast  startup
capability, they do operate over  a range of
power  settings  ranging  from  perhaps  70
percent to full power. Thus,  even  base-load
applications    require    some    turndown
capability. One of the more attractive features
of the COGAS system is its capability of rela-
tively fast startup, 5 minutes to full power for
the gas  turbine power and less  than 1 hour for
the total plant. This capability lends itself to
applications in mid-range load factor (3000 to
6000 hr/yr) in which daily startup would be
common place, with shutdown over weekends
or holidays. A gasifier for use in this system
would have  to have  a fast-start capability.
Since this type  of  power system  typically
operates over a range of power settings from
40 percent to full power, flexibility in gasifier
operation would be necessary.
INTEGRATED    FUEL
POWER SYSTEM
PROCESSING/
   One of the basic tenets of mathematics is
that the whole is equal to the sum of its parts.
The  power  system   has,   in  a   sense,
circumvented  this by putting together  two
parts   of  comparable  efficiency  into  a
combined system of significantly better  per-
formance. It would be indeed fortunate if this
symbiotic relationship could  extend  to  the
joining  together of the fuel  processing  and
power  cycle  portions  into   an  integrated
system.  However,  the serendipity does  not
carry  over; in fact, the requirements at the
interface of the two parts can cause a noticable
reduction in combined performance. There is
also an environmental consideration involving
the production of NOX which could influence
the selection of overall systems configurations.

   One of  the  simplest  of all  integrated
systems is shown in Figure  5. Air for the
gasifier is bled from the compressor, raised to
the required gasifier pressure in a booster
compressor, mixed with the fuel in the gasifier
with the resultant hot fuel gas being supplied
directly to the turbine. Some heat exchange
between the air streams would be possible.
This system utilizes the fuel  sensible heat in
the engine.

   A  much more  complex  configuration
(Figure 6)  results if the fuel sensible heat must
be recovered for use elsewhere in the cycle. In
Figure 6, the gasifier is run at  essentially
atmospheric pressure with the fuel gas exiting
to a boiler/superheater. From there  the fuel
gas passes through a heat exchanger  where
bleed air from the compressor enroute  to the
gasifier is heated, then through a feed water
heater and finally into the booster compressor
were it is  raised to the required pressure for
injection into  the gas turbine burner. The
booster compressor could be  driven,  in part,
by an expansion turbine  in which the heated
bleed air is let down to gasifier pressure. The
feedwater and superheated steam would be
utilized in the waste heat boiler.

   What are the performance differences of
these two  systems? The simple systems  utilize
all the fuel  heating value at combined-cycle
efficiency, while the second utilizes only the
chemical   heating   value    at   combined
efficiencies with the  sensible  heat being used
at steam cycle efficiencies.   The efficiency
differences can be found by inspecting Figure
7, based upon data from Reference 4, in which
the percentage change in efficiency with fuel
temperature is  shown  for  three  different
turbine inlet temperatures. From all indica-
tions, the  system utilizing the sensible heat in
the gas turbine is  the most efficient. This
reinforces the results shown in Figure 2, which
indicated  that large gas turbine participation
was more efficient.
                                                                                   IV-6-5

-------
   There are, however, several considerations
that must be made before a final choice can be
made. Besides the very difficult hardware pro-
blems associated  with  fuel  control systems
handling gases at 35 to 50 atm  and tempera-
tures above  2000F,  there  are  the environ-
mental constraints to  be  considered. The
systems pictured in Figures 5 and 6 are based
upon  both  sulfur  and  particulate removal
within the gasifier; i.e., a method of desulfuri-
zation and particulate removal  that operates
at 1500F and above. Currently, the majority
of sulfur removal systems operate at 600 F or
below, some even requiring below zero gas
temperatures. With the exception of fluidized-
bed gasifiers and perhaps one or two  other
types, the majority of gasifiers require external
desulfurization and thus require some method
of heat recovery from the fuel  gas stream.
   A second environmental constraint is that
 of NOX formation. It is well  established that
 the  formation of NOx is  strongly dependent
 upon the combustion temperature. Combus-
 tion temperature, in turn, is a function of both
 fuel heating value and of fuel and air preheat.
 This dependency is shown in Figure 8 where it
 can  be seen that the combustion temperature
 is a  stronger function of sensible heat than of
 HHV. Using Figure 9, which shows NOX con-
 centrations as a function of temperature, the
 rise in combustion temperature due to fuel gas
 sensible heat will increase; in the extreme, the
 emission of NOx by a factor of about 25 (i.e.,
 from less than 5 to nearly 100 ppm with a fuel
 gas  having an HHV of 120 Btu/ft3). As the
 fuel  HHV increases, the base NOX emission
 factor increases and the multiplying factor due
 to sensible heat decreases  slowly at first and
 then rapidly as NO equilibrium is approached.
 The allowable concentration, using EPA  regu-
 lations for NOX  from coal-fired power-plants
would be  in  the order of 120 to 160  ppm,
depending on engine efficiency. (If the system
 were to be considered as  a gas-fired station,
 the allowable concentrations could be 35 to 45
 ppm. A  brief  discussion of emission  regu-
 lations and their form is given  in Appendix A.)
   The situation;'therefore, is that as the HHV
of the product gas increases, the sensible heat
must be decreased (or vice-versa) in order to
meet the NOX standards. Since the NOX pro-
duction increases  more rapidly with sensible
heat, it would seem to be more advantageous
to  have a  gasifier with a high  cold  gas
efficiency (more HHV in the gas) as was men-
tioned before. This is doubly beneficial since
the HHV of the fuel also has  a noticeable
effect  on the  integrated  system  efficiency
(Figure  10).   (This   effect  is  somewhat
exaggerated  in Figure 10  since  the  steam
system efficiency used in preparing this figure
was relatively low, i.e., ~30 percent.) This is
because of two effects, a reduction in steam
generation due to the decrease in mass flow of
fuel,  and a concurrent decrease  in booster
compressor work.
   There are several ways to increase the HHV
of the product gas. Two of the more promising
are: (1) regeneration of the gasifier air, and (2)
oxygen enrichment of gasifier air. In the first,
the air to fuel ratio needed to attain a given
gasifier temperature is reduced  as the air is
preheated. This means less nitrogen dilution
and  more   Btu/ft3.  Oxygen   enrichment
accomplishes the same thing; i.e., a reduction
in nitrogen dilution thereby increasing HHV.
In fact, the use of oxygen alone to blow the
gasifier would  result  in the  production  of
synthesis gas having  an  HHV  of about 315
Btu/ft3.  Unfortunately,  the  combustion  of
synthesis gas  could result  in  greater NOX
production than the combustion of methane.
   There are, or course, modifications  which
can be made to the combustion process which
could effect a reduction in NOX emission. One
of  these,   off-stoichiometric    combustion
resembles, in theory, the method being used
with some success in gas-fired steam boilers.
At this time, there is no assurance that the
combustion  efficiency during  this staged-
combustion  will  be   comparable to that
currently  obtained in gas  turbines, e.g.,
greater than 99 percent. This is especially true
with  the  low-Btu  fuels whose  combustion
characteristics in gas turbine burners have not
been extensively studied.
IV-6-6

-------
A SUMMING UP

  The  foregoing  discussions  have  briefly
described the advanced power system which
offers the potential  of attaining efficiencies
nearly 50 percent greater than those currently
obtainable yet utilizes equipment which would
be  evolutionary  developments  of the  com-
bined-cycle systems that are currently being
placed on-line by utilities. When used in con-
junction  with  coal/residual oil  gasification,
such systems could  generate electrical power
with a minimum of pollution while using  rela-
tively abundant  high-sulfur fuels. Although
economics have  not been  treated thus far,
prior studies have indicated that the combined
fuel gasification/advanced power system could
generate this pollution-free power at costs less
than conventional steam power plants having
equivalent emission characteristics  and  at
costs which could easily be less than  those
associated with nuclear power (Figure 11 from
Reference 5).
   That this answer to a maiden's prayer is not
without  problems is apparent  even to the
casual reader  of this  paper. Some of these
problems have been touched  on  config-
uration of the power cycle, pressurization level
of the gasifier, trade-offs between sensible and
chemical heating  value, etc.  A myriad  of
practical problems exist in interfacing the two
complex systems.
   However, many of these problem areas can
be resolved only by construction and operation
of a prototype plant of large enough scale to
demonstrate the  concept that clean power can
be  generated  from high-sulfur  fuels  with
acceptable  economics.  Until  a  successful
demonstration(s) takes place, utilities will view
the concept as just thata conceptand will
continue to displace fossil-fired systems with a
greater  dependence  on nuclear power.  The
recent announcements  by the Office of  Coal
Research describing its plans for one or more
demonstration  plants   gives   hope  that   a
reasonable program will soon be underway 
one which will lead to the introduction of com-
mercial systems  by  the latter  part of this
decade.
REFERENCES

1. Giramonti, A. J. Discussions of Steam and
   COGAS  Systems with the  Babcock  and
   W^lco'x  Co.,  Barberton,  Ohio.  UARL
   Report UAR-H246, 1969.

2. Packaging  Sells the  Combined  Cycle.
   Electrical World. September 1,1972.

3. P&WA Specification 526 - Gaseous Fuel.
   Industrial for Turbine Engine. June 1968.

4. Giramonti,  A. J. Advanced  Power Cycles
   for Connecticut  Utility  Systems,  UARL
   Report L-971090-2, January 1972.

5. Robson, F.  L.  Clean  Power  from  Gas
   Turbine-Based      Utility      Systems.
   Combustion. July 1972.

APPENDIX A

Emission Standards

   Current emission standards are based upon
the  amount  and  type of  fuel  burned;  x
pounds/106 Btu for coal, y lb/106  Btu for oil,
and z lb/106 Btu for gas. If the full advantage
is to be taken of the advanced power systems,
a new basis for standards must be  used.

