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 . S. ENVIRONMENTAL PROTECTION A(;EN( V

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                                                PROCEEDINGS

                                                             OF

                                                       SECOND

                           INTERNATIONAL CONFERENCE

                                                             ON

                             FLUIDIZED-BED COMBUSTION
Sponsor: U.S. Department of Health, Education, and Welfare
       Public Health Service
       Environmental Health Service
       National Air Pollution Control Administration
U.S. ENVIRONMENTAL PROTECTION AGENCY
Office of Air Programs
Research Triangle Park, North Carolina 27711

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 The AP series of reports is issued by the Environmental Protection Agency to report the results of
scientific and engineering studies, and information of general interest  in  the  field of air pollution.
Information presented in this series includes  coverage of intramural activities involving air pollution
research and  control technology and of cooperative programs and studies conducted in conjunction
with state and  local agencies, research institutes, and industrial organizations.  Copies of AP reports
are available free of charge - as supplies permit - from the Air Pollution Technical Information Center,
Environmental Protection Agency, Research Triangle Park, North Carolina  27711.
                                    EPA REVIEW NOTICE


These proceedings have been reviewed by the Environmental Protection Agency and approved for pub-
lication.  The contents of this report are reproduced herein as received from the authors. Approval does
not signify that the contents necessarily reflect the views and policies of the Environmental Protec-
tion Agency, nor does mention of trade names or commercial products constitute endorsement or recom-
mendation for use.
                                    Publication No. AP-109
11

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                                       PREFACE

     The Second International Conference on Fluidized-Bed Combustion was held October 4-7,
1970, at Hueston Woods Lodge, RFD No. 1, College Corner, Ohio, under the sponsorship of the
U.S. Department of Health, Education, and Welfare's National Air Pollution Control Admin-
istration (NAPCA). Subsequent to the Conference, NAPCA's functions were transferred: first, to
the Air Pollution Control Office (APCO) of the newly created Environmental Protection Agency
(EPA); later to EPA's Office of Air Programs (GAP). All references to NAPCA in these proceed-
ings, therefore, now refer to OAP.

     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 4.

     The Conference proper, consisting of six sessions, got underway Monday morning with the
official  welcome  extended by OAP's Sheldon Meyers.  Following Mr.  Turner's introductory
remarks,  Prof. D.E.  Elliott  delivered  the  keynote  address,  "Exploiting  Fluidised-Bed*
Combustion."
                      /1
     The six-part Session I,  chaired by Mr. A.A. Jonke,  followed the theme, "Small-Scale
Development of Fluidized-Bed Combustion." Session II, held Monday evening, also consisted of
six presentations, this time on  the general topic, "Control of Combustion Pollutants;" Mr. Alvin
Skopp was chairman.

     Mr. J.W. Eckerd was chairman of Session III, titled "Gasification to Desulfurize Coal, "
which opened Tuesday's activities. This session consisted of seven presentations. "Conceptual
Design and Economic Feasibility" was the theme of both  Sessions IV and V. The former, chaired
by Dr. D.H. Archer, consisted of five papers presented Tuesday evening: the latter, by Mr. T.C.L.
Nicole, consisted of six papers presented Wednesday morning.

     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. R.P. Hangebrauck, the panel consisted of Prof.
Elliott, Dr. Archer, Dr. C.Y. Wen, Mr.  Shelton Ehrlich, and Mr. Henschel.

     All papers presented during the Conference are included  in these proceedings.
*In these proceedings, papers retain local spellings, where applicable.
                                                                                     iii

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                                     CONTENTS

                                                                                  Page

     Preface  	in

     P.P. Turner
     Introductory Remarks: The NAPCA-DPCE
         Program to Control Pollution from Stationary Sources   	O-l-l

     D.E. Elliott
     Keynote Address: Exploiting Fluidised-Bed Combustion   	O-2-1

                                      SESSION I

1.   D.F. Williams
     Pilot Plant Experiments at the Coal Research Establishment (CRE)   	   1-1-1

2.   N.H. Coates and R. L. Rice
     Fluid-Bed Combustion of Various U.S. Coals  	1-2-1

3.   S. Ehrlich
     Combustion of Carbon-Bearing Fly Ash in a Carbon-Burnup Cell	   1-3-1

4.   H.R. Hoy (Presented by D.F. Williams)
     Pilot-Plant Experiments at BCURA	   1-4-1

5.   R.L. Jarry, L.J. Anastasia, E.L. Carls, A.A. Jonke and G.J. Vogel
     Comparative Emissions of Pollutants During Combustion of
         Natural Gas and Coal in Fluidized Beds  	1-5-1

6.   J.J. Svoboda
     Ignifluid Contribution to Air Pollution Control	   1-6-1

                                      SESSION II

1.   D.C. Davidson and A. W. Smale
     The Retention of Sulphur by Limestone in a Pilot-Scale Fluid-Bed Combustor  ....  II-l-l

2.   R.D. Glenn and E.B. Robison
     Characterization of Emissions from Fluidized-Bed Combustion of Coal
         and Control of Sulfur Emission with Limestone   	II-2-1

3.   L.J. Anastasia, EX. Carls, R.L. Jarry, A.A. Jonke, and G.J. Vogel
     Pollution Control Capabilities of Fluidized-Bed Combustion   	II-3-1

4.   G.A. Hammons and A. Skopp
     A Regenerative Limestone Process for Fluidized-Bed Coal
         Combustion and Desulfurization  	II-4-1
                                                                                    v

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 5.   G.P. Curran, C.E. Fink, and E. Gorin
     Coal-Based Sulfur Recovery Cycle in Fluidized-Lime-Bed Combustion        . .  .     II-5-1

 6.   G. Moss
     The Fluidised-Bed Desulfurising Gasifier	      	        II-6-1

                                       SESSION III

 1.   G.P. Curran, C.E. Fink, and E. Gorin
     Production of Low-Sulfur Boiler Fuel by Two-Stage
          Combustion—Application of CO2 Acceptor Process	        .  III-l-l

 2.   H.L. Feldkirchner and F.C. Schora, Jr.
     Coal Desulfurization Aspects of the HYGAS Process	    III-2-1

 3.   A.J. Forney, SJ. Gasior, R.F. Kenny, and W.P. Haynes
     Steam-Oxygen Gasification of Various U.S. Coals   	      	  III-3-1

 4.   E.K. Diehl and R.A. Glenn
     Desulfurized Fuel from Coal by Inplant Gasification	     . .        .   . .  III-4-1

 5.   F.G. Shultz
     Removal of Hydrogen Sulfide from Simulated Producer Gas at
          Elevated Temperatures and Pressures    	         	III-5-1

 6.   C.Y. Wen and S.C. Wang
     Gasification of Solid Particles Containing Carbon	III-6-1

 7.   A.A. Godel
     Manufacture of Activated Carbon by Gasification in a Fluidized  Bed    .       . .  III-7-1

                                       SESSION IV

 1.   A.M. Squires
     Clean Power Systems  Using Fluidized-Bed Combustion    	IV-1-1

 2.   SJ. Wright
     Fluidised-Bed Combustion and the Design of Boilers	  IV-2-1

 3.   R.H. Demmy
     Ignifluid Boilers for an Electric Utility	  IV-3-1

 4.   J.W. Bishop
     The Modular Fluidized-Bed Boiler Concept   	  IV-4-1

 5.   D.L. Keairns and D.H. Archer
     Fluidized-Bed Boilers—Concepts and  Comparisons  	IV-5-1
VI

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                                     SESSION V

1.   D.H. Broadbent
    Development of Fluidised-Bed Combustion for Firing Utility Steam Boilers   ...   . V-l-1

2.   D.H. Archer
    Marketable Designs for Fluidized-Combustion Boilers        .  .            .      . V-2-1

3.   R.V. Seibel
    Design of Atmospheric Fluidized-Bed Combustion Steam Generator   	  V-3-1

4.   R.W. Bryers
    Design Features of Pressurized Fluidized-Bed Boilers	V-4-1

5.   D.E. Elliott and E.M. Healey
    Some Economic Aspects of High-Temperature Steam Cycles  	    V-5-1

6.   L. Reh
    Combustion of Oil or Gas in Fluidized Beds	 V-6-1

                                     SESSION VI

I.   D.B. Henschel (Recorded Minutes)
    Panel Discussion and Summary Session   	VI-1-1

APPENDIX - Attendance List	    A-l
                                                                                   vu

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INTRODUCTORY REMARKS




         Mr. P. P. Turner, NAPCA
KEYNOTE ADDRESS:




        Professor D. E. Elliott, University of Aston, England

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Introductory Remarks
                            1.  THE  NAPCA-DPCE PROGRAM
                                      TO CONTROL POLLUTION
                               FROM STATIONARY  SOURCES

                                P.P. TURNER

                Division of Process  Control Engineering
              National Air Pollution  Control Administration
  By way of review and to lay the ground-
work for this conference, I want to present a
bit  of background regarding our reason for
being here.  The "First  International  Con-
ference on Fluidized Bed Combustion" was
sponsored by the National Air Pollution Con-
trol  Administration  and was  held here in
November 1968. Apparently the Conference
was successful because it was decided that a
similar meeting  would be beneficial  after 2
years of additional research. That decision has
brought us together today for a further ex-
change of information.

  In  his introductory remarks to the first
Conference, Mr. Paul W. Spaite (then Chief
of the Process Control Engineering Program)
explained the growth of U.S.  interest and
effort in controlling air pollution, specifically
controlling the oxides of sulfur and nitrogen,
as well as particulate matter, from stationary
sources.

  It is  generally agreed that fluidized-bed
combustion (FBC) has considerable pollution
control potential. The lower FBC operating
temperature  (1400-1900° F)  may  minimize
the formation and emission of nitrogen oxides
(NOX), and is near optimum for the reaction
of SC>2  with limestone.  Particulates  can be
removed by  cyclones, scrubbers, or precipi-
tators.

  Removal of  SC«2  remains a  very  serious
problem. Table  1 gives some idea of the mag-
nitude of the problem. The figures shown for
1970 and beyond are projected from esti-
mates made by coal and oil interests and by
industry.  The tabulation  shows that power
plants emit most of the  SC<2 that is in the
atmosphere.

  The demand for electric power is increasing
so rapidly that sulfur oxides (SOX) emissions
may increase even if the following techniques
are developed and applied  as soon as possible:
  1. Construction of nuclear power plants.
  2. Substitution of gas or low-sulfur fuel oil
    or coal, where available.
  3. Use  of cleaned low-sulfur-content  coal.
  4. Introduction of  improved  combustion
    methods.
  5. Application of improved stack gas treat-
    ment and sulfur recovery processes.

Even with a national commitment to orderly,
but urgent, application of new technology,
the best that we can hope for  through the
year 2000 is a total SOX  emission rate from
all utilities near the present level. The number
of plants burning fossil fuels is increasing that
much.

  Nuclear power plants emit no SC«2; essen-
tially the same can be said for  plants using
natural  gas. The technology  for removal of
sulfur from fuel  oil appears to be reasonably
well in  hand based on Esso Research in Eng-
land. Further development  of hydroelectric
power will not be a major factor. Importing
                                      O-l-l

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 liquified petroleum gases will most likely in-
 crease to  the limit of economic availability.
 However, these methods of producing energy
 will not be sufficient.  The use  of coal  will
 steadily increase: it is expected to more than
 triple by the  year 2000, before leveling off as
 large nuclear power stations  replace those
 burning fossil fuels. About 65 percent of the
 SOX discharged  into  the atmosphere from
man-made sources comes from the combustion
 of coal. The average sulfur  content of  coal
 now burned  is  2.7  percent:  this  average is
 expected .to  increase to 3.5 percent  by the
 year 2000,  aggravating the situation addi-
 tionally.
Table 1. ESTIMATED  POTENTIAL  SULFUR DI-
   OXIDE POLLUTION IN THE U.S., WITHOUT
                ABATEMENT3
Source
Power plant operation
(coal and oil)
Other combustion of
coal
Combustion of petro-
leum products (ex-
cluding power plant
oil)
Smelting of metallic
ores
Petroleum refinery
operation
Miscellaneous sources
TOTAL
Annual Emission of Sulfur
Dioxide, millions of tons
1967 1970 1980 1990 2000
15.0 20.0 41.1 62.0 94.5
5.1 4.8 4.0 3.1 1.6
2.8 3.4 3.9 4.3 5.1
3.8 4.0 5.3 7.1 9.6
2.1 2.4 4.0 6.5 10.5
2.0 2.0 2.6 3.4 4.5
30.8 36.6 60.9 86.4125.8
   aFebruary 1970 estimates by National Air Pollu-
tion Control Administration, excluding transportation.
    Includes coke processing, sulfuric acid plants, coal
refuse banks, refuse incineration, and pulp and paper
manufacturing.
O-1-2
   NAPCA's  Division  of  Process  Control
Engineering (DPCE) is working on many tech-
niques  for eliminating SC>2 emissions to the
atmosphere.  Without  elaborating on the vari-
ous techniques, Table 2 gives you  some idea
of the  scope  of our program. Note the head-
ing "Fluidized-bed combustion (FBC)" under
item 3.

   DPCE first undertook work on FBC in FY
1968 at  a  funding  level of $400,000; the
current effort is 5 times greater. NAPCA has
already spent about $2.25 million of an esti-
mated  $75  million that will be required to

Table 2. PROCESS DEVELOPMENT FOR ABATE-
               MENT OF SO2

1.  General considerations of individual processes

2.  Precombustion processes

   a.    Coal cleaning
   b.    Coal gasification

3.  Combustion processes

   a.    Fluidized-bed combustion (FBC)
   b.    The Black, Sivalls, and Bryson

4.  Limestone processes
                                                   a.
                                                   b.
       Wet scrubbing
       Dry removal
5. Processes for sulfur recovery from stack gases

  a.   Cat-Ox
  b.   Wellman-Lord
  c.   Esso-Babcock and Wilcox adsorbent
  d.   Magnesium oxide scrubbing
  e.   Formate scrubbing
  f.   Ammonia scrubbing
  g.   Westvaco char
  h.   Molten carbonate
  i.   Sodium bicarbonate adsorption
  j.   Modified Claus
  k.   Catalytic chamber
  I.   Ionics/Stone & Webster
  m.   Alkalized alumina

6. Scrubber development	

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demonstrate  the  best FBC  system. NAPCA
believes that  the  FBC concept is particularly
well-suited  to collaborative  research so that
substantial  portions  of the total R & D ex-
pense can be shared by a number of govern-
ment and industry groups both in the U.S.
and abroad.

  NAPCA's interest in FBC because of  its
pollution control potential brings us to SC>2
formation and emission when coal and oil are
burned. Limestone addition for SC>2 control
should be  especially effective  because flui-
dized-bed combustors operate in the tempera-
ture range best suited for the limestone-SO2
reaction  (1400-1900°F); the  fluidized-bed
combustor  also  should provide good mixing
and  longer residence time for the reaction to
take place.

  The current NAPCA program on  FBC con-
sists of laboratory  and  small pilot-scale  re-
search studies. In addition  to work abroad,
NAPCA projects on the ah-  pollution control
aspects of FBC  are  underway both at Pope,
Evans  and  Robbins and  at  the  Argonne
National Laboratory. NAPCA is also sponsor-
ing new work on FBC combustion and pollu-
tion control  at: Esso Research and Engineer-
ing  Co., Linden, N.J.; The U.S. Bureau of
Mines,  Morgantown, W.  Va.;  Westinghouse
Research and Development Center; Consolida-
tion Coal Co.; Esso  Petroleum Co.,  Ltd.; and
the National Coal Board, England.
  NAPCA's  future  program  includes  con-
ceptual design studies as well as technical and
cost feasibility studies of FBC systems. Since
each NAPCA-DPCE contractor is represented
at this meeting, I will say nothing more about
their programs; rather, I will let them present
the  results of their own work.

  In  closing, just  a word about the various
sessions.  I  have asked each Session Chairman
to  control his  session  carefully. The last
speaker in a given session deserves as much
time  as the first; however, we also wish for
in-depth  discussion. The reason for  limiting
the participants in this conference is to stimu-
late discussion,  both formal  and informal,
much as  is done at the Gordon conferences —
so let us hear from each of you.

  It is now my pleasure to introduce our key-
note speaker who will address the conference
on  "Exploiting  Fluidized-Bed  Combustion."
He  will  point out  the potential of this com-
bustion concept including history, status, and
prospects. He is  well qualified to set the tone
for  this conference. He was with the Central
Electricity Generating  Board  for  9  years
working in the area of FBC heat transfer and
carbonization of coal. For the past 2 years he
has been Professor of Mechanical Engineering
at  the University  of Aston in Birmingham,
England, where he has continued his research
on  FBC concepts. Gentlemen, I present  to
this conference  Professor Douglas E.  Elliott.
                                                                                   O-1-3

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Keynote Address
                                                          2. EXPLOITING
                               FLUIDISED-BED  COMBUSTION
                                 D.E. ELLIOTT
                         University of Aston, England
INTRODUCTION
  Although the original aims of the British
and American fluidised-bed combustion re-
search programmes were different, both pro-
grammes have moved along similar lines. The
British saw fluidisation as a means of reducing
the capital  costs  of power stations;  the
Americans placed much more emphasis on the
possibilities of sulphur removal. However, the
new technology has so many interesting and
potentially commercial facets that we  can
now look forward to a period in which more
and more effort will be mounted by a greater
number of firms and research organisations on
diverging projects. It is therefore useful at this
stage to  take stock of our knowledge, pin-
point  problems  that it is imperative to solve,
estimate the probable return from competing
processes, and put in hand the next round of
developments.

  In terms of the S-curve of Figure 1, charac-
terizing  the development of any technology,
we  are just at the toe. Whether fluidised-beld
combustion peters out (a), as M.H.D. seems to
have done, progresses at a modest rate (b), or
escalates at an ever-increasing rate (c) depends
mainly upon:

  1. Is the technology based  on  good prin-
    ciples and  are the natural  phenomena
    working  with us?  Unless this is  so no
    amount  of effort will produce a work-
    able plant.  In my opinion we are not in
    much danger on this account and most
    of  our  encouraging  results  can be ex-
    plained in sound scientific terms. Only in
    a few areas do we  find conflicting re-
    quirements; in most of these there are
    alternative ways around the problems.
  2. The cnoice and design of the pilot-scale
    plants  which should follow  our large-
    scale rig experiments. It  is in this area
    that we are most vulnerable. A bad deci-
    sion at this stage could mar the progress
    of  fluidised-bed  combustion,  possibly
    irreparably.  We  cannot expect to be
    allowed  to  fritter  away millions of
    pounds  on this work, as has been the
    case with nuclear power, even  though
    several abortive attempts could probably
    be justified on a cost/benefit basis.
  3. The amount of ingenuity we apply and
    whether we are prepared to take more
    than one step at a time. There is always a
    reluctance on the  part of the designer to
    do more than one thing at a time; in this
    day  and age, this could  lead to being
    overtaken by events.

  This Second  International  Conference is
therefore taking place at what must be the
most critical period in the history of fluidised-
bed combustion and some indication of the
probable  outcome will be the progress that
has been made in the last year. Are our prob-
lems still the same as  they were a year ago?
Have we solved any major snags? Have any
new ideas come forth?  Are the economics
turning out as was forecast? In view of the
large number of technical papers which we
shall be  discussing this  week, it would be
pointless  for me  to do  more than  a  brief
survey; I would like to air a number of uncer-
tainties in the hope  that they  will arouse
                                       O-2-1

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 future exchanges. I also wish to outline some
 exciting fields for further research.
 PRESENT STATE (PRIOR CONFERENCE)

 Heat Transfer

   We know enough about the theory and the
 practical  realisation  of  high  heat-transfer
 coefficients to be certain that early claims of
 reduced heat-transfer  surfaces  in fluidised
 systems are fully justified. Little more work is
 needed here  except  for special  conditions
 imposed by some advanced designs.


 Fouling and Corrosion

   Encouraging experience with heat transfer
 surfaces in rig testing over accumulated times
 of several hundred hours supports the conten-
 tions that  this is one  of the most important
 aspects of fluid-bed combustion. Long dura-
 tion trials of 1000 hours or longer will  be
 needed before we can assess  the benefits  of
 reduced maintenance  and the possibility  of
 achieving higher steam conditions, but there is
 good reason for optimism, at least when burn-
 ing coal. Because there is little relevant infor-
 mation, work should start on  the fouling and
 corrosion in heavy oil burning systems.

 Combustion

   I  anticipate  that  we shall be hearing that
 with  medium  velocities using  1/8-in. coal,
 good combustion efficiencies can be obtained
 using an acceptable  recycle ratio of fines, at
 temperatures of 1400  to 1500°F.  The prob-
 lem is not quite as clear for velocities of 8 to
 14 ft/sec favoured  by  plant manufacturers
 because they cut down containment costs and
 produce compact units. Using larger coal en-
 tails a  much higher  carryover; also, recycling
 to the  primary bed may give unreasonably
 high loadings of the separators. Thus attention
 is being drawn to secondary beds for burning
out the elutriated fines. Having successfully
O-2-2
burnt  the residue from anthracite  burning
power stations, which contained only about
20 percent carbon, I was surprised  to  hear
that burning the  fly ash from fluidised-bed
combustors proved difficult in early tests. I
am looking forward to the discussion of this
problem;  I can hardly  believe that it  is  a
critical one.
Carryover

  While on the  subject of carryover, I feel
that too little attention has been paid to what
happens to the  coal  immediately  after its
injection into the bed. There is plenty of evi-
dence to show that, under rapid heating con-
ditions, the volatile emission is much higher
than that shown by standard assays and that
the swelling of the  particles is increased. Con-
sequently, there is a greater tendency for the
particles to decrepitate on carbonising and the
strength of the  chars  is lower; both lead to
greater  degradation of the particles by. attri-
tion in the bed. It is therefore useful to specu-
late  if  we could improve matters by  some
preconditioning of the coal or by making sure
that we do not have too rapid heating.
Coal Feed

  One of the two most serious problems still
outstanding is coal distribution.  I  have seen
tentative  designs  with feed points  10 feet
apart. Do we really think this is practical? If
not, how close must they be? From time to
time, people talk about the rapid  mixing in
fluidised beds; because lateral mixing is rela-
tively slow and to prevent massive local evolu-
tion of volatiles, we must have a good spread-
ing of coal. I will mention possible solutions
later.
Startup and Control

  The second serious problem (the first was
coal feed)  arises because the heat extraction

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rate of submerged tubes only falls a little at
low fluidising velocities and because combus-
tion cannot take place effectively at tempera-
tures much below 1300°F. Thus we  cannot
start a bed containing a full  array of tubes
unless  we provide a burner of near-full rating
beneath the bed. Similarly, we can only vary
the output of such a bed over a very  narrow
range  of output and resort to separate beds
and startup beds without tubes.
Base Design

  Although there may be room for improve-
ment,  base design is not a major problem.
Operating experience with large rigs suggests
that acceptable  designs with  low  pressure
drops have been developed.
Sulphur Removal

  Sulphur removal experience on either side
of the Atlantic  seems to vary. With British
coals in British rigs, adding limestone retains a
substantial proportion of the sulphur. I under-
stand that American results have been rather
disappointing.  Whether  this is  because
American coal is awkward, because American
limestone is unreactive, or because American
rig conditions vary will be established by test-
ing American coal in British apparatus. The
fact that we have one set of encouraging data
suggests  that some way round this probldm
will be forthcoming.


  The economics of the sulphur removal have
yet to be disclosed. Can we  afford  to add
limestone in  the quantities required? Can we
dispose of the resulting ash without polluting
the land? Would  it be  better to  go for  a
regeneration system? What can we  do about
the existing power stations? Can we do for
coal what Esso is trying to do for heavy oil? Is
it more economical to scrap existing boilers,
replacing them with fluid-bed units? Many
questions need to be answered.
Hydrocarbon Pollution

  It is generally assumed that exhaust gases
from  fluidised beds  will  be very  low  in
unburnt  hydrocarbons.  Experience  supports
this; but  again, I return  to the coal feed
arrangements. Too much coal in any one area
will almost certainly yield unburnt volatiles.
On  the credit side however, it is likely that
the heavy clouds of smoke which accompany
load changes in normal plants can be avoided
in the new  system. The fluidised bed has a
high thermal inertia which should allow load
changes to be met without violent changes in
combustion conditions.

  There is however another point which must
be borne in mind. Will there be any very fine
carbon produced which will evade the precipi-
tators, and will the removal of sulphur from
the gases reduce the efficiency of these units?
It is natural to assume that, since  we are start-
ing with a larger size of coal, we must end
with a larger size  of dust which will be more
readily caught. However, remember that it is
not  the mean size  that  matters:  anything
above 5 microns  is easily captured. The sub-
micron  particles  are what  matter: they  are
more  likely to be a function  of the coal
heating period and the attrition in the bed.

Oxides of Nitrogen

  Because fluidised-bed  combustion can take
place  at temperatures of  1500°F or below,
then theoretically the yield of oxides of nitro-
gen should be very low; much lower than with
adiabatic combustion  where flame tempera-
tures of 3000°F are common; and very much
lower  than  is  formed  in petrol  or  diesel
engines.

  However,  there have been conflicting re-
ports  of experimental data.  Some  of these
suggest that the oxides of nitrogen are much
above the theoretical  equilibrium at the flui-
dised-bed  temperature;  others  suggest that
they are only  marginally above this  level. A
possible  explanation is  that  the  particle
                                     O-2-3

-------
temperature  is several  hundred degrees (F)
higher than  the bed temperature. Experi-
ments,  in  which thermocouples have  been
placed in large (1/4-in.) particles, show this to
be  the case. This  explanation  is  also con-
firmed by visual observation of fluidised com-
bustors, especially those running with  large
coal at high velocities.

  In  this  context, it  is  reasonable  to expect
that large-coal/high-velocity systems will be
worse  than  smaller-coal/lower-velocity  beds
because:

   1. There is a tendency to run high-velocity
     systems  at  a  higher temperature  to
     obtain acceptable combustion efficiency.
  2. The scrubbing velocity of the air past the
     larger particles  is higher; therefore, the
     reaction intensity will be higher.
  3. The heat transfer from larger particles to
     their neighbouring large particles will be
     lower than when small particles are in
     the vicinity of other small particles.
  4. The bed volume used in high velocity
     combustors  is  generally  smaller  than
     with low velocity units, hence the spatial
     density of the carbon bearing particles is
     higher and the chance of their coming
     together to create a  temporary  "hot
     spot" is greater.
  5. Observation of fluidisation reveals that
     its quality is much smoother with  small
     particles and, because lower bed depths
     can be used, the bubble  size is smaller.
     Because bubbles contain few particles,
     temperatures  therein   can  rise by the
     combustion either  of  volatiles or of the
     few particles in the bubbles.

  Thus, from the point of view  of the pro-
duction of oxides of nitrogen, it would appear
better to use as small a particle as possible, a
fact that  obviously conflicts with contain-
ment costs.

Costs

  We shall be hearing much more about the
economics  of fluidised-bed boilers later in the
O-2-4
Conference. Because  we have  not met  any
major  snags  over the previous year which
would  be costly to solve, I feel confident that
fluid-bed combustion can be, at one and the
same time, economical and  of great amenity
value. This does not mean that we should be
complacent. On the contrary, because of the
many intriguing possibilities, I  would like to
devote the  remainder of my address to a plea
for expansion of our programmes.

   Three centuries ago Shakespeare wrote:

     "There is a tide in the affairs of men
     which, taken at the flood, leads on to
     fortune."

                              Julius Caesar.

   However to be  more appropriate to the
modern world,  with its rapid  technological
development, the affairs of men ought now to
be  compared  to  a  breaking  wave. Those
familiar with surf riding know that  the water
in a wave is constantly overtaking that ahead
of it; similarly, in research and development,
new  ideas  often overtake the old. Thus the
timing of the effort to ride the wave is much
more critical  than  at any previous time in
history. Shakespeare also knew this:

     "And  we must take the current when it
     serves, or lose our ventures."

   I submit that now is the time  for us to
consider the next round of development of
fluidised-bed combustion, so that we need not
embark on a protracted course of sequential
improvements.
DEVELOPMENT PROSPECTS

  The first improvement on current designs
would be to raise  the  top temperature and
pressure of the steam cycle. A later paper out-
lines"  the economics of the extra tube costs
versus the lowering of the heat rate, so I will
not go into that here. However, this is not the

-------
only way of improving the overall efficiency:
we should  not ignore combined  gas-turbine/
steam-turbine cycles, the Field cycle, or even
liquid metal cycles.

   To  the  British,  efficiency  is extremely
important  because  coal-fired power  stations
will  not be  able to compete with nuclear
energy unless we can lower both fuel costs
and  capital costs. Figures 2 and  3 show the
situation a  year ago.

   Coal  needed something like a  20-percent
reduction in overall costs to compete with the
new generations  of nuclear plants (if they
reach their targets).  Since  1969, coal costs
have risen  by  10 percent: further rises seem
inevitable.  On the  other hand, since nuclear
prices  have gone  up (the cost  of nuclear
power is now about 0.6d/kW), the situation is
no worse for coal. If we can relax the specifica-
tion for coal by  adopting fluidised-bed com-
bustion and increase efficiency, the prospects
for coal are much better.
Secondary Heat Exchange Beds

  The  first  Central Electricity Generating
Board (C. E. G. B.) Research and Develop-
ment design  studies used a second fluidised
bed  above the combustion bed to  act  as an
inert heat exchanger, using economiser tubes
to cool the combustion gases to the air heater
inlet temperature of 750° F. When subsequent
studies showed that the pressure drop penalty
was too great, the idea was abandoned.

  Recent data on extended  (finned) surface
tubing  has led to  a  revival of interest  in
systems  without convective  steam or  water
tubing.  The consequently  possible plant-
layout simplification and heat transfer surface
reduction make this  an  interesting alternative
to existing designs. Note that,  whereas  in
convective systems  the  increased   thermal
performance  of finned tubing is offset by the
increased pressure drop, the  even better per-
formance of  finned tubes in  fluidised beds is
accompanied by a reduction in pressure drop.
Circulation of Solids

  One way of overcoming the startup  and
control problem  is  to remove the cooling
tubes  from the  bed and  place  them  in  a
separate steam raising unit. The heat is trans-
ported from  the  combustor to the tubes by
circulation of bed material via fluidised pipe-
lines or ducts. Since the tubes do not cool the
bed  directly, the  bed is brought  up to
temperature  with   a  comparatively  small
auxiliary  burner  with  no  solids circulating.
Partial loads  would be  controlled by varying
the rate of circulation of solids.

  This system has several distinct advantages
over "normal" fluidised beds:
   1. The  steam raising unit is very compact
     and can be manufactured at the factory.
     For very large stations, the superheater,
     the  reheater, and the  evaporators  would
     be  separate; for  120-MW  units they
     could be transported (as a unit) by truck
     to the site.
   2. Locating the steam raising units near the
     steam turbine eliminates the steam pipes
     and makes it economical to incorporate
     multiple reheat.
   3. Coal feeding can be improved by using
     rotary distributors or by  rotating the
     solids in the bed,  using either air injec-
     tion  or  rotary  paddles.   Another
     possibility that might be investigated  is
     the  combustion of run-of-mine coal at
     relatively low velocities.  Although the
     larger coal would drop to the bottom of
     the   bed, the   continuous  movement
     across the  base  would prevent  the
     particles settling; there is a good chance
     that the system would work. We already
     have  experience  with bricks moving
     around circulating beds under the  action
     of the solids, so that there  should be
     little doubt that 1-in.  lumps would keep
     moving. The saving in coal preparation
     costs  would be  quite  important; so
     would be the improvement in outage and
     maintenance  costs of the preparation
     plant.
                                     O-2-5

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   4. Because  we can more or less dictate the
     temperature of the fluidised solids which
     any particular  tube  material  contacts,
     the design of the tubing can be  further
     optimised; corrosion,  which might occur
     using low-temperature tubes in  a high-
     temperature  bed of very bad  coal ash,
     could  be  eliminated  by lowering  the
     temperature of the solids in that  area of
     the  tube  bundle.  (This   might   be
     particularly relevant when  using heavy
     oil.)
   5. Reducing  the   startup  and   control
     problems and  increasing the rate  at
     which loads can be accepted or shed will
     save live  steam, normally dumped during
     these periods.
   6. There is  a possibility that  introducing
     special heat carriers into  the system will
     improve still further the compactness of
     the  boiler.   This   improvement   of
     compactness  may not amount to very
     much at  low steam  temperatures, but
     could lead to lower maintenance costs
     and worthwhile  cost reductions in high
     temperature units.
   7. Once heat transport is incorporated, it is
     then only a short step to heat storage.


Heat Storage

   Interposing  a   large hot-solids hopper
between the fluidised combustor and the heat
exchanger unit makes it feasible  to run the
combustor for longer periods than dictated by
load requirements. Thus all  equipment  as-
sociated  with  combustion  (namely, coal
conveyors, coal mixers, bunkers, crushers,
feeders, air fans, the combustor itself,  the air
heater,   gas   cleaners  and   precipitators,
chimneys, ash handling gear, and  the founda-
tions)  and civil engineering associated with
this  equipment  can be operated  24 hours a
day. This equipment therefore can be rated at
only a fraction of the station output, depend-
ing on the load factor for which  the plant is
intended. Although  future  base loads are
expected to be met by nuclear power  plants,
there will still  be a requirement for low-load

0-2-6
 fossil-fuel stations: this suggestion is worth at
 least a paper investigation.

   If we take an overall power station cost of
 about £55/kW, the combustion-related costs
 would  be  in the  region of £22/kW and  the
 fixed  cost  dependent on  electrical capacity
 would be about £33/kW. Thus the cost of the
 storage  plant,  excluding  the storage vessel
 itself, would be:

   £33 +/£22 x_L\, where L is the load factor.
        I    247
   For an 8-hour  daily duty, about £15/kW
can be saved, offset by the cost of the storage
vessel. Originally,  a double-vessel system was
envisaged;  but   this was  soon  discarded  in
favour of a compartmented vessel with one
empty  space:the  solids  to be transferred
between the compartments in rotation. This
system was  at a cost disadvantage because the
compartment walls had to be  designed  to
withstand  the  hydrostatic pressure  of the
solids. The latest idea is to sink a single vessel
in the ground,  separating the hot solids from
the cold with  a moveable vertical partition.
The  zones immediately on each side of the
partition  would   be  fluidised,  with  solids
pumped from one to the other depending on
whether the plant  was storing or recuperating
heat.  It may even be feasible to use natural
sandy terrain with the minimum of contain-
ment if the  water  table is well below the level
of the  store.  However, even without  such
ideal  circumstances, it should be possible  to
construct the  storage  unit for a cost in the
region  of   £40/cubic  yard.  This  compares
favourably  with a cost of roughly £12/cubic
yard for an insulated and heated  heavy  oil
storage vessel.

   For a 4-hour daily duty, a 500-MW station
would need 1.3 x  1010 Btu of heat storage,
requiring a  200-ft diameter by 22-ft  deep
vessel,  or  equivalent. The cost  would  be
£/2kW. Figure 4, showing how the cost of the
storage power station would vary for different
load duties, indicates that for low daily usage,
the storage system  is most attractive.

-------
   Apart from cost, such a plant would have
 several advantages over straight sets:

   1. If sulphur recovery is incorporated, the
     extra  equipment  is also  used on  a
     continuous basis with a further saving in
     capital costs.
   2. The  stack  effluent, exhausted on  a
     24-hour basis, for a given electrical out-
     put would be less objectionable.
   3. The station can be brought on line very
     quickly. It  is almost as good as water
     pump  storage, but with the  advantage
     that locating it in the load centres will
     save electrical transmission charges.
   4. The reliability of  the station  should be
     improved.  Even  if some  part of the
     combustion  system  fails,  there  is  a
     period of grace during which  no loss  of
     capacity results. An improvement of say
     5 percent on this account is  equivalent
     to a reduction of a similar 5  percent in
     capital cost in an integrated network.


   Figure  3 shows how improved  efficiency
and  heat  storage can  help coal  in future
competion with nuclear energy.
DIVERSIFICATION

  Up to now, fluidised-bed  combustion has
primarily been aimed at coal combustion; but
in the last year or so, attention has spread to
the burning of heavy oil and  (more lately) Jo
gas combustion. Since the work is still mainly
directed  at large scale  power generation,  it
might pay  us to  review  other  applications,
both  from  pollution and economic  stand-
points.

  If we look at the amount of fuel used, we
find that less  than half goes  to the electrical
industry. Small-scale use of fuel in factories,
domestic heating, and transport add up to a
major portion of fuel  consumed; in many
cases  they  represent higher  sources of pol-
lution  per  pound  of  fuel burnt.   Could
fluidised-bed combustion help in these cases
just  as  we  expect  it  will  in  large  scale
production?

   For small scale factory furnaces there  is
every  reason  to  expect that fluidised-bed
boilers or furnaces for metal processing can be
just  as effective in reducing capital  costs and
pollution. Here, however, I would criticise the
attempts which have been made up to now on
packaged  shell boilers:  If we are to exploit
best the advantages of fluidisation,  we ought
not  merely  to translate previous designs into
fluidised versions. We must survey the  whole
problem and come up with  advanced designs
which take advantage  of all that   the new
technology  can provide.  Our aims  must  be
more  a  efficient  plant,  a  more  easily
controlled system, better steaming conditions
(users have never had the opportunity to buy
anything other than  low temperature  steam
boilers), and  a plant that  can  use a wide
variety of fuels.

   In this context, I would  like to  stress the
future importance of total  energy installa-
tions.  In the  past, such installations have
rarely been economical  because  of the high
incremental cost  of high temperature boilers.
Now we can produce a  1100°F by 2000-psi
plant for virtually the  same cost  as for a
212°F  system;   economics  are   radically
different. The development of new engine
designs, mentioned later, could also have a
large bearing on  this system which everyone
agrees is ideal but seldom appears economical.


Domestic Appliances

   At first  sight,  the application of fluidised-
bed  combustion  to  gas-fired  central heating
systems does not appear to be very attractive:
current boilers  are  fairly highly developed,
have reasonable efficiencies, and are not bad
from a pollution aspect.

   The   only   advantage  of  fluidised-bed
combustion would be a probable reduction in
oxides  of  nitrogen  resulting  from low
temperature   combustion.  Note   that  the

                                     O-2-7

-------
 tendency, if any,  is away  from the common
 fishtail burner  and toward  a more  intense
 (presumably    higher   temperature)   and
 compact  combustion  system  which  may
 increase pollution from oxides of nitrogen.

   Nevertheless, preliminary  designs of small
 gas-fired fluidised-bed combustors which  we
 have undertaken  at Aston show  that  these
 systems, too, can  be very  compact and yield
 better efficiency  than  current  installations.
 Because these systems can be adapted  to burn
 oil easily and, with the addition of a feed
 system, could burn fine coal, they might have
 wide commercial application.
Room Heating

  A great deal of money has been spent in
the U.K. developing smokeless briquettes. The
main  disadvantage  is  high  cost:  roughly
double  that  of the original fine  coal  from
which they are made.

  I often wonder whether it is too  late to
consider   developing   small  fluidised-bed
combustors for  room  and  house heating.
There is no problem in developing a 6-  or 8-in.
combustor; the efficiency,  even  without flue
gas heat recovery, would be above 66 percent.
All that  would  be  required  would be  a
suitable coal-hopper/variable-rate feeder and
an air blower unit. Pre-mixing the coal with
the right  amount of limestone could reduce
sulphur pollution to a minimum.  Such coal
could be  delivered either  by tanker or in
pre-packed bags  to  eliminate much  of the
labour involved in normal coal fires.

  The constantly changing flame  pattern is
one of the main  attractions of an open coal
fire; a similar  advantage would be possible
with a fluidised bed if it were enclosed in a
high-temperature-glass  tube  (these  can  be
made  to withstand  1600°F). The constantly
changing bubble  and particle motion would
be  a  substitute  for  the  coal   flames.
Economically, one could afford to pay quite a
O-2-8
high initial price for such a unit  if it could
burn cheap coal.

  The  disposal  of   the   ash  would  not
constitute a severe problem; it might even be
possible to wash it away through the house
sewage system.


Combustion of Refuse

  Attention  is now being  drawn to better
methods of  disposing  of household refuse.
The low combustion temperatures permissible
in fluidised-bed  combustors, combined with
their thermal inertia which would counteract
the variability of  the  combustible  consti-
tuents, may prove advantageous.

  One problem will be how to deal with the
large range of size and density of refuse. Light
materials such as wood  would float on the
surface  of the bed until  burnt and so should
present  no  difficulty.  The  ideas mentioned
earlier for circulating the bed solids may also
be of use in  keeping larger dense particles in
motion  until  they  are completely burnt or
rejected.

  The major problem  would be the effluent
caused by burning plastics. We might be able
to  reduce the  chlorine pollution  of the
atmosphere by using  techniques  similar  to
those being used for sulphur recovery.


Vehicles

  One  of  the major  pollution problems is
transport  vehicles.  For urban railways, this
has been  solved by  electrification  at high
capital cost; but for long distance rail trans-
port, trucks, and cars, the problem has not
been fully solved. Various devices for cleaning
up  petrol  and  diesel exhausts  have  been
proposed; all  increase  the cost and  decrease
the output of the power units.

  As  a consequence,  special  attempts are
being made to produce  prime movers with

-------
external   combustion   systems;  the  two
principal ones are  steam engines and  hot air
engines. In both these  cases it  is extremely
likely that fluidised-bed combustion not only
could produce a cheaper indirect firing system
but also could use cheaper fuels  and produce
higher temperatures.

  One of the major steam engine problems is
the contamination of the working fluid with
lubricating oil. This oil is difficult to  extract
and,  on passing through the  boiler  tubes,
undergoes   cracking   and   forms   carbon
deposits. Both the deterioration in internal
heating transfer  coefficient and  the presence
of very high temperature combustion gases on
the outside of  the tubes result in periodic
burnout of the tubes. This problem would be
completely eliminated in the fluid-bed boiler;
with  the  tubes  designed to stand full bed
temperature,  carbon  deposits   would  only
result  in loss  of output. Currently, some of
my  undergraduate students at Aston are
building a fluid-bed boiler capable of  1300°F
by 2000-psi pressure operation. They  are also
developing  a  prototype engine, based  on a
normal car engine with modified valve gear.
Fluidised   combustion   offers  similar  ad-
vantages for  other Rankine-cycle working
fluids and for hot air engines.
WE HAVE ONLY SCRATCHED THE
SURFACE

  The above ideas only scratch the surface of
where  fluidised combustion may  lead; other
areas  are industrial  heating  furnaces,  gas
turbines burning ash-laden  fuels, combustion
of waste gaseous products at refineries and
sewage works, and (perhaps  ultimately) the
"combustion" of nuclear fuels.


  Thus we  see  that there is an enormous
development potential, the exploitation  of
which now seems near at hand.


  I believe that the field is so wide that there
need be little conflict of interest either among
research workers  or  firms and I trust that in
the next few days we can  have frank discus-
sions  on  many of these aspects which will
help us all to frame our future programmes.
For my part I  shall  be concentrating on the
flow  properties  of  fluidised  systems since
these  have  a  mixture of  academic and
practical  problems suitable for  study in a
technological university, but which,  never-
theless, could be  extremely important in the
overall   exploitation  of  our   Conference
subject.
                                                                                   O-2-9

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LU
I
3
LU
     .
    O
    LU
                        PROGRAM
                        ESCALATES/
                        PROGRAM PROGRESSES
                          AT MODEST RAT
                        PROGRAM PETERS OUT

                             a)
                                       1.3


                                       1.2


                                       1.1


                                       1.0


                                    a   °-9
                                   x.
                                   ^
                                       0.8
                                   fe
                                   O
                                   0   0.7


                                       0.6


                                       0.5


                                       0.4
                                                  0.3.
                                                                  I         I
                                                              	ORIGINAL ESTIMATE
                                                              	ACTUAL COST
                                                                •  TENDER
                                                               AGR ADVANCED GAS - —
                                                                   COOLED REACTOR
                                                               ATR ADVANCED
                                                                  THERMAL REACTOR
                                                               FR  FAST REACTOR
                                                                        HIGH-COST COAL
                                                               RADCLIFFE
                                                                (COAL)-
                                                                         LOW-COST COAL
                                                                                 ATR
                      TIME
                                                    60
                                                             65
                                                          70
                                                        YEAR
                                                                           75
                                                            80
   Figure 1.  Development of technology.
                                    Figure 2.  Power cost by type versus year.
 io
 o
 O
    0.8
    0.7
    0.6
    0.5
    0.4
    0.3
    0.2
          I     I     I     \
          AGR ADVANCED GAS-
              COOLED REACTOR
          ATR ADVANCED
             THERMAL REACTOR-
          FR  FAST REACTOR
          	COAL WITH TRANS-
             MISSION BONUS
                      FR '85-
                                   20%
                                 AGR '80
                                  •ATR '80
I
I	I
      0.3   0.4   0.5   0.6   0.7   0.8   0.9

                 LOAD FACTOR
  Figure 3.  Power cost by type versus
  load factor.
                                               fe
                                               o
                                               o
                                                   TOTAL COST:
                                                 STORAGE COST X3
                                                 STORAGE COSTX2
                                                 STORAGE
                                                                               T
                                                             ICOST
                                         0
                                                                  COMBUSTION
                                                                  COSTS
                                                    CAPACITY
                                                    COSTS  _
                                                    (CONSTANT)
                                                                           I
                                                  6
                                                                 12
                                                                           18
                                                            24
                                                 DAILY GENERATING HOURS
                                   Figure 4. Variation in heat storage cost for
                                  different daily load factors.
0-2-10

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SESSION I:



      Small-Scale Development of Fluidized-Bed Combustion
SESSION CHAIRMAN:



      Mr. A. A. Jonke, Argonne National Laboratory

-------
                             1. PILOT PLANT EXPERIMENTS
                                                                      AT THE
             COAL RESEARCH ESTABLISHMENT  (CRE)

                             D. F. WILLIAMS

                        National Coal Board, England
   The purpose of the pilot plant work at the
Coal Research Establishment (CRE) has been
to obtain data for costing design studies on a
660-MW  fluidised-bed  boiler,  and  for  the
design of a prototype unit. In  the preferred
concept,  crushed  coal  is fed to a shallow
fluidised bed of its own ash, operating within
the range of 700  900°C. The heat generated
is transferred  to horizontal tubes immersed in
the bed, and the  off  gases  are cooled by
conventional  means.  Elutriated  carbon is
collected, and is burned by recycling to the
original bed or  by injection  into a separate
bed.

   At the  time  of the  first Hueston Woods
Conference,  the initial  small-scale work had
shown that, at a fluidising velocity of 2 ft/sec,
high combustion efficiency could be achieved
at  moderate  temperatures  provided  that
elutriated carbon was recycled to the bed. A
high rate of heat transfer between the bed and
immersed tubes  was demonstrated, and it was
found that the rate was not seriously reduced
by packing the tubes closely together.

   Since that time, a combustor of 3-ft square
cross-section has  been  built  to simulate a
section of a full-scale bed. For this purpose, it
has been provided with insulated walls. The
combustor can accommodate full-size tubes in
the bed  in  realistic  arrays,  which can be
changed to test alternative designs.  Two of
the tubes form  part of high pressure  water
circuit,  in  which bed-tube  heat  transfer
coefficients can  be measured  at the  tube
temperatures that are anticipated for a power
station boiler. The remaining  tubes are used
to extract heat such that the bed temperature
can be  set  at any  required level between
700  and 850°C.  The bed height  is con-
tinuously  monitored,  and  is   controlled
automatically at  any level  up to 4 ft  by
withdrawing  surplus ash through the base.
The type of base and the freeboard height can
be varied.  Ports in the  walls allow pressure
and temperature measurements and solid and
gas samples to be taken to monitor conditions
in and above  the bed. Elutriated fines  are
collected in  two  cyclones in series,  and  the
primary cyclone fines can be recycled to  the
bed, either in total or in part, if required.

   Runs have been carried  out at fluidising
velocities between 2 and 8  ft/sec, with coals
sized 1/16-in by 0 and 1/8-in. by 0, the width
of the combustor having been  reduced to
1-1/2 ft to allow the velocity to be increased
above  4 ft/sec.  The conditions therefore
overlap  those  covered by  BCURA  in their
27-in. diameter combustor. The coal/air ratio
has been varied over wide limits to simulate
the conditions that  might  arise  at different
distances from a coal feed point in a wide
bed.  In these runs, carried  out  without
recycling the fines, the rates of combustion in
the bed and in  the  freeboard were deter-
mined, together with the  amount and size
distribution of the unburnt carbon  and  ash
products from the bed, the  cyclones, and the
gas entering  the stack. The  efficiency of the
cyclones, needed  to  recover the elutriated
fines from a power  station boiler  without
excessive loss of combustible matter, can be
calculated  from  the results.  The  primary
cyclone  on the plant was adequate  for this
                                        1-1-1

-------
purpose,  and   further  experiments   were
performed  to discover how the carbon so
collected could be burned. In some of these
experiments the primary cyclone fines were
continuously recycled to  the bed;  in others
the  fines were  fed  separately  to  the  bed
instead of coal to simulate a fines burnup bed.


  In each run, the bed size  distribution has
been measured and  has been related to the
bed-tube heat transfer coefficient. The effects
of tube and bed temperature, bed density,
fluidising velocity, and tube packing on the
coefficient have also been determined.

  A model of what happens to the ash from
the time it is injected into the bed as part of
the coal feed to the time it is removed as
surplus   from  the  bed  or is elutriated,
including its circulation in the recycle system
if this is in use, has been developed. Given the
feed rate of the ash, its size distribution in the
feed, its resistance to degradation,  and the
efficiency   of   the  cyclone,  the  model
correlates  the  combustor  data  for  the
quantity withdrawn from the bed, size  distri-
bution of the bed, the elutriation rate of each
size fraction, the recycle rate, and the change
in size  through  degradation. The model can
therefore be used to predict the ability to
maintain bed height,  and  the recycle rate or
the rate of  feed to a fines bumup bed  in a
boiler run under a given set of conditions. The
model is also of value in  predicting the bed
size distribution and, hence, the bed-tube heat
transfer coefficient.
   To find how quickly coal mixes laterally in
a  bed  after  injection,  measurements  of
particle mixing rates have been made in a 5-ft
diameter cold bed of ash using a radioactive
tracer  technique.   The  variables  studied
include fluidising velocity, particles size, tube
packing, and  bed height. The  results have
been  incorporated in a mathematical model,
based on the combustion and elutriation rates
measured in the 3-ft combustor, leading to a
prediction of the effect  of coal feed point
spacing on  combustion  efficiency. In  the
model,  the  carbon concentration  gradient
between feed  points, and  the  variation in
excess air level, are calculated and have been
compared with the gradients measured in the
combustor.

   Another  combustor has  been built,  to
investigate  the extent to which  corrosion,
erosion,  or  deposition may occur on tubes
placed in and above the  bed. The tubes are
made up  of different  alloys, and are each
cooled to a different temperature during the
test, which  may  last for from 100 to 1000
hours. The plant has  been  designed  to  run
automatically, except for the replenishing of
coal feed hoppers, for this length of time.

ACKNOWLEDGMENT

   The  work  described  in  this  paper  was
carried out as part of the research programme
of the  Research and  Development Depart-
ment of the National  Coal Board.  Views
expressed are those of the author and not
necessarily of the Board.
1-1-2

-------
                                   2.  FLUID-BED COMBUSTION
                                                                              OF
                                               VARIOUS U.S. COALS
                      N. H. COAXES AND R. L. RICE

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

  Fluid-bed combustion has several charac-
teristics which make it potentially attractive
as a boiler-firing technique. First, the fluid
bed  is an  efficient  contacting medium  for
gas-solids reactions, thus facilitating reaction
of solids with SO 2  to  remove it from  the
combustion   products.   Also,  fluid-bed
combustion takes place at lower temperatures
than does conventional combustion, and thus
should result  in lower emissions of oxides of
nitrogen. Another potential advantage of the
lower combustion temperature is decreased
fireside  corrosion  and deposition. Moreover,
lower-grade fuels (high moisture and ash)  can
be used without certain coal preparation steps
that   are   necessary   for  conventional
pulverized-fuel boilers.
                                     i

  Fluid-bed combustion as a boiler firing
technique first received attention in Europe as
a method of utilizing anthracite fines1'2  and
lignite3'4. Early in the 1960's, experimental
programs  on  fluid-bed combustion  were
undertaken  by England by the National Coal
Board  (NCB)  and British  Coal Utilization
Research Association (BCURA). NCB  has
concentrated on development of atmospheric
pressure  boilers for central  station  power
plants5 ; BCURA has worked on the develop-
ment of shell  boilers6  and  pressurized
combustors to be used in a gas turbine cycle7.
About 5 years ago,  research programs were
begun  in  the U.S. by Pope, Evans  and
Robbins8.9 and the Bureau of Mines1 °>'1.

  More recently in the U.S., The National Air
Pollution Control Administration (NAPCA),
through contracts with several organizations,
has undertaken a broad research program to
investigate fluid-bed  combustion because of
its potential for reducing emissions of  pol-
lutants from  fossil-fuel-fired boilers.  Under
NAPCA contracts,  both Argonne  National
Laboratory12   and   Esso  Research  and
Engineering Company13 have  been inves-
tigating the  use  of limestone  to reduce
emission of SO2 from fluid-bed combustors.
In  1970,  Consolidation  Coal   Company
reported on results of experiments on  coal
combustion in fluidized beds of dolomite and
regeneration of the dolmite14.

  As part of the NAPCA-sponsored program,
the Bureau of Mines conducted research to
assist evaluation of the technical feasibility of
fluid-bed combustion and to obtain engineer-
ing data with which to assess its commercial
potential. An  important consideration in the
development  of fluid-bed  boilers  is  the
amenability of various  U.S.  coals  to  this
combustion technique. Behavior of the ash
may be important, for  example. Coal ashes
that will remain in the combustion chamber
to serve as bed material could be an advantage
or a disadvantage, depending on the method
of sulfur removal.  For instance, if operation
                                       1-2-1

-------
with  limestone  beds  is  the  SO2-removal
scheme,  entrainment  of  ash  in the  POC
(products of combustion) and removal in this
manner probably would be preferred. If fine
limestone injection is used, retention of the
ash  as  the bed material  likely would  be
preferred. Combustion tests were conducted
by the Bureau on a variety of U.S.  coals to
determine   which  are  most   worthy  of
expanded investigations and to provide a basis
for  predicting  coal  types and  extent  of
reserves  that  would  be most suitable for
fluid-bed boilers. Results of combustion tests
are described in this paper.

EQUIPMENT AND PROCEDURE

   Figure  1  shows the 8-foot-high combustor
used  in  the   tests.  A cone-shaped  plate
perforated by 1/8-inch-diameter holes sup-
ports the fuel bed and admits  the fluidizing
air; water passing through a heat exchanger of
3/4-inch  pipe extracts  heat  from the  com-
bustion bed. Figure 2 is a flow diagram of the
system. Coal is fed pneumatically to  the base
of the fuel bed.  Combustion  products are
passed through two centrifugal separators for
removal of most of the entrained solids, and
then to a bag filter for further cleaning. Solids
from the first  separator can be reinjected into
the  combustion  zone.  Combustion gas is
monitored continuously for C>2, CC«2, and
CO.  After each test,  the residue  is removed
from the bottom of the combustor by a screw
conveyor.

   Startup is accomplished  in about  2 hours
by burning natural gas in the combustor, and
then  injecting  coal  or a coal/inert-solids
mixture  into  the combustion chamber. A
2-foot  bed  is  established at about  1200°F,
after which the natural gas is shut off and coal
or coal/inert-solids are fed at a relatively high
rate  until the desired  bed  level is attained.
Thereafter, coal is fed  at a rate  compatible
with the designated superficial air velocity.

   Originally, coal was introduced by a screw
feeder several inches above the air distributor.
1-2-2
Even weakly caking coals could not be fed
this  way, however, because the fuel particles
agglomerated before they  could mix with the
bed  material. When fed pneumatically at the
bottom,  however,  even  strongly caking  fuels
(such  as Pittsburgh bituminous)  could  be
burned satisfactorily without  agglomeration.
COMBUSTION TESTS

  Initial objective of the work was to deter-
mine if ash from the coal could be kept in the
combustion chamber to serve as bed material.
In fluid-bed combustion, the fuel volume is
much less than the bed volume, hence an inert
material (grog) is required. In initial tests with
anthracite and  bituminous  coal, most of the
ash was carried out with the combustion
products (POC), which  kept the ash content
of the bed  from reaching  the desired value
(about  99  percent). Operation  with a high
carbon   content  bed  was  possible,   but
combustion efficiency was poor, as evidenced
by excessive amounts of carbon and CO in the
POC.


  Attempts were  made to  burn the  coal in
various  inert solids, including sand, sintered
fly  ash, and  sintered  bottom  ash from  a
chain-grate stoker, but none of these worked
well.  Subsequently,  mullite,  zirconia, and
fused alumina  were tried  and found to be
physically stable. Fused alumina was used in
most of the tests.


  Six coals have been  tested,  ranging from
high-volatile (A) bituminous to low-volatile
bituminous. Chemical and size analyses of the
coals  are given in  tables 1  and 2. Results of
the combustion tests are  given in (tables  3
through 6.  All tests were made with 1/8-inch
x 0 coal.


  In  the first four tests (runs A-l, B-l, B-2,
and C-l), the coals were burned in -16 +48
mesh fused alumina; all components of the
system operated satisfactorily.

-------
   In   test   C-2,  with   high-volatile  (B)
bituminous  coal  (Indiana No. 5) in a bed  of
zirconia, steady-state conditions were main-
tained for 85  hours without any  operating
problems.

   In  tests  E-2  and F-2,  high-volatile (C)
bituminous coal (Central Illinois, No. 6 seam)
and medium-volatile bituminous coal (Indiana
County,  Pa.,  Lower Kittanning seam) were
burned  in  a bed of -14 +28 mesh alumina.
These coals are relatively high in ash, 22 and
24 percent.  Ash from both of these coals
remained in the bed and had to  be periodical-
ly removed to maintain a constant bed height.
In  both  tests,  recycle rates  from No.   1
cyclone  apparently were high, creating large
temperature differences between the bottom
and top of the combustion bed. This is caused
by the  design of the solids recycle system
which is outside the combustor.  Recycled
solids lose heat in passing through the recycle
system, and at high rates of recycle, this has a
chilling effect  on the lower portion of the
combustion  bed. This  problem was  more
severe in test F-2, and when a reduction  in
fluidizing  velocity   did   not  sufficiently
improve  the operation, material from cyclone
1 was not recycled to eliminate the problem.
In test E-2, coarser bed material was used  to
reduce the amounts of material being recycled
from  the first cyclone.

   Tests  G-l and G-2  were made with (a
low-volatile  bituminous   coal   (Somerset
County,  Pa.): For G-l, mullite was the bed
material; for G-2, A^O^- Test G-l  was made
without  reinjection   of the fines  from the
primary  cyclone, but in G-2 the fines were
reinjected.  In tables  3 and 4, G-l  shows an
excess air of 46.3 percent, which is incorrect
owing to a  leaking  relief valve in the  air
supply  line. This  accounts for  the  high
calculated  carbon  recovery  (145  percent)
owing to the apparent amount of CC«2 in the
POC.   Both  tests   with   the   low-volatile
bituminous   coal  were   terminated   early
because of difficulty with the combustor. In
both  cases, large amounts of  solids  were
deposited in the POC lines, causing excessive-
ly high backpressures. The reason for this is
not yet understood, but it may be important
in determining  the  suitability of coals for
fluid-bed   combustion.  Conventional  coal
analyses, including ash fusion  temperatures,
do  not appear  to adequately describe the
suitability of various coals Tor fluid-bed  com-
bustion.

  Runs  A-2,  A-3, A-4-A, and A-4-B  were
made with Pittsburgh-seam high-volatile (A)
bituminous  coal from West  Virginia,  with
comparable  results   (table  4),  even  with
variations  in  bed   height  and  fluidizing
velocity. The large excess air shown for A-2 in
tables 3 and 4 was caused by  the leak in the
air valve and was not actual, whereas the large
excess shown  in A-4-B was real and was due
to  malfunction  of  the  oxygen  sampling
system.

  Table  3 shows that carbon utilization  is
high,  usually  exceeding 99 percent,  except
when materials from  the first cyclone is not
recycled to the  bed.  In the two  cases when
fines  from the  primary  cyclone were not
recycled to the bed, carbon utilization, at 91
and 86 percent,  was unacceptable for  com-
mercial operation.

  Table 4 gives results of calculations for the
theoretical combustion of the coals tested at
the same air/coal ratios  used in the  tests.
Actual and theoretical gas analyses, compared
in table  3, show generally good agreement
except for oxygen content. In the first four
tests,  the discrepancy was traced to an  in-
correct  span  setting of the oxygen analyzer
for the  range  being used.  As  mentioned
previously, the discrepancy in tests G-l and
A-2 resulted from a leaking relief valve in the
air supply system.

  Table  5 gives mass balances for the  inert
solids for the combustion tests. For the first
four tests (A-l, B-l, B-2, C-l), only an overall
weight change is reported and ash behavior is
not  quantitatively  reported.  However,   it

                                      1-2-3

-------
 appears that for coal A there was  a gradual
 build-up of ash particles in the  bed, and for
 coals B and C virtually all ash was carried out
 in the POC. In subsequent tests, more samples
 were taken to account  for all  of  the inert
 solids and to permit identification of ash grog
 components. When washed coals containing
 less than 14  percent  ash were  burned and
 material from  the primary  cyclone was
 refined, most of the ash was entrained  in the
 POC. In tests E-2 and  F-2, in which high-ash
 coals were  burned, almost half of the ash
 remained in the bed. This was apparent during
 the tests because it was necessary to  withdraw
 material from the  bed during both tests  to
 maintain a constant bed height.

   Mullite,  zirconia, and alumina were used
 successfully as  bed materials. Losses of bed
 material were  low, ranging from  0   to  8
 pounds, which is a rate of loss of 0.1 percent
 per  hour.  The  recovery of inert  materials
 ranged  from 93  to 100 percent, except for
 one test. Considering the mass of material to
 be handled and the many places for solids to
 accumulate  in  the system, these values for
 recovery are acceptable.

   Heat transfer in a fluid-bed boiler will  be
 important  in  establishing  the   commercial
 potential  of  this  combustion  technique.
 Hence,  overall  coefficients were calculated
 directly from heat transfer  data from one
 U-tube  of  the  heat exchanger in  the  com-
 bustor by the relating Q = U0AoAt, where
   Table 6 shows the particle size distribution
of the dust from No.  2 cyclone and the filter
bags, as determined by a Coulter counter. The
cyclone effectively removed the dust down to
about 10-15  microns, with the filter bags
removing the plus 2-micron particles.
CONCLUSIONS

   A fluid-bed combustor was developed and
operated successfully with a variety of coals,
including highly caking types, under nominal
operating  conditions of 3 feet per  second
fluidizing velocity and 1600 °F bed tempera-
ture.  Mullite,  zirconia, and alumina worked
satisfactorily as bed materials, and attrition
losses were low. Carbon  utilization exceeded
99 percent when material from the primary
cyclone was refired; without refiring, utiliza-
tion  was  approximately  90  percent.  When
washed coals containing  less than 14 percent
ash were burned, virtually all of the ash was
entrained in the POC. In the case of two coals
containing  ash  in   excess of  20 percent,
approximately half of the ash remained in the
bed, indicating that for  these coals, the ash
could be the bed material. Overall coefficients
of heat transfer from the bed to a water-
cooled tube averaged about 53 Btu/hr-ft2 -°F
when the tube was immersed  in  the bed,
approximately the same as when the tube was
a few inches above the bed.
    Q  =  Heat removed, Btu/hr
    U0 =  Overall  heat-transfer  coefficient
          based  on  outside  area  of pipe,
          Btu/hr-ft2 - °F
    A0 =  Outside area of pipe, ft2
    At =  Average  temperature difference
          between water and fluid bed, °F
Overall coefficients, table 5, show very little
difference in magnitude when the tube was in
the bed or slightly above the bed, the average
being about  53 Btu/hr-ft2  °F.
BIBLIOGRAPHY

 1. Godel, A.A. Ignifluid, A New System of
   Combustion. Combustion Engineering As-
   sociation Document 7593. p. 1-22, July
    1963.

 2. The Ignifluid Stoker and Latest Improve-
   ments,  with  Special  Reference  to  the
   Combustion of  Untreated  Smalls and
   Anthracite  Duffs. Combustion Engineer-
   ing  Association Document  7781,  16 p,
   March 1964.
1-2-4

-------
3. Novotny,  P. Fluid-Bed Combustion  of
  High-Ash Coals. S.N.T.L. Technical Digest
  (Prague), No. 12. p. 883-891, 1965.

4. New Light on the Fluid-Bed Combustion
  of  High-Ash Fuels.  S.N.T.L. Technical
  Digest (Prague), No. 1. p. 12-15, 1968.

5. McLaren,  J. and D.F. Williams.  Com-
  bustion  Efficiency,  Sulphur  Retention
  and  Heat   Transfer  in   Pilot  Plant
  Fluidized-Bed Combustors.  Combustion,
  v. 41, No. 11, p. 21-26, May 1970.

6. Wright, S.J., H.C. Ketley, and R.G. Hick-
  man.  The   Combustion   of  Coal  in
  Fluidized  Beds for Firing  Shell Boilers.
  Journal of the Institute of Fuel, V. XLII,
  No. 341. p. 235-240, June 1969.

7. Hoy, H.R.   and  A.G.  Roberts.  Power
  Generation   via  Combined  Gas/Steam
  Cycles  and  Fluid  Bed Combustion  of
  Coal. Gas and  Oil  Power, July-August
  1969.

8. Bishop, J.W., E.B.  Robinson,  Shelton
  Erlich, A.K.  Jain, and  P.M.  Chen. Status
  of  the Direct Contact Heat Transferring
  Fluidized  Bed Boiler. Paper 68-WA/FU-4,
  presented   at Winter  Annual  Meeting,
  A.S.M.E.,  New York, N.Y., December 1-5,
  1968.

9. Pope,  Michael  and  John  W.  Bishop.
  Progress in  Fluidized-Bed  Boilers. Power
   Engineering, v. 74, No. 5. p. 46-49, May
   1970.

10. Coates, N.H., P.S. Lewis, and J.W. Eckerd.
   Combustion of  Coal in Fluidized  Beds.
   Paper  70-F-32,  presented   at  annual
   meeting,  Society of Mining Engineers of
   A.I.M.E.,  Denver,  Colorado,  February
   15-19, 1970.

11. Orning, A.A. and C.R. McCann. A Study
   of  Fluidized-Bed Combustion of  Coal.
   Presented at annual meeting, A.I.M.E.,
   New York, N.Y., February 25-29,  1968.

12. Jonke, A.A., R.L. Jarry,  and E.L.  Carls.
   Reduction  of Atmospheric Pollution by
   the Application of Fluidized Bed  Com-
   bustion.  Proceedings of the First  Inter-
   national  Conference  on  Fluidized-Bed
   Combustion,   sponsored   by   NAPCA,
   Hueston   Woods   State  Park,   Ohio,
   November 18-22, 1968.

13. Skopp,  A.  Use  of Fluidized  Beds of
   Limestone-Based  Materials for  Desulfu-
   rizing Flue Gas. Proceedings  of the First
   International Conference  on Fluidized-
   Bed Combustion, sponsored  by NAPCA,
   Hueston   Woods   State  Park,   Ohio,
   November 18-22, 1968.

14. Zielke, C.W., H.E. Lebowitz,  R.T. Struck,
   and  E.  Gorin.   Sulfur Removal During
   Combustion of Solid Fuels in a Fluidized
   Bed of Dolomite.  Journal  of the Air
   Pollution Control Association, v. 20, No.
   3, p. 164-169, March 1970.
                                                                                 1-2-5

-------
            Table 1. ANALYSES OF COALS BURNED IN FLUID-BED COMBUSTION TESTS
Constituent, %
Proximate analysis
Moisture
Fixed carbon
Volatile matter
Ash
Ultimate analysis
Carbon
Hydrogen
Oxygen
Nitrogen
Sulfur
Heating value (Btu/lb)

•*Major constituents of ash
SiO2
AI203
Fe203
Alkali metals in ash
P2°3
Ti02
CaO
MgO
NaO
K20
a,bA

1.2
53.3
36.5
9.0

74.10
5.32
6.13
1.27
2.98



44.1
23.0
18.3

1.0
1.8
5.1
1.9
1.1
1.6
"A,

1.3
50.0
37.5
11.2

73.05
4.72
5.46
1.37
2.90

12,260

45.9
24.3
15.8

-
—
—
—
0.8
2.0
B

2.3
49.7
39.8
8.2

72.27
5.54
6.35
1.47
3.87

12,930

38.5
19.5
—

-
-
_
—
—
—
C

5.2
46.6
38.9
9.3

68.01
5.24
7.00
1.32
3.99

11,940

42.3
19.5
28.0

0.9
1.3
2.9
1.3
0.6
2.1
E

1.2
36.5
40.2
22.1

51.92
4.57
15.15
1.02
4.06

9,290

50.4
19.5
15.5

—
-
—
—
1.4
1.8
F

0.9
55.5
19.4
24.2

63.83
4.11
2.43
0.98
3.50

11,300

47.6
27.4
16.5

-
—
—
—
0.2
1.4
G

0.6
67.4
18.1
13.9

75.31
4.19
1.97
1.23
2.83

13,085

44.3
25.9
21.7

-
-
—
—
0.3
1.7
a-dA2

1.3
54.1
34.2
10.4

73.47
5.23
6.35
1.47
1.80

—

43.7
21.3
19.3

—
-
—
—
0.7
1.7
       aHigh-volatile (A) bituminous, Pittsburgh bed, West Virginia
        Analysis A applies to Run A-1.
       cAnalysis AI applies to Run A-2.
        Analysis A2 applies to Runs A-3, A-4-A, A-4-B.
       eHigh-volatile (A) bituminous, Pittsburgh bed, Ohio
        High-volatile  (B) bituminous, Indiana No. 5 seam, Indiana
       9High-volatile (C) bituminous. Central Illinois No. 6 seam, Illinois
       _ Medium-volatile bituminous. Lower Kittanning seam, Pennsylvania
        Low-volatile bituminous. Upper Kittanning seam, Pennsylvania
        Analysis of material ashed at 800° C.
1-2-6

-------
Table 2. SIZE ANALYSIS OF COALS BURNED IN FLUID-BED
        COMBUSTION TESTS, WEIGHT-PERCENT

Screen
size,
mesh
+ 10
10 + 30
-30 + 60
-60 + 100
- 100+150
- 150 + 200
-200
Test


E-2
26.4
39.4
17.5
4.7
3.1
2.0
6.6


F-2
26.3
39.0
17.5
6.0
3.5
1.9
5.8


G-1
20.3
40.8
20.5
5.8
3.5
1.8
5.5
A-3
A-4-A
A-4-B
23.0
32.5
21.3
7.6
4.1
2.3
8.8


G-2
23.0
27.1
24.8
8.1
4.1
2.2
8.0
                                                             1-2-7

-------
            Table 3. RESULTS OF FLUID-BED COMBUSTION TESTS OF 1/8-INCH x O COALS

Bed size, U.S. sieve
Bed material
Duration, hr
Avg bed temp, u F
Bed height, in.
Superficial vel, ft/sec
Coal fed, Ib
Avg coal rate, Ib/hr
Avg air/coal, scf/lb
Theor air/coal, scf/lb
Excess air, pet
POC comp, pet
CO2
CO
02
S02
Carbon, Ib
In coal
From cyclone
From filter
In POC
Carbon recovery, pet
Carbon utiliz, pet
Sulfur, Ib
In coal
Cyclone solids
Filter solids
Bed residue
SO2 in POC, ppm
By difference
From gas analysis
Oxygen, Ib
In coal
In air
Total input
In POC (CO2, O2, CO, SO2)
As water vapor
Total output
Oxygen accounted for, pet
Run A-1
-16+48
AI203
69
1450
20
2.8
2296
33.3
145
133
8.6

15.2-16.2
0.0-0.5
0.0-0.3
0.16

1702
31
21
1586
96.2
96.9

68.4
2.0
2.2
0.3

2420
1600

141
5903
6044
4278
977
5255
86.9
Run B-1
-16 + 48
AI203
89
1650
20
3.0
2706
30.4
154
132
17.1

14.5-16.0
0.0-0.3
0.0-0.5
0.32

1956
11
7.5
1919
99.1
99.1

104.7
0.3
0.6
0.1

2930
3200

172
7423
7595
5283
1199
6482
85.3
Run B-2
-16 + 48
AI203
77.5
1635
20
3.0
2223
28.7
174
132
31.6

15.0-16.0
0.0-0.1
2.0-3.0
0.33

1606
5.8
2.8
1818
113.7
99.5

86.0
0.1
0.7
0.1

2735
3300

141
6858
6999
5723
986
6709
95.9
Run C-1
-16+48
AI203
OO
30
1635
20
3.2
964
32.1
160
124
29.2

15.0-15.6
0.0-0.1
1.1-2.0
IM.D.

766
1.1
1.3
720
94.3
99.7

38.5
0.3
0.4
0.2

2900
N.D.

67
2740
2807
2103

2507
89.3
1-2-8

-------
 Table 3 (continued). RESULTS OF FLUID-BED COMBUSTION TESTS OF 1/8-INCH x O COALS

Bed size, U.S. sieve
Bed material
Duration, hr
Avg bed temp, °F
Bed height, in.
Superficial vel, ft/sec
Coal fed, Ib
Avg coal rate, Ib/hr
Avg air/coal, scf/lb
Theor air/coal, scf/lb
Excess air, pet
POC comp, pet
C02
CO
02
S02
Carbon, Ib
In coal
From cyclone
From filter
In POC
Carbon recovery, pet
Carbon utiliz, pet
Sulfur, Ib
In coal
Cyclone solids
Filter solids
Bed residue
SO2 in POC, ppm
By difference
From gas analysis
Oxygen, Ib
In coal
In air
Total input
In POC (C02, O2, CO, SO2)
As water vapor
Total output
Oxygen accounted for, pet
Run C-2
-14 +28
ZrO2
85
1600
24
3.0
2121
25.0
195
124
57.0

12.45
0.0
5.0
0.25

1443
2
1
1514
105.1
99.8

84.6
0.7
1.6
6.6

2940 i
2500

149
7338
7487
5739
889
6628
88.5
Run E-2
-8+28
AI203
56
1488
14

2679
47.8
103
92
12.1

15.6
0.0
2.0
0.8

1391
17
6
1266
92.7
98.3

108.7
4.1
1.8
2.1

5820
8000

406
4925
5331
3981
979
4960
93.0
Run F-2a
-14+50
Mu & AI203
57
1543
14
2.4
1700
29.8
135
115
17.7

15.0
0.0
2.2
0.5

1085
87
6
1018
102.4
91.4

59.5
4.4
0.2
2.4

3820
5000

41
4088
4129
3204
559
3763
91.1
RunG-1a
-16+50
Mullite
24.7
1555
16
3.4
711
28.8
194
132
46.3

17.1
0.0
1.0
0.22

535
71
4
701
145.0
86.0

20.1
0.8
0.2
-

2220
2200

14
2448
2462
2004
233
2242
91.0
aWithout recycle from first cyclone.
                                                                                 1-2-9

-------
       Table 3 (continued). RESULTS OF FLUID-BED COMBUSTION TESTS OF 1/8-INCH x O COALS

Bed size, U.S. sieve
Bed material
Duration, hr
Avg bed temp, F
Bed height, in.
Superficial vel, ft/sec
Coal fed, Ib
Avg coal rate, Ib/hr
Avg air/coal, scf/lb
Theor air/coal, scf/lb
Excess air, pet
POC comp, pet
CO2
CO
02
SO2
Carbon, Ib
In coal
From cyclone
From filter
In POC
Carbon recovery, pet
Carbon utiliz, pet
Sulfur, Ib
In coal
Cyclone solids
Filter solids
Bed residue
SO2 in POC, ppm
By difference
From gas analysis
Oxygen, Ib
In coal
In air
Total input
In POC (CO2, 02. CO, SO2)
As water vapor
Total output
Oxygen accounted for, pet
Run A-2
-14+28
AI203
73
1569
18
3.1
1966
26.9
186
129
43.8

15.4
0.0
1.3
0.16

1436
6.7
3.5
1645
115.3
99.3

57.0
1.3
1.2
0.5

2361
1600

107
6506
6613
4805
742
5547
83.9
Run A-3
-14+28
AI203
42
1606
24
2.6
1210
28.8
145
131
10.8

16.7
0.0
2.2
0.1 6a

889
4.7
2.1
879
99.6
99.2

21.8
0.7
0.6
—

1860
1600

77
3126
3203
2675
506
3181
99.3
Run A-4-A
-14+28
AI203
71
1571
16
2.9
2289
32.2
145
131
10.2

16.2
0.0
2.0
0.13

1682
9.6
4.8
1596
95.8
99.1

41.2
1.5
1.4
—

1850
1270

145
5888
6033
4819
958
5777
95.8
Run A-4-B
-14+28
AI203
24
1585
16
3.1
662
27.6
181
131
37.6

16.4
0.0
2.0
0.08

487
2.3
1.5
583
120.7
99.2

11.9
0.5
0.5
—

1460
800

42
2126
2168
1754
277
2031
93
Run G-2
-14+28
AI203
20.5
1576
18
3.0
604
29.4
165
132
24.8

16.4
0.0
2.4
0.16

455
3.2
2.6
491
109.2
98.7

17.1
0.1
0.1
	

2700
	

12
1773
1785
1513
202
1715
96
   aFrom Run A-2 for calculations.
1-2-10

-------
                   Table 4. COMPARISON OF ACTUAL POC ANALYSES WITH ANALYSES CALCULATED FOR
                                 COMPLETE COMBUSTION AT SAME AIR/COAL RATIOS
POC
comp, %
C02
CO
02
N2
SO2, ppm
A-1
Theor
16.6
0.0
1.7
81.4
2400
Actual
15.2-16.2
0.0-0.5
0.0-0.3
—
2420
B-1
Theor
15.2
0.0
3.1
81.4
3000
Actual
14.5-16.0
0.0-0.3
0.0-0.5
—
2930
B-2
Theor
13.5
0.0
5.2
81.1
2700
Actual
15.0-16.0
0.0-0.1
2.0-3.0
—
2735
C-1
Theor
13.7
0.0
4.9
81.10
3025 .
Actual
15.0-15.6
0.0-0.1
1.1-2.0
—
2900
POC
comp, %
C02
CO
02
N2
S02. ppma
C-2
Theor
11.2
0.0
7.8
80.8
2470
Actual
12.4
0.0
5.0
82.3
2940
E-2
Theor
16.1
0.0
2.3
81.1
4730
Actual
15.6
0.0
2.0
81.6
5820
F-2
Theor
15.2
0.0
3.2
81.2
3130
Actual
15.0
0.0
2.2
82.3
3820
G-1
Theor
12.4
0.0
6.7
80.6
1750
Actual
17.1
0.0
1.0
81.7
2220
A-2
Theor
12.6
0.0
6.5
80.7
1870
Actual
15.4
0.0
1.3
83.1
2360
POC
comp, %
CO2
CO
02
N2
S02, ppma
A-3
Theor
16.3
0.0
2.1
81.5
1500
Actual
16.7
0.0
2.2
81.1
1860
A-4-A
Theor
16.4
0.0
2.0
81.5
1510
Actual
16.2
0.0
2.0
81.6
1850
A-4-B
Theor
13.6
0.0
5.8
81.0
1200
Actual
16.4
0.0
2.0
81.5
1460
G-2
Theor
14.6
0.0
4.3
80.9
2060
Actual
16.4
0.0
2.4
81.2
2700
aActual S02 calculated from sulfur balance:  sulfur in gas = sulfur in coal minus sulfur in recovered solids.