   First, as was alluded to in the main text,
fuel  type  should  be defined in  a  different
manner. The power system burns a gaseous
fuel  even though coal, residual oil,  coke, or
garbage is used in the gasifier. In fact, in an
actual system it would be hoped that  the input
to the gasifier could be switched more or less
as the fuel market dictates.

   Secondly, the current standards are based
upon input rather than the output,  which is
the real purpose of the power system. It is now
possible for a power  station  to meet  the
regulations but emit, in absolute numbers,
significantly more  pollutants than would an
advanced power system of equal power. For
example,  a conventional  steam  station  of
1,000-MW  output  running  at  off-design
conditions  to  meet NOX  regulations could
have a heat rate of perhaps 12,000 Btu/kWhr
and, meeting EPA standards, would put out
                                                                                  IV-6-7

-------
  2,400  Ib/hr  or  2.4  Ib/MWhr of NOX. An
  advanced system having a heat rate of 8,000
  Btu/kWhr and meeting the same standards as
  currently written would emit, for the same
  power demand, l,6001b/hr or 1.6 Ib/MWhr of
  NOx.

    As  turbine inlet temperatures increase,
  there  is  a second-order  increase in  NOX
  emissions. While minor, this increase could
  result  in  an emission/106  Btu above  the
  prescribed level. However, the power system
  has  become more efficient, and there could
well  be  a  decrease  in  emissions/output
compared to the  lower temperature  system.
Referring to  the  above example,  suppose a
500F increase in turbine inlet temperature
caused a 10 percent increase in NOX emissions
but also a 20 percent decrease in heat rate.
The system would no longer meet the  current
EPA regulations but would, in fact, emit only
1.4 Ib/MWhr.
   For these reasons it appears that standards
based upon emissions/MWhr would be a more
reasonable choice.
IV-6-8

-------
5
o
5
I
                                                     CARNOT EFFICIENCY
                                             POTASSIUM-STEAM
                                             BINARY SYSTEM
                                              (NUCLEAR FUELED)
            TYPICAL PWR
               SYSTEM
CLOSED-CYCLE
  HELIUM
 TURBINES
                 TYPICAL
                    SYSTEM
                                                        rnr.. COMBINED GAS TURBINE
                                                        cuwta AND STEAM TURBINE    	
                NUCLEAR
                -TEAM
                                                               CURRENT DESIGN
30
     20
                                        2000                   3000
                                 MAXIMUM CYCLE TEMPERATURE, F
 Figure 1,  Comparison of estimated thermal efficiencies for advanced-cycle power stations.
                                                                                 4000
                                                                                    IV-6-9

-------
           COAL OR RESIDUAL OIL
                                                          POWER
                                                          TURBINE
                                                                          ELECTRIC
                                                                         GENERATOR
                                                                          ELECTRIC
                                                                         GENERATOR
                                     PUMP
                           Figure 2. Combined gas-steam turbine system.
IV-6-10

-------
                           COMPRESSOR
                             TURBINE
                                           POWER
                                           TURBINE
                                                         ELECTRIC
                                                        GENERATOR
                                                         ELECTRIC
                                                        GENERATOR
                      PUMP
Figure 3.  Supercharged combined gas and steam turbine system.
                                                                   IV-6-11

-------
       70
       60
0>
i*
as
5   50
  Oil
    Ul
       40
       30
                   STEAM CYCLE EFFICIENCY  38.8%
                                                             TURBINE INLET
                                                             TEMPERATURE, F

                                                                  3100
                                                                 2800
                                                          2200
                        20
                                                                        80
                  40               60
          GAS TURBINE OUTPUT, % of station output

Figure 4.  Performance of exhaust-fired combined system.
100
IV-6-12

-------
    FUEL GAS
     BOOSTER COMPRESSOR
                                    STACK
Figure 5.  Schematic of high-pressure system.
                                                        IV-6-13

-------
FUEL GAS
  FROM
CONDENSER
                                                FROM TURBINE
                                                 WASTE  HEAT
                                                    BOILER
                                                        TO FUEL GAS
                                                       SUPER HEATER
                FROM FUEL GAS
                SUPER HEATER
  TO FUEL GAS
  WASTE - HEAT BOILER
                    Figure 6.  Schematic of low-pressure system.
IV-6-14

-------
o
            TURBINE INLET
            TEMPERATURE,0 F
COMPRESSOR PRES
SURE RATIO, atm
500
         1000           1500
FUEL TEMPERATURE TO TURBINE,F
                                                                      2000
                                                 2500
               Figure 7. Effect of fuel temperature on integrated station efficiency.
                                                                                    IV-6-15

-------
     4400
     4200
  M

  &  4000
     3800
     3600
     3400
               INCREASE FUEL
               TEMPERATURE
                                             INCREASE FUEL HHV
                                            REFERENCE FUEL HHV-120 Btu/scf
                                            REFERENCE FUEL TEMPERATURE=80F
                                            STOICHIOMETRIC FUEL-AIR RATIO
                                            INITIAL AIR TEMPERATURE=825"F
         100
120              140              160
     FUEL CHEMICAL PLUS SENSIBLE HEAT, Btu/scf
180
200
      Figure 8.  Effect of fuel gas chemical and sensible heat on combustion temperature.
IV-6-16

-------
 1000
i
  100
    .0

   0.1
                i     i      i
   o LOW-HEATING-VALUE FUELS

   0 METHANE

_* JP-5
 BASED ON P & WA THREE-ZONE BURNER MODEL
 TURBINE INLET TEMPERATURE=1760F
   A    I      I     I      I     I
     3200 3400  3600  3800   4000  4200  4400  4600

         MAXIMUM COMBUSTION TEMPERATURE,R

 Figure 9.  Effect of combustion temperature
 on nitric oxide emissions.
                                             100  110   120   130   140   ISO    160  170

                                                FUEL HIGHER HEATING VALUE, Btu/scf

                                          Figure 10. Effect of fuel HHV on integrated
                                          station efficiency.
                                                                                    IV-6-17

-------




1.50

125

P
&
o
^ i on
2 ItUU
o
2
3
S -75
>
I
0.50
0.25
0
1. CONVENTIONAL STEAM-ELECTRIC STATION BURNING UNTREATED COAL
2. CONVENTIONAL STEAM-ELECTRIC STATION BURNING UNTREATED COAL
WITH 85% EFFECTIVE SULFUR OXIDE STACK GAS CLEANING
3. COGAS STATION BURNING DESULFURIZED FUEL
- 4. BASE-LOAD GAS TURBINE STATION BURNING DESULFURIZED FUEL
5. NUCLEAR
















1



















2































MM



























WMMI
4






























BMMDB



























M1MMB
1
1

















^^^m
2














	















MMB
3














4















MMBMi
5























4
'

















MMM
2

















.










^BHMI
3














4















5



-


-
-

                 1970 DECADE
    1980 DECADE
LEVEL OF TECHNOLOGY
1990 DECADE
                 Figure 11. Busbar power costs for coal-based power plants.
IV-6-18

-------
       7. A DESIGN BASIS  FOR UTILITY GAS FROM  COAL

                               C. W. MATTHEWS
                          Institute  of Gas Technology
INTRODUCTION
   The  preferred  solution  for control  of
atmospheric pollutants  from the  stacks  of
electric utility boilers is to substitute clean fuel
for polluting fuel. In many cases a clean fossil
fuel such as natural gas is not a practical sub-
stitute for coal in the generation of electricity
because of scarcity and cost. Furthermore, the
available supply of clean fuel may combat pol-
lution more effectively when used to  fulfill
residential and small commercial needs.

   The  combustion  products  of coal con-
tribute one-eighth  of the total  atmospheric
pollutants  emitted in  the United  States,
including one-half of the sulfur oxides and
one-quarter of both the nitrogen oxides and of
the particulates. Sulfur emissions from coal
combustion may be reduced by:  (1) using low-
sulfur coal,  (2) cleaning high-sulfur coal by
physical methods,  (3) removing  sulfur oxides
from coal combustion gases, (4) removing sul-
fur during the combustion step,  (5) producing
de-ashed low-sulfur fuel by solvent processing
of coal, and (6) gasifying coal and removing
sulfur from the gas before  combustion.

   The last method, coal gasification with gas
cleaning  before combustion,  promises the
greatest reduction in sulfur emission. Most of
the sulfur gasified  appears  as hydrogen sul-
fide.  Several   different   commercial   gas
cleaning processes  are available today which
can reduce the hydrogen sulfide content of gas
streams to less than 10 ppm; some processes
can remove hydrogen sulfide to 1 ppm or less.
   This paper discusses the selection of design
criteria for the  gasification  of  coal  and
cleaning  of the  generated  gas before  com-
bustion in an electric utility boiler. The pre-
liminary plant design description is for a large
pilot plant installation that will demonstrate
the feasibility of this concept.

   The gas produced from coal for the boiler is
called utility gas, although producer gas and
low-Btu gas are equivalent names. It is made
by  gasifying coal with  air and  steam at
elevated pressure. Dust and sulfur compounds
are removed from the utility gas before it is
burned in  the   power generating  system.
Heating value of the utility gas will be between
140 and 250 Btu/ft3, depending on the gasifier
and the plant design.

   The low heating value of utility gas limits
the   distance   it   can   be   transported
economically.  When  used to  fuel  electric
power stations, it probably will be generated
onsite. Retrofitting coal gasifiers  to  existing
boilers is  one  of the most  important appli-
cations now .for coal gasification plants. In the
future  installations,  combined  gas turbine-
steam turbine systems will be served by gasifi-
cation plants not only for reduced atmospheric
pollution, but also for greater efficiency in the
generation of electricity. From these plants we
expect savings in investment and  decreased
electricity costs as well as less heat rejection to
the environment  and better conservation of
our coal  resources.
                                        IV-7-1

-------
   We believe that this country has an urgent
 interest in the demonstration of the practica-
 bility of coal conversion to clean utility gas for
 the electric power industry as soon as possible.
 Successful  achievement of this goal in the
 shortest possible time  through government
 and industry support will  provide substantial
 benefits to the country and to the electric
 utility industry.
    The program proposed by the Institute of
Gas Technology 0GT) for proving this concept
includes construction in the near future of a
large pilot plant which will be located near an
existing boiler. The pilot plant will be capable
of feeding from 10 to 50 ton/hr of coal and will
fuel a power plant with a generation capacity
of 20 MW or more. We believe  that with
favorable results  from this  pilot plant com-
mercial plant  designs can be undertaken by
the end of 1975.
IV-7-2

-------
GENERAL DESIGN PRINCIPLES

   The  entire concept hinges on the  coal
gasifier  performance. The gasifier operation
must be reliable; it  must gasify a high per-
centage of feed carbon; and it must be load-
following, that is,  capable of operating  hi
response to  the power system  requirements.
The focal point of our effort is to demonstrate
gasifier  operation.
   We anticipate that coal conversion to utility
gas will be practiced to a great extent in the
coal belt from Illinois  to New York. Because
the coals  from  this  region  have  caking
properties, the  plant  must  be  capable  of
accepting these coals as feed.

   Initially, utility gas will be produced for
existing boiler systems. Operation of the large
pilot plant facility will show the practicality  of
retrofitting coal gasifiers  to  existing boilers.
This is one of the most important applications
now for coal gasification plants. In  the future,
for savings in investment and for decreased
electricity costs, combined-cycle systems will
be served by gasification plants. Flexibility will
be provided in the  pilot plant design to test
advanced    power    generating    system
components.

   The pilot plant gasifier will be large enough
so that scale-up to commercial-size  gasifiers  in
the future can be done with  confidence. For
example, a single gasifier fed at a rate of 10
ton/hr and operating at 300  psia  will fuel a
power station generating about 20 MW  of
power. Operation at 1000 psia will increase its
capacity by more than  3 times. The diameter
of this gasifier permits shop  fabrication and
rail shipment for reduced construction time
and cost.

   Our guiding principles in design of the first
large  pilot plant for  conversion  of  coal  to
utility gas are:

1. Prove the  gasifier design  and  operation.

2. Accept caking coal  as  feed.

3. Demonstrate application of coal gasifica-
tion to existing boilers.

4. Build gasifier large enough so that it can be
directly  scaled to commercial size.
UTILITY GAS FROM COAL PLANT

   Figure 1 is  a block flow diagram for our
proposed utility gas from  coal  plant. This
design is suitable when the gasification plant
is installed  to  fuel a small to medium size
boiler. I will briefly describe the flow scheme
in Figure 1 and then discuss in more detail the
important parts of the plant.

   The coal feed is crushed to the desired size.
Lock hoppers are used to transfer coal from
atmospheric pressure to the elevated pressure
of the gasifier. Heat is recovered from the hot
raw gases, and a small part of the cooled gas is
used to pressurize the lock hoppers. The main
gas stream expands from high to low pressure
through a gas  expander,  thereby generating
power needed to drive the large gasifier air
compressor. Gas vented from the lock hopper
system  rejoins  the  low-pressure  main gas
stream. The combined gas streams are cleaned
of sulfur at low  pressure by the Stretford
process or one that is similar. The Stretford
process  produces elemental  sulfur directly
from hydrogen sulfide. After  sulfur removal,
the gas flows to the boiler. Associated with this
process train are a large air compressor, ash
handling and disposal equipment, waste-water
treatment,  and  possibly oil  stabilization and
storage equipment. When an efficient, high-
temperature sulfur removal system has been
developed, we believe that it will replace the
equipment shown within  the  dashed lines.

DESIGN PRESSURE

   In plants manufacturing pipeline gas from
coal,  maximum methane formation within the
gasifier  is  desired  for  improved  thermal
efficiency. To obtain this, gasifier pressures of
1000 psia or more are preferred. The thermal
efficiency of utility gas plants does not depend
on methane formation within the gasifier, and,
therefore, we have more flexibility in selection
of plant pressure. If necessary  and  this is
what we have done  the plant pressure is
established on  other considerations than the
chemistry of gasification.
   Lock hoppers were selected  to transfer coal
into the plant. A dry solid feed permits a less
complicated gasifier design, and the plant fol-
lowing the gasifier is also less complicated. For
high  plant reliability when  using  dry feed
                                                                                   IV-7-3

-------
 systems,  the  performance of  lock  hopper
 valves will probably set the  upper pressure
 limit. Today, the best commercial lock hopper
 operation  is that demonstrated with  Lurgi
 gasifiers. In these, the pressure  difference
 between lock hopper and gasifier is  about 300
 to 350 psi. Even though we want to gasify at
 pressures up to 1000 psi, we will design for a
 pressure of about 300 psi because lock hopper
 valves are available which  work at  this pres-
 sure.  We intend  to search  for  improved
 methods to feed dry coal into  higher pressure
 systems.

 COAL FEEDING SYSTEM

   Figure 2 presents a simplified illustration
 of the solids handling system,  which includes
 the coal feed system,  pretreatment, gasifica-
 tion, ash removal, and dust removal from the
 gas.

   As  described above, the single-stage lock
 hopper was  selected  to transfer  coal  from
 atmospheric pressure to the elevated pressure
 of the gasifier. This feed system was chosen so
 that the  gasification  plant will be  simple,
 reliable, and less costly.

   The disadvantages  of the lock hopper are
 evident. The most important is the difficulty in
 obtaining reliable operation of the lock hopper
 valves.  During the hopper  cycle, these valves
 alternately seal  against the gasifier pressure,
 and  then open to pass a fine,  dry, abrasive
 solid. Coal dust tends to pack in the valve, cut
 the packing,  and jam in the guides and seat.
 The valve must  seal  fairly  well.  Leakage
 overheats  the  valve  and  lock   hopper;
 introduces dirty, raw  gases into the  hopper
 and its vent gas system; and wets the cool coal
 with moisture, oil, and tar causing the coal to
 bridge and no longer  flow easily.
   A second disadvantage is the need to vent
 gas from the hopper during its operating cycle.
 Loss  of  this  vent gas  decreases  process
 efficiency.  Recompressing  the gas into the
 system is  expensive. In this utility  gas  plant
design, the vented gas  is collected and mixed
with the low-pressure main gas stream before
sulfur  removal.  In  case the valves  leak, the
vent gas system includes a cooler and vapor-
 liquid  separator to reduce contamination  of
 product gas.
   The  IGT  HYGAS process is designed to
manufacture pipeline-quality  gas from coal.
In this process coal is fed to the gasifier in the
form of a slurry which  is pumped to system
pressure.  This is a more reliable  feeding
system than high-pressure lock hoppers, and
the  slurry  pumps  are  capable  of  good
operation to  1000 psi and higher discharge.
You might ask: Why  not use a slurry feed
system for utility gas generation?

   Slurry  feeding  introduces   additional
complexity  in  the  plant. Equipment  for
making the slurry must be provided. When  the
slurry enters  the gasifier,  the  liquid must be
vaporized. Therefore a drying bed is added to
the gasifier, and the hot, raw gasifier gases  are
used to supply the heat needed for drying and
for stripping. The lowered gas temperature
makes efficient recovery of heat from the gas
difficult. Stripped slurry oil must be recovered
efficiently for recycling.  In its  recovery, more
than one stage of quenching may be needed;
the gas stream is cooled to 100F, thus adding
to the cooling water demand; and activated
carbon   towers  or  sponge  oil  scrubbing
completes final  oil  removal  from the gas.
Then, the recovered  oil has to be dewatered
and stripped of dissolved gas before returning
to the slurry tanks.
   We concluded that utility gas plants for the
electric  power industry will be less complex,
less  costly,   easier  to  operate,  and more
efficient when using lock hoppers to feed coal
into the  gasifier.

COAL PRETREATMENT

   Most bituminous coals have the property of
caking   or   agglomerating   when   heated.
Agglomeration of coal  within  the  gasifier
cannot be tolerated because of the possibility
of pugging. So that the utility gas process can
accept the widest variety of coal  as feed,
facilities for  modifying or  destroying  this
property must be  a  part of the process.

   Pretreatment takes place in the presence of
air at 750 to 800 F.  The particle surfaces of
coal are mildly oxidized, destroying the caking
properties. Heat  is  evolved  and  must  be
removed   to  control   the    temperature.
Pretreated char yield is about 90 percent of the
coal feed weight. Off-gas from pretreating
FV-7-4

-------
contains tars, tar acids, carbon oxides, sulfur
dioxide, water  vapor,  and  sometimes small
amounts of oxygen.

   In the utility  gas process, we propose  to
pretreat  coal  at  gasifier  pressure and  to
comingle pretreater off-gas with gases from
the gasifier. Hot pretreater char is fed directly
into the gasifier, thus avoiding thermal losses.
This design eliminates the waste-water  and
gas-treatment  problems associated  with  low-
pressure pretreatment. Furthermore, a smal-
ler pretreater  vessel is needed  for  pressure
operation.  The  overall plant  complexity  is
reduced and so is its capital cost.

   The  heat of reaction in the pretreater is
removed  by   generating   steam  in   heat
exchanger  tubes contained in the fluidized
bed. The amount of steam  generated is more
than the gasifier steam requirement.
THE GASIFIER

   We want to obtain rapid gasification rates
which will permit  a higher throughput for a
given reactor size and, therefore, will result in
a less costly  plant. Rapid, precise control of
the  gasifier  operation  is  needed to  follow
changes in the power demand. Production of a
clean, low-carbon  ash is a primary economic
consideration.