-------
                          Table 5. MASS BALANCE FOR INERT SOLIDS, LB
Test No.
Ash in coal, wt-pct
Bed material
Bed size
Heat-up Period:
Inputs: Grog
Anthracite ash
Test-coal ash
Total
Output: From POC
Bed materials end of heat up
Test Period:
Inputs: Bed material at start
Coal ash
Total
Output: Cyclone
Bag filter
Bed withdrawals
Bed residue
Total
Net weight change in bed:
Grog lost from bed:
Ash retained in bed:
Ave. Overall Heat Trans
Coefficient, Btu/hr/ft2/°F
A-1
9.0
AI203
-16+48
200
-
—
:

200
207
a407
98
101
0
275
474
+75
-
-
C67
B-1
8.2
AI203
-16+48
200
_
-
-

200
111
a422
64
79
0
195
338
- 5
-
-
C58
B-2
8.2
AI203
-16+48
200
-
-
-

200
183
a383
149
0
197
346
3
-
-
C56
C-1
9.3
AI203
-16+48
200
—
-
-

200
102
a302
41
37
0
218
296
+18
-
-
C54
              Does not include ash added during heat-up, but this amount is probably 10% or
                  less of ash input.
              Material withdrawn from  bed during test is considered as  bed residue for this
                  computation.
             cHeat exchanger tube in bed.
1-2-12

-------
OJ
Test No.
Ash in coal, wt-pct
Bed material
Bed Size
Heat-Up Period:
Inputs: Grog
Anthracite ash
Test-coal ash
Total
Output: From POC
Bed materials end of heat up
Test Period:
Inputs: Bed material at start
Coal ash
Total
Output: Cyclone
Bag filter
Bed withdrawals
Bed residue
Total
Net weight change in bed:3
Grog lost from bed:
Ash retained in bed:
Bed residue size, mesh:
+14
-14+28
-28
Ave. Overall Heat Trans
Coefficient, Btu/hr/ft2/°F
C-2
9.3
ZrO2
-14 +28

260
18
16
294
33
261

261
196
457
78
86
- o
275
439
+14
6
20





C38
E-2
22.1
AI203
-8+24

• 128
11
19
158
33
125

125
591
716
238
91
184
158
671
+217
8
225





d63
F-2
24.2
Mu+AI203
-8+24

125
13
27
165
28
137

137
411
548
221
29
159
166
575
+188
7
195





d54
G-1
13.9
Mullite
-16+50

150
.16
102
268
126
142

142
99
241
71
14
—
156
241
+14
3
17





d52
A-2
11.2
AI203
-14+28

175
15
14
204
30
174

174
221
395
75
80
_
211
366
+35
0
35





e52
A-3
10.4
AI203
-14 +28

247
16
16
279
28
251

251
136
387
50
49
54
211
364
+14
0
14





C53
A-4-A
10.4
AI203
-14 +28

150
16
18
184
37
147

147
239
386
94
91
65
136b
-
-
1
-





d51
A-4-B
10.4
AI203
-14 +28

-
—
—
—
-
136

136
69
205
30
34
22
162
248
+48
0
48

21
54
25

d50
G-2
13.9
AI203
-14+28

177
19
31
227
43
184

184
84
268
22
24
—
220
266
+36
1
37

10
80
10

e47
 Material withdrawn from bed during test is considered as bed residue for this computation.
 Calculated bed residue; test A-4-B was a continuation of test A-4-A.
cHeat exchanger tube immersed in bed.
 Heat exchanger tube 2 to 4 inches above bed surface.
eHeat exchanger tube at bed surface.

-------
              Table 6. SIZE ANALYSIS  (COULTER) OF SOLIDS FROM FLUID-BED
                        COMBUSTION TESTS, PERCENT BY WEIGHT
Particle
size, n

+70
-70+50
-50+30
-30+20
-20+15
-15+10
-10+7
-7+5
-5+3
-3+2

+20
-20+15
-15+10
-10+7
-7+5
-5+3
-3+2

B-2

C-2

E-3

F-2
Test
G-1
No.
A-2

A-3

A-4-A

A-4-B

G-2
Cyclone No. 2



4
10
36
23
17
9
1



1
6
33
24
19
15
2
13
5
10
10
13
49




20
15
30
16
11
8




34
16
27
11
6
6







3
8
34
22
18
13
2



3
5
28
26
21
15
2



4
7
34
27
16
11
1



4
8
33
24
19
12




4
6
29
22
22
13
4
Bag Filter

1
9
12
15
40
23

1
6
10
15
45
22
1
2
11
13
18
40
15
3
6
17
16
17
29
12
6
9
24
10
46
4


1
2
5
12
52
28
1
0
4
6
11
47
31

1
5
10
17
54
13

1
7
9
16
58
9

1
4
8
19
58
9
1-2-14

-------
CASTABLE INSULATING REFRACTORY

    4-inch  CARBON STEEb*. ${?•
           INSULATING-—j jj&y
                        flf'/iJ'j

                        H
                     FIREBRICK
                    CASTABLE
                   REFRACTORY

                  THERMOCOUPLI
                 SIGHT GLAS
                THERMOCOUPLE
                                                              .PRODUCTS OF
                                                               COMBUSTION
                                                                  (POC)
                            AIR

                           NATURAL GAS
                                                           *•— WATER
                                                   •-WATER
                                                 ^-IGNITER PORT
                                              BED SUPPORT AND
                                              AIR DISTRIBUTOR,
                                              TYPE 316 STAINLESS
                                              STEEL
                                             COAL-AIR MIXTURE
                             Figure 1.  Fluidized-bed combustor.
                                                                    To O2- CO, and CO2
                                                                         ANALYZERS
          AIR -i


 NATURAL GAS-i

  INERT GAS -i

I
BUSTOR

L J
                                                 CYCLONE
          \ A
                             ECEIVER
                           \
\/
                                                                            CYCLONE
                       COAL PLUS AIR
                                                                      RECEIVER
 SCREW CONVEYOR

*• AIR
                                                                       i MOTOR
       _ ^ ^ - ^ ^  -  -   .MOTOR
RESIDUE "" l-iiu;hIcREW
            CONVEYOR
               Figure 2.  Flowsheet for fluidized-bed combustion system.
                                                                                     1-2-15

-------
                                                3.  COMBUSTION OF
                                  CARBON-BEARING FLY ASH
                                  IN A  CARBON-BURNUP CELL
                           SHELTON EHRLICH

                         Pope,  Evans and Robbins
ABSTRACT
   An empirical model which predicts the
combustion efficiency for carbon in fly ash in
a  fluidized-bed  combustor  is  presented.
Independent variables considered are fuel feed
rate and  analysis, air  rate, bed temperature,
and static bed depth.

   The model predicts increasing efficiency
with increasing temperature, air rate, and bed
depth,  but   decreasing  efficiency   with
increasing fuel rate and inert matter in the
feed.

   The model may be used to  aid in the
design of a carbon-burnup cell.
conventional boiler (e.g., pulverized-coal or
cyclone-furnace)  can  achieve   combustion
efficiency   approaching   100   percent,
fluidized-bed  boilers  would not  be com-
mercially feasible  unless the  fuel lost as
carbon in fly ash were markedly reduced.

   A method proposed by John  Bishop1-2
involves firing the  collected fly  ash in an
auxiliary  combustor.  The  name  "carbon-
burnup  cell"  was  coined for this auxiliary
combustor.

   NAPCA requested that Pope, Evans and
Robbins  conduct  a  series  of  statistically
designed tests to model the performance of
the carbon-burnup  cell. This paper describes
these tests and their results.
INTRODUCTION

   The   National  Air Pollution  Control
Administration (NAPCA), Division of Process
Control Engineering, has sponsored work o'n
various applications of fluidized-bed com-
bustion to  reduce  noxious emissions from
power generating plants. One such application
is the fluidized-bed boiler in which coal (or
possibly residual  fuel  oil)  is  burned in  a
fluidized  bed of noncombustible  particles.
These particles may be selected from a wide
range  of suitable  substances, some of which
would be reactive with sulfur.

   When  coal is the fuel to be fired in  a
fluidized-bed  boiler, an unacceptably high
fraction  of the energy of the input fuel
appears as  a high-carbon  fly  ash.  Since  a
BACKGROUND

A Description of the Problems

   The problem of carbon loss from a coal-
fired fluidized-bed  combustion  device was
recognized by Winkler3 and the fly-coke from
Winkler gas generators was used to fire nearby
boilers4.  Even where  the  combustor was
oxidizing,  as  in  GodePs  Ignifluid5, the
complete  combustion of carbon values could
not be achieved in a single pass through the
fluidized-bed combustor.

   Conventional    boilers,   using   either
pulverized fuel or  a cyclone furnace, are
capable of complete utilization of the fuel in
a single pass. The reasons for this performance
                                       1-3-1

-------
are clear—the combustion temperature is very
high,   approaching   the   adiabatic   flame
temperature,  and the furnace is very large.
The reasons for the high conversion efficiency
of the conventional boiler lead,  however, to
certain fundamental deficiencies—large size,
high cost, and the inability to practice on-site
capture  of  sulfur—which the  fluidized-bed
boiler is intended to overcome.

    Development of a commercially feasible
fluidized-bed  boiler, therefore,  requires the
development  of  techniques  and hardware
which   would  permit  a   compact   and
economical boiler to be constructed  which
would  still  achieve  a high combustion  ef-
ficiency.

    The   requirements   for   combustion
efficiency have been understood for some
time  and are described in a rule-of-thumb by
combustion engineers  as the three T's—time,
temperature,-and turbulence. A fluidized-bed
boiler in its simplest form cannot meet all of
these requirements.

    A number of  means have been proposed
for  achieving  the desired  degree  of com-
bustion efficiency: increasing the time  factor
by  recycling;  or increasing the  temperature
factor  by  operation  in  an  agglomerating
mode.   Turbulence   has    already   been
maximized by use of the fluidized bed  and
attempts to increase this factor would not be
fruitful. However, the increase in excess air is
similar in effect to  turbulence,  causing more
oxygen to be made available at the particle
surface.

    While  it  was  known that  combustion
efficiency would be favored by high tempera-
ture,  deep beds  (or  very  low superficial
velocities), and high excess air,  nothing  was
known  of the relative importance of these
variables or how other factors might influence
combustion.

   A series  of statistically designed experi-
ments might answer some of the questions
about the  performance  and design  of a
1-3-2
carbon-burnup cell. Certainly the system was
too complex to treat analytically. A statistical
approach  could  minimize  the  number of
experiments and  maximize  the  information.
To this end, and with the aid of a consultant
skilled in the design of experiments and their
interpretation, we set out to design a series of
experiments.
    We had  some  difficulty  in  finding an
individual who possessed the necessary train-
ing and experience and who was also capable
of relating to engineers and the realities of an
engineering  apparatus. We did find Professor
Arthur Hoerl of the University  of Delaware,
who did meet our requirements. Much of our
understanding of the carbon-bumup process is
due to Professor Hoerl's skill.

The Variables
    An important first  task in  the  program
was to list all of the potential variables and
then  to categorize  them as independent (or
control variables),  dependent  (or  response
variables),  or design variables.  It has  been
clear  for  some time that  certain important
variables   which   have   previously   been
categorized  as  control variables are, in  fact,
responses, and so one purpose of this paper is
to set this matter straight.
    Tables 1 through 6 identify and categorize
these variables. Note that the division of the
variables  into  six  groups  is  arbitrary-all
control and design variables will influence all
responses.
 Table 1. COMBUSTION AIR FEED VARIABLES
A.   Control variables

     1.   Mass flow rate, Ib/hr ft2
     2.   Temperature above ambient, ° F
     3.   Pressure, atmospheres
     4.   Composition (e.g., % oxygen and % hu-
            midity)
B.   Design variables
     1.   Distributor design
     2.   Injection point of secondary air
     3.   Arrangement of secondary injection sys-
            tem
C.   Response variables
     None

-------
         Table 2.  FUEL FEED VARIABLES

 A.    Control variables

       1.    Mass flow rate, Ib/hr ft2
       2.    Temperature above ambient," F
       3.    Composition (e.g., % carbon, % inert di-
               luent, and % moisture)
       4.    Particle size distribution

 B.    Design variables

       1.    Injector design
       2.    Injector location
       3.    Recycle rate

 C.    Response variables

       None

 D.    Natural material variables

       1.    Ignition temperature
       2.    Porosity
       3.    Ash particle size
       4.    Ash sintering temperature
       5.    Degree of graphitization of carbon
       6.    Ash/carbon matrix characteristics
       7.    Attrition characteristics
       8.    Swelling characteristics
       9.    Surface area
      10.    Shape factor
      11.    Density
      12.    Emissivity
      13.    Diffusivity


            Table3.  BED VARIABLES

A.    Control variables
      1.    Mass, pounds (weight)/square toot of grid
              area
      2.    Static depth, ft
      3.    Composition (e.g., nominal particle size,
              density, thermal properties, and hard-
              ness)
      4.    Temperature history (e.g., is the steady
              state temperature approached from a
              higher or lower initial temperature)
B.
Design variables

1.
           Combustor  shape  (e.g.,  constant or
              conical cross section)
                                                      2.    Internals (e.g., arrangement of heat trans-
                                                               fer surfaces and baffle screen)
                                                      3.    Circulation of hot or cold particles
                                                      4.    Heat sink (e.g., surface area, effectiveness,
                                                               and temperature)
                                                      5.    Solids removal method
                                                      6.    Cross-sectional area, ft2

                                                 C.   Response variables

                                                      1.    Nominal or average bed temperature," F
                                                      2.    Bed temperature distribution; e.g., T(x,y,z)
                                                      3.    Effective thermal conductivity—axial and
                                                               horizontal, Btu/ft2hr°F/ft
                                                      4.    Diffusivity—axial and horizontal, ft2/hr
                                                      5.    Expanded bed height, ft
                                                      6.    Bed density distribution; e.g., p (z)
                                                      7.    Bed circulation patterns
                                                      8.    Equilibrium particle characteristics (e.g.,
                                                               particle size and shape)
                                                      9.    Residence time distribution of fine parti-
                                                               cles
                                                      10.    Fines capture rate through agglomeration
                                                      11.    Ash particle size increases through agglom-
                                                               eration
                                                      12.    Viscosity of bed
Table 4. COMBUSTION-GAS STREAM VARIABLES

A.   Control variables

     None

B.   Design variables

     1.    Arrangement of above-bed heat transfer
              surfaces

C.   Response variables

     1.    Mass flow rate, Ib/hr (or superficial veloc-
              ity ft/sec)
     2.    Final composition (e.g., % ©2, N2, CO2,
              CO, HxCy, SOX, NOX, and H2O)
     3.    Composition distribution; e.g., O2(x,y,z)
              and S02(x,y,z)
     4.    Temperature distribution
     5.    Final or average temperature," F
     6.    Electrical properties (for electrostatic pre-
              cipitation)
     7.    Solids loading, grains/scf

-------
      Tables. SOLIDS STREAM VARIABLES

 A.   Control variables

      None

 B.   Design variables

      1.    Freeboard (disengaging height)
      2.    Above-bed baffles
      3.    Secondary air injection

 C.   Response variables

      1.    Mass flow rate, Ib/hr
      2.    Final composition (e.g., % C, H2, S, N2,
             Cl, and Na+K)
      3.    Size and density distribution
      4.    Resistivity (for electrostatic precipitation)
      5.    Carbon content vs size distribution
        Table 6. REACTOR VARIABLES

 A.   Control variables

      1.    Position of cooling coils
      2.    Coolant temperature

 B.   Design variables

      (Covered under fuel, air, bed, and products of
        combustion)

 C.   Response variables

      1.    Steam, generated or superheated
      2.    Erosion of surfaces
      3.    Corrosion of surfaces
      4.    Slagging of surfaces	


 Variables Actually Studied

    While  the listing (Tables 1  through 6) is
 extensive,  it  cannot  include  all  of  the
 variables; it would have been impossible even
 to attempt to study all of those listed.

    It was necessary, therefore, to select a
 limited  number  of variables  which  would
 allow some understanding of the process and

1-3-4
hopefully  answer the basic question-wili a
carbon-burnup cell work?

    Table 7, below, lists the control variables
selected, the range  suggested  at  the start of
the study,  and the range actually investigated.
Note that for the purposes of a test program,
temperature,   excess   air,  and   superficial
velocity  are  considered   control  variables
despite the fact that they are responses.

Table 7. CONTROL VARIABLES  SELECTED FOR
   CARBON-BURNUP CELL INVESTIGATION
       Variable
                                                  1. Carbon content of fly
                                                       ash,%

                                                  2. Particle size of fuel
3. Static bed depth, in.

4. Nominal bed tempera-
     ture, ° F

5. Excess air, %

6. Superficial velocity, fps
   Range     Range
  Initially   Actually
 Suggested   Studied
 20 to 80   23 to 65

As received As received
and 70%
through 200
mesh

   6 to 18   10 to 22

1500 to    1700 to
   2100       2150
Oto50

5 to 20
10 to 90

6 to 15.3
    In  addition  to  the  control  variables
indicated in Table 7, it was found desirable to
add some special experiments which  might
point out design considerations. These special
tests included the addition of coal and opera-
tion with two bed particle-size distributions.

    The  effect of coal addition and particle
size could not  be included in the modeling
process.  Neither was found  to  be of  major
importance over the range studied.

    Table 8 lists the data for those tests which
were useful in the model.

-------
Table 8. EXPERIMENTAL TEST CONDITIONS AND RESULTS
Run
No.
C-302-1
-2
-3
C-303-1
-2
-3
-4
-5
C-306-1
-2
-3
-4
C-307-1
-2
-3
-4
C-308-1
-2
-3
C-310-1
-2
C-311-1
-2
-3
-4
-5
-6
C-312-1
-2
-3
C-313-1
-2
-3
C-314-1
-2
-3
C-315-1
-2
-3
Bed
Temp,
°F
1800
2140
1780
1750
1980
1750
1860
1970
2100
1750
2120
1800
1800
2100
2140
T750
1900
1900
1900
2040
2000
2000
1930
2000
1980
1950
1800
2000
1990
1980
1980
1980
1980
1980
1960
1980
2010
2000
1980
Air
Rate,
Ib/hr
760
760
820
330
330
500
500
500
600
860
775
600
600
330
600
330
600
600
600
700
700
700
700
600
600
800
800
700
700
700
600
700
800
600
700
800
600
700
800
Bed
Depth,
in.
18
18
18
18
18
18
18
18
18
18
18
18
17
18
18
18
18
18
18
18
18
18
18
18
18
18
18
18
14
10
18
18
18
10
10
10
18
18
18
Carbon
Rate,
Ib/hr
40
53
44
21
24
39
37
39
41
56
53
43
41
26
44
29
38
42
43
46
55
52
63
79
51
75
96
47
46
51
37
43
48
39
46
57
46
49
52
Inert
Rate,
Ib/hr
33
52
28
20
22
29
30
32
65
78
77
62
65
39
66
44
49
58
59
59
70
106
127
171
107
157
206
93
92
105
26
31
36
31
36
45
35
37
40
Part. Cooling CE,b
Size3 Coil %
+ 90
+ 83
-I- 86
+ 65
+ 66
+ 65
+ 74
+ 78
86
70
89
73
+ 76
+ 90
+ 91
+ 63
+ 81.5
+ 81.3
H- 80.6
+ 84.6
+ IN 75.8
+ 77.0
+ IN 66.7
+ IN 58.0
+ 75.8
! + 64.5
+ IN 56.2
+ 84.2
+ 83.8
+ 76.0
+ 91.0
+ 87.2
+ 86.0
+ 84.8
+ 82.0
+ 76.0
+ IN 83.3
+ IN 84.5
+ IN 85.2
C02,
11.5
14.2
10.8
9.8
11.5
12.0
13.2
14.5
14.3
11.0
14.6
12.4
13.2
16.8
16.0
13.4
13.4
14.5
15.0
13.2
13.8
16.0
16.6
18.1
15.9
15.8
16.4
13.8
13.5
13.2
13.3
12.9
12.8
13.4
13.0
13.0
15.4
15.0
14.6
02,
9.1
6.4
9.7
10.8
8.0
8.5
7.4
6.2
6.2
9.3
5.8
8.2
7.5
3.5
4.8
6.9
7.1
6.2
5.5
7.1
6.4
4.5
3.6
2.1
4.5
4.7
3.9
7.0
7.2
7.5
7.5
8.0
8.1
7.4
7.6
7.8
5.5
5.8
6.0
S02,
ppm
350
800
550
320
600
320
500
700
800
750
800
750
500
600
650
700
600
750
800
700
720
700
800
850
820
(820)
(820)
900
900
900
1160
1160
1160
1050
1050
1050
1250
1250
1250
NO,
ppm
560
670
570
470
480
410
570
600
720
560
770
520
410
500
550
370
520
500
520
460
420
570
470
470
500
510
350
500
480
480
630
610
600
560
560
560
560
560
540
                                                           1-3-5

-------
Table 8 (continued).  EXPERIMENTAL TEST CONDITIONS AND RESULTS
Run
No.
C-316-1
-2
-3
C-317-1
-2
C-318-1
-2
Bed
Temp,
°F
2000
1990
1980
1980
1980
1900
1950
Air
Rate,
Ib/hr
600
700
800
700
700
700
700
Bed
Depth,
in.
10
10
10
22
18
22
18
Carbon
Rate,
Ib/hr
50
58
63
52
51
61
56
Inert
Rate,
Ib/hr
36
36
46
43
43
43
41
Part. Cooling
Size3 Coil
+ IN
+ IN
+ IN
_

+
+
CEb,
74.5
76.5
72.0
82.2
81.5
72.5
79.8
C02,
14.8
14.3
13.8
15.5
15.3
15.8
16.0
02,
5.8
6.0
6.4
4.8
5.3
5.0
4.5
S02,
ppm
1250
1300
1250
1200
1275
1050
1050
NO,
ppm
520
510
520
570
570
550
550
 a+ = -16 +22 Tyler mesh
  - = - 8 +16 Tyler mesh
  Combustion efficiency

RESULTS

    For the purpose of modeling, the control
(or independent)  variables  were slightly  re-
arranged and included the following:
            Variable             In
       Bed temperature        °F
       Air rate                Ib/hr
       Static bed depth        in.
       Carbon rate            Ib/hr
       Inert rate              Ib/hr

   These  independent  variables  were  con-
sidered  in  various  alternative  regression
modes.
   Empirical regression models of the type
were  successively  fitted  to  the  data  and
evaluated  through  their  corresponding  F-
ratios. the F-ratio is a statistical measure of
the variance in the  response data which  is
explained  by  the  model  divided by  the
residence or unexplained  variance. If there
were no relationship between the independent
variables and the response, the F-ratio would
tend toward unity. Therefore, the higher the
F-ratio,  the  more  reliable  the  model is in
explaining the trends in the data.
  The inclusion of  bed temperature as an
independent variable is unusual: it should be a
1-3-6
response to air  rate, feed rate, and percent
carbon. However, a  satisfactory model could
not  be  developed  without including bed
temperature as a pseudo-independent variable.

   The fly-ash feed variable was examined in
two ways:

   Group 1    Fly ash feed rate,     Ib/hr
              Carbon content,      %

   Group 2    Carbon feed rate,     Ib/hr
              Inert feed rate,       Ib/hr
Although these two groupings are identical, it
was  found that using the  cross-product of
carbon-feed   rate    times   inert-feed   rate
improved the reliability of the model.
   Various alternative  forms were  used for
representing the  combustion efficiency (CEV
     1. CE.
     2. Log(l-CE/a).
     3. LogCE.
   The purpose of this step was to develop a
response  form  consistent with the physical
case; i.e.,   combustion  efficiency  cannot
exceed 100 percent. The form

              Log(l-CE/a)

with a = 95 to 100 percent would be the
preferred form. Table 9 is a summary of the

-------
 F-ratio for CE, Log (CE), and Log (1-CE/a),
 with various values of a.

 Table 9. F-RATIO  FOR  VARIOUS FORMS OF
 REPRESENTATIVE COMBUSTION EFFICIENCY
Table 11. OBSERVED AND CALCULATED
   COMBUSTION EFFICIENCY
Form of Combustion Efficiency (CE)
CE
Log
Log
Log
Log
Log
Log

(1-CE/95)
(1-CE/101)
(1-CE/110)
(1-CE/120)
(1-CE/145)
(CE)
F-Ratio
74.5
46.2
69.2
72.1
74.2
75.1
73.8
  These  results show that the simple form,
CE, gives the best result except where a takes
on unreal values above 120 percent. The form
CE was therefore chosen for the response.
  The final  form of the model selected was:

     Combustion Efficiency = b0 +  bj (bed
       temperature, °F)
     + b2 (air rate, Ib/hr) + b3 (bed depth,
       inches)
     + b4 (carbon rate, Ib/hr) + b5 (inert rate,
       Ib/hr)
     + bg  (carbon rate, Ib/hr) (inert rate,
       Ib/hr).

  Table 10 lists the values for the coefficients
bQ through bg.

  Table 10. REGRESSION COEFFICIENTS FOR
      COMBUSTION EFFICIENCY MODEL


Term
Constant
Bed temperature, °F
Air rate, hr/lb
Bed depth, in.
Carbon rate (CR), hr/lb
Inert rate (IR), hr/lb
(CR) (IR).hr2/lb2
Designation
of
Coefficient
bo
b2
b3
b5

Value of
Coefficient
-13.78
0.051 93~1
0.04620
0.3831-1
-0.8737
-0.1905
0.0027
  Table 11 shows the calculated and observed
values of combustion efficiency.
Run No.
C-306-1
-2
-3
-4
C-307-1
-2
-3
-4
C-308-1
-2
-3
C-310-1
-2
C-311-1
-2
• -3
-4
-5
-6
C-312-1
-2
-3
C-313-1
-2
-3
C-314-1
-2
-3
C-315-1
-2
-3
C-316-1
-2
-3
C-317-1
-2
C-318-1
-2
Obsrvd
86
70
89
73
76
90
91
63
81.5
81.3
80.6
84.6
75.8
77.0
66.7
58.0
75.8
64.5
56.2
84.2
83.8
76.0
91.0
87.2
86.0
84.8
82.0
76.0
83.3
84.5
85.2
74.5
76.5
72.0
82.2
81.5
72.5
79.8
Calc
88.9
71.7
89.1
72.1
73.3
90.0
88.8
69.0
82.0
78.3
77.5
87.3
78.3
78.6
68.0
59.5
73.4
67.7
53.8
82.3
81.0
75.1
89.0
88.4
88.8
83.9
81.6
78.4
82.7
84.4
85.5
76.0
73.8
73.8
82.2
81.4
71.2
76.2
Result
-2.9
-1.7
-0.1
0.9
2.7
0.0
2.2
-6.0
-0.5
3.0
3.1
-2.7
-2.5
-1.6
-1.3
-1.5
2.4
-3.2
2.4
1.9
2.8
0.9
2.0
-1.2
-2.8
0.9
0.4
-2.4
0.6
0.1
-0.3
-1.5
2.7
-1.8
0.0
0.1
1.3
3.6
                                                                                 1-3-7

-------
CONCLUSIONS

   The results of tests  and the form of the
model indicate that combustion efficiency is
favored by  increasing temperature,  air rate,
and  bed depth, and by decreasing fuel feed.
These   results  are  consistent   with  the
predictions made at the beginning of the test
program, with  one  exception; i.e., the model
predicts increasing  efficiency with increasing
superficial velocity. This was unexpected and
in fact the results tend to show no effect for
velocity.  Combustion  efficiencies  of  90
percent can  be achieved  in burning fly ash
from a fluidized-bed boiler in a single pass
through  a carbon-burnup cell. This would
then indicate combustion efficiencies  of 99
percent  for  a  commercial  boiler which
included a carbon-burnup cell.  It would be
expected that a full-size burnup cell (at least
10 ft2) would give better performance and
that  additional  recycle  of  flyash would
further improve burnup.
BIBLIOGRAPHY

1. Bishop,   J.W.  et al.  Status of  Direct
  Contact Heat  Transferring Fluidized-Bed
  Boiler. Paper No. 68-WA/FU-4  delivered
  to Fuels Division  of the  ASME  Winter
  Annual Meeting, December 1-5, 1968.
2. Bishop, J.W. U.S. Patent No. 3,508,506.
3. Winkler, F. U.S. Patent No. 1,687,118.
4. Von Portatius,  B.  Freiburger Forschung-
   shefte. A 69, 5 (1957).
S.Godel,  A.A.  Revue  Generate  de  Ther-
  mique. 5,349(1966).
 1-3-8

-------
                            4.   PILOT-PLANT EXPERIMENTS
                                                                 AT BCURA

                                   H.R. HOY

                        National Coal Board, England
INTRODUCTION

  The objective of the programme of experi-
mental work at BCURA is to provide data for
designing  steam  generators ranging  in  size
from  small industrial units (e.g., 8000 Ib/hr)
to large central power station units (e.g., 660
MW).  A large proportion of the effort has
been  directed toward pressurised fluidised
combustion mainly  for combined gas/steam
power generating plants. The BCURA re-
search programme is complementary  to that
in progress  at the  National Coal  Board's
Central  Research  Establishment at Chelten-
ham.

  Since  the   last  conference  at Hueston
Woods:

  1. Research  on   small  industrial  steam
    generators has been taken to the proto-
    type stage.

  2. A  programme  of  research  to  obtain
    specific data for the design of industrial
    steam generators has been completed.

  3. A pressurised pilot-scale combustor has
    been operated  (initially  to  assess  the
    suitability of the combustion gases for
    gas turbines and subsequently to obtain
    process data for design studies).

  4. A  major  programme   of  pilot- and
    laboratory-scale investigations on sulphur
    oxide retention has begun.

  These  activities have  been supported by
"cold"  model work and  rationalised  by
mathematical models.
EXPERIMENTAL TECHNIQUES AND
EQUIPMENT

  The main pilot-scale  equipment has been
described previously; it comprises:

  1. A firetube boiler capable of generating
    8000 Ib/hr of steam at 100 psi, with a
    48-inch diameter cross-section fluid bed.

  2. A  fully cooled pilot-scale  combustor
    with  a  27-inch  diameter  bed cross-
    section.

  3. A  48  by  24 inch  bed  cross-section
    pressure combustor capable of operating
    at pressures up to 6 atmospheres.

  The principal features of the combustors
are shown in Figures 1, 2, and 3.

  A common feature of all is that a major
part of the heat transfer surface is contained
in the bed.

  A  fairly  wide  range of fuel  sizes  and
fluidising velocities  have been  investigated.
High velocities are  attractive from the point
of view of minimising containment costs but
the higher heat transfer rates result in lower
tubing costs; the optimum situation remains
to be proved.
INFORMATION OBTAINED

Startup

  A number of techniques have been used,
ranging from preheating the combustion air to
                                       1-4-1

-------
burning gas in a bed with higher-than-normal
coal content. The latter technique, one form
of  which was first  reported by Pope, Evans
and Robbins, is particularly effective when
lighting up  a  heavily  cooled  bed.  Before
adopting this, we  did  not  achieve self-
sustained combustion in the prototype boiler.
The main snag with this startup system, unless
it  is  carefully applied,  is smoke emission;
however, a wholly satisfactory technique was
evolved before the small industrial  boiler
programme was completed.

Combustion

  The  principal loss of combustible material
is in the form of solid carbonaceous material
elutriated  from  the  bed.  As  might  be
expected, losses are greater at high than at
low fluidising velocities;  better combustion
efficiencies  are currently  attained under the
latter conditions.  Recirculation  of carryover
material is  essential even at  low velocities
(e.g.,  2 ft/sec) to  enable high (99 percent or
greater) combustion efficiency to be attained.

  Experiments have been carried  out so far
on  1/16-in. x  0, 1/8-in. x 0, and 1/4-in. x 0
coal; the relevant fluidising velocities for these
have been 2 ft/sec (pressure combustor), 4-8
ft/sec, and 10-12 ft/sec, respectively.

  The  carbon-rich particles of the elutriated
material are  mainly (though not universally)
concentrated  in the finer size fractions (e.g.,
<50 Aim) and more than one pass through the
bed are needed to achieve a reasonable degree
of  burnout.  The amount of material to  be
recirculated can be large enough to affect the
heat balance of the bed.

  There are  a number of ways in which the
problem can be alleviated. In the first place,
careful  design  of the coal preparation equip-
ment  can keep the amount of fines  (e.g.,
material smaller than 75 Aim) in the coal feed
to a low level. Reducing the amount of fines
from about 18 percent to about 8 percent on
the pressure  combustor  has, for example,

1-4-2
improved combustion efficiency by two units
of percentage.

  Classification of the carryover material into
the  appropriate  size  ranges,  rejecting the
mainly  ash  material, and recycling only the
carbon rich stream, is one way of reducing the
thermal  load on  the bed  while, at the same
time, improving combustion efficiency.

  An   attractive   alternative   now  being
explored is  using a separate burnout bed for
reacting  carbon  in  the elutriated material.
Present  indications  are that low  fluidising
velocity  and/or higher bed temperatures (e.g.,
>900°C) are needed to  achieve a high degree
to burnout of the material elutriated from the
combustor   operated   at  high   fluidising
velocities.

  More  data is needed  before a final choice
could be made between recycling elutriated
carbon to the main bed or to a separate bed;
the solution  adopted may well depend upon
specific features of the application.

Coal Distribution

  The attainment of a high combustion ef-
ficiency  (and  a  minimum  of  combustion
above  the bed)  is also strongly dependent
upon rapidly achieving uniform  distribution
of coal over the cross-section of the bed.

  Diffusion of a  coal stream admitted at the
bottom of the bed is more rapid  in a vertical
than in  a horizontal direction-as might be
expected-and from this point of view alone a
large number of  coal nozzles are needed to
ensure good lateral dispersion of the coal ink
big fluidising bed.

  The extent to  which lateral dispersion can
be assisted by introducing coal streams at high
velocities in a   lateral direction has  been
explored  in a  5-ft  diameter  cold  model
(situated at  the National Coal Board's Central
Research Establishment).  The rate of mixing
was deduced from the distribution of radio-

-------
activity in samples  extracted from  the  bed
after  a pulse of irradiated coal had been fed
through a nozzle into the bed.

   It  was  found   that  mixing  could  be
described in terms  of diffusion  theory  and
that,  over  the  limited  range  of variables
investigated,  the  coefficient of diffusion was
proportional to the  fluidising velocity and to
the momentum  of the  coal  particles. In-
creasing  the  momentum  of  the  particles by
the use of high velocity air streams was found
to  be difficult,  however, due to a  "slip"
between the coal and air streams.

   In  general  it is also difficult  to increase
momentum  to an effective  extent without
also increasing the quantity of transport air.
Other experimental work has shown, how-
ever,  that  combustion  of  coal fines  and
volatiles  above the  bed  is increased  if the
quantity  of  transport air exceeds about  2
percent of the total combustion air.

   The general conclusion is that other means
need to be found  for increasing the amount of
lateral diffusion achievable from a single coal
inlet point. Work to this end is in progress at
BCURA and at CRE.
Measurement of Coal Flow

  A number  of methods are available for
trimming the flow of air/coal mixtures so as
to  achieve  reasonably  uniform  division of
flow between streams travelling along a multi-
tude of pipelines.

  The problem is  to measure the amount
flowing; because there will be a large number
of pipelines, an inexpensive device is needed.
We have tried  a number  of systems; experi-
mental work continues, but we are far from
being able to say, to within even ± 10 percent,
how much  coal is flowing in  a pipeline.
Reliable  detection of whether or not coal is
flowing, however, can be achieved by modern
proprietary ultrasonic devices.
Heat Transfer

   The initially predicted high values for heat
transfer to  surfaces immersed  in fluid beds
have   been   confirmed.   Horizontal  tubes
arranged on staggered pitches give the highest
values—and   correlations  (one  example is
shown in Figure 4) between bed particle size
and heat transfer coefficients valid to within ±
15 percent have been evolved.

   Higher coefficients are achieved to hot than
to cool surfaces but there is more  scatter in
the  data for the  former,  most  probably
because  of  the   difficulty  of accurately
measuring the temperature of metal surfaces
at high temperatures.

   Higher than expected  heat  transfer rates
were achieved in the prototype boiler mainly
because  the design was  based on  measure-
ments made to "cold" tubes. Some  tubes had
to be cut out to eliminate over-cooling which
caused unsatisfactory effects on combustion.
ASH BEHAVIOR

   The  coals tested have generally  had ash
contents in the range of 15-25 percent. No
difficulties  have   been  experienced  with
clinkering with bed temperatures maintained
below about 950°C.

   The ability  to maintain bed level  depends
on the nature of the ash (its friability) as well
as the quantity available. An empirical test
(evolved by the  National  Coal Board) has
been used to assess the degradability of coal
ashes;  most of the United  Kingdom  coals
tested  so  far  appear  to  have ashes  with
reasonably  good  resistance to degradation.
The  penalty of adding  material (e.g.,  coal
washery shale) to maintain bed level—where a
coal ash was too friable—would not, however,
be serious.

  In the experimental rigs  operated at low
fluidising velocities, about 20 percent of the
                                      1-4-3

-------
input ash is extracted from the bed; however,
at  high  velocities, most is carried out  and
normally only enough is retained to maintain
bed level.
  While it is realized that only tests with an
actual   turbine   will   provide   convincing
evidence,   experience   to   date   is  highly
promising.
   It   has  been  found  that  by  suitably
 arranging  closely  spaced horizontal  tube
 surface at the top of the bed, more ash  is
 retained  in  the  bed—and more combustion
 occurs within the bed.
 Deposits and Corrosion of Surfaces
 in the Bed

   Specimens of boiler superheater alloys have
 been  exposed  to  the  conditions in  the
 pressure combustor bed for several hundred
 hours. The results have generally been similar
 to  those  obtained in the  specially designed
 corrosion tests in a special rig at the NCB's
 Central Research Establishment.
 Suitability of Combustion Gases for
 Gas Turbines

   The  exhaust  gases  from  the  pressurised
 combustor  at  750-800°C  are passed  over a
 static cascade of blades to assess the extent of
 erosion,  corrosion,  or   deposition.  The
 cascade, a nozzle segment from an industrial
 gas turbine, has now been  in  operation for
 several  hundred  hours. The dust loading of
 the gases has generally been about 200 ppm,
 with most of the dust less than 20 /urn in size.
 Because most of the particles greater than 10
 ftm are in the form of "platelets," it is very
 difficult to obtain a realistic size distribution
 of the ash.

   The nozzle blades are  lightly coated  with
 an ash  layer which appears to reach equi-
 librium  after operating about 20  hours. The
 layer can be limited  to a negligible amount by
 conventional  on-line cleaning  methods. No
 signs of erosion have been observed.

1-4-4
Part-Load Operation

  Operation over a wide load range poses a
number of problems in a system where bed
temperature is the major factor controlling
heat generation in the bed  and where  the
range of temperature over which the bed  can
be operated is relatively narrow.

  The  main  application   envisaged   for
pressurised fluidised  combustion  is for  a
combined gas/steam turbine power plant. In
this system the combustion and heat transfer
characteristics  of  the fluid  bed have to  be
matched to the pressure, temperature, and
mass flow characteristics of the gas turbines.

  Two  main  methods  of controlling heat
absorption  in  the  bed have been considered
and  some  features have  been  investigated
using "cold" models.

  Reduction of load by shutting off sections
of the bed appears feasible for  atmospheric
pressure  combustion;  but,  until  a  large
combustor has been tested, the full extent of
the operating problems will not be known.

  It is  difficult to achieve satisfactory load
control  for a  combined cycle, using this
system;  one  problem  is  to maintain  the
air/fuel  ratio at a  satisfactory level over  the
whole range.

  The alternative method of controlling heat
absorption is to alter the amount of tubing
immersed in the  fluid bed by changing bed
level to match the load demand.

  Model work has shown rapid change of bed
level  to be feasible in principle.  Figure  5
indicates that this method can be used to
maintain a satisfactory  air/fuel  ratio over a
wide load range.

-------
FUTURE WORK
   2. Coal distribution.
   Sufficient data has been obtained from the
pilot plants to enable a major scale-up of the
atmospheric  process   to  be  faced  with
confidence.  This does not, however, mean
that all essential background work has been
completed; there is still a need for a selective
investigation of such factors as:

   1. Coal characteristics.

   2. "System  response"  for designing auto-
     matic control systems.

   3. Methods for completing combustion of
    material elutriated from the bed.

   4. Alternative  and possibly more advanced
    methods for load control that could be
    incorporated in  the scale-up plant at a
    later date.

   Much of the data obtained for operation at
atmospheric pressure  is relevant to com-
bustion under  pressure. Operation of a large
bed at atmospheric pressure is an essential
stage in scaling up the pressure process.

   Operating  experience on  the  pressure
process  suggests that  there  is  sufficient
technical justification to proceed  with  the
design  of the  next stage,  namely, a plant
incorporating a small industrial gas turbine to
demonstrate the long  term  feasibility  of
running  turbines on  the  gases  from  this
combustion system.

   Design studies on the pressure process have
also shown the need for further investigation
at the present scale  of operation of such
factors as:

   1. Ash removal.
   3. Specific features of the proposed load
     control system.

   Another programme of research recently
begun  at  Leatherhead—sulphur retention  by
lime in both pressurised and non-pressurised
fluid-bed  systems-is described in a separate
paper.
ACKNOWLEDGMENTS

  The  work  described in  this paper  was
carried out as part of the research programme
of the  Research  and  Development Depart-
ment of  the National Coal  Board.  Views
expressed are those of  the author, not neces-
sarily those of the Board.

BIBLIOGRAPHY

  The following papers have been published
by BCURA staff since the last Hueston Woods
Conference:

 1. Barker, M.H.,  A.G.  Roberts,  and  S.J.
   Wright. Fluid-Bed Boilers. Paper presented
   at Steam Plant Convention of Institute of
   Mechanical Engineers. London, 1969.

 2. Hoy, H.R.  and J.E.  Stanton.  Fluidised
   Combustion   Under  Pressure.   Paper
   presented at Joint  Meeting of  Chemical
   Institute of Canada and Division of Fuel
   Chemistry  of  the   American  Chemical
   Society. Toronto, 1970.

 S.Wright,  S.J.,  H.C.  Ketley, and  R.G.
   Hickman. The Combustion of Coal in
   Fluidised  Beds for Firing  Shell Boilers.
   Journal of the Institute of Fuel. XLII, No.
   341, June 1969.
                                                                                  1-4-5

-------
 AMBIENT  AIR INLET
                           FAN
ASH TO    CARRYOVER
WASTE     TO WASTE
                    Figure 1.  27-inch diameter fluidised-combustion rig.
                                                     STEAM
                                                     OUTLET
                                                           ACCESS
                                                                 WATER INLET
                                                               AIR
                                                         "^"•DISTRIBUTOR
                                                              PLATE
                                                                         FEED
  Figure 2.  Prototype BCURA industrial fluid-bed shell boiler (4-pass design.  Steam out-
  put 8000 Ib/hr.  Diameter 8 ft.  Height 13 ft.).
1-4-6

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                         14
15
                                                 17
                                                           18
                                               WATER INLETS & OUTLETS
                                               1ST STAGE CYCLONE
                                               RECIRCULATION CYCLONE
                                               BALANCING AIR SUPPLY
                                               STARTUP GAS BURNERS
                                               BED REMOVAL PIPE
                                               PRESSURE SHELL
                                               WATER COOLED LINER
                                               COMBUSTOR CASING
                                               2ND STAGE CYCLONE
                                               AIR INTAKE
                                               CASCADE
                                               ALKALI SAMPLING PROBES
                                               MIXING BAFFLE
                                               NOX SAMPLING POINT
                                               WATER SPRAYS
                                               DEPOSITION PROBE
                                               DUST AND GAS SAMPLING PROBE
                                               TO PRESSURE LETDOWN VALVE
                                               RECIRCULATION CYCLONE DIP-LEG
                                               COAL INLET
                                               AIR DISTRIBUTOR
                                               ASH OUTLET
                                           DETAIL SHOWING ARRANGEMENT
                                                    OF TUBES
                                                 IN FLUIDISED BED
Figure 3.  Section through 48" x 24" pressurised  fluid-bed combustion rig.
                                                                   1-4-7

-------
      200
u
LU
o
O
QC
LU
LL
CO _
Z iT
U
LU
>
Z
o
o
100

 80


 60



 40
                        TUBE TEMPERATURE 212°F (100°C)
       20
                   I
                    I	I
       0.004
      100 —
   O
   cc
           0.006         0.01             0.02             0.04
                             MEAN PARTICLE SIZE (in.)
         Figure 4.  Fluid-bed heat transfer coefficients.
                                                                       0.06
                                                                                    0.1
                                                                                       1600
                                                     PERCENTAGE OF
                                                     TUBE SURFACE
                                                     IN BED
                            % EXCESS AIR — OVERALL
                            % EXCESS AIR — BED
                                      50       60
                                        LOAD, % mcr
            Figure 5.  Calculated part-load characteristics: combustor with
            variable bed level.
                                                                              100
1-4-8

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  5.  COMPARATIVE EMISSIONS OF POLLUTANTS
                                     DURING COMBUSTION OF
    NATURAL GAS  AND  COAL IN FLUIDIZED BEDS

              R.L. JARRY, L.J. AN AST ASIA, E.L. CARLS,
                      A.A. JONKE, AND G.J.  VOGEL

                      Argonne  National Laboratory
ABSTRACT

   The  feasibility of employing a fluidized-
bed combustor for the combustion of natural
gas is  being  investigated.  Because  of the
relatively low temperature (1600° F) usable in
this method of combustion, NO generation is
held  to  the  level dictated by the nitrogen
fixation  equilibrium, about 60 ppm. Com-
bustion efficiencies in the range 94-99 percent
were obtained; however, a combustion ef-
ficiency  of greater than  99 percent  can
undoubtedly  be achieved by  more efficient
distribution of natural gas in the fluidized
bed.

   The reaction of artificially generated SC*2
with limestone during natural gas combustion
at 1600°F resulted in SC>2 removals of 95 to
99 percent at  a low Ca/S mole ratio of 1.5.
These excellent results were probably due to
more efficient distribution of the injected
SC>2 in  the combustor and to the ash-free
environment.

   In other work, it was noted that when the
bed of partially reacted limestone was heated,
desorption  of  SC"2  started  at slightly  above
1700°F.  At  1850°F  with  1 vol percent
oxygen in the  flue gas, SC>2 concentration in
the flue  gas was 5000 ppm. The reason for
this high degree of SC>2  desorption  at a
relatively low temperature (a temperature of
about  2200°F was thought  necessary to
obtain this effect) may be the presence within
the bed  of partial reducing conditions  at the
low excess-air level employed.
   Higher than equilibrium concentrations of
NO have  been observed during fluidized-bed
coal  combustion. The source of the nitrogen
was shown to be the nitrogenous content of
the coal itself-argon was substituted for the
nitrogen content of the combustion air and
no significant effect on NO concentration was
observed.

   A concomitant  reduction  in  NO  con-
centration with a reduction in SO2 concentra-
tion upon the addition of limestone has been
noted. This is probably related largely to the
presence of sulfated limestone.

   The use of metal oxides (A12O3, CO3O4,
and ZrO2> as catalysts for the decomposition
of NO was not effective. In fact, when €0304
was tested, the NO concentration in the flue
gas increased.

   The   results   of   three  independent
analytical procedures showed that nitric oxide
is the only nitrogen oxide species present in
the flue gas in a significant quantity during
the fluidized-bed combustion of coal.
INTRODUCTION


   Natural gas fuel is increasingly utilized in
power generating plants as a means of elimina-
ting the SO2 emission problem; however, NO
concentration in the flue gas is high due to
the reaction of oxygen and nitrogen at the
elevated temperature  of the gas flame. The
                                     1-5-1

-------
 NO concentrations observed are in the range
 160-1140  ppm,  indicating that combustion
 temperatures are as high  as  2000-3000°F.
 Because of increasing  concern  about the
 emission of nitric oxide from  electric power
 generating stations,  the feasibility of natural
 gas combustion in a fluid ized-bed combustor
 has  become of interest.   Combustion  of
 natural gas in a fluid-bed combustor could be
 efficiently   performed  at  relatively  low
 temperatures (e.g., 1600°F), and conceivably
 only the  equilibrium NO  concentration of
 <100 ppm would be in the gas.

    To test  the  feasibility of fluid-bed com-
 bustion of natural gas as a means of reducing
 NO  emissions,  experiments   have  been
 performed  at  Argonne in  a 6-in. diameter
 bench-scale  combustor. The results of these
 experiments are presented  here, along with
 ancillary    investigations    concerning  the
 behavior of a partially sulfated limestone bed
 during  natural  gas  combustion. The data
 obtained   with  partially  sulfated limestone
 beds suggests answers to such questions as:

     1.  How does the location of SO2 genera-
        tion in  the fluid  bed  affect SO2
        removal?

    2.  Why is  SO2 released from partially
        sulfated   limestone  at  temperatures
        considerably  below  those  expected
        from equilibrium considerations?

    The final section of this paper discusses
general  experience pertaining to NO forma-
tion  and  removal  from  the  flue  gas  in
fluidized-bed  experiments  with  coal and
natural  gas flues.
APPARATUS AND MATERIALS

   The 6-in. diameter bench-scale combustor
is described in detail in a recent annual report
on this program.1 Briefly, as shown in Figure
1,  the bench-scale equipment consists of a
combustor (about 6 ft long), an air preheater

1-5-2
(also 6  ft long, but only 6 in. in diameter),
vibratory screw feeders for delivering the coal
or  limestone additive  into an air transport
stream  for injection into the bed, cyclone
separators  and a glass  fiber final  filter to
remove  entrained solids from the flue gas, and
instruments for gas metering and gas analysis.

    In the startup procedure for these bench-
scale experiments, the fluidizing air was pre-
heated  to about  1000°F;  the  heated  air
increased the bed temperature to 700-800° F.
A  bed  temperature  of about  1200°F  is
necessary for the ignition and sustained burn-
ing of natural gas. For the natural gas com-
bustion  experiments, the bed temperature was
raised  further  by  burning  either coal or
propane so that the  methane could  ignite.
Subsequently, natural  gas was injected into
the fluidized  bed at the coal injection point
(about   4  in.  above  the  fluidizing-air
distributor   plate).   Upon   ignition   and
sustained  burning  of the  natural  gas, the
fluid-bed  temperature  was raised to  the
operating range of 1600-1850°F.

    The natural  gas, from a commercial gas
line, was supplied  by Northern Illinois  Gas
Co., who  also  supplied a  typical analysis
(Table  1)  accomplished by  gas  chromato-
graphy.  The nitrogen  present as  elemental
nitrogen  was  added  to adjust  the  heat
content value of the natural gas.

    The fluid bed  for the natural gas com-
bustion  experiments was the final bed from a
previous  coal  combustion  experiment  in
which a fine-particle-size limestone had been
reacted  with SO2-  The  calcium and  sulfur
contents of this bed were 1.54 and 0.95 wt
percent, respectively;  the remainder of the
bed was essentially refractory alumina.

    In several experiments in the natural gas
series, SO2 was added to the gas phase. Purity
of the SO2 was >99 percent. The limestone
No. 1359 used had an average particle size of
25 nm and contained 97.8 wt percent CaCO?
and 1.3 wt percent

-------
NATURAL-GAS COMBUSTION
EXPERIMENTS

NO Level and Combustion Efficiency

    The NO concentrations,  measured in the
flue gas for natural  gas  combustion experi-
ments  at  a temperature of  1600°F, varied
from  60 to MOO ppm  at  an oxygen  con-
centration in the flue gas of about 3.5 vol
percent.' (The oxygen concentration in the
flue gas reflects the percent  excess air in the
feed and the combustion efficiency.) With 3.5
vol percent oxygen in the flue gas at 1600°F,
the nitrogen fixation equilibrium data would
give a NO concentration of 54 ppm. This is in
reasonable  agreement  with  the  observed
levels.  This data indicates that significantly

     Table 1. COMPOSITION ANALYSIS OF
	NATURAL GAS	
                               Analysis3
                                by Gas
                            Chromatography,
Constituent
CH4
C2H6
C3H8
C4H10
C5H12
C02
He
02
Miscellaneous
N2
Heat content, Btu/cu ft
Specific gravity
vol%
88.15
5.38
1.65
0.43
0.10
0.36
0.17
0.05
0.07
3.64
1035
0.6
aNorthern Illinois Gas Co.
                                              lower NO  concentrations should be achieved
                                              in a fluidized-bed natural gas combustor than
                                              in a conventional combustor.

                                                 Combustion efficiency of the natural gas
                                              at 1600°F was about 99 percent with oxygen
                                              concentrations in  the  flue  gas  of 3-4 vol
                                              percent. More efficient distribution of natural
                                              gas in the fluid bed could no doubt increase
                                              combustion efficiency above  99 percent. The
                                              efficiencies were  calculated  from  CO and
                                              hydrocarbon   analyses   (obtained  by  gas
                                              chromatography) of grab samples of flue gas.
                                              The  concentrations of  CO and hydrocarbon
                                              (expressed as  CHX)  in  the  flue gas  at 99
                                              percent combustion efficiency averaged about
                                              900 and 250 ppm, respectively. As would be
                                              expected,  the  combustion  efficiency was
                                              sensitive to oxygen concentration in the flue
                                              gas;  at  1-2  vol percent oxygen, the  com-
                                              bustion efficiency averaged 94% and the CO
                                              and CHX  concentrations averaged 6000 and
                                              2600  ppm, respectively. These results  are in
                                              general agreement  with those obtained in  a
                                              Russian application2 of fluidized-bed  com-
                                              bustion of gas to heating furnaces.
                                              Desorption and Reaction of S02 with
                                              Limestone

                                                 During the course of a natural gas com-
                                              bustion experiment (Figure 2), SO2 began to
                                              desorb  from  the  partially  sulfated  bed
                                              material  when  the  bed  temperature  was
                                              M720°F. Desorption increased rapidly and
                                              reached a peak concentration in the  flue gas
                                              of 2450 ppm at a temperature of 1820°F, at
                                              which point the oxygen concentration in the
                                              flue gas was 1.3 vol percent. When the oxygen
                                              level was increased to 4.4 vol percent,  SO2
                                              concentration in the flue gas decreased to 200
                                              ppm. Of potential interest for elucidating the
                                              mechanism of this desorption reaction are the
                                              high  concentrations of CH4 and CO at the
                                              stage where  SO2  desorption  was near its
                                              highest  value (see Figure 2). The presence of
                                              the reducing gas CO might contribute signifi-
                                              cantly to desorption.

                                                                                  1-5-3

-------
     Both combustion temperature and oxygen
  concentration are variables that could affect
  the  desorption  of SCK  In  another experi-
  ment, with a constant 1.8 vol percent oxygen
  concentration in the flue gas,  the SC»2 con-
  centration was 2250 ppm at  1840°F and 800
  ppm   at  1790°F.  When the combustion
  temperature was maintained at 1850°F, Table
  2 shows that the S(>2 concentration varied
  inversely  with oxygen  concentration in the
  flue gas over a rather narrow oxygen con-
  centration  range.   Figure   3  shows   the
  dependence of SC»2 concentration in the flue
  gas on oxygen  concentration.  It shows that
  doubling the oxygen concentration (from 1.0
  to 2.0 :vol percent) reduces the SC»2 desorp-
  tion  by about half,  and that a  fourfold
  increase in oxygen concentration (from 1 to 4
  vol percent) decreases the desorption of SC>2
  tenfold. Integration of the SC^-time desorp-
  tion  curve  of Figure 2  indicates  that  ap-
  proximately 30 percent of the sulfur original-
  ly contained in the  fluid-bed  material was
  desorbed  during the  short  period of the
  desorption phase of the experiment.

    The relative ease of desulfurization of the
 bed material is potentially  of high signifi-
 cance. It suggests that the  regeneration of
 sulfated limestone and the recovery of sulfur
 values   might  be  readily  achieved in  an
 auxiliary reactor operating as  a combustor at
 slightly  higher  temperatures  and at a  low
 oxygen  concentration.  The  use of highly
 reducing conditions of regeneration apparent-
 ly is not necessary.

    A series of natural-gas combustion experi-
 ments was performed to investigate the effect
 of the absence of fly ash on the extent of SC«2
 removal. Sulfur dioxide was injected into the
 fluidizing air stream near the  preheater at a
 flow rate sufficient to result in a flue gas
 concentration of 3000  to 4000 ppm in the
 absence  of additive.  Limestone  was then
 injected into the fluid bed, and the extent of
 SC»2 removal was observed.

    Figure 4 shows the results  for an additive
 injection  period of  a  typical  natural  gas

1-5-4
 Table 2. VARIATION IN S02 CONCENTRATION
 IN FLUE GAS AS A FUNCTION OF OXYGEN CON-
           CENTRATION AT 1850 °F
02 in





Flue Gas, vol %
1.8
1.5
1.0
0.5
1.8
SO2 in Flue Gas, ppm
2700
2820
4280
4900
2050
 combustion experiment (NG-3)  at  1600°F.
 The starting SC>2 concentration of 3520 ppm
 was reduced  to about 160 ppm,  an SO?
 removal greater  than 95 percent. Because of
 uncertainty in the zero point of the infrared
 detector for SC>2,  the actual SC>2 removal
 may have been as high as 99 percent. The CaO
 utilization  for limestone No.  1359  (average
 particle size,  -v25 p.m) was 66 percent. Both
 the extent of SC>2 reduction  and  the lime-
 stone  utilization   are  higher  than were
 obtained during  combustion of coal, especial-
 ly since a rather low Ca/S mole ratio of 1.5
 was used. A similar coal combustion experi-
 ment at the same Ca/S mole ratio would be
 expected to  result  in  removal of only  60
 percent of the SO2-

    The SC>2 removal by reaction with lime-
 stone appeared to depend on  where SO2 is
 injected into  the bed. In two experiments at
 1600°F with limestone No. 1359 and Ca/S
 mole ratios of 1.8 and 2.1, SO2 was injected
 directly into  the combustor; SO2  removals
 from  the flue gas were 95 and 92 percent
 respectively. The major difference between
 the sets of experiments discussed in this and
 the preceding paragraph was the location in
 the apparatus where SO2 was injected For
 the  experiment  (discussed  above)  which
resulted in 95 to 99 percent removal of SO-j
tile S02 was injected into the fluidizing air at
the preheater, upstream of  the combustor
For the other experiment, which resulted in a

-------
lower SC>2 removal, the SC>2 was injected into
the combustor near the point where natural
gas was injected into the reactor. These results
suggest that the reason for a greater degree of
SO2 removal was more efficient mixing and
homogenizing of the  gas mixture  related  to
the mechanics  of injecting the SC>2 into the
fluidizing air.

   The possibility that added fine particulate
material  affects SC>2 removal during natural
gas combustion  was investigated.  The  high
levels of SC>2 removal which were obtained in
the natural gas experiments were for cases in
which no fly ash was present. Tested for their
        effect on SC>2 removal were (1) fine alumina
        (<44 /im) and (2) fly  ash produced in a coal
        combustion experiment and  collected in a
        cyclone separator. Table 3 lists experimental
        conditions  and  results  obtained.  In  this
        experiment, SC>2 was added at a 1600° F com-
        bustion temperature until flue gas concentra-
        tion was  3840 ppm. At this point, limestone
        No. 1359 was added  at a rate to achieve a
        Ca/S  mole ratio of  1.8, resulting  in an SC>2
        removal of 95 percent. Fine alumina was then
        injected into the reactor at a rate of 1.4 Ib/hr,
        approximately three times the usual rate of
        fly-ash production during a coal combustion
        experiment. The presence of the fine alumina
             Table 3. EFFECT OF ADDITION OF FINE PARTICLES ON SO2 ADSORP-
             TION BY LIMESTONE IN A NATURAL GAS COMBUSTION EXPERIMENT3
Added
Solids
-
—
Fine AI2O3
Fly ash
Fly ash
Fly ash
Solids,
Ib/hr
-
-
1.4
0.9
0.9
1.7
Limestone
No. 1359,
Ib/hr
0
0.8
0.8
0.8
0.8
0.8
Ca/S
Mole
Ratio
0
1.8
1.8
1.6C
1.6°
1.4C
S02in
Flue Gas,
ppm
3840
180
180
-V400
'WJOO
^620
% SO2 % CaO
Removal Utiliz.
- -
95 53
95 53
'MJO M56
'vgO 'v/SS
^84 'veO
             Experiment conditions:
               Starting bed

               Bed height

               Fuel


               Additive

               Superficial gas velocity

               Temperature

             bSTPat70°F, 1 atm.
             °Fly ash contained 2.4 wt % S.
29.1 Ib sintered alumina (30 mesh)

15 in. static and 24 in. fluidized

Natural gas. Northern Illinois Gas Co., 0.82
scfm,b plus SO2 added at 1.7 scfh.b

Limestone No. 1359, 25 (im, 97.8 wt % CaC03

2.7 ft/sec
1600°F
                                                                                    1-5-5

-------
particles did not affect the reaction of SC>2
with limestone. Fly ash was then injected into
the  combustor  in three separate injection
periods: during  the initial two periods, at a
rate twice that  of fly-ash production during
coal combustion; and during the third period,
at a rate 3.5 times the normal fly-ash produc-
tion  rate  during coal  combustion.  The
removal of SC>2 was lowered to 90 percent
during  the  periods  at  the  lower  fly-ash
injection  rate,  and  to 84 percent  during
fly-ash  injection at the higher rate. In this
experiment,  the effect of the presence of fly
ash on  SO2 removal, although significant, did
not  completely account for  the observed
difference between the natural  gas  experi-
ments and the coal combustion experiments
(in which the extent of SO2 removal was only
 about 65 percent). However, the effect on SC>2
removal of fly ash produced on  site may be
quite different (e.g., the actual effect may be
greater) from that indicated by these experi-
mental results.
 Conclusions—Experiments with Natural
 Gas Fuel

     1. Fluidized-bed  combustion of  natural
 gas  at 1600°F results in NO concentrations in
 the  flue gas that are essentially those expected
 from the nitrogen fixation equilibrium.

    2. The extreme simplicity of applying the
 fluidized-bed  concept  to  combustion of
 natural gas in  a  fluid bed is  noteworthy.
 Because  ash is absent and because the fuel is
 easily  distributed  throughout the  fluidized
 bed, scale-up would be simpler than for coal.
 Only small-scale experimentation would be
 needed to optimize the operating conditions
 for high combustion  efficiency.  Hence, if a
 market exists  for this  concept, it  would
 represent a means for the early demonstration
 of fluidized-bed combustion on  a prototype
 scale.

   3.  The reaction of injected SO2  with
 limestone  during  natural   gas  combustion

1-5-6
resulted in a high degree of SO2 removal from
the flue gas. Most probably, the 95 percent or
higher  SO2 removal is largely due to better
SO2 distribution in the combustor and to the
absence of fly  ash in this method of com-
bustion.

    4. By operation  of  a natural gas com-
bustor  at 1850°F  with  a low oxygen con-
centration (1 to 2 vol percent) in the flue gas,
SO 2 might be desorbed from sulfated lime-
stone. This mode of combustion is effective in
desorbing SO2  at a  temperature  several
hundred  degrees below that at which desorp-
tion is expected, and is thought to be a result
of partial reducing conditions in the bed.
NITROGEN OXIDES IN THE FLUIDIZED-
BED COMBUSTION OF COAL

Source of Nitrogen

   Nitrogen  oxides,  principally  NO,  are
formed during the combustion of fossil fuels
and  are important  contributors to  air pol-
lution.  Although  the  quantities  of NO
observed  in   the  flue   gas  from  high-
temperature conventional combustors can be
accounted  for  by the equilibrium  of the
nitrogen fixation reaction, this is not the case
for fluidized-bed  coal combustors, in which
the concentrations of NO have been observed
to be far in excess of those expected on the
basis  of  the equilibrium.  At 1600°F,   a
common  temperature for  fluid-bed  com-
bustion, the equilibrium concentration of NO
should vary from 50-200 ppm, depending on
the oxygen concentration. The oxygen con-
centration, in turn, would depend on the level
of excess  air employed. However, in actual
fluid-bed  coal  combustion,  NO  levels of
400-800  ppm in  the  flue  gas have been
measured.


   One possible  explanation of these high
NO concentrations is that the temperature of
the  combusting  particles  might  be  higher
(e.g.,   1800-2100°F)  than  the  measured

-------
 temperature,   1600°F.   Another   possible
 explanation for higher than equilibrium NO
 concentrations is that the nitrogenous content
 of the coal (1-1.5 wt  percent in U.S.  coals)
 reacts  with oxygen in the fluidizing  air to
 form the excess NO. That coal is indeed the
 source  of  nitrogen was  demonstrated by
 substituting an  inert component, argon, for
 the nitrogen component of the fluidizing gas.
 Several fluid-bed experiments were performed
 substituting argon, and it was found that the
 change from nitrogen to argon did not  affect
 the  NO  level significantly, showing clearly
 that most of the  NO originated  from the
 nitrogenous content  of  the  coal. In  one
 experiment, the initial NO concentration of
 580 ppm  did not  vary  during the  argon
 substitution by  more than ±20  ppm.  In
 another,  the initial  NO concentration of 400
 ppm actually  increased  to 450  ppm  when
 argon  was  substituted for nitrogen in the
 fluidizing gas.

    This  conclusion was substantiated by the
 results of the previously described natural gas
 experiments. In  these  experiments,  in which
 the only source of nitrogen was the  fluidizing
 air, the observed NO concentration in the flue
 gas, <100 ppm, was that expected  from the
 nitrogen fixation equilibrium.
Removal of NO by Limestone Additive

   Another interesting finding is that the NO
concentration   was  reduced concomitantly
with the SO2 removal when limestone was
added. The original NO concentration in the
flue  gas was  lowered by  as  much as 40
percent.  That  the  removal  of  NO  was
influenced  by  the  addition of limestone and
therefore possibly related to the reduction in
SO2 concentration was indicated by a rise in
the NO  concentration when the  limestone
addition was terminated.

   The behavior of the  NO concentration
during the addition  of limestone to  a fluid-
bed combustor was  observed in a series  of
coal  combustion  experiments.  The  most
significant reduction  in NO concentration (to
200-300 ppm from  an initial  level  of about
600 ppm) was achieved when the initial fluid
bed was composed entirely of fresh limestone.
An  interesting  observation  was  that  NO
removal did not occur immediately  after the
addition of limestone but  only after  SO2
removal  had occurred. Possibly, the agent
responsible  for  NO removal  is the sulfate
formed by  the  reaction between  SO2 and
limestone.

    Information   has  also  been   obtained
showing that the extent of NO reduction is
inversely proportional to the Ca/S mole ratio.
Figure  5,  in  which the  reduction of NO
emission is plotted  against the Ca/S  mole
ratio,  shows that as  the  Ca/S  mole  ratio
increased,  the  extent  of  NO removal  de-
creased. A mechanism for the effect is not yet
known.

    An analysis of a series of fluid-bed experi-
ments showed that the degree of reduction in
NO concentration was also related to the
velocity  of the  fluidizing gas.  These results
(Figure 6)  show that NO removal in the flue
gas is inversely proportional to superficial gas
velocity, suggesting that gas residence time in
the bed is a factor in NO decomposition.

    A proposed mechanism for the removal of
NO concomitantly with the removal of SO2
upon the addition of limestone in a  fluid-bed
combustor is the equilibrium reaction NO2 +
SO2 <*SO3 + NO, which has been studied by
British workers.3 If this reaction occurs, a
decrease in SO2 concentration  could result in
a corresponding decrease in NO concentration
and an increase in NO2 concentration. This
hypothesis  was  tested  by determining the
NO2 content of the flue gas during a fluid-bed
combustion   experiment.   Three   different
methods were used to assay the NO2 content
of the flue gas:

    1.  The phenoldisulphonic acid method.
    2.  Mass spectrometric analysis.
    3.  The Mast NO2 analyzer.
                                     1-5-7

-------
    Analyses were  performed on samples of
 flue  gas  taken  both before and after  the
 addition of limestone in a coal combustion
 experiment.  The  results  from  all  three
 methods  indicated that essentially no NC>2
 was present in the flue gas samples taken. For
 example,  mass  spectrometric  analysis
 indicated that the NO/NO 2 ratio was at least
 50/1 and probably much higher. On the other
 hand,  the   Mast  in-line  NC>2  analyzer
 consistently indicated no NC>2 or at the most
 concentrations only in the ppm range. Of the
 three  methods, the  phenoldisulfonic  acid
 method  seemed the  least precise;  by dif-
 ference (NOX - NO = NO2>, both negative and
 positive  values  were obtained for the NO 2
 concentration.
 Investigation of Catalytic Decomposition

   Catalytic decomposition was considered as
 a means of  controlling  NO emission. The
 literature in the  field of  catalytic control of
 nitrogen oxides4'5 indicates that low reaction
 rates,  instability  of the catalyst at fluid-bed
 temperatures, and catalyst poisoning by sulfur
 oxides are  among  the  factors  mitigating
 against on-site use  of particulate catalysts.
 Nevertheless,  the attractiveness  of this  ap-
 proach for lowering NO emission, the fact
 that these catalysts  had not been previously
 evaluated  in  a  fluidized  bed,  and  the
 simplicity of the  testing procedures led to the
 performance of scoping tests. The literature
 indicated that oxides  of aluminum,  cobalt,
 and zirconium  appeared to  be the most
 effective of the decomposition catalysts.

  Experiments   were  conducted  in   the
bench-scale  fluid-bed combustor by injecting
solid  metal  oxides  into  the fluidized-bed
combustor at the same location at which  the
coal and limestone additive were injected. In
the first experiment, the  solid catalysts used
were (in order of their use) aluminum oxide
(A^Os), zirconium oxide (ZrO2>, and cobalt
oxide  (00304,  cobaltous-cobaltic  oxide).
Except for zirconium oxide, which contained

1-5-8
1-2  percent  hafnium,  these  materials were
reagent  grade with purities greater than 99.9
percent.

   Injection  of  AbO3 or  ZrO2 into  the
fluidized bed at rates  of several pounds  per
hour did not change the NO concentration in
the flue gas. Injection of each material was
continued for about 1 hour.

   The  injection  of  00304  (cobaltous-
cobaltic  oxide)  at a  rate of about  4 Ib/hr
caused the NO concentration in the  flue gas
to increase  rather than decrease (the  latter
would be expected if  NO were catalytically
decomposed).   The    NO   concentration
increased from  about  700 ppm before injec-
tion  to  about  1400  ppm during injection.
When 00304 injection was stopped, the  NO
concentration in  the  flue gas dropped back
nearly to its original level. Upon reinstitution
of catalyst injection, NO concentration in the
flue gas immediately increased.

   In a  second  experiment with 00304,  NO
and  SO2 concentrations in the  flue gas prior
to catalyst injection were 600 ppm and 1000
ppm, respectively. When 00304 was first
injected at a rate of about 4.5 Ib/hr, the  NO
concentration increased to about 1000 ppm,
an increase of about 400 ppm. 00304 might
act as a catalyst for the oxidation of nitrogen
compounds  in  coal to NO, similarly to  the
catalytic oxidation of ammonia. In all experi-
ments, the addition of metal oxide catalysts
had  no  discernible effect on SO2 concentra-
tion  in the flue gas.
Conclusions—Nitric Oxide Behavior in
Coal Combustion Experiments

   1. The concentrations of NO observed in
fluidized-bed combustion of coal were higher
than  those predicted  on  the  basis of  the
equilibrium. A major source of nitrogen was
the nitrogenous content of the coal itself.

  2. Upon the addition of limestone, nitric
oxide concentration in  the flue gas is lowered

-------
concomitantly with SC>2 concentration. This
reduction in NO level appears to be related to
sulfation of limestone.

  3. NO is the predominant species if not the
only nitrogen oxide species in the flue gas.

  4. The  use of a  metal oxide  (A12O3,
0)304, or  ZrO2>  as a  catalyst  for  the
decomposition of NO was  not effective. In
fact, the NO level in the flue gas was increased
rather  than decreased in  the presence  of
€0304.

  5. Evidence was  obtained  that the equi-
librium reaction NO2 + SO2 = NO + 803 was
not involved in  the reduction of NO con-
centration.

  6. For a Tymochtee dolomite, the extent
of reduction of  NO  concentration during
limestone addition was shown to be inversely
related to  the Ca/S mole ratio; for limestone
No.  1359, it was inversely related  to  the
superficial fluidizing gas velocity.
BIBLIOGRAPHY

 l.Jonke, A.A., E.L. Carls, R.L. Jarry, M.
   Haas, W.A. Murphy, and C.B. Schoffstoll.
  Reduction of Atmospheric  Pollution by
  the Application  of Fluidized-Bed Com-
  bustion,   Annual    Report.   ANL/ES-
  CEN-1001 .July 1968-June 1969.


2. Rubtsov, G.K. and N.I. Synomyatmkov.
  Investigation   of  Gas  Combustion  in
  Fluidized  Beds  as Applied to Heating
  Furnaces. Russian  Metallurgy and Mining
  No. 2: 50-57, 1964.


3. Annual Report,  The British  Utilisation
  Research Association: p. 38, 1966.

4. Shelef,  M.   and  J.T.  Kummer.  The
  Behavior  of  Nitric   Oxide in Hetero-
  geneous   Catalytic   Reactions.  Paper
  presented at the  Technical Meeting of the
  Eastern Section  of the Combustion Insti-
  tute,  September 29   October  1, 1969,
  W.Va.  University,  Morgantown,  W.Va.,
  1969.

5. Bartok,  W.  et  al.  Systems  Study  of
  Nitrogen  Oxide  Control  Methods  for
  Stationary    Sources.   Interim   Status
  Report, GR-l-NOS-69, Esso Research and
  Engineering  Co.,  Government  Research
  Laboratory, May 1, 1969.
                                                                                 1-5-9

-------
02  N2   Ai  An
  PREHEATER
                     ADDITIVE
                      FEEDER
                                   Y
                                   H
                     TRANSPORT     M
                        AIR-CX}—'
RECYCLE-
MATERIALS
FEEDER
                                                                                                           TO
                                                                                                           FILTER
                                                                                                           AND
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                                                                                                           EXHAUST
                                                                                                     SECONDARY
                                                                                                     CYCLONE
                                                                                                     TRANSPORT
                                                                                                        AIR
                                                      COMBUSTOR
                      Figure 1.  Equipment for investigating 302  emission.

-------
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   3200
   2800
 52
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1200
    800
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h-

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                                r
                                 i
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                    12345
                                        TIME, hr

  Figure 2. Flue gas composition (dry basis) and combustion temperature, experiment NG-2.
                                                                                1-5-11

-------
   6000
   5000
   4000
 Q.
 CO
 <
 O
   3000
   2000
    1000
       01234
          OXYGEN IN FLUE GAS, vol %
    Figure 3. Desorption of S02 during
    natural gas combustion as a function
    of oxygen concentration at 1820-
    1850°F.
                                         Q.

                                         CO
                                         <
                                         tr
                                         O
                                         O
                                         O
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                                             3200
                                             2800
                                             2400
                                             2000
                                             1600
1200
 800
                                              400
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       Ca/S MOLE RATIO IN FEED
        H	-I
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                                                                         SO2
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          25 urn LIMESTONE NO. 1359
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                                             1600
1500
              4    6
             TIME, hr.
10
Figure 4.  Operating conditions and
flue gas composition, experiment NG-3,
part 2.
1-5-12

-------
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O
UJ
O
O

O
Q
       40
       30
20
       10
                                                              I
         0                           1                          2
                                      Ca/S MOLE RATIO
       Figure 5.  Nitric oxide reduction with -325 mesh (<44 urn) Tymochtee dolomite.
GAS
VELOCITY,
ft/sec
02.7
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A8.6
A8.6
02.7
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25
25
25
25
600
1400
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NO
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NO
NO
TEMPERATURE: 1600 °F
COAL FEED: 4.5wt%S
ADDITIVE: LIMESTONE NO. 1359
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          Figure 6.  Effect of gas velocity on nitric oxide reduction in flue gas.
                                                                                    1-5-13

-------
                            6.  IGNIFLUID CONTRIBUTION
                                                                       TO
                                  AIR POLLUTION  CONTROL

                             J. J. SVOBODA

                       Babcock-Atlantique, Paris
INTRODUCTION
  The Ignifluid combustion process, charac-
terized by the  fact that the total amount of
ash  retained comes out as clinker, is opera-
tional on an industrial scale.

  Boilers associated with this kind of furnace
are strictly conventional; those at Casablanca
Power Station in Morocco are typical (Figure
1).

  One of the main advantages of the Ignifluid
process  is  that it makes it possible to burn
coals which have a wide  variety of charac-
teristics; e.g., volatile matter, ash, and grain
size.


  Table  1  gives  the characteristics of the
Ignifluid units we shall refer to in this paper;
Table  2 gives the grades of  coal (from
anthracite to bituminous) fired in them.


  The operating  results obtained in all of
these different industrial  installations  can,
therefore, very often be considered as having
a general character.
                 Table 1. CHARACTERISTICS OF IGNIFLUID PLANTS

Location
La Taupe
Casablanca
Solvay
La Pochette
Steam Output
M.C.R., Ib/hr
77,000
254,000
110,000
110,000
Pressure,
Ib/sq in.
840
1,300
290
670
Superheat
Temperature, °F
915
1,000
570
825
Operating Mrs
October 1970
5,000
15,000
8,500
53,000
                             Table 2. COAL USED


Location
La Taupe

Casablanca

Solvay

La Pochette



Type of Coal
Middlings
from Auvergne
Anthracite
from Djerada
Bituminous
from Lorraine
Anthracite
from La Mure
Characteristics of Coal
G.C.V., Dry
Basis, Btu/lb
11,000

11,200

13,300

11,200


V.M., %
10.0

6.5

35.0

5.5


Ash, %
22

23

7

21


Sulphur, %
1.0

2.0

0.7

0.5

                                   1-6-1

-------
   It  is not the purpose of this paper to deal
with  the  Ignifluid process  from  the  com-
bustion  point of  view (e.g., its  efficiency,
flexibility  in  operation,  and  availability).
These subjects have already been discussed:
by Mr. (A.A.) Godel at the first International
meeting 2 years ago, as well as in more recent
publications.