   The gasifier is designed to gasify coal with
air   and  steam   in   a  fluidized   bed.
Simultaneously, the coal ash is agglomerated
into larger and heavier  particles for selective
separation from the bed. The principle of ash
agglomeration and separation was discovered
by A. Godel1 and developed into the Ignifluid
boiler. The concept was described by Jequier
et al.2-3  following  laboratory and pilot plant
gasifier development  at Centre d'Etudes et
Recherches des Charbonnages de France. We
have adapted and modified the Jequier  design
in development of this reactor  concept.
   We   call   the   gasifier    the    ABR
(Agglomerating  Bed  Reactor).  The   ABR
concept resolves the main disadvantage of coal
gasification in a fluidized bed rich in carbon:
How can  low-carbon-content ash be selectively
removed  from the  bed?  Advantages of fluid-
bed gasification  are retained. These are:
   1. Bed  temperature can be  uniformly and
;.,.,   readily controlled.

   2. High  reaction   rates  can  be  attained
     because of excellent gas-solids contact and
     large surface area of the solids.

   3. Coal fines from mining and crushing can
     be used in the  feed.

   4. The  mass  of carbon in  the fluid  bed
     ensures reducing conditions at all times.

    The ABR is fluidized  by a mixture of air
 and steam. Gasification takes place at about
 1900F  in the  fluidized bed.  Part  of the
 fluid izing gas enters through a  grid which is
 sloped toward one or more cones contained
 in the grid. Heavier particles migrate along the
 sloped grid toward the cones.  The rest of the
 fluid izing gas  flows upward at high velocity
 through the throat at the cone apex, creating a
 submerged jet within the cone. The tempera-
 tures  generated  within the jet are somewhat
 greater than in the rest of the bed. As carbon
 is gasified in and near the jet, ash is heated to
 its softening point.  The  sticky ash  surfaces
 cling  to  one another, and ash  agglomerates
 grow in the violently agitated jet. When heavy
 enough, the agglomerates fall counter to the
 high-velocity gas in the throat  and are thus
 separated from the  fluid bed.

    To protect the ash lock hoppers from the
 hot agglomerates, they are filled with Water
 which is boiling from the heat contained in the
 ash. The steam generated reduces the amount
 of external steam needed for the ABR. When
 filled, the hopper is  flushed into filters  to
 recover a wet cake  of ash for final disposal.
 Filtered water returns to the  lock hoppers.

 TAR  AND DUST REMOVAL

    Above the ABR fluid bed we have designed
 for a gas residence time of 10 to 15 seconds;
 the gas temperature will be between 1500 and
 1900F. By "soaking" the gas at high temper-
 ature, tars and oils which may be evolved are
 thermally  cracked   to  gas   and   carbon.
 Elimination or reduction of tarssand oils in the
 raw gas will reduce heat exchanger  fouling
 and    will    simplify   by-product   and
 waste-stream cleanup and treatment.
                                                                                   IV-7-5

-------
    Most of the dust contained in the gasifier
 gases is removed by cyclone separators and
 returned directly to the ABR bed. Very fine
 dust is separated in the second stage of dust
 removal and is  returned  to the gasifier  by
 injection  beneath the gasifier  cones. Within
 the cones the carbon contained in the fine dust
 is gasified. The fine ash sticks to the  heavy
 agglomerates and is removed from the system.
 Although  cyclone  separators  are shown  in
 Figure 2,  we plan to investigate other high-
 temperature solids separators. We will provide
 space and plant flexibility in  the  utility gas
 pilot plant for large-scale testing of alternative
 separators. Efficient removal of hot dust is
 important in retrofitting  utility gasification
 plants to  existing boilers to prevent erosion,
 contamination,  and plugging in the raw gas
 heat exchangers. When utility gasification is
 applied to combined-cycle power generation,
 even  greater gas  cleanliness  is needed  to
 protect the gas turbines.


 SULFUR  REMOVAL SYSTEM

    Most of the sulfur produced by coal gasifi-
 cation  appears  in  the form  of hydrogen
 sulfide. Because the pilot plant produces low-
 pressure, low-temperature fuel for a boiler, we
 can use the Stretford or  a similar process for
 product gas sulfur cleanup. The  Stretford
 process is  commercial;  it is  effective  when
 scrubbing low-pressure gas; it can produce a
 cleaned gas  containing  as little  as 1  ppm
 hydrogen  sulfide; and  the process converts
 hydrogen sulfide directly to elemental sulfur,
 avoiding the  need for a Claus plant. For this
 application the  Stretford process is easy to
 operate and  is inexpensive.

   Figure  3 is a simplified flow diagram for
 the Stretford process. The scrubbing liquor is
 an  aqueous  solution of sodium  carbonate,
 sodium vanadate, and ADA (the sodium salt
 of anthraquinone 2:7 disulfonic acid).  Gas
enters the  scrubbing tower at less than 140F
 and usually at a pressure of less than 75 psia.
 We have selected 120F  and 25 psia as pilot
 plant conditions. Absorbed hydrogen sulfide is
oxidized by the solution to fine, suspended
sulfur particles.  After  completion  of  the
oxidation reaction, the reduced solution is re-
oxidized by blowing with air at atmospheric
pressure. The fine sulfur concentrates in the
froth during air blowing and is collected from
the solution as an elemental  sulfur product.

   The preferred  system for  sulfur removal
may  change  depending  on the gasification
plant capacity and the final use for the gas.
For large gasification plants sulfur removal at
high pressure using processes such as Selexol,
Purisol, or Alkazid may  be more economical
than   low-pressure   sulfur  removal.   In
conjunction with combined-cycle plants, yet-
to-be-developed,   high-temperature   sulfur
removal  is desirable for  improved  plant
efficiency and for decreased cost.  For small-to-
medium-sized power  plants backfitted with
coal gasification systems, we believe that  the
Stretford process  will be widely applied  for
sulfur removal from the gas.

ENERGY RECOVERY

   Raw  gas   leaves  the   gasifier   at   a
temperature between 1500 and 1700F and at
a pressure of 300 psi or higher. The main  gas
flow from this section of plant enters the sulfur
removal system at 25 psia and 120F. Sensible
heat contained in the gas represents about 20
percent of the heat available from combustion
of the coal feed. The energy recovery section
(Figure 4) is designed to recover as much of
this energy as possible. Since most of the heat
recovered must be used in the power cycle,  the
design of this  section of the plant will  be
strongly influenced by the heat levels that can
be used in the power cycle.

   After withdrawing from 1 to  5 percent of
the main gas stream for use as lock hopper
pressurizing gas, the pressure of  the main  gas
stream is broken by expansion through a  gas
expander. We want the expander exhaust  gas
condition to be suitable for feeding directly to
the Stretford scrubbing tower. In our design,
the condition is at  25 psia and 120F with  the
gas water saturated. Condensation should  not
IV-7-6

-------
occur in the expander. Having defined the gas
condition at the expander exhaust, the desired
moisture content of the gas is obtained by ad-
justment   of  the  main   gas   separator
temperature.  With a 300-psia gasifier,  this
temperature  is 224F;   with   a   1000-psia
gasifier, it is 300F. The  gas expander inlet
temperature  is adjusted  to  give the  desired
exhaust temperature. Power recovered by gas
expansion  is  used to  drive  the gasifier air
compressor.  The  air  compressor   power
requirement is  about 19 bhp/ton-day of coal
feed in a 300-psi plant and 26 bhp/ton-day in
a  1000-psi plant.
   Condensed  water  from  the  main  gas
separator   and from the lock  hopper gas
separator  is fed into a low-pressure stripping
tower to remove dissolved gas. The stripped
gas  rejoins the main  gas flow  entering the
Stretford  scrubbing tower.  Stripped waste
water is cooled and sent  to either  biological
treatment  or  active  carbon  treatment  for
phenol  removal.

   The advantages of the proposed design are:
(1) the  gasifier air compressor  is driven by
process energy, (2) the main  gas stream is not
water-cooled to obtain the desired 120F, (3) a
minimum of waste water is produced, and (4)
the conditions  of heat recovery to the power
cycle are well defined so  that an efficient
recovery system can be designed.
 COOLING    WATER
 TREATMENT
AND    WASTE
   Cooling water requirements are minimal in
 this utility gas plant design as a result of some
 of the process choices that were made. These
 are: (1) use of a dry feed system, (2) pretreat-
 ment at pressure, (3) gas expander exhausting
 at 120F, (4) gasifiying at higher pressure, and
 (5)  use of the  Stretford process  for  sulfur
 removal. In addition, air cooling is used where
 feasible.

   Waste-stream     treatment     problems
 associated with  coal gasification plants can be
 serious.  We  reduced the  severity of these
                     problems  by choosing the dry feed system,
                     pressure pretreatment, single-stage high-tem-
                     perature gasification with ash agglomeration,
                     raw gas "soaking" at high temperature, and
                     the Stretford process.

                     SYSTEM PERFORMANCE

                        Table  1  shows   the   calculated  plant
                     performance for 300-psi and  1000-psi utility
                     gasification  plants. Note that although more
                     methane is formed in the 1000-psi gasifier, the
                     product gas heating  value  is  not greatly
                     different for the  two cases. Higher-heating-
                     value gas  can be produced if the  gasifier is
                     designed as a two- or three-stage unit  with
                     gasification  temperatures increasing progres-
                     sively in each stage. We do not believe that this
                     increased  complication is  warranted.  Also,
                     higher-heating-value  gas can be made if the
                     pretreating unit off-gas is diverted from the
                     gasifier. This  will involve additional  plant
                     equipment for condensing, separating,  and
                     waste-stream treating of the off-gas and its
                     components. This, too, is considered to be an
                     undesirable  plant addition.

                     Table 1. CALCULATED PLANT PERFORMANCE
Gasifier pressure, psia
Product gas
Heating value, Btu/scf
Composition, vol %
CO
C02
H2
H?0
CH4
N2
Total
Thermal efficiency, %
To all products
To gas only
To steam only
300
Wet
140

17.8
9.2
12.1
8.5
4.3
48.1
100.0
Dry
153

19.4
10.0
13.3
__
4.7
52.6
100.0

86.5
73.2
12.7
1000
Wet
150

12.5
13.6
11.6
8.5
7.1
46.7
100.0
Dry
164

13.7
14.8
12.6
__
7.8
51.1
100.0

88.2
73.9
12.6
                       The thermal efficiency of the plant is very
                     good in both cases, with the 1000-psi plant
                     being slightly higher. The thermal efficiency is
                     determined  by comparing product  heating
                     values, heat contained in net steam, etc., to the
                     heating value of the coal feed. Compare the 86
                     percent efficiency of the utility  gas process
                     with the 66 percent efficiency of coal gasifica-
                     tion plants producing synthetic pipeline gas.
                                                                                   IV-7-7

-------
    The  gasification  plant   output   should
 respond at nearly the same rate as the power
 plant responds to electrical demand changes.
 The  ABR gasifier design should be operated
 nearly all of the time; its temperature should
 not  fluctuate  drastically or serious  internal
 refractory damage may result.  However, the
 throughput can be cut back significantly, and
 in this way  the  plant  may serve as  a load-
 following plant in addition to its use  in base-
 load  operation.   If the  plant  is  initially
 operating at design capacity, the first  move in
 decreasing output is to decrease steam and air
 flow to  the  gasifier. This  flow change can
 reduce plant output by a factor  of 3 or 4 very
 rapidly.  If a further reduction is needed, the
 ratio of air to steam  is reduced, causing a slow
 decline in gasifier temperature.  By lowering
 this temperature from 1900 to 1300 or  1400F,
 an additional reduction in output of 10 to 20 is
 obtained. The  gasifier temperature should be
 held within  a few  hundred  degrees of its
 normal  operating temperature  most of the
 time to avoid thermal shock and consequent
 cracking and spalling of the gasifier  internal
 refractory. Therefore, if the boiler plant is shut
 down temporarily, we prefer that the  gasifier
 operation continue at minimum rates and that
 the produced  gas is  burned in  a flare or
 burning  pit.
CONCLUSIONS
   We have described the preliminary design
of a coal gasification plant for manufacturing
clean utility fuel  gas  for the electric power
industry.  Its operation will demonstrate the
economics and reliability of such a plant when
used to fuel an existing boiler. Modifications
in this design will adapt the gasifier concept to
combined-cycle power generation and to the
manufacture of clean fuel for other industrial
uses.
   We expect  that the  plant  cost  and  the
product energy cost will be less by  a significant
amount than those costs for an equivalent-size
synthetic  pipeline gas  pla,nt. Proof of this
design will provide electric utilities with a
realistic method for conversion  of coal  to a
clean fuel.
REFERENCES

1. God el,  A. A New Combustion Technique.
   Eng. Boiler House Rev.  71:145-153, May
   1956.
2. Jequier, L., L. Longchambon, and G.  Van
   de Putte. The Gasification of Coal Fines. J.
   Inst.  Fuel.  33:584-591. 1960.
3. Jequier, L.  et al.  Apparatus  for  Dense-
   Phase Fluidisation, U.S. Patent 2,906,608.
   September  29, 1959.
IV-7-8

-------
                     LP
                   QUENCH
                                     VENT GAS
                  PRESSURIZING GAS
                    COAL
                  CRUSHING
                                       1
                          LOCK
                        HOPPERS
  GASIFIER
PRODUCT GAS
     SULFUR
   GAS
  CLEANING
(STRETFORD)
                                        EXPANDER
                                                                           ~1
   HEAT
 RECOVERY
                   WASTE-
                   WATER
                 TREATMENT
                            OIL
                        FABILIZA
                            AND
                         STORAGE
TION
E



    AIR
COMPRESSOR
             I

             \ (FUTURE)
            REPLACE WITH HIGH-TEMPERATURE
            PARTICULATE AND SULFUR
            REMOVAL
                                                 ASH
                                               HANDLING
                                                 AND
                                               DISPOSAL
    Figure 1,  Coal gasification pilot plant block flow diagram for clean utility gas.
                                                                               IV-7-9

-------
                                                             RAW GAS TO
                                                             TREATING
  COAL REED
LOCK HOPPER
 PRETREATMENT
 FOR BITUMINOUS
 COALS
    STEAM
    GENERATION
                                                DUST      \  /
                                                REMOVAL    W
                                                                2ND STAGE
                                                                DUST
                                                                REMOVAL
        AIR AND STEAM
                                                                SOUDS FEEDER
AIR AND
STEAM
                                             ASH LOCK HOPPER
                                             (WATER - FILLED)
                          Figure 2. Agglomerating bed reactor.
IV-7-10

-------
                     CLEANED GAS.
GAS FEED.
                      SCRUBBING
                       TOWER
                                AIR
                                                               SULFUR
                        Figure 3.  Stretford process.
                                                                          IV-7-11

-------
RAW GAS FROM GASIFIER ^
1500  1700 F, 1
HIGH PRESSURE /^
HEAT /
EXCHANGER V

t
HEAT
RECOVERY
TO
POWER
CYCLE ^.


\.
GAS
^XPANDER
DM ^^^^^^"
JL
S^\ GASIFIER
K A AIR
\^J/ COMPRESSOR
J _^ 	
^~~y* s~ 	

M^ i^nni co
^\^ IrUULtK
MAIN
GAS
SEPARATOR
GAS-TO-SULFUR
25 psia,
WATER SATURATED
, 	 GAS TO LOCK
l*~ HOPPER, 120 F
1 HIGH PRESSURE
TOR J
l^ WATER TO
^^ STRIPPING, 120 F
                                                               WATER TO STRIPPING
                                                                  200-300 F '
                           Figure 4.  Energy recovery section.
IV-7-12

-------
SESSION V:
  Pilot Plant Design, Construction, and Operation

SESSION CHAIRMAN:
  Mr. H.B. Locke, National Research Development
                              V-0-1

-------
     1. THE DESIGN, CONSTRUCTION, AND OPERATION

            OF THE  ABINGDON FLUIDISED BED  GASIFER

                     G. MOSS AND D. E. TISDALL

                    Esso  Research Centre, England


ABSTRACT

  A detailed description is given of the design and construction of the desulphurising fluidised-
bed gasifier which was built and operated at the Esso Research Centre at Abingdon under the
terms of GAP Contract CPA  70-46.

  The unit was operated under gasifying conditions for a total of 450 hours during the commis-
sioning  period. Information is  presented concerning the  operational problems. which were
encountered and the remedial steps which were taken.

INTRODUCTION

  The information in this paper  supplements   combustion of fuel oil. The Abingdon gasifier
that given in reference 1, which provides pro-   is the first of its kind and incorporates a num-
cess data relating to the retention of sulphur in   ber of unique features which were designed
fluidised beds of lime during the in situ partial   specifically for pilot scale operation.
                                  V-l-1

-------
 THE DESIGN AND CONSTRUCTION OF
 THE GASIFIER

    Three  primary decisions  determined the
 size, configuration, and mode of operation of
 the gasifier. These were  as follows:

  1. The internal configuration of the regener-
     ator of the gasifier was to  be  identical
     with that  of the batch units which had
     been used in the exploratory phase.

  2. The gas produced was to be burned within
     a standard packaged boiler fitted with a
     suitably modified combustion system.

  3. A monolithic form of construction was to
     be used in which all  vessels and transfer
     lines were to be formed as cored holes in a
     solid block of refractory  concrete.

    The first decision eliminated  one area of
 uncertainty  because it  was known  from
 experience that the  batch units functioned
 satisfactorily under regenerating conditions. It
 also set a limit to the capacity of the unit based
 on  what  was  then  known  concerning  the
 capacity of the batch units. At the time that
 the decision was made these had only been
 operated at a superficial gas velocity  of  4
 ft/sec. At an assumed SO2 concentration of 10
 percent by volume in the regenerator off-gas,
 this gave a sulphur handling capacity of about
 8.6 Ib/hr. In  the case of  a  2.2  percent by
 weight sulphur fuel,  oil this limited the fuel
 throughput  to  391  Ib/hr, giving an  energy
 throughput of 7.1 x 106 Btu gross/hr or 6.7 x
 10 6 Btu net/hr. The gasification of 391 Ib fuel
 oil/hr at 900 C and a gas velocity of 4 ft/sec
 with 20 percent of stoichiometric air indicated
 a cross-sectional area for the gasifier of 4.8 ft2.

   The decision to  use  a  packaged boiler
 rather than  a  flare  for the second  stage  of
 combustion was influenced by a  number  of
 considerations. A flare might have had an
 unfortunate impact on local public relations;
 it was in any case considered desirable to use
 the gasifier to fire a standard  piece of equip-
 ment in order to demonstrate it as a  practical
 proposition. The capacity  of the boiler which
was available was 10 x 106 Btu/hr delivered as
pressurised hot water.  This appeared to be
quite suitable for use as an energy sink for a 7
x  106  Btu/hr  gasifier;  it was decided to
dissipate the output to  the  atmosphere via a
pressurised heat exchanger and  an  atmos-
pheric  evaporative  cooler.  What  was  not
realised at the time, was that it was possible to
operate the batch units at up to 8 ft/sec super-
ficial gas velocity. Consequently, although on
the original  design  basis  the  heat disposal
equipment provided a reasonable margin of
spare capacity, it subsequently turned out that
heat disposal was a factor restricting the range
of  operating  conditions  which  could  be
explored.

   The use of the packaged boiler enabled flue
gas recycle, to be used to  control the gasifier
temperature to levels lower than those dictated
by  adiabatic operating  conditions. This was
advantageous since steam would most likely
have  been used for  this  purpose  in other
circumstances, involving the use of  an addi-
tional utility.

   So far as  the gasifier was concerned,  the
choice of monolithic construction imposed its
own logic upon the geometrical  configuration.
The  reasons   for  choosing  monolithic
construction  were: (1) the absence  of joints
between the various components enabled them
to be grouped in a very compact arrangement;
(2) the position of the bed transfer ducts in the
heart of the block enabled the sensible heat of
the transferred bed material to be conserved to
the same degree as it would be in a large scale
unit; and (3) this was a very simple and cheap
form  of construction.