   It  is generally agreed that the fluidized-bed
combustion  process  may  be  an  efficient
means of overcoming atmospheric pollution.
It is interesting to examine the position of the
Ignifluid system with regard to this important
problem.
GOAL-TO LIMIT FLY ASH
   Historically speaking, in the course of the
 15 years devoted to the development of the
 Ignifluid process, our efforts from this point
 of view are aimed solely at limiting the solid
 fly ash at the stack. The reason is that French
 coals, like  most European coals,  have the
 great  advantage  of  containing very  little
 sulphur. On the other hand practically all the
 Ignifluid furnaces  have  been installed  in
 France or abroad for firing coal with a high
 ash content.
   For units with a steam output of 150,000
Ib/hr or less, mechanical dust collecting has
proved efficient enough. This  is mentioned
only as a reminder.

   The interesting results to be considered are,
therefore,   those   obtained with the  elec-
trostatic precipitator at  the 60-MW Power
Station at Casablanca in Morocco, fitted with
Ignifluid furnaces burning anthracite  fines
with more than 20 percent ash.

   Various   measurements which have  been
carried out  make it certain that the efficiency
guarantee (240  mg/Nm3 or 0.105 grains/std
cu ft) is easily reached (160 gr/Nm3 or 0.070
grains/std cu ft).

1-6-2
   At the present  time, very exhaustive tests
are being carried out in Casablanca; we hope
they will provide us with more useful data.

   These results are mainly due to the intrinsic
characteristics  of the  Ignifluid  grate  and
particularly:

   1. The  large-sizing  particles  of  the fly
ash—compared  with   the   fly   ash  from
pulverized   coal    furnaces,   considerably
facilitates the work of the  precipitators by
natural settling. The dust  content at  the
outlet  of  the  precipitator  is practically in-
dependent of the dust  content  in the gas at
the inlet (for high-ash coals) because the fly
ash increase,  according to  the  boiler load,
occurs through  the addition  of coarse grains.
As an  example,  the  above-mentioned  effi-
ciency is obtained with a very high dust con-
tent before the precipitator, of 80 g/Nm3 or
35 grains/std cu ft.
  2. The high carbon content in the fly ash is
also  a favourable  factor. It  may seem para-
doxical that a high unburnt carbon content in
the fly ash (70 percent  at Casablanca where
the coal fired is anthracite) should be con-
sidered as an asset  for the combustion furnace
from which it comes. The reason is that all fly
ash is retired, and that all the ash is discharged
in the form of slag with a low carbon content,
and therefore with high efficiency.

  Other factors,  irrevelant  to  the Ignifluid
process, have a good effect:

   1. The S content in the coal at Casablanca
is higher than in French and European coals.!
It amounts to approximately 2 percent: 0.82
percent as  pyrites; 1.06 percent as organic
matter; and 0.12  percent as sulphate. The
general  belief that  precipitator  efficiency
increases with the sulphur content is at least
in favour of this "impurity." From this point
of  view,  American bituminous coals,  with
their high sulphur content, should not offer
any difficulty for dust removal.

-------
   2.  The gas  temperature  in  the  precipita-
tors,  which  is approximately 150°C (300°F),
is apparently  fairly  close to  the  optimum
value.

   3.  Finally,  other factors (e.g.,  steam and
alkali content) are doubtless to be taken into
consideration.
SO2/NOX INTEREST IS RECENT

  As mentioned above, only recently has our
attention  been  drawn to the  problems of
atmospheric  pollution  due  to release  of
sulphur dioxide (SC>2); and, on another scale,
to nitrogen oxides (NOX).

  Moreover, the  important project we have
been  studying for  an American  company,
which Mr. (R.H.) Demmy will tell us about
later (in Session No. IV), is based on the use
of refuse  bank coal with a high ash content
(40 percent) and only 1 percent of sulphur.

  However, the anxiety expressed throughout
the world  on the consequences of atmos-
pheric pollution  by  S(>2, and the interest
shown in  fluidized-bed combustion methods,
have  led  us not only to examine existing
possibilities of the Ignifluid system as it  is
operated in industrial installations, but also to
consider certain  possible alterations or ad-
ditions  to make  them suitable  for more ef-
ficient prevention of  atmospheric  pollution.

  The measurements, the results of which are
given  below,  were  therefore  taken during
industrial operation,  and not  by means of
laboratory tests.

  The different tests (Table  3) were carried
out at La Taupe, Solvay, and Casablanca, in
three Ignifluid furnaces, as well  as  in older
boilers  equipped with various furnaces:  a
chain grate stoker  at La Taupe, a  cyclone
furnace at Solvay, and a PF slag tap furnace at
Casablanca. Coals fired  are  those given in
Table 2.

  The results should be considered compara-
tive, rather than absolute, especially since the
S content is generally low.
NOX Formation
  The most interesting result, systematically
revealed during the various tests, concerns the
nitrogen oxides NO and NC>2 contained in the
flue gases of Ignifluid furnaces.
  As shown in  Table  4, the discharge  of
nitrogen oxides is two  or three times lower
than  that  measured in  all  the other com-
bustion processes.

  These results corroborate the observations
that have been made in many existing installa-
tions, mainly in the U.S.A., and show that the
formation of nitrogen oxides is a function of
combustion temperature, excess air, and gas
velocities.
                         Table 3. CHARACTERISTICS OF OTHER PLANTS
Location
La Taupe
Casablanca
Solvay
Type of
Combustion
Chain grate stoker
PF slag tap furnace
Cyclone furnace
Steam
Output, Ib/hr
77,000
190,000
220,000
Pressure,
Ib/sq in.
840
700
2,400
Superheat
Temperature,
915
860
1,000
"F



                                                                                    1-6-3

-------
                         Table 4. NOX MEASUREMENT IN FLUE GAS
Location
La Taupe
Casablanca
Solvay
Type of
Combustion
Ignifluid
Chain grate stoker
Ignifluid
PF slag tap furnace
Ignifluid
Cyclone furnace
Heat
Release, TO6 Btu
90
84
237
202
107
245
N0xin
Flue Gas, ppm
163
470
735
940
470
1720
Ratio3
NOX/106 Btu
2.68
8.45
10.70
17.80
7.55
18.50
         aNOx in cuff
   The  values of these  three  parameters  in
Ignifluid combustion combine  to reduce the
formation  of nitrogen oxides to a minimum:

    1.  Combustion  in an  Ignifluid bed  is
        completed in two stages —

        a. Combustion in a  reducing atmos-
        phere in the Ignifluid bed itself, where
        a  moderate  temperature,  slightly
        higher than  the  mean ash  sintering
        temperature, prevails.

        b. Combustion of the CO gas from
        the Ignifluid furnace in the area of the
        secondary air inlet (one or two rows).

        It is  moreover well known  that the
        two-stage  combustion is an  effective
        way  of  preventing  nitrogen oxides
        from being formed. It is an idea which
        has already been applied to the com-
        bustion of fuel oil, in particular in the
        Los Angeles District (B.W.  Co. patent
        No. 821,214).
    2.
Combustion  in  the Ignifluid system
takes place with a small amount of
excess air because, in the final stage, it
consists of burning a gas (CO), which
is  easier than  burning a solid fuel.
From this point of view, the Ignifluid
system  can readily bear comparison
       with the most efficient furnaces; e.g.,
       the cyclone furnace. This low volume
       of excess  air  could be reduced still
       more if we added certain alterations,
       to be described later.

   3.  The  long residence  time  the  fuel
       particles spend in the bed is obviously
       an  essential  characteristic  of the
       Ignifluid process.

  Finally, without wishing to draw too quick
conclusions, it seems feasible to think that the
nitrogen oxides content  is an increasing func-
tion  of the heat release, as witnessed by the
curves of Figure 2. This phenomenon, already
observed,  seems to be explained by the fact
that  the peak temperature in the combustion
chamber  increases with  the size of the units.


S02  Release

  With regard  to the release of SO2, measure- '<
ments carried out  as in the previous case, with
Draeger tubes, show  that the  result is ap-
preciably  the  same  whatever the  type of
furnace (except for stokers).

  The ratio   SO2       in which  SO2  is
             S x CO2
expressed  in ppm, S  (sulphur in the coal) in
percent,  and  CO2  in  percent,  is situated
between 50 and 60 (Table 5).
1-6-4

-------
              Table 5. MEASUREMENT OF SULPHUR IN FLUE GASES WITHOUT LIME
Location
La Taupe


Casablanca

Solvay

La Pochette
Type of
Combustion
Ignifluid

Chain grate stoker
Ignifluid
PF slag tap furnace
Ignifluid
Cyclone furnace
Ignifluid
S in the
Coal, %
1.00
0.55
1.00
2.00
2.00
0.7
0.7
0.5
S Retained
in Clinker, %
8.0
6.0
11.0
5.2
2.7
2.0
0.2
13.0
SO2in
Flue Gas, ppm
635
400
890
425
1120
1500
650
320
Ratio3
SO2/S-C02
55.0
52.0
70.0
50.0
57.0
52.5
50.0
43.0
      aSO2 in ppm;  C02 in %; and S in the coal in %
   This observation really shows nothing un-
usual, taking into account the furnace oper-
ating conditions.

   It occurred  to us to  use  additives (e.g.,
calcium  oxides  or  magnesium  oxides),
projected in the Ignifluid bed  in a granulated
condition.

   A comparatively short test performed in an
installation (La Taupe),  although  giving
restricted results, led  us to  expect some
possibilities of development.

  These  first results (Table  6) show that
sulphur retention increases proportionately
with the content  of lime injected, but it is
however  still very low,  compared with  the
total amount of sulphur.
   These restricted results may  be explained
 by higher temperatures at the rear of the grate
 (about  1350-1400dC)  where  an  oxidizing
 atmosphere prevails. With these temperatures,
 the sulphates created  inside the  bed de-
 compose releasing SC>2, but the lime remains
 in the clinker slags.

 VARIOUS S02 POSSIBILITIES

   However,  if SC>2 retention must be con-
 sidered, as is the case with a large  number of
 American  bituminous  coals,  various pos-
 sibilities seem to be available for the Ignifluid
 system, in  which  chemical reactions  occur
 either in the reducing zone or in the oxidizing
i zone.

   Bearing  in mind Professor (A.M.) Squires'
 views on the subject, we believe it would be
                    Table 6. SULPHUR RETAINED IN CLINKER WITH LIME3
                     Quantity of Lime
Without lime
With lime
Twice stoichiometric
quantity
Threefold stoichiometric
quantity
6
8
11
          S Retained
          in Clinker, %
                     3Test at La Taupe
                                                                                   1-6-5

-------
 advisable to perform the desulphurization in a
 reducing atmosphere, where a high conversion
 of CaO additives to CaS can be achieved.

   By  recycling flue gases  taken  at  the
 economiser outlet and reinjected  in the last
 compartments  of the  grate,  an  entirely
 reducing area can be maintained in the lower
 part  of the furnace below  the secondary air
 injection nozzles (Figure 3) at a temperature
 slightly higher than the sintering temperature
 (1050 to 1100°C or 1930 to 2000° F).

   The  recycled flue gas can also be used with
 primary air to regulate the bed temperature
 while maintaining the minimum gas flow rates
 for fluidization.

   As another consequence of CaO  addition in
 the  bed, a corresponding fluxing  action will
 result in the formation of suitable clinkers.

   However,  in  the  complex fluidized bed
 with  heterogeneous conditions, it  is difficult
 to predict the extent to which sulphur will be
 fixed as sulphide or, perhaps, as sulphate until
 experimental data is available.

   During tests  to be  performed, parameters
 other than recycling factors are to be taken
 into  consideration; e.g., the  additive grain
 sizing and injection points (undergrate, that
 is, in  primary air with coal).

   Another  approach  to  the  SC>2 retention
 problem is a process described by  Mr. Godel
 during the  first  International Conference in
November 1968,  by application of the U.S.
 patent No. 3,431,892.

   The  figure  (Figure 4) presented  by  Mr.
Godel was comparatively  complex; the instal-
lation involved  many  superposed  beds. The
present  arrangement (Figure 5) is limited to a
single secondary bed, made  of desulphurizing
additives in a finely granulated state.

   The simpler arrangement is an interesting
combination of  the Ignifluid process with

1-6-6
slagging  bed  at  high temperature  and the
fluidized  combustion  system  at  a  lower
temperature;  this one  is controlled by the
immersion of  boiler tubes in the fluidized
bed.

   The  main  advantage of  this  association
consists  of the possibility of reinjecting fly
ash,  consisting  of  sulphur-free  ashes  and
unburnt particles, into the slagging bed from
which all the ashes  are finally eliminated  in
the form of clinkers.

   The  secondary bed  can  work in  either
atmosphere: reducing or oxidizing. Because of
the contrast  between kinetics for formation
of CaS and CaSC>4, it appears to be better to
operate  with  a reducing bed where a high
conversion of CaO to CaS can be achieved.
   According to this program, it is possible to
make a valid test, by erecting (along with an
Ignifluid conventional boiler) a secondary bed
with only 5  to 10 percent  of the flue gases,
coming out  of  the boiler used  for  desul-
phurization.  The desulphurized flue  gases,
containing a  large quantity  of CO, would be
burnt  above  the  bed (Figure 6)  and  re-
admitted, after having been  cooled, above the
level of the secondary air inlets.
  As a  conclusion,  we may  say that the
Ignifluid grate  presents features that can be
used to efficiently control dust and nitrogen
oxides emissions.


  To overcome the problem of the retention
of sulphur in solid form, many tests of various,
kinds are required.


  Although such investigation is not justified
in France,  we would be willing to draw up a
test program,  and then carry it  out in col-
laboration  with the  organisations concerned
in one of their Ingifluid plants, on behalf of
any other countries where this need occurs.

-------
                  Figure 1.  Roches Moires power station at Casablanca.
    20
    15
 X
o
   O IGNIFLUID

— • OTHER PLANTS
     50
                         100
                                       150

                              HEAT RELEASE, 106 Btu
                                                                  200
                   Figure 2. Nitrogen oxides content versus heat release.
                                                                               250
                                                                                    1-6-7

-------
            0  D   D  0   D   D
                                                                                FLUE GAS
                                                                            RECIRCULATING FAN
                    PRIMARY AIR
                                                    STEAM AIR HEATER
                      Figure 3.  Desulphurization in a reducing atmosphere.
1-6-8

-------
                           FEED WATER
                             HEATER

• .'^i<."' •&.'*-' '.'-\zJ ">^ •'.'?<-"•
-. -.*.*..:.,'•..'.' '•'..-. .*.'..'; .*'..:',•.*.'.•
                                     TO HIGH-
                                     TEMPERATURE
                                     FLUID BED
                           SUPERHEATER
                         EVAPORATING
                         TUBE NEST
                                                                          ^
TO HIGH-
TEMPERATURE
FLUID BED
                                 RADIATING
                                 PANELS
                                 SECONDARY AIR
                                      I
                                     9.	-ll--N-.,
                                      PRIMARY AIR
             Figure 4.  Former multi-superposed-bed arrangement.
                                                                                 1-6-9

-------
   SECONDARY AIR
  SECONDARY BED
       I.F. BED
• IMMERGED BOILER
     TUBES
                                    PRIMARY AIR

                    Figure 5.  Present single-secondary-bed arrangement.
1-6-10

-------
AIR
AIR
                   IMMERGED WATER TUBES




                        Figure 6.  Ignifluid conventional boiler and secondary bed test arrangement.

-------
SESSION II:




        Control of Combustion Pollutants
SESSION CHAIRMAN:




        Mr. Alvin Skopp, Esso Research and Engineering

-------
                          1. THE RETENTION  OF SULPHUR
                                                       BY LIMESTONE
         IN A PILOT-SCALE FLUID-BED COMBUSTOR

                  D. C. DAVIDSON AND A. W. SMALE

                       National Coal Board,  England
INTRODUCTION

  The advantages of burning coal in fluidised
beds have  been  widely  publicised1'2  and
much  experimental activity  has been
reported. The fluid bed is an excellent gas/
solids contactor and is, therefore, not only
an ideal solid fuel  combustion system, but
also offers attractive possibilities for removing
sulphur oxides from combustion gases.

  It has been  stated3 that limestone or
dolomite addition is the simplest means of
removing sulphur oxide from flue gases in
existing  boilers, both solid fuel and oil fired,
but that this  simplicity is unfortunately ac-
companied by inefficiency due principally to
very limited residence time in what has been
claimed4  as the effective temperature range
(950 - 1230°C; 1770 - 2250°F). This problem
is  largely overcome by fluidised-bed com-
bustion  which is sensibly isothermal. How-
ever, the typical operating  range  of 800
900°C (1472 - 1652°F) is, according to Reid's
predictions,4   too  low for  rapid reaction
between limestone and SC»2.

  This  paper discusses experimental work
carried  out  at the Coal Research Establish-
ment of the National Coal Board; some of this
has already been reported1  in outline. It
refers  to a  current  research  programme
sponsored  by  the  National  Ah" Pollution
Control Administration  (NAPCA),  whose
permission to quote results from this work is
gratefully acknowledged.
EXPERIMENTAL

  The combustor used for the work is shown
in outline in Figure 1; a more detailed descrip-
tion of its function is given by McLaren and
Williams.1  The bed is 6 inches in diameter;
bed depth is normally controlled at 2 feet. It
was possible to recycle fines from a primary
cyclone if desired.

  The experimental programme has included
six coals to date, three from the U. K. and
three from the U. S. A. A description of these
fuels as fed to the combustor is given in Table
1. In all but one  case, the same limestone
(U.K.) was used and in all cases the feedstock,
both coal and limestone, was nominally -1/16
inch. However, there were differences in the
size analyses of the various batches of lime-
stone; these are shown in Table  2. In one
series of tests with Peabody  No. 10 coal, a
U.S. limestone (Ref. BCR-1359)  was used.
Supplied by NAPCA, this had previously been
used  in related  studies at  the  Argonne
National Laboratories.

  Table  3  is a summary of experimental
conditions used for the tests described.

  During all  runs, samples  of solids were
collected and an  extensive  programme  of
analysis  carried out to obtain mass balances
for all  important  components.  It is  not
proposed to discuss these  in  the  present
context  except to note that hi the earlier
work there  was  frequently a discrepancy in
the sulphur balances amounting to some 20
                                      II-l-l

-------
                                Table 1. DETAILS OF COALS
Coal
% Compared
(as received)
Moisture, %
Ash, %
V.M., %
F.C., %
Sulphur, %
Size Analysis
+ 10
10 + 30
30 + 60
60 +120
-120 +240
-240
U.K.
Goldthorpe
5.0
17.0
29.7
47.7
2.05
0.5
40
24
17
9
10
Welbeck
4.2
18.2
29.7
47.9
1.25
0.0
26
40
17
7
10
Park Mill
2.1
16.5
31.9
49.5
2.4
0.1
60
19
10
5
6
U.S.A.
Arkwright
2.1
13.9
34.4
49.6
2.25
0.0
44
26
15
7
8
Farmington
1.8
9.8
28.3
50.2
2.25
	
48
25
14
8
5
jArgonne)
Peabody No. 10
9.1
11.4
38.2
41.3
4.1
0.4
52
22
13
7
6
                       Table 2. PARTICLE SIZE OF BRITISH LIMESTONE
Sieve Size
+ 10B.S.
-10+30
-30+60
- 60 +120
-120+240
-240
As used with: Goldthorpe and
Farmington
0
27
20
18
19
16
Arkwright
0.7
65
19.5
9.3
4.6
0.5
Welbeck, Park Mill,
and Peabody No. 10
0.5
52.0
19.0
10.0
7.5
11.0
percent of the  input. This was particularly
noticeable  at  the higher levels of limestone
addition.  In the  later tests (Welbeck, Park
Mill, and Peabody No. 10), the better sulphur
balances obtained were ascribed both to the
II-1-2
greater attention that was paid to measuring
the quantity of material in the bed  at  the
beginning of the test period, and an improve-
ment  in  sampling procedures  for  solids
products.

-------
                     Table 3.  EXPERIMENTAL CONDITIONS INVESTIGATED

Bed temp., "C
Bed height, ft
Fl. velocity,
ft/sec
Recycle
Ca/S ratio
Goldthorpe
800
2

2
Yes
0-2.2
Welbeck
800
2

2 and 3
Yes and No
0-2.8
Park Mill
800
2

3
No
0-2.8
Arkwright
800
2

2
Yes
0-1.8
Farmington
800
2

2
Yes
0-1.8
Peabody
No. 10
700 and 800
2 and 3

3
No
0-3.0
 RESULTS

   This paper considers only flue gas sulphur
 dioxide and the extent to which it is reduced
 by the addition of limestone under various
 test conditions.

   Figure 2 plots the data. The zero condition
 is represented by the combustion of the coal
 without limestone addition; subsequent runs
 with  limestone  are plotted in terms of the
 percentage by which the SC>2 in the flue gas is
 reduced against  the quantity of lime injected.
 This   latter quantity,  expressed  as the
 Calcium/Sulphur  (Ca/S)  ratio,  does not
 include the calcium natural to the coal. Thus,
 the figure  reflects directly the effect of the
 limestone  added and all  data refers  to a
 common origin, the  "no limestone" condi-
 tion.

   The results with Welbeck coal,  which has
 been  used over  the widest variety of experi-
 mental  conditions,  cannot be clearly repre-
 sented in Figure 2. These are therefore shown
 in more detail in Figure 3, together with the
 Goldthorpe results already shown in Figure 2.
DISCUSSION

  The results are discussed in terms of the
following variables:

  1. Quantity of limestone.
  2. Quality of limestone.
  3. Effect of fluidising velocity.
   4.  Effect of recycle.
   5.  Coal caking properties.
   6.  Coal sulphur content.
 Quantity of Limestone

   Considering all the data with the British
 limestone at a fluidising velocity of 3 ft/sec
 without recycle of fines, these form a group
 with the following average values:
     Ca/S  Ratio
          1
          2
          3
  SC>2 Retained %

Average    Range
  48
  75
  94
40-54
68-83
90- 10
1   At a Ca/S ratio of 2, the likely practicable
maximum, the figure of 75  percent of SC«2
retained  by the limestone  means  that  38
percent  of the  calcium  was  utilised. The
results for the  U. S. limestone with Peabody
No.  10 coal shows 63 percent of SC«2 retained
at a Ca/S ratio of 2 which is a calcium utili-
sation of 32 percent. This is  poorer than the
results with British limestone  and is con-
firmed by the  results  at a Ca/S ratio of 3
which show  a  progressively poorer perform-
ance by  the  U. S. limestone.  These differing
results at high  Ca/S ratios  could imply small
difference in the reactivity of the limestone
but may  be due to other causes; e.g., particle
size.
                                     II-1-3

-------
 Quality of Limestone

   All of the limestones contain more than 97
 percent  calcium carbonate which removes one
 possible  source of variation. The differences
 in particle size analysis for the various batches
 of U. K. limestone could also be  a source of
 variation in performance. However,  because
 the  earlier tests on Goldthorpe coal using a
 finer limestone  also used  recycle,  a direct
 comparison between the different U. K. lime-
 stone batches  is not  possible.  The  poorer
 performance  of  the  U.   S.  limestone,
 detectable principally at high Ca/S ratios, may
 be due to its finer particle size compared with
 the  equivalent batch of U. K. limestone. The
 tests concerned  were all  at 3  ft/sec  without
 recycle and it is likely that under these condi-
 tions much  fine limestone is lost from the
 system  without being utilised. A  further test
 using the U. S. limestone with recycle could
 resolve this point.


 Effect of Fluidising Velocity

   The tests with Welbeck  coal provide  the
 only direct evidence on  this variable; from
 Figure 3 it is apparent that there is a small
 loss  in  efficiency  caused by  raising  the
 fluidising velocity from 2 to  3  ft/sec.  This
 amounts to  about 10  percent of the input
 sulphur content for Ca/S ratios between 1 and
 2. If this factor is applied to the Goldthorpe
 data over the same range of limestone/sulphur
 stoichiometry, then it could be predicted that
 Goldthorpe  would  still  respond  marginally
 better to  limestone  addition  than did  the
 other coals. However, it must be remembered
 that  the limestone used for Goldthorpe was
 finer than for Welbeck.
Effect of Recycle

  Again,  the  only  relevant data  is  from
Welbeck coal at 2  ft/sec;  it is clear from
Figure 3 that no effect was detected. This
seems slightly improbable at first sight but, as
the limestone was approximately 70 percent +
II-1-4
60 B.S. (250 micron), it could be argued that
the small proportion of fine limestone lost by
elutriation  would  not  significantly help  the
sulphur  retention  figure  even if it were
recycled. The  effect  of  recycle  at  higher
fluidising velocities could  be expected to be
more marked.
Coal Caking Properties

  The Goldthorpe coal is weakly caking and
originally its  superior  performance  to
Farmington and Arkwright was thought to be
at least partly on account of the highly caking
properties  of  the  two latter fuels. Welbeck
(also a weakly caking coal) gave results closely
similar to  Farmington and Arkwright which
appears to contradict this assumption. How-
ever,  if a  10  percent  bias is applied to the
Arkwright  and  Farmington data,  to  equate
them  with  the  3  ft/sec  results, they  then
become closely comparable with Park Mill-a
caking coal—and appreciably lower than the
Welbeck results.

  Thus, using a common limestone with  all
data "corrected" to 3 ft/sec, the caking coals
fall into a  compact  group inferior in per-
formance  to the other coals tested. Develop-
ing  a  hypothesis  to  explain this behavior
presents serious problems.

  Arguments  based   on  incomplete  com-
bustion leading to  incomplete  release  of
sulphur as S(>2 should lead to  better  than
expected efficiencies in terms of flue gas SC>2
analysis. Agglomeration in the bed, leading to
envelopement  of limestone  particles  and
hence their incomplete sulphation, is possible;
however, no agglomerates have been observed
in any of the experiments. Simple delay in the
release of  sulphur oxides  from the  bed is
possible; however,  since both Arkwright and
Farmington were tested both with recycle of
fines and with a fine limestone, there is plenty
of opportunity for the recycled limestone to
contact the sulphur oxides. However, it could
be argued  that by contrast to the situation

-------
within the fluid bed, the gas/solid contacts in
the gas stream are insufficient to make up for
lower  sulphur retention in the bed. Further
experiments  using caking  coals  with and
without  recycle are needed  to  explore this
suggestion more fully.
  A reduction in  sulphur retention is  ap-
parently related to coal properties in that the
more strongly caking coals are less responsive
to limestone addition. An adequate explana-
tion for this has yet to be developed and
further tests are needed.
Coal Sulphur Content

  There is no directly observable effect due to
the total sulphur content of the fuel. Welbeck
has the lowest sulphur content  of the coals
tested (1.25  percent a.r.) and the results for
this coal correspond quite closely to those for
the Peabody  No.  10  coal  which has  the
highest sulphur content  (4.1 percent a.r.).
Information currently  available on the forms
of sulphur in the coals does not lead to any
meaningful  correlation,  but  this is  to  be
examined in more detail in other experiments.
CONCLUSIONS

  At a  fluidising  velocity of  3 ft/second
without recycle of fines, sulphur retentions of
better than 90 percent can be achieved but
require  approximately  three  times  the
stoichiometric quantity of limestone.

  An increase in fluidising velocity from 2 to
3 ft/second decreases sulphur retention by
approximately  10  percent  of the  sulphur
input. At  these modest fluidising velocities
the effect of recycle is small but may  be
increased significantly: at higher velocities, or
with finer limestone.
                                              BIBLIOGRAPHY
 1. McLaren, J. and  D.F. Williams. Jnl. Inst.
   of Fuel. 42: 303-308, August 1969.

 2. Hoy, H.R. and I.E. Stanton. Amer. Chem.
   Soc. Div. of Fuel Chem. 14, No. 2: 58-77,
    1970.

 3. Slack, A.V. and H.L. Falkenberry. Jnl. of
   Engng. for Power, Trans. ASME, Series A.
   92: 5-10, January 1970.

 4. Reid,  W.T. Jnl.   of Engng. for Power,
   Trans.  ASME,   Series  A.  92:   11-15,
   January  1970.

ACKNOWLEDGEMENT

  The  work described in  this paper  was
carried out  partly as part  of the  research
programme of the Research and Development
Department  of  the National Coal Board and
partly as  part of a joint National Coal Board/
National  Air Pollution  Control Administra-
tion research programme. The views expressed
are those  of  the authors and not necessarily
those of  either  the Board or the Administra-
tion.
                                                                                  IM-5

-------
        SO2 SAMPLING
           POINT
S02 SAMPLING
  POINT
     FINES
     CATCHPOTS
                                                                       'FINES
                                                                       CATCHPOT
            COOLING
             WATER
            RESERVOIR
                                        VIBRATOI
                                                                                 AIR
                                                       ASH
                     Figure 1.  6-inch diameter fluidised-combustion rig (CRE).
11-1-6

-------
   100
   90
   80
    70
    60
    50
    40
    30
    20
    10
   100

    90

    "

    n
    60

    50
uj   40
u.
?   30
    20
    10
                                                 1 GOLDTHORPE COAL 4 BRIT. LIMESTONE
                                                 2 WELBECK COAL+ BRIT. LIMESTONE
                                                 3 ARKWRIGHT
                                                 4 FARMINGTON
                                                 5 ARGONNE COAL + BRIT. LIMESTONE
                                                 6 PARK MILL COAL4- BRIT. LIMESTONE
                                                 7 ARGONNE COAL-)-US LIMESTONE
                           N.B. GOLDTHORPE, ARKWRIGHT, AND FARMINGTON TESTS 
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                2. CHARACTERIZATION OF EMISSIONS
    FROM FLUIDIZED-BED COMBUSTION  OF COAL
                  AND CONTROL  OF SULFUR EMISSION
                                                 WITH LIMESTONE

                   R. D. GLENN AND E. B. ROBISON
                        Pope, Evans and Robbins
INTRODUCTION
  This paper presents a  brief review of
emission studies in the area of fluidized-bed
combustion, conducted  by Pope, Evans and
Robbins for the National  Air  Pollution
Control  Administration.  The  studies, in
general,  involved  a  characterization of
emissions from fluidized-bed combustion of
coal,  sujfur emission control by  limestone
injection  into the bed,  and sulfur emission
control by combustion of coal in a limestone
bed. This paper also presents an estimate of
the potential for sulfur emission control by
continuous regeneration of a limestone bed
system based  on  some preliminary  experi-
mental data.


CHARACTERIZATION OF EMISSIONS

  The initial investigation to characterize the
emission from the coal-fired  fluidized bed
indicated  levels  of sulfur dioxide,  sulfur
trioxide,  nitric oxide,  hydrocarbons,  and
particulates,  and the extent to which these
could be modified  by purely  operational
changes; e.g., changes in excess air  rate, coal
feed rate, and  bed temperature. A  summary
of the results of this investigation follows.


Sulfur Oxides
balance was retained by the ash. A very small
percentage, 1  percent or  less, appeared as
sulfur trioxide in the flue gas. In the absence
of a sorbent, the sulfur emission was substan-
tially unaffected by changes in the operating
variables.

Nitrogen Oxides

  Emission of nitric oxide (NO) was found to
be the dominant form in the NOX group. The
emission level, although low in comparison to
the level of other coal-fired units, was still an
order of magnitude greater  than the  equi-
librium value predicted from the bed tempera-
ture.  Furthermore, the  NO level did not
change when  bed temperature was changed.
This  result led to the conclusion that NO
emission was controlled by local temperatures
higher than the measured bed temperature.

  The  NO  emission  responded  only  to
increase in the oxygen content in the flue gas
(increase in excess air).  In one test an NO
concentration  of 280 ppm,  observed  at  1
percent oxygen in the flue gas, increased to
340 ppm at 4 percent oxygen. This variation
is shown in Figure 1. The average concentra-
tion of a test  series conducted at 3 percent
oxygen in the full-scale unit was 275 ppm.

Hydrocarbons
  About 95 percent of the sulfur in the coal      The full-scale boiler could be operated with
appeared as sulfur dioxide in the flue gas; the    as little as 5 percent excess air without visible
                                     II-2-1

-------
smoke in the flue gas; however, hydrocarbon
concentrations were  found to be as high as
 1500 ppm (methane) at this excess air level.
This high concentration was reduced sharply
as the excess air  rate was increased. Hydro-
carbon burnup was essentially complete at 24
percent excess air (4 percent C>2 in the flue
gas). The variation in shown in Figure 2.
Particulates

   In  the test system, particulate emissions
were  a function of the cyclone collector ef-
ficiency. Analysis indicated that 90 percent of
the  fly  ash  that passed  the cyclone  was
smaller than 20 microns in particle size.
EFFECT OF OPERATIONAL CHANGES
Effect of Limestone Particle Size

   When limestone (BCR 1359) was added to
the bed in a comparatively coarse particle size
(-7 +14  mesh),  the sulfur absorption  was
rather poor. Under the most favorable condi-
tions,  the limestone utilized  in  the sulfur
capture was  limited to  about 20  percent.
Emission  of SC>2 was reduced by about 20
percent at a stoichiometric ratio of  1.0 and
about 40  percent at a stoichiometric  ratio of
2.0. A dolomitic limestone (BCR 1337) of the
same size  was  more  effective  when  the
stoichiometric  feed  ratio was  based  on the
calcium fraction  alone; however, on a total
weight basis, it was equally poor.

   At this  point in time, it became evident to
many researchers that utilization of  sorbent
was limited by the formation of a sulfate shell
on the  surface  of the sorbent  particle.
Improvement  required  an  increase  in  the
surf ace-to-mass ratio  of the particle; i.e., a
reduction  in particle size.

   When  limestone  was  ground  to  a  fine
powder (-325 mesh) and injected into the

II-2-2
bed,   the  sulfur  capture  was  markedly
improved.  The  reduction  in  SO2  emission
with increase in stoichiometric rates is shown
in Figure 3. The trend shows SC«2 reductions
of 40, 66, and 84 percent at respective Ca/S
ratios of 1, 2, and 3. The corresponding lime-
stone utilizations are 40, 33, and 28 percent.
These results were obtained with a high sulfur
(4.5 percent) coal burning in a 10- to 12-inch
deep fluidized bed operating in the tempera-
ture range of  1500    1600°F, and with  a
superficial  velocity of  12-14  ft/sec.
Comparable  results  were  observed  with  a
medium (2.6 percent) sulfur coal.

  The lower  line  on  Figure  3  shows  the
greater reactivity obtained with a  dolomitic
limestone; however, the stoichiometric ratios
are calculated for calcium only. If the weight
of  magnesium  is included  and results  are
expressed  in practical terms, such as pounds
of stone fed per pound of sulfur removed,
dolomite has no advantage over limestone.

  In a short series of tests with calcined lime-
stone, relatively poor results were obtained, as
indicated by the upper line of Figure 3.

  A detailed study with the particular lime-
stone  used  (No.  1359) showed  that fine
grinding  would  be  necessary  for  favorable
sulfur capture.  This  limestone is one of  the
most durable, however, and such fine grinding
may not be necessary with less durable lime-
stones  which may decrepitate from  thermal
shock during the initial calcination stage.

  Figure 4  shows the effects  of  limestone
particle size, bed depth, and temperature on
SC>2 reduction. All of the data in this series
was obtained while burning a 3-percent sulfur
coal in a fluidized bed of sintered ash.  Lime-
stone was  injected at a stoichiometric ratio of
2.6.

  At particle sizes below about 400 microns,
the beneficial effects of reducing the particle
size  of the  limestone  injected are clearly
evident. At particle sizes above  400 microns,

-------
some  other  factor—probably  increased
residence time  in bed—more than offsets the
decrease in surface area, and a slight improve-
ment is obtained with increase in particle size.
From  a practical  viewpoint, however, only
high reductions are important.

   If  SOX emissions  from  a fluidized-bed
boiler are to  be controlled by the injection of
limestone, it must  be a very finely divided
limestone.
Effects of Bed Depth and Temperature

   Figure  4  also  shows  the  effect of  bed
depth. As expected, an 18-inch bed shows a
higher absorption than does a 10-inch bed.

   Also shown is the effect of bed tempera-
ture; the beneficial effects of operating  at a
temperature of  1550°F are clearly indicated.
The poorer results obtained at 1675°F and at
1800°F  bring to mind some of the  poor
results that have been obtained by injecting
powdered  limestone into  a conventional
pulverized-coal-fired  furnace which  must
operate at even higher temperatures. One of
the  outstanding advantages of fluidized-bed
combustion is  its ability  to  operate
economically  at temperatures of  1500 to
1600°F. This makes the process particularly
well  adapted for  sulfur  emission  control
through the use of limestone.
LIMESTONE INJECTION

  The  injection  of  limestone  for sulfur
control  had no  observable effect on hydro-
carbon emission. In several instances  the ad-
dition of limestone reduced NO emission but
the  effect  was  not reproducible.  Sulfur
trioxide emission was reduced to zero when-
ever limestone was injected.

  Results  of the limestone injection  tests
indicate that a 4.5-percent sulfur coal can be
made the equivalent of a 1-percent sulfur coal
with the injection of 1359 limestone at a rate
of 27 lb/1000 Ib of coal. A 2.6-percent sulfur
coal can be  converted to a 1-percent  sulfur
equivalent  with  the  addition of  1359 lime-
stone at a  rate of 10 lb/100 Ib of coal. The
corresponding Ca/S ratios are 1.9 and 1.2.
LIMESTONE BED

   As an  alternative to  injecting  pulverized
limestone,  limestone  was used  as  a bed
material. The effect on  SO2  emission was
remarkable.

   Tests conducted with a limestone  bed  of
-10 +20 mesh  particle size and a 3-percent
sulfur coal indicated that SO 2 emission could
be almost completely eliminated for a period
of 2 to 3 hours. After  this period, SO2 began
to appear in the flue gas above the bed with
concentrations  increasing  with time. After
about 4 hours,  the SO2 emission represented
about 30  percent  of trie input  and the con-
centration  of sulfur in the bed rose to 7.4
percent. This concentration corresponds to a
limestone utilization of about 16 percent. A
once-through process in a limestone bed was
clearly uneconomical, and attention was given
to the regeneration of the limestone for reuse.
Emissions monitored during both adsorption
and regeneration are shown in Figure 5.

   In this  test the  sulfur content of the coal
was again 3 percent, corresponding to a flue
gas concentration  of 0.24 percent SO2- As
before,  for the  first  2 hours,  substantially  all
of the SO2 was removed. The concentration
in the flue gas then rose gradually and, after a
little over 3 hours, the SO2 emission rose  to
about a third of what it would have been if
coal ash rather than limestone had  been used
as the bed material.

   The  operating  conditions were then
changed. The coal rate was increased so that
the temperature, which had been 1550°F, in-
creased  to 1950°F. The oxygen concentration
in the flue gas, which had been 3 percent, was
                                     II-2-3

-------
dropped to 1 percent. A rapid desorption of
SC>2 took place, and 90 percent of that which
had been absorbed during more than 3 hours
of operation, was desorbed  in a few minutes.
The  SC>2  concentration of  the  flue  gas
reached  a peak of 8.1 percent—a value which
is more  than 30 times higher  than  the  0.24
percent  that would have been  experienced if
the coal had been burned without the use of
limestone in the bed.
estimated from the SCb removal efficiency of
the regenerated limestone bed. Experimental
observations of SC>2 emission after regenera-
tion of the bed showed that the emission rate
increased linearly  with time. The total sulfur
in the bed also increased simultaneously in a
linear  fashion.  The  rate  of sulfur emission
could, therefore, be related to the total sulfur
in the  bed  for a system operating  with a
3.0-percent sulfur coal.
 REGENERATIVE PROCESS

   The results observed in this test suggested a
 regenerative process for using a limestone bed
 to control SC>2 emissions from a fluidized-bed
 boiler. Absorption might be carried out in the
 fluidized  bed of the main part of the boiler.
 Desorption, and the production of a flue gas
 that is very high in SC>2 concentration, may
 be carried out in another section. The product
 shell  limitation—the  mechanism  by which
 absorption of SC>2 on the surface  of a lime-
 stone particle reduces its effectiveness for the
 absorption  of  additional SC>2—can  be
 circumvented by repetitive use of the particle
 surface. The high concentration of SC>2 in the
 small volume of flue gas from the desorption
 section can be removed and/or recovered by
 any of the  variety  of  processes  that have
 previously been proposed for the treating of
 conventional boiler flue gas.

  These processes,  presently high in cost, will
 become economical as a result of the high
 S(>2 concentration and low volume of flue gas
 to be treated.

  A patent application has been filed on this
 "SC>2 Acceptor Process."

  The performance of a  continuous absorp-
tion  and  regeneration  system  has  been
   A sulfur balance around an absorption cell
employing recycle of regenerated limestone
follows the relation:
                                      (1)
where L is the bed mass, xs the sulfur fraction
in the bed, t time, GQ the sulfur input rate, XQ
the sulfur fraction  in  the regenerated lime-
stone, fL the fraction of the bed recycled per
unit time, and Ge the flue gas sulfur emission
rate.
   Relating  Ge  to  the sulfur in the  bed,
equation (1) becomes:

        = G0 + (x0 - Xs) L fL - KxsL    (2)
where  K  is  the  proportionality  constant
between  the ps sulfur emission rate and the
total sulfur in the bed.
  Integrating equation (2), the sulfur fraction
of the bed varies as follows:
            _ GQ + xQLfL - [GQ + XQLfL - XQ (KL - LfL) ] e - (K + fL)t
                           KL +
                              (3)
II-2-4

-------
  The transient  term indicates that xs ap-
proaches equilibrium rapidly  as the  recycle
rate is increased.
  At  equilibrium,
disappears and
              the  transient  term
     xs =
       x0LfL  Ge
       + I)  ~KL
(4)
  If the  sulfur fraction  of the  regenerated
limestone (XQ) is small, the term xgLfL can
be neglected. The sulfur emission rate Ge is
then related to the sulfur input rate, the gas
constant, and bed recycle rate by the relation:
  K(G0)
= 	  or
  K + fL     GO
                              K
                                       (5)
                                         The value of K determined by experiment
                                       was found to be 0.32 hr"^ (Ib sulfur emission
                                       per hr per Ib sulfur in the bed).
          The measured values  of
                          — £, — §•, and L
                          dt    dt
were  0.37  lb/hr2,  0.0215 hr1, and 55 Ib,
respectively.

  Figure 6  is a plot of the calculated SC>2
removal versus the number of bed changes per
hour. It is indicated that, with only  one bed
change per  hour, a 76-percent  reduction of
SO2 emissions may be obtained. With two
bed  changes per hour,  an  87-percent reduc-
tion is predicted, and with three bed changes
per hour, a 91-percent reduction is predicted.

  These  results  indicate a  very favorable
potential for sulfur  emission control by  lime-
stone regeneration in a continuous operation.
Further investigation is planned.
                                                                                   II-2-5

-------
    400
    300
LU
Q
X
O
O
cc
200
    100
   1500
                        1.0
                                         I
                                    2.0             3.0
                                  OXYGEN IN FLUE GAS. %
                                                                         4.0
5.0
                  Figure 1.  Effect of oxygen concentration on NO emission.
                                                                          4.0
                  1.0              2.0              3.0
                                 OXYGEN IN FLUE GAS, %
         Figure 2.  Effect of oxygen concentration on hydrocarbon emission.
                                                                                         5.0
11-2-6

-------
o
o
LU
cc
01
Q

O
Q
GC
U-
    20
O
Q  40
LU
X
O
Q
tr  60
    80
    100
                                                 2.0
                                             Ca/S RATIO
                    Figure 3.  SC>2 reduction - - effect of limestone-to-sulfur ratio.
                                             I STATIC BED
                                                DEPTH
                                             •OlO in.
                                                >18 in.
T
I
               200       400      600       800      1000     1200

                               LIMESTONE PARTICLE SIZE, microns
        1400
                                                                                 1600
                           1800
 Figure 4.  S02 reduction - - effect of limestone particle size, bed depth, and temperature.
                                                                                          II-2-7

-------
   0.8
   0.7
   0.6
   0.5
se
w  0.4
UJ
o
Q
oc
   0.2
    0.1
 u- 2000
 O

 5 1800
   1600
                   8.1% S02

                   ~1     T
       S02 INPUT EQUIVALENT
                   \
                             XT'
                I  /  \l     I

                  /   \   %°2


               .J\f\    TEMP
                I      iV/T"
      0


 Figure



II-2-8
          234

        TEST PERIOD, hours
5. S02 in flue gas - - regenerative

   process using a limestone bed.
        1.0       2.0      3.0
    BED CHANGES PER HOUR BY RECYCLE

Figure 6. Calculated S02 removal by

         regenerative process.
                                                                                 4.0

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            3.  POLLUTION  CONTROL  CAPABILITIES

                        OF  FLUIDIZED-BED COMBUSTION

                       L. J. ANASTASIA, E. L. CARLS,
              R. L. JARRY, A. A. JONKE, AND G.  J. VOGEL


                       Argonne National Laboratory
ABSTRACT

  The combustion of a high-sulfur coal has
been  studied  in  a 6-inch dia fluidized-bed
reactor. Basic additives containing CaO (e.g.,
limestone and dolomite) have been used for
SC>2 control during combustion. The parame-
ter having the greatest effect on SC>2 emission
has been found to be the Ca/S ratio  in the
additive and coal feed streams. At Ca/S mole
ratios of 2.0 to 2.5, about 80 to 90 percent of
SC>2 emission can be prevented.  When coarse
additive particles  are  used  as  both  the
fluidized bed  and additive,  the optimum
temperature range for SC>2 removal appears to
be 1480 to 1550°F.
INTRODUCTION

  Fluidized-bed combustion is being studied
at Argonne National Laboratory as a method
of reducing the  quantity of atmospheric
pollutants  (oxides of  sulfur  and nitrogen)
released during the combustion of fossil fuels.
This work  is being carried out under contract
with  the  National  Air  Pollution Control
Administration.

  The concept of fluidized-bed combustion
involves burning fuel in a fluidized bed of
solids.  In  an industrial application,  boiler
tubes  would be immersed in the bed.  The
rapid motion of the fluidized particles at the
heat  transfer  surfaces, together  with high
heat-transfer rates between gas and particles,
 makes the fluidized bed  a  highly efficient
 heat-transfer medium.  If the excellent heat
 transfer  characteristics of fluidized-bed
 combustors allow combustor  size to  be
 decreased, lower capital and operating costs
 than for conventional units may be expected.
 Another advantage of the fluidized bed in this
 application is  that it  is  a  highly  efficient
 contacting medium for carrying out gas-solids
 reactions; therefore, use of a bed material that
 reacts  with  gaseous  pollutants  generated
 during combustion offers good prospects for
 controlling air pollution.  Also, combustion
 can be carried out at lower temperatures in a
 fluidized bed  than by  conventional  com-
 bustion methods.

   Disadvantages of the fluidized bed include
 a gas velocity limited by the extent of entrain-
 ment of unburned carbon from the fluidized
'bed; therefore, relatively large bed areas and a
 large number of fuel introduction points are
 required. The rapid motion of the fluidized
 particles results in some decrepitation of bed
 particles; the viscous drag of the gas is suf-
 ficient to entrain  fine particles from  the
 reactor, necessitating off-gas cleanup of both
 additive and fly ash particles.

  To control sulfur oxide emissions, lime-
 stone  or  dolomite  additive may  be  fed
 continuously to the fluidized-bed combustor,
 where calcination  to lime  (CaO)  occurs
 simultaneously  with combustion of the fuel
 and in situ reaction of sulfur oxides with lime
 to form calcium sulfate.
                                      II-3-1

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   Two alternative modes of operation  for a
desulfurizing  fluidized-bed  combustor  have
been considered in which the major differences
pertain to additive particle size and composi-
tion of the fluidized bed.  In  both modes of
operation, crushed and washed coal (-6 and
-14  mesh)  was  used;  at  the fluidizing-gas
velocities employed (3 and 9 ft/sec), essential-
ly all the coal ash generated by combustion of
these  coal particles  was elutriated from the
fluidized bed. In the first mode of operation,
pulverized additive (<325 mesh) was used for
SC>2  control,  and  the fluidized bed  was
refractory alumina having a relatively coarse
particle size (30 mesh). Since  most of the ash
and additive  was elutriated  from  the  bed
during a run,  the fluidized bed at steady state
was composed largely of alumina particles. In
the second mode of operation, large additive
particles used for SC>2 control were of suf-
ficient size to remain in the fluidized bed. At
steady state  in  this mode of operation, the
fluidized bed was composed  essentially of
partially sulfated additive material.

EQUIPMENT, PROCEDURE, AND
MATERIALS

   Figure 1 is a schematic diagram of the
fluidized-bed  combustor system.  The  com-
bustor is a  6-inch dia stainless  steel vessel
equipped with  electric  heaters  and an  air
cooling system  for temperature  control. To
start a run, preheated air is passed  into the
combustor through   a  bubble-cap  air  dis-
tributor  mounted  on  the  bottom flange.
Electric  heaters raise  bed temperature to
1000°F, at which point coal is fed to the bed
at  controlled  rates. Ignition of the   coal
increases the bed temperature to the desired
operating temperature (i.e., M600°F). When
combustion conditions have stabilized,  feed-
ing of additive  and/or recycle of elutriated
fly-ash/limestone   mixture  is  started.
Variable-drive volumetric  screw  feeders
mounted on  scales allow solids feed rates to
the  system  (coal,  additive,  and  recycled
solids) to be metered.  The  solids are fed
pneumatically to the fluidized bed at a point
just above the gas distributor.
II-3-2
   Flue  gas from  the combustor  is  passed
through  two  high-efficiency   cyclone
separators in series and a  fibrous-glass final
filter to remove entrained solids. Downstream
from the  cyclones, approximately  5 percent
of the  total flue  gas is diverted  to a  gas
analysis system. The water content of the flue
gas sample is reduced to 3000 ppm (by con-
densation  and refrigeration)  to  prevent  the
moisture from interfering with gas analysis.
Continuous analyses of the  dried gas for NO,
SC>2, and  oxygen are  carried  out by infrared
analyzers  and  a  paramagnetic  oxygen
analyzer. Gas chromatography provides inter-
mittent  analyses for  CC<2. Periodically,  the
bed  and  overhead solids  are  sampled  for
chemical  analysis  and  to obtain  material
balances.

  A Central Illinois bituminous coal mined in
Christian County has been used in all experi-
ments reported here. The coal  contains 4.5 wt
percent  S and 12.3 wt percent ash and has a
heating  value of 1 1,550 Btu/lb. The additive
materials studied include:

    1.  Limestone No. 1359,  Stephens City,
       Virginia (97.8 wt percent CaCC«3,  1.3
       wt percent MgCC«3).

   2.  Limestone  No.  1360,  Monmouth,
       Illinois (78.0 wt percent CaCC«3, 22.0
       wt percent
   3.  Dolomite No. 1 337, Gibsonburg, Ohio
       (53.4 wt percent CaCO3, 46-5 wt Per~
       cent
   4.  Tymochtee dolomite, Huntsville, Ohio
       (49.3 wt percent CaCO^, 36.6 wt per-
       cent MgCO3).

In several experiments  with milestone  No.
1359, the fuel was natural gas supplied by a
public utility through commercial gas lines.

  Table  1  summarizes operating  conditions
and variables studied in  combustion experi-
ments.  (Experiments  pertaining to  NO

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Table 1. OPERATING CONDITIONS AND VARIABLES STUDIED IN
          FLUIDIZED-BED COMBUSTION EXPERIMENTS
Fuel
Coal




Coal


Gas


Coal


Coal



Coal


Number of
Experiments
9




4


3


4


5



1


Starting
Bed
Material
Alumina




Alumina


Alumina


Alumina


Partially
sulfated
limestone

Partially
sulfated
limestone
Additive
Type
Limestone No. 1359




Limestone No. 1359


Limestone No. 1359


Tymochtee
dolomite

Limestone No. 1 359

Limestone No. 1360
Dolomite No. 1337
Limestone No. 1359


Average
Particle
Size,
ptm
25
600

1400

25
103

25


<44


615

630
540
490


Combustion
Temperature,
°F
1600
1600

1600

1550-1650


1600


1600


1600

1600
1480-1800
1400-1600


Superficial
Gas Velocity,
ft/sec
3
3

9

3


3


3






3


Variables Studied
a) Gas velocity
b) Recycle of elutriated
solids
c) Additive particle
size
a) Combustion temperature
b) Additive particle size
c) Ca/S mole ratio
a) SO2 removal with natural
gas combustion
b) Additive regeneration
a) Additive type
b) Additive particle size
c) Ca/S mole ratio
a) Limestone fluidized
bed
b) Combustion temperature

a) Combustion temperature



-------
 removal,  discussed  in  another paper,  are
 omitted  from  this  table.)  The  variables
 studied  include gas  velocity, recycle  of
 elutriated  solids, combustion  temperature.
 additive  type,  particle  size,  fluidized-bed
 material, and the Ca/S mole  ratio in the feed
 streams. As discussed below, the Ca/S mole
 ratio in the additive and coal feed streams
 appears to have the greatest effect on SO->
 control.
 S02 EMISSION CONTROL WITH
 FINE-PARTICLE ADDITIVE

   The  most thoroughly studied additive for
 SC»2  control  has been limestone No.  1359
 because:

    1.  It has a  minimum  amount of waste
        burden (e.g., its MgCO3 content is 1.3
        wt percent).

    2.  In preliminary  tests1  this stone
        showed  a  high chemical  reactivity
        with SO2-

   Pulverized  limestone No.  1359,  with an
 average particle size  of 25 (im, was studied at
 addition rates equivalent to Ca/S mole ratios
 in the feed up to ^3; Figure 2 summarizes the
 results. The SC<2 content of the flue gas was
 reduced by about 80 percent at a Ca/S mole
 ratio of 2.5.

   Figure 2 includes SC»2 removals for runs in
 which recycled solids  (a mixture of partially
sulfated  limestone  and  fly ash)  were
introduced into the combustor along with the
additive and coal feeds. For runs in which
recycled material was added, the contribution
of unsulfated  calcium in the recycle material
was not included in  the calculated Ca/S ratio.
The  absence  of any improvement  in  SC»2
removal,  when  using fresh  limestone  plus
recycled fly-ash/additive mixture, is puzzling
since  the recycled  solids  contain significant
quantities of  unreacted  CaO.  Also,  as
discussed below,  separate laboratory tests on
II-3-4
the reactivity and capacity of fresh limestone
and  the recycle material detected no signifi-
cant  difference in  reactivity  for  the two
materials.

   The effect of superficial gas velocity (3 and
9  ft/sec)  for limestone No. 1359,  with  an
average particle size of 25  jum. is shown in
Figure 3. At the higher gas velocity, there was
less SO2 reaction with fine additive particles.
The  difference  is not large but does appear to
be significant;  this result is not unexpected
since at the higher gas velocity, elutriation
rates  and,  therefore,  bypassing  of  solid
particles are increased. At  a  superficial gas
velocity  of 3 ft/sec, the residence time for
additive particles, calculated from the weight
of the  fluidized bed and the  additive feed
rates, is about  1 hour; however, experimental
evidence indicates that, even at the low gas
velocity, actual residence time for the bulk of
the additive particles is much less.

  The effect of combustion temperature in
the range  1550 to  1650°F for 25-jum  lime-
stone No. 1359 particles (Figure 4) appears to
be small; the data indicates that SO 2 removal
at 1600°F  is  slightly better than at  either
1550 or 1650°F. Figure 4 also shows that a
90-percent reduction in SO2 emission occurs
at a Ca/S mole ratio
  The overall utilization (conversion of CaO
to  CaSO4)  of 2S-nm  limestone No. 1359
ranged from 25 to 40 percent (Figure 5). This
utilization was calculated from the relation-
ship between SO2 removal and the Ca/S mole
ratio  in the  feed  streams.  As  might  be
expected from the relationship between Ca/S
ratio  and maximum CaO utilization (solid
curve, Figure 5), higher utilization of CaO was
obtained at the lower Ca/S ratios.

  Chemical analyses of various effluent solids
were  performed. Significant differences  in
CaO utilization and  in the extent of calcina-
tion  for the  different solids streams were
noted.  The  results  (Figure 6)  show  that
material collected in  the  primary cyclone,

-------
which had the lowest extent of calcination,
also had the lowest CaO utilization. Material
retained in the fluidized bed, which had the
highest extent of  calcination,  also had the
highest extent of CaO utilization. Secondary
cyclone solids were intermediate with respect
to both extent of calcination and CaO utiliza-
tion. Since most  of the entrained solids in the
flue gas might be  expected to have  similar
residence tunes in the reactor, the difference
in  CaO  utilizations for materials from the
primary and secondary  cyclones  can  be at-
tributed to particle size. Larger particles were
collected  in  the  primary  cyclone;  finer
particles, collected  in the secondary cyclone,
had a higher specific surface area and reacted
more quickly and more thoroughly than the
larger  particles  collected  in  the  primary
cyclone. For additive particles retained in the
bed,  residence  times are  long and a  high
degree of utilization and calcination can be
achieved. The  Ca/S mole  ratios for cyclone
samples  and  fluidized-bed  samples  taken
during a single experiment  with 25-^tm  lime-
stone No. 1359 are shown in Figure 7. In the
fluidized-bed samples, the  Ca/S mole ratios
were  relatively  constant at  1.5,  which
represents a  CaO utilization of  ^70 percent.
CaO  utilization   of  elutriated  material
collected in the secondary cyclone ranged
from 30  to  50 percent; CaO utilization for
material  in  the primary  cyclone  was 20
percent or less. The data in Figures 6 and  7
suggests  that calcination  of  CaCO3  and
utilization of the resulting CaO are dependent
upon the same factors, which include particle
size and residence time in the combustor.
  In an attempt to increase the utilization of
CaO,  elutriated solids were  recycled  in  the
fluidized-bed  combustor in  several experi-
ments.  Material  collected  in  the primary
cyclone,  having  a  CaO utilization  of  20
percent, was recycled. As indicated in Figures
2 and 5, SO2 removal  was not  discernibly
improved by  adding recycled  material.  To
explain these results,  laboratory  tests were
carried out  to compare the rates of sulfation
 of fresh limestone and recycled solids. Figures
 8, 9, and 10 show the results of these tests.

    The reaction rates for fresh limestone and
 for recycled solids  (mixture  of fly ash and
 partially sulfated limestone) were essentially
 the same, both at the start of the experiment
 and  after 2 hours  of reaction (Figure  8).
 These tests were performed at an SO2 partial
 pressure of 3.2 mm  Hg, which is considerably
 higher than SO2 pressures in the combustor
 (<0.8  mm  Hg).   Additional tests were
 performed  (Figure  9)  at  an  SO2  partial
 pressure of 0.4  mm Hg: again, the reaction
 rates for the fresh limestone and the recycled
 solids were  similar,  both initially and  after 2
 hours of reaction. Thus, the apparently lower
 reactivity of the recycled  fly-ash/limestone
 mixture  in  the fluidized-bed  combustor
 cannot  be attributed to a  difference in re-
 action  rate  as a function  of  SO2  partial
 pressure.  The   percentage  of CaO  utilized
 during  the   laboratory  sulfation tests  was
 determined for recycled  solids and fresh lime-
 stone (Figure 10). As might  be expected from
 the similarity of reaction rates, the total per-
 centage of CaO utilized was the same for both
 materials. Thus, the  ineffectiveness of recycle
 operation in reducing SO2 levels cannot be
 accounted for  by a reason  intrinsic  to the
 reaction of recycled material with SO2-

    Electron microprobe studies2 were carried
! out in an attempt to explain the anomalous
 behavior  of  the  recycled  solids  in  the
 fluidized-bed combustor. Cross sections of
 typical  particles, collected in  the  primary
 cyclone during combustion experiments with
 natural gas and with coal, were examined for
 homogeneity of  sulfur and  calcium  content.
 Figures 11 and  12  show the results  of the
 examinations.  Some differences in particle
 composition  noted  are  related to  the  fuel
 used.  With coal as the fuel, sulfur concentra-
 tion  usually  decreases at increasing distance
 from the surface of the particle; on the other
 hand,  sulfur  concentrations  in  particles
 collected  during the combustion of natural
 gas appear to  be more  uniform throughout
                                      11-3-5

-------
the particle.  In gas combustion experiments,
SC>2 was introduced into the coal feed port
from an external source and SC>2 concentra-
tion in the fluidized bed may have been more
uniform than during the combustion of coal.
Three of  the four particles  from  the gas
experiment (Figure 12) had a calcium content
quite close to that expected for CaCC>3;only
one particle had a calcium and sulfur content
expected for CaSC>4.

   The  laboratory tests  have indicated that
high utilization  of the limestone is  possible
and that  the reactivity  of recycled solid  is
essentially  the same  as that of fresh lime-
stones. A lack of further reaction of recycled
solids in the  combustor may be due to insuf-
ficient  gas-solids contacting, which, in turn,
may result from short residence times. When
both  recycled solids and fresh limestone are
fed to the combustor, the conditions in the
combustor may favor greater initial reactivity
of fresh limestone. Perhaps better utilization
of the recycle material could be achieved with
separate combustor beds for recycled additive
and fresh additive feed.
Effect of Additive Material

   The effectiveness of pulverized Tymochtee
dolomite  was compared  with  that of  lime-
stone No.  1359 for  SC>2  control during coal
combustion in a fluidized  bed of alumina.
Figure 13 shows that the difference between
the  additives  is small, but  that Tymochtee
dolomite yields higher 862 reductions at Ca/S
mole ratios smaller  than 2.  It is generally
thought that  the MgO content of calcined
dolomitic  additives  helps keep the particle
pore structure open so that more of the CaO
can react with SC>2- However, the presence of
MgO in dolomites is a serious  disadvantage
because MgO  does  not react  with SO2 and
greatly  increases  the waste  burden  for
dolomitic  additives;  e.g., the  stoichiometric
quantity to react with the sulfur in 1 Ib of 4.5
wt  percent S  coal  is 0.29 Ib Tymochtee
dolomite,  compared  with  only 0.15 Ib  lime-
stone No.  1359.
II-3-6
  The waste burden for any of the additive
materials  can  be considerably  reduced  in
operational schemes which include an additive
regeneration step. In such  a scheme, sulfur
values could be recovered from the spent ad-
ditive,  and  regenerated  additive  can  be
utilized for SO2 emission control so that the
net waste burden of the system would only be
due to fresh makeup additive.
SO2 EMISSION CONTROL WITH
COARSE-PARTICLE ADDITIVE

  Experiments were performed with additive
particles  too coarse  to  elutriate  from the
fluidized  bed;  fluidized  beds of  partially
sulfated additive  of a  composition close to
that expected at  steady-state operation were
used.  Table  2 shows the experimental results
for  limestones  No.   1359  and  1360,  and
dolomite No. 1337. Data for coarse limestone
is compared with results for pulverized lime-
stone  No. 1359 (25 jum)  in Figure 14. The
results for coarse additive particles tend to
cluster in  a  region of smaD area, indicating a
relative independence of the type of additive
material tested and, for limestone No. 1359, a
relative  independence of particle size up to
615  jum.  Although the  coarse  additive
particles  have lower reactivity toward  SO2
than  do  fine  particles,  their  increased
residence time and the resultant large mass of
material  in   contact  with  the  gas stream
resulted in the same degree of SO2 removal as
with  finer (more  reactive)  particles, which
elutriate readily  from  the fluidized bed and
hence have shorter contact times with the gas
stream.  The observed similarity of  SO2
removals for a variety  of  stones and  particle
sizes may be fortuitous or may indicate that
the Ca/S  mole  ratio in the feed streams  is
highly significant.

  The effect on  degree of SO2 removal of
combustion  temperature in the range 1400 to
1800°F was determined for coarse limestone
No. 1359 and dolomite No. 1337 (Figure 15).
The maximum SO2 removal of ^90  percent

-------
     Table 2.  SO2 REMOVALS3 ACHIEVED
WITH COARSE-PARTICLE ADDITIVE MATERIALS
Additive

Material
Limestone No. 1359
Limestone No. 1359
Limestone No. 1359
Limestone No. 1359
Limestone No. 1360
Dolomite No. 1337
Average
Particle
Size,
/urn
615
615
615
490
630
540
Ca/S
Mole
Ratio
2.6
2.3
2.5
2.5
2.3
2.2
S02
Removal,
%
79
83
79
86
74
81
 aAt combustion temperature of 1600°F; with gas ve-
 locity of "v>3 ft/sec.

 occurred in the temperature range 1480 to
 1550°F. In contrast, for pulverized limestone
 No.  1359 (Figure  4), SC>2 removal was little
 affected by variation of combustion tempera-
 ture  in  the  range  1550  to 1650°F. The
 different behaviors of coarse and fine lime-
 stone may be explained by recognizing that
 reaction of the pulverized limestone with SC>2
 depends on  the  extent  of  calcination.
 Although finely powdered limestone  readily
 elutriates from the bed, it has a high specific
 surface area and reacts at high rates so that
 the extent of calcination may be relatively in-
 dependent  of bed temperature in the range
 1550 to  1650°F. Thus the optimum tempera-
 ture for reaction  with SC>2 from pulverized
limestone  is not necessarily the same as that
for coarse  limestone.
CONCLUSIONS

  Control  of  SCH  emission by reaction at
1400  to  1800°F with basic additives (lime-
stone  or  dolomite) containing  CaO has been
studied  in a  6-inch  diameter  fluidized-bed
combustor. In the combustion of a high-sulfur
Illinois   coal,  SOo  emissions have  been
controlled effectively by feeding additive into
the combustor. For  each of  the additives
tested, Ca/S mole ratio in the  feed had the
greatest effect on 862 emission. Emission of
about 80 to 90 percent of the sulfur has been
prevented by feeding additive  and coal  at a
Ca/S mole ratio of 2.0 to 2.5. For fluidized
beds of limestone No. 1359 or dolomite No.
1337, the optimum combustion temperature
for SC<2  removal  appears to  be 1480 to
1550°F. Variables that have minor effects on
SO2  removal  are  additive  type,  additive
particle size, and superficial gas  velocity.
BIBLIOGRAPHY

  l.Borgwardt,  R.  National  Air  Pollution
    Control  Administration.  Private  com-
    munication, 1968.

  2. Natesan,  K. Argonne National Labora-
    tory. Electron Monoprobe Studies.
                                                                                  II-3-7

-------
00
      PREH EATER
                          ADDITIVE
                           FEEDER
                                        I
                         TRANSPORT     M
                            AIRHXr—'
                          COAL
                          FEEDER
                                                                                                           SECONDARY
                                                                                                           CYCLONE
                                                                                                   RECYCLE-
                                                                                                   MATERIALS
                                                                                                   FEEDER
TO
FILTER
AND
VENTILATION
EXHAUST
                                                                                                           TRANSPORT
                                                                                                             AIR
                                                           COMBUSTOR
                           Figure 1.  Equipment for investigating S02  emission in combustion experiments.

-------
CO

(D
UJ
D
g
O
 CM
O
CO
90
80
70
60

50
40
30
20
10
n
I I
1 1 1
8/0
h 7*
@f A TEMPERATURE: 1600°F
~~ 0 COAL FEED: 4.5 wt % S
/ ADDITIVE: LIMESTONE
I
—
—
—

— f FLUIDIZED BED: ALUMINA"
~~ ^ GAS
ft VELOCITY,
— / ft/sec
/ O 2.7
~l • 2.7
/ & 8.6
7 A 8.6
1 1 1
LIMESTONE
NO.
1359,
pm
25
25
25
25
RECYCLE
No
Yes
No
Yes
I I I

—
—
I
              2345

                Ca/S MOLE RATIO

Figure 2   Effect of calcium/sulfur mole
ratio on S02 removal with fine limestone
additive.
s?
C/J
O
cc
 (M
                                 I     I
                   ft/sec
                       I
                 2345

                  Ca/S MOLE RATIO
                                               S?
                                               00
                                               UJ

                                               Of
                                               O
                                               cc
                                               UJ
                                               cc
                                                                    23

                                                                 Ca/S MOLE RATIO
                                               Figure 4. Effect of combustion temperature
                                               on S02 removal with fine limestone additive.
   Figure 3.  Effect of gas velocity on SO2
   removal with fine limestone additive.
                                                                                      II-3-9

-------
o
O
(0
O
100

 90

 80

 70

 60

 50

 40

 30

 20

 10

  0
                 I     I     I    I
GAS
VELOCITY,
ft/sec

( 0 2.7
\ • 2.7
\ £ 8.6
\A 8.6
LIMESTONE
NO.
1359
pm
25
25
25
25

RECYCLE

No
Yes
No
Yes
MAXIMUM CaO UTILIZATION
FOR REMOVAL OF 100% OF_
   )o (no recycle)
             25 \tm ADDITIVE
                 I     I     I    I     I     I
                                               O
                                               $
                          O
                           <0
                          O
100

 90

 80

 70

 60

 50

 40

 30

 20

 10
                                       I    I     I     I     I    I
                                   _• FLUIDIZED-BED SAMPLES
                                    A PRIMARY CYCLONE SAMPLES
                                   .O SECONDARY CYCLONE SAMPLES
                                                                                     8   -
                                                       -   a
                                                                             oo
                   Ca/S MOLE RATIO
                                 30   40    50    60   70   80   90   100

                                        EXTENT OF CALCINATION, %
Figures. Utilization of CaO as a function of   Figures.  Relationship of CaO utilization to
calcium/sulfur mole ratio.                      extent of stone calcination.
HI
_l
Q_
z
o
1
UJ
O
CO
\
ID
U
/.u
6.0
5.0
4.0
3.0
2.0
1.0
0
I I I I II I
/\
/ X ^
a _^_ / <>• PRIMARY
/ ^ A
I ' I
/ Ca/S MOLE RATIO IN FEED
— / ^-* '9 	 SECONDARY
/ .— -• *^^«
/ • 	 •— • •*
/ /n ^ 0 ^ c 0 ruin
v / / W0o
« I I I I I
CYCLONE

CYCLONE —
3IZED BED
-i-
                                             56

                                              TIME, hr
   Figure 7.  Calcium/sulfur mole ratio in fluidized  bed and in elutriated fines during a coal
   combustion experiment, 25-um limestone No. 1359.
II-3-10

-------
I
O
I
       40

       30

       20
10
 8

 6

 4
                                          I
                •  FRESH LIMESTONE

                O  RECYCLED SOLIDS
            O
               I
              I
              20   40
                   60

                 TIME, min
80   100   110
  Figure 8.  Reaction rates at an SC>2 partial
  pressure of 3.2 mm Hg for fresh  limestone
  and recycled solids.
 c
 3


 I
g
I
3.0

2.0



1.0
0.8

0.6

0.4



0.2



0.1
              I     I     I
                            1     T
                  O  FRESH LIMESTONE NO. 1359

                  &  RECYCLED SOLIDS
                  A UTILIZATION ACHIEVED FOR FRESH LIMESTONE,
                     LABORATORY TESTS

                  D UTILIZATION ACHIEVED FOR RECYCLED SOLIDS,
                     LABORATORY TESTS

                  O TOTAL UTILIZATION ACHIEVED FOR RECYCLED
                     SOLIDS IN BOTH FLUIDIZED-BED COMBUSTION
                     EXPERIMENTS AND LABORATORY TESTS
                      100
                                 0.5      1.0       1.5

                                        TIME, hr
                                                 Figure 10. CaO utilization for fresh limestone
                                                 No. 1359 and for recycled solids.
              20    40   60   80   100  120  140

                        TIME, min
 Figure 9.  Reaction rates at an SO2 partial
 pressure of 0.4 mm Hg for fresh limestone and
 recycled solids.
                                                                                     II-3-11

-------
     I             I
*d ; 165 pm
           20          40



     DISTANCE FROM SURFACE, pm
                 60
                        I

                 * d = 140 pm
           20
                       40
                                    60
      DISTANCE FROM SURFACE, pm 	^


(«) d is dimension of particle in the direction of the trace.

Arrow 1 is for Ca in CaCO-j in standard samples.
                                                                   *d - 70 pm
                                                            I
                                                      I
    20           40



DISTANCE FROM SURFACE, pm
                                              *d- 100 pm
                                         20
                                                                        40
                                                                                    1	*•
                                                                 60
                                                                                    60
                                     DISTANCE FROM SURFACE, pm 	»~



                                   Arrow 2 is for Ca in CaS04 in standard samples.

                                   Arrow 3 is for S in CaS04 in standard samples.
  Figure 11.  Microprobe traces showing calcium and sulfur levels in typical elutriated par-

  ticles collected in the primary cyclone during combustion of coal.
JI-3-12

-------
                     : 65
0          20          40           60

       DISTANCE FROM SURFACE, \>m
0          20          40

     DISTANCE FROM SURFACE, pn
                                    T
60
                        I            I            T
                              *d : 100 pm
                                                                        Ca
                                                                        I
                        1
                                                               *d = 50
                                                      Ca
                                                ^/VV^V/WA^^^/N^V
      20          40

DISTANCE FROM SURFACE,
                                                                                   3	*-
                                                I
            0          20          40          60

                   DISTANCE FROM SURFACE, pm
                               I
                                                                                   60
(«)  d is dimension of the particle in the direction of the trace.  Arrow 2 is for Ca in CaSO4 in standard samples.
Arrow l is for Ca in CaCOg in standard samples.             Arrow 3 is for S in CaSO^ in standard samples.


  Figure 12.  Microprobe traces showing calcium and sulfur levels in typical elutriated par-
  ticles  collected in the  primary cyclone during  combustion of natural gas.

                                                                                      II-3-13

-------
    100
     90 -
     80 —
*    70
CO
O
£    60
     50
     40
cc
 CM
     30
     20
     10
                    t   LIMESTONE NO. 1359,
                    '     25 gm
                                               52
                                                    100
                                                    90
                                                     80
                                                     70
                                                     60
E    50
1
i    40
DC
gT   30


     20


     10
                                                                   I           I
                                                          POINTS:  COARSE ADDITIVE AND
                                                                  ADDITIVE FLUIDIZED BED
                                                           ' CURVE:  FINE LIMESTONE NO. 1359
                                                                    (25 Hm> AND ALUMINA
                                                                    FLUIDIZED BED
                                                                   I
                                                                               I
                    1            2

                  Ca/S MOLE RATIO
  Figure 13.  Control of S02 emission with
  Tymochtee dolomite No. 1359 limestone.
    100

     90

     80

     70
3?
i    eo
>
§    50
oc
g"   40-

     30

     20

     10
     1300
                                                       0123
                                                                  Ca/S MOLE RATIO

                                                Figure 14.  SC>2 removal  with coarse additive
                                                and inert fluidized beds  and with fine additive
                                                and additive fluidized beds.
                                                                    I               I
                                                                    A  LIMESTONE NO. 1359
                                                                       AVERAGE SIZE 490 \m
                                                                       Ca/S - 2.5

                                                                    O  DOLOMITE NO. 1337
                                                                       AVERAGE SIZE 540 gm
                                                                       Ca/S - 2.2
           6-in. DIA BENCH-SCALE UNIT
           GAS VELOCITY:  ~3 ft/sec
           BED HEIGHT: ~2 ft FLUIDIZED
           3-4 vol % 02 IN OFF GAS
                      1
                     1400
                                     1500            1600
                                          TEMPERATURE, "F
                   1700
                                                                                  1800
          Figure 15.  Effect of fluidized-bed combustion temperature on 802 removal.
II-3-14

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

                      G. HAMMONS AND  A. SKOPP

                      Esso Research and Engineering
                            Linden, New Jersey
ABSTRACT

  The desulfurization efficiency attained on a
3-inch  ID  fluidized-bed  coal  combustor
operating with a limestone bed is reported as
a function of the independent variables of the
system. A kinetic model was derived for use
in  design  studies;  the reactor length-to-
diameter  ratio  (L/D)  was  suggested as a
scaling  parameter in this model. High L/D
ratios were found  to  give  relatively poor
•desulfurization  efficiency.  It  was  demon-
strated  experimentally that the stone main-
tains a  relatively high capacity with repeated
cycles of combustion and reductive regenera-
tion. NOX  emissions  decreased  as a com-
bustion run progressed and the average SO2
concentration in the batch reactor increased.
This indicated a possible interaction between
SOX and NOX which resulted in lowered NOX
emissions. Small scale experiments  using a
simulated flue gas showed that SO 2 and NO
did react to an appreciable extent  in  the
absence  of  oxygen. A  different mechanism
appeared  to be operative in  the presence  of
oxygen  but  a  decrease in  NOX was still
observed.
INTRODUCTION

  Esso Research and Engineering Company is
conducting an experimental program for the
National Air Pollution Administration (under
NAPCA contract CPA 70-19) on a system in
which finely  ground coal is combusted in a
fluidized bed of limestone. The Esso study is
a  part of  NAPCA's overall program to
examine  fluidized-bed  combustion  as a
possible new boiler (power regeneration)
technique.  The air pollution control potential
of a fluidized-bed combustor is great because
of  the good  gas-solid  contacting  and,
consequently, by appropriate  choice of bed
material, the high SO2  removal efficiency
which can  be attained. Limestone appears to
be one of the most promising materials for
capture of SO2- The  sulfur  in the coal is
oxidized to SO2, which is then captured by
the lime as CaSO4- The system proposed by
Esso  involves  transferring the  partially
sulfated stone  from the combustor  to a
separate regeneration  vessel  in which  the
following reaction occurs:
 CaSO4
CO

H2
CaO + SO2 +
C02

H2O
(1)
The  regenerated stone (CaO) can then be
returned to the combustor for further use.
This  method of operation naturally reduces
the fresh limestone requirements substantial-
ly. The off gas from the regeneration unit has
a high SO2 concentration and can be utilized
as a feed to a byproduct sulfur or sulfuric acid
pknt. Figure 1 is a schematic diagram of this
process.
  Finely  ground  coal Ov200  fji)  is being
utilized  in  the  experimental  program  to
                                     H-4-1

-------
facilitate separation of the coal ash from the
limestone bed material by entrainment of the
coal ash from the bed.

   Esso's fluidized-bed  combustion  program
has four primary objectives:

    1.  Investigate reactor  combustion effi-
       ciency  when  feeding finely  ground
       coal.

    2.  Investigate  flue  gas  desulfurization
       attainable  as  a function of  the  in-
       dependent variables of the system.

    3.  Determine  the  potential  for high
       temperature reductive regeneration of
       the sulfated bed material.

    4.  Determine the level of NOX emissions
       from  the  Esso  fluidized-bed  com-
       bustor  and   investigate  possible
       catalytic reduction of NOX by lime-
       stone.

This presentation is concerned primarily with
the air pollution control aspects of the pro-
gram; i.e., the last three objectives.
EXPERIMENTAL EQUIPMENT

  Two  fluidized-bed units—a combustor and
a regenerator—are being utilized in the Esso
experimental study. Operation  in these units
during  regeneration studies is on  an inter-
mittent basis: a batch of limestone is charged
to the combustor; coal is continuously fed to
the combustor until the desired level of lime-
stone sulfation is achieved; the combustor is
cooled; and the partially sulfated material
from  the  combustor  is  transferred  to the
regenerator. Following  regeneration,  the
solids are  returned to the combustor, and
another cycle is begun.