   The  major disadvantage with monolithic
construction  is that it does not readily lend
itself to modification. It was necessary to be
sure  that   transfer   lines   would  work
satisfactorily  before they were cast in  refrac-
tory concrete. An additional disadvantage was
the reliance on the durability of the  concrete
when subjected to  temperature stresses. In
fact some minor cracks  did develop but these
were found to be self-sealing under  gasifying
conditions.
V-l-2

-------
   When the choice of monolithic construc-
tion was made it was found expedient to adopt
a  rectangular  cross  section,  because this
enabled the casing to be constructed from flat,
edge-stiffened panels. It was then found that
the overall height of the unit was dependent on
the size of the gasifier cyclone; the decision
was finally made to use two smaller cyclones
instead of  one  large one.  This  gave  the
configuration shown in Figure 1. It can  be
seen that the gasifier bed is  rectangular in
plan and that the regenerator  is between the
two gasifier cyclones to one side of  the
gasifying reactor.

   The bed transfer system which was selected
was an adaptation of a system proven in com-
mercial practice, where  it  is  employed  to
control the rate  at  which  bed  material
descends through a series of stacked beds. The
adaptation was necessary in order to enable
the  system to be utilised to  transfer bed
material between two beds in parallel and at
the same height.  In this system the bed trans-
fer duct is almost vertical, but  incorporates a
horizontal  section  at  its  lower end.  In  the
absence of any external agency the duct simply
fills with static  bed material. When a pulse of
gas is introduced into the horizontal section,
however, bed material flows down  the duct
under the influence of gravity. The adaption
involved utilising the two fluidised beds as lift
pumps for the bed material,  each bed dis-
charging the material into a cavity at the top
of a transfer duct. The two transfer ducts can
be seen in dotted outline in Figure 1 and they
are situated on either side  of the regenerator in
plan.  The existence of these  transfer ducts
adjacent to each of the  cyclones suggested that
they might also be used as  cyclone drains. This
was in fact done  and in operation half of the
gasifier cyclone  fines are returned  to  the
gasifier itself and  half are passed on  to the
regenerator. The  fines  leaving the regenerator
bed are trapped  by an external  cyclone and
drained from the system.

   It was necessary to test  a novel bed transfer
system of this complexity before casting it in
concrete. This was done by building the full
scale cold rig illustrated in Figure 2. The bed
material used in this rig was a crushed brick of
roughly the same density and size distribution
as the lime. Although the  gas velocities were
matched in the fluidised beds, there  was a
strong element of conservatism in the  opera-
tion of the transfer  system; no allowance was
made  for  the  very considerable  volumetric
expansion  of injected gas  under  hot  condi-
tions. It was also anticipated that the  higher
gas viscosity at the normal operating tempera-
ture would improve the flow properties of the
bed material.  Tests run with  this rig soon
showed that the  introduction of the cyclone
fines   into  the  transfer   ducts  via   simple
branches led to severe bridging problems.  It
was deduced that this bridging was caused by
fines being blown back up the transfer lines to
block the interstices between the particles of
descending bed material.

   A way out of this difficulty was found by
devising the mixing cavity shown in Figure 3.
The shape of this cavity is such that a pocket
of gas  is trapped  within  it, leaving  a  free
surface of bed material above the horizontal
section of the transfer duct.  The theory behind
this design is that when the  transfer duct  is
activated, bubbles of gas rise into the pocket
and displace gas already there; this gas re-
enters  the free surface of  bed material  and
proceeds up the transfer duct. In this way the
free surface can act as a filter for fines. In
practice the device  which was built of plexi-
glass worked very well. Lenses of fines were
trapped  by the  coarser  solids,  and  these
inclusions moved down towards the horizontal
transfer duct where they  disappeared. This
modification solved the bridging problem.

   Other  design problems  related to  the
thermal expansion   of the refractory  block
when   it   was   brought  up   to  operating
temperature. There was also the question of
thermal insulation   to be dealt  with.  The
construction of the unit allowed the transverse
thermal expansion  in  a very simple fashion.
The plates forming the casing of the gasifier
were lined with 3 in. thick slabs  of 50 lb/ft3
castable insulating material. When the casing
                                                                                     V-l-3

-------
 was  assembled the inner surface was  lined
 with 1/4 in. thick expanded polystyrene sheet;
 the refractory concrete block was cast within
 this  lining.  The  expanded  polystyrene  was
 subsequently  melted  out  leaving  a   gap
 between the refractory block and the insula-
 tion. The vertical expansion of the block posed
 more  difficult  problems  because  it   was
 necessary to connect the gasifier to the burner
 in a gas-tight manner, but which allowed 3/8-
 in.  relative  vertical  movement.  Since there
 were two gasifier outlets it was necessary to
 employ a Y-shaped bifurcated duct consisting
 of a mild steel casing enclosing a refractory
 concrete lining in which the  gas  passages are
 cast (Figure 4). A layer of calcium silicate slab
 insulation   is   sandwiched   between   the
 refractory concrete  and  the  steel casing to
 minimise heat losses.

    The duct is suspended immediately above
 the  cyclone   outlets  with   an  expansion
 allowance between  the  corresponding faces.
 The  expansion joint is sealed by a stainless
 steel bellows to provide a gas-tight assembly.
 The duct is supported at the other end  by a
 roller which   accommodates the horizontal
 expansion of the duct and burner; it is sealed
 to  the  boiler face  with   a  compressible
 insulating seal.

    The   burner  design  presented   some
 problems because there was  no information
 available concerning the combustion charac-
 teristics of this hot gas. Some simple burners
 had been tested on the early batch units which
 showed that the gas would  burn easily. By
 introducing some premix air to the burner  it
 was possible to burn the gas with  a steady
 smoke free flame and low excess air.

   The  7 x 106 Btu/hr  burner used for the
continuous gasifier is illustrated in Figure 5. It
consists  of two sections--a premix zone  in
which about 10 percent of the combustion air
may be  introduced  and  a  main section  in
which the balance of the air is added. In both
of  these  sections an  inner  stainless   steel
assembly is used which is insulated from the
cold combustion air  introduced  around the
assembly. The insulation is necessary to main-
tain the temperature of the hot gas duct which
otherwise might become obstructed with con-
densing material from the hot gas.

   These inner insulated assemblies are fixed
to the outer casing at one end and are free to
expand along  the axis  of the  burner  at
operating temperature.  The  gas  issues from
the burner to mix with the main combustion
air through a stainless steel orifice sized to give
a pressure drop between 3 and  4-in. water
gauge.

   For safety reasons it was decided to use a
continuous  pilot flame  on the experimental
plant; here problems arose because this flame
must be stable at normal running conditions
which means a high gas and air rate, unlike a
conventional burner which lights off a pilot at
a  low  flame  setting.   The  problem  was
overcome  by placing a  small stainless steel
deflector plate to shield  the flame of the pilot
from the  main air.

   The only other problem with the burner
arose from gas turbulence at the  entry to the
burner orifice nozzle which threw  out deposits
in the burner inner duct. This was overcome
by smoothing the gas flow into the orifice by a
suitable entry  duct; no further  deposits were
observed.

   Because the overall height of the gasifier is
considerably greater than the centre height of
the boiler furnace, it was  necessary to place
the gasifier in a pit in order to line up the two
components. A general plant layout is shown
in Figure 6.

   The  air distributor  of  the   gasifier  is
provided  with  horizontal  nozzles made by
drilling six radial holes of 0.177-in. diameter
through each   of  32 stainless  steel  capped
tubes. The distributor and its plenum form a
removable box structure built of mild steel;
the top of the box,  from which the nozzle
assemblies project, is covered with refractory
concrete.
V-l-4

-------
   There  are  a  number  of  penetrations
through the walls of the gasifier to provide
access for thermocouples, manometer probes,
fuel  injection tubes, bed drains and the  gas
injectors used  to  activate the  bed transfer
system.  These  penetrations  were  all  cored
before the block was cast. In addition, a pre-
cast  quarl for  the startup  burner  was also
placed  in  position  before  the  refractory
concrete was cast.

   Figure 7 shows the lower core assembly of
the gasifier during an early stage of construc-
tion.  The cores for the transfer system were
made of plexiglass and were filled with wax in
order to avoid  the possibility of damage and
filling with concrete during the casting opera-
tion. Figure 8 shows the unit after the first two
lifts  had been  cast.

   G.R. Stein Refractories Ltd. advised on the
method of construction and  built the unit
using their refractory concrete Durax C.1600.
This  material  contains  about  50  percent
alumina, 42 percent SiO2, 5 percent  CaO,
with  traces  of   Fe2O3   and   MgO.  The
maximum  operating temperature is 1600C
with  a melting point of 1710C.

   In the first instance the only internal metal
components of the gasifier  were the cyclone
outlet tubes. These proved to be unsatisfactory
and  were subsequently replaced  by tubes of
self bonded silicon carbide.

THE ASSOCIATED SYSTEMS

   The pilot plant flow plan is shown in Figure
9. It  will be seen  that the startup burner is
fired by propane. This burner is used to heat
the gasifier to its working temperature and  is a
standard commercial burner with a variable
output,  delivering 700,000  Btu/hr at  the
maximum firing rate. Propane is also used to
fuel the pilot burner fitted to the boiler. The
main fuel system utilises three small metering
pumps which draw fuel oil from a circulating
stream in a ring main  and deliver it to  the
three fuel injectors of the gasifier. A  small
amount of air is injected with the oil in order
to prevent coking in  the  injector tubes. A
switch from  fuel  oil  to  kerosine is  also
provided.  This enables the consumption of
propane  during the warm-up  period  to be
reduced while  avoiding  the  introduction of
sulphur into the gasifier. The flue gas recycle
system which is used to control the tempera-
ture of the gasifier is also shown in the flow
plan. The flue gas recycle  stream is first
cleaned in a cyclone, then passed through an
orifice plate flowmeter and a control valve to
the inlet side  of the first gasifier blower. A
second control valve throttles the air supplied
to this blower;  by making suitable  adjust-
ments to these two valves it is possible to vary
both the total supply of gas to the plenum of
the gasifier and the composition of the gas in
terms of the proportions of flue gas and air
which it  contains.
   The operating temperature of the regener-
ator tends  to  be self regulating  when  no
oxygen is present in the tail gas, because CaS
can  yield  two  oxidation  products  CaSC>4
and  CaO + SO2. The first  reaction releases
much more heat than the second reaction; but
as the temperature rises the  second reaction
tends to predominate so that  in  effect the
calorific value of the sulphur fed to the regen-
erator tends to fall.  In  order to  hold the
temperature at a specified level however the
bed  transfer system  is  arranged  to auto-
matically increase the transfer rate when the
temperature falls below the set point. Because
of the temperature difference between the two
beds, the consequent adjustment to the rate of
heat transfer  from  the regenerator to the
gasifier  brings  the regenerator temperature
into line. The rate of SO2 release  at any  set
regenerator temperature depends on the rate
at which air is fed to  the regenerator. For
experimental  purposes  this  is  a  manual
adjustment;  when   completely   automatic
control is used, the air rate to the regenerator
may be controlled to hold  the O2 concen-
tration in the regenerator tail gas at a constant
level. The return of solids from the regenerator
to the gasifier  is controlled by the pressure
drop across the regenerator bed.
                                                                                    V-l-5

-------
    Equipment was not installed to deal with
 the 10 percent SC>2 stream from the regenera-
 tor because this did  not form part  of  the
 development  programme at this  stage. The
 SO2 stream from the pilot plant was therefore
 connected into the boiler stack, thus creating a
 flue gas identical to that resulting from direct
 combustion of the test fuel.

    The   instrument   flow  plan  for   the
 installation is shown  in Figure   10.  This,
 however,  does not show  the packaged pres-
 surisation system which  maintains constant
 boiler water pressure  and temperature.

    Considerable attention has  been given to
 safety measures; the plant is protected by a
 number of sensor systems which detect both
 hazards  and   conditions.  Signals  from  the
 sensors will,  according to a predetermined
 selection, either shut  down the whole plant
 and operate an alarm, or give an alarm and a
 visual indication of the trouble, or merely give
 a visual indication.

    During test runs the unit is  operated on a
 24-hour shift basis, with one professional and
 two non-professional personnel in each shift
 team.

 OPERATIONAL PROBLEMS

    Three    operational   problems  became
 apparent  during the first  attempts to run  the
 unit for a prolonged period:

  1. The formation of lime/coke deposits in the
    ducting between  the gasifier and  the
    burner.

  2. A tendency to form too large a  proportion
    of CaSO4 within the  regenerator.

  3. Poor containment  of bed fines.

   The  deposits  in the  gas  ducting were
 formed locally in areas of high turbulence. In
 the first instance deposits tended to choke  the
 inlet ports of the cyclones which were origin-
 ally square  edged. This problem was greatly
 alleviated by chamfering these edges to give a
 smoother  flow transition. Two other  critical
 areas were found where the vertical cyclone
exit channels intersected the converging hori-
zontal channels within the Y-shaped duct. The
design of the ducting has since been modified
to provide a smoother  gas passage at these
elbows.
   The deposits laid down in the gas ducting
could be removed without shutting down the
unit by a controlled burn-out  procedure.  A
more serious form  of deposit was, however,
found within the  regenerator and within the
transfer  line from  the  regenerator  to the
gasifier. The deposits in these areas built up
relatively slowly, but could only be removed
when the unit was shut down. It was deduced
that the  cause  of these deposits, which con-
tained no carbon,was the excessive formation
of CaSO4 within  the regenerator.
   The oxidation  of CaS in the regenerator  is
analagous to the  oxidation of carbon in the
gasifier in that two  reactions  occur and they
appear to occur sequentially. In  the gasifier
the carbon  deposited  on  the  lime is first
oxidised  to  CO 2  which  is  subsequently
reduced to a large extent to  CO on passing
through the bed. Within the regenerator there
is a tendency to form CaSO4  near the  distri-
butor which subsequently reacts with CaS to
form CaO + SO 2.  It has been observed by
Curren,  Fink  and  Gorin2 that  during the
course of this reaction  a transient liquid  is
formed which can cement particles together.
This is thought to be the mechanism by which
deposits were laid  down within the regenerator
and  in the  regenerator to gasifier transfer
duct. These deposits were largely composed of
calcium sulphate; a photograph of the deposit
which was removed from the regenerator at
the completion of Run 3 is shown in Figure 11.
This deposit was wedge shaped and grew from
the wall of the regenerator opposite the outlet
of the bed transfer duct from the gasifier. As
can  be seen in  Figure 12  the  distributor
nozzles  under  the  deposit were  themselves
blocked by coarse bed material; it seems likely
that the  initial  presence  of a  dead  zone
induced deposit formation  in this area.  The
design of the regenerator distributor has since
been modified and it is now more similar to
that of the gasifier.
V-l-6

-------
   The formation of deposits  in  the  regen-
erator  is  best   avoided  by  reducing the
tendency to form CaSC>4. This may be done by
introducing the bed material from the gasifier
into the regenerator at a level well above that
of its distributor.
   The effect of this change  would  be to
ensure that both the fresh bed  material from
the gasifier and the incoming air are brought
up to the working temperature before  they
meet. Under these circumstances there should
be a greater .tendency to form unstable CaSO3
in a single step; less CaSC>4  should be avail-
able for the liquid phase decomposition. In the
case of the experimental  unit  this has been
attempted by lowering the level of the regen-
erator distributor.
   The gasifier cyclones were not found to be
very effective in retaining the fines which were
produced by attrition and decrepitation. In
normal operation they passed solids at a rate
amounting to about 2 percent by weight of the
fuel used. It is  likely that this  poor  per-
formance was largely due to the rough surface
finish of the cyclones, but it was  considered
that the  best  way to prevent solids  from
entering  the  boiler would be  to reduce the
amount of fines produced.
    Decrepitation  occurs  during  calcination
and may be reduced by lowering  the bed
replacement  rate. The batch results indicate
that increasing the depth of the bed will not
impair desulphurization efficiency. An indi-
cation of the importance of bed losses incurred
shortly after the entry of stone into the gasifier
was the  fact that at the  end  of Run 3 the
vanadium content of the bed was three times
higher than  could  be accounted for on the
assumption of a  uniform bed  life.
   There were two  mechansims which might
have contributed  to this effect-decrepitation
and elutriation during addition. During these
runs the bed material was fed into the gasifier
through its lid and fell in a  stream  past the
entry of one  of the cyclones. It was thought
likely that some  of the finer  material  was
swept out of the unit before it entered the bed.
In order  to reduce  this possibility the stone
feed system was modified; stone now enters
the unit through the wall of the gasifier in the
vicinity of the  bed surface  and at a  point
remote from the cyclone inlets.  Attrition  in
fluidised beds increases rapidly as the superfi-
cial gas  velocity  rises.  The  highest gas
velocities within the bed occur at the distribu-
tor nozzles and are required in order to ensure
bed  stability.   The  problem  here  was  to
combine bed stability with a low nozzle  efflux
velocity. The adopted solution was to use two
stage  nozzles  in which  the kinetic energy
imparted to the gas  by the  pressure drop
through the first nozzle was dissipated prior to
the low velocity entry of the  gas into the bed
via the second  nozzle.
   During the first phase of operation the unit
was run for  450 hours under  gasifying condi-
tions; the main structure  does not appear to
have suffered any significant  deterioration.  A
major objective of future work is to reduce the
quantity of fines leaVing  the bed to a  level
which will enable us to envisage the construc-
tion of a full  scale  unit which would, not
require hot  cyclones. If this could be  done
there  would be a  considerable saving on
investment and also an improved operability
due to the decrease  in  deposit  formation
arising from the simplification  of  the gas
ducting. Finally,   it seems   reasonable  to
conclude that the operational problems which
have so far been encountered do not appear to
be unduly severe or intractable.
REFERENCES

1.  Craig, J.W.T., G. Moss, J.H. Taylor,  and
   D.E.  Tisdall.  Sulphur  Retention  in
   Fluidised  Beds of Lime Under Reducing
   Conditions. (Presented at 3rd International
   Conference on Fluidized-Bed Combustion.
   Hueston Woods. October 29 - November 1,
   1972.) (See   Session   HI,  Paper  4   this
   volume.)
2.  Curren, G.P., C.E.  Fink,  and E. Gorin.
   Phase II Bench-Scale Research on C.S.G.
   Process,  Research  and Development  No.
   16. Consolidation Coal  Company,  Library,
   Pa. Prepared for Office of Coal Research,
   U.S.  Department   of  the  Interior,
   Washington, D.C. under Contract  Number
   14-01-0001-415. July 1969.
                                                                                    V-l-7

-------
                CONNECTION BETWEEN CYCLONESv
                EXPANSION BELLOWS TO
                ABSORB VERTICAL EXPANSION
                ON GAS OUTLET
                REMOVABLE LID
                                                                             BURNER
        FUEL  SUPPLY

   SECTION A-A

              BED
             DRAIN

   OUTER METAL CASING
   INSULATING REFRACTORY^

   CASTABLE REFRACTORY
   SECTION B-B
                           / / /// //// / ///////
3.25 ft
     FUEL
   SUPPLY   
                                                             CYCLONE FOR
                                                               GASIFIER
                                                                REGENERATOR
                                                                CYCLONE FINES
                                                                FED INTO BED
                                                                TRANSFER PIPES
V-l-8
                                            GAS PULSE

                        Figure 1. Layout of continuous gasifier unit.
                                                                                7.75ft
                                                      r? "I RETURN TUBES FOR
                                                          BED RETURN
                                                          FROM
                                                          REGENERATOR
                            m
                             JISTRIBUTOF
                                                                                4.75 It