  Figure  2 is a flow diagram of the Esso
fluidized-bed combustor (FBC). The reactor is
a 3-inch ID incoloy tube. Four continuous
flue  gas  analyzers are used: NDIR SC>2 and
CO analyzers, and polarographic NOX and ©2
analyzers.

  Figure 3 is a  flow diagram of the fluidized-
bed regeneration unit. The reactor is a 2-inch
ID alumina-ceramic tube. An NDIR analyzer
is used for continuous measurement of SO2 in
the off gas.
EXPERIMENTAL RESULTS

Range of Variables Examined

  Table  1  shows  the range  of  variables
examined in the Esso program. Only one coal
and one  limestone have been utilized in this
study.


  Table 1. RANGE OF VARIABLES EXAMINED
       Variable
                Combustor

Bed temperature
Settled bed height
Superficial gas velocity
Excess air
Average stone particle diameter
Average coal particle diameter
Coal
Limestone
                 Regenerator
 Bed temperature
 CO
 C02/C0
 Superficial gas velocity
    Range
1500-1800°F
4-16 in.
2-4 f ps
3-50%
460-930/1
200 At
Bit. (A) (3% S)
N-1359
2000°F
10mol?
2
2fps
Reactor Combustion Efficiency

  Table  2  shows  the effect  of the  unit
operating conditions upon the reactor com-
bustion efficiency. Perhaps the most  unusual
effect was on bed height. Increasing  the bed
height might be expected  to cause a longer
solid  residence time  and, consequently, a
higher combustion  efficiency. The opposite
II-4-2

-------
 effect was observed in this study; i.e., as bed
 height  was  increased, combustion efficiency
 decreased. This effect is  believed to be  the
 result of the slugging nature of the small Esso
 FBC.  This  slugging  apparently  causes  a
 reduced  coal  particle residence  time as bed
 height  is increased.  Hence,  the  increased
 velocity  of the  coal particles  through  the
 bubble phase  in  deep beds apparently  more
 than offsets the effect of the additional bed
 height.
Table 2. EFFECTOR INDEPENDENT VARIABLES
        ON COMBUSTION EFFICIENCY
        Variable
Increasing this variable
Bed temperature
(1500-1800° F)
Settled bed height
(4-16 in.)
Superficial gas
   velocity
(2-4 fps)

Excess air
(3-50%)

Average stone particle
   diameter
(460-930 ju)
  Increases combustion
     efficiency
     (91-97%)

  Decreases combustion
     efficiency
     (94-91%)

  Decreases combustion
     efficiency
     (97-91%)

  No effect above 10%
     excess air

  No effect
Desulfurization

  The  desulfurization  reaction  was  inves-
tigated over the range of conditions previous-
ly shown in Table 1. Figure 4 shows typical
batch desulfurization  data  obtained. The
capacity  of the  stone at  20  percent SC>2
breakthrough has been chosen  as a relative
measure  of the desulfurization efficiency at a
given set of conditions. Table  3 shows the
effect of  the reactor  operating conditions
upon the desulfurization  efficiency of the
Esso FBC. The effect of bed height upon the
                        desulfurization  efficiency again  apparently
                        reflects the slugging nature of the small Esso
                        fluid  bed. Increasing  the excess  air  level
                        decreased  the stone  capacity.  One  possible
                        explanation which has been  advanced1  for a
                        similar effect is as follows: at high excess air
                        levels  (high  C>2  content  in  the  bed), the
                        following reaction sequence can occur:
                                                      CaO + SO2
                                                 CaSO3
                                                      CaSO3 + 1/2  O2
                                         (2)
                                                                 (3)
At high C>2 levels, reaction (3) occurs quickly
and  apparently  causes a  pore blockage, re-
stricting the accessible CaO. However, at low
excess air levels, reaction (3) does not proceed
as quickly so that less pore blockage  by the
CaSO4 is experienced.


Table 3. EFFECT OF INDEPENDENT VARIABLES
            ON STONE CAPACITY
     Variable
Increasing this variable
Excess air (3-30%)
Decreases stone capacity
  (23-12%)
                        Average stone particle    Decreases stone capacity
                           diameter (460-930/1)     (32-18%)

                        Superficial gas velocity   No effect
                           (2-4 fps)
                        Bed temperature
                           (1500-1800°F)

                        Settled bed height
                           (4-16 in.)
                      Not effective above 1600 F
                      Decreases stone capacity
                        (20-7%)
                        aCapacity at 20% SO2 breakthrough.


                        Kinetics of Desulfurization

                          A kinetic model derived in a previous Esso
                        study2  has  been  utilized  to  correlate  the

                                                              II-4-3

-------
desulfurization results from the present work.
Figure  5 gives  the  model assumptions and
development.  Apparent rate constant k(X) is
a function of CaO utilization X.

   Figure 6 shows the experimental kinetic
data from the Esso FBC program. Constant
k(X)  is  observed to be  a function of bed
height,  as well  as CaO utilization. This bed
height effect  is  believed  to  be a result of
poorer   gas-solid  contacting  (because  of
slugging) as bed height is increased.

   A  comparison was made  between the
kinetic  data obtained in this study  and data
obtained in a  previous Esso study2 in which a
simulated flue gas was desulfurized.  Figure 7
shows that the two sets of kinetic data agree
closely. The implication from this comparison
is that a large fraction of the SC>2 generated
during coal combustion in the Esso FBC is
released near the reactor inlet.

   In  tins study, the data  shown in  Figure 6
has been replotted at  constant values  of the
L/D (aspect) ratio. We believe that by main-
taining   a constant  L/D  ratio in  scale-up,
similar  contacting   efficiencies  can  be
expected.  Figures  8 and 9  show  that de-
creasing the  L/D ratio at a constant set of
other  operating conditions  increases the
percentage of available  CaO which can be
utilized  in achieving a  given  SO2  removal.
Hence, it appears that  high values of the L/D
ratio  will give conservative estimates  of the
performance   of large  scale  equipment,  hi
which the L/D ratio is  less than unity. For
design   purposes, the  desulfurization data
shown in Figure 8 at  an L/D of 3.7 should
provide a conservative estimate of commercial
reactor performance.

Stone Regeneration

  The   partially  sulfated  stone  was re-
generated by  reduction  at 2000° F. A 2/1
volumetric ratio of CO2/CO was maintained
in the reducing gas to minimize sulfide forma-
tion.  The reducing gas  contained  10 mol
II-4-4
percent CO. Equation  (1) is the primary re-
action occurring in the regenerator.


   Figure 10 shows the SO2 concentration in
the  regenerator  off gas.  The dashed line
indicates the SO2 concentration that would
be  attained  at thermodynamic equilibrium.
The occasional low values of SO2 concentra-
tion in the off gas are believed to represent
poor  gas-solid contacting  in the regenerator
rather than  any kinetic  limitations.  The
system can  be expected  to  attain  thermo-
dynamic equilibrium in a commercial system.
  Cycling  of the  bed  material  between
combustion  and regeneration gave the stone
capacity data shown in Figure 11. Initially the
capacity of  the regenerated stone was higher
than  that  of  the fresh stone.  This  phe-
nomenon  is possibly  a result  of a  more
favorably crystal lattice, as compared to the
fresh stone,  formed on eliminating 803 from
the sulfated material. Sorbent activity slowly
decreased on  repeated  regeneration. Also
shown in  Figure  11 is cyclic stone capacity
data obtained in a previous Esso study2 using
a simulated flue gas for SO2 sorption.
NOX Emissions

  Figure 12 is a typical record of NOX emis-
sions from the Esso FBC. As a run progressed,
and the  SO2  level in  the  bed and flue gas
increased, the NOX emissions decreased. The
indication is that there  is some interaction
between  the SOX and NOX which causes ,the
NOX emissions to decrease.


  In order to investigate what type of inter-
action  might be occurring between NOX and
SO2,  a  series of  experiments  (using  a
simulated gas with varying concentrations of
NO and  SO2)  was conducted in the 2-inch
regeneration reactor.  Table  4 shows  the
results of this exploratory series of  experi-
ments.

-------
              Table 4. NOX-SO2 INTERACTION IN A FLUIDIZED BED3 OF LIMESTONE
Bed Material

Gas phase
Partially sulfated
stone from combustor
Alundum

Gas phase
Partially sulfated
stone from combustor
Alundum
NOX Cone., ppm
Before S02
Intro.
After S02
Intro.
Transport Gas - N2
900
J840
J830
820
900
(440
\180
820
Transport Gas - Air
800

710
620
760

710
380
S02 Cone., ppm
Inlet

1290
( 785
\1510
1000

1290

785
1090
Outlet

1290
(300
(480
1000

1180

0
1090
           aBed temperature-1630°F.
  The indication is that the NO and S(>2 can
react in a partially sulfated bed in the absence
of 62; no reaction occurs over an inert bed in
the absence of  O2- One  possible reaction
mechanism which can be occurring is:
   CaO + SO2 —

   CaSOs + 2NO

CaS03-(NO)2 —
   CaSO3          (4)

   —• CaSO3 • (NO)2 (5)
-CaSO4+N2+l/2O2(6)
However, in the presence of O2, reaction oc-
curs between NOX and SO2 even over an inert
bed. Although no reaction mechanism has
been identified,  the  oxidized  species, NO2
and 803, could possibly be involved.
  Further studies will be made of the exact
nature of  SOX    NOX interaction in  the
presence  of 62- Methods of utilizing  this
interaction  to effect low NOX emissions from
a commercial reactor will also be investigated.

BIBLIOGRAPHY

  l.Coutant, R.W. et al. Investigation of the
    Reactivity of Limestone and Dolomite for
    Capturing SO2 from Flue Gas. Summary
    Report  to NAPCA  under Contract PH
    86-67-115, August 1968.

 2. Skopp,  A. et al. Fluid Bed Studies of the
    Limestone Based Flue Gas Desulfurization
    Process.  Final  Report under NAPCA
    Contract PH 86-67-130, August 1969.
                                                                               II-4-5

-------
FLUE GAS
& COAL FLY ASH
t
FLUID-BED
COMBUSTOR
CaO + SO2 + 1/ 02
— -3» CaS04
t I *
LSULFATED
SORBENT


HIGH S02 GAS

FLUID


| | REGENERATED
FRESH FLUIDIZING SORBENT
SORBENT AIR
+ COAL

-^
t
REDUCING
GAS
PLANT
-BED REGENERATOR
DISCARDED
*" SORBENT
CaS04+CO — » SO2-I
C02 + CaO
            Figure 1-.  Regenerative limestone system proposed by Esso Research.
 AIR
                                                     REFRIGER-
EACTOR  FILTER
      I         AIR
     /    CON DENS
        SCALES
NDIR SO2
ANALYZER

NDIR CO
ANALYZER


»/v\rr
                                                                 POLAROGRAPHIC
                                                                 NOX ANALYZER
                                                                 POLAROGRAPHIC
                                                                  02 ANALYZER
                                                                   INTERMITTENT
                                                                   GAS SAMPLER
                     Figure 2-  Esso fluidized-bed combustion unit.
11-4-6

-------
                                    ANALYZER VENT
              COMBUSTION
                ZONE

           THERMO-
           COUPLE
ALUMINA
REACTOR
              lso:
                 STEAM-
                          COMBUSTION AND
                            FILTER OVEN
                        BY PASS
                         LINE
                     AIR  /       I   WATER CONDENSER
                   CONDENSER     H" AND KNOCKOUT
                                                   I
                                                   ?H,O
5555
_»   N2_f   C02_f   AIR—*
                                                            NDIR SO-,
                                                           ANALYZER
                                                             AND
                                                           RECORDER
                Figure 3. Flow diagram for the fluidized-bed regeneration unit.
   2500
   2000
 £ 1500
 Q.
 CO
 O
 8
 i 1000

 CM
 8


    500
            I      I      I

             S02 EMISSIONS IN THE

             ABSENCE OF REMOVAL
            1
                      MATERIAL BALANCE:        U
                        S02 REMOVED FROM GAS : S02 ABSORBED BY
                        SORBENT
                      YIELDS:  In (CH/CO) = -k(X) . Ho/U
                      MODEL ASSUMPTIONS:
                         COMPLETE MIXING OF SOLIDS
                         PLUG FLOW OF GAS
                         ALL S02'RELEASED AT REACTOR INLET
                  2345
                RUN TIME, hours
Figure 4.  Typical  batch desulfurization data.   Figure 5.  Esso kinetic model development.
                                                                       II-4-7

-------
   1.8
   1.5
   1.2
 = 0.9
5. 0.6
.*
   0.3
   0.0
    5.0
                                                   4.0
                                                T 3.0
                                                  c
                                                 o
                                                   2-0
                                                    1.0
                                                     0
                                                                D SIMULATED FLUE GAS DATA3
                                                                O PRESENT EXPERIMENTAL DATA
                                                                  WITH COAL COMBUSTION
      0
                                           30
                                                      10
                   15           20
                 CaO UTILIZATION, %
25
                10           20
               X, CaO UTILIZATION, %
      T; 1600 "F
      EXCESS AIR -10%
      (dp) COAL - 200 \i
      (dp) STONE : 930 p                                  "         aData from reference 2.
Figure 6.  Kinetic constants from Esso data.  Figure 7. S02 formation occurs primarily near
                                               reactor inlet.
                                                       T-1600"F
                                                       U = 3-4 FPS
                                                       H0 = 4 in.
                   (dp) COAL ; 200 p
                   EXCESS AIR ~10%
                   (dp) STONE : 930 |
                                                 O
                                                     i.o
                                                    0.9
                                                    0.8
                                                    0.7
                                                     0.6
                                                     0.5
                                                         "T—|~T
                                                           X:9
     0.0  0.1  0.2  0.3 0.4 0.5 0.6  0.7 0.8 0.9 1.0
                    u
                    ao, seconds
      L/D i 3.7      u |(dp) STONE : 930 p
      EXCESS AIR ~10%l(dp) COAL : 200 M
   Figure 8.  Kinetic results at L/D = 3.7-
        0 0.1 0.2 0.3 0.40.5  0.6 0.7 0.8 0.91.0
                     LJ
                     —o   seconds
        L/D: 10.3    U |(dp) STONE: 930 p
        EXCESS AIR ~10% I (dp) COAL = 200 p
Figure 9. High L/D gives lowstone utilizations.
II-4-8

-------
    10
O
en
LJJ
QC
X
<
                            i    r
           EQUILIBRIUM
                            U : 2 FPS
                            CO  10 MOL
                            C02/C0 : 2
                            T r
                        I     I    I
     1
                                                 0.5
                                              D Z
                                               CM
                                              O
                                    gs  0.4
                                              O
                                              flj
                                              O
         4567
         CYCLE NUMBER
                                     8
                                                  0.2
                                                  0.1
                                                          SORPTION 1600 °F
                                                          REGENERATION 2000°F
                                                          U- 2 FPS
                                                            500 H N-1359       -
                                                            SIMULATED FLUE GAS£
                                                          930 fi N-1359
                                                          PRESENT STUDY   	
                                                           VITH COAL COMBUSTION
                                                            
-------
        5.  COAL-BASED SULFUR RECOVERY CYCLE

                                                                           IN
                    FLUIDIZED-LIME-BED COMBUSTION

                    G. P. CURRAN, C. E. FINK, AND
                             EVERETT GORIN

                       Consolidation  Coal Company
ABSTRACT

  Sulfated dolomite  was produced  by the
combustion of coal  in  a fluidized  bed  of
dolomite continuously fed with raw dolomite,
A brief experimental study is reported here of
a two-stage process for regeneration of the
€3804; i.e., its conversion to CaO, by use of
coal as a  reductant. The composition of the
off gases  from this process is controlled so
that just  enough CO + H2  is available  to
reduce the SC>2 to elemental sulfur.
  The results of a computer study on
management  of  the  off  gas  to  obtain
maximum yield of sulfur are  also presented.
INTRODUCTION

  A previous paper1 from these laboratories
presented data on the efficiency of removal of
sulfur in the combustion of coal within  a
fluidized bed of dolomite. The sulfur is fixed
on  the  dolomite in  the form of calcium
sulfate. Regeneration of the calcium sulfate to
reform calcium oxide also was investigated to
obtain data on  acceptor life and dolomite
makeup requirements.

  The regeneration in the previous study was
conducted by the use of partial combustion
with air of CO  as a fuel gas.  The essential
regeneration reaction is endothermic and may
be written,
CaSO4 + CO = CaO + CO2 + SO2        (1)
        A H= +53,720 cal/g mol at 25°C.
Hence, the  need  for partial combustion  to
supply heat.

  Although the procedure was successful as a
laboratory expedient, it would be much more
practical in a commercial situation to use coal
as a reductant and  fuel; i.e.,
CaSO4 + C
                                   (2)
               AH= +27,520 cal/g mol
  Attempts to use char from high-sulfur
Pittsburgh seam coals for this purpose were
unsuccessful.  It  was  found  that under
reducing conditions and at the high tempera-
tures (>1900°F) required to effect  reaction
(2),  ash  fusion  was an insurmountable
problem.


  Accordingly, to avoid this problem a two-
stage regeneration process was devised, using
coal as fuel. This paper describes the process
and presents preliminary experimental data
which demonstrates its feasibility.

  This paper also  discusses  the  thermo-
dynamics  of recovery  of elemental sulfur
from the regeneration off gases, and outlines
                                     II-5-1

-------
the  net savings  in dolomite  and  energy
requirements by use of the process.
PROCESS DESCRIPTION

  Figure 1 is a schematic sketch of the overall
process. In the first stage, CaSO4 is reduced
to CaS at a relatively low temperature (about
1850°F). In principle, CaSO4 can be reduced
to CaS by a series of reactions which can be
expressed as the overall reaction,
     CaSC«4 + 2 C = CaS + 2 CO2
              AH =+42,440 cal/gmol
               (3)
In practice, however, additional carbon must
be burned  to produce  CO. The  overall  re-
action then can be expressed as,
                  X(2-A)-4(1-A)
                                  02 (4)
       = CaS + X CO2 +
 AX
(1-A)
CO
where: I^A  = mol ratio of C/CaSO4 required
              for process heat balance, and
        A  = CO/ (CO +  CO2>, mol ratio in
              exit gas to satisfy the reaction
              kinetics.

  Another feature of the first-stage process is
that it is so conducted that sufficient reducing
gas  is produced to reduce the SO2 formed in
the  second stage to sulfur by reactions such
as,
                = 1/2S2 + 2CO2       (5)

            AH = -48,900 cal/g mol

Thus, AX/(1-A) = 2 and reaction (4) becomes
 CaSO4 + 3.890 C + 0.890 O2

II-5-2
              (4a)
       = CaS + 1.890CO2 + 2CO

                 AH = O

The  high  CO/CO2  ratio in  reaction (4a)
ensures that the CaSO4 reduction will occur
rapidly. With the use of coal, combustion of
its  hydrogen  content  produces  additional
heat, which is needed to preheat the incoming
coal and air. With respect to the coal-air re-
action, the  first stage  is  a  fluidized-gas
producer in which ash slagging will not occur
at nominal bed  temperatures of  1850°F
because the combustion  takes place in the
presence of heat sink  created by the en-
dothermic reduction  of CaSO4 to CaS - re-
action (3).


  In the second stage,  the acceptor, now in
the form of CaS, is  contacted with air in a
fluidized bed at temperatures in the range of
1900-1950° F. No external fuel is used. The
sulfur  is rejected by a series of  reactions
which  may be  expressed  as the overall re-
action,


    CaS + 3/2 O2 = CaO + SO2          (6)

     AH = -108,970 cal/g mol
                      A portion  of the
                      oxidized to CaSO4 v*
                                                                 incoming  CaS  first  is
                                                  CaS + 2 O2 = CaSO4
                                                            (7)
                Then the CaSO4 reacts with the residual CaS
                via,


                     3 CaSO4 + CaS = 4 CaO + 4 SO2    (8)


                to give CaO and SO2- At a nominal pressure
                of 1  atm, the incoming air must be diluted
                with recycled tail gas from the sulfur recovery
                section in order  to provide  a AP SO2 driving
               -force for reaction (8). A mol ratio of about

-------
1 /1  recycle-gas/air  is  required.  The  exo-
thermic reaction provides  the  preheat  duty
for the incoming air and recycled gas.

  The combined off gases from both stages
are treated in  the  sulfur recovery section
where  the  SC>2 produced in the second  stage
is reduced  to elemental sulfur by CO (and H2
formed  by  the  water-gas  shift  reaction)
produced in the first stage.  Since both CaSC>4
and CaS are present in the first stage, some
SC>2 inevitably will be formed by reaction (8).
In the presence of CO and H2, this SO2 will
be reduced to H2S, COS, and 82 with the last
material predominating. This situation is not
detrimental, since all sulfur compounds are
nearly  completely  converted  to  elemental
sulfur in the recovery section.

  The purpose  of the work reported here was
to study briefly  the  salient features of the
two-stage  regeneration  process  and  the
severity of the possible processing problems
outlined below.
First Stage

    1.  At the  very  low air/fuel  ratios
       required, the incoming raw coal may
       form coke.
    2.  To demonstrate  freedom  from  ash
       slagging at reducing conditions using
       Eastern steam coal.
    3.  To determine  the operating tempera-
       ture  and  coal/CaSO4/air  ratios
       required to generate sufficient  CO
       to satisfy reaction (4a).
    4.  To determine  whether deposits form.


Second Stage

    1.  To determine the  O2/CaS  ratio re-
       quired for optimum rejection of sulfur
       as SO2, as a function of temperature
       and APg02 driving forces.
    2. Previous  experience  with  the  CO2
       acceptor  process,  in  which  CaS  is
       converted to CaO and SO2 in the re-
       generator under conditions  similar to
       those of the second stage, showed that
       when CaS and CaSO4 exist simulta-
       neously  in  the acceptor,  a transient
       liquid  formed  which  led  to massive
       deposits and/or cementing together of
       the acceptor particles.  In this work, it
       was vital to determine  the nature and
       extent of deposit formation.
EXPERIMENTAL

   Generally, the equipment and experimental
procedure were the same as reported previous-
ly.1  Figure 2 is a schematic diagram of the
equipment.

   Briefly, the reactor had a  4-inch I.D. and
was  made from Type 316 stainless steel. The
fluidized-bed height was  controlled  at 36
inches by an overflow  weir. The original re-
actor  was  modified by  replacing the per-
forated disc, baffle,  and  plenum chamber at
the  bottom with a cone having an included
angle of 40 degrees. All solids were fed to the
apex of the cone. The same reactor was used
for  all steps of the process,  including coal
preoxidation; i.e.,  no attempt was made to
operate all the steps simultaneously.

   After reaching  the programmed flows and
bed temperature, solids feeding was continued
nominally  for three bed inventory changes
and  the data was recorded  over  a  1-hour
balance period.  By  adjusting the  electrical
power input to each of the three sections of
the reactor heater, the maximum temperature
gradient across  the fluidized bed was held to
within 20° F.

   The  data workup  was straightforward,
being based on measured  values for the input
                                     II-5-3

-------
and  output streams,  the dry  exit gas rates,
product gas analyses, and the acceptor com-
positions as determined by  a special assay for
the CaS and CaSC>4 content. In the first-stage
runs, considerable sulfur was rejected to the
gas through  reaction (8). At  the conditions
used, most of the SO2 was reduced to $2
which condensed  as fog in the gas recovery
system and escaped with the product gas. The
amount  of  elemental sulfur formed  was
obtained by a forced sulfur balance.

   The acceptor  was  the  same Tymochtee
dolomite (Western  Ohio), sized  to  16 x 28
mesh, used in the previous work.
RESULTS AND DISCUSSION

Preoxidation of Coal Feed

  All work reported here was carried out
with Ireland Mine coal, a  high-sulfur, highly
caking Pittsburgh seam coal.  Early  work on
feeding  raw coal  to stage  1 showed that this
was  not practical due to a small amount of
coke  formation.  Accordingly, it was found
necessary to  preoxidize  the feed  coal to
reduce its caking propensity.

  A level of preoxidation of 5.3 wt percent
(defined as pounds of oxygen reacted with
100  pounds MF coal) was  found sufficient to
prevent   coke  formation.  The preoxidation
was  carried  out  by continuous  feed of raw
coal  (28 x 100 mesh)  to  a fluidized bed at
700°F. A mixture of air and nitrogen was used
as fluidizing gas. The composition of the pre-
oxidized coal is:
  Hydrogen

  Carbon
  Nitrogen
  Oxygen (by diff.)
  Sulfur
  Ash
II-5-4
 4.52, wt percent, MF
basis
72.44
 1.29
 6.71
 4.34
10.70
                         Gross  Btu,  MF basis, by  Dulong  formula,
                         12,995 Btu/lb.
Stage 1 Runs

  After preliminary runs were made to deter-
mine  the  extent of preoxidation required, a
series of runs were made with the above pre-
oxidized coal. Five runs were made with fresh
sulfated acceptor at 1825 and 1875°F and at
various  input  acceptor,  coal and  air rates.
Another run, to help  determine any kinetics
effects of  acceptor activity loss on the rate of
CaSO4 reduction, was made with an acceptor
exposed previously  to  seven  combustion-
regeneration  cycles.   Run  conditions  and
results are shown in Table 1.

  The objective of these runs was to deter-
mine  conditions under which three simulta-
neous  conditions could be satisfied: nearly
complete reduction to CaS, a ratio  of CO +
H2/CaS = 2 in the product, and process heat
balance.  For  each run,  a  complete heat
balance was calculated with the results shown
in the last row of Table  1. The heat balance
was based on air and preoxidized coal fed to
the process at 100 and 700°F, respectively.

  Comparisons  of   Runs A7-A7A  and
A10-A10A  show  that,  by  increasing   the
temperature, the  desired ratio,  (CO -*• H^)!
(total S in acceptor product), can be achieved
easily. Substantially  complete reduction of
CaSO4 occurred in all  the fresh acceptor runs,
showing that the reduction reactions are rapid
at temperatures of 1825°F and above, even
with (CO + H2) concentrations as low as  9.9
percent in the  exit gas,  as in Run  A8.  The
acceptor retention times in these runs ranged
from 0.8 to 1.1 hours. In run A9, made with
deactivated  acceptor,  the  somewhat lower
level  of CaSO4  reduction may have been
caused by the  decreased  retention  time
brought about by the  increased acceptor feed
rate  which  was used in order to keep  the
CaSO4 input roughly comparable with that in
the other  runs. Thus,  deactivation has  no

-------
                  Table 1. CONDITIONS AND RESULTS FOR FIRST-STAGE RUNS3
Run Number
A7
Temperature, ° F 1 825
b
Acceptor Feed i
Feed rate, Ib/hr 9.44
Comp., mol & of total Ca
CaO 10.50
CaS 0
CaS04
89.50
Coal Feed Rate, Ib/hr 4.79
Inlet Gas, SCFH
Air
N2 purges
Exit Gas Rate, SCFH Dry Gas
Exit Gas Comp. at Top of Bed
H20, mol %
H2
CH4
CO
C02
S2
H2S
COS
N2 (diff.)
Product Acceptor
Rate, Ib/hr
Comp., mol % of total Ca
CaO
CaS
CaS04
Unreacted Char, Ib/hr
Outlet Fluidizing Velocity, ft/sec
Bed Density, Ib/ft3
Bed Weight, Ib
Solids Retention Time, hr
.. _ „ Lb C Gasified
% Carbon Burnout 100LbCFed
% S in Acceptor Rejected to Gas
CaS

„ , 0 . CO + H2
Mnl RTfift
IVHJI naiiu, _, _ f\^r*f\
C3o * CaSOA
Process Heat Balance,
Btu Absorbed/lb Mol CaS04 Fed

152
3.6
231

12.72
3.88
0.48
9.79
18.81
1.03
0.29
0.03
52.97

6.27

29.8
69.6
0.6
1.99
2.19
24.6
6.15
0.98
60.3
21.5
99.2
2.20

-67,700
A7A
1875
b

9.44

10.50
0
89.50
4.79

152
3.6
235

12.06
4.65
0.20
12.30
17.55
1.05
0.34
0.03
51.82

6.25

31.3
68.1
0.6
1.88
2.27
23.9
5.98
0.96
63.3
23.3
99.1
2.84

-48,400
A8
1825
b

11.26

10.50
0
89.50
4.79

152
3.6
232

13.80
2.71
0.46
7.21
21.51
0.96
0.25
0.03
53.07

7.54

25.3
73.8
0.9
1.99
2.20
24.5
6.13
0.81
60.7
16.5
98.8
1.26

-27,400
A10
1825
b

11.01

10.50
0
89.50
4.93

125
3.6
207

15.07
3.98
0.55
9.29
20.61
1.06
0.26
0.02
49.16

7.35

27.0
72.2
0.8
2.31
1.96
27.0
6.75
0.92
54.8
18.4
99.0
1.57

-4840
A10A
1875
b

9.44

10.50
0
89.50
4.93

125
3.6
203

13.76
5.88
0.52
13.02
16.74
1.18
0.19
0.03
48.68

6.27

29.8
69.6
0.6
2.33
1.96
27.0
6.75
1.08
53.5
21.6
99.2
2.67

+7800
A9
1825
c

23.72

65.50
0
34.50
4.22

139
3.6
218

13.61
2.20
0.43
6.58
23.86
0.81
0.10
0.02
52.39

20.08

69.4
28.6
2.0
1.52
2.06
30.7
7.68
0.38
68.1
10.6
93.5
0.94

-19,200
aFuel was 5.3% preoxidized Ireland Mine coal, at an 8-psig system pressure.
 Fresh  sulfated dolomite.
cSulfated dolomite after 7 cycles of regeneration.
                                                                                          11-5-5

-------
 obvious effect on the rate  of CaSC>4 reduc-
 tion.

   No deposits of any kind were found, nor
 was there evidence of ash slagging in any of
 the runs.

   In  none  of  the  runs was process  heat
 balance  achieved simultaneously  with  the
 desired ratio, (CO + H2>/(total sulfur) = 2.
 Four  of the six runs  were strongly  exo-
 thermic. Table  2 shows calculated process
                             conditions  which  lead  to thermoneutrality
                             and a (CO + H2)/(total sulfur) ratio of 2. For
                             the process calculations, a carbon burnout of
                             70 percent (versus 60.3 percent in Run A7)
                             was assumed, to allow the effects of a deeper
                             fluidized bed. The process calculations show
                             that the  following CaSO4/coal/air ratios will
                             be required:

                                 Coal/CaSO4 = 72 Ib coal/mol CaSO4fed.

                                 Air/Coal   = 27 SCF/lb coal.
                    Table 2. CALCULATED PROCESS CONDITIONS FOR HEAT
                                  BALANCE IN STAGE 1a
aBasis:  1 Ib mol CaS04.
bHeat formation at 25° C.
 (CaS04> in. - (CaO + CaS +


Preoxidized Coal
Sulfated Acceptor
Inert
CaO
CaSO4
Air
Heat of Reaction15
Total

MAP Char
Ash
Product Acceptor
Inert
CaO
CaS
CaS04
Product Gas
H2O
C02
CO
N2
H2
S2
Heat Losses
Total
Temp,°F
Lbs
Input
700
1800
1800
1800
100
72.45
56.66
6.56
136.14
145.4

Output
1850
1850
1850
1850
1850
1850
1850
1850
1850
1850
1850
1850
15.74
7.74
56.66
18.68
56.13
0.83
21.84
83.71
32.38
111.0
0.83
9.68
Mols

-

0.1173
1.000
5.032


—
—

0.3331
0.7781
0.0061
1.212
1.902
1.156
3.962
0.412
0.151
A'H, Btu

17,600
27,210
2,440
61,070
1,400
120,570
230,290

10,850
3,520
27,990
7,170
18,040
390
43,440
40,480
15,220
52,900
5,080
2,330
2,880
230,290
                                            + CO2 + CO + H20) out.
II-5-6

-------
The breakdown of the total oxygen required
at process conditions is:
         Air
         Coal
         CaSC>4
33   percent
 4   percent
63   percent
  Small  deposits caused by  the CaSO4-CaS
transient liquid occurred in all the runs. Their
location, appearance, and composition were
the same as the deposits which were formed
in the CC>2 acceptor process regenerator.3  In
Run B2  the extent of deposit formation was
0.6 percent of the acceptor fed.
The  air requirement  corresponds to 22 per-
cent of stoichiometric air for combustion of
the 5 percent preoxidized Ireland Mine coal.
Stage 2 Runs

  The feedstock used for these runs was, for
experimental  convenience,  prepared by
reducing the CaSC>4 to CaS, using CO both as
a fuel and reductant at  1750°F. Part of the
CO was burned with air to provide the pre-
heat duty for the incoming gas and solids. In
the second stage, a portion of the incoming
CaS is- oxidized via reation (7). The CaSO4
reacts with residual  CaS via reaction (8),
thereby rejecting  sulfur.  The overall reaction
is highly  exothermic and provides sufficient
preheat duty for the incoming air and diluent
recycled tail gas.  During the program, about
twice as much diluent  (N2> was  used as is
required for process heat balance in order to
provide a conservative AP SO 2 driving force
with respect to reaction (8). Run  conditions
and results are shown in Table 3.

  In Run B2 the input ah- rate was varied to
determine the ©2 requirement which gives the
optimum  sulfur  rejection at 1950°F. The
optimum  conditions appear to be about 95
percent of the theoretical ©2 needed to reject
all  the  sulfur. At higher  ©2  inputs,  the
product acceptor contains CaSO4_

  It is not  clear whether the lower level of
sulfur rejection in run  Bl at  1900°F was
caused by the lower reaction rates for re-
action (8) and/or (7) or by inhibition due to
the lower AP SO2  driving force. More data is
needed on the kinetics of the second-stage re-
actions.
                         After  repeated  combustion-regeneration
                       cycles, it  is likely,  as  in the  CO2  acceptor
                       process, that the extent of deposit formation
                       will decrease drastically, and that  the deposit
                       problem can  be tolerated by shutting down
                       periodically to descale the reactor walls.
                       Sulfur Recovery

                         Sulfur is recovered by so blending the re-
                       generation off gases from the two stages that
                       the  ratio of (CO + H2 + H2S + COS)/(SO2) is
                       equal to or slightly greater than 2. The gas is
                       first passed  through a catalytic reductor to
                       effect reaction (5) and the corresponding re-
                       action,
                                   SO2= 1/2S2 + 2H2O.
                                       (9)
                       Small amounts of "trimming" air may be
                       added to the feed gas so that the product gas
                       from  the  reductor step continues a proper
                       feed  to  a multistage Claus plant.  The equi-
                       libriums in reactions (5) and (9)  are so fa-
                       vorable that substantially no CO or H2 remain
                       in  the product gas. The Claus feed gas  now
                       contains a ratio of (H2S + COS)/(SO2) = 2.


                         Figure 3, a  schematic flow diagram of the
                       sulfur recovery section, also shows a typical
                       feed gas composition.


                         A  thermodynamic  analysis of the  sulfur
                       recovery  section was made with  the  basic
                       premises given below.
                                                           11-5-7

-------
               Table 3. CONDITIONS AND RESULTS FOR SECOND-STAGE RUNS
Run Number
System Pressure, psig
Bed Temperature, °F
Acceptor, Mol % of Total Ca
CaS
CaSO4
CaO
Feed rate, Ib/hr
Inlet Air, SCFH
Inlet N2, SCFH
N2 Purges, SCFH
Exit Gas Rate, SCFH
Exit Gas Composition, Mol %
N2
S02
Outlet Fluidizing Vel., ft/sec
% Sulfur Rejected
Mol % Total Ca in Product
CaO
CaS
CaS04
Input O2, % of Theoretical
Outlet S02 Partial Press., atm
AP Driving Force, atm
Bed Density, Ib/ft3
Solids Retention Time, hr
B1
5
1900
73.3
2.4
24.3
4.64
70.5
147.0
3.6
215.0
95.76
4.24
2.42
82.1
86.5
7.3
6.2
92.2
0.0560
0.0680
21.2
1.22
Equilibrium is assumed to be established in
all stages with respect to the following re-
actions:
CO+1/2S2 = COS (10)
H2+1/2S2 = H2S (11)
2H2 + SO2=1/2S2 + 2H2O (9)
B2(l) B2(ll) B2(lll)
555
1950 1950 1950
72.3 72.3 72.3
1.8 1.8 1.8
25.9 25.9 25.9
4.95 „ 4.95 4.95
64.4 76.4 85.5
130.6 127.0 129.0
3.6 3.6 3.6
194.0 202.0 208.0
95.21 94.63 95.52
4.79 5.37 4.48
2.23 2.33 2.39
79.4 93.8 80.1
84.8 94.8 85.2
14.7 4.1 2.0
0.5 1.1 12.8
79.4 94.4 102.8
0.0632 0.0709 0,0592
0.0978 0.0901 0.1018
22.7 22.4 21.7
1.25 1.21 1.15
2CO + SO2=1/2S2 + 2CO2 (5)
3Sg = 4S6 (12)
S8 = 4S2 (13)
The equilibrium relationships in the above si>
reactions, at a given temperature and pressure
along with the four elemental balances
completely define the system in each stage.
II-5-8

-------
   Experimental data4 obtained in our labora-
tories  and  by others5  shows  that, at least
under laboratory conditions, it is very easy to
establish equilibrium  in  all of the above re-
actions  even  at  very  low  temperatures,
provided   an  active  alumina  catalyst  is
employed.

   It is necessary, in order to maintain catalyst
activity, to operate above  the  dew point of
the  sulfur  to prevent  its deposition on  the
catalyst. Each stage was so operated, there-
fore, that  the vapor pressure of sulfur at the
outlet  temperature was  20  percent greater
than its partial  pressure in the outlet gas.
Sulfur is  condensed  from  the product  gas
from each  stage before feeding  the  gas to the
succeeding stage.

   A computer program  was set up using an
iterative trial-and-error procedure to solve for
sulfur recovery and product gas compositions
with  the  above  restraints  imposed  on  the
system.

   The free energy data for reactions (5), (9),
(10), and  (11) were taken  from  the  Janaf
tables.6  Experimental  data7  was  used  to
define the  equilibrium constants for reactions
(12) and (13).

   The outlet temperatures from each stage,
for the illustrative example given, are shown
in  Figure  3.  The  total sulfur  recovery
potential is 98.2 percent.  Addition of a fourth
stage,  operated   at  270°F, increases  the'
potential sulfur recovery to  99.3 percent.
Equilibrium acceptor
  activity
Second stage
Fluidized boiler
Sulfur recovery section,
   1 st reactor
=  0.41, equivalent
     to 8 percent
     makeup rate
=  1950°F
=  1800°F

=   900°F
Table  4. COMPARISON OF RELATIVE ENERGY
AND RAW MATERIAL REQUIREMENTS IN FLUI-
              DIZED BOILERS

                          With Re-
                          generation
                          and Sulfur   Once-
   Requirement             Recovery   Through

Energy as Coal Equivalent3

   Coal fed to regeneration      9.16       —
   Sensible heat for fresh
     dolomite               0.39      1.92
   Heat to calcine fresh
     dolomite               0.61      3.01

Less

   Sensible heat regenerated
     acceptor
   Residual char from re-
     generation burned in
     boiler
   Steam and boiler feed
     water duty from sul-
     fur recovery             	     	
Net                        5.53      4.93
   0.21       —
   2.22       —
   2.20       —
Energy and Raw Material Requirements
  The relative energy  and raw  material
requirements  were  evaluated  for the  two
types  of  fluidized-boiler  operations using
once-through dolomite and acceptor regenera-
tion with  sulfur recovery,  respectively.  The
evaluation  is based on the  following process
conditions:
Sulfur in coal            =  4.5 percent
Removal of SC>2 from
  stack gas             =  90 percent
Cost - iflon Coal Burned

   I ncremental coal at
     $6/ton                 +33       +30
   Fresh dolomite at
     $2.50/ton              +28      +138
   Worth of recovered sul-
     fur at $25/ton           -85        —

   Total cost                -24      +168
   Relative cost                0      +192
alb/100 Ib coal burned in boiler.
                                                                                      II-5-9

-------
   Table 4 compares the two systems. The net
difference in material costs between the once-
through  and regeneration  processes, about
$1.90 per ton coal burned in the boiler, is a
powerful incentive to continue the develop-
ment of the two-stage regeneration process.

BIBLIOGRAPHY

  1. Zielke, C.W., H.E. Lebowitz, R.T. Struck,
    and E. Gorin. Jour. Air Pollution Control
    Assoc. 20: 164-169, March 1970.

  2. Skopp,  A.A. Proceedings  of  First Inter-
    national  Conference  on   Fluidized-Bed
    Combustion, Hueston Woods State Park,
    Oxford, Ohio. November 18-22, 1968.
3. Curran, G.P., C.E. Fink, and E. Gorin. U.
  S. Dept.  of Interior, OCR  R&D Report
  No.  16, Interim Report No.  3 on CO2
  Acceptor Process.

4. Consolidation  Coal  Laboratories.  Un-
  published Data.
5. Pasternak,  R.  Brennstoff-Chemie,  50:
  200, 1969.

6. Dow Chemical Co., Janaf Tables. Clearing-
  house for Federal Scientific and Technical
  Information, U. S. Dept. of Commerce.

7. Preuner, G. and W. Schupp. Zeit. Physik.
  Chem.  68: 129, 1909.
II-5-10

-------
                                                    STEAM
1800°F
Ca SO4
FROM BOILE
FLUIDIZEC
BED
COAL
700 <>F
AIR
100 °F
CaO
TO BOILE
"I
t


1st
STAGE
REACTOR
~1850°F

t
1

r ./"A .

• HO
ASH
CaS
I
2nd
STAGE
REACTOR
~1950°F
I {^

co2
so2
100 op
V/^X
H2O
SULFUR
RECOVERY
ELEMENTAL
SULFUR
TO BOILER
,
,
t A Am C02 RECYCLE BLOWER \J)
I *AIR N 2±S
R 100°F
       Figure 1. Flow diagram of two-stage regeneration of CaSO^ to CaO:
         COAL  ACCEPTOR
  FUEL
  FEED
HOPPERS
  (2)
SOLIDS
FEEDERS
D-®
 CARRIER
   GAS
           JAL  AUUtriUK    I.
           I      i        i
           •S    AC5E,1TOR^-----T
                     FEED
                    HOPPERS
                      (2)
                                                           MOISTURE
                                                          CONDENSER!
                                                            •-COOLING WATER
     DUST
   RECEIVER

REACTOR
                   FLUIDIZING
                      GAS  "^
      CONDENSATE
     RECEIVER WITH Err:
      SIGHT GLASS

 ACCEPTOR FINES
UNBURNED CARBON
   & FUEL ASH
                             •COOLING WATER
                                                      WATER
                           SIGHT
                           GLASS
                                                                 BPCV  T0 GAS
                                                                      METERING
                                                                      & SAMPLING
                     ACCEPTOR |
                      RECEIVER
                                  USED
                                  ACCEPTOR
                    Figure 2.  Simplified flow diagram of equipment.
                                                                                11-5-11

-------
        "TRIMMING" AIR
FEED GAS

Gas
co2
CO
N2
H2
H20

so2

2
COS
H2S

Vol
15.
6.
64.
2.
6.

4.



%
30
37
44 ]
27
68

01







0.022
0.204
,


STAGE
1
900 °F




(REDUCTOR)
I
' '













STAGE
3
330 °F








                                                                 STAGE
                                                                   2
                                                                 450 °F
                                                                                      n
              TOTAL S RECOVERY
                    98.2%
     TAIL GAS
TO FLUIDIZED BOILER
                                                    Gas
                                                    H2S

                                                    SO,
             Vol. %
             0.068

             0.034
11-5-12
             Figure 3. Sulfur recovery from off gases - - two-stage regeneration of

-------
                                          6.  THE  FLUIDISED-BED
                                  DESULPHURISING GASIFIER

                                GERRY MOSS

                      Esso Petrolleum Company, Ltd.
                                    England
ABSTRACT

  Explanations  are offered of the optimum
sulphur-absorption  temperatures which  are
observed  when  fuel  oil  is  combusted  in
fluidised beds of lime particles, under both
oxidising and reducing conditions. In  the
gasifying case, information is given concerning
the effect  on  desulphurising efficiency  of
variations in stoichiometric ratio, bed replace-
ment rate, and  the mean particle size of the
bed material.

  Consideration  is given to  some of  the
control  problems  which  arise  when
continuous  gasification is  attempted;  pre-
ferred solutions are described.
INTRODUCTION

  In a  paper  presented at the  First  Inter-
national Conference on Fluidised-Bed  Com-
bustion  it was shown that sulphur-containing
oils may be desulphurised during  combustion
within a fluidised bed of lime particles  under
both oxidising and reducing conditions.  In the
first case the sulphur is fixed as CaSC«4;  in the
second case it is fixed as CaS. It was reported
that in  both cases  the  optimum absorption
temperature  was in  the region of 850°C. It
had also been found that regeneration of the
reacted  bed  material was  possible in  both
cases and that  high concentrations of SO 2
could be  obtained  at temperatures in  the
region of 1050°C, by reducing CaSO4 and by
oxidising CaS. Several possible applications of
this principle were described and some experi-
mental results were presented relating largely
to gasifying conditions.

  This paper reviews the progress which has
been made in this field since then at the Esso
Research Centre at Abingdon.
OPTIMUM TEMPERATURE FOR SULPHUR
RETENTION

  Thermodynamic  considerations  indicate
that sulphur should be fixed efficiently by
lime under both oxidising and reducing condi-
tions at  temperatures up  to 1100°C. Re-
generation is accounted for by the decomposi-
tion of calcium sulphite, which is produced
(in  one case)  by the  reduction of calcium
sulphate, and (in the other case) by the oxida-
tion of calcium sulphide. Figure 1  shows the
effect  of  temperature  on  the equilibrium
partial pressure of SO2, for the reaction of
oxygen with calcium sulphide; in other words,
the  SO 2  equilibrium  concentration  with
CaSO3- It is quite easy to  obtain  this  curve
experimentally  and  the  important  point
which  emerges is that very high  SO2 con-
centrations  may be obtained at temperatures
above 1000°C.

  Even  at 865°C,  the equilibrium  SO2
concentration is as high as would be obtained
by burning a 4 percent by wt sulphur fuel oil
with 5 percent  excess air. It follows from this
that, under oxidising conditions at tempera-
tures above about 850°C, sulphur  cannot be
acquired by lime via  calcium sulphite but
must be fixed as  calcium sulphate by direct
                                       II-6-1

-------
reaction  of  calcium oxide  with  sulphur
trioxide.

   Figure  2  shows the pronounced optimum
temperature which is obtained when sulphur
is fixed in  a fluidised bed  of lime under
oxidising conditions as calcium sulphate. It is
important to note that this curve relates to
lime in  the size range 100-200 microns which
was  31  percent reacted to calcium sulphate.
When  a  fresh bed   was  used,  the initial
absorption efficiency was in  the region of
90-95 percent  over  the  whole  temperature
range of the experiment. It follows from this
that  we are dealing.with a question of kinetics
rather than  thermodynamics; indications are
that  the reaction is ash-diffusion limited. Even
so some  further  explanation is required to
account for the optimum temperature.

   As we have already seen, although there are
two  possible  routes to  the  formation of
calcium sulphate, one of them  (via  calcium
sulphite) may be ruled out on thermodynamic
grounds at  temperatures  much higher than
850°C.   This leaves the direct  reaction
between 803 and  lime  to  be  considered.
Figure 3 shows the equilibrium partial pres-
sure  of 863 in the combustion products of 4
percent  S fuel oil  burned with 5  percent
excess air,  as  a  function of  temperature.
Although  the  formation  of 863 is favoured
by low temperatures, the rate of formation at
low temperature can be extremely slow in the
absence of a catalyst. It follows from this that
the  rate  of formation of 863 and, con-
sequently, its concentration could well reach
a  maximum level in the region of 850°C.
Under ash-diffusion controlled conditions this
could result in the optimum SC>2 absorption
temperature which has been observed.

   Turning now to gasifying conditions, there
is in this case a very much simpler explanation
of the optimum  absorption temperature. It
will be appreciated that even under gasifying
conditions,  the  air  which enters the  bed
contains its normal  quota of oxygen. Con-
sequently  regenerating conditions can occur
II-6-2
immediately adjacent  to  the  distributor. As
we have seen regeneration is favoured by high
temperatures. Therefore  it appears  that the
fall off in  the  absorption efficiency of the
gasifier at high  temperatures is due to an in-
creased  tendency to regeneration  at the dis-
tributor, with reabsorption of SCH higher in
the bed resulting in an internal refluxing of
SO-; within  the fluidised  bed. Table 1  shows
the concentrations of SCb in parts per million
which may be observed in gas samples drawn
from within the bed at various heights above
the distributor.  The first set of readings relate
to a temperature  of 850°C; the second, to a
temperature of  950°C. It is clear  from these
results that the  sulphur content of the lime in
a bed  sample   can, under some  conditions,
vary according to the position within the bed
from which the  sample is taken.
Table 1.  CONCENTRATION OF SO2 AT STATED
     HEIGHTS ABOVE DISTRIBUTOR8 AT
     30 PERCENT STOICHIOMETRIC AIR
Height, in.
1/2
4
8
At 850° C, ppm
20
10
10
At 950° C, ppm
1000
50
20
 Stone sulphur content 6 percent by weight.
INFLUENCES OF OPERATIONAL VARIA-
BLES ON GASIFIER PERFORMANCE

Stoichiometric Ratio
                                    i
  At the first Conference, information w/as
presented  concerning  the  effect  of
Stoichiometric ratio  on  the desulphurising
efficiency of  a  gasifier. At that  time the
gasifier was being operated with some degree
of pre-combustion; about 25-35 percent of
the incoming  air was  being bumed before
reaching the distributor.   This  procedure
results in a bed of uniform sulphur content at

-------
 temperatures  in the  region  of 850UC and
 enables the sulphur balance to be checked by
 stone  analysis. Since then the apparatus has
 been operated without underfiring  and it has
 been found that this method of operation is
 markedly superior.

   Figure 4  shows  the effect of changing
 stoichiometric ratios  on  desulphurising ef-
 ficiency, with and without underfiring, with 4
 percent  by  wt of sulphur in  the 800-micron
 mean-particle-size bed material. It can be seen
 that, in  the absence of underfiring, the desul
 phurising efficiency is maintained  even at a
 stoichiometric ratio of 20 percent; whereas,
 when the bed is underfired, there is a marked
 drop in  desulphurising efficiency at stoichio-
 metric ratios lower than 30 percent. Further-
 more the results which were obtained without
 underfiring  were obtained  at a superficial gas
 velocity  of 4  ft/sec, as against  the 3  ft/sec
 used with  underfiring. This enhanced per-
 formance naturally has a considerable impact
 on the size of the gasifier and therefore on the
 investment required for a given fuel through-
 put. The reason for the poor performance  in
 the  presence  of underfiring  is not entirely
 clear, but the  particles get very  sooty under
 these conditions at low stoichiometric ratios
 whereas, in the absence of underfiring, they
 remain clean.
Bed Material Replacement Rate

   A very  important factor bearing on the
running costs of the desulphurising gasifier is
the rate at which the bed material has to be
replaced in  order  to maintain desulphurising
efficiency. In order to obtain information on
this subject, we  ran  several series of cyclic
tests each at a different bed replacement rate,
both  with  and   without  underfiring.  The
results of just one of these test  series are
shown in  Figure 5. In  this  case  the  bed
replacement rate was 0.5 wt of lime per unit
weight of sulphur; i.e.,  about a third  of the
stoichiometric  requirement. It  can be seen
that under these conditions the desulphurising
efficiency rapidly bottomed out at about 60
percent.  The effect  of varying  bed replace-
ment rates on desulphurising efficiency can be
seen in Figure 6. Although these results were
obtained using underfiring,  the  efficiencies
measured without underfiring fall on the same
line. It is clear from this curve that virtually
100-percent  desulphurising  efficiency  is
obtainable with a bed replacement rate which
meets the stoichiometric requirement of the
sulphur;  i.e., about  1.75 wts of lime per unit
weight of  sulphur.  This result relates to a
maximum  stone  sulphur  content  of
approximately 4 percent by weight.
Particle Size

  The effect of particle size on desulphurising
efficiency has also been  examined. No dis-
cernible difference in performance was found
when the same stone was compared at mean
particle  sizes  of  810  and  635  microns.
Indications are that, due to the highly porous
nature of the lime particles, the internal sur-
face contributes most of the reactive surface
area; consequently, there is little to be gained
by making the mean particle size smaller than
the  diameter  at  which bed containment
becomes a problem.

  An interesting  feature of the system is its
ability to fix vanadium as  well  as sulphur.
Vanadium may be fixed with virtually 100-
percent efficiency under both oxidising and
reducing conditions.
DESIGN OF A CONTINUOUSLY
OPERATING GASIFIER

  With the feasibility of the desulphurising
gasifier established, consideration was given to
the  design problems  posed by  a practical
system; an experimental unit  is  currently
being constructed which will deliver 7 MM
Btu  of desulphurised fuel  per  hour on  a
continuous basis.
                                     II-6-3

-------
   Figure 7 is a schematic flow diagram for a
continuous unit. On the left-hand side is the
gasifying bed in which lime is converted into
calcium sulphide; on the right-hand side is the
regenerator  in which  calcium  sulphide  is
oxidised by  air to lime and SO2- bed material
is  circulated  between  the  two units.  If the
gasifier is to function  efficiently, it is neces-
sary  to maintain  the temperature  of the
gasifying bed in the region of 850°C and the
temperature of the regenerating bed in the
region of 1050°C.  Should the temperature  of
the regenerating  bed  fall  below  800°C,  re-
carbonation will occur, lowering the tempera-
ture  of the  regenerator.  If  the gasifying
temperature  rises   above  900°C,  the  desul-
phurising  efficiency will be  reduced. In the
case of the regenerator, too low a temperature
will  result  in  the formation  of calcium
sulphate and a buildup of sulphur in the bed
material; too high a temperature may tend  to
deactivate the lime.

   A method of controlling the temperatures
of  the two  beds  has been chosen  which
involves no  heat losses and  no heat-transfer
surfaces. In the case of the gasifying bed, it is
proposed  to  control  the  temperature  by
recycling  some of the flue gas formed  by
complete combustion back to the inlet of the
air blower.  This necessitates a  separate  air
supply for the regenerator. In the case of the
regenerator,  the temperature  will  be
controlled by varying  the circulation rate  of
the bed material. The rate of heat release  in
the regenerator is  determined by  the rate  at
which air is supplied:  it will normally match
the rate at which  sulphur is acquired by the
gasifier. The effect  of increasing the flow rate
of bed  material will be to reduce its sulphur
content  and, hence,  the amount  of heat
generated within  the regenerator per unit
weight of lime. Since  the lime enters the re-
generator  at  a temperature of 850°C  and is
heated  to a temperature  of 1050°C,  there
must  clearly be a specific bed transfer rate for
any particular set  of operating conditions  at
which the regenerator  temperature is in equi-
librium at the desired level. It follows from
this that an accurate and reliable method for
controlling bed  transfer rates is  an essential
requirement for successful operation.

  The  method  which  has  been  chosen is
gravitational  dense phase transfer.  Each bed
container  is provided  with  a  catchment
pocket  slightly above the surface of the bed;
each  of these pockets  communicates via an
almost vertical duct with the side wall of its
neighbor at a point slightly above the level of
the distributor.  There is a short horizontal
section  at the bottom of each duct so that, in
the absence of any external stimulus, the two
transfer  ducts simply  fill  with static bed
material. An arrangement  of this type  is
shown in Figure 8 which is a schematic layout
of the first bed transfer test rig. Here you can
see the two transfer ducts and their horizontal
delivering  sections.  Each  horizontal  con-
nection has a suitable perforated tube running
parallel  to its axis. When gas is introduced via
one  of  these tubes,  the  particles in the
horizontal  duct are fluidised  and expelled;
they are replaced by gravitational flow down
the vertical duct. As soon as the gas ceases to
flow,  so do the solids. By pulsing  the gas flow
and varying the  frequency of the pulses, it is
possible  to closely  control the rate at which
bed material is transferred. Very little gas is
required to operate the system. In the original
test rig it was possible to shift about 2 pounds
of bed material per cubic foot of gas. In view
of the  possible  variation of sulphur content
throughout the depth of the bed, the transfer
of material from the top of each bed to the
bottom of  the other is advantageous from an
operational point of view. The dense packing
in each duct, together with the continuous gas
bleed which is required to prevent blockage of
the  perforations  in the  activating  rtubes,
ensures   that  the   leakage  of gas  between
compartments is minimal.

  The experimental gasifier has been designed
as a monolithic structure in castable  refrac-
tory which is insulated, from the walls of the
mild steel casing which contains it,  by a layer
of vermiculite fondu. The overall dimensions
II-6-4

-------
of the unit are about 4 feet square by 8 feet
high.  In  order  to  check the operation of the
bed transfer system and the cyclones, the full-
scale  cold  rig (shown in Figure  9)  has been
built in mild steel  and perspex. The large tank
represents  the  gasifying compartment:  the
central cylinder flanked by the two cyclones
represents the regenerator. The perspex ducts
allow observations to be made of the flow of
bed material which can of course  be measured
by  weight. The bed material which is  being
used in these  tests is crushed brick with the
same size distribution and bulk density as the
lime. The results  which have been  obtained
have been used as a basis for the experimental
gasifier which  is  now  being constructed  as
part of a contract with NAPCA for develop-
ment  of  the  chemically  active   fluid-bed
gasifier.
                                                                                     II-6-5

-------
 CM
o
0.5



0.4



0.3



0.2



0.1
                  EQUILIBRIUM SO,
          MAXIMUM S02 LEVEL
        ATTAINABLE USING AIR

                                                    70
                                                3?

                                                Q
                                                LU
                                                CD
                                                DC
                                                oc
                                                D

                                                a.
                                                _i

                                                CO
                                                    60
                                                    50
      900
                                   1100
                    1000
                   TEMPERATURE, °C
Figure 1.  Effect of temperature on the

equiliDnum 862 partial  pressures for the re-

action of oxygen with calcium sulphide.
 X
 CO

 5
 CO
LU
OC

a>
c/>
LU
DC
a.
                                                     800
                                                                                      950
           850          900
           TEMPERATURE, °C

Figure 2.  Optimum temperature for

sulphur oxide absorption by lime (31%
of bed reacted).
                                                oc
                                                O
                                                DC
                                                D

                                                OL
                                                _i

                                                CO
                                                    100
                                                    80
                                                    60
                                                 40
                                                     20
                  5% BY WT S IN BED  —


            • WITHOUT UNDERFIRING

               (Gas velocity 4 ft/sec)


            O WITH UNDERFIRING    —

               (Gas velocity 3 ft/sec)

               I       I       I
      600       800      1000      1200

                  TEMPERATURE, °C
                                          1400
                                                  15      20      25      30      35

                                                            STOICHIOMETRIC AIR, %
                                                                                           40
Figure 3.  Equilibrium partial pressure of S03
in combustion products of 4% S fuel oil
burned  with 5% excess air.
                                                 Figure 4.  Stoichiometric air rate with and
                                                 without underfiring.
o
z
LU
LL
U.
LU
•^
O
t
or
m

oc
I
i
100
80
60
40


20

0
*U 1 1 1
-V.^
•
—


_

1 1 1
4 P 12
I I !
•• •
• •»
—


—

I I I
16 20 24
<*
EMOVED,
oc
oc
i
D
CO
_l
LU
2

                                                    100
                                                     80
                                                     60
                                                     40
                                                     20
                                                                   TEST CONDITIONS

                                                            JLPHUR LOADING  0.03 Ib/lb BED/hr

                                                           AIR RATE    33%OFSTOICHIOMFTRIC
                                                           GAS VELOCITY           2.5 ft/sec
                                                           FEED PART. SIZE        850 - 1200'u_
                                                           BED TEMP             800 - 900 °C
                                                               I
                                                                             I
                                                                                 I
                     RUN NUMBER

 Figure 5.  Bed ageing effect replacement rate

 0.5 wt CaO/wt S.
                                                                     1.0             2.0

                                                                WT CaO PER WT SULPHUR

                                                         Figure 6. Effect of makeup rate on

                                                         desulphurising efficiency.
II-6-6

-------
SULPHUR-FREE SULPHUR DIOXIDE
FUEL GAS RICH STREAM
t !



FUEL >*•

Ca S
t
Ca O (LI ML)
Ca S

SOLIDS
TRANSFER
ra n

Ca S
!
Ua u
A ,>
AIR 	 taJ 	 ».
                         GASIFIER
REGENERATOR
                       Figure 7.  Chemically active fluid-bed gasifier.
CYCLONE OUTLET
                                                                             CYCLONE OUTLET
 P : PULSED-AIR FEEDER
 V : VIBRATORY FEEDER
                                         Al R SUPPLY-
                          Figure 8.  Test rig for bed-transfer trials.
                                                                                      11-6-7

-------
SESSION HI:




        Gasification to Desulfurize Coal
SESSION CHAIRMAN :




        Mr. J. W. Eckerd, USBM, Morgantown

-------
1. PRODUCTION OF  LOW-SULFUR BOILER FUEL
                           BY TWO-STAGE COMBUSTION —
        APPLICATION OF CO2 ACCEPTOR PROCESS

                    G. P. CURRAN, C. E. FINK,  AND
                                  E. GORIN
                       Consolidation Coal Company
ABSTRACT

  A modification of the CC>2 acceptor proc-
ess  is presented as a method for the produc-
tion of low-sulfur boiler fuels; i.e., low-sulfur
char and/or low-sulfur producer gas. Partic-
ular emphasis  is given to total gasification of
Eastern coals for the production of low-sulfur
producer gas, which could be utilized for the
production of clean power in new combined-
cycle plants. Experimental data is given on
acceptor  life  under projected operating
conditions
INTRODUCTION

  Considerable interest is developing in the
production  of clean  power from  coal by
means of a combined-cycle power plant. The
first  step necessary in such a plant is the,
generation of high-pressure  producer gas by
the gasification of coal. Recently1  construc-
tion of a 170,000 KW plant of this type in
Germany was  announced  by  Steinkohlen-
Electrizitat AG. The plant will adapt the well
known Lurgi pressure  gasification process to
production  of high-pressure producer  gas.
Surprisingly,  although  no process is provided
for desulfurization of  the gas, the incentive
for construction of the plant is stated1 to be
its lower investment cost in  comparison with
that  of  a conventional power plant. If the
economics can be confirmed, the combined-
cycle  plant  opens up the prospect of  pro-
ducing clean power from coal at a possible
lower cost in new plants than in conventional
plants  equipped  with stack gas  scrubbing
equipment.

  The most critical factor in the development
of such a new power cycle is the availability
of a satisfactory process for the production of
low-sulfur, high-pressure  producer gas.  The
Lurgi process, which must include facilities
for desulfurization of the gas, is the only such
process commercially available today. Eastern
bituminous coals, however, for the most part,
offer difficulties as feedstocks to the Lurgi
process because of coking problems. Further-
more, it  would  be desirable to  purify  and
desulfurize the gas while it is hot, as part of
the gasification process, in order  to improve
thermal efficiency of the cycle. High thermal
efficiency in the combined-cycle plant  can
only  be obtained if the gas can be cleaned and
desulfurized hot.

  The CC>2 acceptor process is now under-
going intensive development2 as a means of
producing high-Btu pipeline gas from coal. It
is  the purpose of this paper  to show  that
many of the features of this process and the
experience gained in  its development can be
utilized to speed the development of an anal-
ogous process for the production of clean,
high-pressure producer gas.

  A  contract, with  the same title as  this
paper, has been signed with NAPCA to accel-
erate research on this application of the CC«2
acceptor  process. The work reported here
                                     III-l-l

-------
concerns both some precontract work carried
out by Consolidation Coal Research and some
initial results of work carried out under the
contract.
DESCRIPTION OF THE PROCESS

  The proposed  process is in reality a two-
stage, air-blown, fluidized gas producer. The
simple air-blown, pressurized, fluidized gas
producer  itself is not  a  developed process;
many problems have been associated with its
development. These potentially can be over-
come by use of the present process.  The
advantages of the present process are:
  1. In situ desulfurization of the generated
     gases.
  2. Improved  efficiency  of  carbon  utili-
     zation.
  3. Improved operability  by control of the
     ash-fusion problem.
  The first advantage  stems from the well
known sulfur-acceptor properties of the lime
acceptor.  The second and third advantages
arise because of the two-stage nature of the
process. The final combustion of  carbon  is
conducted,  separate  from  the  gasification
reaction, in the presence of the endothermic
lime calcining reaction.

  Figure  1  shows one version of the process
presently  being analyzed as part of the feasi-
bility study under the NAPCA contract. The
coal is  preoxidized  in a  separate  air-blown
fluidized vessel operated at 750-800°F.

  A  computer  program,  consistent  with
thermodynamic  restraints imposed on the
process, gives the heat and material balance of
the system  illustrated in Figure  1. The pro-
gram is broad enough to encompass variations
in the process; e.g., producing low-sulfur char
as a co-product with the low-sulfur producer
gas.  The char-to-gas ratio  can be varied over
wide limits.

  Table  1 is a condensed  version of the heat
and  material balance  for one-spot condition.
III-1-2
This  corresponds  to the  situation  where
complete  conversion  to  producer gas  is
effected; i.e., no char byproduct occurs. The
heat and material balance relationships given
for the gasifier correspond  to gasification of
65  percent of  the fixed  carbon in  the feed
coal.

  The process  consists of two main process
steps— gasification and acceptor regeneration-
connected in series. The gasifier is operated at
15-20  atmospheres  (atm)   pressure at
1 700-1 775°F. Preoxidized coal is fed to the
gasifier along with excess  air. The char bed is
fluidized by steam and recycle gas from the
regenerator operation after it has been proc-
essed  for sulfur  recovery  in  the modified
Claus plant  that is described later.  Calcined
dolomite from the regenerator is fed to the
top of the gasifier.

  The endothermic heat of gasification in the
gasifier is supplied, both  by the  partial com-
bustion of carbon with air, and  by the exo-
thermic heat of absorption of the CC"2 by the
calcined dolomite.

   It is necessary, in order to supply chemical
reaction heat to the gasifier, that  a driving
force for absorption  of CC>2 by the lime be
present.  Because the acceptor falls through
the fluidized char in a "showering" action, it
is  only necessary that the driving  force be
present in the bottom part  of the bed. At the
gasifier temperature  chosen in the Table  1
illustration,  1700°F,  such a driving  force for
CC>2 does exist at the bottom of the gasifier:
PCCH  = 2-9 atm versus the equilibrium value
of 1 .6 atm.

  Another essential  feature of  the gasifier
operation is  simultaneous desulfurization of
the  producer gas in situ,  via the  acceptor
reaction:
CaO + H2S = CaS + H2O
                                        (1)

-------
                      Table 1. SUMMARIZED HEAT AND MATERIAL BALANCE3
Material
Input
Coal
C
H (as H2)
O
S
Ash, Ib
Steam
Coal moisture
Dry air
Moisture
Claus plant tail gas
CO
CO2
N2
H2O
H2
Acceptor
MgO-CaO



Output
Char
C
H
Ash, Ib
Gas
CH4
H2
CO
C02
H20
N2
NH3
H2S
Acceptor*3
MgO-CaO
MgO-CaCO3
MgO-CaS
Preoxidizer and producer
°F

60





1200
60
398
398
1200





1885





1700



1700








1700
mol

—
5.812
2.381
0.475
0.1341
12.30
3.745
0.354
7.174
0.031

0.105
1.958
4.284
0.075
0.002

7.0861





1.223
0.054
12.30

0.277
3.288
3.523
1.956
2.617
9.951
0.047
0.0050

(0.0303) 6.0318
(0.0045) 0.8904
(0.0006) 0.1285
Material

Char
C
H
Ash (Ib)




Dry air
Moisture
Lift gas
CO
CO2
N2
H2O
Acceptor
MgO-CaO
MgO-Ca03
MgO-CaS
MgCO3-CaCO3
(make-up)

Char
C
Ash (Ib)
Gas
CO
C02
N2
H20
H2
S02
S2
H2S
COS
Acceptor
MgO-CaO


Regenerator
°F

1700







398
398
280




1700






1885


1885









1885



mol

—
1.223
0.054
12.30




5.422
0.023

0.011
0.124
0.281
0.002

6.0318
0.8904
0.1285
0.0354



0.122
12.30

0.172
2.024
4.565
0.075
0.004
0.0300
0.0487
0.0005
0.0006

7.0861


a  Basis:    100 Ib dry coal
           15 atm system pressure
           14 atm clean gas delivery pressure
b  Numbers in parentheses are for acceptor discarded in order to maintain acceptor activity.

                      Gross heating value of cold gas    _. _„.
   Cold gas effaency = Gross heatin^va,ue of coa, feed = 74'7/0'
                                                                                            III-1-3

-------
   The gas composition given in Table 1 pre-
sumes that  equilibrium is established in reac-
tion (1). The potential extent of desulfuriza-
tion under  the operating conditions given in
Table  1 is thus 96 percent; i.e., the sulfur in
the producer gas product is only 4 percent of
that in the feed coal.

   Acceptor regeneration  is the second prin-
cipal process step. The acceptor is regenerated
by burning  fuel char in an air-fluidized  bed.
The fuel char is the gasification residue and its
combustion provides the heat to carry out the
endothermic calcination reaction:
            CaCC>3 = CaO + CO2
                            (2)
The regenerator thus serves as a receptacle for
gasification  residue that can be efficiently
used  as fuel.  This procedure  significantly
improves  the total carbon gasification  effi-
ciency  over  that  of an  air-blown fluidized
producer.

   Another feature of the regenerator opera-
tion is that combustion of char is conducted
concurrently with the  endothermic  calcina-
tion reaction (2).  This feature minimizes the
difficulties due to ash fusion that are usually
encountered  in  combustion  of  char  in  flui-
dized beds  and that are particularly severe
when  the  operation  is  conducted  under
pressure.

   Another important process that  occurs in
the regenerator is  rejection of sulfur via  the
overall roasting reaction:

        CaS + 3/2 O2 = CaO + SC>2       (3)

Reaction (3) is  in reality the  composite of
two separate reactions:
CaS + 2 O2 = CaSC-4

3CaSO4 =
                                        (4)

                                        (5)
The  chemistry  of the  sulfur cycle in the
regenerator2 c has been extensively studied in
IH-1-4
                                    work done for the Office of Coal Research in
                                    development of the CO2 acceptor process.

                                      The  material balance of Table 1 provides
                                    for a driving force sufficient to eliminate SO2
                                    via reaction (5); i.e., PSO? = 0-065 atm versus
                                    the equilibrium value of PSO-> = 0-090 atm at
                                    the  regenerator   operating  temperature  of
                                    1885°F This  driving force is sustained by
                                    operating at slightly reduced  conditions in
                                    order to generate enough  CO to reduce SO2
                                    by the reaction:
                                                      2CO + SO2= 1/2S2 + 2CO2
                                                                           (6)
                                   Note that more than half the sulfur issuing
                                   from the regenerator is in the form of ele-
                                   mental sulfur.

                                      There appears to be a theoretical difficulty
                                   in this method  of operation; i.e., at the CO
                                   levels used, calcium sulfate would not seem to
                                   be  present to  effect reaction (5) because it
                                   would be reduced by the reaction:
                                           CaSO4 + 4 CO = CaS •*• 4 CO2
                                        (7)
  In  practice,  however, an  environment of
oxidation sufficient to form CaSO4 is present
throughout most of the regenerator bed, and
reaction (7) occurs too slowly to reduce all of
the calcium sulfate at the bed outlet, perhaps
because of a slow rate of diffusion of CO into
the  acceptor  pores. Verification of  this
method of operation is described later in this
paper.

  Sufficient CO is also present in the regener-
ator off gases to permit reduction of the SO2
to elemental sulfur by a modified multistage
Claus process.  The process is essentially'  the
same  as  that  described in  another  paper4
presented to this conference. Table 1 gives  the
composition of the Claus plant tail gas.

  Important considerations  in  the develop-
ment  of a process of this  type  are acceptor
life and  cost.  Some information has been
developed  experimentally  under conditions

-------
similar  to  those proposed  for this process:
however, more is required. The data available
to date will be  presented later in the paper.
Table 2 gives the estimated acceptor makeup
rate  and cost  under the conditions proposed
for the process. The makeup cost is approxi-
mately 1.3^/MM Btu of product gas.


Table 2. DOLOMITE MAKEUP COST AND SULFUR
                REJECTION

Makeup rate, mol MgCO3'CaCO3/mol
     MgO-CaO circulated, %                0.5

Lb raw dolomite/100 Ib dry-coal feed         7.41

Makeup cost, il\ O6 Btua at $3.25/ton         1.27

Sulfur distribution
   Product gas, %
   Acceptor, %

Spent acceptor composition

   CaO, mol, %
   CaC03
   CaS
 3.7
96.3
85.55
12.63
 1.82
a  Gross heat of combustion of product gas.

   The  literature does  not provide  suitable
kinetic  data  to  properly size  and define the
optimum operating conditions in the gasifier.
Some limited integral rate data  is presented
later in  this paper; however, more is required
to further develop the process.
EXPERIMENTAL PROCEDURE

  The equipment  and operating procedure
have both been fully described in the report
to the Office of Coal Research.20 Briefly, the
gasifier was a 70-in. long, thin-walled vessel,
with a 4-in. ID. The fluidized bed height was
controlled at 40-in. by an overflow weir. The
vessel was contained, along with the electrical
heaters and insulation, inside a pressure shell.
CC>2 was used as  pressure-balancing gas. The
38-in.  long regenerator was of similar  con-
struction, with a  2-in. ID; the  fluidized-bed
height was controlled  at 22 in. by the over-
flow weir. These two  reactors were designed
to operate at temperatures up to 1950°F and
pressures  up  to  300 psig. All vessels and
internals were Type 310 stainless steel except
the regenerator thermowell, which was Hastel-
loy X. The acceptor, sized to 16 x 28 mesh,
was dolomite from the Tymochtee formation
in Western Ohio.

   Calibrated rotary  feeders  controlled the
flow rates of three of the solids: (1) gasifier
feedstock (char or coal), (2) regenerator fuel
char,  and  (3) acceptor  feed  to regenerator.
These solids were fed  from dual lockhoppers
through pneumatic transfer lines. The accep-
tor circulated  continuously between the two
vessels. Acceptor flowed from the regenerator
to  the gasifier by gravity, through a water-
cooled  standleg.  The  acceptor  particles
showered down through the fluidized bed  of
char and segregated as a separate fluidized bed
in the 2-1/2 in. ID "boot"  of the gasifier. The
acceptor  was  then  withdrawn through  a
water-cooled standleg  and a rotary  device
similar to the solids  feeders,  and  was dis-
charged  periodically  from  dual lockhoppers
situated below the rotary device.

   A fixed investory  of acceptor, correspond-
ing to 18 pounds of raw dolomite, was used in
each run. No  fresh makeup was added. The
acceptor discharged  from the bottom of the
gasifier was sampled periodically; its activity
was determined by a special assay.  Thus, the
relationship between activity and the number
of  calcining-recarbonation  cycles undergone
by the acceptor could be determined.

   In  the regenerator, the heat required  to
calcine  the acceptor  was  supplied  by the
combustion of char,  which was  partially gasi-
fied material from a previous run.  The fuel-
char ash was elutriated from the fluidized bed
of acceptor and was removed by an external
cyclone.

   The inlet gas to both  vessels contained
recycled  product  gas, which  was  used  to

                                    III-1-5

-------
 simulate the partial pressures that would exist
 near the top of the tall fluidized beds in the
 commercial process. Hydrogen and CC>2 were
 added to the gasifier recycle gas  to adjust the
 exit  partial  pressures to the desired  values.
 This inlet  dry gas was then passed through a
 saturator in which  the water temperature was
 controlled  to give  the  desired inlet-steam
 partial pressure. The  regenerator inlet gas was
 primarily recycled gas and air. CC>2 or N2 was
 added to adjust the  outlet  CC>2 partial pres-
 sure  to  the desired  value.  The  product gas
 from both vessels was analyzed frequently by
 gas chromatography.
 RESULTS AND DISCUSSION

   The experiments reported here were carried
 out under conditions similar to those for the
 gasifier of the process under discussion,  al-
 though  they were originally  designed  for
investigation of the process as a "CO maker."
The acceptor life  and integral rate data gen-
erated are therefore similar, but not identical,
to what might be expected in practice  with
the process described above. In general, con-
ditions (from the point of view of acceptor
life) were  more severe than those projected
for the process under discussion; i.e., higher
temperatures  were used in the  gasifier and
regenerator, and  lower steam pressures were
used in the gasifier.