-------
Figure 2. Full-scale cold rig.
                                                        V-l-9

-------
                                                     FINES FROM CYCLONE
           SOLIDS FROM GASIFIER
           DESCENDING SOLIDS
                                                                        GOOD LEG SEAL
                                                                 FLUIDISED REGION
                                                                INLET TO REGENERATOR
                         Figure 3. Successful cyclones fines return zone.
V-l-10

-------
                 SLAB INSULATION

               REFRACTORY LINING

            EXPANSION ALLOWANCE


                   BELLOWS SEAL


               REFRACTORY LINING
CYCLONE OUTLET CONNECTION
         GAS OUTLETS PIPES -
           BURNER

             \



1 1






1 1
           SECONDARY
               AIR
PREMIX
  AIR
                                                             SECTION A - A
                                              SUPPORT YOKES

                                             SUPPORT FRAME
                                     PLAN
                  Figure 4.  Cyclones to burner manifold.
                                                                         V-l-11

-------
                                                 r ."n \ . . j^'T'i "Tnrr v1i \ iTlrTl
                                                 m\\\\\\\\\\\\\ft\\
                                                  /Cv\\\\\\\\\\\\\\
                                                  *-' t*f^ * * ' f * t j * ' -*->'~
                                                  \\\\\\\\\\\\\\\\
                              SECONDARY

                                 AIR
                          Figure 5.. Main gas burner.
                                                           PRIMARY

                                                             AIR
V-l-12

-------
        REGENERATOR CYCLONE
    ELECTRICAL
     CONTROL
     CABINET
   16ft
 FUEL INJECTOR (3 off)
         AIR SUPPLY TO
         GAS I HER
                         GASIFIER
                             \
                                     \
   TO
CHIMNEY
                                                        COOLER
                                                      HEAT EXCH.
     BOILER
    AIR SUPPLY TO
    REGENERATOR
 CIRCULATING OIL
SUPPLY FROM
30- BY 9-ft TANK
                                     SIDE SECTION
ELECTRICAL
CONTROL
CABINET
SWING JIE
PILLAR C
1
 OJ 11 	

MECHANICAL r
CONTROL
CONSOLE
"" GASIFIER
IRANE , 1 	 Tj
/XB*
I/ \
1 x_
iS5
BLOWERS
FOR
GASIFIER

OFFICE
| HEAT

BOILER

rn c
EXCH. 1

Q




32
r

T
 ft uiru
                                     PLAN
                                                           STACK
                           Figures.  General plant layout.
                                                                                 V-l-13

-------
           Figure 7.  Gasifier lower core assembly during early construction stage.
V-l-14

-------
Figure 8.  Gasifier lower core assembly following casting of first  two lifts.

-------
                                                            COMBUSTION AIR BLOWER
                                                           LIME DRAIN
                                                                  REGENERATOR AIR BLOWERS
                                                                            N2 FOR SOLIDS
                                                                               TRANSFER
FUEL INJECTION
    AIR
    (3)
                    PROPANE FOR
                     START-UP
                               Figure 9. CAFB pilot plant flow plan.

-------
                             FLUE GAS
    BOILER GAS ANALYSIS
f/c


M
d/p
P/s
                                                                                                    -  S02 TO STACK
                                                                                           REGENERATOR GAS ANALYSIS
                                   BED DENSITY
                     BED LEVEL!
                      P/s  d/p   M   d/p
                                   M
                                 .nr.gr-
                         DISTRIBUTOR

                           A P
                                        L..
                                                  GASIFIER
TR
FREQUENCY CONTROLLER


MANOMETER


DIFFERENTIAL PRESSURE CELL


PRESSURE SWITCH


TEMPERATURE RECORDER


TEMPERATURE CONTROLLER
                                                              GASIFIER-REGENERATOR

                                                                    A?
                                                                  REGENERATOR INLET
                                                                  SOLIDS TRANSFER
                                                                                                             N2
  GASIFIER PLENUM
    GAS ANALYSIS
	STATIC PRESSURE LINE

\ TRANSMITTED SIGNAL

	MANUAL CONTROL


^"" "*t ?\^
A^J
^""" *






___l M

T
; !r
L

\IR



                                                                                             REGENERATOR OUTLET

                                                                                              SOLIDS TRANSFER
                                                                                             >
                                                                                                                AIR
                                                                                               FLUE GAS RECYCLE
                                   Figure 10.  CAFB pilot plant instrument flow plan.

-------
Figure 11.  Deposit removed from regenerator after Run 3.

-------
Figure 12.  Distributor nozzles blocked by coarse bed material.

-------
                   2. DESIGN OF FLUIDIZED-BED MINEPLANT
                         M. S. NUTKIS AND  A. SKOPP
                        Esso Research and Engineering
ABSTRACT

   Fluidized-bed combustion of coal offers potential both as an efficient compact combustion-
boiler system and an air pollution emissions control system. Esso Research and Engineering
Company, under contract to the Office of Air Programs of the Environmental Protection Agency,
has designed a  system  capable  of fluidized-bed  coal  combustion and desulfurization  with
continuous  limestone regeneration.

   The fluidized-bed miniplant will operate at pressures up to 10 arm and with an input of
approximately 6.3 x 106 Btu/hr. This energy input is equivalent to a power plant rated at 635kW
(0.63 MW).

   In the miniplant, combustion and heat transfer to tubes take place in a reactor containing a
fluidized-bed of limestone at 1500-1700F, providing good heat transfer and an efficient desulfuri-
zation reaction between the sulfur dioxide and limestone. The calcium sulfate produced during
desulfurization is transferred to an adjacent fluidized-bed reactor and contacted with a reducing
gas at 1900-2050F. This regenerates the lime for reuse in the combustor and produces a by-
product off-gas stream of concentrated sulfur dioxide. Thus, regeneration minimizes the limestone
feed requirements and the  calcium sulfate disposal problems.

   The fluidized-bed miniplant design incorporates a 12.5-in. ID combustor and a 5-in. ID regen-
erator vessel with continuous transfer of solids between these two refractory lined reactors. In the
combustor, the fluidizing air enters a plenum, passes through the distributing grid, up through the
fluidized bed of solids and the combustion products discharge through two refractory cyclones in
series.

   Superficial bed velocities and pressures in the combustor  and regenerator are automatically
controlled. The pressure differential between the two vessels can also be automatically controlled.
   Heat extraction and temperature control in the fluidized-bed combustor are accomplished by
vaporizing demineralized water in 10 independent loops located in discrete vertical zones of the
reactor. The water flows to these loops are controlled by valves whose positions change to maintain
bed temperature in  each of the zones.
   Coal and makeup limestone to the combustor are fed continuously from a system designed for
controlled solids feeding under pressure. Solids transfer between reactors and discharge of solids
from the system (i.e., from the regenerator reactor) are accomplished using a pulsed gas transport
technique controlled by pressure differentials across and between these fluidized beds.

                                        V-2-1

-------
 INTRODUCTION
    The fluidized-bed coal combustor provides
 a new boiler technique where coal  is com-
 busted in a bed of particles maintained in a
 state of fluidization by the air required for
 combustion. The  use of limestone or other
 suitable sorbent as the bed material in such a
 system permits the capture and  removal of
 sulfur dioxide simultaneously with the com-
 bustion  process.

    Fluidized-bed  boilers   (FBB)   offer  the
 potential of an efficient and compact boiler
 combustion  technique  also  capable  of pro-
 viding   pollution   control.  Some   of   the
 advantages and economic factors are:

  1. Capability to combust lower quality fossil
    fuels in a fluidized bed.

  2. Immersion of  the boiler tubes directly in
    the fluidized bed achieves improved  heat
    transfer  rates  compared to conventional
    boilers.

  3. The higher volumetric heat release rates in
    a  fluidized-bed combustor  will  permit
    reduced  boiler unit sizes.

  4. Since efficient  combustion can be achieved
    at comparatively low  bed  temperatures
    (i.e.,  1500-1700F),  boiler tube corrosion
    and fouling should be reduced.

 Pressurized fluid-bed combustion  offers  even
 greater benefits in size reduction,  efficiency,
 and load control.

    Within the fluidized-bed boiler, limestone
 is calcined  to lime  which reacts with sulfur
 dioxide and  oxygen in  the  flue gas to form
 calcium sulfate. When used on a once-through
 basis, relatively high limestone feed rates are
 required to the fluidized-bed boiler if sulfur
 dioxide removal in excess of 90 percent is to be
 maintained.

   In order to reduce  these high limestone
 feed rates, a  system was proposed by Esso
 Research and Engineering  whereby the cal-
 cium sulfate  formed would be regenerated
back to calcium oxide in a separate fluidized-
bed reactor (i.e., regenerator) by reaction with
a reducing gas at a temperature  of approx-
imately 2000F. The regenerated lime would
be  returned to the fluidized-bed combustor,
where  it would again react with the sulfur
dioxide.
   In  a study completed by  Esso for  the
National Air  Pollution Control Administra-
tion,1 the following essential features  of the
proposed regenerative-limestone FBB system
were demonstrated:

 1. Removal  of over 90  percent of the SO2
    formed by combusting  coal in  fluidized
    beds of lime.

 2. Reductive regeneration  of  the  sulfated
    lime to yield an off-gas containing 7 to 12
    mole percent SO2. This  is a sufficiently
    high concentration  to  permit  its  con-
    version to H2SO4 or elemental sulfur with
    conventional technology.

 3. Good  activity  maintenance  of the  lime
    cycled back and forth between combustion
    and regeneration. The make-up require-
    ment for fresh limestone in a commercial
    plant was estimated to be about 15 percent
    of  that required  for  once-through  use of
    this material.

   These experimental results were  obtained
at  atmospheric pressure conditions.  Since
completing this study, engineering and  cost
analyses carried out by Westinghouse2 for the
Environmental Protection Agency (EPA) have
indicated   a   much   greater  commercial
potential for  a pressurized FBB  system when
used in conjunction with a combined  gas-
s