  Tables 3 and 4 show the  conditions and
results from the  runs made during the pro-
gram.  Table  3 shows the  conditions  and
results of Run Kl, which  was made:  (1)  to
demonstrate the  operability of the process,
(2) to produce fuel char,  and (3) to  obtain
some practical gasification  kinetics data. The
conditions  and results for  a typical acceptor
life study run are shown in Tables 3 and 4 for
the gasifier and regenerator, respectively.
                      Table 3.  GASIFIER RUN CONDITIONS AND RESULTS

Cycle
System pressure, atm
Acceptor
Size, mesh
Activity
Circulation rate, Ib/hr (raw basis)
Gasifier
Temperature, ° F
Feedstock


Feed rate, Ib/hr (MF basis)
Product char rate, Ib/hr
Recycle fluidizing gas, scfh
Inlet steam, scfh
H2, scfh
C02, scfh
N2, scfh
Total inlet gas, scfh
Inlet fluidizing velocity, ft/sec
Acceptor layer
Bottom of char bed
Run No.
K1
	 	
18
—
—
—
—

1770
Colstrip 1 1
Subbituminous
Dried at 400° F
5.74
3.20
161.4
2.5
0
51.2
0.5
215.6

—
0.165
L2
15
18
Tymochtee-6 Dolomite
16x28
0.39
8.18

1775
Disco char
Carbonized
Dried at 950° F
2.95
1.86
232.3
26.5
0
54.3
7.2
320.3

0.540
0.246
III-1-6

-------
Table 3 (continued) GASIFIER RUN CONDITIONS AND RESULTS

Inlet gas partial pressures, atm
H20
H2
CH4
CO
CO2
N2
Others
Product gas, scfh
N2 purges, scfh
Condensate, scfh
Total outlet gas, scfh
Outlet fluidizing velocity, ft/sec
Outlet gas partial pressures, atm
H2O
H2
CH4
CO
C02
N2
Others
Product gas composition, mol %
H2
CH4
CO
CO2
N2
Others
Product char particle density, Ib/ft3
Char bed density, Ib/ft3
Char bed weight, Ib
Retention times
Solids, hr
Vapor, sec
^ ... . Ib carbon gasified ...4
(Ib C in bed x mm)
Carbon burnout, %
Equilibrium CO2 partial pressure, atm
C02 partial pressure driving force, atm
Char product
H,wt%
C
N
O
S
Ash (diff.)
Run No.
K1

0.21
2.09
0.56
7.86
6.57
0.68
0.03
116.6
4.9
20.4
293.4
0.225

1.25
2.63
0.70
9.86
2.95
0.50
0.11

15.54
4.14
58.37
17.01
4.75
0.19
52.7
25.0
6.21

1.94
15.1
43.8
32.2
2.53
—

0.35
83.23
0.41
0
1.00
15.01
L2

1.49
1.24
0.08
6.01
6.24
2.90
0.08
103.7
12.7
27.3
352.5
0.271

1.39
1.63
0.11
7.89
4.23
2.64
0.11

9.51
0.65
45.98
24.37
19.15
0.34
71.8
27.4
6.87

3.69
11.4
25.7
36.2
2.62
1.61

0.30
76.17
0.64
0
0.29
22.60
                                                              III-1-7

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                    Table 4. REGENERATOR  RUN CONDITIONS AND RESULTS
                                                               Run  No. L2
                    Cycle
                    System pressure, atm

                    Acceptor
                      Size, mesh
                      Circulation rate, Ib/hr (raw basis)
                      Activity

                    Regenerator
                      Temperature, °F
                      Fuel char

                         Size, mesh
                         Feed rate, Ib/hr
                      Overhead solids rate, Ib/hr
                      Recycle gas, scfh

                      Inlet gas
                         N2, scfh
                         CO2, scfh
                         Air, scfh

                      Total inlet gas, scfh
                      Inlet fluidizing velocity, ft/sec

                      Inlet gas partial pressure, atm
                         02
                         CO
                         C02
                         N2
                      Outlet gas
                         Product gas, scfh
                         N2 purges, scfh

                      Total outlet gas, scfh
                      Outlet gas partial pressures, atm
                         02
                         CO
                         CO2
                         N2

                      Recycle gas composition, mol %
                         02
                         CO
                         C02
         15
         18

Tymochtee dolomite
       16x28
          8.18
          0.39
       1940
     Colstrip
Subbituminous char
      28x150
          1.58
          0.82
        182
         45
         28
        100

        355
          1.15
          1.06
          0.44
          3.73
         12.77
        197
         15

        364
          0
          0.89
          4.69
         12.42
          0
          4.75
         25.01
         70.24
III-1-8

-------
              Table 4 (continued). REGENERATOR RUN CONDITIONS AND RESULTS
                      Equilibrium CO2 partial pressure, atm

                      C02 partial pressure driving force, atm
                         Inlet
                         Outlet

                      Run duration, hours
                      Run duration, cycles

                      Lb/fuel char fed during run

                      Carbon burnout, %

                      Overhead material
                         H, wt %
                         C
                         N
                         0
                         S
                         Ash (diff.)

                      Input char
                         H, wt %
                         C
                         N
                         0
                         S
                         Ash (diff.)
                                                               Run No. L2
                 7.60


                 3.87
                 2.91

                26.5
                25.3

                41.9

                63.2
                 0.29
                54.97
                 3.39
                41.35
                 0.62
                77.55
                 0.37
                 0.0
                 0.64
                21.44
Operability
  No serious operability problems were en-
countered in using the CC>2 acceptor process
to make  CO,  using  sub-bituminous coal as
gasifier  feed-stock. In Run Kl, raw Colstrip
sub-bituminous  coal  that had been dried at
400°F was fed  to the gasifier at 1770°F: no
coking occurred. The char product particles
were  not rounded or swollen. In  contrast,
slight preoxidation  (4 percent) of the sub-
bituminous coal was required to avoid coking
in the hydrodevolatilizer of the conventional
CO2 acceptor process, where  H2 partial pres-
sure  was much higher  and  CC>2  partial
pressure was much lower.

  However, the process was not operable when
preoxidized Pittsburgh seam bituminous coal
was used as feedstock.
  Two  attempts were made to feed Ireland
coal that had been preoxidized to levels of 5.5
and  13 percent, respectively, into a bed of
Disco char: no massive coking occurred with
either feedstock, but the coal became plastic
as it entered the bed, and cemented together
the char particles. The agglomerates that were
formed,  although tenuous, plugged the  boot
and forced a shutdown.

  In the actual process as proposed (Figure
1),  coking is prevented by feeding the preoxi-
dized coal  into  the gasifier  entrained in a
stream of gas containing free oxygen, allowing
additional pretreatment  to take place at the
point of introduction into the gasifier. The
efficacy of this method was demonstrated on
a 30 ton-per-day pilot-scale unit at the Con-
solidation Coal Co. Research Center in  1950.
                                    III-1-9

-------
   During Run Kl, where the vapor retention
time was 15 seconds, there was no overt tarry
material in the gasifier product gas; however,
there was some residual naphthalene. Most of
the  naphthalene condensed  as  fog  and was
collected in glass-wool filters installed down-
stream of the primary and secondary product
gas  coolers.  Some of the naphthalene  con-
densed in the primary cooler.


   No deposits of any kind  appeared in the
gasifier  during Run  Kl. In the acceptor life
study runs where Disco char was used as feed-
stock,   however,  slight gasifier  deposits
occurred, as  they had in the earlier work at
1600°F.
Gasification Kinetics

   The  gasifier was operated  with recycled
product gas to simulate an upper section of a
tall fluidized bed. The feedstocks were either
raw coal or partially devolatilized Disco char
(produced from Pittsburgh seam coal).  Thus,
the  measured  masification rates,  shown  in
Table 2, include  carbon-bearing gases formed
during  pyrolysis and  devolatilization.  Com-
parison of the carbon burnout for Run Kl
with that for an earlier run, in which Colstrip
coal was fed to the conventional CC>2 gasifica-
tion process,2 c shows that at the CC>2 and CO
partial  pressures  used, essentially no gasifica-
tion of fixed carbon occurred in Run Kl and
that the  yield  and  product  distribution
represent only the pyrolysis and devolatiliza-
tion reactions  at  the  modified process con-
ditions.

   The  relatively  poor kinetics in Run  Kl is
apparently due to very strong inhibition by
CO  at  the high CO/CO2 ratios employed in
this run.  The equilibrium constant in  the
"Boudouard" reaction is:
            (Pco)2
                    = K= 100
at the conditions used. The Kexp at the outlet
conditions of Run Kl  is about 33. It is clear,
therefore,  that  total  inhibition  by  close
approach to equilibrium  is not in itself the
reason for the poor kinetics.

   The integral rates  observed with Disco char
are representative  of what may be achieved
with  bituminous  coal or  char  feedstocks.
They are marginally acceptable for the proc-
ess  under  discussion. Because   the  steam
partial pressure  would be higher in the pro-
ducer-gas operation, higher gasification rates
are anticipated.
Operation of the Regenerator

  The regenerator was operated at a nominal
level of 4.5 percent CO in  the  outlet gas to
ensure that no difficulty would  be caused by
the CaS-CaSO4  transient liquid. To  achieve
this  conservative CO  level with  fuel  chars
having a wide range-of particle diameter—28 x
150 mesh versus 48 x 65 mesh used in earlier
work   —a  higher  carbon  investory  was
required  in the regenerator bed.  Therefore,
the carbon  burnout was lower  than that at
comparable  bed expansions  in the  earlier
work.
Char Desulfurization and Sulfur Cycle

   In  Run Kl the total  sulfur rejected from
the Colstrip  coal in the gasifier was 35 per-
cent.  This low level is typical of the conven-
tional CO2 acceptor process using low-ranking
Western coals. The high lime content of the
ash in these coals causes  "self-acceptance,"
with  the  result that most of the  sulfur  is
carried to the regenerator in the form of CaS.
In the life study runs with  Disco  char, which
has low lime content in  the ash, sulfur rejec-
tion was at  the high level  of 83-84 percent.
The sulfur content of the  product char was
only 0.29 wt. percent versus  1.16 wt. percent
in the feed. The  above desulfurization was
effected with a  hydrogen partial pressure of
IH-1-10

-------
1.6 atm. The results are roughly in  accord
with prior  work3 on desulfurization of low-
temperature chars from bituminous coals.

   The sulfur content of the gasifier off gases
was not determined. However, if equilibrium
were approached in the acceptor reaction,
          CaO + H2S = CaS +
the H2S content of the gas would have been
about 0.02 mol percent.

   In the life study runs, the rejected sulfur
was equivalent to  the conversion to  CaS of
1.8 mol percent of the  CaO  content of the
acceptor on each pass through the gasifier, a
level  considerably  higher  than in any run
during the conventional CC»2 acceptor process
studies.  Analyses of  the acceptor, sampled
periodically from the gasifier,  showed that all
of the sulfur had been rejected in the regen-
erator, even at the relatively  high level (4.5
percent) of CO in the exit gas.
Acceptor Activity

   Figure 2 shows acceptor activity, measured
preiodically during the runs, as a function of
the number of calcining-recarbonation cycles.
The activity is defined as the mol percent of
CaO that  is active as a CO2 acceptor in the
recirculating inventory. Equilibrium activities
(R) calculated from the data in Figure 2, are*
shown below as a function of fresh acceptor
makeup rate. The method for calculating the
equilibrium  activity  has  been  given  pre-
viously.20

            Makeup, %    R
                1
                2
                3
0.143
0.244
0.320
  Note that the  figure used  for  acceptor
makeup in  Table  2  was only half of  the
figures above, derived from the experimental
                      data presented here. The lower value is justi-
                      fied  by the fact  that the process is projected
                      to operate under milder conditions than those
                      under  which the above  data  was obtained.
                      Correlations developed  in  prior work20  on
                      development  of the  CO2  acceptor  process
                      were used to justify the  lower makeup cost.

                        The  relative  influences  of gasifier  and
                      regenerator bed  temperatures, of steam and
                      H2 partial pressures, and  of  the  extent of
                      conversion  of CaS on acceptor activity remain
                      to be determined over wider ranges of these
                      variables than were  used in the brief study
                      reported here.
BIBLIOGRAPHY

 1. Rudolph, P.F.H.  Coal Combustion for
   Present and  Future Power  Cycles. (Pre-
   prints  of  Division  of Fuel Chemistry,
   ACS.)  Presented  at  Toronto,  May 23,
   1970.

 2. Office of Coal Research, U.S. Department
   of  the  Interior.  Research and Develop-
   ment Report No. 16:
   a.  Interim  Report  No.  1 —  Feasibility
          Study.
   b.  Interim  Report No. 2 — Low-Sulfur
          Boiler   Fuel  via Consol CO2
          Acceptor Process.
   c.  Interim Report No. 3 — Bench Scale
          Research  on  CO2  Acceptor
          Process—Books 1, 2, and 3.
   d.  Interim   Report No.  4  -  Current
          Commercial Economics.
 3. Batchelor, J.D. and E Gorin. Desulfuriza-
   tion of Low-Temperature Char. Ind. Eng.
   Chem. 52: 162. 1960.

 4. Curran, G.P., C.E. Fink, and E.  Gorin.
   Coal-Based   Sulfur  Recovery  Cycle  in
   Fluidized-Lime-Bed  Combustion.  Paper
   presented at the  Second International
   Conference  on  Fluidized-Bed  Combus-
   tion, Hueston Woods, Ohio, October 6,
   1970.
                                  III-l-ll

-------
                                                               PLANT
                                                               STEAM
                    GASIFIER  PLANT
                     STEAM   STEAM
 TO COMBINED
 STEAM AND GAS-*
TURBINE CYCLES
           AIR
           Figure 1.  Simplified flow diagram of clean gas at 14 atm pressure.
    1.0


    0.9


    0.8


    0.7
   \v
%   0.5
tr
O

LU
O
O
0.4



0.3


0.2



0.1



 0
      0
                       10
15        20       25

   NUMBER OF CYCLES
                                                    30
35
                                                                         40
                 Figure 2. Acceptor activity versus number of cycles.
IH-1-12

-------
                                     2. COAL DESULFURIZATION
                                                              ASPECTS OF
                                              THE  HYGAS PROCESS
                       H. L. FELDKIRCHNER AND
                             F. C. SCHORA, JR.

                         Institute of Gas Technology
INTRODUCTION

    One obvious alternative method for avoid-
ing air pollution  in the generation of energy
from  coal is to convert the coal to gas, de-
sulfurize  the  gas,  and  utilize the  low-
pollution, sulfur-free gas to  produce energy.
Among the gasification processes now under
study is the Hygas process being developed at
the Institute of Gas Technology (IGT) under
the joint sponsorship of the U.S. gas industry
(via the  American Gas Association) and the
U.S. Government (via the U.S. Department of
the Interior, Office of Coal Research). The
former is seeking to  secure a supplemental
supply of gas equivalent to natural gas so that
it  will be assured of a future gas supply; the
latter is  vitally concerned with ensuring the
availability  of  adequate energy  supply and
stimulating the utilization of our large coal
reserves.  This process not only converts coal
to  gas,  but produces  byproduct sulfur  in
elemental form.

   The Hygas process has been under develop-
ment  at  IGT since  1946. Until  1964, this
work  was supported almost entirely  by the
American Gas Association (AGA), at which
time  the U.S. Department  of the  Interior
joined with the AGA in developing the proc-
ess concept via a large-scale pilot-plant pro-
gram. This pilot  plant, now in the final con-
struction and shakedown  phase of develop-
ment,  is  designed to produce about  1.5
million cu ft/day of gas from 80 tons of coal.
It is  hoped  that an  80-million  cu ft/day
demonstration plant will  be completed by
1974. Work, some of it still in progress, has
been  done in several smaller development
units.
PROCESS DESCRIPTION

  The  Hygas process has  been discussed in
detail in a number of recent publications.6 -9
Figure  1 gives the basic steps. This process is
operable with high- as well as low-rank coals,
at a  thermal efficiency of  about 70 percent.
However, it is at present necessary to pretreat
agglomerating coals,  such as  the Eastern
bituminous  coals,  before  hydrogasification.
Pretreating  ensures  that  the coal will not
agglomerate  and hinder solids flow in the
hydrogasifier. Nonagglomerating coals, such
as those  in the lignite and subbituminous
class, do not require pretreatment. A substan-
tial  fraction of the sulfur may be removed,
from some high sulfur coals, by pretreatment.

  Hydrogasification is carried out in a two-
stage fluidized-bed reactor. As now planned,
coal or char is fed to the gasifier in a light oil
slurry to avoid the use of lock hoppers. Small
quantities  of light  oils are formed in the
process; they vaporize in a  drying zone at the
top  of the  reactor, leave  the reactor along
with the raw product gas, and are recycled for
reuse. Temperatures in the hydrogasifier range
from about  1300°F to about 1800°F; reactor
                                      III-2-1

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                                          Table 1. TYPICAL ANALYSES OF FEEDSTOCKS

Coal rank
Mine
County
State
Coals (before pretreatment)
Ultimate analysis, wt%
Carbon
Hydrogen
Nitrogen
Oxygen
Sulfur
Ash
Total
Proximate analysis, wt %
Moisture
Volatile matter
Fixed carbon
Ash
Total
Chars (after pretreatment)
Ultimate analysis, wt %
Carbon
Hydrogen
Nitrogen
Oxygen
Sulfur
Ash
Total
Proximate analysis, wt %
Moisture
Volatile matter
Fixed carbon
Ash
Total
Coal seam
Pitts. No.8
HVBA
Ireland
Marshall
W. Va.


67.90
4.91
1.18
6.38
4.38
15.30
100.00

3.9
32.6
48.8
14.7
100.0


69.90
3.09
1.46
10.78
3.33
11.44
100.00

1.4
24.2
63.1
11.3
100.0
Ohio No.6
HVBB
Broken Aro
Coshocton
Ohio


74.40
5.64
1.23
9.64
3.36
5.73
100.00

2.1
39.2
53.1
5.6
100.0


75.00
4.18
1.30
10.80
2.31
6.41
100.00

1.4
26.8
65.8
6.0
100.0
III. No.6
HVBC
Crown
Montgomery
Illinois


70-10
4.88
1.10
9.82
4.00
10.10
100.00

13.3
33.0
44.9
8.8
100.0


70.30
3.58
1.20
10.11
3.33
11.48
100.00

1.4
22.7
64.6
11.3
100.0
W. Va.
Ind. No.6
HVBC
Minnehaha
Sullivan
Indiana


73.20
5.23
1.29
8.73
2.34
9.21
100.00

12.0
32.7
47.2
8.1
100.0


73.00
3.43
1.55
10.11
2.03
9.88
100.00

0.5
22.7
67.0
9.8
100.0
Block 5
HVBA
Kanawha
Kanawha
W. Va.


79.00
5.84
1.34
6.04
0.80
6.98
100.00

1.5
35.7
55.9
6.9
100.0


76.10
3.97
1.51
9.96
0.77
7.69
100.00

0.6
23.2
68.6
7.6
100.0
111-2-2

-------
USED IN HYGAS PROGRAM
Coal seam
Sewell
MVB
Dante
Nicholas
W. Va.
83.20
5.04
1.52
5.58
0.62
4.04
100.00
1.4
29.4
65.2
4.0
100.0
80.20
3.65
1.59
9.85
0.59
4.12
100.00
1.8
19.0
75.2
4.0
100.0
Sewell
LVB
Lockgelly No. 2
Fayette
W. Va.
87.40
4.85
1.52
2.86
0.60
2.77
100.00
1.1
20.6
75.6
2.7
100.0
85.10
3.32
1.49
7.84
0.58
2.67
100.00
1.0
16.2
80.2
2.6
100.0
Pocahontas No .4
LVB
Stotesbury No. 10
Raleigh
W. Va.
85.10
4.20
1.12
2.81
0.73
6.04
100.00
0.7
16.1
77.2
6.0
100.0
85.30
2.23
0.84
5.24
0.59
5.80
100.00
1.2
14.0
79.1
5.7
100.0
Laramie
No. 3
Subbit.
Eagle
Weld
Colorado
75.10
2.56
1.33
15.90
0.38
4.73
100.00
3.7
35.2
56.5
4.6
100.0
73.10
4.27
1.40
15.62
0.27
5.34
100-00
0.3
31.1
63.3
5.3
100.0
—
Subbit.
Colstrip
—
Montana
67.9
4.52
1.01
17.86
0.81
7.90
100.00
3.0
36.6
52.7
7.7
100.0

—
—
—
—
—

__


—

—
Lignite
Glenharold
Mercer
N. Dakota
66.0
4.37
0.56
20.00
0.99
8.08
100.00
5.0
40.8
46.5
7.7
100.0

—
_•_
—
—
—

	 	


—

—
Lignite
Savage
Sidney
Montana
64.8
4.17
0.95
21.22
0.68
8.18
100.00
4.3
39.3
48.6
7.8
Tf)Q&

—
—
—
—
—

	


—

                                                                             III-2-3

-------
pressures  are  up  to  about 1500 psig. Nor-
mally, sulfur removal is greatest in this proc-
essing step.

   After the two stages of hydrogasification,
the partially gasified  char is gasified further
with  steam at  1800-1900° F  in an electro-
thermal gasification stage. The hydrogen and
carbon oxides produced in the electrothermal
gasifier are fed directly to the hydrogasifier.
The char  from the electrothermal gasifier is
then used as fuel for conventional power and
steam generation.  The sulfur content of this
spent char is normally minimal.

   The raw hydrogasifier effluent gases are
water-quenched  and the light oil and  excess
water are  removed along with ammonia  and
coal fines. Next,  the gas passes through an
acid-gas-removal system that  removes carbon
dioxide and sulfur  compounds   primarily
hydrogen  sulfide.  After a final caustic  and
water  wash, the  gas  is catalytically  meth-
anated to reduce  the  carbon  monoxide
content of the gas to less than 0.1 mole per-
cent.  The final gas has a sulfur level well
below  1  ppm,  a dew point of about -40°F,
and a heating value of over 950 Btu/scf.
SULFUR REMOVAL

  As was  mentioned earlier,  most  of the
sulfur is removed from the  coal or char in
three of the  Hygas process steps: (1) coal
pretreatment,  (2) hydrogasification, and (3)
electrothermal  gasification.  These  steps are
discussed in greater detail below.

  Information on the amounts and types of
sulfur  compounds  removed in these three
stages is somewhat limited at this time. This
limited  information  is the result of  the
research program .on the Hygas process having
been directed  primarily toward developing a
process  for producing high-Btu gas from coal,
with emphasis on hydrogasification, electro-
thermal gasification,  and  pretreatment—in
that order. Sulfur removal aspects have been
considered, but to a lesser degree. In addition,
a wide range of coal feeds has been studied,
which limits  the  amount of data on any one
coal and makes it  difficult to draw conclu-
sions for some feedstocks.
  Table 1 lists the coals that have been used
as feedstocks thus far in this program. Most of
the work has been with the  Pittsburgh No. 8
seam, HVBA char; consequently, more  is
known about  sulfur removal from this feed.
For example, over 60 pilot-plant hydrogasifi-
cation  tests  have  been conducted on the
Ireland  mine coal,  whereas  only a few have
been conducted with Illinois No. 6 and Ohio
No. 6 coals.

  Table 2 lists the pyritic and organic sulfur
contents reported12 by the U.S.  Bureau of
Mines for each type of coal studied. This data
should only  be considered  as approximate,
however, because we did not analyze  in our
laboratories  samples of coals tested in the
development unit.
Pretreatment


  Agglomerating  coals  are pretreated  by
making them contact air in a fluidized bed at
about 800°F. This process has been described
in  detail in  earlier  papers.1'3  The  sulfur
removed is primarily  as sulfur oxides,  al-
though there are undoubtedly small amounts
of organic sulfur and other sulfur compounds
present.  The  pretreater  off gas could  be
treated with  one of  the  many developing
sulfur-oxide-removal  processes  to   recover
sulfur, or  it  could, after concentration  of
SO2, be  combined with the hydrogen-sulfide-
containing  gases removed  from  the  hydro-
gasifier effluent  and sent to  a Claus plan>t  for
elemental sulfur production. The large pilot
plant now  being  completed will contain a
Claus-type plant.
  Figure 2 shows the  results of typical pro-
treatment tests for the coals listed in Tables 1
and  2. Although only a  limited amount of
III-2-4

-------
data is available and  the  data exhibits con-
siderable scatter, it is clear that sulfur removal
in pretreatment increases with  increases in
total  sulfur content. Accurate measurement
of the fraction of sulfur removed is difficult.
especially for  low-sulfur coals, because of the
small  differences  in the sulfur  contents of
most of the coals and chars. Inaccuracies in
material balances, variations in  pretreatment
levels, and  sampling errors all contribute to
the overall error in this measurement.

   Figure 3  shows the  effect of organic sulfur
content  of  the  coal  on  sulfur  removal.
Apparently pyritic  sulfur  is   more  easily
removed  from  the bituminous  coals  than
organic  sulfur.  The  higher  sulfur removal
measured for the lower rank subbituminous
coal is not considered significant in view of its
very low total sulfur content.

   We  expect  that more information will be
obtained from the results of the  operation of
the large pilot plant now being completed. In
general,  however, we  can expect to remove
from 20 to 40 percent of the sulfur present in
bituminous coals in the pretreatment step.
Table  2. TYPICAL SULFUR  BREAKDOWNS RE-
   PORTED FOR TYPES OF COALS TESTED8
Coal
Pittsburgh No. 8
Ohio No. 6
Illinois No. 6
Indiana No. 6
W. Va. Block 5
W. Va. Sewell
W. Va. Sewell
Pocahontas No. 4
Laramie No. 3
Montana
N. Dakota
Montana
Type
HVBA
HVBB
HVBC
HVBC
HVBA
MVB
LVB
LVB
Subbit.
Subbit.
Lignite
Lignite
Pyritic
sulfur
wt%b
2.20
1.80
2.84
1.15
0.05
0.06
0.04
0.09
0.06
0.39
0.02
0.37
Organic
sulfur
wt%b
2.27
1.61
2.17
0.87
0.55
0.51
0.50
0.51
0.22
0.41
0.85
0.36
Total
sulfur
wt%b
4.48
3.48
5.09
2.06
0.61
0.57
0.55
0.60
0.28
0.82
0.88
0.77
aReported by U. S. Bureau of Mines, 1C 8301 -12
bDry basis.
H y drogasif ication


   More  data  is available on  sulfur removal
during hydrogasification than during pretreat-
ment  because  this  program  has  been con-
cerned  primarily with  the  hydrogasification
step. Details on this step have been published
previously.7-8  When  coals and  chars  are
hydrogasified  with  either  hydrogen-steam
mixtures  or synthesis gas, sulfur is removed
primarily as hydrogen  sulfide and  carbonyl
sulfide,  which  can  be removed  from  the
hydrogasifier effluent by a number of well-
developed processes.  One such approach, for
example,  is  to scrub out r^S, COS, and €03
by the hot  carbonate process and  then con-
vert the H2S to elemental  sulfur in a Clans
plant,  as was  discussed  above.  COS is  hy-
drolyzed to H2S in the gas-cleanup operation.
The Claus plant feed could  contain (on a  dry
basis)  from  about 3 mole percent (for lignite
feeds)  to about  32 mole percent  (for bitu-
minous coal chars and high-hydrogen-content
feed gases to the hydrogasification section), if
the pretreater off gas were combined with the
H2$,  and  CO2  were  removed   from  the
hydrogasifier effluent.
  Figure 4 shows that, as in the pretreatment
test  data, considerable scatter exists in the
sulfur removal data for hydrogasification. The
major  trend  observed is  that the degree of
char sulfur removal varies directly with the
MAF char-gasification level. In contrast to the
pretreatment results,  sulfur is removed more
readily from chars from coals with the highest
organic-sulfur contents.  The  lignite  feeds,
which  were not pretreated and which hydro-
gasify  much  more  readily than higher rank
feeds, gave results comparable with the  other
feeds.  In addition, the results with  North
Dakota lignite compare favorably with  those
for  Montana  lignite,  even  though  their
organic/pyritic  sulfur ratios differ markedly.
The  West Virginia coals, all of which have a
high organic/pyritic sulfur ratio, allowed even
higher sulfur removal. This is just the opposite
                                     III-2-5

-------
of the pretreatment  results, but is in  agree-
ment with the coal hydrodesulfurization work
of Vestal and Johnston.1 J

  The large amount of scatter in  the data
shown in  Figure  4 suggested that other vari-
ables (e.g.,  temperature,  pressure, and gas
composition)  might also  be important. The
only approach available at that point seemed
to be regression analysis becuase the test pro-
gram to date had been concerned with hydro-
gasification variables,  not with those variables
influencing  sulfur  removal. We attempted,
therefore,  to  correlate  the  results   with
Ireland-mine coal char (the majority of testing
was  done with this feed material) using mul-
tiple linear regression analysis.

  There were three main  sets of data:  those
using  hydrogen-steam  feed-gas  mixtures,
those  using hydrogen-methane steam, and
those using  synthesis gas.  The results of the
regression  analysis  showed that  the   only
statistically  significant  variable  for the syn-
thesis gas "and hydrogen-methane-steam feed-
gas  runs  was  temperature. The  hydrogen-
steam  run results showed a dependence on
both MAF-char gasification and feed-gas/char
ratio. We  then decided that these results were
subject to too great an experimental error for
conventional techniques. We did try to  force-
fit the results by standard techniques; Figure
5 shows that  the predicted results agreed, in
general,  quite well  with  the  experimental
data. It  appears that  further work will be
required  to establish precisely  the variables
affecting  sulfur  removal. It is  possible,
however,  to estimate  sulfur removal from the
coals tested to within about ±10 percent.
Using  this  information,  we  can  expect  to
remove anywhere from 40 percent  to  essen-
tially all of the feed-char sulfur content  in the
hydrogasification step.
Electrothermal Gasification

   At the present time, less data is available
from our electrothermal gasification program.
Again,  details of  this program  have been
III-2-6
published  elsewhere.4-s  Because  the  gas
produced in the electrothermal gasifier passes
directly into the hydrogasifier, it is important
to be able to  predict its sulfur content. Any
sulfur compounds evolved here will ultimately
end up in the hydrogasifier effluent.

  The  data  available  correlates better for
these tests (Figure 6) than for those discussed
previously. It  should be pointed  out that the
variable  is  "char" gasified  (in  the  electro-
thermal  gasifier),  not  "MAF-char" gasified.
Also, the sulfur removal data is based on the
sulfur  content of  the  feed  to  the  electro-
thermal gasifier; this feed has had most of its
sulfur removed in the two previous processing
steps—pretreatment and hydrogasification.

  The closer agreement of the data from the
various  feeds   tested may stem  from  their
having similar properties:  they are all highly
gasified materials. Earlier  work2  has shown
that the reactivities of  coals, ranging in rank
from lignite  to  anthracite, approach one
another at high levels of gasification.
Economics of Sulfur Removal

  Figure  7 shows a  typical breakdown of
sulfur removal from the coal in each stage of
the Hygas process. It is difficult to extract the
true cost of sulfur removal from available data
because  of the uncertainties in  the  future
value  of sulfur. It is also difficult to separate
sulfur-removal costs from  the total costs of
gas purification. However,  if elemental sulfur
could be sold  for  $25/long ton, the Hygas
process economics might  be more  favorable
for higher sulfur coals; the cost  of sulfur
removal  would then be less than the sulfur
byproduct credit. For low-sulfur  coals and a
lower sulfur price, the reverse would be true.
These  figures  assume  a   final  gas cost  of
5-6^/therm  (based  on  AGA  accounting
procedures),  a  coal cost of  16^/million  Btu,
and present  economic conditions in  the
United  States.  A detailed  study1 °  on sulfur
recovery in the manufacture of pipeline gas
from coal has recently  been presented.

-------
BIBLIOGRAPHY

 1. Channabasappa,  K.C. and  H.R. Linden.
   Fluid-Bed  Pretreatment  of  Bituminous
   Coals and  Lignite—Direct Hydrogenation
   of Chars to Pipeline Gas. Ind. Eng. Chem.
   50: 637-44, April 1958.

 2. Feldkirchner, H.L.  and  H.R.  Linden.
   Reactivity of Coals in High-Pressure Gasi-
   fication With Hydrogen and Steam. Amer.
   Chem. Soc. Div. Fuel Chem. Preprints, p.
   191-208,  September 1962;  also  I&EC
   Process Design Develop. 2:  153-62, April
   1963.

 3. Kavlick, V.J. and B.S. Lee.  Coal Pretreat-
   ment in Fluidized Bed. Amer. Chem. Soc.
   Div. Fuel  Chem. Preprints. 10: 131-39,
   September 1966; also Advances in Chem-
   istry  Series No.  69, p.  8-17.  American
   Chemical  Society,  Washington, D.C.,
   1967.

 4. Kavlick, V.J. and B.S. Lee.  High Pressure
   Electrothermal  Fluid-Bed Gasification of
   Coal  Char.  Paper  presented  at  64th
   National Meeting of the A.I.Ch.E., New
   Orleans, March 16-20, 1969.

 5. Kavlick, V.J., B.S. Lee and F.C. Schora.
   Electrothermal  Coal  Char Gasification.
   Paper presented  at the Third Joint Meet-
   ing of the I.I.Q.P. and the A.I.Ch.E., San
   Juan, Puerto Rico, May 17-20, 1970.

 6. Lee, B.S.  Synthetic Pipeline Gas From
   Coal by the  HYGAS Process. Paper pre-
   sented  at  the  American  Power Con-
   ference, Chicago, April 21-23,  1970.
 7. Pyrcioch, E.J. and H.R. Linden.  Produc-
   tion of Pipeline  Gas  by  High-Pressure
   Fluid-Bed  Hydrogasification  of  Char.
   Amer.  Chem.  Soc. Div. Gas Fuel Chem.
   Preprints, p. 59-69, September  1969; also
   as Pipeline  Gas by High-Pressure Fluid-
   Bed Hydrogasification of Char. Ind. Eng.
   Chem.  52: 590-94, July 1960.

 8. Pyrcioch, E.J., B.S. Lee and F.C. Schora.
   Hydrogasification  of Pretreated Coal  for
   Pipeline Gas Production. Amer.  Chem.
   Soc.  Div.   Fuel  Chem.  Preprints.  10:
   206-23, September  1966; also Advances
   in  Chemistry  Series  No.  69,  104-27.
   American Chemical Society, Washington,
   D.C., 1967.

 9. Schora, F.C. and B.S.  Lee. Hydrogasifi-
   cation  Process. Paper presented  at 65th
   National Meeting  of the A.I.Ch.E., Cleve-
   land, May 4-7, 1969.

10. Tsaros, C.L.,  J.L. Arora,  W.W.  Bodle.
   Sulfur  Recovery  in the  Manufacture  of
   Pipeline Gas From  Coal. Amer.  Chem.
   Soc.  Div.   Fuel  Chem.  Preprints.  13:
   252-69, September 1969.

11. Vestal, M.L. and  W.H. Johnston. Chem-
   istry and Kinetics of the Hydrodesulfuri-
   zation  of Coal. Amer. Chem.  Soc. Div.
   Fuel Chem. Preprints.  14:  1-11, May
   1970.

12. Walker, F.E. and  F.E. Hartner. Forms of
   Sulfur  in U.S. Coals. U.S. Bureau of Mines
   Information Circular 8301. Washington,
   D.C., 1966.
                                                                               III-2-7

-------
                      FUEL GAS
                                     HYDRO-
                                     GASIFIER
           PRETREATER
                      RAW
                   HYDROGEN-
                      RICH
                      GAS

•.'••'.'.'•'•'"•'
. •...".
(
1^>




(HIGH-
BTU GAS


                                                 GAS PURIFICATION
                                                 AND METHANATION
              STEAM
                                           ELECTROTHERMAL
                                              GASIFIER
                                                                 ELECTRICITY
                                   CHAR1
                                           AIR:







,1,1,1
POWER
GENERATION
I'l'i' 1






                                                                              ASH
    50
    40
    30
 DC
 Q_
 a
 DC
 D
 Q
 LU
 o
cc
D
20
    10
    0
                    I      f
                    Figure 1. Schematic flow sheet of Hygas process.

                                                50
          O BITUMINOUS COALS
          * SUBBITUMINOUS COAL
      012345
SULFUR CONTENT OF COAL FEED, wt % (dry basis)
Figure 2.  Sulfur removal from coals during
pretreatment increases with increasing sulfur
content.
III-2-8
                                                 tr.
                                                 o.

                                             cc
                                             u.
                                                    40
                                                     30
cc
Q
O
£   20
O
                                                     10
                                                    ^*u**i 'v •' • '•; \ •'
                                                         ^^        •'•
          ° BITUMINOUS COALS
         . A SUBBITUMINOUS COAL
                                                  40    50   60    70    80    90
                                                       ORGANIC SULFUR CONTENT OF
                                                        COAL FEED, % of total sulfur
                                       100
                                            Figure 3.  Sulfur removal from coals during
                                            pretreatment decreases with increasing
                                            organic sulfur content.

-------
    100
  OC
  o:
  u.
 o
     30   	
       10   20   30   40   50    60

             MAP CHAR GASIFIED, %

o PITTS NO. 8 HVAB CHAR  OCOLORADO SUBBIT CHARS
                        AND COAL
• NORTH DAKOTA LIGNITE OMONTANA SUBBIT COAL
6OHIO NO. 6 HVBB CHAR AMONTANA LIGNITE

^7 INDIANA HVBC CHAR

DILLINOIS NO. 6 HVBC CHAR
OW. VA HVBA
        MVB CHAR
        LBV
  Figure 4.  Char sulfur removal increases
  with char gasification in hydrogasifier.
 Or
 CC
 D
 O
    100
     90
     80
     70
     60
     50
     40
      20      30       40       50
                CHAR GASIFIED, %
                                        60
   O IRELAND MINE HYDROGASIFIED CHAR

   AIGT-PRETREATED CHAR

   D INDIANA HYDROGASIFIED CHAR
   tfHV BIT. HYDROGASIFIED CHAR

   O FMC CHAR

 Figure 6. Char sulfur removal in electro-
 thermal  gasifier increases with increasing
 char gasification.
                                                 O


                                                 OC
                                                 OC
                           Q
                           LU

                           I
                           O
                           <
                           O
100


 90



 80


 70


 60



 50


 40
                                                                       I
                                                             I
                                40    50    60    70    80    90    100

                                   OBSERVED SULFUR REMOVAL, %
                             o HYDROGEN-STEAM

                             A SYNTHESIS GAS

                             O HYDROGEN-METHANE-STEAM
                            Figure 5.  Agreement between measured
                            sulfur removal data and forced-fit data.
                           SULFUR
                           IN RAW
                            COAL
                          TO HYGAS
                          PROCESS
                           Figure 7. Typical sulfur removal in each
                           stage of the Hygas process.
                                                                                   HI-2-9

-------
                          3.  STEAM-OXYGEN GASIFICATION
                                                                             OF
                                               VARIOUS U. S. COALS

                       A. J. FORNEY, S. J. GASIOR,

                   R. F. KENNY, AND W. P. HAYNES

                            U. S. Bureau of Mines
ABSTRACT

  The Bruceton process for making a supple-
mental natural gas from coal consists of gasifi-
cation, purification,  and methanation steps.
The advantages of the system are that caking
coals can be used in the fluid-bed gasifier, and
more than half of the methane can be made in
the gasifier  instead  of in  the  methanator.
Furthermore,  the final product contains no
sulfur and has a heating value exceeding 900
Btu/cu ft.
INTRODUCTION

  The need for a supplemental natural gas is
becoming  more  evident each  year. The
reserves-to-production ratio, now about 14:,!,
has been declining. If the cost of natural gas
continues its present upward trend,  a
substitute natural gas will be able to compete
with the natural product in the immediate
future.

  One source of this substitute gas is to make
it from coal because of its widespread abun-
dance and low  cost. The steam-oxygen proc-
ess for making supplemental gas from coal not
only makes the product inexpensively, but
also decreases the sulfur pollution to a tolera-
ble  level. The  process has been previously
reported in the literature.1 -2 -3
  All of the work described here was done at
the Bruceton, Pa., laboratories of the Pitts-
burgh Energy Research Center of the Bureau
of Mines.
THE OVERALL PROCESS

  The Bruceton process for making a supple-
mental natural  gas is shown schematically in
Figure  1. The main components are:  (1) the
gasifier; (2) the shift converter, which changes
the gas ratio from approximately IK^ ICO to
3H2: ICO; (3) the purification system, which
removes most of the CO2 and all of the sulfur
compounds from the gas; and (4) the catalytic
methanator, in which 3H2 +  ICO reacts to
form CH4 + H2O. Because the shift and  puri-
fication systems are commercial, research at
the Bureau is concentrated on the gasifica-
tion and the methanation steps. This paper
will concentrate on the  gasification step; it
should  be emphasized, however, that the gas
is 400-500 Btu/cu ft before methanation and
afterwards should be of pipeline quality—at
least 900 Btu/cu ft. Because of the nickel
catalyst used  in  the methanation  step, the
sulfur level cannot exceed  1 ppm; therefore,
the final product may be  considered sulfur-
free.

   The gasifier is shown  schematically  in
Figure  2 and photographically in Figure 3.
                                      III-3-1

-------
The Bructon gasifier combines the operations
of pretreatment, carbonization,  and gasifica-
tion in one reactor. The coal,  70 percent
through 200 mesh,  is dropped  through the
pretreater with oxygen and steam (or CC>2) at
400°C, where it is rendered noncaking. The
decaked coal then falls into the carbonization
zone  and finally into the gasification zone,
where it is gasified with steam plus oxygen at
900-1000°C. The gasifier is operated  at  40
atmospheres pressure.

  Advantages of the Bruceton gasifier are: (1)
it can  handle  both caking and noncaking
coals, (2)  gases from pretreatment add to the
product gas from the gasifier, and (3) it makes
a purified product gas with a very high (20-30
percent) methane content (the balance is H2 -
CO).

  This high  methane  content  means  the
volume of gas is less than if it were H2 + CO
only. Thus, the  size of all process vessels and
lines downstream  from  the  gasifier  are re-
duced  as  much as 30-40 percent;  the  final
methane reactor is reduced about 50 percent
from the size it would need to be if only  H2 +
CO were in the feed gas.

   The Bruceton  process has been estimated
to  have an overall thermal efficiency of 63
percent. The selling price for the product gas
has been  estimated  at  54^/M cu ft,  a  price
that is competitive with that of other coal-to-
gas processes.4


   Table 1  shows the results of selected tests
with four  coals. The free-swelling index (FSI)
is  used to  indicate the caking or coking prop-
erty of each coal. Pittsburgh seam, a highly
caking coal, has a maximum FSI of 8  to 9;
Illinois No. 6,  a  weakly caking coal, has an
FSI of about 4.5; and North Dakota lignite
and Montana subbituminous coal have an FSI
of 0 (i.e., they are noncaking).
                 Table 1. RESULTS OF TESTS WITH VARIOUS COALS GASIFIED
                                  AT 40 ATMOSPHERES
Test
Free-swelling index
Carbon conversion, %
Product gas, H2+CO
+CH4, scf/lb
MAF coal
CH4, scf/lb MAF
Gas analysis, %
H2
CO
CH4
CO2
Tar, % feed
Type coal
Pittsburgh
Seam
8 to 9
67
17
3.9

34
19
16
32
2.9
Illinois
No. 6
4.5
72
18
3.7

36
19
14
31
3.4
North
Dakota
Lignite
NCa
88
25
4.2

35
23
15
29
2.0
Montana
Subbit.
NCa
71
16
3.0

38
11
11
40
3.3
Goals

65
17
4.7



-
            aNC = Noncaking.
III-3-2

-------
  The carbon conversion for the tests shown
in Table  1  ranges from 67-88  percent; our
goal is 65 percent. This degree of conversion
permits the  complete utilization of carbon in
the overall process because the  char and tar
can be burned to raise the steam needed for
compression and other steps in  the process.
Other  goals required  to  make  the  process
economical are listed in Table 1. It is desirable
to make about  17 scf of product gas per Ib
MAP coal  because  this  results  in  a final
high-Btu gas volume  of 8 to 9 scf per Ib MAP
coal. We  want to make 4.7 scf CH4 per Ib
MAP coal, but are  slightly low on this goal.
The gas  analyses are  similar  even though
different coals are used.
  Table 2 shows that the gas leaving the gasi-
fier contains most of the sulfur in the coal.
Some sulfur is left in the char and some in the
tar; however, both will be burned.
  It is desirable—even necessary-that the sul-
fur  content- of both the char and tar be below
1 percent so they can be burned without dis-
charging unacceptable  amounts  of  SOX into
the air. Table 3 shows that this limitation is
achieved with all coals tested except Illinois
No.  6. This coal, therefore, and similar high-
sulfur  coals may present a problem. If the
carbon  conversion is increased to lower the
                  Table 2. DISTRIBUTION OF SULFUR DURING GASIFICATION3
Sulfur distribution, Ib
Coal
Char
Tar
Gasb
Total S in gas, %

Pittsburgh
Seam
26.0
3.1
0.5
22.4
86
Type
Illinois
No. 6
78.0
11.2
1.1
65.7
84
coal
N. Dakota
Lignite
16.0
4.4
0.4
11.2
70

Montana
Subbit.
28.0
4.4
0.3
23.3
84
              aBasis: per ton of coal feed to gasifier.
              bBy difference.


                    Table 3. SULFUR CONTENT OF COALS AND PRODUCTS
Sulfur content, wt %
Coal
Char
Tar
Gas
Type coal
Pittsburgh
Seam
1.3
0.5
0.8
0
Illinois
No. 6
3.9
1.4
1.7
0
N. Dakota
Lignite
0.6
0.6
1.0
0
Montana
Subbit.
1.4
0.9
0.5
0
                                                                                   III-3-3

-------
 sulfur in the char, the amount of char is
 decreased.  This  means  low-sulfur supple-
 mental coal must be burned to raise the steam
 required  for  the  overall process.  Another
 possibility is that the coal could be washed to
 achieve a lower level of sulfur before gasifica-
 tion.

 CONCLUSION

   The  high-Btu  gas made in the Bruceton
 process is low enough in sulfur that it causes
 no problem with air pollution requirements.
 The chars and tars from most coals can be
 burned to raise steam and can be made so the
 total sulfur in the feed  to the power plant is
 less than 1  percent. Some high-sulfur coals
 may present a problem, however, because the
 sulfur in the char and tar may be too high.
 More research will be required to resolve this
 problem.

 BIBLIOGRAPHY

 1. Field, J. H. and A. J. Forney. High-Btu Gas
   via Fluid-Bed Gasification  of Caking Coal
  and Catalytic Methanation. Proceedings of
  Synthetic  Pipeline Gas  Symposium.
  American Gas Association, Pittsburgh, Pa.
  p. 83-94, November 1966.
2. Forney, A. J., S. Katell, and W. L. Crentz.
  High-Btu Gas From  Coal via Gasification
  and Catalytic Methanation. Proceedings of
  American  Power  Conference.  Chicago,
  Illinois, April 1970.
3. Forney, A. J., S. J. Gasior, W. P. Haynes,
  and S. Katell. A  Process to Make High-Btu
  Gas  From  Coal.  BuMines  Technical
  Progress Report 24. 6 p. April 1970.
4. Tsaros, C. L. and T. J. Joyce. Comparative
  Economics  of Pipeline  Gas  frorri  Coal
  Processes. Proceedings of Second Synthetic
  pipeline Gas Symposium. American Gas As-
  sociation,  Pittsburgh,  Pa.  p.  131-146,
  November 1968.
HI-3-4

-------
                COAL
   STEAM        f-*-|
   OXYGEN	*H
       SPRAY
       TOWER
                                                  SHIFT
                                                CONVERTER
                                             HOT
                                          CARBONATE
                                           SCRUBBER
         GASIFIER
   STEAM
   OXYGEN
       TAR
       AND
       /DUS
PIPELINE
 GAS
               RESIDUE
                                                                     Hjs
                                                                     cos
                                                                     C02
 FINE
SULFUR
REMOVAL
                                                      METHANATOR
            Figure 1.  Flow diagram of system used to make high-Btu gas from coal.
                COAL FEED HOPPER
  STEAM GENERATOR
     PRETREATERr
  WATER
 OXYGEN

  WATER
STEAM GENERATOR
    GASIFIER
 OXYGEN
                                                  OXYGEN     CHROMATOGRAPH
                      ASH HOPPER'  '             ANALYZER


                 Figure 2.  Diagram of a 40-atmosphere fluid-bed gasifier.
                                                                                  III-3-5

-------
Figure 3.  40-atmosphere gasifier.

-------
                                        4.   DESULFURIZED FUEL
                                                              FROM COAL
                                  BY INPLANT GASIFICATION


                     E. K. DIEHL AND R. A. GLENN
                      Bituminous Coal Research, Inc.
ABSTRACT

  The growing demand for electrical power
will  continue  to  draw on  fossil fuels as  a
source of energy. At the same time, the need
to control  and reduce environmental  pollu-
tion  will dictate  the consumption of fossil
fuels  in a  manner that will minimize  their
contribution to atmospheric pollution.
  For coal, a  solution lies in the conversion
of the solid fuel into a clean, sulfur-free fuel
gas. This fuel gas can be subsequently utilized
in a variety of  downstream systems, many of
which  promise overall improvements  over
conventional power-generating systems.
  Bituminous  Coal Research, Inc. (BCR) is
developing, for the Office of Coal Research,
U.S.  Department of the Interior, a two-stage,
super-pressure,  oxygen-blown coal gasifier for
the  production  of pipeline  gas.  Results
obtained thus  far  have led to the conclusion
that  a similar air-blown system would be ad-
vantageous  for the production of a fuel gas
suitable for desulfurization and pollution-free
combustion in power plants.
  This paper describes the BCR gasifier and
suggests several alternate uses of the clean-fuel
gas so that a major reduction in atmospheric
pollution can be realized with continued use
of coal in the growing energy market.

INTRODUCTION

  As  a result of the increased concern  over
environmental pollution, strict limitations and
control  requirements  are being imposed on
processes emitting atmospheric contaminants.
This is  especially  true  of  fossil-fuel-fired
electric generating stations that may emit fly
ash and/or  sulfur oxides.  Flue gases can be
processed to  minimize fly ash emission at
reasonable costs.  Currently, the development
of  equally  satisfactory  methods for the
control  of  sulfur  oxide  emission is the
principal objective of major research programs
being  conducted  by  both  Government and
industry.

  The electric power  generating industry, in
particular, finds  itself on the horns of a
dilemma. The demand for  electricity  within
our society  continues  to grow at a minimum
rate of 7 percent each year, with a doubling
time of about 10 years; in 1968 and 1969, the
growth   rate  was 9 percent.  Generating
capacity to  meet this demand  is  being
constructed, planned, and projected on a scale
never before experienced.

  Although  there  is  a wide  difference of
opinion  concerning how this  mushrooming
generating capacity will be fueled, there is
general agreement that for the next several
decades  coal will remain a major source of
energy  fuel. Thus the  dilemma  for the
industry: huge coal  resources will continue to
be  tapped  to keep pace  with  the  electric
power demand, but environmental pollution
control will  dictate their consumption in a
manner that will assure minimum atmospheric
contamination.
                                      IH-4-1

-------
   While there may still be some minor gains
to be  realized  in  conventional  fuel-to-
electricity cycles,  research is rapidly shifting
toward  the development  of  new uncon-
ventional concepts. Basic  to the development
of these new concepts is the realization that
they  must  include features that reduce  or
eliminate environmental pollution problems.
Further, new  concepts, in  order  to  be
accepted and  adopted  by the power genera-
tion industry,  must be economically competi-
tive with conventional systems equipped with
pollution control devices.
   In the recent  years, terms such as magneto-
hydrodynamics, electrogasdynamics,  ad-
vanced-cycle power systems,  fuel cells, flu-
idized-bed combustion, and gasification  have
become  commonplace wherever  technical
people  gather to  contemplate  the  future  of
power generation. It is obvious  that uncon-
ventional concepts are receiving serious atten-
tion,  and that they will,  in time, take their
place in the  total power generation picture.

   For  coal-fired  installations,  a promising
approach is the conversion  of  the coal  to
sulfur- and  dust-free gas prior to its ultimate
use as an energy fuel. The careful selection of
specific operating  parameters and the appli-
cation of new  technology  can  lead to the
production of gas  having a gross heating value
ranging from about 200 Btu/scf to as high as
950 Btu/scf.

   Such a system holds promise as a means of
generating power  not only without air pollu-
tion,  but also with  decreased  overall cost,
especially if recovered sulfur can  be  sold as a
credit against the cost of the process. In addi-
tion,  a reduction  of nitric oxide emission is
conceivable  because  of  the  more precise
combustion  control  possible  with  gaseous
fuels.
GAS PRODUCERS

  The only commercial fixed-bed process for
the gasification  of coal at elevated pressure
111-4-2
available today  fulfills the requirement for
gasifying coal with  high  thermal efficiency.
However, this process is not suitable for the
gasification  of  highly  caking coals and is
limited  to the use of lump fuel. The permissi-
ble gas-flow  velocity  limits  the  size of the
individual gasifiers,  as well,  and for a large
utility  power  plant,  costly  multiple  units
would be  required.  The production  of tar
inherent in  this  process  also adds to  the
process cost because purification of both the
liquid and the gaseous effluents then becomes
necessary.

   Coal-gasification processes have  also been
developed, using fluidized beds. The Winkler
process,  operating at atmospheric  pressure,
has been used in large units burning predomi-
nantly brown coal. Operation  of the process
at elevated pressure has been  carried  out in
pilot  plants only, and  with  only noncaking
fuels.  Fluidized-bed  gasification  processes
leave  a residue  with a high carbon content
that requires separate utilization  and  thus
complicates  the process  economics.  The
operability  of  fluidized  beds  at  elevated
pressure using caking coals has not, moreover,
been established. This system  requires devel-
opment work on  a large  scale, and may im-
pose limitations  on such a process.

   The simplest  process is the  gasification of
coal in the entrained state. Very large units
can be built and complete carbon conversion
assured by  operating under  slagging condi-
tions. These  processes use coal at  any rank,
regardless of caking properties.  Many com-
mercial  plants,  operating  at  atmospheric
pressure, have been  built to  use  peat, brown
coal,  and bituminous coals as raw  material.
Operability of this type of process at elevated
pressure has been proved in pilot plants.
BCR Two-Stage Gasifier

   BCR is  developing, under sponsorship  of
the Office of Coal Research,  a two-stage
process for the gasification  of coal for the
production of high-Btu pipeline gas.1 -2 »3 The

-------
process  combines  the advantages-high  gas
yield and low exit temperature-of the more
complex  fixed- and  fluidized-bed processes
with the simplicity of entrained processes and
their ability to use any coal.
   The basic component of the process under
development by BCR is the two-stage gasifier
(Figure  1).  Fresh coal and  steam are intro-
duced into the upper section (Stage 2) of the
gasifier, where the fresh coal is contacted and
entrained by a rising stream of hot synthesis
gas produced in the lower section (Stage 1).
The fresh coal  is rapidly heated and partially
gasified to methane and additional synthesis
gas in Stage 2. Residual char, swept out of the
gasifier along with the raw gas, is separated,
then returned  to  Stage 1. Here the char is
completely  gasified under  slagging conditions
to  produce both  the synthesis gas and  the
heat required in Stage 2v
   For the production of pipeline gas, Stage  1
is blown with oxygen and steam, and  the
product  gas from Stage 2 undergoes purifica-
tion,  partial  shift  reaction,  and catalytic
methanation to raise the heating value above
900 Btu-/scf. The gasifier is operated at about
lOOOpsi.

   Experimental results obtained so far in the
development of the two-stage, super-pressure
process have led  to  the  conclusion  that, if
blown with air rather than with oxygen,  the
two-stage principle  would also  be  advanta-
geous for the production of a fuel gas suitable
for desulfurization and pollution-free combus-
tion in  power plants. Thus,  the air-blown
gasifier could  be  used to generate fuel  gas
from which sulfur, as well  as particulates,
could  be  removed  by  existing cleanup
systems. The  resultant clean,  sulfur-free  gas
could then  be  utilized in  a variety of power-
generating systems.

Air-Blown Gasifier

  Using information generated in the pipeline
gas project,  studies have been made to verify
the technical feasibility of using the two-stage
gasifier to supply  a fuel gas. Figure 2 is one
example  of a  flow diagram  and  material
balance for a conceptual fuel-gas system.  A
basis of 100 Ib of moisture and ash-free coal is
used for the material balance,  the gasifier is
operated at 20 atmospheres, and the gasifier
outlet temperature is 1700° F.

  Steam  and air are preheated to  1700°F.
For the purpose of the example, it is assumed
that progress in materials  development  and
the  small  pressure  differential  between
products  and feeds can make  possible  this
preheat step.

  Table 1 gives  analyses of the feed  coal and
resultant fuel gas; Table 2 shows a simplified
heat balance.

Mathematical Model

  The  material and heat  balances  for the
example  system  were estimated  using  a
mathematical model  developed  at  BCR.4
Figure  3  shows  the general scheme of the
model.  In Stage 1,  all of the materials are
mixed thoroughly and the makeup  char  is
gasified. The resulting  mixture  comes to
water-gas  shift  equilibrium at  the  Stage  1
temperature  that results from  the  Stage  1
energy balance based on an assigned heat loss.
This gas moves  to Stage 2 and contacts the
coal, where the following reactions occur:
       C+2H2-»CH4

       C + H2O -»• CO + H2
(1)

(2)
  In the gasifier model, several general assump-
tions  were made pertaining  to the thermo-
dynamics of the process:
  1. The  water-gas  shift  equilibrium  ade-
    quately  describes the gas composition
    from Stages 1 and 2.
  2. The enthalpies of mixing of the gases are
    zero.
  3. All  gas fugacities  and  activity coeffi-
    cients are unity.
  4. The effect of pressure on the heat of
    reaction is minor.
                                    III-4-3

-------
                      Table 1. ANALYSES OF COAL AND PRODUCT GAS
Coal analysis
Constituent
Carbon
Hydrogen
Oxygen
Nitrogen
Sulfur
Ash
Moisture
Total
H.V., gross
H.V., net
H.V., net,
MAP coal
Percent
70.1
4.9
6.6
1.4
3.0
12.7
1.3
100.0
1 2,800 Btu/lb
1 2,300 Btu/lb
14,350 Btu/lb
Fuel-gas analysis
Constituent
CH4
CO
C02
H2
H2O
N2
H2S
Total
H.V., gross, dry


Mo I percent
3.3
22.7
7.0
24.9
10.9
30.7
0.5
100.0
210 Btu/scf


Table 2. HEAT BALANCE, AIR-BLOWN GASIFIER3    Effects of Major Operating Variables
     Heat input
Heat output
Constituent
Coal, H.V., net
Coal, sensible
heat
Steam
Air
Char
Total
MBtu
1435.2

22.8
87.0
90.8
34.4
1670.2
Constituent
Gas, H.V., net
Gas, sensible
heat
Slag
Char
Loss
Total
MBtu
1332.1

263.4
12.4
58.6
3.7
1670.2
aBasis:  100 Ib MAF coal; 20 atm; exit gas at 1700°F;
       steam/coal ratio 1.06.

   Kinetic  equations for methane  yield and
carbon  oxide  formation,  together with
appropriate material  and  energy  balances,
were programmed for a digital computer. The
resultant simulation was used to obtain both
the  foregoing balances and the relationships
that follow.

III-4-4
  Using the mathematical model, the system
was tested for the effects of changes in steam-
coal ratio,  operating  pressure, and  gasifier
outlet temperature. The resulting influence of
these variables on  the heating value of  the
product gas, the producer-gas volume, and  the
thermal efficiency of the system is shown in
Figures 4, 5, and 6.

  The strongest effect is that of the gasifier
outlet temperature of the heating value of  the
product  gas. If a higher temperature is  re-
quired by a certain fuel-utilization system,  the
heating  value of the gas can be increased by
changing one or more of the other operating
variables.
  Although the system study presented here
is by no means exhaustive, it shows the opera-
tional flexibility of the gasifier. Thus, depend-
ing upon downstream use of the fuel, such
operating  variables as  steam-coal ratio,

-------
temperature, and pressure may be selected to
optimize costs, to provide for convenience of
operation, or to reflect downstream needs.
Sulfur Removal

   Because  sulfur  appears  in  the  gasifier
product gas as H2S, its removal is not a major
problem. Reduction of both gas temperature
and pressure, if required, can be accomplished
without undue penalty  to the system. Such
processes  as the  Benfield and iron oxide  fit
the system quite well.
Current Status

  For the  further  development  of an  air-
blown two-stage gasifier for commercial-scale
use, the next  step  is considered to be  the
operation of  the  system  at  the pilot-plant
level, A-technical evaluation and cost estimate
of  a  5-ton/hr pilot  plant was recently com-
pleted  with  the   assistance of  Koppers
Company,  Inc.,  as  an  outside engineering
subcontractor.  In  addition to  the basic gasi-
fier, the pilot plant would include coal prepa-
ration and handling  facilities,  gas desulfuriza-
tion  units,  effluent waste  treatment, and
other general facilities needed for the conver-
sion of coal to sulfur-free, low-Btu fuel gas.

  Cost evaluation is continuing. It has been
expanded in scope to include an estimate of
the cost of a commercial plant as well as tne
ultimate cost of the clean fuel  gas.
FUEL GAS UTILIZATION

   The clean,  desulfurized, high-pressure fuel
gas could be utilized through several alternate
routes. Perhaps the least efficient use of the
gasifier product would be its combustion in a
conventional steam boiler. Although the ash-
free quality of the gas would eliminate the
need  for  boiler-cleaning equipment,  and
freedom from sulfur would solve a major air
pollution problem,  it  is  doubtful if these
benefits  would ordinarily offset the coal con-
version process. It is conceivable, nevertheless,
that gasification of available coal could be the
best solution in a particularly vital  energy
situation. Low-pressure gasification could, for
example, supply fuel gas  directly to  a con-
ventional boiler. If more efficient gasification
would be attained at  higher  pressures,  the
gasifier   product gas  could  be  expanded
through a gas turbine before being burned.
   A more efficient use of the gasifier product
would be its  combustion  in a supercharged
boiler. With the addition of  a topping gas
turbine to utilize the work available  in the
high-pressure flue gas, the system shows some
advantages over conventional operation. The
cycle efficiency can  be improved  if, in addi-
tion to driving a combustion air compressor,
the topping turbine is coupled to a generator.
While  steam  remains  the major  generating
fluid, the gas turbine can add to the net gen-
eration and at the same time eliminate other
fan power requirements.

   Although present-day supercharged boilers
do not operate much above 5-7 atmospheres,
it is conceivable that higher pressures could be
utilized.  Little economic gain, however,  can
be realized  above 7 atmospheres.  Neverthe-
less, Figure  5 shows that the gasifier can be
operated  at  5  atmospheres without  appre-
ciably  affecting gas  quality  or thermal effi-
ciency.


Combined Cycles

   Over the past several years, many examples
of  conceptual combined  cycles  have been
studied.  A  recent paper5  by  Robson and
Giramonti evaluates  a number of combined
cycles and establishes a timetable based on gas
turbine   development.   Known  by various
names, most  cycles  have in  common  the
major generation of power by  a gas turbine,
which  then  exhausts  into  a steam boiler at
essentially atmospheric  pressure. Net genera-
tion includes that obtained from a  steam
turbine.
                                     I1I-4-5

-------
   There are many  variations  on the  com-
 bined-cycle concept. Some publicized  cycles
 limit themselves  to  current commercial gas-
 turbine  inlet  temperatures,  approximately
 1800°F.  Others,  however, expressing firm
 belief in  advancing  gas  turbine technology,
 assume  higher  temperatures  in  order  to
 demonstrate  the  potential of the combined
 cycle.

   Direct coal-fired gas turbines, the object of
 some past and current research,  are faced with
 a major problem  area: erosion  and corrosion
 from coal  ash and certain ash  constitutents.
 The coal gasifier, on the  other hand, because
 it produces a clean product gas, is an excellent
 means for using coal as a source of energy for
 the combined cycle.


  Figure  7  shows a  conceptual  1000-MW
 combined-cycle  power  generating  system.
 This  system,  conceived6 by  Curtiss-Wright
 Corporation,  uses  the BCR two-stage gasifier
 to generate fuel  gas at 22 atmospheres and
 2000°F.  (Note that the cycle incorporates
 liquid-metal  heat  exchange to reclaim  the
 sensible heat in the product gas before it is
 cleaned.)  Further,  the  gas   turbine  inlet
 temperature is shown as  3196°F. While both
 of these innovations are not considered to be
 currently  commercial,  they are not believed
 to be beyond grasp. Sufficient incentive for
 their development may  be fostered by  the
 promise that the combined cycle holds.
  The  Curtiss-Wright  plan  shows  that  a
system heat rate of 6740 Btu/KWH can be
obtained using 12,800-Btu/lb coal. Although
this gain in cycle efficiency may be somewhat
optimistic,  because  it  assumes  successful
development  of the innovations mentioned
above, it does indicate that the combined-
cycle concept deserves further attention.

  With  gasification and gas  cleanup at  ele-
vated pressures, the size and possibly the cost
of the gas producer would be significantly less
than with atmospheric pressure systems. In
III-4-6
addition,  the power cycle's increased  effi-
ciency, coupled with the air pollution control
aspects of the system, should help to supply a
solution to the growing need for clean power.

  Engineering development of the BCR two-
stage  gasifier for the production  of pipeline
gas is entering the pilot-plant stage. Current
projections indicate that a commercial plant
can become a reality in about  10 years.
  In the meantime, it is believed  that use of
the two-stage system as a producer-gas gen-
erator for inplant production of clean fuel gas
from  coal  can be realized  earlier. Much will
depend  upon  the  incentive  provided  by
current  evaluation of the types of end uses
described.

  New power-generation  systems will  be
developed to help meet the growing require-
ments  of  our energy-consuming way of life.
Coal,  transformed into a clean, non-polluting
fuel gas, can continue to help fuel those new
systems.
ACKNOWLEDGEMENT

  Parts of  this  paper  are  based on  work
carried out at Bituminous Coal Research, Inc.,
with support  from the  Office of Coal Re-
search,  U.S.  Department  of  the  Interior,
under Contract No. 14-01-0001-324.
  The assistance of  Dr. R. L.  Zahradnik,
Carnegie-Mellon University, and of Dr. E. E.
Donath,  Consultant,  is  gratefully acknowl-
edged.
BIBLIOGRAPHY

1. Donath,  E.  E. and R. A. Glenn. Pipeline
  Gas  From Coal by Two-Stage Entrained
  Gasification Proc. AGA Production Con-
  ference. Buffalo, New York, 1965.

2. Glenn, R. A. Progress in the Development
  of the Two-Stage Super-Pressure Entrained
  Gasifier.  Am.  Gas Assoc. Proc. Synthetic

-------
  Pipeline Gas  Symposium,  New York: p.
  67-75, 1966.
3. Glenn, R. A. and  R. J. Grace. An Inter-
  nally-Fired Process Development Unit for
  Gasification of  Coal  Under  Conditions
  Simulating Stage Two of  the  BCR Two-
  -Stage Super-Pressure Process. Presented at
  the  1968  AGA  Synthetic  Pipeline  Gas
  Symposium, Pittsburgh, Pa., 1968.
4. Zahradnik,  R. L. and  R.  A.  Glenn.  The
  Direct Methanation of Coal. ACS Div. Fuel
  Chem. Preprints 13 (4):  52-70,  1969.

5. Robson,  F.  L.  and  A.  H.  Giramonti.
  Advanced-Cycle Power Systems Utilizing
  Desulfurized Fuels. ACS Div.  Fuel Chem.
  Preprints 14(2): 79-96,  1970.
6. Curtiss-Wright  Corporation.   1000  MW
  Combined-Cycle Power Generating System.
  (Curtiss-Wright brochure).
                                                                              lll-4-l

-------
COAL

STEAM
                STAGE 2
GASIFIER
                                         CYCLONE
                         RECYCLE
                         SOLIDS
                                   OXYGEN
                                     AND
                                    STEAM
                                                                    GAS
                                                                PURIFICATION
                                                                    AND
                                                                METHANATION
                                                                 FINAL
                                                             ^PIPELINE GAS
                              SLAG
          Figure 1. Simplified flow diagram for two-stage super-pressure gasifier.




COAL
660 «F
100 Ib, MAF -^
14.8 IbASH
1.5 Ib H2O
GAS AIR-
*n/"\ft oc
1 /UU r
435 Ib

STAGE
2




N




s


*i r
HEAT
EXCHANGER


AIR
1700 °F
228 Ib













	 1^


— STEAM




^ 435 Ib. DRY GAS
| 6932 scf


v

STEAM
1700 °F
105 Ib















CHAR
1115 °F



SLAG
14 Ib
2924 °F
P = 20 atm
III-4-8
                      Figure 2.  Material balance for air-blown gasifier.

-------
                                                 99
COAL-
STEAM •
AIR
                                                 97
                 PRODUCT GAS
         STAGE 2

THE OUTLET STAGE 2 TEMPERA-
TURE IS CHOSEN.
KINETIC MODEL: METHANE AND
CARBON  OXIDE  YIELDS  FOR
STAGE  2 ARE CALCULATED
FROM  THE STAGE 1  PARTIAL
PRESSURES.  WATER-GAS SHIFT
EQUILIBRIUM IS THEN ASSUMED
AT STAGE 2 OUTLET TEMPERA-
TURE.
                                             < 0
                                             -u.  95
                                               ui
                                                  93
                                                                  T830°F
                                                                   20120F
                                                7400
                                                7300
                                    CHAR
                                  AND ASH
                                    CO O
                                    < o
         STAGE 1

ALL MATERIALS MIX THOROUGH-
LY AND  COME TO WATER-GAS
SHIFT EQUILIBRIUM.  THE MIX-
ING  TEMPERATURE  RESULTS
FROM THE SIMULTANEOUS SOL-
UTION OF THE ENERGY BALANCE
AND  THE EQUILIBRIUM  REAC-
TION.
                                             £•57100
                                             g«
                                                7000
                                                6900
                     ASH
                                                 210

                                             oc
                                             O
                                                 180
                                                   1.0
                                                      1.2
                        .1.4
                                                                                       1.6
 Figure 3.
 model.
   General scheme of mathematical
               STEAM/COAL
 Figure 4.  Effect of steam/coal ratio
(PV20atm).
                                                                                   I1I-4-9

-------
      99
      97
X ii
  Hi  95
      93
_     STEAM/COAL^
    7400i
    6900
     210
     200
     190
     180
        5            10           15           20
                GASIFIER PRESSURE, atm
        Figure 5.  Effect of operating pressure
        (T = 1830° C).
                                               1700     1800      1900      2000      2100
                                                    GASIFIER OUTLET TEMPERATURE, °F
                                                 Figure 6.  Effect of gasifier outlet
                                                 temperature (P = 20  atm).
III-4-10

-------
COAL-12,800 BTU/lb
    265 tons/hr
      GASIFIER
                F : 2012
            F=900
            = 318.5
            = 589
F : 975 F : 500
1 NO. 1 LIQUID MULTI- N0- 2 LIQUID
| WILT ALII LA [ 	 » fN ONF<5 ^^ MLTALMLAT >
1 EXCHANGER EXCHANGER
\ /
/ *
1 LIQUID
1 METAL H EAT — ,
1 EXCHANGER
F = 1374 I ,
A'R W-U72
1 *
1
GAS UQUID F -
< METAL HEAT <
EXCHANGER
i
F r 1374 3600 PSI -
P = 297
W = 749
F : 1000
P : 3600
W = 459
COMBUSTOR -|Fs3196
F : 276 F i 230
HEAT
EXCHANGER
140 ,,
1000 *F - 1000 RH -
STEAM J


1" Hg


" TURBINE I 1 «.-.,.-,.,»,«,,
///J COND.
— 6-66^)-^
 753 MW
                             F = 900
                             P = 318.5
         GENERATOR
COMPRESSOR
P/P - 22.0
                              W = 2220
                                   F = 1453
                                   P = 16.26
                                   W = 2371
  GAS
TURBINE
                                  AIR IN, F = 68
                                      W = 2061
                              F=717
                              P s 15.60
                            F = 500
                            P - 15.07
                              F:350
                              P = 14.85
 MAIN
BOILER
AUXILIARY
 BOILER
 EVAP-
ORATOR
                                                                                                                            STACK
                                                                                                 WATER
                                                          COOLING STEAM F = 600
                                                                         P : 400
                  753.0 MW GAS TURBINE
              ->•  253.6 MW STEAM TURBINE
                                                                                    F = TEMPERATURE, °F
                                                                                    P ; PRESSURE, psia
                                                                                    W = FLOW, Ib/sec
                 1006.6 MW TOTAL
                                    Figure 7.  1000-MW combined-cycle power generating system.

-------
                 5.  REMOVAL OF HYDROGEN SULFIDE
                     FROM SIMULATED PRODUCER GAS
                           AT ELEVATED TEMPERATURES
                                                    AND PRESSURES

                               F. G. SHULTZ
                           U. S.  Bureau of Mines
INTRODUCTION

  The  Morgantown Energy Research Center
of the U. S. Bureau of Mines has been investi-
gating the possibility  of removing hydrogen
sulfide  from  simulated producer gas near the
temperature  at  which the  gas  leaves  the
producer. The temperature range of interest is
1000 - 1500°F. Experimental work has been
limited to these temperatures and to pressures
between near zero and  100 psig.
Background

  We have previously reported1  that hydro-
gen sulfide can be removed from this type of
hot gas by passing the gas stream through an.
absorbent bed consisting of sintered pellets.
These pellets consist of mixtures of fly ash
from  a  coal-burning  power plant  at  Fort
Martin, W.  Va.,  and either technical-grade
ferric oxide or "red mud," a residue of the
aluminum  refining process. These  sintered
pellets were made by wetting the mixture of
the two materials, forming the mixture into
nominal 1/4-inch spheres (pellets), drying the
pellets, and then  firing them in an electric
furnace at 1900 - 2300°F for 10-15 minutes.
Temperatures required to sinter the materials
depended on the composition of the pellets and
were higher for a composition containing red
mud.  Sintering was  considered  satisfactory
when  the  pellets had been  fired enough that
the sharp edge of a knife blade would not cut
them, but not enough to vitrify the surface.
  The absorbent pellets used were previously
reported1 as having PbS absorption capacities
of approximately 12 grams of sulfur per ,100
grams of absorbent at 1000°F and 35 weight
percent sulfur at 1500°F, when used at space
velocities  of 2000, bed  depths of 15 inches,
bed diameters of 1 inch, and absorption pres-
sure near atmospheric. (Space velocity  is the
number of volumes of gas at standard condi-
tions passing through a unit volume of absorb-
ent per hour.)

  Various other  metallic oxides, including
those of zinc,  molybdenum, tin, cerium,
lanthanum,  and yttrium, were individually
sintered from mixtures of the oxide and fly
ash.  None of these oxides proved to be an
effective  hydrogen  sulfide absorbent  when
used in-this form.
Apparatus and Methods

  These absorbents were tested by passing a
simulated producer gas through a fixed bed of
the absorbent, contained in a stainless-steel
pipe, in an  electrically heated furnace. Figure
1 shows the basic test system. The test appa-
ratus consisted of compressed-gas cylinders
with individual rotameters for blending a gas
similar in composition to producer gas; the
furnace contains the absorbent and plumbing
                                      III-5-1

-------
suitable  for  passing the  gas through  the
absorbent.  Ports  allowed  for  analytical
determinations in absorber influent and efflu-
ent  gas  streams.   Influent  concentrations
normally used were 900 - 1000 grains H2S per
100 cu ft of gas (1.5 percent); effluent con-
centrations gradually increased from zero to
100 grains per  100 cu.  ft., at which concen-
tration the tests were terminated.
had  condensed  and  been deposited down-
stream of the absorber to increase both the
flow  resistance  and  the pressure  in  the
absorber. Tests made at equal pressures, with
and without tar  vapor in the gas, indicated
equal absorption  capacities (tests 110  and
112). These capacities were all determined at
gas-flow  rates  equal to  a space  velocity of
2000.
DISCUSSION
Effect of Pressure and Gas Velocity
   One unresolved question about the previ-
ous work was the effect that tar vapors, which
would be present in an actual producer gas,
would have on the ability of the abosrbent to
remove hydrogen sulfide. A  tar-vapor gen-
erator was built by partially filling an exter-
nally heated 2-inch-diameter cylindrical vessel
with  a bituminous coal tar made from Pitts-
burgh-seam coal  in the Morgantown Energy
Research Center's  flash  carbonizer. The
nitrogen from the flow  system was passed
over  the hot tar to pick up the vapors, then
transported through a heated line (900°F) to
the  bottom  of the  absorber, where it was
mixed with  the other producer-gas  constit-
uents and then  passed through the absorp-
tion bed.  Gas producers yield about 130,000
cu ft of gas per ton of coal,2  and an average
tar production of about 15 gal. per ton.3 This
is approximately 0.4 ml of liquid tar as vapor
per cu ft  of gas, or 12,500 ppm by weight.
Because nitrogen only was used to sweep  the
tar-vapor generator, the vapor generator was
found to  produce 0.8 ml of  tar vapor at
600°F. The vapor generator was operated at
600 - 700°F for test work.

   Table 1  shows that, at first,  the effect of
tar vapor seemed to be to increase the capac-
ity of the  absorbent. The reason for this in-
crease was unknown  at  the time.  In subse-
quent tests, Table 2,  it was  noted that  the
presence of tar vapors made little difference
in absorption  capacities  until sufficient  tar
III-5-2
  The effect of increased pressure on absorp-
tion was investigated to 10 psig, the limit of
the apparatus:  Figure 2 illustrates the effect
of increasing  pressure.  Within  this  pressure
range, higher absorption pressures caused an
increase in absorption capacity.
  Figure  3, relating the effect  of increasing
space velocity  to  absorbent capacity, shows
that a space velocity of 2000 is near the maxi-
mum that would  be permissible at  1000°F,
whereas at 1500°F removal was still effective
at a space velocity of 16,000. These capacities
were  all determined at  3-psig pressure in the
absorption bed.
Table  1. EFFECT  OF COAL-TAB VAPORS
    SULFUR ABSORPTION CAPACITIES OF
    SINTERED JAMAICAN RED MUD FROM
           REYNOLDS METALS CO.
ON



Test
No.
99
100
101
102a
103a
104a


Tpct
temp..
°F
1000
1250
1500
1500
1250
1000
Capacity of absorbent.
g sulfur/100 g absorbent

Absorbed
from H2S
10.7
27.6
44.1
56.5
37.5
12.0
Regenerated
from S02
12.1
24.6
40.6
52.3
34.0
10.9



sulfur
balance,%
-11.6
10.9
7.9
7.4
9.3
9.2
aTest gas contained tar vapors.

-------
         Table 2.  RESULTS OF H2S ABSORPTION TESTS USING ALCOA MOBILE RED MUD
Test
No.
105a
106
107a
108
109
110
111
112a
113
114
Absorption
Temp.,
°F
1500
1500
1250
1500
1500
1500
1500
1500
1500
1500
Pressure
psig
0
0
0
0
0
3
0
3
6
10
Weight of
sulfur absorbed
g sulfur/1 OOg absorbent
6.8
7.4
5.1
7.0
13.3
23.1
14.1
23.6
29.1
32.8
Weight of sulfur
regenerated
g sulfur/1 OOg absorbent
4.1
8.2
4.1
5.6
11.9
22.2
14.8
23.8
27.9
28.5
Error in
sulfur
balance
39.6
9.8
19.6
20.0
10.5
3.9
4.7
0.8
4.1
13.1
         aTest gas contained tar vapors.
Sintering Conditions

  The effect of  sintering  temperatures on
capacity was thought worthy of investigation.
It was found, however, that these absorbents
have  a fairly narrow  temperature  range  in
which sintering will occur within a reasonable
time period. Hurricane Creek (H. C.) Jamai-
can red mud will sinter in 2 hours at 2200°F,
and in 10 minutes at 2300°F, but it quickly
develops   a  vitreous  surface at  2400°F.
Materials incorporated into a fly ash mixture
generally sinter best at 1700  2000°F, but
again  a  temperature  range  of  only about
200°F  can  cause  the difference  between
excessively long sintering times and the rapid
formation of a glassy surface for any specific
composition.
Durability of Absorbents

   All previous work had been done using a
1-inch-diameter stainless steel pipe, type 304
or 316, as an absorption tube. This unit was
too small to permit easy unloading of absorb-
ent pellets. The inside walls of the pipe react
with the  gases used  in  this work, causing a
scale formation that gradually (but partially)
envelops the pellets in contact with the pipe
wall. After  several tests  the pellets could  be
removed only by  redding out the pipe  and
crushing the pellets.
  To  assess the  durability  of absorbents
better, a split-shell furnace with a bore suffi-
cient to hold a 2.5-inch-diameter  pipe  was
substituted  for the original furnace. In  the
1-inch-diameter unit,  some absorbents disin-
tegrated into dust, some fused into a mass,
and  others appeared fairly durable. The more
durable absorbents tested in the 1-inch unit
were: (1)  sintered Fort Martin fly ash, (2) a
sintered mixture of 25 percent Fe203 and 75
percent Fort Martin fly ash, and (3) a sintered
mixture of 40 percent H.C. Jamaican red mud
and  60 percent Fort Martin fly ash. Although
the  fly ash alone demonstrated good dura-
bility in the smaller unit, no additional work
was  done with it because it had relatively low
sulfur-absorption  capacity.  Both the other
absorbents were used in the larger diameter
absorber;  however, each was  tested through
six H2$ absorption-air regeneration cycles.
  Because the larger unit permitted  easy
unloading, weights of absorbents before  and
after testing could be determined. After test-
ing,  the red-mud-containing  absorbent  had
lost  3 percent of its weight, probably as dust;
10 percent of the pellets had  fused into  two
separate lumps; and a large amount of pellet
fracture had occurred. The Fe20e-containing
absorbent  did  not lose  weight, no fusion
                                    III-5-3

-------
between pellets had resulted, and only a small
amount  of  pellet  fracture was observed.
Absorption capacity was somewhat lower in
the larger diameter unit: only 16 weight per-
cent at a space velocity of 500, a temperature
of  1500°F, and  a  LO-inch bed  depth. The
reason for this lower absorption capacity is
being investigated.
Absorption at Higher Pressures

  The effect of higher pressures (up to 100
psig) was investigated in the 2.5-inch-diameter
unit with a 5-inch bed depth of the sintered
25 percent Fe2C>3-75 percent fly-ash absorb-
ent that had been used in ten previous  tests.
(Table 3 shows the results.)  Premixed gas in
pressurized  cylinders  was used  for the
100-psig test. Absorption tests were made at
1500°F and pressures of 15, 30, 45, 50, and
100  psig. Space velocities used  ranged  from
1100 to 4400. Capacities varied from 15.3 to
16.3 weight percent, except for a45-psigtest
in which flow was erratic. This capacity was a
20  percent  increase above that obtained at
1500°F and 3 psig with a 5-inch bed depth,
but no trend of capacity change as a function
of pressure was indicated in this region. The
effect of pressure  on absorption capacity for
this system  becomes asymptotic at 10 to 15
psig,  within the  region  of space  velocity
utilized.

  The 25-percent  Fe2C>3 and fly-ash absorb-
ent has been tested through  15  cycles of
absorption and  regeneration (Table 3):  the
only evidence of degradation has been some
fracturing of pellets and the fusion of 2  per-
cent of the pellet weight into one lump during
the series of absorptions at elevated pressure.
The maximum  temperature at which  this
absorbent is durable is 1500°F; at  1600°F,
fusion between pellets becomes severe.
           Table 3. H2S ABSORPTION CAPACITIES OBTAINED DURING 15 ABSORPTION-
                REGENERATION CYCLES WITH A 25 PERCENT Fe2O3-75 PERCENT
                              FLY-ASH SINTERED ABSORBENT3
Absorption
temp.,
°F
1000
1500
1250
1500
1250
1000
1000
1250
1500
1000
1500
1500
1500
1500
1500
Gas
flow,
scfh
15
15
15
15
15
15
15
15
15
15
30
45
40
66
16.5
H2S
g/1 00 ft3
1030
890
900
960
910
945
990
1080
1040
1010
1100
1050
1050
1050
393
Absorption
pressure,
psig
3
3
3
3
3
3
3
3
3
3
15
30
45
50
1000
Space
velocity
2000
2000
2000
2000
2000
2000
1000
1000
1000
1000
2000
3000
2670
4400
1100
Duration
of test,
hr
1.5
3.5
2.5
3.5
3.0
2.25
3.25
4.33
6.5
4.75
4.0
2.5
2.25
1.75
18.75
Sulfur
absorbed,
wt %
6.2
12.0
9.0
13.5
11.0
8.4
6.1
8.9
12.8
9.1
16.3
15.3
12.2
15.7
15.8
           aFirst 6 tests made with 2.5-inch bed depth; last 9 tests, with 5-inch.
III-5-4

-------
Regeneration of Absorbent

  Regeneration of these absorbents with air
has been found to proceed best when initiated
at about 1200°F. Sulfur dioxide is liberated
in concentrations of 5-8 volume percent until
about 60-70 percent of the absorbed sulfur is
burned  off at air-input space velocities  of
2000. The  concentration of  sulfur dioxide
then rapidly declines but never becomes zero,
even after a regeneration time of 80 hours. An
SO 2  concentration  of  about  50 pprn  is
reached after 24 hours of regeneration, and
the absorbent is then  reused for H2S absorp-
tion. During regeneration a small quantity of
elemental sulfur is also liberated.
  Water vapor is formed during the absorp-
tion step  and condenses in lines downstream
of the absorber. No doubt some 803  forms
during regeneration and combines with this
condensed water, because sulfuric acid  has
been identified in the effluent system.
CONCLUSIONS

   Conclusions  deduced  from work to date
follow:
   1. With the type of sintered absorbent de-
     scribed here, H^S can be absorbed from
     a producer gas at near the exit tempera-
    ture of the gas, thus preserving the sensi-
    ble heat of the gas for producing work.
  2. Absorbents  containing more than  37
    percent  Fe2C>3  degrade  more  rapidly
    than  those  with lesser  amounts  of
  3. Degradation  consists  of either  pellet
     fusion or disintegration into dust.
  4. The maximum temperature at which this
     absorbent can be used is 1500°F.
  5. Tar vapors in the producer gas do not
     adversely affect  the  capacity  of the
     absorbent for H2S during the absorption
     time intervals used in this work.
  6. Absorptive capacity increases with pres-
     sures up to 15 psig.
BIBLIOGRAPHY

l.Shultz,  F.  G. and  J. S. Berber. JAPCA.
  20(2): 93-96, February 1970.

2. Lowry,  H.  H. Chemistry of  Coal Utiliza-
  tion,  Vol.  II. National Research Council
  Committee. New York, John Wiley & Sons,
  Inc.: 1868 p., 1945.

3. Walters, J. G., C. Ortuglio, and J. Giaenzer.
  BuMinesBull. 643: 91 p. 1967.
                                                                                  III-5-5

-------
                          GAS SAMPLE
      H20 OUT
                                        VENT
 1
H20 IN
J-.ABSORBER  DRAIN
                         f
           TOS02
       ABSORPTION TRAIN
 MANOMETER
            ,*,* *
            i^V

                          FURNACE
                          CONTROL
     TUBULAR GLOBAR
  3\ TYPE FURNACE

   N\BSORBENT
                                   0-30" H20
        AIR, 100 psig
      'ROT. 10-100 scfh
                 SAMPLE
      sj)nxh
                                              12345
                                             PRESSURE IN ABSORPTION BED, psig
                                       Figure 2. Sulfur absorption capacity
                                       as a function of absorption-bed
                                       pressure for Alcoa mobile red mud.
                                        401	—
                                                                                         To
                           DRIP LEG
       ROT.
     ,0.3-3.0
    FM scfh
      ROT.
      1 -16
     kscfh
   ROT.
 3.06 - 0.6
Fl } scfh  f p|


 I FILTER
      )CO^
     (CO
  |HoS
             26%
                      1.5%
                     17%
                                         2000
    Figure 1.  Flowsheet for removal of sulfur
    from hot producer gas.
                                                             epop_      10000
                                                            SPACE VELOCITY
                                                           14000
                                      Figure 3-  Sulfur absorption capacity
                                      as a function of space velocity for
                                      Reynolds H. C. Jamaican red mud.
III-5-6

-------
               6.   GASIFICATION OF SOLID PARTICLES
                                             CONTAINING CARBON

                        C. Y. WEN AND  S. C. WANG
                           West Virginia  University
INTRODUCTION

  Fluidized-bed gasifiers have been shown to
possess technical and economical advantages
over conventional  coal-gasifying devices in
solving air pollution problems resulting from
the use of high-sulfur coal. However, the com-
plex phenomena of chemical reaction, heat
and mass transfer on solid particles, and the
interaction  of gas and solid flow patterns
within the  fluidized  bed  have  hampered
design  development. Researchers  in recent
years have, however, provided insight into the
rate of processes occurring on a single particle
and with the phenomena associating bubbles
in fluidized  beds. It is the purpose  of  this
paper to present the method by which realis-
tic  mathematical model describing a single-
particle phenomenon can be formulated  and
how such information  can be  incorporated
into a fluidized-bed model that accounts for
the bubble growth and coalescence in order to
arrive at a more reliable design for fluidized-
bed gasifiers.
SINGLE-PARTICLE SOLID-GAS  REAC-
TION MODELS

  The reaction  models  for  single-particle
solid-gas reaction  systems  can be generally
divided into three categories: (1) unreacted-
core-shrinking  model,  (2)  homogeneous
model, and (3) zone-reaction model.

  If the porosity of the  unreacted solid is
very small  and the solid is practically impervi-
ous  to gaseous reactants,  the reaction will
occur at the interface between the unreacted
solid and  the  porous-product layer.  Under
such conditions the unreacted-core-shrinking
model is applicable. Figure 1 shows the solid
and gaseous reactant concentration profiles
for this model. This model is also applicable
when the chemical reaction rate is very rapid
in comparison with the  diffusion rate of the
gaseous reactant. The zone of reaction in such
a case is  narrowly confined to the interface
between the unreacted solid and the product
solid.

  On  the other hand, if the  solid is porous
enough for the gaseous reactant  to diffuse
freely into the  interior of  the  solid,  the
unreacted-core-shrinking model is no longer
applicable. In such a case the reaction may be
said to occur, in a macroscopic sense, homo-
geneously throughout the solid to produce a
gradual and uniform variation in solid-react-
ant concentration  in the  particle.  The
homogeneous  model describes this situation
closely.

  In most actual cases, however, the solid-gas
reactions cannot be completely  represented
by  either of the  first two models mentioned
above, which represent  two extreme cases in
the solid-gas reaction systems. The majority
of  the  solid-gas  reactions probably  fall
between these two extremes.
  The  zone-reaction  model1-2  is a  more
general model for solid-gas reactions because
it  displays an  intermediate  between  the
unreacted-core-shrinking and homogeneous
models, as shown schematically in Figure 2.
In  fact,  the  unreacted-core-shrinking  and
homogeneous models can be derived as special
cases from the zone-reaction model.3
                                      III-6-1

-------
   In the  zone-reaction  model, the  kinetic
behaviors  depend  on  the solid structures.
Figure 2 also shows three types of solid struc-
tures:
   1. Volumetric  reaction model:  the  solid
     particle is composed of small grains of
     solid  reactaiit  embedded  in  inert  ma-
     terial.
   2. Grain  model:  the solid particle is com-
     posed  of small grains, each  of which
     reacts  according  to  the unreacted-core-
     shrinking model.
   3. Pore model: the reaction occurs at the
     surface of the pores.

   The effects of temperature on the reaction
rate and  on the concentration profiles of
gaseous reactant (A) and solid reactant (S) are
summarized schematically in Figure 3. In the
lowest temperature  region, V,  of  the zone-
reaction  model,  the concentration distribu-
tion of both gaseous and solid reactants are
uniform—a  limiting  case in which the homo-
geneous model is applicable. The reaction rate
in  this temperature region is  controlled by
that of  individual  grains or  pores.  In  the
highest temperature region,  I,  on  the other
hand,  the  gaseous  reactant concentration at
the  boundary between the reaction zone and
the  product layer is practically zero, so that
the reaction rate is controlled by the diffusion
of gas through the product layer.

   Note that in temperature regions I and III
of  the  zone-reaction  model,  the  gaseous
reactant  concentration profiles in the product
layer are similar to  those for the unreacted-
core-shrinking model as shown in Figure 3(b).

   Because  the  unreacted-core-shrinking
model has  wide application for many investi-
gators, and because the  mathematical treat-
ment is relatively simple  in comparison with
the zone-reaction model,  a brief discussion of
the  unreacted-core-shrinking model is given
below, with emphasis on the effects of heat
and mass transfer on reaction rate in terms of
the  effectiveness factor. The detailed mathe-
matical  treatments  can be  found  else-
where.3 >4'5
III-6-2
  The effectiveness factor is defined as  fol-
lows:

        actual (overall) reaction rate	
    s    reaction  rate obtainable  when  the
         reaction site  is exposed to the gas
         concentration and  temperature  of
         the bulk-gas phase

Because  the denominator  is a  constant, the
effectiveness factor is, in effect, a dimension-
less surface-reaction rate.

  Figures 4 and 5 show plots of effectiveness
factor versus fractional solid reactant conver-
sion,  and compare unsteady-state heat trans-
fer  solution to that of pseudo-steady-state.
Figure 4 indicates a case in which the pseudo-
steady-state analysis  could lead to an erro-
neous  conclusion.  The  chemical-reaction-
controlling  region  would  never have been
realized  if  the pseudo-steady-state  analysis
had  been used. The unsteady-state analysis,
on the other hand, shows that chemical reac-
tion could be rate-controlling when the initial
temperature of the particle is sufficiently low.
Figure 5 depicts the effects of the heat capac-
ity of the unreacted core and the heat of reac-
tion on thermal instability. The occurrence of
thermal instability is less likely when the heat
capacity is high.

  Figure  6  illustrates experimental  results
that  were  computed on  the  basis  of  the
unreacted-core-shrinking model. The corre-
sponding  numerical  solutions (solid, dotted,
and dashed lines) are also shown in the figure
for  comparison.  The experiment  was per-
formed in a thermabalance (shown  in Figure
7) by burning a single solid sphere in a stream
of heated air. The solid pellet was prepared by
mixing  the desired  proportion  of  activated
charcoal with aluminum oxide, which serves
as an inert porous medium. Sodium  silicate
(waterglass), diluted with the proper amount
of water, was used as a binder in forming the
pellets.  The particle was heated at  700°C in
an inert atmosphere prior to the combustion
test  to remove moisture  and to ensure no
weight loss from inert solid during the test.

-------
"Extinction" occurred at about 85 percent of
solid conversion. Note that the unsteady-state
heat transfer analysis  (solid line)  describes
more closely the experimental result than the
pseudo-steady-state  analysis (dashed line) at
the initial and transition stages. The values of
the parameters  used  in the correlation are
independently estimated  from   empirical
correlations.
  Figure 8  is an example of a triangular dia-
gram  that  shows  the reaction  path of  an
endothermic reaction. The  corresponding
effectiveness factor versus  conversion plot is
shown in the Figure 8 inset. The main advan-
tage of a triangular diagram is that it shows
the magnitude of each resistance during the
history of a reaction.
  In  the following discussion we  consider
only the case in which solid particles react
with fluidizing  gas while  maintaining  their
original size because of the formation of inert
solid product. The gasification of  coal, the
roasting-of sulphide ores, and the reduction of
iron ores are examples of this situation. The
following stoichiometric equation can be used
to represent these reactions.

  aA(gas) + S(solid) ->• gaseous
               and/or solid products     (1)

  The proposed calculation method6 assumes
that solids  follow  the shrinking-core model
and that the  overall conversion  rate is con-
trolled  by a chemical reaction step; in the
unreacted-core-shrinking model, the  reaction
is confined  at the  surface of the core, which
recedes from the  outer surface  toward the
interior  of the   particle.  When   diffusion
through  the  product  layer becomes  rate-
controlling, or when other single-particle reac-
tion models are used, the conversion-versus-
time expressions must be changed accordingly
for our calculations.
  The reaction of a gaseous component by a
first-order  irreversible  reaction can  be given
as:6
     1
   47rrc2a
dNA
dt
  1
47rrc2
dNc
dt
                                        (2)
                                    where rc is the radius of unreacted core and
                                    kc is the rate constant for the reaction.

                                      When the  reaction is carried  out in the
                                    bulk-phase  reactant-gas concentration, CA,
                                    the  extent  of conversion, Xg, of a particle
                                    having radius R is given by:
                                          l=lJ£=l-(l  xs)i/3

                                    where time for complete conversion, T, is:
                                                                    (3)
                                                                           (4)
                                    When the resistances of the chemical reaction
                                    step and of diffusion through the ash layer are
                                    comparable,  the rate constant, kc, is replaced
                                    by k, defined by:
                                       _L=1 '
                                        k   kc
                                                            eA
                                                                           (5)
                                      Next, let us consider a reactor with a con-
                                    stant feed rate of both  solids and gas, the
                                    solids being of  uniform  size and complete
                                    mixing. Because the conversion of an individ-
                                    ual particle of solid depends on its length_pf
                                    stay in the reactor, the mean conversion, Xs,
                                    of the exit stream of solids is given by:
                                          1-XS=\     (l-Xs)E(t)dt      (6)
                                                 A=0

                                    where the exit age distribution function for a
                                    reactor of complete mixing is:
                                                      1    -tyf
                                               E(t)=r-e                 (7)
                                                      t

                                    When chemical reaction is the rate-controlling
                                    step in a shrinking-core particle, substitution
                                    of equations (3) and (7) into equation (6) and
                                    subsequent integration yield:
                                                           -   T
                                             6(i)3 [l-exp(-T/t)l
(8)
                                                                                    III-6-3

-------
Gas and Solid Flows Based on the Bubble As-
semblage Model

  Let  us  summarize  the essentials of the
bubble assemblage model7 as follows:

  A fluidized bed may be approximately
represented  by "N"  numbers of compart-
ments  in series. The height of each compart-
ment is equal to the size of each bubble at the
corresponding bed height. Each compartment
is considered to consist  of the bubble phase
and the emulsion phase.

  The bubble phase is assumed to consist of
spherical clouds;  the  cloud diameter.can  be
derived from:
                        2umf/emf
                      - umf/emf
                                      (9)
The size of the bubble diameter along the bed
height is approximated by:
 where do is the bubble diameter just above
 the distributor. The rising velocity of a bubble
 is given by:

       % = Uo-Umf+O.TlKgdb)1/2 (11)

 From an arithmetic average of bubble sizes,
 the height of the i-th compartment can be
 expressed as:
                       (2+m) J'1
                                     (12)
 where m = 1.4 ppdp (uo/umf).
  The voidage distribution assumes that up to
Ljnf, the voidage, e, is uniform, while above
Lmf, e increases linearly along the bed height.
Then  the  number of bubbles, n, in  the i-th
compartment is given by:
             _ 6St
           n ~
                         e — emf
                          1-emf
                                            Therefore, the volume of bubbles, clouds, and
                                            emulsion in the i-th compartment can be cal-
                                            culated, respectively:

          Vbi = n-(Ahj)
                                                                                (14)
        Vei = StAhi- Vbi- Vci
                                                                                 (16)
                                            The overall interchange coefficient of the gas
                                            between the bubble and the emulsion phase
                                            based on a unit volume of gas bubbles may be
                                            given by the following experimental relation:
                                                         (Kbe)b = 11/db
                                    (17)
                                            The  solid  interchange coefficient between
                                            both phases is assumed to be given by:

                                                   iv   \   -  7 (l-emfumf ub
                                                   (Kbe)bs -
                                            In order to describe the distribution of solids,
                                            let us define, for convenience, the following
                                            quantities:
                                            7C   =  volume of solids dispersed, in clouds
                                                    _ and wake _
                                                          volume of bubbles

                                            7e   =  volume of solids in emulsion _
                                                          volume of bubbles
                                            The upward motion of the solids, as a part of
                                            the wake of the bubbles rising from  the  i-th
                                            compartment to the (i+l)-th  compartment,
                                            sets  up a circulation in  the bed with down-
                                            ward  movement of solids in  the emulsion
                                            phase from  the (i+l)-th  compartment to the
                                            i-th compartment. When  the solids are fed to
                                            the bottom of the bed  at  a constant volu-
                                            metric flow rate, W, the total upward  flow
                                            rate, Wj,, from the i-th compartment is given
                                                                                  <,„
The total downward flow rate, We, from the
(i+l)-th compartment to the i-th compart-
ment is given by:
 III-6-4

-------
CAS

Cgbi
           >   C-      •-  I  I  l    "'I    V—W
           >t - sbi
If the solids are fed at the top of the bed and
withdrawn from the bottom, W must be  re-
placed by -W in the above equations. These
relations  mentioned above are  shown sche-
matically in Figures 9 and 10.
NOMENCLATURE

   The following  list of terms defines expres-
sions used throughout this paper, in both the
text and the figures.

Term  Definition

a      Stoichiometric coefficient
•^     CAo^eA(To)Cpe/aCSoke>  ra^°  °^
         mass to thermal diffusivities in ash
         layer
CA    Concentration of gaseous reactant A
         in ash layer, C^c
^Ac   Concentration of gaseous reactant A
         to unreacted-core surface
C?Am  Concentration of gaseous reactant A
         at boundary between reaction and
         diffusion zones
       Concentration of gaseous reactant A
         in bulk gas phase
       Concentration of gaseous reactant A
         at   outer  surface  of  particle,
         mole/L3
       Gas-phase concentration in the bubble
         in the i-th compartment
       Gas-phase concentration in the emul-
         sion in the i-th compartment
       Heat  capacity  of  unreacted core,
         H/MT
       Volumetric heat capacity of ash layer,
         H/L3T
       Concentration of solid reactant S
       Initial  concentration of solid reactant
         S
Cgs    Concentration of solid reactant S  at
         outer surface  of particle, mole/L3
db     Bubble diameter, cm
dy    db in the i-th compartment, cm
                                              ED

                                              Eks
                                              Ekv
                                              g
                                              G
                                              Ah
                                              H
                                              AH
       Diameter of cloud, cm
       Bubble  diameter  just above the dis-
         tributor, cm
       Particle diameter, cm
       Effective diffusivity of component A
         in ash layer, L2 /0
       Activation  energy  of reaction  rate
         constant
       Apparent  activation  energy,  H/mole
       Activation energy  in  temperature
         range where diffusion  is controlling
       1/2 Ekv
       Activation energy  in  temperature
         range where reaction rate is control-
         ling
       Acceleration of gravity, cm/s2
       pcCpcT0/aCs0(-AH), ratio of enthal-
         py  of unreacted core to heat of
         reaction
       Distance from the distributor, cm
       Convective heat transfer coefficient,
         H/L20T
       Height of the i-th compartment, cm
       Radiational heat transfer coefficient,
         H/L20T4
                            -1
                                                         
-------
M
                                             W
NRe
'm

R
t
T
ub
ubi
ue
umf
uo

Vbi

Vci

Vei
   Number of bubbles  in i-th compart-
     ment; also order  of reaction  for
     gaseous reactant
N3Gramme moles of A and B
           modified Nusselt  number of
     convective heat transfer
           /^e> modified Nusselt number
     for radiational heat transfer
   Reynolds number
   RkmA(T0)/DeA(T0)>  modified  Sher-
     wood number
   Distance from the center of sphere (or
     from solid  surface) to  reaction
     surface
   Radius of unreacted core, L
   r at boundary between reaction zone
     and diffusion zone, L
   Radius of a particle, cm
   Gas constant, H/mole T
   Cross sectional area of bubble phase in
     the i-th compartment, cm2
   Cross sectional area of the bed, cm2
   Mean residence time of particles, s
   Temperature
   Temperature at unreacted-core surface
   Initial particle temperature
   Temperature at bulk gas phase
   Temperature at outer surface of parti-
     cle
   Velocity of rising bubble, cm/s
   Velocity of rising bubble  in the i-th
     compartment, cm/s
   Velocity of  bubble  with  respect to
     emulsion ahead of it, cm/s
   Velocity of emulsion phase, cm/s
   Minimum fluidizing velocity, cm/s
   Superficial gas velocity, cm/s
   TC/TO
   Volume of  bubble phase  in the i-th
     compartment, cm3
   Volume of cloud region  in the i-th
     compartment, cm3
   Volume of emulsion  phase in the i-th
     compartment, cm3
                                             W
                                              ei
                                             X
                                             13
                                             €
                                             emf
       Volumetric feed and outflow rate of
         solids, cm3
       Volumetric upward flow rate from the
         i-th compartment, cm3/s
       Volumetric downward flow rate from
         the i-th compartment, cm3/s
       Fractional conversion of solid reactant
         S
       Mean fractional conversion of solid
                                             PC
                                             Pp
       Void fraction in a bed as a whole
       Void fraction in  a bed at minimum
         fluidization
       Effectiveness factor based on surface
         reaction
       Effectiveness factor based on volume
         reaction
       Density of unreacted core, M/L3
       Density of solid, g/cm3
       a^g0)CA/DeA(T0), modified
         Thiele modulus
       Time for complete  conversion of a
         single particle, s
BIBLIOGRAPHY

1. Ishida, M. and C. Y. Wen. Chem. Eng. Sci.
  26:1031,1971.

2. Ishida, M., C. Y. Wen and T. Shirai. Chem.
  Eng. Sci. 26:1043, 1971.

3. Wen, C.  Y.  Ind.  Eng. Chem.  60(9): 34,
  1968.

4. Wen, C. Y. and  S.  C. Wang. Ind. Eng.
  Chem. 62(8):30, 1970.

5. Yagi, S. and D. Kunii.  Fifth International
  Symposium on  Combustion, p. 231. Rein-
  hold, New York, 1955.

6. Yoshida, K. and C. Y. Wen.  Chem. Eng.
  Sci. 25:1395, 1970.

7. Kato, K. and C. Y. Wen. Chem. Eng. Sci.
  24:1351, 1969.
III-6-6

-------
                     SOLID
                     REACTANT
                                                                       VOLUMETRIC
                                                                      REACTION MODEL
Figure 1. Concentration profile for the
unreacted-eore-shrinking model.3
                                                Figure 2.  Concentration profile and
                                                solid structure in the particle.1
                                                                1/T
                                                                   - 1
    •'AO
           \
£c,
^
                    Ek  -.Efc/2
                      s    v
                     A°
^c
y.
                                   AO
                                                 AO

                                                              -
                                                                -r=5
                                                                AO
            .(a) ZONE-REACTION MODH.
                                                    |b) UNREACTED-CORE MODEL
Figure 3. Schematic diagram representing characteristic behavior of solid-gas reaction
systems under various temperature regions.1
                                                                               IH-6-7

-------
              j3= 0.005

              NSh-lOO
              RTn
                   25
              s ; 0.139
              (NNu'c  - 10
             n : 1

             A=0
             -EL = 0.95
             To
                                       100
                                        60

                                        40
20
                                                              I    I   I    I    I
   HEAT TRANSFER*:^

  "UNSTEADY STATE	 ** - .,,

   PSEUDO-STEADY-STATE       ""~*
                                        0.2
                                          0      0.2     0.4      0.6     0.8     1.0
                                             SOLID REACTANT CONVERSION, percent
Figure 5.  Effect of heat of reaction and heat capacity of unreacted core on thermal instability
in terms of G.4
III-6-8

-------
TO TRANSDUCER


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SOLID CONVERSION, X
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Figure 6.  Comparison of experjmental result
and computer simulation for oxidation of
carbon in porous solid, using unreacted-
core-shrinking model.4
                    Q
                  ?**""»<   GAS
          Figure 7.  Thermobalance.
                                     IH-6-9

-------
                0.2    0.4    0.6    0.8

                 SOLID CONVERSION, X
                                                   db- - m+ do - 1-4Pp dp —
     0      0.2

     HEAT TRANSFER
       CONTROL
                   0.4
                                                                   Uo
\ SM  MASS TRANSFER
       CONTROL
   Figure 8.  Triangular diagram showing the    Figure 9.  Schematic diagram of i-th compart-
   reaction path and the magnitude of each      ment in the bubble-assemblage model for
   resistance in an endothermic reaction.4      fluidized bed.7
III-6-10

-------
              . . /?"7\ • • /i*"**. • /r" i^ •• /?"*A»* /T"*A* i*/5"VV * " ' '
             •••.- Q - Q;: Q/.'SQ v.Q:::©.% :.-
             iononnonnnnoonn
Figure 10. Main features of solid movement and gas flow as visualized.6
                                                II1-6-11

-------
       7.  MANUFACTURE  OF  ACTIVATED  CARBON
                                                    BY  GASIFICATION
                                               IN  A FLUIDIZED  BED


                                 A. A. GODEL
                           Societe Anonyme Activit
 INTRODUCTION

   Today, I propose abandoning momentarily
 the study of combustion problems in fluid-
 ized beds in order to discuss chemically the
 manufacture of activated carbon.

   Activated carbon  is  a most  powerful
 absorbent.  It  is  used  to purify  both air
 (removing malodourous gases)  and liquids
 (purifying water and blanching sugar.)

   The principle of the  manufacture of acti-
 vated carbon which concerns us here consists
 of subjecting  carbonaceous  matter (e.g.,
 charcoal, peat lignite, coal, and briquettes) to
 the action of steam at about 800 - 1000°C.

   In  fact,  this treatment represents a true
 gasification process involving selective oxida-
 tion of tarry matter contained in the carbo-
 naceous matter. It is sometimes considered-
 in a rather simplified manner—as a cleansing
 operation  for  existing  pores or, for  pores
 created by the operation within  the internal
 structure of the carbon, pores and capillaries
 which thus become "absorbent."

   In  1948, our factory at Vernon was  the
 first in France  to manufacture  activated
 carbon by fluidization, in  accordance with a
continuous-flowing thin layer.1 This process
was  later licensed to the French Societe
"Carbonisation & Charbons Actifs" which
applied  it in two industrial  furnaces in its
factory at Parentis (Landes).
NEW TECHNIQUE

  However,  it's not  about  this continuous
activation  technique which I should like to
speak today, but rather about an entirely new
technique, still  in  the  experimental  design
stage, that is not yet developed industrially.

  Contrary to the preceding process, the new
process  is characterized by  treating  carbo-
naceous matter in deep fluidized beds under
the action of pure stream at high temperature,
by discontinuous "batch process" operation.

  As  indicated  previously,  this  operation
results in the gasification of a more or less
important  fraction  of the raw material. To
obtain certain grades of activated carbon, it is
sufficient to  gasify a quarter or a third of this
raw material; for others, the process must be
carried further—to two-thirds or even three-
quarters.  The  manufacture  of activated
carbon is thus linked with gasification. Yield,
that is the ratio between the weight of the
activated carbon obtained and the weight of
raw material charged,  varies not only accord-
ing to the nature of the carbon obtained, but
also  according to  the  processing  methods
used:  the goal of the new process is to con-
siderably improve technical efficiency.
                                      III-7-1

-------
   In general,  the  production  of  activated
 carbon  and gasification result from  reactions
 of steam  at red  heat on carbon; these are
 pressure- and  temperature-dependent balan-
 cing  reactions situated  between  two  equa-
 tions:
 C + H2O = CO + H2                     (1)
                    (28.8 K calories  per mole)

 C+2 H2O = CO2 + 2 H2                  (2)

                   (14.8 K calories per mole)


   Equation (1) is dominant above  1000°C
 (1830°F),  requiring practically  no  excess
 steam;  equation  (2)  is dominant  below
 800°C  (1470°F),  requiring,  by comparison,
 steam consumption  significantly greater than
 the stoichiometric ratio. It is for  this reason
 that I have linked here the manufacture  of
 activated carbon with the gasification process.
 I  shall  limit my discussion,  however, to the
 manufacture of activated carbon and  shall
 simply  mention,, in  passing, possibilities for
 the  application  of  this  same  process  to
 integral fuel gasification.

   In this new technique, we operate  by batch
 process  and not by  thin-layer  fluidized-bed
 continuous flow:  we do  this to  avoid both
 activated carbon  degradation during process-
 ing by  bituminous gases which are given off,
 and the inconvenience which occurs inevita-
 bly during continuous fluidized flow due  to
 the mixture of crude and activated fractions.
   Having adopted the high-temperature batch
 process, in order to obtain the best  weighted
 yield and  best gas production, it is obvious
 that" we must operate in cyclic  phases. To
 obtain,  during one reaction phase, the impor-
 tant calorie transfer necessary to maintain the
 temperature at 1000°C (1830°F), in spite  of
 reaction endothermicity,  it  was particularly
 essential to operate by fluidization in order to
 make maximum  use of  the excellent  heat
 transfer ratio existing by contact with the
 internal heating surfaces. In the new process,
 these heating surfaces will be constituted by
III-7-2
fixed heat-accumulating masses preheated to
high temperature in a preceding phase.

  As you may note, this technique resembles
somewhat that of the production of water gas
in fixed  layer; however,  instead  of  using
treated  carbon for  the accumulation of calo-
ries,  which  would  only lead to insignificant
phase durations (a few minutes only), we use
important  fixed  heat-accumulating  masses
which permit extending phase durations to
more than half an hour  or an hour. Under
these circumstances, fluidized-bed tempera-
ture  may be made to oscillate, at will, around
the   average  reaction  design  temperature,
whether for equation (1)  or (2); although
each  absorbs  an unequal, but considerable,
quantity of heat,  the  oscillations  are of the
order of 1000°C maximum.

  It has always been difficult to achieve such
heat transfers by traditional processes; e.g., an
admixture  of flue  gas with steam at high
temperature, or the use of retorts heated from
the  outside  by  flames,  thus resulting, in
practice, in  too low temperatures  giving low
yield.

  The following figures will clarify the  ques-
tion  of  the required heat supply—the gasifica-
tion   of 12  (the  atomic weight)  grams of
carbon,  according to reaction 1 already men-
tioned  as  applicable  at  1000°C  (1830°F),
requires the following heat supply:
   1. For endothermicity, 28.8 calories.
  2. For  additional  superheating of the
     steam,  between 450 and  1000°C  (840
     and 1830°F), 6.3 calories.

  Although theoretical,  these  figures  show
clearly  that endothermicity dominates  and
that  it  is easy to  complete superheating of
steam from 450 to 1000°C (from 480 to
1830°F)  by  direct introduction  into the
fluidized bed. The  steam is produced by the
recuperating boiler and introduced into the
base of  the  reactor  at  precisely  450°C
(840°F). To work in cyclic phases, a produc-
tional unit will preferablly be provided with)

-------
 two reactors, either coupled or separate, but
 which operate in alternative phases: that is,
 one  gasification  phase  and  one  reheating
 phase of heat-accumulating refractory masses.
 It is, of course,  possible to have production
 units consisting of three or more reactors.

   Each  reactor  is  equipped  with piles of
 refractory material forming heat-accumulating
 masses which will be  placed in good contact
 with the fluidized bed. These masses will be
 chosen  on  the basis  of their refractoriness,
 their non-adherence  of clinkers, and  their
 good thermal conductivity.

   With  regard to  thermal  conductivity,
 certain carborundum  agglomerates which are
 highly  refractory  (1650     1700°C-3000-
 3100°F) have  thermal conductivity 5  or 10
 times greater than that of current refractory
 materials, or almost equal to that of stainless
 steel.
TWO EXAMPLES

   To illustrate the method of construction
and operation  of  the new process,  I  shall
describe two examples, although other combi-
nations may be envisaged  to  correspond to
various uses.
Twin-Reactor Furnace

   In the first example, represented by Figures
1  and 2, the two reactors  in  the  activated
carbon and  gasification production unit are
juxtaposed  in a  single rectangular furnace
divided into two identical, but distinct, parts
(A and B); each reactor (A and B) contains a
nest of  parallel, vertical,  heat-accumulating
refractory slabs (1). These slabs swell out at
the base (2), leaving only a narrow slot (3)
used  for introducing  reaction  steam;  the
fluidized bed is supported by this steam injec-
tion during the entire blowing period which
begins when carbonaceous material is charged
into reactor A at the beginning of the reheat-
ing phase. The  material is charged pneuma-
tically at  (4). reaching  a level slightly higher
than that of the slabs.

   At the end of the reheating phase, the steam
is stopped, thus provoking instantly the flow
of carbon by (3) in the hopper. The activated-
carbon product is evacuated pneumatically by
(5) toward storage (not shown).

   During  the reaction (or  gasification) phase
in reactor A, where  gas is produced, air is
introduced at (6') in reactor B; this gas thus
burns  in  the upper  chamber  of  reactor  B
which is then in reheating phase. The combus-
tion gases reheat the slabs (!') and escape  at
the base at (3'). They are finally  evacuated
from  the  reactor  by (7')  to  a recuperating
waste heat boiler (not shown) which produces
superheated  steam  at  450°C  (840°F) for
activation; this  superheated steam is intro-
duced in  reactor  A  at (8). Slab  thickness,
usually  greater than  their spacing, is calcu-
lated  so   as  to  obtain  required  calorific
capacity corresponding  to phase and reacting
duration.

   The set  of reactors A and B are put into
operation  quite simply:  when empty, they are
reheated simultaneously to service tempera-
ture by gas or fuel oil  burners  (9) and (9').
When  a   temperature of about  1200°C
(2190°F)  is reached in the reactors and when
the boiler is under pressure, the burners are
stopped and the set of reactors is placed  in
cyclic reversal operation by injecting, at the
beginning  of  the  phase,  alternately in  one
reactor  then in the  other, the carbonaceous
material to  be activated. The steam must be
injected from the beginning, in the lower part,
and during the active charging and gasification
operations.
Cylindrical Double-Reactor System

  In  the  second  example,  illustrated  by
Figures 3 and 4, the two reactors A and B are
separate and cylindrical, permitting the use of
                                     III-7-3

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 cyclone dust collectors for treating gases given
 off during reaction, and also operation under
 a pressure of several bars.

   The heat-accumulating mass placed in each
 reactor (A and B) is composed of a refractory
 block  (1), run through by a nest of 20 cm
 diameter tubes (2). These tubes narrow out at
 the base so that the orifice (3) at their ends
 may  let  pass  the  reaction  steam injected
 through (8); the steam backs up the fluidized
 bed  which has been previously charged into
 reactor A to a level slightly higher than that
 of the heat-accumulating block.

   Steam  blow-in  through (8) must begin in
 reactor A at the very beginning of the phase,
just  before  introducing  the carbonaceous
 material   pneumatically   through  (4);  it is
 stopped at the end  of the activation-reaction
 phase when  the carbon  flows down immedi-
 ately, to be evacuated pneumatically through
 (10) toward storage (not shown).

   Reaction-gas dust collecting takes place in a
 cyclone (12) in which the lightest particles are
 deposited; these  particles,  consisting  mainly
 of activated  carbon, may be either re injected
 into  reactor A  or collected directly (as
 shown).

   The gas thus  treated  is fed through (11)
 into a combustion chamber (15) where air is
 introduced through  (14); the  emitted flue gas
 is introduced through (16') to the bottom of
 reactor B and,  therefore, during the next
 phase through (16)  to reactor A; flue gas thus
 passes from the bottom  to the top of reactor
 B and  through  a bed  of finely granulated
 refractpry material—on  top  of  the nest of
 tubes in the accumulating  block—introduced
 pneumatically through (6') at the beginning
 of the reheating phase. Obviously, this  granu-
 lated material contributes no calories of itself,
 but constitutes an auxiliary fluidized bed that
 increases  the transmission  of calories to the
 walls of the nest tubes during reheating. The
 granular refractory  material may be a sulfur
 acceptor.
III-7-4
  The flue gas evacuated from reactor B passes
through  cyclone (13)  which retains the flue
dust eventually emitted from the refractory
material. Flue  gas  is finally evacuated  to  a
waste heat boiler, not shown.

  At the  end  of the reactor B  reheating
period, which coincides with the end of the
reactor  A activation  period, gasification is
stopped and  so is the  circulation of flue gas.
The  refractory  material  then  flows freely
down through the lower orifices of the tubes,
from where it is evacuated at the base through
(10') and conveyed pneumatically into a silo
(not shown)  where it  awaits transfer during
the  next reactor  A  heating  phase. When
reactor B is  reheated, it  is ready to receive
first  the  steam, and then the load  of carbo-
naceous material: the cycle thus continues.
System Operation

  To put  reactors A  and B  into operation,
procedure is as indicated in the first example
(Figures  1  and  2): they are  heated simulta-
neously by gas or fuel oil  burners at (9) and
(9').

  In  both  described  installations, air and
steam valves are  ordinary, but the valves for
evacuating  activated carbon, gas, and flue gas
are  water cooled.

  Valves operate automatically in accordance
with programming  which takes into account
heating  requirements. These  requirements
may  vary  considerably   according  to  the
nature of  the  carbons treated, or  rather,
according  to the nature of the ash of these
carbons:  if the ash is extremely fusible, it may
be evacuated in melted form, in which  c'ase, it
would be advisable to adopt  a high reaction-
temperature solution.  If the ash is not ex-
tremely fusible, temperature levels lower than
ash melting point will be used and the ash will
be  evacuated in powder form together with
the activated  carbon:  ash  will  then  be
separated,  using commonly  used  industrial
processes: e.g., dust screening and washing.

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  Coming back to the second type of installa-
tion which  I have just described (Figures 3
and  4), the following figures  are provisional
estimates.

  The inside  dimensions  of  the unit,  illu-
strated by Figure 4 and which corresponds to
a production capacity of 6 to 12 tons per day
(even 18 tons per day, for production varies
considerably with the quality of the activated
carbon needed) are:
   1. Section of each cylindrical reactor, 6 sq
     m.
   2. Inside diameter, 2.80 m.
   3. Height of the cylinder, 3.50 m.
   4. Total height of heat-accumulating refrac-
     tory material in each reactor, 1 m.
   5. Total volume of heat-accumulating re fac-
     tory material in each reactor, 4 cu m.

   Production of  superheated  steam by  the
waste heat  boiler is superabundant; for this
reason, excess gas can be tapped through (17).

   Operation under a pressure  of several bars
may be desirable, either in increase the pro-
duction of activated carbon or to improve the
purification of the gas in view of its use in gas
turbines.

   Another  way of increasing the production
of activated carbon and gas is to operate with
a highly expanded fluidized bed.

   In fact, it is conceivable that if the installa-
tion of  the reactors must  be adapted to
integral fuel gasification, the  compressed gas
produced,  perfectly desulphurized, might be
used in turbines for the production of energy
in a mixed- gas-steam  cycle, without atmos-
pheric pollution.
CONCLUSION

  I  must state,  however, that the process
which I have just described to you was specifi-
cally designed for the production of activated
carbon, with the  simple aim of being able to
adapt it eventually to integral fuel gasification
for the production  of synthesis gas, rich in
hydrogen and carbon  monoxide,  without
either  nitrogen or carbon dioxide.  The gas
produced, being of a highly reducing nature,
should be easily desulphurized by traditional
scrubbing methods or by absorption in dry
conditions  by sulfur acceptor, according to
new processes such  as Professeur  Squire's,
thus permitting effective recuperation of pure
sulfur.

  In  conclusion,  I should like to state that
activated carbon, the subject of  the manu-
facturing process  which I  have just described,
is one of the most precious auxiliaries for the
protection  against  both  atmospheric  and
water pollution agents.
BIBLIOGRAPHY

1. Chemical Engineering, July 1948.

2. Swiss patent No. 250, 891.

3. French patents No. 942, 699 and 951, 153.
                                                                                   III-7-5

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                                              SECTION BB
LEGEND
1.  REFRACTORY SLABS
2.  SLAB BASES
3-  INTER-SLAB SLOTS
4.  CARBONACEOUS
   MATERIAL INLET
5.  CARBON OUTLET
   (TO STORAGE)
6-  AIR INLET
7.  GAS OUTLET
   (TO BOILER)
8.  SUPERHEATED
   STEAM INLET
   (FROM BOILER)
9.  BURNER  (GAS
   OR FUEL OIL)
                                   Figure 1.  Twin-reactor furnace.
III-7-6

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KEY



STEAM •



GAS  -
       BOILER
1

REACTOR A
8]
t ^L
i &


1..- "
REACTOR B
a';; i /
                        Figure 2.  Twin-reactor furnace, flow sheet.
                  Figure 3.  Cylindrical double-reactor system, flow sheet.
                                                                                    111-7-7

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-ij
00
                                   REFRACTORY BLOCK
                                   TUBE NEST
                                   ORIFICES
                                   CARBONACEOUS
                                   MATERIAL INLET
                                   (NOT USED)
                                   GRANULATED REFRACTORY
                                   MATERIAL INLET
                                   (NOT USED)
STEAM INLET
BURNER (GAS OR FUEL OIL)
CARBON OUTLET (TO STORAGE)
TREATED GAS FROM CYCLONE
CYCLONE (REACTOR A)
CYCLONE (REACTOR B)
AIR INLET
COMBUSTION CHAMBER
GAS TO REACTOR
EXCESS GAS TAP
                                           Figure 4.  Cylindrical double-reactor system.

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SESSION IV:




        Conceptual Design and Economic Feasibility
SESSION CHAIRMAN:



        Dr. D.H. Archer, Westinghouse

-------
                                   1.  CLEAN POWER  SYSTEMS
                                                                      USING
                               FLUIDIZED-BED COMBUSTION

                               A. M. SQUIRES
         The  City College of The City University of New York
INTRODUCTION

  A major opportunity for clean power from
coal involves the combustion of coal at high
pressure in presence of a desulfurizing agent,
and  the generation of power by a combina-
tion of gas- and steam-turbine cycles.

  This paper reviews briefly three candidate
combustion  technologies for use in the fore-
going combination, and considers power cycle
arrangements suited  for each of  the three
approaches to combustion.
HIGH-PRESSURE SULFUR CAPTURE

  In a system for combustion at high pres-
sure, an  agent derived from  dolomite  may
advantageously capture sulfur during a first
coal-processing step. For this first step there
are three cases to consider.
  1. Complete combustion, using air in excess
    of stoichiometric.
  2. Partial combustion, using between about
    a third and half of the stoichiometric air
    and  yielding a fuel gas containing CO
    andH2.
  3. A carbonization yielding low-sulfur coke
    as well  as fuel gas, with heat for the
    carbonization  supplied  by  a  partial
    combustion  which  consumes  on  the
    order of 10 percent of the  stoichiometric
    air.
gaseous sulfur  species may be absorbed by
agents derived from dolomite:
  Work is underway at The City College on
the  kinetics  of three  reactions  whereby
   [CaCO3+MgO] + SO2 + 0.5 O2 =
       [CaSO4+MgO] + CO2

   [CaCO3+MgO] + H2S = [CaS+MgO]
         H2O + CO2

   [CaO+MgO] +H2S= [CaS+MgO] +
             H2O
(1)


(2)


(3)
Our primary interest is in the application of
the reactions in systems where the solid  is
used cyclically, a step for desorption of H^S
from the solid  being present in  the cycle,
thus:
   [CaS+MgO] + H2O + CO2 =
       [CaC03+MgO] + H2S
(4)
The equilibrium for reaction (4), the reverse
of reaction (2), is a strong function of temper-
ature; formation of H2S is favored by a lower
temperature. To obtain a gas containing H2S
at a  concentration suitable  for economic
conversion to sulfur in a Claus system, reac-
tion (4) should be conducted at a temperature
in the general vicinity of 600°C and at ele-
vated pressure.

  This paper is confined to a consideration of
each of the three aforementioned combustion
technologies in a  version  which employs an
appropriate reaction selected  from  (1), (2),
and (3) for the capture of sulfur, along with
reaction (4) for the release of sulfur from the
solid.
                                     IV-1-1

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Complete Combustion Approach

  Hoy and Stanton1  have described a large-
scale  experiment for fluidized-bed combus-
tion of  coal at 6  atmospheres and 800°C.
They  propose a pressure  of 15  atmospheres
for the commercial  embodiment of this tech-
nique. About 70 percent of the heating value
of coal would be transferred to boiler tubing
within the bed;  the remaining 30  percent
would be available  as  sensible  heat in  the
800°C effluent gas, which would be expanded
in a gas turbine. The authors hope that alkali
and dust content of the gas will be acceptable
for its direct use in the turbine, so that equip-
ment  need not be provided to cool  the gas,
scrub  it of alkali dust, and reheat the gas to a
suitable turbine-inlet temperature.

  The CC>2 partial  presure in gas leaving the
complete-combustion  fluidized bed  is suffi-
ciently great that CaCC>3  in dolomite added
to  the   bed  would not  decompose.  Half-
calcined  dolomite  present in  the bed would
react  readily with  SC»2, as in reaction (1).
Kinetics  for   this  reaction,  apparently on
account  of half-calcined dolomite's porosity,
are favorable at temperatures even as low as
600°C.2'3'4
  The team at The City College has demon-
strated  that  a way  exists   to regenerate
[CaCOs+MgO] from  [CaSC>4+MgO]  while
liberating H2S for sulfur manufacture.4 The
sulfate would  be reduced  to sulfide at about
800°C, thus:
   [CaSO4+MgO] +4H2 =
       [CaS+MgO] +4H20
(5)
The sulfide would be converted to the carbo-
nate  by reaction with  steam  and carbon
dioxide at  about 600°C,  as in reaction (4).
Reactions (5) and (4), as well as reaction (1),
would be conducted at elevated pressure, suita-
bly the 15  atmospheres which Hoy and
Stanton proposed.

  In our recent paper,4 we offered a scheme
employing methane reforming as a source of
IV-1-2
       hydrogen for reaction (5). In the context of
       Hoy  and Stanton's proposal,  reaction  (5)
       might advantageously be carried out in a small
       auxiliary fluidized bed operating under reduc-
       ing conditions.  Solid would be  transferred
       from  the main  fluidizerf-bed boiler  to  the
       auxiliary bed  and thence to a  fluidized bed
       for sulfur desorption by means of reaction
       (4).

          It is important to note that the maximum
       temperature for the dolomite-derived solids
       would be 800°C  in this  scheme.  There is
       evidence5 that [CaO+MgO]  does not sinter
       and lose surface area at this temperature. We
       are examining  the question of the sintering of
       other  dolomite-derived  solids, and have
       grounds to hope that none of the solids will
       undergo serious sintering under this complete
       combustion approach. We would expect solid
       reactivity to remain excellent through a large
       number of cycles. The practical importance of
       this fact will,  of course, depend upon one's
       degree of success in separating  the dolomite-
       derived solid from coal ash to be discarded.
Partial Combustion Approach

  If partial combustion of  coal were  con-
ducted  in  a fluidized bed of  fully-calcined
dolomite, sulfur would be absorbed by reac-
tion (3). Reaction (4) could then be used to
desorb sulfur. The resulting carbonated solid,
[CaCO3+MgO],  could be  returned  to  the
partial  combustion  bed,  to  be  calcined
therein.

  If all of the carbon in coal is reacted with
oxygen  in air to form CO, a surplus of heat is
developed. The surplus heat could be used to
raise steam in boiler tubes passing through the
partial combustion bed or to heat combustion
air. Alternatively, heat from the formation of
CO could be used to support the endothermic
reaction of steam with carbon, steam being
supplied to the bed  along with air. Adding
steam to the bed would, of course, increase
the latent  heat loss from the overall  power-

-------
generating system. This is not quite as bad as
might appear at first glance. Steam supplied
to the partial combustion step will eventually
reappear as steam at the gas turbine,  and the
total flow of gas through the turbine will be
larger than if steam had not been supplied to
the partial  combustion step. This means  a
larger net production  of power from the gas
turbine,  and the cost of the gas  turbine,
including air compressor, will  be significantly
less than  if steam were not used. This reduc-
tion in  capital cost will provide an economic
offset  against the  reduction in efficiency
chargeable to the latent heat loss. The higher
the gas-turbine inlet  temperature and  the
higher the pressure, the less advantageous it
will be  to use the excess  heat from partial
combustion to raise steam  for a steam cycle
and the more advantageous to use this heat to
decompose steam by reaction with carbon.

  Partial combustion has potential advantages
over complete combustion:
   1. About half as much gas would  be pro-
    duced, and so the fluidized bed would be
    smaller.
  2. The  gas,' containing potential energy as
    well as  sensible heat, would represent a
    larger proportion  of the heating  value of
    the  coal  and would  permit a  second
    combustion step in which the gas would
    be  heated to a much higher temperature
    at the inlet of the  gas turbine.
The drive for higher performance in  aircraft
engines  will  continue,  and experience from
such engines can be expected  to maintain the
historic  upward trend of temperature  of gases
at the inlet of industrial gas turbines. As this
temperature  rises, the  second advantage of
partial combustion may appear progressively
more important.6

  Partial  combustion  has  difficulties,  how-
ever:
  1. Sufficient carbon  must be present in the
    bed  to   support  the   CC>2-carbon  and
    H2O-carbon reactions. Carryover  from
    the  bed will contain carbon which must
    probably be burned  elsewhere,  perhaps
     best in an  auxiliary  bed for complete
     combustion.
   2. The temperature must be significantly
     higher, probably at least 900°C, to cater
     to the kinetics of the CC>2-carbon reac-
     tion.  The higher temperature  increases
     the danger  that  alkali  in  the gas will
     injure  the  gas turbine, and  therefore
     increases the likelihood that gas will have
     to be cooled and scrubbed.

   If gas must be cooled and  scrubbed to deal
with alkali from partial combustion, but not
from a total combustion at 800°C as Hoy and
Stanton  hope, then the advantages of partial
combustion might  disappear in a total eco-
nomic balance,  at  least for  gas-turbine inlet
temperatures currently available.

   When a  temperature  is  selected for the
partial  combustion  bed,  attention  must be
paid to the problem of sintering of the  dolo-
mite-derived  solid.  Preliminary work at The
City College confirms Haul's  finding5  that
some  sintering  of  [CaO+MgO]  occurs at
temperatures in  the vicinity of 925-975°C.
There are indications, however, that the solid
stabilizes at  a surface area on the  order of
about 5 m2 /gram. It must be emphasized that
this is  a preliminary conclusion, but  it would
appear reasonable to expect  that a  tempera-
ture below about 975°C should be compatible
with a  relatively long life for the dolomite-
derived solid in a partial combustion bed.
   If the gases from such a bed must be cooled
and scrubbed, there may  be an advantage of
an  ash-agglomerating fluidized bed  operating
at a higher temperature and not incorporating
an agent to capture sulfur. Gases from the bed
at higher temperature  will be lower in  H2O
and CC>2 content,  and hence will be admir-
ably suited for  desulfurization  by  reaction
with half-calcined dolomite, as in reaction (2).
The gases should be cooled  several  hundred
degrees before  the  desulfurization  step. We
have been surprised to discover that reaction
(2) is considerably faster, at a given tempera-
ture and H2S partial pressure, than reaction
(3). Accordingly, a  panel bed filter7 is well

                                    IV-1-3

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suited for carrying out reaction (2). Perhaps
such a filter, with the help of a filter aid, can
remove alkali dust from the gas to meet the
cleanliness required for the gas-turbine inlet.
It will be  important to learn whether or not
this is so.  The filter could operate at a tem-
perature below 800°C,  and the life  of the
filter solid should be excellent.

Carbonization Approach

   One of the great inventions is the proposal
by Gorin, Curran, and Batchelor8  to desul-
furize  coal  char  by cooperative  action  of
hydrogen and an  "acceptor" for sulfur, such
as calcined dolomite. Under the proposal, the
acceptor  and coal  char would differ suffi-
ciently in particle size and density so as to be
readily separable.

   This proposal's originality  is appreciated if
it is recalled that the scientific basis for the
idea was available as early as 1923  from the
work of Powell9  of the U.S. Bureau of Mines.

   The inventors' apparent major interest was
to provide a low-sulfur char for use in making
"formcoke" for metallurgy.

   Theodore1 ° elaborated the idea to produce
a low-sulfur coke for power generation while
converting  volatile  matter in  the coal  to
pipeline gas.

   I have  examined  the idea's potential for a
system intended simply to provide low-sulfur
fuels  for  power.11'12-13  Figure  1 depicts
broadly a  scheme we are studying. The coal
would be carbonized at elevated pressure and
the resulting coke desulfurized by cooperative
action of hydrogen and the acceptor. Heat for
the distillation of volatile matter  from coal
would be supplied by partial combustion of
the volatile  matter, typically consuming  11
percent of the stoichiometric air for complete
combustion  of the coal. Hence,  equipment
volumes would he even less than in the partial
combustion  approach. The quantity of lean
fuel gas is relatively small, and equipment for
its cooling and scrubbing, if necessary, would
be correspondingly small. The  lean fuel gas
would drive a  gas  turbine,  and a satellite
steam power station would receive waste heat
from the gas turbine's exhaust as well as from
the coal desulfurization  process. An installa-
tion  would operate at a  steady coal feed and
would provide baseload power corresponding
to about a third of the  heating value of the
coal. The installation would furnish low-sulfur
coke for production of variable  power, either
in nearby equipment or at a distance.
POWER  PRODUCTION  FROM  LOW-
SULFUR COKE

  I will  defer a discussion of the scheme it-
self, and first  take up alternatives  for  the
production  of power  from  the low-sulfur
coke.
Alternative 1

  The  coke might be  burned  in existing
power stations. Grinding the coke for pulver-
ized boilers might  be costly, but using  the
coke  in  stations having  cyclone combustors
might  be relatively easy, if the coke were
made from  coal having ash suitable for such
combustors.  Some gaseous  or  volatile fuel
might  be needed to  support combustion of
the  coke,  but  many  utilities  have now
switched  to   low-sulfur  oil,  and  would
welcome  the opportunity to replace even a
part of this  expensive oil  with a cheaper fuel.
A modification  of  Figure 1  might provide a
suitable gaseous or  liquid  fuel produced from
the coal itself.
Alternative 2

   Elliott14  has pointed out the advantages of
the fluidized-bed  boiler  for  stations whose
purpose  is  to  meet  peakload requirements.
The coke would be an excellent fuel for such
stations,  relieving  their owners of the neces-
IV-1-4

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sity  of providing equipment to  cope with
sulfur. It  would appear that the cost  advan-
tage might belong to a design at high pressure,
such as that proposed by Hoy and Stanton.
Alternative 3

   The  efficiency of  Hoy  and  Stanton's
scheme would be improved if gas from the
fluidized-bed  boiler were heated to a higher
temperature  ahead of the gas turbine by a
combustion  of  fuel derived  from  volatile
matter. This idea is depicted broadly in Figure
2. The figure assumes, perhaps too pessimis-
tically, that  the gas  from  the fluidized-bed
boiler must be cooled and scrubbed to remove
dust. If this pessimistic assumption is correct,
the availability of a gaseous fuel to reheat the
boiler offgas would be welcome.
Alternative 4

  The coke might  be gasified in a partial
combustion  process, providing fuel gas to
drive a gas turbine. A wider range of alter-
native gasification  processes  might be con-
sidered for  the low-sulfur coke than for raw
coal.  For example,  a  process using an  ash-
agglomerating fluidized bed  might  seem more
attractive for the low-sulfur coke than for
coal.  If  the  coke  has a  sufficiently  large
particle size,  it would be  a good fuel for
Secord's  slagging,  gravitating-bed  gasi-
fier.15-16
  This  alternative  becomes attractive if a
major part of the coke is to be burned to pro-
vide baseload power, and if advantage is to be
taken of the higher gas-turbine inlet tempera-
tures expected to become available.
Alternative 5

  The coke might be burned in a fluidized-
bed boiler using total combustion and heating
steam to temperatures well beyond those now
in  use  in conventional  power  equipment.
Elliott17  has called attention to this oppor-
tunity. In light of present knowledge, devel-
opment of a device for partial combustion of
coal or coke must be viewed as more uncertain
than development of a fluidized bed for total
combustion. It must also be admitted that the
expectation that "ultra-high" gas-turbine inlet
temperatures  (i.e., beyond 1100°C) will
become available is probably not quite yet a
certainty.  Under  these  circumstances,  it
would be a mistake to overlook the oppor-
tunity which the fluidized-bed boiler offers to
improve steam cycle efficiency by raising the
steam temperature.

   If the historical line of development of the
steam cycle is  followed,  even  a modest in-
crease in  steam temperature will be  accom-
panied by a sharp increase in steam pressure. I
have called attention to a steam cycle employ-
ing an unusually large amount  of super heat
and affording an  efficiency comparable  to
that of the conventional steam  cycle for  a
given maximum steam temperature,  but not
requiring  such  high pressures.11'18'19 De-
tailed cost examination is still needed to judge
the worth of the new cycle, which should not
be  altogether  overlooked when  considering
higher steam temperatures.
ACCEPTOR TECHNIQUE FOR DESULFUR-
IZING COAL

  Powell's historic  work9, confirmed  and
extended by Batchelor, Gorin, and Zielke,2 °
teaches that the acceptor technique for desul-
furizing coal char will work best if the char is
carbonized quickly and moved promptly from
the carbonization step  to the desulfurizing
environment. This fact is easily understood
when it is remembered that cokes sinter exo-
thermically at  temperatures  above  about
600°C. If the  coal char is allowed to sinter,
sulfur is locked into the coke structure, the
sulfur is inaccessible to reaction with hydro-
gen, and "deep desulfurization" of the coke is
                                    IV-1-5

-------
 impossible. In thermodynamic terms, sulfur in
 a  coke  soaked for a  long time  at  a high
 temperature is characterized by a free energy
 such that the pseudo-equilibrium between the
 sulfur and hydrogen is highly unfavorable for
 production of H2S.

   The  agglomerating  tendencies of bitumi-
 nous coal  have posed problems for designers
 of processes  to carbonize such coals in fluid-
 ized beds. Many American  coals, particularly
 in  the  East,  are  so  strongly  caking that
 attempts to  carbonize  the  coals in a simple
 one-step fluidized-bed carbonization process
 have often ended  in failure caused by the
 formation of massive agglomerates. Processes
 have been devised to avoid  massive agglomer-
 ates by providing several fluidized carboniza-
 tion stages at progressively higher tempera-
 tures.  Other  processes would inject raw coal,
 at a relatively low rate, either into a dilute
 stream  of  hot recycled  char  or into  a
 mechanically mixed mass of such char. Either
 dodge is poor: the first because of the length
 of time during which  the  char  is soaked at
 high temperature; the second because of the
 danger of back reactions returning sulfur to
 the recycled char.

   In  an  acceptor  desulfurization process,
 there  will be an  advantage if the coal is
 ground quite fine, so that  desulfurization is
 not  hindered  by diffusion of  the gaseous
 reactants, hydrogen and hydrogen sulfide.

   Figure 3 is a fuel desulfurization scheme we
are  studying.  Coal  is  supplied  in a finely
divided  condition,  but coke particles  are
allowed  to grow large in a coke-self-agglom-
erating fluidized bed. Because the agglom-
erated coke is almost completely nonporous,
each fresh patch  of  coke  (produced  by
capture of a tiny sticky particle of coal) must
be desulfurized promptly after it forms upon
a coke  pellet. To accomplish this, there is a
rapid exchange of coke  between the agglom-
erating zone  and a desulfurizing zone super-
posed  upon   the agglomerating  zone.  The
desulfurizing  zone is a fluidized bed of finely
divided calcined dolomite.  The fluidization
IV-1-6
velocity in the agglomerating zone is too high
to permit the dolomite  acceptor particles to
sink deep into this zone, and the fluidization
velocity in the desulfurizing zone is too low
for coke  pellets  ejected into this  zone  to
remain there permanently.

   Because of the small  size of the acceptor
particles,  they  may  advantageously  be
calcined  in  the "recirculating fluid bed" de-
scribed  recently  by  Reh.21 Such a  bed is
indicated  schematically  for the calcination
zone of Figure 3.

   Partial  combustion  of volatile   matter
supplies heat  to  calcine the spent acceptor.
Partial combustion not  only provides a lean
fuel gas  to drive a gas turbine but  is also
advantageous  because the temperature in the
acceptor calcination zone may be lower, for a
given total  operating  pressure, than  the
temperature which must be specified if com-
plete  combustion of a fuel is used to  provide
heat to this zone. The maximum temperature
necessarily reached by the solid in the calcina-
tion zone of  Figure 3,  for operation at  21
atmospheres,  would  be  in the vicinity  of
950°C. Acceptor reactivity and life should be
good.

   Back  reactions  can  restore sulfur  only
superficially to the desulfurized coke product,
for coke  pellets  produced in a  coke-self-
agglomerating bed will be almost completely
nonporous and hence  only superficially  in
contact with gases present  in the carboniza-
tion step.

   Only  a pilot   plant  can  establish  with
certainty  the  operability  of  the coke-self-
agglomerating  zone  of Figure  3.  We  are
conducting fluidization  tests at atmospheric
conditions which  are beginning to elucidate
the principles governing behavior of  a fluid-
ized bed of large particles coated with a sticky
substance. The work indicates that such a bed
can carry  an astonishingly large  burden  of
sticky matter. This result, viewed  together
with a number of successes elsewhere in the
operation of  agglomerating beds, makes  us

-------
confident of the operability of the coke-self-
agglomerating zone.
CONCLUSION

  The reader may now object that this paper
has presented' too many proposals, and he
may  wonder  what  I  consider the  best to
develop.  I would reply that paper research is
cheap and the importance of the subject justi-
fies detailed engineering study and cost evalu-
ation  of  almost  any  plausible  idea which
offers itself. The needs of the power industry
are so varied that a number  of systems may
prove commercially viable.

  Without such engineering studies, a choice
among the systems must be largely subjective,
and I can make it only by asking myself, on
which systems would  I be happy  to devote
substantial amounts of time and  energy?  I
would, today,  pick three systems.

   1. The acceptor desulfurization scheme of
    Figure  3; it  offers  the overwhelming
    advantage of providing low-sulfur fuel to
    existing plant in  a  form which can be
    stored and shipped.

   2. A bed  for complete combustion, either
    of coal in presence of half-calcined dolo-
    mite or  of low-sulfur  coke  from the
    acceptor desulfurization process.

   3. An ash-agglomerating partial combustion
    bed  followed by  a panel  bed  filter at
    lower temperatures, with the  bed  con-
    suming either coal or low-sulfur coke; if
    coal is  used, the panel would be charged
    with half-calcined  dolomite.

ACKNOWLEDGEMENTS

  Work at The City College on reactions (2)
and (3) is supported by  Research Grant No.
AP-00945 from  the National Air Pollution
Control Administration, Consumer Protection
and Environmental Health  Service. Work on
the panel bed filter is supported by Research
Grant AP-00692 from the same sponsor.

  Contributions  toward  matter reported
herein  have been  made by Professors Robert
A.  Graff  and Robert Pfeffer  and  Messrs.
Melvyn  Pell, Lawrence  A.  Ruth,  Leon
Paretsky,  S.  Narayanan,  Ralph Levy,  F.
Farouk, Basil Lewris, Barry Hertz,  Steven
Weinstein, and   Gary Weil. Messrs.  John
Bodnaruk  and George Dilorio helped greatly
with experimental arrangements.
BIBLIOGRAPHY

 1. Hoy, H. R. and J. E. Stanton. A.C.S. Div.
    Fuel Chemistry Preprints 14 (2): 59, May
    1970.

 2. Coke, J. R.  Ph.D. thesis (mentor: M. W.
    Thring), University of Sheffield, England,
    Dept. of Fuel Tech., May 1960.

 3. Bertrand,  R. R., A. C. Frost, and A.
    Skopp. Fluid Bed Studies of the Lime-
    stone Based Flue Gas Desulfurization
    Process, report from Esso  Research and
    Engineering  Co.  to  NAPCA, October
    1968.

 4. Squires, A. M. and R. A. Graff. Panel Bed
    Filters for Simultaneous Removal of Fly
    Ash and Sulfur Dioxide: III. Reaction of
    Sulfur Dioxide with Half-Calcined Dolo-
    mite. Paper  presented  at meeting of Air
    Pollution Control Association, St. Louis,
    Missouri, June  1970.

 5. Haul, R. A. W. Z. Anorg. Allgem. Chem.
    281: 199,1955.

 6. Robeson,  F. L.  and  A.   J. Giramonti.
    A.C.S. Div. Fuel Chemistry Preprints 14
    (2): 79, May 1970.

 7. Squires, A. M.  and R. Pfeffer. J. Air Poll.
    Control Assoc. 20 (8): in  press, August
    1970.
                                   IV-1-7

-------
 8. Gorin, E., E. P. Curran, and J. D. Batch-
    elor. U.S. Patent 2,824,047, February 18,
    1958.

 9. Powell, A. R. J. Am. Chem. Soc. 45: 1,
    1923.

 10. Theodore, F. W. Low Sulfur Boiler Fuel
    Using the  Consol CC>2 Acceptor Process:
    A  Feasibility Study. Report from  Con-
    solidation Coal Company to Office  of
    Coal Research, November 1967.

 11. Squires, A. M. A Role for Fluidized Com-
    bustion in Clean Power Systems. Paper
    presented  at First  International Confer-
    ence  on Fluidized-Bed Combustion,
    sponsored by NAPCA,  Hueston Woods,
    Ohio, November 18-22, 1968.

 12. Squires, A. M.,  R. A. Graff, and M. Pell.
    Desulfurization  of Fuels with Calcined
    Dolomite:  I.  Introduction  and   First
    Kinetic   Results.  Paper  presented  at
    Washington,  D.  C.  meeting of A.I.Ch.E.,
    November  1969,   to  be  published  in
    A.I.Ch.E. Symposium Series.

 13. Squires, A. M.  U.  S. Patent 3,481,834,
    December 2, 1969.
14.  Elliott, D. E. Can  Coal Compete?  The
    Struggle for Power.  Inaugural lecture,
    The University of Aston in Birmingham,
    England, November 20, 1969.
15. Secord, C. H. U.S. Patent 3,253,906, May
    31, 1966.


16. Hoy, H.  R., A.  G.  Roberts, and D. M.
    Wilkins. Inst. Gas Engrs. J. 5: 444, 1965.
17. Elliott, D. E. J. Inst. Fuel. 43: 258, 1970.


18. Squires,  A. M. Clean Fuel Power Cycles,
    ASME paper 67-WA/PWR-3, 1967.


19. Squires,  A.  M. U.S.  Patent 3,436,909,
    April 8,  1969.

20. Batchelor, J. D., E.  Gorin, and C. W.
    Zielke. Ind.  Eng. Chem. 52: 161, 1960.

21. Reh, L.  Highly Expanded Fluid Beds and
    Their Applications.  Paper presented  at
    Puerto Rico meeting of A.I.Ch.E.,  May
    1970, and to appear in Chem. Eng. Progr.
IV-1-8

-------
- AIR A|
W ^ COMPR
i
AIR AT
R GAS
ESSOR TURB|NE 	 »- BASELOAD
I

HIGH PRESSURE
I
COAL DIS1
PARTIAL CC
OF VOLAT
DESULFURIZA1
. 	 , " I

UUMliUoTIQN '
A STEAM POWER
LEAN STATION
ILLATION
5MBUSTION
LE MATTER
riON OF COKE
^ COAL f
*
HIGH-PRESSURE STEAM
LOW-SULFUR COKE TO POWER STATIONS AT A DISTANCE
(VARIABLE POWER)
                  Figure 1.  Scheme for providing low-sulfur fuels for power.
         COMBUSTION
         OF FUEL GAS
             COAL DESULFURIZATION
                  PROCESS
COAL
AIR
     Figure 2. Lean fuel gas from coal desulfurization used to raise the temperature of
     combustion products from a coke-fed pressurized fluid-bed boiler.
                                                                                  IV-1-9

-------
                                                                           LEAN FUEL GAS
                                                                               IN SULFUR)
 AIR
 DESULFURIZING
    ZONE
CaO
CaS
                                                                    H2S
                                                CaC03
                                                 CaS
                                                CaCOo
                         SULFUR
                        DESORPTION
                                          CLAUS
                                          SYSTEM
                                   STEAM AND CO2-
                                                                 AIR
                            COAL HYDROCARBONIZING
                                 AND COKE SELF-
                             AGGLOMERATING ZONE
                                      ELEMENTAL
                                       SULFUR   I
                                      TO MARKET W



                                 RECYCLE OF PORTION
                                 OF LEAN FUEL GAS
                               (CONTAINING HYDROGEN)
                 LOW-SULFUR
                 COKE PELLETS
             Figure 3. Acceptor process converting coal  into low-sulfur fuels.
IV-1-10

-------
                            2.   FLUIDISED-BED  COMBUSTION
                                                                      AND THE
                                                 DESIGN  OF  BOILERS

                                  S. J. WRIGHT
                        National Coal Board, England
INTRODUCTION

  Conventional electricity-generation boilers,
whether  fired by pulverised coal or oil, are
characterised by very large combustion cham-
bers. The size is dictated: first, by the need to
allow sufficient time for combustion despite
temperatures in the range  1200-1600°C; and
secondly, by the need to allow  molten and
potentially  corrosive ash  constitutents  to
resolidify before coming into contact with the
banks of convective superheater tubing. The
result is a supporting structure more than 200
ft high, for a 660-MW boiler, costing some 14
percent of the total boiler  cost with, in addi-
tion, extensive  civil  works to  accommodate
the substantial point loads.

  Fluidised  beds have two properties  which
allow a  radical rethinking of boiler design
concepts.

  1.   The   rapid  movement  of  particles
within  the fluidised bed gives rise to relatively
high rates of heat and mass transfer  between
the solids and the fluidising gas. The result is
that coal can be burned in a fluidised bed of
some  2-3 ft deep at volumetric heat release
rates in  excess  of anything  achievable by
conventional methods  and at  mean bed
temperatures between 800 and  850°C. At
these relatively low temperatures: ash consti-
tuents  do not  melt; almost all  the  alkali
metals,  known  corrosion promoters,  are
retained in the bed; and, by the  addition of
limestone to the bed, a substantial proportion
of the sulphur content  of the fuel can be
                                        FV-2-1
retained as CaSO4- Thus the need for a large
combustion volume is eliminated.

   2.  The rapid  movement of the particles
within the  fluidised bed gives rise to relatively
high  heat-transfer coefficients between  the
bed and surfaces immersed in it, up to 5 times
greater than those achievable in conventional
gas-tube heat exchangers. The environment of
a  fluidised-combustion  bed  allows  steam-
raising tubes  to be placed in the bed. Thus,
substantial  savings can be made in the length
of tubing required to extract a given amount
of  heat,  and  possibly, the less corrosive
environment  will  allow less expensive alloys
to be used in the superheater and reheater
sections.
DESIGN OF FLUIDISED-BED BOILERS

  The  most important decision in the design
of a fluidised-bed boiler operating at  atmos-
pheric  pressure is the choice of fluidising
velocity.  The  fluidising velocity  fixes the
oxygen  availability, and  hence heat-release
rate, per  unit  cross-sectional  area of bed
which,  in turn,  fixes the total cross-sectional
area required for a given steam duty.  There-
fore, considered in isolation, the capital cost
of the  boiler falls with increasing fluidising
velocity because of savings  in containment
costs.

  However, two factors oppose the incentive
to increase fluidising velocity.

-------
   1.  As the fluidising velocity is increased, a
 progressively  coarser size  spectrum  of bed
 particles must be used in order to limit elutri-
 ation  losses.  As the mean particle size  of a
 fluidised bed   increases,  the  heat-transfer
 coefficient  within it decreases:  e.g., as the
 mean  particle size increases  from 0.4  mm
 (0.016 in.) to 1.0 mm (0.039  in.), the coeffi-
 cient  falls about 35  percent.  The result is a
 tendency for the bed depth, and hence pres-
 sure  drop and  running costs,  to increase  in
 order to  accommodate the increased  area  of
 tubing required in the bed; this is in addition
 to the increased capital cost of tubing.

   2.  As the  fluidising  velocity  increases,
 elutriation rates, and hence carbon loss from
 the  bed,  tend to increase due to the increased
 agitation and degradation within the bed, and
 residence time in the combustion zone tends
 to decrease. Therefore, larger and more elabo-
 rate  recycle  systems are required to  achieve
 any particular combustion efficiency with a
 consequent   increase in  both capital  and
 running costs.

   Thus  the  choice  of  fluidising  velocity
 becomes an  optimisation of conflicting eco-
 nomic incentives.

   If,  on  the other hand, the boiler is pressur-
 ised, an additional degree of freedom is intro-
 duced in that, for a given velocity and a given
 cross-sectional area,  the heat  release may  be
 increased by increasing  the  pressure,  and
 hence oxygen  availability.  Therefore, for a
 given steam output the containment cost can
 be reduced by  increasing the  pressure. There
 is no heat transfer penalty  because velocity,
 and hence  bed  size  spectrum,  remains  con-
 stant; however,  bed height, and hence  pres-
 sure drop, increases  because more heat must
 be  extracted from  a bed  of fixed cross-
 sectional area,  and  there is  a limit to the
 density of tube packing. The apparent insen-
 sitivity towards operating under pressure is,
 however, offset by the increased plant  com-
 plexity imposed by  the pressure and also  by
 the  availability  of  industrial  gas turbines
 designed  for a  suitable pressure ratio and
 capable of being modified  without incurring
 substantial  development charges. Although
 the remainder of this discussion will  consider
 only atmospheric pressure boilers, much of it
 will be relevant to the case of operation under
 pressure.

   Having  chosen  a fluidising velocity, the
 cross-sectional area of the unit  is  fixed for
 a given  steam  output.  Because  the  bed
 temperature  will   be  limited  to   between
 800 and  900°C  by ignition  and ash fusion
 criteria,  respectively  (at least for English
 coals),  the  required  distribution  of  heat-
 transfer surface between the fluidised bed and
 conventional surfaces above the bed may be
 calculated. Typically some  60 percent of the
 heat release must be extracted directly  from
 the bed if a temperature of about 850°C is  to
 be maintained.

  There is a cost incentive to  place  those
 parts of  the  steam/water circuit where the
 more exotic materials have to be used in the
 fluidised bed. If carbon steel has a cost index
 of 1.0 (under English conditions), 2 percent
 Cr  has  2.4,  9  percent  Cr  has 5.4,  and
 Austenitic  316  has  13.4.  Therefore, the
 incentive  is to put the superheater and  re-
 heater in  the  fluidised bed and make up any
 deficiency  with  evaporator  surface leaving
 the water walls and any  gas-space surface  to
 the remaining vaporisation  and  the econo-
 miser.  However,  under part-load conditions,
 the downsteam ends of in-bed reheaters and
 superheaters can  get relatively hot, requiring
 the use of austenitic steels where ferritic steels
 are adequate at the continuous-rated output.
 Thus, particularly at relatively  low fluidising
 velocities  with consequently  high heat-trans-
 fer coefficients,  there may be an economic
 incentive to place the downstream sections  of
 the reheater and superheater  in the gas-space
 above the bed and  avoid the use of expensive
 steels.

  Finally, there is an incentive (at least under
English conditions)  to maximise  factory
IV-2-2

-------
construction, thereby  minimising  site work
and  the  consequent scope  for construction
errors—all of which reduce the time of con-
struction  and  hence  the  interest  charges
incurred. Interest during construction can be
as much as 26 percent of the capital cost of a
2000-MW  power station. Fluidised-bed  com-
bustion presents, for the first time, the possi-
bility of designing a large boiler as a number
of  semi-independent  units  which  can  be
factory-assembled, leaving the minimum of
connecting-up on site. For English road trans-
port conditions, the maximum transportable
size  is about 14 ft wide x 40 ft long x  16 ft
high; however, for a coastal  power station,
sizes  as large as 25 x  44 x  22.5 ft can be
brought in by sea. As an example, using road
transport  criteria, a 120-MW boiler operating
at a  fluidising  velocity of 8 ft/s can be de-
signed as five  factory-assembled  beds  (two
evaporator, two reheat, and  a superheat) with
a secondary superheater and an ecomoniser in
the gas-space above the bed. This arrangement
is estimated to save 9 months in construction
time. With high interest charges, the incentive
to  utilise  factory  assembly  can  be strong
enough to decide a fluidising velocity at the
expense of say a slightly lower combustion
efficiency.
CONCLUSION

  The foregoing  shows that many  of the
potential advantages of fluidised-bed combus-
tion,  for reducing the capital cost of boiler
plants, are  self defeating unless the design is
optimised in respect to each variable.
  The design should:

  1. Obtain  maximum  heat-transfer  coeffi-
    cients in  the  bed and  the  bulk  of the
    boiler  surface  operating  at  metal
    temperatures  in  excess  of  400°C
    immersed in the bed.

  2. Obtain  the  maximum combustion  effi-
    ciency by maximising the residence  time
    of gas and particles in  the bed and the
    combustion space above it.

  3. Obtain  maximum  factory  assembly  in
    packaged units.

  4. Reduce structural  and  civil  work  to a
    minimum.

  5. Reduce to  a minimum the lengths  of
    high-pressure pipework connections.

  Design aims 1  and 2 argue  for low fluidising
velocities; aims  3, 4, and 5 argue for  high
fluidising  velocities. Thus, if full advantage is
to be taken of the potential of fluidised-bed
combustion  for  reducing  the cost of power
station plant, an economic optimum must  be
sought whereby the fullest advantage is taken
of all possible savings.


ACKNOWLEDGMENTS

  The  work described  in  this  paper  was
carried out as part of the research programme
of the Research and Development Depart-
ment  of  the  National  Coal Board.  Views
expressed  are those of the author, not neces-
sarily of the board.
                                                                                  IV-2-3

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                                            3.  IGNIFLUID BOILERS
                                                                             FOR
                                             AN ELECTRIC UTILITY
                                  R. H. DEMMY

                                UGI Corporation
INTRODUCTION*

  The UGI  Corporation operates an electric
utility (Hunlock)  in  Northeastern Pennsyl-
vania, serving an area of 440 square miles near
Wilkes-Barre, along the Susquehanna River.
UGI's Hunlock Generating Plant (88,700-kw
net) burns anthracite from the Northern and
Middle  Pennsylvania  anthracite  fields  on
Barley stokers and in pulverized-fired boilers.
This  plant's latest addition,  completed  in
1959, is  a  Foster Wheeler  400,000 Ib/hr
pulverized  boiler  utilizing anthracite bank
material  with  a  quality  of approximately
9000 Btu/lb. Stack emission is controlled by
electrostatic and  mechanical  precipitators
installed by Research-Cottrell.

  UGI's Planning Department  indicated that
new  generation was  required in the early
1970's, both to meet new sales and to replace
old facilities. To  satisfy this projected need
for new generation, we investigated the possi-
bility of burning fuels adjacent to our service
territory. We found that Babcock-Atlantique
of France had a boiler utilizing the Ignifluid
combustion  process which would be capable
of burning the  anthracite culm banks of our
area (there  are 100,000,000 tons within 40
miles of our  projected plant site).

  Small samples of the culm bank material
were sent to France: first, four 5-lb bags; later
40  tons  in  55-lb  bags.  When preliminary

The technical assistance of P.A.Mulcey, fuel consultant of
Dallas, Pennsylvania, is acknowledged in the development of
this paper.
laboratory  tests indicated  that  a  full-scale
combustion  test  should be  attempted,
approximately 800 tons  of 40-percent ash,
7500   8000 Btu/lb culm bank  fuel was
shipped to an area in the French Alps near
Grenoble,  where  the Cartonneries  de  la
Rochette operates a cardboard factory.
OBSERVATIONS OF PRODUCTION TESTS

  Today, 1 would like to discuss with you our
observations in both Casablanca,  Morocco,
where the largest Ignifluid boiler is located,
and in La Rochette, France, where the com-
bustion tests were performed.
In Casablanca

  Arriving in Casablanca late  in the after-
noon, we were concerned that  we would be
unable to visit the plant that  day. We pro-
ceeded immediately  to the  power  plant,
approximately 1 mile from the center of Casa-
blanca.

  As we approached, we were shocked to see
that the stack indicated no power generation.
Because we  were  approaching  an anthracite
plant we had anticipated what we have experi-
enced in the United States in the way of
anthracite stacks; seeing the  stack completely
without any visible emission, we were quite
upset that we had travelled such a distance to
find the plant not operating.
                                        IV-3-1

-------
   When we  arrived in  the control room, the
 Moroccan operators told us that the plant was
 at full load, a fact which was difficult for us
 to accept. But the recording charts indicated
 that the plant, with a  maximum capacity of
 60 mw, was  operating at its normal capacity,
 approximately 55 mw, and that the boiler had
 been operating at  approximately the same
 output most of the day.
   The stack emission control for this boiler is
a Lurgi electrostatic precipitator preceded by
a  mechanical precipitator. The precipitator
was to have been  a three-cell  unit  but the
Moroccans, without sufficient money to buy
all three cells, purchased two internals and the
shell for the third.  Results were so successful
that they  have  not added the  third electro-
static section.
   The dust removal from these units is similar
to what we are used to in the States; a hydro-
vac or similar system is used. The one differ-
ence is that all of the dust which is collected
by the precipitators is reinjected  into the
furnace. There is no fly ash  pond. All ash
from  the Ignifluid  boiler  is  removed  as
bottom ash, and sluiced  into the Atlantic
Ocean which, at the site of the power plant, is
400-ft deep.
  The burnout of the carbon is quite com-
plete: 4 to 6 percent carbon is left in the ash.
  Electrostatic  precipitation is  successful
because the  fly ash of an Ignifluid boiler is
high in carbon and therefore low in resistivity
which is the opposite of the pulverized boilers
operated  in  the States.  Therefore,  a clean
stack  is  not a  desire but a fact.  For an
operator  of an  anthracite-fired plant or  a
bituminous coal low-sulfur plant which would
again have high resistivity, it was indeed hard
for us to  believe our own eyes when we saw a
boiler at  full  load  with  an optically clear
stack.
In La Rochette

  Visiting the site of our production tests in
the French Alps at La Rochette, we found
that the Ignifluid boiler there has no electro-
static  preci- pitators,  but two  mechanical
precipitators. The stack emission  at that site
was similar to  the  best emission  we have at
Hunlock Plant, utilizing electrostatic precipi-
tators installed last year.

  La Rochette's coal is transported by rail
from Antwerp, Belgium, and dumped directly
into the silos at La Rochette.

  The  developer of  the  Ignifluid  process
(described more fully at the end of this paper)
is  Ms. Albert  Godel, of  Socie'te' Anonyme
Activit, who spoke at the First International
Conference in  1968, and is still  very active
and interested in  this process. Babcock-
Atlantique has purchased  the patents from
him and  are directing the  development pro-
gram.

  A steam flow of 100,000 Ib/hr from a grate
area approximately 3  ft wide  by 25 ft long
(compared to 100,000 Ib/hr from a standard
anthracite spreader stoker 20 ft wide by 28 ft
long) indicates the improvement in the state-
of-the-art of burning anthracite.

  The tests  were  under the  supervision of
CERCHAR, the French equivalent of the U.S.
Bureau of Mines and Bureau  of Standards,
combined.

  The continuous  removal of the ash from
the Ignifluid boiler is accomplished  by an
inclined  grate  rising at  an angle, of 10-12
degrees. The level is maintained in Ifte fire bed
by having variable air pressure supplied to the
various zones of the fire by  separate air ducts.

  In  this boiler  there are six  zones;  the
highest pressure  is where the  raw fuel is
supplied to the boiler by a Redler feeder. The
pressure  decreases  as  the fuel   progresses
toward the front of  the  furnace; the  area
IV-3-2

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where the  ash rises out  of the fluidized bed
requires additional air for final combustion of
the carbon in the ash.
  During  our  test,  the  combustion  was
stopped due  to  a plugging of the coal feeder
and  we were able to witness a reignition of
boiler. The coal  is supplied to the small grate
in a manner similar  to  the way a spreader
stoker would receive its coal; the  fire  is
started in the same manner,  with  excess air
supplied  until  the  fire  starts effervescing.
When this effervescent condition is more or
less  uniform  over the entire grate, which is
not  moving  at  the  time,  the bed  is then
fluidized by the introduction of sufficient air
to cause the  particles to rise from the grate
and  be balanced by the upward flow of the
primary air and the action  of gravity on the
particles.


  Air distribution in the boiler does not seem
to be critical.  In this boiler, ambient  air
(approximately  50 percent primary  air and 50
percent  secondary  air)  is  introduced  just
above the fluidized bed in the same zone as is
the reinjected fly ash. When the bed is fluid-
ized, the furnace looks like a gas furnace with
an incandescent  white heat.
  An inspection of the internals of this boiler
indicated  the tubes to have no slag accumu-
lation. The 2-year-old grate  looked brand
new,  undoubtedly because of the introduc-
tion of primary air to keep  the  grate from
reaching high  temperatures.  The  ash carbon
burnout  for our tests, as was predicted  in
model tests in the laboratory, was not as high
as those  in Casablanca (4-6 percent) and La
Rochette  (6-8  percent): the carbon in the ash
of the Pennsylvania culm was 17-20 percent.
In  the final design of the proposed boiler,
burnout to a level of less than 10 percent will
be  accomplished  by a post-combustion fur-
nace similar to the type utilized for expanded
ash for lightweight aggregate purposes.
PROPOSED NANTICOKE STATION

   After  evaluating the combustion tests in
France and completing the overall engineering
evaluation,  it  was  determined  that  a
300,000-kw station  could  be built utilizing
two  Ignifluid  boilers and  one  turbine-
generator. A  unit size of 300,000 kw  was
selected due to the availability of units of this
size already operating. To further enhance the
availability, two boilers are being used. Since
the major cause of loss of generation is boiler
failure,  the use  of two  boilers  will allow
operation at  half load while  repairing the
failed boiler. Steam conditions at the throttle
of  the  turbine   will  be  1800  psig
1000°F/1000°F.  The  turbine  will also  be
supplied  with the  option  for  extraction:
adjacent  to  the  power  plant  site,  the
Commonwealth of Pennsylvania is planning to
install a  pilot  plant  for acid  mine-water
demineralization which will require approxi-
mately 300,000 Ib/hr  of low-pressure steam.

   The fuel for this plant is above ground and
available. No deep mining of coal will be re-
quired for the proposed 30-year life of the
plant.  The  fuel will  come  from existing
anthracite culm  banks (refuse  from  the
preparation of anthracite),  100,000,000 tons
of which are located within 40 miles of the
plant site. This  fuel  will be upgraded  to
approximately  8000 Btu/lb,  a  heating value
which the boiler has proven itself capable of
handling.  Material not taken to the power
plant will  be  left  at  the  bank sites and
regraded  into  usable contours so that all of
the present areas which are not usable as real
estate today will be available for real estate
development.   The  high  ash  content  (40
percent)  of the fuel  would normally cause a
considerable problem  in handling; however,
tliis is not anticipated  as the ash has proven
itself to be excellent for utilization in cinder
blocks and has potential use  as a lightweight
building aggregate.

   We believe  that this plant will  meet the
ecological requirements of today's society.

                                    IV-3-3

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   1. We do  not anticipate participate emis-
     sion problems.
   2. Since the fuel has less than 0.5% sulfur,
     there will be no SCH emission.
   3. NOX  emission will be reduced. Tests in
     France showed that the NOX emission of
     an  Ignifluid  boiler was  approximately
     half that of a pulverized boiler firing the
     same  coal.
   4. Thermal discharge  to the Susquehanna
     River will  be controlled  by the utiliza-
     tion of cooling towers.
THE IGNIFLUID PROCESS APPLIED TO
UTILITY BOILERS*
the stoker,  as  is the case with conventional
stokers or spreaders.

  An Ignifluid-fired boiler burns any type of
small fuel: sized  from 1/8 in.-O to 3/4 in.-O,
from anthracite to bituminous coal, coking or
not. and  regardless of the ash content. The
efficiency is as high as with pulverized firing.
  Mechanically, the Ignifluid unit is similar to
the chain or traveling-grate stoker, including
as it does the same drive and return sprockets,
traveling   grate  or  chain,  and  wind-box
arrangement  consisting  of  six  or  more
compartments.
   The  Ignifluid process  is  a fluidized-bed
method of combustion of small coal which is
maintained  in stable aerodynamic suspension
under the effect of an ascending air current
which gives the layer  the appearance of an
incadescent  boiling liquid, hence  the  trade-
mark Ignifluid.

   The process  utilizes a  standardized fluid-
ized combustion burner  at the base of the
boiler. Consisting of a very narrow inclined
stoker of less  than one-tenth of  the  total
section of the boiler, it supports the fuel bed
which is fluidized by a controlled primary air
current; it also  extracts the clinker from the
bed in its ascending rearward motion.

   Secondary air, blown over the top of the
fluidized  bed, completes combustion. All the
grit carried  over is cycled back on the surface
of the fluidized bed, with the result that all
the ash from the coal  is extracted as clinker
by the chain-grate.

   A high rate of combustion is obtained (over
300 Ib of coal per sq ft) because combustion
takes place  in the whole 2-ft-deep mass of the
turbulent bed and not  only on the surface of
*The assistance of F.W. Kuehn, engineering consultant of
 AUentown, Pennsylvania, is acknowledged for this portion
 of the paper.
Essential Differences

  The essential differences are:
  1. The grate width is only about 10 percent
     that of the traditional anthracite stoker.
  2. The grate is tilted 8-12 degrees from the
     horizontal, in  the direction of travel of
     the grate,  as  contrasted  to  the  level
     stoker.
  3. There is  a different under  grate  air-
     pressure pattern  from  front  to back-
     substantially a complete reversal—with
     the maximum (main  duct)  pressures
     about 50 percent higher.
  4. The  coal  feed is different: instead  of
     being carried by  the  stoker from a feed
     hopper into the furnace under a leveling
     gate, coal is dropped by a Redler con-
     veyor into  a  steep-angled pipe which
     intersects  the  furnace front wall a few
     feet  above  the  grate.  Air  injected
     through the bottom of  the  feed  pipe
     tends to spread  the  coal  outward and
     forward as it  falls toward rJie fluidized
     bed  which, in  turn, has  an inherent
     spreading  proclivity  due to the boiling
     action.

  On a stoker,  the wind-box pressures usually
range drom 1/2-in.  water in the first compart-
ment, to a maximum of 4-5 in. in the fourth,
and  grading off to  3 in. in the sixth. In the
IV-3-4

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Ignifluid unit, pressure is highest (about 8 in.
of water) in the first compartment; it grades
down uniformly to 1  in. in the sixth, where
the coal bed is fluidized in normal operation,
as distinguished from start up (which will be
explained).

  As might be  surmised, with a chain width
only  10 percent  that  of  the  stoker,  the
burning or reaction is much more intense.
Startup

  To  understand  the  action,  let  us  run
through  a startup. A small amount of coal is
fed into the unit so it lies on the grate to a
depth  of 6-12 in. at the front. The  induced
draft fan is started. Next, oil or gas torches are
lighted and  directed down onto the surface of
the  coal until  it is uniformly ignited. The
forced draft  fan is  started and  undergrate
(called primary air) air pressure is slowly built
up  in  the compartments as indicated, with a
straight-line sloping pattern from  maximum
pressure  in  the first (or front) compartment
to a minimum in the last.

  As the flow and pressures are increased, a
point is reached where the coal lifts off the
grate and  suspension firing begins.  At this
point,  the  fluidity action is rather unstable,
and the air  flow must be increased further at
constant pressures  to somewhere under  the
point in velocity where most of the particles
would have left the fluidized bed (become gas
borne). Fortunately,  because there is a wide
band of velocities possible between the low or
unstable point  at  which the bed tends to
collapse, and  the high point where  the bed
would  be carried off with the gas, the opera-
tion is  inherently very stable.

  Under normal operation  then  (from full
load down  to  about 15 percent of rating), the
top surface  of the fluidized bed is level, look-
ing like boiling lava. Depth above grate ranges
from a maximum of about 20 in. at the front
(over the first air compartment)  to  nothing
but the clinker thickness over  the  fifth air
compartment of a six-compartment unit.

  When the coal particles are all consumed,
the remaining bits of fine ash become sticky
and agglomerate with others forming dense
clinker  which  falls  on to  the  grate  and is
carried over the rear sprocket in a ribbon 1-3
in. deep.

  Blankets of dead coal cover and protect the
sloping  refractory  walls along  the sides and
front  of  furnace,  below  the  water-cooled
walls. The dead coal  layers are maintained
because there is no air flow in these areas to
promote and support combustion.

  Vertical water jackets, extending from the
top grate  line down to the refractory at  base
of the crickets outside the wind boxes, form a
barrier between the moving hot clinker on the
grate  and  the dead coal; they also support the
latter.

  An air  seal, at  the interface  between the
bottom of structural bridge across the furnace
and the top  of the moving chain grate just
ahead of  the  first undergrate  air compart-
ment, prevents any loss of siftings in this area,
especially  during startup.

  Very little loss of fuel in the form of sift-
ings  is experienced through the grate,  it is
claimed, because of the higher undergrate air
pressures  used  (as compared to  conventional
stokers) and to the small air openings. Individ-
ual stoker chain links are  notched in  the
middle to facilitate separation of the clinker
at the rear sprockets.

  Secondary  air,  which  in larger  units is
usually preheated to higher temperatures than
the primary air, is introduced into the furnace
through side  walls  a  few   feet above   the
fluidized bed to complete combustion of the
fuel rich gases.

  An inclined air nozzle under the rear arch is
used  to blast any fine coal particles which

                                     IV-3-5

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 drop onto the surface of the clinker bed back
 into the active combustion zone.
 Control

   Control  is  essentially  conventional,  and
 typical of later  automatic stoker control in
 the States except for a unique fuel-bed thick-
 ness  probe  and  controller.  Primary  air  is
 controlled by deviation  of a master steam-
 pressure controller from a set point,  steam
 pressure being sensitive to load changes. As
 primary air flow increases or decreases under
 direction of the  master controller, it changes
 the rate of  coal  combustion and, therefore,
 the  depth  of the  fluidized bed.  A  water-
 cooled fuel-bed  thickness probe, inserted in
 the bed just above the  grate near the front of
 the  furnace,  monitors the thickness of the
 bed,  and passes  its signal continuously to a
 fuel-bed thickness controller. The controller,
 set to  hold  the  fluidized bed at a constant
 thickness,  passes  on  the  deviation-from-
 normal  signal to the  fuel-bed controller.
 Secondary air flow  is trimmed  by desired-
 percentage CC>2 in furnace exit gases. Judging
 from  typical air duct sizing, secondary  air
 appears to be  about 50 percent of the total
 air. Induced-draft fan dampers, in turn, hold a
 constant furnace  draft of minus 5 mm.
Operating Advantages

   Operating  advantages  claimed  for  the
Ignifluid unit over  the  stoker, which taken
together account for its ability to efficiently
burn low-grade coals containing  up to  50
percent ash, include:

   1. Automatic spreading of coal inherent.
   2. Clinkers  sink  through fluidized  bed to
     grate.
   3. Intimate and intense reaction of air and
     fuel.
   4. Fluidized bed completely pervious to air.
     (With a bed in a fluid  state, it is obvious
     that the cleanliness of the coal is not
     important;  in  fact, the  coal can  be
     unwashed.)
   5. Loss of sittings through air holes in grate
     chain link is not as great as in conven-
     tional stoker firing because  of  higher air
     pressures  required by the process which
     in turn means smaller air holes through
     the grate. (Claimed riddling  or  sifting
     losses are  as low as 0.5 percent.)

   Unusually high thermal  efficiencies (up-
wards of 88 percent) are claimed, even with-
out air preheating  when  using a medium
volatile (6 percent), reasonably well prepared
(16.4  percent  ash),  fairly dry  (6.6 percent
moisture) anthracite  of proper sizing (zeruto
5 mm). Boiler efficiency of 81.74 percent was
guaranteed for  the UGI units.

    As the coal becomes poorer (i.e., higher in
ash), a significant higher loss to be  reckoned
with is that in the sensible and latent heat loss
in the clinker,  part of which will be  reclaimed
with a post-combustion furnace.

   Forced outage rate for the 55-metric ton
evaporation per hour Ignifluid burning unit at
La Rochette  (given  to  Socie'te Babcock  &
Wilcox by  the owner-operator  paper
company, Cartonneries  de La Rochette) for
the year August,  1966 to August,  1967 was
2.14 percent:  0.3 percent  was due to the
stoker alone; 1.84 percent was for such things
as  repairs to  ash-reinjection system,  coal-
handling  system  breakdowns,  and replace-
ment of belts on the forced draft  fan.
CONCLUSION
                                  I
  The  ecological  demands of  the United
States today are forcing electric utilities not
to consider coal-fired plants due to the severe
problems of electrostatic precipitation of high
resistivity  ash  which  results  from  either
pulverized-anthracite or  low-sulfur-bitumi-
nous  firing.  With  insufficient  natural  gas
available, the utilities are leaning toward low
sulfur oil and nuclear material as the future
IV-3-6

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sources of energy for electrical generation. If
the  coal  industry is to participate  in  the
electric utility energy market, a modification
in the method of burning fuel is indicated. We
believe that the fluidized bed as proposed by
Babcock-Atlantique  is  the  only  method  of
fuel combustion available today in sizes which
can  be  upgraded  to the large  steam produc-
tion requirements  of  the  electric  utility
industry of the United States. Although this
method is available in commercial size today,
we fear that the delays inherent in developing
commercial pressurized  fluidized  beds will
prohibit their commercial development in the
United States.  However,  if  the  Ignifluid
process is introduced into the  United States,
we will have fuel technologists knowledgable
in fluidized combustion  capable  of guiding
the development of the next step forward in
this technology; namely, pressurized fluidized
beds.
                                                                                     IV-3-7

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                                                    4.  THE  MODULAR
                                         FLUIDIZED-BED  BOILER
                                                                   CONCEPT
                                J. W. BISHOP

                          Pope, Evans and Robbing
ABSTRACT

   The development of the Pope, Evans and
Robbins modular fluidized-bed boiler concept
is described.  The  early and unsuccessful
attempts to design a fluidized-bed boiler along
more "conventional" lines are also described
and  the  lessons drawn  from these failures
presented. A preliminary conceptual  design
and  cost estimate is presented for an electric
utility boiler having  a capacity  of 300,000
pounds of steam per hour.
INTRODUCTION

  The purpose of this paper is to describe the
experimental and design history leading to the
presently proposed modular design concept of
a fluidized-bed boiler, along with some details
of the most recent concept of a 300,000 Ib/hr
small utility unit.
EARLY HISTORY

  In 1965, as part of our program to extend
the capacity of packaged coal-fired boilers,
the late  LeRoy  F.  Deming  visited various
European combustion and boiler developers,
including  BCURA.  It  was  obvious  that
fluidized-bed firing and direct  transfer of heat
from the  bed  to the heat transfer surface
represented the most promising system for
providing a high-capacity coal-burning boiler.
Imbedded-Tube Concept

  The  first effort  was to  develop a boiler
concept as a  guide for the ensuing experi-
mental work.  Using BCURA's initial project
of 1,000,000  BTU/hr/ft2  of bed area, the
boiler concept of Figure  1 was developed.
This  concept consisted of a  continuous
imbedded-tube bank unit with a longitudinal
steam drum and three lower headers. Coal was
to have been screw-fed at two locations in the
hope that  the fluidizing air issuing from the
grid would lift and distribute the coal.

  The atmospheric  steam boiler of Figure 2
was then built, originally with a longitudinal
screw-feeding  mechanism and  a perforated
plate grid. Of course, the air did not distribute
properly and the coal merely caked above the
screw.
Water-Cooled Column

  We  then  digressed  to  a  water-cooled
column wherein the air distribution problem
was solved:  first, with a "rock sandwich" grid
(i.e., a layer of rock between two perforated
plates); later,  with the  "button grid"  cur-
rently in  use. This system directs air down
onto the  grid  plate, precludes dead spots,
prevents sifting into the plenum, and permits
operation at over 2000°F, a condition which
has not been approached with other grid plate
designs. Relying on BCURA's experience, we
                                       IV-4-1

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 screw-fed the coal into the column with air
 jets blowing the coal  off the screw, but we
 were unable to  duplicate the  BCURA opera-
 tion.  We  then  developed  a  system  for
 dropping coal into a pneumatic transport line.
 Coal-Distribution Problems

   The boiler of Figure 2 was placed in opera-
 tion,  but problems of coal  distribution were
 encountered in the presence of the imbedded
 tube  bank. Any  agglomeration  of particles
 between tubes rapidly formed clinkers which
 stuck between the tubes.  Although we could
 remove these clinkers, by lifting the boiler off
 the grid  plate  plenum  assembly,  it  was
 obvious  that this  could not be done with an
 actual boiler. Erosion was heavy on the lower
 bank of tubes. The lessons learned here were:

   1. Avoid use of imbedded tube banks if
     possible.
   2. If imbedded tube banks are required,
     consider  the  effect on fuel distribution,
     access for cleaning,  and  the  minimum
     height above the distributor to preclude
     erosion. (Design for heat removal is only
     one of a number of considerations.)

   The  imbedded tube  bank  system  was
 abandoned.

   The  design  of the  atmospheric-pressure
 boiler had been predicated on a heat transfer
 coefficient  between  the  bed  and the  im-
 bedded  tubes  of  50 Btu/hr/ft2/°F.  It  was
 found, however, that this coefficient was not
 useful in reproducing either the furnace  exit
 gas temperature or  the bed temperature. It
 was necessary to almost double the value of
 the measured  coefficient to reproduce  the
 performance with a given imbedded  surface
 area. A sizable fraction of the released energy
 was removed  from the bed by mechanisms
 other than direct contact heat transfer. It was
 necessary (as in the case of the later BCURA
 shell-and-tube boiler  experience) to cut  and
 remove tubes so as to reduce the heat removal
 rate so that, in turn, a deeper bed could be
IV-4-2
operated in which good combustion could be
achieved. The lessons learned were:

   l.A measured heat transfer coefficient is
     useful in predicting the heat flux into a
     single tube.
   2. Bed  temperature  and exit gas tempera-
     ture cannot be predicated on the basis of
     wetted surface area and a heat transfer
     coefficient  in a water-cooled enclosure.
     An   empirical  correlation  has  to  be
     derived as a prediction tool.
Open Modular Concept

  Because  so  little  "wetted"  surface  was
needed  to   control  bed  temperature  if
"viewed"  surface were  available, the design
concept of Figure 1 was seen to be unwork-
able, and the problems inherent in that design
unnecessary.  The simplest way to  avoid coal
distribution and maintenance problems was to
go to the open modular concept of Figure 3.

  By selecting  the distance  between  tube
walls, the  required heat  exchange  surface
(wetted and viewed) could be provided, still
retaining a relatively large maintainable open
combustion space.  If fuel and air  to  each
module  were controllable, it would be possi-
ble to achieve a larger turndown.

  The first checkout of this system was per-
formed  on the boiler of Figure 2 by halving
the width (to 18 in.) and removing the tubes
on one side. Pending construction of an actual
single-module unit,  the modified boiler of
Figure 2 was operated to develop coal feeders
and light-off procedures.
Coal Feed and Flue Gas Travel

  The system which provided the most even
coal distribution was  an up-through-the-grid
method. However, with single-screened coal,
any moisture resulted in plugging of the feed
tube. Also, erosion at the turns was excessive.

-------
 Therefore, a straight-in  coal feed system was
 optimized and adopted  for use on the single-
 module  boiler (FBM). Excessive carryover of
 bed material was noted in all FBX test runs.
 This was attributed  to the  gases taking a
 horizontal path above the bed, entraining the
 material  without the allowance  for  particle
 recovery  provided  by  a  straight  vertical
 express gas path. This assumption was verified
 in later  FBM tests: a 6-percent ash coal held
 the  bed level; i.e., made up for attrition elutri-
 ation losses. The lessons  learned included:

   1. Do  not specially prepare the coal used in
     the development work; use coal identical
     to that to be used in practice.
   2. Do  not consistently feed as-received coal
     pneumatically if an  upward directed turn
     is  required,  at  least not where forced
     draft air is the motivating force. (Com-
     pressed air has not  been used because of
     economics.)
   3. Avoid horizontal movement of the flue
     gas  until it has passed through a screen
     such as a convection bank, superheater,
     or tubular air heater.
   4. Bed particles will carry over if gases take
     a horizontal turn  even  at  heights well
     above the theoretical disengaging height.
   5. A vertical express route for flue gas pro-
     vides for simple trouble-free design.
Initial FBM Test Results

  The FBM  design  (see Figure 4), based on
the  lessons learned in  the  FBX, utilized  a
direct coal injection system and incorporated
an  integral system  for fly  ash  reinjection.
Both coal  and fly ash were  fed  to the same
boiler bed.
  Figure 5 graphically presents a heat balance
during typical  operation. Note the high loss
from unburned carbon in  the fly ash.. Utiliz-
ing,  in  a  separate  operation, a  smaller
fluidized-bed column (the  FBC),  insulated so
as to reduce heat loss to the walls, collected
fly  ash  was refired,  and the  carbon loss was
reduced to an acceptable level. See Figure 6.
   The  result of  this  experiment  was  the
Carbon-Burnup  Cell—a separate  combustion
zone  or  segregated  region  for which  the
primary fuel is fly ash.

   During the initial  test phase,  the present
system of  light-off was  developed.  By
"puddling" a flame into one section of the
bed  surface,  ignition is accomplished,
followed  by  propagation  of the fire to all
portions of the  module through  openings in
the dividing tube wall.


   It is realized that acceptable carbon burnup
has also been achieved in a single fly ash rein-
jected combustion space by utilizing extensive
uncooled  refractory surfaces above the  bed
and  imbedded  heat exchange surface area,
utilizing the same  principle as the refractory
arch in  a conventional anthracite-fired boiler.
In other words, if the elutriating  unburned
carbon  leaving  the bed  can be held long
enough in a hot  non-radiant heat-removing
zone, the fly ash  carbon will burn out. A
column employing 80 percent  of  its total
height and volume as a refractory-lined, non-
heat-transferring,  elutriated-carbon  burnup
section  has demonstrated  acceptable carbon
loss. The combustion of carbon in this exten-
sive  freeboard area  results in a temperature
rise in that area, further enhancing the carbon
combustion  potential.  However,  such a
system  can only be applied to boiler design
employing a large  open "furnace" above the
bed, thus defeating the objective  of compact
combustion and lowered cost.
  In addition to the above size and cost con-
sideration, the integral (one combustion zone)
concept requires  multiple recycle of fly ash
with a consequent increase in dust loading to
the  collectors.  For  this  reason,  present
particulate emission standards are currently
forcing existing pulverized coal and spreader
stoker fired units to restrict recycle of fly ash.
The once-through separate carbon-burnup cell
precludes  the   "multiplication"  of  dust
loadings.
                                     IV-4-3

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Early Modular Boiler Concepts

   In addition to the initial concept of Figure
3, other preliminary designs based on initial
experimental work  were prepared, and  per-
formance predicted  for  efficiencies  ranging
from  85 to 93 percent. These early concepts
have included:

   l.A   250,000-pound-per hour  packaged
     boiler designed for the production of
     250-psig saturated steam.

   2. A   250,000 pound-per-hour  packaged
     boiler designed for the production of
     600-psig, 750°F steam
   3. A   2,000,000  pound-per-hour boiler
     designed  for  the  production  of
     2000-psig. 1000°F steam (and for reheat
     to 1000°F).
   4. An  8,000,000  pound-per-hour boiler
     designed  for  the  production  of
     2400-psig. 1000°F steam (and for reheat
     to 1000°F).
Further Development in the FBM

   One of the important  design features that
required checking  was  the effect  of bed
communication  between  adjacent modules,
particularly between a boiler cell operating at
a  relatively  low temperature and  a  carbon-
burnup cell (CBC)  operating at a  high tem-
perature.

   The FBM was modified and a small insu-
lated region was added. (See Figure 7.) While
this small column is  not a realistic CBC, which
would be much larger, it was the right size for
the quantity of fly ash generated by the FBM.

   To ensure that good  communication and
equal bed height were obtained, an open area
of 90 square inches was initially provided be-
tween the FBM and  the CBC.

   Initial  operation of the FMB/CBC was
poor. With  the  FBM  at 1600°F, the CBC
could  not reach 2000°F,  the  temperature
IV-4-4
required  for adequate  burnup.  The  major
cause was the rapid interchange between beds
through  the  90-square-inch opening.  When
there is  too much intercommunication be-
tween the two zones, the desired temperature
difference cannot be maintained.  Reduction
of the opening of 2 square inches was found
to be adequate  for light-off  and bed-level
equalization.  This  small opening permitted
the desired temperature to be achieved.

  A  second energy  loss resulted  from the
carryover of hot bed material into the hori-
zontal duct. It was also impossible to deter-
mine burnup performance since the collected
ash  was heavily  contaminated  with  bed
material.

  This carryover problem was solved by the
addition  of a BCURA-conceived baffle screen.
However, the screen acts  as  a  remarkably
effective  heat sink. A high (approximately 40
Btu/ft2hr °F) heat  transfer coefficient has
been computed for this device.

  A  vertical coal feeder has been tested. The
uniformity of coal distribution can be deter-
mined  by  measuring  oxygen  content  at
various points above the bed. Oxygen in flue
gas may be held to within ±0.2 percent across
the width and length  of the bed with coal
feeders as developed in this work. The tech-
nique lends itself to a simple coal delivery
system.  Figure 8  shows  the  present  FBM
boiler setup  in  the  Alexandria, Virginia,
laboratory.

  The lessons learned from this work include:

  1.  Very small (but properly located) open-
     ings  between regions permit  eq{ial bed
     levels and interbed light-off. ;    '
  2.  Tubular  baffle  screens,  properly posi-
     tioned,  serve  as effective  heat  ex-
     changers,  almost as good as  "wetted"
     surface, and may also be used in lieu of a
     high freeboard.
  3.  Downward-directed vertical  pneumatic
     coal  feeders provide  for  an even fuel
     distribution  over a relatively large area

-------
     and  greatly  simplify  the  problem  of
     boiler coal supply.
   4. Cold air, supplied by a forced draft fan,
     can provide all the motive force required
     to inject coal into a fluidized bed.

   Carbon  loss  from  the FBM/CBC,  when
employing  a  water-cooled  baffle  screen
directly  above  the  bed, is  as high as that
found in a typical spreader stoker operation
(see Figure 9). The present CBC is simply too
short* and the baffle screen is too effective as
a  heat exchanger. By replacing the water-
circulated  baffle  with uncooled alloy  steel
rods, satisfactory  combustible loss (0.8 per-
cent, see Figure 10) has been attained. Laying
out and equipping the unit with total freedom
could  have provided  a  design for a  thermal
efficiency in the 90-percent range. Although
the  present  FBM/CBC  will not achieve  an
overall thermal efficiency much  greater than
80 percent, much has been learned from  its
operation;  contemplated  revisions  do not
include a departure from the basic modular
concept.
PRESENT DESIGN CONCEPTS

  Present design concepts are:

   1. An open fluidized-bed combustion space
     should be utilized so as to minimize coal
     distribution problems  and to  provide
     adequate access for boiler maintenance.

  2. Coal  injectors should  be served with
     motive  air supplied   directly  from a
     forced-draft far.

  3. Heat transfer  surface  should  be propor-
     tioned betveen  direct-contact, viewed
     surface, and baffle surface so as to trans-
     fer  directly from the bed 50-60 percent
     of the  energy released in burning  the
     fuel.

* The CBC, added 2 years after the FBM was installed, had to
 be located beneath  the steam drum. The distance between
 grid and horizontal discharge is only 40 in.
   4. Flue gas should  follow  a long vertical
     express route, including  convection and
     economizer banks.

   5. The  design of the boiler should  permit
     all boiler cells to communicate with the
     CBC, especially  if a regenerative SC>2-
     capture method is to be used.

   Based on  these concepts, a 300,000-lb/hr
shippable unit  has  been conceived for an
eastern utility.  (See  Figures 11 and 12.) The
modules (cells) run parallel to the steam drum
and connect to a single CBC. At least until
additional experimental work  can be done on
coal-feed techniques, a 4.5-ft spacing for coal
feed points seems advisable. The design shows
eight fuel-receiving points,  each serving two
adjacent locations. Coal  and  limestone feed
(where used for  SC>2 abatement) are com-
bined. Primary  superheaters are shown in the
bed,  but  may  also be arranged  to serve  as
baffle screens above the bed. When a module
is shut down, superheating stops, providing an
automatic  attemperation mechanism.  Some
external attemperation will still be required.
Fly ash from the boiler cells is collected in the
primary cyclone  and  fed  to the  CBC  via
parallel circuits to four  mushroom feeders.
There is also provision for auxiliary coal addi-
tion.  No point in the CBC bed is farther than
2 feet from a feed point.  Control of the CBC
is based on maintaining constant bed tempera-
ture.  The CBC is cooled by the secondary
superheater,  taking advantage of the  higher
available  temperature differentials  in the
section. Preliminary predicted data for this
unit is outlined in Table 1.

   Referring to  Figure  11, it should be noted
that two of the vertical bed-immersed  tubes
extend into the bed,  completely surrounded
by the bed particles. As suggested by Seibel of
Erie City, a few rows of tubes (say 2  or 3)
could  extend from the walls, forming a re-
stricted tube bank,  enhancing heat transfer
but not interfering  with the  dispersion of
coal feed  or the maintenance accessibility of
the boiler. This would permit widening the
cells (reducing  the required number of cells
                                     IV-4-5

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               Table 1.  PRELIMINARY DESIGN DATA FOR 300,000-POUND-PER-HOUR
                      FLUIDIZED-BED UTILITY BOILER (1500-PSIG DESIGN,
                                  1270-PSIG OPERATING, 925°F)
Service
Final superheater in CBC
(In-bed)
(Above-bed)
Primary superheater
in boiler cells
Evaporation
(In-bed)
(Particle recovery
screen)
(Convection)
Economizer
Regenerative air
heater
Totals
Duty,
MBtu
11.2
5.4
64.0
98.4
74.9
25.1
48.0
35.4
362.4
AT diagram,
°F
2000*^2000
860 -»• 925
2000 -M 270
860*- 834
1600*+1600
575 -*• 834
1600*+1600
575*+575
1600**1600
575** 575
1 600 -M 270
575** 575
1270-> 709
525*- 385
709 -> 274
70-* 550


Tubing data
Surface
ft2
223
203
1,640
1,930
1,820
2,980
9,480

18,276
Length, ft
(O.D., in.)
430
(2)
390
(2)
3,130
(3)
2,460
(3)
1,460
(3)
1,280
(2)
5,700
(2)
18,100
(2)

29,030 (2)
3,920 (3)
               Bed data

                 Grid area, ft2
                 Air rate, Ib/hr ft2:
                    superficial velocity, fps
                 Boiler section
                     364
                           CBC section
               Fan data

                 Draft loss,
                   in. w.g.
Grid
Bed
        16
      794:11.7
 Convection
& economizer   Air heater
 45

730:12.8

       Total
                                       31
               Fan power @ 76,560 SCFM, 31 in. w.g. = 498 HP
               Heat balance data, MBtu:  Fuel input = 358.6; stack loss 16.6;
                 unburned combustible = 7.2; radiation and misc. = 7.4; boiler efficiency = 91.8%.
IV-4-6

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for a given capacity) and/or permit the reduc-
tion of bed  level  (and draft  loss).  Arrange-
ments incorporating this concept are  under
study  and will be incorporated in the design.
  A preliminary cost estimate has been pre-
pared  based on post-development conditions
but  using  1970 costs.  This  estimate, sum-
marized in Table 2, shows a basic boiler cost
                    Table 2. PRELIMINARY COST ESTIMATE FOR 300,000-LB/HR
                                 FLUIDIZED-BED UTILITY BOILER
                    1. Basic boiler
                      a) Steam drum                   106,500 Ibs
                      b) Drum headers                   6,400 "
                      c) Drum internals                  4,100 "
                      d) Tubes                         51,600 "
                      e) Lower headers                  28,800 "
                      f) Casing                          3,000 "
                      f) Contingency                    59,600 "

                           Total weight                260,000 Ibs

                      Estimated cost @ $1.32/lb        $343,200
                      Add 50% for special construction   171,600

                      Total basic boiler

                    2. Boiler appurtenances
                      a) 409 ft2 of air distributor       $ 40,000
                      b) Economizer (30,000 Ib)          30,000
                      c) Air heater (120,000 Ib)          55,000
                      d) Boiler trim                     25,000
                      e) Coal/ash/additive supply        200,000
                      f) Ducts/plenum/breeching         40,000
                      g) Insulation (30,000 Ib)            60,000
                      h) F.D. fan and  drive               40,000
                      i) Collectors (2)                  30,000
                      j) Electrostatic  precipitator        290,000
                      k) Controls and instrumentation    120,000

                      Total boiler appurtenances
                    3. Normal design; installation and
                        contingency
                      a) Installation (includes miscel-
                            laneous piping and electrical   350,000
                      b) Contingency                  200,000
                      c) Normal design                 120,000

                      Total design, installation, and
                         contingency

                    Total normal installation at existing plant
                    Add 25% for first prototype unit

                    Estimated prototype facility cost	
                    $514,800
                    930,000
                    670,000

                  2,114,800
                    528,700

                 $2,643,500
                                                                                           IV-4-7

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 of  $514,800  and  a  total  plant  cost of
 $2,114,800. For the first installation, we have
 applied  the  multiplier of  1.25  to  the esti-
 mated cost for post-development conditions.
 The cost for the first unit is, therefore, esti-
 mated at $2,643,500.
 CONCLUSIONS

   During the past several years, a number of
 design arrangements have been proposed for
 fluidized-bed boilers. Three true  boilers have
 actually been built: two by Pope, Evans and
 Robbins; followed  by BCURA's  sheU-and-
 tube  unit.  PER's "FBM" is the only unit
 currently in service.  In two cases it was found
 necessary  to remove some  or  all  of the
 imbedded tubes,  since at  the time these two
 designs  were  completed, the  removal  of
 energy  from the bed by mechanisms other
 than  direct  contact heat transfer was  not
 anticipated. All three of these actual boilers
 have  exhibited unacceptable carbon losses.
 The use  of a  segregated fly  ash carbon-
 recovery section  operating at high tempera-
 ture (2000°  F)  has been demonstrated as a
 means  of recovering the  unburned  carbon
 without resorting to low velocities and exces-
 sively large bed areas (pressurized  fluidized-
 beds excepted).

   The use of extensive closely spaced steam
 generating tube  banks immersed in the bed
 (particularly horizontal banks  requiring
 forced circulation) are felt to be  an unneces-
 sary  complication  for  most fluidized-bed
 boilers.  The  arrangement  of tubes in single-
 row banks  so as  to form  walls has been
 demonstrated in  the  FBM  to  be a useful
 design concept. The modular approach is the
 only concept that has thus  far been exten-
 sively  tested  under actual  operating  con-
 ditions.
  In summary, it is felt that the design con-
cepts  proposed  here  offer  the following
advantages:


     1. Maximum flexibility.
     2. Simplest coal and sulfur acceptor feed
       capability.
     3. Maximum  accessibility  for  mainte-
       nance.
     4. Good turndown capability.
     5. No requirement  for an  induced draft
       fan.
     6. The use of natural, rather than forced,
       circulation.
     7. Compact construction by the use of
       high-mass  gas flows  and heat release
       rates.
     8. Low bed particle carryover; ability to
       operate with low-ash coal.
     9. Segregated fly ash combustion region;
       capable  of  operation  at optimum
       burnup  conditions, without  com-
       pounding  the  particulate  emission
       problem. (This feature  alsp  obviates
       the need  for large furnace-like  free-
       boards for the  consumption of un-
       burned carbon.)
    10. Capability of using a "built-in" lime-
       stone regeneration process.
    11. Low capital costs.


  Regardless of  the  direction taken  by
fluidized-bed boiler  development, it is  felt
that experience should be effectively dissemi-
nated and considered. There can  be  no good
reason  for reintroducing  design   features
already  shown to be  deficient—unless a new
"twist" has  been  conceived to obviate  the
previously  demonstrated  deficiency. Boiler
design  features  must be based  .upon  data
derived  from large-scale experimental equip-
ment duplicating, or very closely simulating,
the proposed arrangement.
IV-4-8

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Figure 1.  Plastic model, packaged fluidized-bed boiler.

-------
                                                 STEAM OUTLET
                                       Dot
                              COMBUSTION ZONE  0
                                    3ft.    "    0
                                                                      OVERALL LENGTH 7 ft.
                                                                     BOILING WATER JACKET
                                                                    FLUIDIZED BED
                                                                       THERMOCOUPLE
                                                                            AND
                                                                  MANOMETER CONNECTIONS
                                                                   GRID

                                                                   COAL FEED SCREW
     Figure 2.  Imbedded-tube atmospheric-pressure boiler with 3 x 6 ft combustion zone.

IV-4-10

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                                                                           TO DUST
                                                                           COLLECTOR
                                                                                                   FEEDWATER
                                                                                                   INLET TO
                                                                                                   ECONOMIZER

                                                                                                   FEEDWATER
                                                                                                   INLET TO
                                                                                                   DRUM
                                                        FLUEGAS
                                                        COLLECTION
                                                        BREECHING
                   REINJECTED
                   FLY ASH AND
                   ADDITIVES
   CONVECTION
   BANK
                                                                                             HIGH. LOW
                                                                                             COAL LEVEL
                                                                                             SWITCHES
                                                                     LIGHT-OFF
                                                                     BURNER
PNEUMATIC
COAL FEEDER
TUBES
CROSS HEADER
                                           CROSS HEADER
                                          COAL FEEDER
                                             TUBES
AIR SUPPLY
PLENUM
                         AIR DISTRIBUTION
                         GRID
                   OXIDIZING
                   FLUID BED
                   (SHOWN IN
                   ONE CELL ONLY!
              Figure 3-  Industrial modular boiler (original concept).

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                       Figure 4.  Fluidized-bed boiler module (FBM).
     BOILER ABSORPTION
      4.450 KBtu (54.8%)
    AIR PREHEATER CYCLE
    540 KBtu (6.6%)
       — FUEL INPUT —
        8,130 KBtu (100%)
   LOSSES
RADIATION AND
UNACCOUNTED
 140 KBtu (1.7%)
                          CARBON LOSS
                          1.240 KBtu (15.2%)
 /
                             FLUE GAS
                             2,300 KBtu (28.2%)
Figure 5.  FBM heat balance without carbon
recovery (boiler convection and economizer
heat transfer deleted).
BOILER ABSORPTION
4,450 KBtu (59.7%)
                          AIR PREHEATER CYCLE
                          500 KBtu (6.7%)
    LOSSES
RADIATION AND
UNACCOUNTED
480 KBtu (6.4%)
                                                 CARBON LOSS
                                                 70 KBtu (0.9%)
                       /
                      /
   FLUE GAS
   2.460 KBtu (33.0%)
                            	 FUEL INPUT	
                              7,460 KBtu (100%)

                      Figure 6.  FBM heat balance with carbon
                      recovery (boiler convection and economizer
                      heat transfer deleted).
IV-4-12

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                                                                   COOLING WATER INLET
                                                                       STOP-SLOP SCREEN
                                                                  COOLING WATER OUTLET
                                                                       TWO 1 x 2 m.
                                                                   INTERCOMMUNICATION
                                                                    SLOTS (NOT SHOWN)
                                                                     ACCESS DOOR
                                                                      MUSHROOM FEEDER
                                                                           (FLY ASH)

                                                                       AIR DISTRIBUTION
                                                                             GRID
                                                                        CBC PLENUM
                 0'      T     2'     3'      4'     5'

                           GRAPHIC SCALE

                      Figure 7. Present boiler arrangement.
Figure 8-  Test boiler, Pope, Evans and Robbins Alexandria, Virginia, laboratory.
                                                                                IV-4-13

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      BOILER ABSORPTION
       6,186 KBtu (82.~8%)~
               LOSSES

             RADIATION
             100 KBtu (1.2%)
        COMBUSTIBLES
        265 KBtu (3.4%)
                            ASH AND B.D.
                            12 KBtu (0.2%)
AIR PREHEATER CYCLE
512 KBtu (6.9%)
                                    GAS
                               907 KBtu (12.2%)
       	   FUEL INPUT   	
         7,470 KBtu (100%)
 Figure 9.  FBM test B-5 heat balance (with
 integral  CBC water-cooled stop-slop screen).
                                               BOILER ABSORPTION
                                                4,926 KBtu (73%)
    LOSSES

RADIATION
179 KBtu (2.7%)
                                                            ASH AND B.D.
                                                            60 KBtu (0.9%)
                                         COMBUSTIBLES
                                         56 KBtu (0.8%)

                                      AIR PREHEATER CYCLE
                                      300 KBtu (4.5%)
                                                                            7_
                                                             •FLU EG AS
                                                              1,529 KBtu (22.6%)
                                        	  FUEL INPUT   	
                                         6,750 KBtu (100%)
                                Figure 10.  FBM test B-14 heat balance (with
                                integral CBC, uncooled stop-slop screen).
 COMBUSTION AIR
 FROM F.D. FAN
           EXIT GAS
—I    (-*• TO PRECIPITATOR
  f    |    OR SCRUBBER
 MAIN DUST
 COLLECTOR
 TO CBC
 REINJECTION
 PREHEATED
 AIR DUCT
                          RUNAROUND REDLER CONVEYOR
                           STEAM DRUM
                   SATURATED SfEAM
                   FEEDERS TO SUPERHEATER
                   PRIMARY SUPERHEATER
                   INLET HEADER
                                                               COAL/ADDITIVE
                                                               DROP LINES
                                                                           CBC DUST
                                                                           COLLECTOR
                                                 SUPERHEATER
                                         OUTLET HEADER
                                       PLENUM
                                                                                 CBC
                                                                                 ECONOMIZER
                                                                                 SURFACE
                                                                         INLET FROM
                                                                         PRIMARY SUPER-
                                                                         HEATER OUTLET
                                                                          SECONDARY
                                                                     «--*"* SUPERHEATER
                                                                 FINAL
                                                                 SUPERHEATER
                                                                 OUTLET
                                                                           FLY ASH FROM MAIN
                                                                           DUST COLLECTOR
     Figure 11.  Factory-assembled, coal-fired, fluidized-bed utility boiler (300,000 Ib/hr,
     1270 psig, 925 °FTT) side view.
IV-4-14

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                                           REDLER CONVEYOR
                                                         FORCED-DRAFT AIR HEADER



                                                         COAL FEEDER



                                                                   SIGHT PORT



                                                                     ECONOMIZER

                                                                     COAL SUPPLY TUBE
                                                                     CONVECTION BANK

                                                                     SATURATED
                                                                     STEAM FEEDER
                                                                     TO SUPERHEATER
r                                                                      PRIMARY
                                                                      SUPERHEATER
                                                                      INLET HEADER
                                                                      PRIMARY
                                                                      SUPERHEATER

                                                                      PRIMARY
                                                                      SUPERHEATER
                                                                      OUTLET HEADER
                                                                        LOWER HEADER
                                                                        (ONE OF THREE)
Figure 12. Factory-assembled, coal-fired, fluidized-bed utility boiler (300,000 Ib/hr,
1270 psig, 925 °FTT) front view.
                                                                              IV-4-15

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                                5.  FLUIDIZED-BED BOILERS-
                                                         CONCEPTS AND
                                                          COMPARISONS

                   D. L. KEAIRNS AND D. H. ARCHER
                    Westinghouse Reserach Laboratories
INTRODUCTION

   Fluidized-bed  boilers offer  economic and
social benefits of compact, cheap,  and effi-
cient power generation systems; clean air; and
effective use  of natural  resources  and raw
materials for  fuel. Westinghouse, as  part of a
contract with  the  National  Air Pollution
Control  Administration,  is  preparing  and
evaluating boiler designs in order to provide a
technical and  economic measure of the poten-
tial of fluidized-bed combustion boilers for air
pollution control. These concepts must then
be  compared between themselves and  with
conventional boiler or power systems.

   Both  new and existing concepts are being
considered.  Fluidized-bed boiler  concepts
have been  generated by considering various
fluidized-bed  configurations  and operating
conditions. Operating conditions are selected
based on the evaluation of fluidized-bed com-
bustion  data, fluidized-bed  technology as
applied  to other processes, and  boiler and
power  system  requirements.  The  critical
fluidized-bed  combustor  parameters are cal-
culated from heat and material balances, heat
transfer  relations, kinetics of combustion and
desulfurization, and fluid flow relations given
the configuration, operating conditions, pollu-
tion control limitations, and customer specifi-
cations.  These  parameters  include  cross-
sectional area  of the bed, heat transfer surface
requirements,  bed height, bed composition,
pressure drop through the system, solids and
gas  distribution,  and auxiliary  equipment
specifications.
  The  projected investment and operating
cost and  pollution control capabilities of the
fluidized-bed boiler power systems considered
can be compared with each other and  with
other power systems. The best power system
is the one  which meets  customer specifica-
tions and pollution control limitations at the
minimum cost.

  A pressurized boiler power system is  used
to  illustrate  a  fluidized-bed  concept  and
comparison with other power systems.  The
high-pressure  fluidized-bed boiler  power
system utilizes  the potential advantages of
fluidized-bed  combustion over conventional
combustion systems and has several potential
advantages over an atmospheric fluidized-bed
boiler system:

  1. Boiler  size  is reduced. The potential for
     shop fabrication of utility boilers holds
     promise for  large  cost  savings—both
     materials costs and field erection costs.
  2. Power cycle efficiency can be increased.
  3. Problems of solids  handling and good
     distribution of fuel and air are reduced
     since the fluid bed area is small.
  4. Scale-up problems are minimized.
  5. Heat transfer surface is reduced by using
     a combined gas steam-turbine cycle.

  Preliminary design calculations for a high-
pressure fluidized-bed boiler power system are
presented to illustrate the development of the
concept and the identification of critical
fluidized-bed  combustor parameters.  The
design  and  evaluation of this  concept are
                                       IV-5-1

-------
proceeding based on these parameters. Tech-
nical and economic evaluations may result in
parameter  changes in order  to  achieve  an
optimum design.  Power plant efficiency and
cost for the pressurized  system have been
calculated  based  on projected boiler effi-
ciency and cost.  The projected high-pressure
system performance is compared with a con-
ventional coal-fired power plant.
                       Fluidized-bed boiler power  systems have
                    the potential for burning a broad range of low
                    grade fuels, including solid wastes.

                       This provides a means for the effective use
                    of natural  resources  and  raw  materials.
                    Thermal discharge to cooling water may also
                    be reduced  if a  pressurized  combined gas
                    steam cycle proves economic and operable.
DESIGN BASIS

  The development and  evaluation of  flu-
idized-bed boiler  concepts  must  consider
pollution control projections; market projec-
tions;  power  cost analyses;  fluidized-bed
combustor operating and design parameters;
and  power system operation,  maintenance,
and reliability.
                    Market Projections

                       The impact of fluidized-bed combustion
                    with sulfur absorption in the bed on air pollu-
                    tion abatement  and on the economics of
                    steam  and  power  generation  has  been
                    studied.1  Results of the market studies have
                    been  used to  establish  fluidized-bed  boiler
                    specifications. These  results are summarized
                    elsewhere.1
Pollution Control Projections

  Projected  uncontrolled  emissions from
power plants  and  projected air-quality goals
for SC>2,  NOX, and  particulates have been
considered.1 Power plant design requirements
for the year 2000, based on  these projects,
might be:
        Emission

     SO2

     NOX

     Particulates
Control Level
Leaving Stack

lOOppm

100 ppm

O.OlOgr/scf
Control targets, for small .particles (< 5 micro-
meters) have  not been projected.  However,
regulations are anticipated which will require
significant  reductions from  current emission
levels.  Information indicates  that  fluidized-
bed combustors, operating with a coarse coal
feed, may reduce small particle production by
'v 60 wt percent.
Power Cost Analyses

  Investment  and operating costs projected
for fluidized-bed boiler power systems must
be  compared  between themselves and  with
conventional and  proposed power systems.
The fluidized-bed  boiler  power system must
have an energy cost which is competitive with
those of nuclear plants and fossil fuel plants
with  air  pollution control. Present  power
plant investment  costs are $186/kW for  a
600-MW coal plant without pollution  control
and $269/kW for a 1100-MW nuclear plant.
Operating and Design Parameters

  Operating  and design  variables for  a
fluidized-bed boiler system must be selected.
Table  1  is a list of some of the important
variables which must be considered. A given
concept may not require consideration of all
the parameters listed. However, careful con-
sideration of these parameters, based on the
analysis of available  data and technology, is
important  if an optimum  design  is to  be
achieved.
IV-5-2

-------
                  Table 1. OPERATING AND DESIGN PARAMETERS FOR DESIGN OF
                                      FLUIDIZED-BED BOILERS
                                    INDEPENDENT VARIABLES
                Operating variables
                                                           Design variables
                Fuel
                Sulfur absorbent
                NOX control
                Bed temperature
                Pressure
                Gas velocity
                Excess air
                Particle size
                Particle flow
                Bed height
Reactor configuration
Arrangement of boiler functions (e.g.
   economizer, superheater,
   regenerator, solids removal,
   carbon burnup cell)
Heat transfer surface (tube size,
   orientation, arrangement)
Air distribution
Solids handling
Water steam flow
Sulfur absorbent regenerator
                                     DEPENDENT VARIABLES
                Operating effectiveness

                Pollution control
                Combustion
                Sulfur sorbent utilization
                Fluidization
                Heat transfer
                Elutriation
                Attrition
                Corrosion
                Erosion
                Agglomeration

                Economics

                Tubing and  fabrication
                Structures
                Controls
                Auxiliaries
                Water circulation
                Maintenance
                Steam headers
                Construction time
                Solids preparation and distribution
                Air preparation and distribution
Control
Turndown
System efficiency
Solids distribution
Air distribution
Particle diffusivities
Reliability
Pressure drop
Heat release
Operation, Maintenance, and Reliability


   Any new  boiler design  must  consider the
maintenance,  operability,  and  reliability
requirements of  the  system. The importance
of these factors is clear in light of the present
day brownouts  and blackouts. The ability to
achieve  part-load  operation  with  rapid re-
sponse time is an important consideration for
intermediate load generation which must vary
in power production throughout the day.

                                        1V-5-3

-------
 A Basis for Evaluation

   These constraints and guidelines provide a
 basis for the development, analysis, and evalu-
 ation   of  fluidized-bed  boiler  concepts.
 However, they  must be  viewed with  flexi-
 bility.  An example is the  projected pollution
 control  standards.  Present  data  on the
 removal of sulfur dioxide in a fluidized-bed
 combustor  desulfurizer is not  adequate to
 permit  the design of units to remove greater
 than  approximately 90-95 percent  of the
 sulfur. Thus, for a coal containing 4.3 percent
 S, it is not possible to design for control  levels
 below  M80 ppm with reliability until further
 data is available.
DESIGN CONCEPTS

   Several boiler systems  have been built,
operated, or proposed which incorporate
fluidized-bed combustion. Concepts go back
to 1928 when Stratton developed  a spouting
fluidized-bed boiler.2  The early boiler con-
cepts incorporating fluidized-bed combustion
were  generally developed  for  burning low-
grade fuels. These designs did not recover heat
or consider sulfur removal in the fluidized-bed
combustor.  In the last decade,  designs which
incorporate  heat recovery and sulfur removal
in the  fluidized-bed  combustor  have been
conceived  in England and in the  United
States.

   In order to utilize the potential  benefits of
fluidized-bed combustion in minimizing air
pollution while minimizing  energy cost, new
and  existing concepts have  been considered.
Atmospheric and pressurized  systems have
been  considered  which incorporate  various
fluidized-bed  configurations  (single beds,
stacked beds, unmixed beds, and  segmented
beds) and  system  component arrangements
(solids  feeding  and  removal,  final  carbon
burnup,  sorbent  regeneration,  heat transfer
surface, and boiler functions).
IV-5-4
  One concept is the high-pressure fluidized-
bed boiler combined gas steam cycle system
which  offers both large potential savings in
investment and operating costs and effective
pollution  control. The concept illustrates the
development and analysis of a fluidized-bed
boiler system. Preliminary design calculations
have  been  completed  for the  pressurized
utility boiler system shown in Figure 1.


  Power  system performance  studies and
boiler cost and performance analyses for the
system presented in  Figure  1  show that the
system should operate at 10 arm, using sub-
critical steam conditions. Preliminary design
calculations have been  made  for a 600-MW
pressurized utility boiler system operating at
10 atm with once-through subcritical steam.
The  operating conditions are  summarized in
Table 2.

  Figure  2 is a process flow diagram for the
pressurized utility boiler. The system consists
of a primary fluidized bed combustor desul-
furizer,  cyclones for particulate removal,  a
carbon  burnup  cell,  CBC,  to achieve high
combustion efficiency, and a limestone re-
generator. The sulfur  removal system being
considered would feed regenerated stone from
the regenerator to the primary fluidized bed.
The  regenerator is  based  on  the  system
developed  by  Esso.4   The high  efficiency
primary multicyclone  removes 92.5 percent
of the solids  from  the flue gas  leaving the
boiler. The collected  solids are  fed to the
CBC, the combustion  efficiency of which  is
90 percent. The flue gas from the CBC is com-
bined with the  effluent  from the  primary
cyclone  and  sent  through  the  secondary
multicyclone, which removes  97 percent of
the solids, to the gas turbines. The  combus-
tion efficiency is approximately 98.5 percent;
1 percent of the carbon fed is assumed to be
carried out in the flue gas and the CO concen-
tration in the flue gas is M).2 percent. Table 3
summarizes the material balance for the boiler
and  limestone regenerator. Figures 3  and  4
summarize the energy balance for the boiler
and limestone regenerator.

-------
        Table 2. PRESSURIZED FLUID-BED BOILER POWER SYSTEM
                        OPERATING CONDITIONS
Fluidized-bed combustor/desulfurizer
Pressure
Temperature
Fuel


H.H.V
Excess air
S02 removal
agent

10atm
1750°F
Pittsburgh
No. 8
seam coal
1 3,000 Btu/lb
10%

Limestone

Estimated limestone
feed to achieve
95% removal

Primary bed
material
Solids elutriation
Ash
CaO/CaSC>4

8 times
stoichiometric
CaO to react
with S02
S02 removal
agent

100%
M).5% of bed
weight/hr
Carbon burnup cell
     Pressure
10atm
Limestone regenerator
     Pressure        10atm
     Temperature    2060° F
Steam conditions
     Steam flow
     Final feed-
        water temp
     Superheater
        outlet press.
     Superheater
        outlet temp
     Reheat steam
3,500,000 Ib/hr
578° F
2500 psi

1000°F
3,200,000 Ib/hr
Gas turbine system
     Air temp
        leaving
        compressor   636  F
     Flue gas temp
        to gas
        turbine(s)    1600°F
     Flue gas temp
        to upper stack
        gas cooler    831"F
Temperature
                 Regeneration
                    of CaSC>4
Reheater
   inlet temp
Reheater
   outlet temp
Reheater
   inlet press.
Reheater
   outlet press.
                  Flue gas temp
                    to lower stack
                    gas cooler
                  Flue gas temp
                    leaving system
2000 F
                    ~90%
                                     650 F
                                      1000°F
600 psi

580 psi
                    525° F

                    275° F
                                                                              IV-5-5

-------
Table 3. MATERIAL. BALANCE FOR PRESSURIZED UTILITY BOILER

Stream
No.a
1-1







2-1
2-2
2-2A
2-3




2-3A




2-4

Description
Ohio seam,
Pittsburgh #8 coal






Air
Air to FBCD
Air to CBC
Flue gas from FBCD




Flue gas from CBC




Solids to CBC from
primary cyclone(s)
Flow rate
Solids
tons Air

207







	
	
29.4




19.9





27.2
Gas
moles/hr

	 ,






160,000
151,400
8,600
158,900




8,600





	
Composition
Solids weight %

C 71.2
H 5.4
0 9.3
N 1.3
S 4.3
Ash 3.5
100.0

	
	
Ash 59.9
C 29.9
Spent
limestone 10.2
100.0
Ash 81.9
C 4.0
Spent
limestone 14.1
100.0

Same as 203
Gas mole %

	






N2 77.5
O2 20.4
H2O 2,1
100.0
Same as 2-1
Same as 2-1











	
Temperature
°F

	






636
636
636
M750




-V2000





M 700

-------
Table 3 (continued). MATERIAL BALANCE FOR PRESSURIZED UTILITY BOILER
Stream
No.a
2-5
2-6
2-7
2-8
3-1
3-2
3-3
3-4
Description
Flue gas from
primary cyclone(s)
Combined flue gas
to secondary
cyclone(s)
Flue gas to gas
turbine(s)
Waste solids
Regenerated lime-
stone to FBCD
Spent limestone
from FBCD
Regenerated stone
from regenerator
Waste stone
Flow rate
Solids
tons/hr
2.2
22.1
0.66
(0.15gr/SCF)
21.4
139.3
145.6
123.3
12.3
Gas
moles/hr
158,900
167,500
167,500
	

	
	
	
Composition
Sol ids weight %
Same as 2-3
Ash 79.7
C 6.6
Spent
limestone 13.7
100.0
Same as 2-6
Same as 2-6
CaO 77.4
CaC03 20.4
CaSO4 2.2
100.0
CaO 73.5
CaSO4 26.5
100.0
CaO 97.2
CaS04 2.8
100.0
Same as 3-3
Gas mole %

N2 74.2
C02b 15.0
CO 0.2
H2O 8.7
02 1.8
NO "v/400 ppm
S02 M70 ppm
100.0
Same as 2-6
Same as 2-6

	
	
	
Temperature
°F
M700
'vieoo
1600
1600

M750
^2000


-------
00
                Table 3 (continued). MATERIAL BALANCE FOR PRESSURIZED UTILITY BOILER
Stream
No.a
3-5
4-1

5-1
5-2
Description
Makeup limestone
Coke

Air
S02 rich flue gas
Flow rate
Solids
tons/hr
28.4
9.2

	
2.2
Gas
moles/hr
	
	

6,070
6,800
Composition
Solids weight %
CaC03 97.0
C 98.0
Ash 2.0
100.0
CaO 72.7
CaS04 18.2
Ash 9.1
100.0
Gas mole %
	
	

Same as 2-1
N2 70.1
CO2 22.4
CO negligible
S02 7.5
100.0
Temperature
°F




^2000
                 aSee Figure 2.
                 ''Includes CO2 produced from calcining makeup limestone.

-------
  Figure  5, a temperature-enthalpy diagram
for  the  boiler,  graphically  represents  bed
temperatures, water and steam temperatures,
estimated boiler tube-wall temperatures, and
the  energy breakdown of the system. Table 4
summarizes the  heat transfer surface require-
the preliminary  design since reliable  data is
not  available  to design  for sulfur removal
greater  than  90-95 percent.  Techniques for
reducing nitrogen oxides have not been con-
sidered  in the initial design. Design specifica-
tion  for the particulars removal system after
            Table 4.  PRESSURIZED FLUID-BED BOILER OPERATING CONDITIONS
                                 AND DESIGN PARAMETERS
Function
Heat transferred,
Btu/hr x 10*
Overall heat
transfer coeff ,
Btu/hr-ft2-°F
Surface
requirements, ft2
Inlet tube wall
temperature
estimates, °F
Outlet tube wall
temperature
estimates,0 F
Pre-evaporator
FBCD
6.3

47
12,000

660

730
CBC
1.2

47
1.900

750

750
Evaporator
8.8

47
17,200

730

730
Superheater
14.1

43
36,400

820

1110
Reheater
6.1

40
16,700

870

1150
Total
36.5

—
84,200

-

-
ments for the  boiler and  estimates of  the
tube-wall temperatures.  The maximum tube
wall temperature  estimate is 1150°F. These
calculations are based on a bed-to-tube heat
transfer coefficient  of 50  Btu/hr-ft2-°F,
which is considered pessimistic.

   Table 5 summarizes the performance of the
pressurized-bed  boiler power system. The  net
power production is 635 MW, with the steam
cycle producing MJ5 percent  of the power.
The boiler efficiency, 88.6  percent, is based
on conventional  procedures  for calculating
boiler performance. A term has been added to
account for the loss from sensible heat of the
spent  limestone leaving the bed. Pollution
control requirements for  the  preliminary
design  concentrated  on sulfur  removal. A
basis of 95 percent removal was selected for
the fluidized-bed boiler is based on particulate
loading requirements  for the gas  turbine.
Results from the BCURA high-pressure boiler
erosion tests have been used as  the basis.4
Their results indicate that particulate loadings
of 0.15 gr/scf may be permitted in the gas to
the turbine. Additional tests and analyses are
required to determine the effect  of coal ash
and spent limestone or other SC>2 sorbent on
turbine blade life. Final dust removal would
be added after the gas is cooled to achieve the
0.01 gr/scf stack emission level.
   Details  of the  fluidized-bed combustor
 desulfurizer design  have been considered
 based on the operating conditions and design
 parameters. A shop-fabricated modular design
 is  being considered  for the  high-pressure
                                     1V-5-9

-------
boiler in  order to  achieve  reduced  boiler
erection time and added savings in construc-
tion  costs. Rail shipment requirements are
approximately 12 x  16 x 40 ft. With a vertical
unit, the active fluidized-bed cross-sectional
area  can  be  50-60  ft2.  The number  of
modules and the bed depths will  be  deter-
mined by  the heat transfer  surface require-
ments,  heat transfer  surface configurations,
operating and  structural  considerations, and
cost. Table 6 shows that the effect of tube
bundle  design on the volume requirement  is

  Table 5. PRESSURIZED FLUID-BED BOILER
       POWER SYSTEM PERFORMANCE
Power

  Steam turbine
  Gas turbine
  Plant requirement
  Net power

Efficiency

  Plant heat rate
  Plant efficiency
  Boiler system losses3

     Dry gas (based on stack
       temperature)
     Hydrogen and moisture
       in fuel
     Moisture in air
     Unburned combustible
     Radiation
     Sensible heat of solids
     Unaccounted for losses
          Total losses

Boiler efficiency

Pollution control

  S(>2 in stack gas
  Paniculate loading
     to gas turbine(s)
  538.4 MW
  113.1 MW
   16.1 MW
  635.4 MW
8975 Btu/kWhr
 38%
    3.88%

    4.14%
    0.08%
    1.51%
    0.15%
    0.11%
    1.50%
   11.37%

   88.6%
  160ppmb

    O.l5gr/scf
aBased on heats of reaction for coal-02 and
 M 3300 Btu/lb coal.
^Emission  is less  than that from boiler due to air
 added to gas turbine system for cooling.
important  for large  utility boilers. If a single
vertical deep-bed  concept were  adopted for
each  module  and if the tube bundle design
consisted of horizontal  1-in. tubes on a 2-in.
triangular pitch, the boiler would require five
shop-fabricated  vessels  with  20-30 ft  deep
beds, each  serving a separate function. If a
stacked bed design  were adopted with each
module  containing  each  of the  respective
boiler functions, six vessels would be  requir-
ed; the extra vessel would be required because
of the increased space  requirements  for gas
distributors and freeboard.  The number of
vessels required for  each concept is doubled
by going to 2-in. tubes with a 4-in. pitch. The
boiler would then require 10-12  vessels. With
this design, shop fabrication must be weighed
with  field  erection  of a larger unit(s).  Con-
struction  costs,  external  piping,   solids
handling,  and  control  problems must  be
studied in order to determine the best system.
The optimum heat transfer arrangement will
be selected based on steam-side pressure drop,
tube costs, header design, and fluidization, in
addition  to the effect of boiler tube configu-
ration on bed volume.
  Variation  in pressure drop  through the
boiler system has little effect on plant per-
formance for the pressurized power system. A
1-percent  increase  in  the  pressure  drop
through  the  boiler  system results  in  a
0.16-percent  decrease  in  power  and  a
0.09-percent  increase  in  heat rate.  This
permits the use of deep beds, and high pres-
sure drop  distribution plates and dust collec-
tion  equipment without  significant  loss in
efficiency.

  Under  contract to  Westinghouse,  Foster
Wheeler  is providing (for NAPCA) proposal
drawings of both a pressurized and. an atmos-
pheric utility  boiler.  Foster Wheeler  will
describe their work in a companion paper.
                 DESIGN COMPARISON

                   Fluidized-bed boiler  designs  and  power
                 systems must be  compared between them-
IV-5-10

-------
 selves and with conventional boiler and power
 systems. The pressurized boiler power system
 concept is  used  to illustrate  the types of
 comparisons which must be considered.
 Comparison of Fluidized-Bed Boiler Designs

   Fluidized-bed boiler concepts are compared
 on the basis of pollution control effectiveness,
 cost, efficiency, and operating characteristics.
 The variable space which must be considered
 in developing a fluidized-bed boiler design was
 discussed  and  summarized  in Table  1.  To
 achieve an optimum design, the alternatives
 available for each variable will require analysis
 and comparison.

   Selection of a fluidized-bed configuration
 for the pressurized  boiler is an example of a
 critical  design task which is being evaluated.
 Several  configurations offer promise: vertical
units with single deep beds, vertical units with
stacked beds (each fulfilling a  separate boiler
function),  open horizontal beds, and  seg-
mented horizontal beds. Each concept can be
considered for shop fabrication or field erec-
tion.  A comparison  of these  configurations
 must be made based on operating character-
 istics (turndown, solids handling, air distribu-
 tion,  erosion,  etc),  heat  transfer  surface
 design, sulfur removal and sorbent regenerator
 design, freeboard requirements,  fluidization,
 combustion  efficiency, and cost.  For
 example,  single deep (on the order of 30 ft)
 beds offer several potential advantages: solids
 handling  problems can  be reduced (feeding
 and removal),  the  number of air distributor
 plates is   reduced,  available space is  used
 effectively, heat transfer surface out  of the
 bed is minimized,  higher  combustion effi-
 ciency may be achieved as the result  of longer
 residence  time*,  and boiler size and cost can
 be  reduced.  Problem areas can also be visu-
 alized:  coal  distribution through  the  bed,
 temperature  profile through the bed,  vibra-
 tion, startup  and turndown capability, and
 tube maintenance. Stacked vertical beds offer
 a method of providing  the  four boiler func-
 tions in  each module  which  may  simplify
startup and  turndown  of the system.  The
horizontal  configuration simplifies structural
* Deep beds may be able to achieve the high combustion
 efficiencies with coarse particles that have been projected
 for fine particles in shallow beds.
                     Table 6. FLUIDIZED-BED VOLUME REQUIREMENT FOR
                               HEAT TRANSFER SURFACE
Tube
arrangement
Pre-evaporator
volume, ft3
Evaporator
volume, ft3
Superheater
volume, ft3
Reheater
volume, ft3
Total
volume, ft3
Triangular configuration
1-in.diam,
2-in. pitch
1,280
1,580
3,340
1,530
7,730
1-in. diam.
3-in. pitch
2,880
3,560
7,530
3,450
17,420
2-in. diam,
4-in. pitch
2,560
3,160
6,700
3,070
15,490
Rotated-square
configuration
2-in. diam.
4-in. pitch
2,950
3,650
7,720
3,540
17,860
                                                                                  IV-5-11

-------
design problems and offers alternate methods
for achieving turndown and  recycling solids.
This configuration does not use the volume as
efficiently as does the vertical unit  and may
also limit  the  flexibility in  heat  transfer
surface design. Fuel distribution may be more
difficult  over the larger area in a horizontal
design. If lower gas velocities are used, carbon
carryover may  be reduced; although deeper
beds with vertical units may provide  the same
result. These comments illustrate the types of
comparative analyses  which  must  be con-
sidered before selecting a configuration.
  The achievement of high combustion effi-
ciency is  another major concern. A vertical
deep-bed  design  may permit  solids  to  be
recycled back  to the bed as  a means for
achieving  high  combustion efficiencies. With
shallow beds and gas velocities greater than
2-4 fps, the CBC  concept conceived by Pope,
Evans and Robbins5  has marked advantages.
With  a deep bed,  the possibility exists for
achieving  high  efficiencies at high velocities
by  recycling material. The  tradeoffs  which
must be considered are solids handling prob-
lems, operating cost, and control. It does not
appear practical to separate the  CBC from the
pressurized  boiler unit  since this would  in-
volve additional transport of solids in and out
of high-pressure vessels. Several  CBC concepts
have been conceived to accomplish this goal.
Since  the  CBC represents a small fraction of
the total bed volume (see Table 4), it is not
practical to allow a single bed in each module
for the CBC. A single CBC could be designed
for one module  and  solids transferred from
the other modules; however, this increases the
solids handling problem  and is not  recom-
mended. Although a single high-pressure field-
erected unit might reduce the problems with a
single CBC, it is not recommended as a solu-
tion. Thus, design  configurations  have been
considered which incorporate  the CBC in a
section  of one  of the beds in  a stacked bed
design; e.g., in a quadrant of the bottom bed.
Solids are then transferred to the CBC from
the upper  beds through an internal cyelone(s).

IV-5-12
  The design of the sulfur  removal system
illustrates another area where critical analysis
is  required.  Processes for high-temperature/
high-pressure regeneration  of the  sulfur
removal  agent  must be  considered. Given a
regeneration  process, the  location  must be
considered. If  the regenerator  is external to
the fluidized bed, solids must be transported
at high temperature and pressure to the re-
generator and the regenerated solids must be
returned  to  the boiler. Incorporating the
regenerator in the high-pressure fluidized-bed
boiler may greatly reduce the solids handling
problem,  minimize heat losses, and reduce
space requirements. An additional  process
scheme  which may offer advantages  is to
combine the carbon burnup cell (CBC) and
the  regenerator  into  a single  unit.  Both
systems operate near 2000° F and the energy
requirements for the   regenerator can  be
supplied by  the solids to  the CBC. (See
Figures  3 and 4.) If the CBC is operated with
30  percent  excess  air,  the  excess  air may
reduce the SC<2 concentration in the flue gas
from the regenerator-CBC to  a level which
cannot be handled by a sulfur recovery plant.
The  carbon  monoxide  requirements for
efficient operation of the regenerator must be
studied  as well as the potential problems of
matching the operation of the CBC, regen-
erator,   and  the  fluidized-bed  combustor/
desulfurizer.

  Solids feeding (tangential,  bottom,  side,
and top), solids removal, and arrangements of
boiler  functions  and  heat transfer surface
configurations (vertical, horizontal, spiral, and
serpentine) are being studied and evaluated. A
range of each  operating condition  must be
considered and the tradeoffs equated: e.g.,
the effect of bed temperature  on corrosion,
deposits,  SO2 absorption,  NOX  emissions,
combustion  efficiency,  heat  transfer surface
requirements, and ash fusion; and the effect
of particle size on kinetics, combustion effi-
ciency, heat transfer, elutriation, and prepara-
tion costs. Higher plant efficiencies may also
be possible since steam temperature may not
be limited by corrosion and deposits.

-------
   These examples are presented to illustrate
 the types of analyses and comparisons which
 must be made  between fluidized-bed boiler
 designs. They are by no means complete.
Comparison with Conventional Systems

   The high-pressure boiler requires  the anal-
ysis of a total power system. Thus,  compari-
son with conventional systems  must  ulti-
mately be based on overall power plant per-
formance.  The high-pressure  FBB power
system is compared with a conventional boiler
and power system on  the basis of  common
heat transfer surface, efficiency, cost, operat-
ing characteristics, and pollution control.

   Heat Transfer Surface. The estimated heat
transfer  surface for the pressurized  boiler is
compared with the surface  requirements  for
the Hammond Unit No. 4  which has essen-
tially identical steam conditions. The surface
requirements  for  the  pressurized boiler  are
approximately 30 percent of those for a con-
ventional  coal-fired boiler. Table 7 presents
the comparison.

  Efficiency. The  projected pressurized FBB
efficiency of 88.6 percent compares favorably
with the 89 percent of the Hammond Unit
No. 4. The pressurized FBB power efficiency
is 38 percent, 1  percent greater than that of
the Hammond unit. Higher efficiencies could
be achieved with the pressurized boiler system
by increasing the gas temperature to the gas
turbine. This could be achieved by modifying
the system  to produce some fuel gas which
can be  combusted at high temperature and
mixed with the low-temperature fluidized-bed
combustor  gas.  If  the  combustor  were
operated at  a  higher temperature, the  sulfur
removal process would not be effective.

  Cost. The projected  plant  cost  (on an
installed basis for plant operation in 1975) for
the pressurized power plant is $158/kW. This
is based on a boiler efficiency of 89 percent, a
fluidized-bed boiler/desulfurizer system cost
of $40/kW of steam turbine power, and a 15-
                 Table 7.  COMPARISON OF HEAT TRANSFER SURFACE AREA





Function
Pre-evaporator

Evaporator
Superheater
Reheater

Heat transferred, Btu/hr x 1 08

FW Hammond
unit No. 4
(3,600,000
Ib steam/hr)
7.8b

9.1
14.5
6.3
37.7
Pressurized
fluid-bed
boiler
(3,500,000
Ib steam/hr)
7.5

8.8
14.1
6.1
36.5
Heat transfer Surface, ft2


FW Hammond unit No. 4

Reported
)
> 25,900
)
81,600°
1 13,000
220,500
Total surface
estimated3
(
81,000 {
\
118,000
109,000
308,000


Pressurized
fluid-bed
boiler
13,900

17,200
36,400
16,700
84,200

Surface
reduction
with pressurized
boiler
%
)
62
)
69
85
73
aConverts surface reported as projected surface to actual tube surface.
^Based on temperature leaving economizer of 593° F.
cSurface includes both projected and actual.
                                                                                 IV-S-13

-------
 percent reduction in indirect costs. A boiler
 system cost of  $40/kW (steam) does not
 appear unreasonable for a modularized design,
 which also  permits indirect cost savings on
 construction and interest during construction.
 A  conventional  600-MW  coal-fired  power
 plant cost in 1970 dollars is estimated to be
 $186/kW, without pollution control equip-
 ment for sulfur removal.

   Turndown. The modular  design of a pres-
 surized boiler power system is  projected to
 have  the capability of  achieving turndown
 ratios of 4:1 or greater. This compares favor-
 ably  with turndown ratios of  conventional
 coal-fired plants.

   Pollution  Control. The  pressurized-boiler
 power system  offers advantages for control-
 ling  sulfur  and  particulate emissions. Al-
 though control of nitrogen oxide emissions
 has not been considered in the present design,
 techniques are  being studied to  lower them.
 The pressurized-boiler was  designed  for 95
 percent sulfur removal. Data is limited which
 permits design calculations to be made at the
 95-percent  level.  Thus, refinement  of the
 designs to achieve high SC>2 removal may be
 required.  Particulate removal from the  pres-
 surized  system  has not  been considered
 beyond the  gas turbines. Particle size distribu-
 tions of emissions from fluidized-bed combus-
 tors indicate that the emission of particles <5
 micrometers may be significantly  less  than
 those from  conventional units. Conventional
 coal-fired power  plants using a similar  coal
 have  SO2, particulate, and NOX emissions of
 approximately  3200  ppm,  0.2  gr/SCF, and
 500 ppm, respectively.

   Table 8 summarizes  a comparison of the
 preliminary   design  of  a  pressurized-boiler
 power system with a conventional coal-fired
 power  system.  Based  on  the  preliminary
 design and analysis of a pressurized fluid-bed
 boiler power plant, such a system has the
 potential  for providing  not only  compact,
 cheap, and efficient power, but also clean air
 and effective use of our natural resources.
IV-5-14
CONCLUSIONS
   Fluidized-bed  boiler concepts  are  being
generated and  evaluated for cheap and effi-
cient power generation with  air  pollution
control. A broad range of fluidized-bed com-
bustor operating and design parameters are
being studied in order to achieve an optimum
design. The concepts are developed  and com-
pared with other fluidized-bed boiler systems
and  with  conventional power systems, based
on:  projections for emission regulations of
SC»2, NOX, and particulates; the power gen-
eration market; boiler and power plant costs;
and  boiler and power  system operation,
maintenance, and reliability.


   A high-pressure fluidized-bed boiler power
system illustrates the development and evalua-
tion of a  concept. Power system  operating
conditions have  been  specified, preliminary
engineering calculations have been performed,
critical  fluidized-bed  combustor parameters
have been identified, and proposal designs and
evaluations  are underway.  Preliminary com-
parisons have  been made with conventional
power  systems. Data  indicates that  sulfur
emissions  can be  reduced 95  percent. The
fluidized-bed boiler also offers advantages in
controlling particulates and may reduce NOX
emissions. The amount of heat  transfer sur-
face  required by a pressurized boiler operating
in conjunction with a  combined gas/steam
turbine  power cycle is estimated to be •V'SO
percent  of that required  by a conventional
boiler for the same  steam  conditions and
flow. The  preliminary cost estimate for the
fluidized-bed  boiler  power system  is
$158/kW. The cost estimate for a conven-
tional  coal-fired  power  plant,  allowing
$10/kW for SO2 removal, is $106/kW. The
projected  efficiency  of the  fluidized-bed
boiler is 88.6 percent, which  compares  favor-
ably with  a conventional plant  of  the same
capacity.  Thus, the high-pressure fluidized-
bed  boiler  power  systems  shows  potential
technical  and  economic  advantages over  a
conventional coal-fired power plant.

-------
  Table 8. COMPARISON OF PRESSURIZED-
        BOILER POWER PLANT WITH
       CONVENTIONAL POWER PLANT

                                Pressurized
                   Conventional fluid-bed boiler
Function           coal-fired plant  power plant
Boiler tube surface, ft2   308,000
Efficiency
  Boiler, %
  Plant, %

Cost, $/kW

Turndown

Pollution control
  SO2, ppm
  Particulate

     Total, gr/SCF
   NOX, ppm
89.0
37.0

186a

4:1


<\,3200


-V0.2

15

>500
84,200


88.6
38.0

158

4:1


160


<0.15b



<250
a600-MW plant without pollution control equipment.
bParticulate removal beyond the gas turbine require-
 ments has not been considered in the initial design.

BIBLIOGRAPHY

1. Archer, D.  H., D. L. Keairns, and W. C.
  Yang.  Marketable  Designs for Fluidized
  Combustion Boilers. Paper presented at the
  Second  International   Conference  on
  Fluidized-bed Combustion, October 1970.

2. Stratton,  J.F. O. Power. 68:  486, Septem-
  ber 1928.
3. Skopp, A., J. T. Sears, and R. R. Bertrand.
   Fluid Bed Studies of the Limestone Based
   Flue Gas Desulfurization Process, Final
   Report,  GR-9-FGS-69,   Prepared  for
   NAPCA. 1969.

4. Hoy, H. R. and J. E. Stantan. Amer. Chem.
   Soc. Div. Fuel Chem. Prepr. 14, 2: 59, May
   1970.
                                              5. Bishop, J.W., E. B. Robison, S. Ehrlich, A.
                                                K. Jain, and P. M. Chen. Papers presented
                                                at the  annual meeting of the ASME,
                                                December 1968.
ACKNOWLEDGMENT

  The  work  discussed  in  this  paper  was
carried  out under  the  sponsorship of the
National Air Pollution  Control Administra-
tion,  Department of Health, Education, and
Welfare. Mr. P. P. Turner has monitored the
work for NAPCA.
   The  authors  wish to thank:  Mr. N. E.
Weeks  for  providing  fluidized-bed  boiler
power system  performance  and  economic
analyses; Mr.  J. R. Hamm for this analysis of
fluidized-bed  boiler power  system concepts;
Foster   Wheeler  Corporation  (Mr.  R. J.
Zoschak, Mr.  R. W. Bryers,  and Mr. J. D.
Shenker) for  providing, under contract to
Westinghouse, information  on the fluidized-
bed boiler design; and the NAPCA contractors
working  on  fluidized-bed  combustion  for
their cooperation and support.
                                                                                 IV-5-15

-------
                                          DUMP
                                                                     DUMP
                                                                                 AIR
                                     POWEh
                                    TURBINE
COMPRESSOR
TURBINE
        I
       TO
   CONDENSER
                                                               BOILER	-*	j
                                        1. DIAGRAM IS FOR 3500-
                                          PSIA SUPERCRITICAL
                                          BOILER. FOR 2400-
                                          PSIA BOILER, #8
                                          HEATER IS REMOVED.
                                        2. FOR NON-INTERCOOLED
                                          GAS TURBINE,
                                        3. STATES (jj) to gg  IN
                                          STEAM CYCLE REPRESENT
                                          LP END LOSS.
                                        4. STATE  @  IN STEAM
                                          CYCLE REPRESENTS
                                          FINAL FW TEMPERATURE
                                          TO BOILER.
                                                       1
                                            n
                                               ___%id
 I	-S-	!
                                                    STACK
 Figure 1.  High-pressure fluid-bed boiler power system (block diagram, combined cycle).
IV-5-16

-------
   NOTE:  TABLE 3 DEFINES CIRCLED NUMBERS.
            FLUE GAS
    TO SULFUR RECOVERY PLANT
AIR.
                      SPENT LIMESTONE
           LIMESTONE
          REGENERATOR
FUEL
                     REGENERATED STONE
                                                                        FLUE GAS TO GAS
                                                                           TURBINE (S)
                     WASTE    MAKEUP   COAL
                     STONE   LIMESTONE
AIR
                 Figure 2.  Process flow diagram for pressurized utility boiler.
                                                                                IV-5-17

-------
                                                                  FLUE GAS TO GAS TURBINE (s)
                                                                 SENSIBLE HEAT: 21.2 x 108 Btu/hr
                                                                 COMBUSTIBLES:  0.41 x 108 Btu/hr
   FLUE GAS FROM FBCD:
     SENSIBLE HEAT 20.1  x 108 Btu/hr
     COMBUSTIBLE GAS~0.4 x 108 Btu/hr
     CARBON         ~2.57 x 108 Btu/hr
   LOSSES FROM FBCD & CBC
   3.3 x 108 Btu/hr
  SPENT CaO/LIMESTONE
                                                                         WASTE SOLIDS
                                                                 COMBUSTIBLE: 0.42 x 108 Btu/hr
                                       COMBUSTIBLE SOLIDS
                                        2.37 * 108 Btu/hr
                                                                        FLUE GAS
                                                                        SENSIBLE HEAT
                                                                        1.1 x 108 Btu/hr
                                                                        CARBON
                                                                        0.23 x 108 Btu/hr
                                             CBC
                                      HEAT OF COMBUSTION
                                         2.1 x 108 Btu/hr
        FBCD
HEAT FROM COMBUSTION
   51.2 x 108-Btu/hr

 HEAT GENERATED FROM
  CaO-S02 REACTION

    1.4 x 108 Btu/hr
HEAT TRANSFERRED
   TO STEAM
 35.3  x 108 Btu/hr
 REGENERATED CaO/LIMESTONE
                                COAL
                            53.8 x 108 Btu/hr
                                                                    HEAT TRANSFERRED TO STEAM
                                                                    1.2 x 108 Btu/hr
                                                                     AIR
                                                                0.3 x 108 Btu/hr
                                                                                    AIR
                                             AIR  6.3 x 108 Btu/hr

                                        BASIS: 77?F
                                                                               ,6.6 x 108 Btu/hr
                    Figure 3. Energy balance for pressurized utility boiler.
IV-5-18

-------
                 LOSSES
             0.26 x 108 Btu/hr
        AIR.
      FUEL
                           T
                     SO  RICH FLUE GAS     f
                                           
-------
SESSION V:




     Conceptual Design and Economic Feasibility (Continued)
SESSION CHAIRMAN:



     Mr. T. C. L. Nicole, National Coal Board, England

-------
                 1.   DEVELOPMENT  OF FLUIDISED-BED
                                     COMBUSTION FOR FIRING
                                      UTILITY STEAM BOILERS


                             D. H. BROADBENT
                        National Coal Board,  England
  The  national Coal  Board's interest  in
fluidised-bed combustion stems from a desire
to maintain its coal sales to its largest outlet,
electricity generation, against growing compe-
tition from other sources of energy.

  Figure 1  is  a projection  of the United
Kingdom energy market up to the end of the
century. It reflects  the  most optimistic fore-
casts  available for the part to be played  by
nuclear energy and  natural gas, yet shows an
increasing use of fossil  fuels  from 1975 on.
Hence the Board's  justification for research
and development of fluidised-bed combustion
as a possible means of reducing the capital
and running costs of fossil-fuel-fired boilers.

  We  have  compared  our energy forecasts
with those for the U.S.1  which show a similar
trend  up to  2000 A.D.; but in the U.S., there
is the added incentive  of pollution control
which makes fluidised-bed combustion even
more attractive.

  Work at both National Coal Board labora-
tories  (B.C.U.R.A. and C.R.E.) has proceeded
for 7 years and costs are currently running at
£1/2 ($1-1/4) million per annum. We have 94
scientists working on the project, backed by
an engineering workshop and  other support-
ing staff.

  There are  11 major research tools:  7 hot
and 4 cold;  10 at atmospheric pressure and 1
 pressurised  to  6  atmospheres. Results  from
 these rigs have  been  used  to  constnict  a
 mathematical model from which the param-
 eters of a variety of boiler designs have been
 established.

  In addition, some £200,000 has been spent
 over the last 2 years in obtaining designs of
 20-,  120-, and  660-MW  boilers  from manu-
 facturers  of international standing. In all
 cases, at least two designs have been obtained
 for each boiler size from different contractors
 to ensure unbiased designs. N.C.B. owns these
 contractors' designs exclusively.

  The programme of development shown in
 Figure 2 covers both pressurised and atmos-
 pheric units and  is planned at a rate which
 will enable building of a  660-MW unit in as
 short a time as  possible without risking exces-
 sive scaleup factors. The programme  covers  a
 19-year period,  7 of which have gone.

  The next major capital investment is on the
Grimethorpe boiler, planned  for one of the
Board's power stations in Yorkshire. It is 20
MW with a 12- x 36-ft bed  and will take 12-18
months to build and commission. The boiler,
designed in  detail, was sent out to tender to
six firms: five offers were  received, each gave
a design to specification and also their own
version.  Results  of bids  were  remarkably
close, and the  Board  is now in position to
place a  firm contract. Eighteen months of
                                      V-l-1

-------
experimentation is planned for the Grime-
thorpe boiler, by which time we would have a
120-MW design  out for  tenders. A 120-MW
boiler takes some  3 years to build and com-
mission and  would then require 2 years  of
testing.  During this 2-year period, a 660-MW
design would be  put  out for tenders, and
would take 5 years for construction and com-
missioning.

   Total cost of this programme is  some £45
million  (at   1969   prices) over the  19-year
period.  Our market surveys indicate  that
there would be considerable  offsetting  of
these costs  before the  end  of the  19-year
period  by virtue   of specific  experimental
work we  will carry out for  other  people,
commercial  exploitation of  boilers  in the
20-MW range possible from 1973, and similar
exploitation of boilers  in the  120-MW range
from 1978.

   If pressurised fluidised-bed  combustion is
proceeded  with alone,   it  would  follow a
similar pattern; but the time scale would also
be influenced by time requirements for suita-
ble industrial gas turbine developments.

  We would aim to run the pressurised and
atmospheric  programmes concurrently in
which  case the information crossflow will
ensure some considerable cost sharing.

  The question now is when does the N.C.B.
stop going it alone in the U.K. and  start to
attract financial support from others. We have
decided to seek partners for the Grimethorpe
stage.
BIBLIOGRAPHY

1. Spaite, P.W. and R.P. Hangebrauck. Sulfur
  Oxide  Pollution:  An Environmental Qual-
  ity  Problem Requiring  Responsible
  Resource  Management. Paper to the  19th
  Canadian  Chemical  Engineering  Con-
  ference. Edmonton, October 1969.
V-l-2

-------
   700
   600
U
§  500
to
"5
8  400
in
c
o
c
o
E  300

t
cr
1
ft  200
   100
1966
COAL
OIL
NUCLEAR
NATURAL GAS
TOTAL
MTCE
174.7
111.7
 10.2
  1.1
297.7
                                                                                      306
                             „	. ENERGY GAP._	
                             TO BE FILLED BY FOSSIL FUEL
                                 (BY DIFFERENCE!:
                                                                                       49
                                                                                     290
                                                    NUCLEAR ENERGY
    1965
                1970
                            1975
                                        1980         1985
                                           YEAR
                                                   1990
                                                               1995
                                                           2000
  Figure 1.  Projections showing the total fossil fuel "energy gap" up to the year 2000.



i

i
O
i
<



6.0


b.O

4.0
3.0
2.0
1.0


! "I I i ' '••
^^_


^—

1 1





- C02N0SMWUCT CONSTRUCT
PROTOTYPE 1 20"MW PROTOTYPE
BOILER COMMISSION BOUJH' COMMISSION
\ AND DEVELOP nc^?no
\ 20-MW / DEVELOP
_ LOAD FACTOR \ PROTOTYPE roOTOTYPE

*"" RASIC RESEARCH AND
1 1 	 1 	 1 	 1 	 1 	 1 	 1 	 1 '
', *
1 1 1 1

^^_
DESIGN AND CONSTRUCT
660-MW POWER STATION



LOAD FACTOR
40% 40%
X
ANALYTICAL SUPPORT FOR PROTOTYPES
	 1 ~\- \~ \
1 1
63 64 65 66 67 68 69 70 71 72 73 74 75 76 77 78 79 80 81
^^

	
	
	

1
82 8
                                      CALENDAR YEAR ENDING
  Figure 2.  Cost and time scale of developing fluidised-bed combustion for central power
  station boilers (present day prices).
                                                                                     V-l-3

-------
                                      2.   MARKETABLE DESIGNS
                                                                            FOR
                     FLUIDIZED - COMBUSTION BOILERS

          D. H. ARCHER, D. L. KEAIRNS, AND W.  C. YANG
                   Westinghouse Research Laboratories
INTRODUCTION

  Westinghouse, under contract to NAPCA, is
providing  market studies, hardware designs,
performance evaluations, and economic data
for  fluidized-bed boilers suited to industrial
and to utility applications. Recommendations
are also being furnished regarding the develop-
ment  program  and  pilot  plant  operation
required  to  realize  practical boiler systems
which are effective  in air pollution control
and economic in operation.
MARKET STUDIES

  The industrial  and utility boiler  market
surveys, which are nearing completion, have
two purposes:

  l.To determine  the  possible  impact  of
     fluidized-bed  combustion  with  sulfur
     absorption in the bed both on air pollu-
     tion abatement and on the economics of
     steam and power generation.
  2. To establish functional specifications for
     the fluidized-bed boiler designs that are
     responsive to customer needs.
Industrial Boiler Installations

  Table 1 shows historic data for the indus-
trial  water-tube boiler market; statistics are
being projected so that boiler sales can be
predicted through  1980. Certain trends are
already clear:
                                       V-2-1
   1. The  total capacity of industrial  boilers
    installed each year is increasing.
   2. The  number of new  coal-fired installa-
    tions are  sharply decreasing and gas-fired
    installations are increasing.
   3. The  average  capacity of new boilers has
    increased from  about 60,000 Ib/hr to
    75,000 Ib/hr in the past 7 years. (Coal-
    fired  boilers  are  30-60  percent larger
    than the average.)

   The operating life of an  industrial boiler is
about 30  years. A conservatively low estimate
of the total  industrial  boiler  capacity  now
extant is  1.2  x  109  Ib/hr; total annual fuel
costs for these might be about $2.5 billion.

   The  possible  pollution  control  and  eco-
nomic benefits  of developing a  successful
fluidized-bed combustion industrial  boiler are
difficult  to assess.  Purchasers of  industrial
boilers currently are avoiding 862 and partic-
ulate (but not NOX)  emission problems by
using natural  gas or low sulfur oil. But the
supplies of such  fuels  are  limited  and their
costs can  be expected to rise.  (These points
will be expanded later.) In certain areas of the
U.S.  it has been announced that  available
natural gas is inadequate for new  industrial
customers. There may well be a growing need
for low-cost packaged industrial boiler in sizes
up to 250,000 Ib/hr (perhaps even larger) that
will burn  coal in such  a way as to  minimize
SO2,  NOX, and  particulate emissions.  Our
preliminary design work indicates  that the
fluidized-bed boiler is indeed more compact
than the conventional coal-burning boiler and

-------
                Table 1. ANNUAL SALES OF INDUSTRIAL WATER-TUBE BOILERS
                	IN THE U.S.	

                                                 1963    1965    1968   1969
Total capacity sold, Ib/hr x 1CT6
Coal-fired
Oil-fired
Gas-fired
Other3
50.5
10.0
12.3
20.2
8.0
76.9
12.1
17.4
34.6
22.8
67.0
4.1
15.6
39.6
7.7
78.0
2.1
13.3
47.8
14.8
                Total number sold
                  Coal
                  Oil
                  Gas
                  Other3

                FOB costs of a packaged industrial boiler

                  Coal  (up to 60,000 Ib/hr)
                       (above 60,000 Ib/hr
                         field erection)

                  Oil  (up to 250,000 Ib/hr)

                  Gas  (up to 250,000 Ib/hr)
   879   1055
   125    104
   263    296
   409    561
    82     94
908
 32
194
617
 65
  $1.25 per Ib/hr

  $3 to $4 per Ib/hr

  $0.80 to $1.00 per Ib/hr

  $0.80 to $1.00 per Ib/hr
                Includes bagasse, black liquor, bark, and waste.
thus can be  packaged in sizes up to 250,000
Ib/hr. Emissions of SO2 can be reduced by a
solid absorbent—such as CaO or dolomite—in
the bed, or it may well be  cheaper to use a
desulfurized  char fuel.  Use  of such a  fuel
would  eliminate  (or  minimize) the use of a
solid absorbent and the problem of disposing
of it or regenerating it. Fluidized-bed combus-
tion may also be effective in reducing NOX
because of its lower operating temperature.
Particulates  may  also  be  more readily  re-
moved from  stack gases; they are, in general,
larger  than  particles  from  pulverized  fuel
combustion;  because they  are  not sintered,
wear of cyclones and mechanical collectors
should be minimized; and their high  carbon
content may improve the operation of elec-
trostatic precipitators.

  Fluidized-bed industrial boilers may well be
effective and economic in air pollution con-
trol if natural gas and low  sulfur oil rise in
V-2-2
cost and if supplies of  a  desulfurized char
become readily available. The costs of operat-
ing an industrial boiler with a sulfur absorbent
in the bed,  regenerating the absorbent, and
recovering (or disposing of) the sulfur will be
assessed  to determine the economic feasibility
of using a high sulfur coal (or oil).
Utility Boiler Installations

  Table 2 predicts the total generating capac-
ity and new installations of generating capac-
ity  in the  U.S.  The new  installations are
broken down into:

  1. Base load plants—with a  load factor
     greater than  80 percent and employing
     fossil fuel and a boiler.
  2. Intermediate  load  plants—with  a  load
     factor around 45 percent and employing
     fossil fuel and a boiler.

-------
                           Table 2. ANNUAL SALES OF UTILITY STEAM
                            GENERATORS IN THE U. S. AND ELECTRIC
                                 UTILITY POWER GENERATION

                                                      1970    1975    1985
                    Total installed capacity, gigawatts

                    New installations, gigawatts/yr
                       Base load—coal, oil, gas
                       Intermediate load—all fossil
                       Peaking load-all fossil

                    Electric utility power generation
                       Coal, %
                       Gas, %
                       Oil, %
                       Nuclear, %
       340    530   1000
        30
      10.1
       6.2
       5.8
  40
12.7
13.3
 2.3
                                                      1960   1968
      66.3
      26.0
       7.6
       0.1
61.9
27.6
 9.4
 1.1
                    Installed costs of a coal-fired utility steam generator

                       Generator and support, per kW      $35
                       Generator and support (including
                         draft feedwater, control, coal
                         and ash handling, dust collection,
                         piping, and other auxiliaries),
                         per kW                       $60
 5.2
23.4
 7.2
   3. Peaking load plants-with a load factor
     of  less than 20  percent  and operating
     with gas turbines.

   The most obvious market for fluidized-bed
combustion boilers are those  now predicted
for the base and intermediate load fossil-fired
plants; this market is  expected to increase by
70 percent-from 16 gigawatts  in 1970 to 28
gigawatts in 1985. If the fluidized-bed boilers
are effective in  air pollution  control and if
their  capital and operating costs are suffi-
ciently  low, they may capture an additional
share of the generation market now ceded to
nuclear-fired base-load plants and to  oil-fired
intermediate-load plants. Higher plant effi-
ciencies  might  be possible with increased
steam   temperature   and   with  pressurized
boilers  with  combined  gas/steam  turbine
cycle.
   A gigawatt of power generating capacity is
roughly equivalent to 10 x 106  Ib/hr of steam
generating capacity and a coal consumption
rate of 400 tons/hr. The present utility boiler
market is therefore larger than the industrial
boiler market by a  factor of  2  in  terms of
capacity and by  a  factor of 7 in  terms of
money.  Utility boilers  also  present a more
serious problem in  air pollution abatement.
Purchasers of the smaller industrial boilers are
turning to natural gas and  low sulfur oil fuels.
Because  of  coal's low  cost  and availability
(despite recent problems of mining it in suffi-
cient  quantities), coal remains a prime fuel for
the utilities.  Most  of the coal available to
power plants in the eastern United States has
quantities  of  sulfur  beyond  air  pollution
limits. A  1000-megawatt  (1-gigawatt) plant
burning 3.0-percent sulfur fuel products 24
tons/hr of SO2- A fluidized-bed utility boiler.

                                       V-2-3

-------
which can economically absorb this pollutant
(permitting the recovery of sulfur) while mini-
mizing  the production  of NOX  and partic-
ulates, can be of great benefit  in reducing air
pollution since  about half of  the SC"2 emis-
sions and a quarter of the NOX emissions are
attributed to electric power generation by the
utilities.
Functional Boiler Specifications

   In  addition  to  examining  the past  and
future  demand  for  industrial  and  utility
boilers, the market survey has also considered
customer requirements for the boilers  they
purchase. Figures  1  through 5  show capacity
and steam conditions for future industrial and
utility boiler markets.

   Figure  1 shows that the largest portion of
the industrial  boiler market  will be supplied
by  units  in  the  capacity  range  of
150,000-250,000 Ib/hr, a size that, not coin-
cidently, is the  largest that can now be  fac-
tory assembled and shipped as  a package. If a
fluidized-bed  boiler  can be more  compact
than  a  packaged  gas- or oil-fired boiler of
conventional design, larger industrial  boilers
may become more popular.

   Figures 2 and 3 show  that the steam con-
ditions of pressure and temperature in general
exhibit  a broad range of preference: about
two-thirds of the boilers have steam pressures
below 600 psi and temperatures below 750°F.

   Figure 4 shows that the largest share of the
fossil-fired utility boiler market will be sup-
plied  by units in the 400-600 megawatt range;
a number of waste-heat recovery boilers in the
100-200 megawatt  range  may also be pur-
chased  to operate  in  conjunction with gas
turbines in combined cycle  power  plants.
Figure 5 shows that super-critical boiler units
operating at around 3600 psi and sub-critical
units operating at around 2400  psi steam pres-
sure will  continue to be popular with  cus-
tomers.
V-2-4
  Table  3  summarizes  the information on
functional  requirements used in establishing
specifications  for the industrial and utility
fluidized-bed boiler designs. The final designa-
tion of specifications is discussed in papers
dealing with the detailed boiler designs.
Fossil Fuels for Boilers

  Table  4 shows the proved recoverable fuel
reserves available for possible use in the U.S.
boilers.  Actual  economically   recoverable
reserves of these fossil fuels may be as much
as 2-3 times the proved values; but still the
message appears clear. Coal is the most plenti-
ful  source of  fossil energy. Domestic oil and
gas  reserves are  in  relatively  short  supply.
Their  prices  can be  expected  to increase
especially if foreign  developments hinder the
free flow of oil to this country. Slightly over
half of the coal energy of the U.S. is in de-
posits  east of the Mississippi  River: most of
this coal has a sulfur content of 2.5-4.0 per-
cent by   weight.  Ultimately  industrial and
utility  boiler  operators will probably  be re-
quired to utilize coal or a coal-derived fuel.

  Table  5 and Figure 6 summarize predicted
costs of  fossil fuels. Major factors—including
production  costs, alternate market opportu-
nities, and government regulation—influencing
the  market   price  of each  are different.
Political  and  environmental constraints,  al-
ready shaping the price, supply, and demand
of each fuel, will continue to affect fuel price
trends  over the next  15 years.
  Table 6 gives transportation cpsts for fuels.
Costs indicate that  coal  and gas cannot  be
used economically in large utility boilers far
from the source of supply.

  Table 7, based on a careful consideration of
all  the  technical, economic, and  political
factors  involved in the choice of fuel supply,
indicates that utilities will use fossil fuels for
generating  over two-thirds of  the electric

-------
           Table 3. FUNCTIONAL SPECIFICATIONS FOR FLUIDIZED-BED COMBUSTION
                  BOILERS DEVELOPED BY THE BOILER MARKET SURVEY
        Characteristic
        Steam conditions
                               Industrial boiler
                 Utility boiler
Capacity
Maximum
Most frequent
Minimum

350-500 x 103 Ib/hr
150-250 x 103 Ib/hr
25 x 103 Ib/hr

1300MW
500-700 MW
100 MW
                                           Size, MW  <300
Pressure, psig
Temperature, ° F
Performance
Efficiency, %
Special requirements
Turndown
Overpressure & flow %
Dynamics, %/minute
Startup, shutdown
Pressurized operation
TV viewing
Packaged construction
Multifuel capability
150-600
(sat-775)

86

1/3
-
—
-
No
No
Preferred
Not usually demanded
1800 2400
950-950 1000-1000

89-92

1/4 (1/2)
5 (on each)
5
Automatic
Yes
Yes
Unfeasible at present.
Desired
3600
1050-1050









but desired

power  in  1980.  A  similar  prediction for
industrial boilers is much more difficult at
this time. In  any case, the development of an
adequate solution  to  the air pollution abate-
ment problem is a very important factor in
determining  the choice  of fuels and in the
economics of steam and power generation.

Table 4. PROVED RECOVERABLE RESERVES OF
           FOSSIL FUELS IN 1968
Fuel
Coal, U.S., tons x
Gas, U.S., MCF x

109
109
Proven
reserves
265
287
Annual
consumption
0.5
19.9
Oil, U.S.,bblsx109
   U.S.                31
   Balance free world    381
   Communist block      59
   Alaskan field (est)     20-30
4.78
7.57
2.08
Alternate Means of SC>2 Pollution
Control in Boilers

   A number of stack-gas processing systems
have been proposed for SO2 removal. These
systems have been considered for possible use
with conventional coal-fired boilers in electric
power plants.  Some  of these systems are
ready for commercial application or nearly so;
some  require  additional  development  and
testing. Table 8 estimates their effectiveness
and cost.


   In general, the capital costs of the systems
are from a third to half the cost of the utility
steam  generator itself. Monsanto's catalytic
oxidation costs  equal  the $35 per kilowatt
cost  of the boiler; TVA's dry lime process  is
about a quarter of this amount.  While most of
the  processes   claim  the  90-percent  SC>2
removal  required  by  today's  air pollution
abatement goals, it is not clear how many of

                                     V-2-5

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               Table 5. PRICE TRENDS OF FOSSIL FUELS FOR POWER GENERATION


Fuel
Coal (mine mouth)
Steam coal
Gas
Residual oil (sea terminal)
Low-S (0.3%)
High-S (2-1/2%)
1968
Price,
d/106 Btu

17
25

32
32

Price Indices (1968=100)
1970

218
108

230
144
1975

205
287

267
141
1980

243
347

291
148
1985

291
386

295
156
  Table 6.  COMPARATIVE TRANSPORTATION
COSTS FOSSIL FUELS FOR ELECTRIC UTILITIES
Fuel
Cost3
Coal (unit train)
       Under 300 miles
       Over 600 miles

Gas (48-in. pipeline)

Residual oil
       Large tankers
       Large river barges
 3.34
 2.92

 1.50
 0.40
 0.60
aCost in ^/million Btu/100 miles.
them can economically accomplish the 95-98
percent  removal which  may ultimately  be
required  as the use of coal in power genera-
tion increases. These stack-gas cleanup proc-
esses  are unlikely to be economic for indus-
trial boilers.  Scaled down  to  a size corre-
sponding with 250,000  Ib/hr  steam boiler,
their  cost  would  be  in  the range  of
$1.50-$3.00 per Ib/hr, greater than the cost of
the boiler itself.

  If  fluidized-bed  combustion  boilers  are
more effective in air pollution control  and
lower in cost than  conventional  boilers plus
stack-gas  cleanup equipment,  they  may
         capture a sizeable portion of  the  predicted
         market.
Pollution Control Targets

   Projections of power generation by fossil
fuels have been used to estimate annual emis-
sions of SC>2,NOX and particulates to the year
2000.  These estimates are based on the fol-
lowing assumptions:

   1. Sulfur content of coal increases from the
     current  2.7 percent to 3.5 percent in the
     year 2000.
   2. Coal and oil generation are combined.
   3. Nitrogen oxides are produced at a rate of
    800 Ib/billion Btu for coal and oil.
   4. Heat rates are 10,300 Btu/kWh for coal
    and oil and  12,500 Btu/lb for coal.
   5. Flue-gas/coal ratio by weight is 12.5.

   Tables 9,  10, and  11  show the degree of
pollutant  reduction  or  removal  required to
achieve acceptable air standards in  1970 and
to maintain the same total emissions (despite
increasing quantities  of fossil fuels used in
power  generation)  in  1985  and 2000. Al-
though 90-98 percent  reductions of SC«2,
NOX, and  particulates  may  be  beyond the
state of present  art, such reductions are the
targets of fluidized-bed boilers. As such, they
provide a basis for design and a direction for
future development.
V-2-6

-------
Table 7. RELATIVE SHARE OF ELECTRIC UTILITY STEAM GENERATION
Fuel
Fossil
Coal
Gas
Oil
Nuclear
Total
1969
kWh x 109
1283.0
793.1
354.6
135.3
19
1302
%
98.5
60.9
27.2
10.4
1.5
100
1970
kWh x 109
1366.0
858.8
371.5
135.7
63
1429
%
95.6
60.1
26.0
9.5
4.4
100
1975
kWh x 109
1741.0
1101.7
461.7
177.6
461
2202
%
79.1
50.0
21.0
8.1
20.9
200
1980
kWh x 109
2028.0
1252.5
536.0
239.5
933
2961
%
68.5
42.3
18.1
8.1
31.5
100
1985
kWh x 109
2416
1396.1
625.4
394.5
1648
4064
%
59.5
34.4
15.4
9.7
40.5
100

-------
                Table 8. ASSESSMENT OF FLUE GAS DESULFURIZATION PROCESSES

Process
Plant
size,
MW
Process cost
Capital
$/kW
Operating
rf/106 Btu
Years to
full-scale
plant

%SO2
removal
Chance
of
success
    Ready for commercial application
      C.E. dry lime-wet scrub           500      11      4.2        3 sold       85     Good
      Monsanto cat-ox                 500      36      0.2a       ready        90     Good

    Nearly ready for application
      NAPCA-TVA dry lime            500      8      4.3        ready       40     Fair
      Wellman-Lord                   500      20b    3.9a-b      1 year       90     Fair
In development
S & W-lonics electrolysis
Esso-B&W dry adsorb
A.I. molten carbonate
Consol Coal potassium formate

1200
800
800
1300

19b
17"
15b
17b

0.6a-b
0 a-b
4 a,b
5 a.b

5 years
6 years
7 years
7 years

90


90

Fair
Unpredict
Unpredict
Unpredict
    aAssumes byproduct credit but no credit for lower stack, eliminated precipitator, etc.
    '•'Optimistically low cost, estimated by process developers.
   Table 9. PROJECTED S02 EMISSION LEVELS
      FROM POWER GENERATION PLANTS
Table 10. PROJECTED NITROGEN OXIDES EMIS-
  SIONS FROM POWER GENERATION PLANTS
Year
1970


1985


2000


Control
level3
I
II
III
I
II
III
I
II
III
Average S02
leaving stack, ppm
2000
2000
220
2220
1230
140
2600
500
55
S02 reduc-
tion, %
0
0
89
0
45
94
0
81
98
aControl levels are: I—uncontrolled; II—equivalent to
 maintaining total SO2 emissions from power plants
 at the 1970 level; and  Ill-equivalent to S02 emis-
 sions  if all the 1970 generating capacity was based
 on using M>.3 percent sulfur fuel.
Year
1970


1985


2000


Control
level3
1
II
III
1
II
III
1
II
III
Average NOX
leaving stack, ppm
520
520
260
520
320
160
520
130
65
NOX reduc-
tion, %
0
0
50
0
39
69
0
75
88
aControl levels are: I-uncontrolled; II—equivalent to
 maintaining total emissions to the atmosphere from
 power plants at the 1970 level; and III—equivalent
 to reducing the total NOX emissions to 50 percent of
 the 1970 level.
V-2-8

-------
BOILER DESIGN

   The results of the boiler market studies and
the  air   pollution  abatement  considerations
have  been  used  in establishing  functional
specifications for one industrial and two util-
ity boilers. A number of boiler concepts have
been  reviewed and evaluated to  determine
their  compatibility  with these specifications
and the general and economic requirements of
the market.
Functional Specifications for a
Fluidized-Bed Industrial Boiler

   Results of these considerations, with regard
to the  utility boiler, are presented in a com-
panion paper. This paper deals with the indus-
trial boiler. Table 12 gives its functional speci-
fications.  A  250,000-lb/hr capacity  was
chosen because of the market for this size and
Table 11. PROJECTED PARTICIPATE EMISSIONS
     FROM POWER GENERATION PLANTS
Year
1970


1985


2000


Control
level3
I
II
III
I
II
III
I
II
III
Emission,
gr/SCF
3.7
0.5
0.037
3.7
0.30
0.023
3.7
0.12
0.009
Reduction, %
0
87
99
0
92
99.4
0
97
99.8
aControl levels are: (-uncontrolled; 11-equivalent to
 maintaining total emissions rfom  power plants at
 1970 control level of 'V 87-percent removal;b and
 Ill-equivalent to maintaining total emissions from
 power plants at a level corresponding to 99-percent
 removal in 1970.

bNAPCA.  Control Techniques  for  Particulate Air
 Pollutants.  Publication  No.  AP-51.  Washington,
 D. C., January 1969.
because  it may  be the largest  fluidized-bed
boiler shippable  as a  single module. A steam
pressure of 600  psi with  superheat from the
saturation temperature of  489°F  to 750°F
has  been  selected; these values  have  been
chosen because  two-thirds  of the industrial
boilers operate at  or below these conditions.
The fuel for this boiler is a crushed Pittsburgh
coal, undried, with 4.3 percent sulfur. (Pulver-
ized coal has been ruled  out because of the
expense and  operating problems involved with
coal mills.)

   Reduction of SC>2 emissions by 90 percent
is to be accomplished  by  lime (CaO, CaCC>3,
or dolomite) additions to the bed. Alternate

Table 12. FUNCTIONAL SPECIFICATIONS FOR A
     FLUIDIZED-BED INDUSTRIAL BOILER
Characteristic
Specification
Steam capacity, Ib/hr              250,000

Steam conditions
   Pressure, psig                   600
   Saturation temperature, °F       489
   Superheated to, °F              750
   Water to economizer at,u F       250
   Water leaving economizer at, °F    350

Fuel
   Coal                         Ohio Pittsburgh
                                No. 8 seam
   Crushed to                     -1/4 in. xO

Air pollution control targets
   S02 reduction by limestone
     absorption, %                90
   NOX
   Partciulates reduction, %         95a

Boiler efficiency, %                 >85

Packaged construction
   Factory fabrication              Yes
   Shippable                      Yes
   Dimensional limits of
     primary modules, ft           12 x 16 x 40

aAII ash and 3 percent of lime assumed to be elutri-
 ated.

                                        V-2-9

-------
absorbents are not considered because their
cost is in general higher and their effectiveness
is not yet  demonstrated.  A  regeneration
system is  specified  for the lime absorbent
because large  quantities of this  material will
probably be required to achieve a 90-percent
reduction in SC>2 emissions. It does not seem
economically  or technically feasible to dis-
pose  of large amounts (up to half  the coal
tonnage)  of spent absorbent.  Regeneration
with  sulfur recovery does not appear easy or
cheap but it remains to be evaluated. A more
economic solution for industrial applications
may be to use a desulfurized coal char in a
fluidized-bed boiler. Such a boiler is probably
the only economic means for using char while
minimizing NOX and particulate emissions.

   Reduction of the particulate ash and lime-
stone  emissions in  the stack gases  is to be
achieved by a combination of cyclones and, if
required, an electrostatic precipitator.

   Finally,  an overall boiler efficiency greater
than  85 percent has been specified in order to
obtain reasonable fuel costs.
Fluidized-Bed Industrial Boiler,
Operating Conditions

   Figure 7 shows that the fluidized-bed boiler
comprises three main sections—the combustor
and  SC>2  absorber,  the carbon burnup cell
(CBC),  and the lime regenerator. The  func-
tions of the  last two sections might be com-
bined if the amount of carbon elutriated from
the combustor can be minimized. When large
amounts of  carbon carryover must be con-
sumed,  combining  the CBC and the  regen-
erator produces a gas mixture dilute in SO2-
This bed can  be operated intermittently, vary-
ing temperature and flue gas composition, to
produce periodic  surges of SC»2 from the
combined  CBC/regenerator. But  in  either
case,  capture or recovery  of  the  sulfur  is
complicated.

   About 85  percent (11 ton/hr) of the total
coal required  is fed into the primary  fluidized-
V-2-10
bed  combustor (FBC) with coal combustion
efficiency of 87  percent. The remaining 13
percent unburned coal is elutriated, along with
the ash and 3 percent of the spent limestone.
Ninety-five  percent  of the elutriated solids
from the FBC is recycled through the primary
cyclone back to the CBC for further combus-
tion. In addition to the recycled coal, fresh
coal (about 15  percent of total coal required)
is also fed to the CBC. Consequently, about
75 percent of the total coal is burned in the
FBC; the other 25 percent is burned  in the
CBC. The  ash  and spent limestone recycled
back to the CBC are again elutriated along
with additional  ash  and   spent  limestone
produced in the CBC and eventually collected
in a secondary  cyclone and electrostatic pre-
cipitator.

   The spent limestone from the combustor is
regenerated in the limestone regenerator. The
total SC«2  removed  in the  regenerator  is
assumed to be 90 percent of that picked up in
the  combustor. The average composition of
the recycled limestone is assumed to have 5
mole percent sulfate. To maintain reactivity,
10 percent  of  the recycled limestone is re-
jected and replaced by fresh calcined lime.

   Table  13 details operating  conditions for
the boiler.  Bed  temperatures chosen are based
on the experience of all those  in the NAPCA
program: they are high enough to obtain good
combustion and heat transfer, yet low enough
to obtain good performance  from the lime
absorbent.  High  gas velocities (15  ft/sec) in
the  beds have been specified to achieve high
boiler  compaction. A large top size of the
crushed coal (1/4 in.) has also been specified
to minimize elutriation and to ease coal distri-
bution  problems:  fine  particles(  tend  to
devolatilize  and burn closer to the point of
their introduction. Coarser particles  are ex-
pected to  burn more uniformly  throughout
the bed.

   Minimizing  elutriation has  also  been  an
important consideration in choosing a particle
size  distribution  for  the lime absorbent. A

-------
          Table 13. OPERATING CONDITIONS FOR A FLUIDIZED-BED INDUSTRIAL BOILER
Condition
Temperature, "F
Coal combustion efficiency, %
Air excess, %
Superficial fluidizing velocity, ft/sec
Fluidized-bed
combustor
1650
87
10
15
Carbon
burnup cell
1900
90
30
15
Limestone
regenerator
2000 (bed)
-
-
_
Boiler
—
—
-

         SC>2 removal, %                       _

         Regenerating gas                      _

         Particulate removal efficiency, %         -

         Stack gas temperature, °F               -

         Air preheat temperature, °F             —

         Boiler efficiency, %                    -

         Heat losses
            Dry flue gas, %
            Evaporation of water formed by burning hydrogen, %
            Evaporation of water in fuel, %
            Heating water in air, %
            CO and unburned C, %
            Radiation, %
                     90

                   C02/C0 = 2

                               95

                              350

                      —        None

                               86.2
                                5.65
                                4.37
                                0.30
                                0.14
                                3.00
                                0.36
         Limestone feed
            Stone—BCR 1359 as the primary bed material with characteristics
            Particle size-1000-5000 microns with average diameter ^2500 microns
            Feed requirements—^90% of 802removal w'tn 6 times stoichiometric feed ratio
            SOo/CaO reaction heat generation-3 x 10* Btu/ton CaS04 produced and the stone enters
                 boiler at CaO at 1900° F after regeneration
            Recycle rate-3% elutriated with the fly ash and 10% rejection in the regeneration
lime  flow rate has been specified to yield a
molar Ca/S ratio of 6. Data from NAPCA
contractors indicates that this ratio should be
adequate to obtain a 90 percent reduction in
S02.

   Excess air amounting to 10 percent is fed
to the combustor; and 30 percent to the CBC.
The regenerator is operated with an air defi-
ciency so that the molar CO2/CO hi the SC>2
containing flue gases is 2.

   The temperature of the stack gases leaving
the boiler has  been specified as 350°F; water
temperatures entering and  leaving the econ-
omizer, as 250 and 350°F. These temperature
designations are based on conventional boiler
practice.

                                      V-2-11

-------
 Preliminary Design Computations
 and Considerations

   Material  and heat balances, together  with
 heat transfer computations, have been carried
 out  for  the  fluidized-bed  industrial  boiler;
 some of the results of this work are presented
 in Table  14 and Figures 7 through 1 1. Figure
 7 shows the large quantities of lime absorbent
 and particulate solids which must be handled
 in the system; it seems important to closely
 integrate  both the lime regenerator  and  CBC
 with the  fluidized-bed  combustor and sulfur
 absorber. This integration can minimize the
distances  over  which  solids  must be  trans-
ported.

   Figures 8 through 11 are heat balances, in
the form  of temperature-enthalpy diagrams.
for various fluidized-bed  boiler functional
arrangements. These diagrams show tempera-
tures  throughout the  boiler and the  corre-
sponding heat quantities transferred from the
burning solids  or  combustion  gases to  the
water and/or steam  expressed in terms of Btu
per Ib of  fuel burned. About 20 percent of
the heat released increases the water tempera-
ture from 250°F  to  the saturation  value,
        Table 14.  HEAT TRANSFER COMPUTATIONS FOR FLUIDIZED-BED BOILER DESIGNS
Characteristic
Heat load, Btu/hr x 106
Submerged in bed
Above bed
Superheater
Convection pass
Total
Log mean temperature difference, °F
Submerged in bed
Above bed
Superheater
Convection pass
Heat transfer coefficients, Btu/ft2/hr/°F
Submerged in bed
Above bed
Superheater
Convention pass
Heat transfer surface, ft2
Submerged in bed
Above bed
Superheater
Convection pass
Total
Shop-assembled
conventional
boiler

78.0

26.1
143.4
247.5

1910

1410
796

29.2

30.4
20.9

1400

609
8630
10639
FBC boiler
Vertical tubes in bed
extending into
freeboard

136.8
63.4
43.6
20.8
264.6

1161
742
1280
291

50
30
30
20.4

2350
2840
1140
3510
9840
Horizontal tubes
in bed

136.8
10.8
43.6
73.3
264.5

1161
1075
1280
495

50
30
30
30

2350
335
1135
4950
8770
V-2-12

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489°F;  50 percent vaporizes the water: and
15 percent superheats the steam to 750°F.

  Figure  8 corresponds to  a  fluidized-bed
boiler with vertical evaporator  tubes extend-
ing  through the  bed and  freeboard.  (These
tubes might be distributed more or less uni-
formly or clustered in banks or platens.) The
tube surface above  the bed  cools  the  gases
from the  1650°F bed temperature to  925°F.
A convection  pass further  cools the gases to
672°F;  and  an  economizer  recovers  the
balance of the heat from the stack gases. The
superheater  is located in the CBC: sufficient
coal is  fired  with  the  char to provide the
necessary  heat. Dotted lines show a possible
alternate  of bringing water to saturation in
the bed and using the convection  pass as an
evaporator.

  Figure 9 corresponds to  a boiler with hori-
zontal  tubes in the fluidized  beds.  In  this
particular situation,  the CBC does  not super-
heat all of the steam; a portion (24 percent)
of the heat released by combustion gases as
they are cooled from 1650 to 1530°F is trans-
ferred  to the steam in a  convection super-
heater.

  Figures 10 and 11 show two of many alter-
nate arrangements  of  different  functional
units of an industrial boiler. In Figure 10 an
air  preheater is used which, in this example,
transfers the same amount  of energy that ,the
economizer does and preheats the air from 80
to 430° F  For each Btu of heat transferred to
the combustion air  in a preheater, an  addi-
tional Btu must be transferred to the steam in
the  fluidized-bed  combustor if the  bed tem-
perature  is  to remain constant. The econ-
omizer is  put  into the fluidized bed, and the
convection pass surface is  packed  above the
fluidized bed to take advantage of higher heat
transfer coefficients. The economizer is elimi-
nated in Figure 11 and the heat transfer sur-
face above the bed  preheats the water.  The
superheater  is put in the  CBC with a small
portion in  the fluidized  bed  of the main
carbon combustor; in another alternative, the
superheater is in the fluidized-bed combustor
and the CBC serves as an evaporator. Numer-
ous other arrangements can also be proposed:
the optimum design will depend not only on
the economic consideration but also on steam
quality, velocity in the tubes, and  the turn-
down  characteristics of  the boiler plant.
Consideration of all the possible alternatives is
a way to an optimum design.

   Preliminary  heat  transfer computations
have been carried out on the vertical and hori-
zontal tube designs  whose heat  balances are
represented  in  Figures  8 and 9. Table 14
compares the  results of these computations
with similar results for a conventional oil-fired
packaged boiler.  The fluidized-bed  boilers
transfer 7 percent  more heat and have 10 per-
cent less transfer surface. Heat transfer coeffi-
cients  above the bed and in the  superheater
have  been  chosen  conservatively  low  (30
Btu/ft2/hr/°F): they may  well be40Btu/ft2/
hr/°F  or higher. In  this  instance, a fluidized
coal-fired boiler will have  30 percent less heat
transfer surface than a comparable convection
gas- or oil-fired  boiler.

   For  NAPCA, and  under subcontract  to
Westinghouse, Erie City is providing proposal
designs of an  industrial  boiler  whose func-
tional  specifications,  operating  conditions.
and material and heat balances have been dis-
cussed. Erie City will describe its work in a
companion paper.
PROBLEMS IN BOILER DESIGN
AND DEVELOPMENT

   Proposal designs of one industrial and two
fluidized-bed boilers are now being prepared.
The primary purpose of these designs is to
evaluate  the  effectiveness and  economics of
such boilers  in generating  steam and  power
without  air pollution. Their secondary pur-
pose is  to reveal areas in which additional
knowledge or more development is needed to
produce a commercial fluidized-bed boiler.
                                    V-2-13

-------
   Although  designs are still far  from com-
 plete,  some problem areas have already been
 encountered.  A  companion paper  discusses
 those  involved  with  a  pressurized  utility
 boiler; this paper is concerned with  the prob-
 lems  involved  with  the design  of an  atmos-
 pheric industrial boiler. These  problems  fall
 into two areas—those involved primarily with
 operation and those with design.

 Operational Problems

   The most important choice in determining
 the  mode of operation  of a  fluidized-bed
 boiler  is that of gas velocity (and of the corre-
 sponding bed particle size). A high gas veloc-
 ity, together with a coarse coal and limestone,
 has been  chosen for the  industrial  boiler to
 minimize  the  cross-sectional bed area. Hope-
 fully  the  boiler will then be most compact.
 Although coarse coal may also minimize coal
 distribution problems, coarse particles result
 in lower heat  transfer  coefficients. High air
 velocities  result  in:  greater elutriation  of
 carbon from the bed; higher attrition in  the
 limestone absorbent; and, perhaps, erosion of
 the heat transfer surfaces. The problem is to
 determine  an  optimum gas velocity-particle
 size  choice; present performance and  eco-
 nomic data do not appear adequate to make
 this choice.

   A second  operational problem is that  of
 boiler  turndown; several procedures are possi-
 ble:

   1  "Slumping" the bed; i.e., merely decreas-
     ing the air flow and depending on bed
     cooling, bed contraction, or defluidiza-
     tion  to reduce  heat transfer and thus
     steam generation.
   2. Decreasing air flow and simultaneously
     draining a portion  of the solids from the
     bed to reduce heat transfer area.
   3. Shutting off the air  flow to whole sec-
     tions of the boiler.

   Selecting the procedure  that will provide
 the  greatest  flexibility  in turndown,  yet

V-2-14
 minimize operating problems and boiler costs,
 is a problem.


 Design Problems

  The  most  important  design  problem  is
 designating the heat transfer surface arrange-
 ment.  Vertical tubes permit natural circula-
 tion, but horizontal tubes with forced circu-
 lation  may be more compact.  The choice  of
 tube orientation affects  the air distributor,
 coal feeding,  and tube manifolding systems.
 Also the  choice  of tube  diameter  and pitch
 has considerable effect on  the design and
 operation of the  boiler. Small tube diameters
 and  close  spacing  produces  a  compact
 boiler; however, they may also  result in poor
 fluidization  and uneven  fuel  distribution.
 Tube erosion, clinkering, and excessive carry-
 over of carbon may also occur.

  A second design problem is the choice and
 location  of  the  various boiler  functions—air
 preheater, economizer, evaporator, and super-
 heater—in the boiler. It is not  clear than an
 economizer  alone is  the  most effective and
 economic heat trap; it may be better to com-
 bine an  economizer and an air  preheater.
 Other  questions involve the  location  of
 evaporators and superheaters in the beds and
in the convection  passes: Which should be
 placed  where? And should the superheater be
 distributed throughout the boiler to facilitate
turndown by sections?

  A third design problem concerns using hori-
 zontal  baffle  tubes  at the surface of the
 fluidized bed  to  minimize the  splashing and
spouting of particles. Baffle  tubes  are effec-
tive in minimizing the carryover of particles.
And heat transfer coefficients for these tubes
are  apparently almost  as high as for sub-
merged tubes. But the  reduction of particles
above the bed may reduce heat transfer rates
in the  freeboard  above the bed surface. Too
 severe  cooling of the bed gases may  inhibit
 combustion processes in  the freeboard and
result  in increased  carbon  monoxide and
 hydrocarbon losses.  Data is  lacking  for the

-------
design of an optimum baffle tube and free-
board section of a fluidized-bed boiler.
Other Problems

  Two other general problems encountered in
the design of a fluidized-bed  industrial boiler
have already been mentioned.

  1. Designation of an optimum sulfur clean-
     up system. Lime has been designated as
     the absorbent, but should it be a once-
     through or a regenerative process? How
     should  byproduct  sulfur be recovered?
     Might it be more realistic in the long run
     to use a desulfurized char fuel?
  2. Designation  of an optimum design and
     operating conditions for NOX and partic-
     ulate emissions control. Data on  simul-
     taneous  control  of  SC>2,  NOX,  and
     particulates  is inadequate  to choose an
     optimum  method.  Might  a  two-stage
     combustion process be used to minimize
     NOX formation?  Can an  electrostatic
     precipitator  be  used  effectively  to
     remove fly ash from fluidized beds since
     carbon may  provide some electrical con-
     ductivity of the carryover?
CONCLUSIONS

  Market studies  for  industrial and utility
boilers,  together with  studies of fuel availa-
bility and  of stack gas cleaning  processes,
indicate that a coal-burning, sulfur-absorbing
fluidized-bed  boiler  can  have  a significant
effect on both air pollution abatement and on
the economic generation of steam and power.
Functional  specifications, including target
values for SC>2, NOX, and particulate reduc-
tions, have  been drawn up for an industrial
boiler. Operating conditions have  been sug-
gested, preliminary engineering computations
have  been performed, and proposal designs
and evaluations are underway. Some  problem
areas require further data and development to
confirm the practicality and effectiveness  of
fluidized-bed boilers.
ACKNOWLEDGEMENT

  The  work  discussed  in  this  paper was
carried  out under the  sponsorship of  the
National Air Pollution Control Administra-
tion,  Department of Health, Education, and
Welfare. Mr. P.  P. Turner monitored the work
for NAPCA.

   The authors  wish  to  thank for their
cooperation and support: Messrs. H. L. Smith
and S. J. Jack for their work on the analysis
and projections of the utility boiler market;
Erie City Energy Division of Zurn Industries
(Messrs. R. V. Seibel and W. D. Schwinden)
for their  work, under  contract to Westing-
house, on  industrial  boiler market analyses
and industrial fluidized-bed boiler design; and
the NAPCA contractors working on fluidized-
bed combustion.
                                                                                  V-2-15

-------
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                                                                                                BOILER OPERATING TEMPERATURE RANGE. °F
                                                                                                 Figure 3.  Projected distribution of
                                                                                                 U.S. industrial boiler sales  — oper-
                                                                                                 ating  temperature.

-------
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       100                            BOILER CAPACITY, megawatts
        Figure 4.  Projected U.S. utility boiler sales -- steam generating capacity.
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 1970   72   74    76    78    80    82   84
            YEAR OF INSTALLATION
Figure 5.  Projected distribution of U.S.
utility boiler installations-steam pres-
sure.
                                                            I  I   I  I  I   I  I  I   I  I  I  I  I   I
                                                             No. 2 Diesel Oil
                                                          — 0.3°o S Residual Oil-
                                                             0.3% S Crude Oil-
                                                          — Natural Gas
                                                                      iV|% S Residual Oil   —
                                                                        % S Crud* Oil
                                                                        ligh S Crude Oil   —
                                                                        •ligh-S Residual  Oil
                                                                      ^Mine-Mouth Coal    —

                                                          I   I  I  I   I  I  I   I  I  I  I  I  I  I
                                                        1970
                                                                                          1985
             1975        1980
                   YEAR
Figure 6.  Projected fossil fuel prices
for U.S.  electric industry -- 1971-1985.
                                                                                        V-2-1

-------
                                                PRIMARY
                                                CYCLONE
                                   FLUE GAS: 11824 Ib-mole/hr
                                PARTICULATE: 1.71 ton/hr
                                     /  ASH: 62.0 w/o
                                     < COAL: 22.8 w/o
                                     I  L.S.: 15.2 w/o
   FLUE GAS:
PARTI CULATE:
      e  ASH:
     < COAL:
      I  L.S.:
               8810
    Ib-mole
      hr
2.44 ton/hr
33.4
58.5
8.1  w/o
                                                            SOLID:
            SOLIDS:
             (COAL:
                     LIMESTONE:
                     1.63 ton/hr
                     CaS04: 11.4
       FLUIDIZED-
           BED
       COMBUSTION
          WITH
         SULFUR
        REMOVAL
                SPENT LIMESTONE:
                  6.32 ton/hr
                CaS04:  33.7 w/o
                 I
    AIR:
                r
    COAL:  11 ton/hr
 LIMESTONE:
 5.60 ton/hr
 CaSO4:  11.4 w/o
  w/o
      58.5 w/o)
                                                CARBON
                                                BURNUP
                                                 CELL
                                      COAL:  1.85^0.
     AIR:  2900
Ib-mole
  hr
                                        TO SECONDARY CYCLONE
                                             & ELECTROSTATIC
                                             PRECIPITATOR
                                         FLUE GAS: 3014
                                                          hr
                                                                     PARTI CULATE:  1.59 ton/hr
                                                                          /  ASH:  64.0 w/o
                                                                          < COAL:  20.2 w/o
                                                                          <  L.S.:  15.8 w/o
                       SPENT LIMESTONE:
                         1.85 ton/hr
                       CaS04:  34.4 w/o
                             FLUE GAS: 361.4
                                (S02: 7.42mole
                              SOLIDS: 0.14 ton/hr
              SPENT LIMESTONE:
                 8.17 ton/hr
              CaS04: 33.8 w/o
                                  LIMESTONE
                                 REGENERATOR
  7.23 ton/hr
                                COKE: 0.48 ton/hr
                                         Ib-mole
                                                                              AIR: 321
                                                                                         hr
                                          6.96 ton/hr
                                        CaS04: 13.2 w/o
CaS04: ll.4w/o^

             0.97 ton/hr    0.70 ton/hr
                        CaS04:  13.2w/p
                                             MAKEUP
                                            LIMESTONE
                          DISCARDED
                              Figure 7.  Overall material balance.
V-2-18

-------
  3000
o
LU
CC
  2000 —
  1000
                                      3560/
                                  Btu/lb FUEL
       SENSIBLE HEAT FROM LIMESTONE   0.65'
       SENSIBLE HEAT LOSS FROM ASH    31.2
       CaO/SO2 REACTION            246
                                       I   I    I   I   I   I   I   I   I   i356?/
                                                                Btu/lb FUEL
                                     SENSIBLE HEAT FROM LIMESTONE    0.65
                                    hSENSIBLE HEAT LOSS FROM ASH     31.2
                                     CaO/SO, REACTION             246
                                SUPERHEATER
                                    (SH)
       SURFACE ABOVE
            BED
       CONV.
       PASS
       ECON-
                                                      ONVECTION PASS
                                                                              SUPERHEATER
                                                                       SURFACE    (SH)j
                                                                      ABOVE BED
                       350  EVAPORATOR

                         I
                                                                         EVAPORATOR
                                                                          I   I   I
             23456789  10 11
           ENTHALPY 1000 Btu/lb FUEL BURNED
        Figure 8.  Temperature-enthalpy
        diagram for vertical bed tube design.
                                                        1   23  4  56  7  8  9  10 11
                                                         ENTHALPY 1000 Btu/lb FUEL BURNED
                                                    Figure 9.  Temperature-enthalpy
                                                    diagram for horizontal bed tube design.
  3000
uJ 2000
cc.
LU

I
   1000
                                         1900
          ABOVE BED
        PRE-
       H EATER

         672/
         Vfe
             FBC
       i
    7501
       .ECONOMIZER
SH  489|       489
                         . -ECONOMIZER
                      f°f,5° EVAPORATOR
     ~6~ 1   2  3  4  5  6  7  8  9  10  11
           ENTHALPY 1000 Btu/lb FUEL BURNED
      Figure 10. Arrangement of functional
      units of a fluidized-bed boiler — al-
      ternative No. 1.
                                                3000 —
                                                 200( —
                              iS
                              i
                              LLI
                              Q.
                              I
                                                 1000
                                            12
                                                                      SUPERHEATER
                                                                      (EVAPORATOR)
                                       123456  789  10  11
                                         ENTHALPY 1000 Btu/lb FUEL BURNED
                                      Figure 11. Arrangement of functional
                                      units of a fluidized-bed boiler — al-
                                      ternative No. 2.
                                                                                    V-2-19

-------
                                3.   DESIGN OF ATMOSPHERIC
                              FLUIDIZED-BED  COMBUSTION
                                               STEAM GENERATOR
                               R. V. SEIBEL

         Erie City Energy Division, Zurn Industries, Inc.
INTRODUCTION

  If steam  were an end product, studies of
capital expenditures,  operating costs,  and
maintenance costs for various steam-generator
configurations would have precise meanings
to all. Since steam is not an end product and a
steam generator is never the only process in a
plant, it follows that  the worth of a steam
generator to any particular user is a complex
subject. In one case, it does not have the same
meaning to all users.

  Basically, a steam generator converts chem-
ical energy (available from oxidizable organic
materials) into heat energy (in a form that is
transmittable to a remote process).  Heat, in a
convenient form, is the product of a steam
generator. Producing heat  for a process by
converting chemical energy is the work  of a
steam generator. How  much of the available
chemical energy is converted? What procedure
is best for making this conversion? How much
does it cost to make  this conversion? How
much reliability  is demanded in making the
conversion?  All  these questions encompass
the  study of steam-generator design and appli-
cation.
CONVENTIONAL STEAM GENERATORS
  The four questions above (with the second
and third combined for convenience) can be
used as the basic thoughts in the study of
fluidized-bed  combustion techniques  as
applied  to  industrial applications in other
terms in atmospheric design. This paper dis-
cusses basic facts and experience with conven-
tional steam generators, and then summarizes
initial work and thoughts on a fluidized-bed
combustion (FBC) steam generator.
Quantity of Chemical Energy Converted

  For all fuels, a terminal temperature of the
combustion products (selected by experience)
is the determining factor of boiler efficiency,
defined as heat output divided by heat input.
For  natural gas, oil, and coal, the final  or
"stack" temperature for optimum boiler effi-
ciency ranges  from 300 to 350°F. Sensible
and latent heat losses vary, depending on the
fuel; at  the same input and  under identical
conclusions: natural gas has  the lowest effi-
ciency; oil, the next; and coal, the highest.

  Table 1 summarizes heat  losses for three
different fuels where the final combustion gas
temperature is  300° F, and the sensible heat
base is 80° F. Practical cases are shown where
conventional design allows for varying excess
combustion air and unburned combustible
losses. No unburned combustible is generally
found with the oxidization of natural gas and
residual fuel oil. Practical designs of combus-
tion  methods of coal do not allow complete
                                      V-3-1

-------
 combustion of coal.  The 3-percent combus-
 tible loss for coal in Table 1 represents spread-
 er stoker firing for a bituminous coal.

   Table  1  shows  that  boiler efficiency is
 greater for configurations  of  heat  transfer
 surface on fuel burning equipment that allows
 lower  final combustion product temperature,
 lower  combustible  losses, or reduced excess
 combustion air.
Process and Cost of Energy Conversion

   Conventional  steam  generator designs
promote  rapid  combustion  reactions.  The
rapid combustion reactions produce very high
temperatures; in  fact, flame  temperatures
closely approach those produced at adabatic
conditions, which are a factor in combustion
chamber  construction. It has become com-
mon  practice to water-cool the combustion
chamber  to  control maintenance from  the
high flame temperatures. Heat absorbed in the
combustion chamber water-cooling system is
recovered by generating steam.

  In burning low ash fuels, the size of the
combustion chamber is dependent on the fuel
and the combustion equipment. A slow burn-
ing fuel or a low turbulence burner may dic-
tate  a larger residence time  at high tempera-
tures to complete combustion.  Where ash in
the fuel is molten at higher temperatures, the
combustion furnace chamber water-cooling is
used to lower (through radiant heat transfer)
the  combustion product  gas  temperature
below the ash fusion temperature. In  conven-
tional design,  flue  gas temperatures leaving
the combustion chamber range from 2000 to
2500°F.

  Convection heat transfer  is  the principal
way  to  reduce  the  furnace  combustion
products exit  temperature  to  the final or
stack temperature. Although convective heat
             Table 1. COMPARISON OF HEAT LOSSES AND BOILER EFFICIENCY FOR
                                THREE DIFFERENT FUELS3
                Heat losses, % of heat input
    Natural    Residual   Bituminous
     gas      oil (#6)      coal
           Sensible heat in dry flue gas

           Heat loss due to water in (and formed from)
      4.2
4.5
           aAssumptions:   combustion product final temperature—300°F;
                         natural gas excess combustion air-10%;
                         oil excess combustion air—15%;
                         coal excess combustion air—25%; and
                         type of coal firing-spreader stoker.
5.0
the fuel
Heat loss due to moisture in combustion air
Unburned combustible
Radiation and unaccounted 'for
Total heat losses
Heat output (boiler efficiency)
10.6
0.1
—
1.5
16.4
83.6
6.0
0.1
—
1.5
12.1
87.9
4.2
0.1
3.0
1.5
13.8
86.2
V-3-2

-------
transfer depends somewhat on the local con-
ditions of gas viscosity tube configuration and
some  other  minor factors, it depends mostly
on velocity. However, two  factors limit ob-
taining very high coefficients.


  One limit stems from conventional design
practice  which has been changing over the
years  in  an attempt  to optimize power and
heat absorption rates for gas and  oil fuels:
power requirements for air-moving equipment
are  a  direct  exponential function of velocity.
The other limit is the presence of particulate
material  in the flue gas, with erosion of con-
struction materials a necessary consideration.
For examples  of the  above, velocities in the
convective zones for  oil- and  gas-fired steam
generators have increased to 120   140 ft per
second; for spreader  stoker firing, velocities
are  normally  limited to  75  ft per second.
Convection  heat transfer rates can be much
greater for a  particulate-free  gas stream, or
when the erosiveness of the particulates  is
less.

  The temperature of  flue gas leaving the
steam  generator  (that  is, upstream of the
separate  heat  trap) is largely  determined by
practical design factors that are not apparent
to  the  casual observer;  optimization  of
supports, drum ligaments, fabrication  capa-
bility, and assembly costs are just a few of the
many factors that must be considered by the
boiler designer. It  is  normal, however, to be
able to  achieve gas  temperatures  of about
700°F in industrial  boiler  practice. Useful
heat remaining in the  combustion product gas
stream can  be roughly computed  from the
above and corresponds  to a 400° F tempera-
ture change.  Often  only one heat trap  is
necessary to  produce  the  desired  thermal
efficiency of the>steam generator system.

  An important philosophy, here, is that the
steam generator is a combination of combus-
tion chamber  and convection zone. Utiliza-
tion of all the heating surface configuration
that is  physically  possibly  adds little more
cost than only partial use. Heat traps, separate
items, are an additional expense. An increase
in heat duty  assigned to a heat trap has a
direct effect on  the  cost  so  that, if boiler
heating  surface is sacrificed to where heat
traps  must  be used,  overall  costs must  be
increased. For example, if there is no convec-
tion surface in  parallel with furnace tubes
from  the steam  drum  there is  only  partial
utilization of the steam drum physical capa-
bilities.
  In  the  discussion of heat traps, there are
two types of combustion air heaters and also
flue  gas  contacting  feedwater  heaters  or
"economizers."  There are  regenerative and
tubular air heaters; economizers may be "bare
tube" or  "extended surface." Where  a com-
bustion air heater is used, the heat absorbed is
directed  to  the combustion chamber and
convective boiler  surface;  where  an econ-
omizer is used,  there  is no increase  in heat
input to the combustion chamber.
   The coefficients  that can be obtained and
the  heat  transfer  temperature  differences
favor the application  of economizer, but a
limiting factor is the amount of heat that can
be absorbed in an  economizer. Conventional
design avoids  obtaining  saturated tempera-
tures in  the  feedwater. Normal  industrial
applications have   feedwater  inlet tempera-
tures of about 220°F,  allowing a temperature
rise of over 100°F in the economizer which is
sufficient to achieve  the desired boiler effi-
ciency.
   Combustion air heaters find a limit on the
maximum  possible air temperature  as  it
affects the mechanical design. It is common,
for example,  to limit  stoker  combustion air
temperatures to 400°F. Another limit is the
"cold  end" problem  where  metal  parts  in
contact with flue gas are at the dew point. It
is common  to find a separate heater used to
limit the lowest metal temperature to some-
thing above the flue gas dew point.
                                     V-3-3

-------
 Operational Reliability

   It is impossible  for an  unscheduled shut-
 down  not to affect the economy of plant
 operation. Maintenance may be a trade off for
 operating economy or for  lower capital cost,
 but designs that do not have predictable reli-
 ability are unacceptable.

   It is very  difficult to  describe normally
 expected  reliability. Reliability can vary from
 indefinite periods of operation approaching a
 full  year (typical  in  petroleum refining) to
 periods of 12 hours on weekdays only (as in a
 carpet  mill).  Reliability  is  not  necessarily
 related to maintenance. For example, whereas
 a petroleum  refinery  may  be willing to com-
 pletely rebuild a boiler once a year if the reli-
 ability  is re-established,  a  carpet mill  would
 expect  virtually  no  maintenance costs for
 several years.

   If reliability were affected by a mainte-
 nance problem (such as burner wear or clean-
 liness) it is a normal standard  to have on-line
 replacement  capability to  restore perform-
 ance. If the problem were continual, full-load
 capability may be expected even with main-
 tenance being performed.  It is  normal to
 expect  no maintenance  problems with the
 steam  generator  in regard to tube life or
 material considerations. The typical industrial
 boiler customer prefers to have no steam-side
 maintenance, including circulation pumps or
 other items that could affect steam-generator
 reliability.
burning  techniques. In addition, if there  is
reduction of sulfur gases in  the combustion
products, a lower final temperature may be
possible.

   How the chemical energy is converted to
useful  thermal  energy is somewhat different
in an FBC boiler. Combustion takes place at a
controlled temperature (now thought to be
1650°F  for  optimum  pollution  control).
There is a great deal of discussion on the pre-
diction of submerged heating surface. Some
experimenters feel that the overall coefficient
of 50  Btu/ft/°F, commonly assigned to the
submerged heating surface,  is  too  conser-
vative.  Suffice it to say that the mechanics of
predicting bed  temperature are  not well de-
fined. Radiation heat transfer capability from
flue  gases at the bed temperature as they
traverse the combustion chamber signify that
further temperature reduction of combustion
products  must be  through  convective heat
transfer.

   Using the previous outline, heating surface
in the  bed and in the walls  of the fluid-bed
chamber requires a system of steam-generat-
ing tube headers and  steam drun^  comparing
to the  water-cooled  furnace in the  convec-
tional  unit. Further utilization of the steam
drum and separators, by adding parallel con-
vective heat transfer surface,  should be possi-
ble and is a factor in the temperature reduc-
tion of the combustion gases. It seems reason-
able that  an  equivalent reduction to conven-
tional  designs can be made;  so, a single heat
trap would be sufficient.
FLUIDIZED-BED COMBUSTION STEAM
GENERATOR

  The above analyses set the stage for a dis-
cussion of the application of fluidized-bed
combustion (FBC) to industrial boilers.

  More chemical energy can be converted in
an FBC boiler than in conventional types. The
outlook  is that excess air values and  com-
bustible losses may be less than in other coal-
V-3-4
   Erie City has begun to develop FBC boiler
designs to  test certain  features.  A  market
survey of industrial boilers has confirmed ,a
firm  market position for steam  generators
having  250,000  Ib/hr capacity,  producing
750°F steam at 600 psig. To begin our anal-
ysis,  we used  these  requirements  as  design
basis. Also, we established that the physical
configuration  of the  FBC  steam  generator
would lend itself to shop assembly  and to
shipment by railroad.

-------
   We  tried two different fluidized-bed de-
signs.  One  utilized modular shop-assembled
boiler  units, with a fluid-bed chamber shipped
separately from  the boiler convection surface.
The two would be joined together in the field.

   In   all  arrangements,  we  proposed that
combustion gases would flow in the freeboard
area to one end  of the chamber. They would
be  collected and passed through the convec-
tion zone.  In  one case, we considered  sloped
or  vertical  surface  penetrating the grid  and
passing through the freeboard area. In another
case,  we considered only  heating surface in
the bed in the  combustion chamber.

   In still another trial, we set the goal as a
totally shop-assembled unit. Where there are
sloped tubes  in  the bed  and the bed free-
board, we were able to duplicate conventional
boiler  performance  if  there  was sufficient
heating surface to have only one heat trap to
produce the desired boiler efficiency. Table 2
compares performance values of the two-piece
modular arrangement and  of the single-piece
unit. Although the advisability of the flow to
one end may be questioned, it does have very
real advantages. Heat transfer surfaces may be
installed in the  freeboard area and  (most
important)  a particulate removal system may
be  installed to protect against excessive un-
burned carbon recirculation through both the
convection  pass and the heat trap. In accumu-
lating all gases, convection pass  heat transfer
coefficients may be improved using conven-
tional configurations of heating surface.

   Figures 1 and 2 show two arrangements of
FBC  boiler components  demonstrating  the
use of an  intermediate particulate removing
system.

   So far, we have not mentioned transfer of
heat to produce steam to elevate the useful
energy that is  transmitted  to the process. It is
fundamental that the superheater must be as
carefully designed in FBC boiler applications
as it is in conventional designs. Compared to
conventional design,  the  fluid-bed  system  is
largely  an  on-off device; consequently  com-
pletely  submerged superheater surface  does
not  seem  appropriate when  low  loads are
considered.  A  combination of submerged in
the bed and conventional surface appears to
be better. Figure 3 shows one way to protect
against  excessive superheater submergence in
the bed at low steaming rates: steam flow  is
biased toward  the submerged surface to pre-
vent overheating at low steaming capacity.
 CONCLUSION

   Reliability  of the  FBC industrial  boiler
 appears to be the heart of the design problem.
 Although we have not designed the details of
 a  fuel injection system or grid, we have set
 three goals:

   1. The  fuel injection system  must have
     on-line replacement capability.
   2. The grid  system must  have positive cool-
     ing to protect against temperature excur-
     sions approaching bed  temperature.
   3. The grid  must  be replaceable without
     major alterations to the steam generator.

   We feel that the fuel feed systems and grid
 design are the principal hurdles to industrial
 steam generator  customer  acceptance of an
 FBC  boiler design.  We feel that the other
 objections can be overcome by a design that
 sacrifices shop assembly.

   At  this point,  we  are not at all confident
that  the  FBC  atmospheric steam generator
will win  acceptance  over  the present alter-
natives without full utilization of the poten-
tial for  pollution control.  With pollution
control, however, lower cost and more readily
available fuel supplies can be used. It is in this
area, that air-fluidized-bed  combustion tech-
niques will have  worthwhile economic com-
pensations.
                                                                                   V-3-5

-------
              Table 2. COMPARISON OF DESIGN VALUES FOR TWO FLUIDIZED-BED
                             COMBUSTION BOILER DESIGNS
Characteristic
Fuel
Steam flow, Ib/hr
Steam pressure, psig
Steam temperature, °F
Boiler efficiency, %
Heat input, M Btu/hr
Heat absorption M Btu/hr
Combustion chamber & superheater
Convection pass
Economizer
Total heat absorption
Gas temperatures, °F
Leaving combustion chamber
Leaving convection pass
Leaving economizer
Combustion air flow, Ib/hr
Excess air, % of theoretical
Flue gas flow, Ib/hr
Fluid bed press, drop, in. wg
Freeboard press, drop
Convection press, drop
Total press, drop, in wg.
Gas velocity, ft/sec
Through furnace
Entering convection pass
Superheater press, drop, psi
Heat transfer surface, ft2
Submerged in bed
Surface above bed
Convection pass
Superheater
Total heat transfer surface, ft2
Approximate height/length/
width, ft
Modular design
(no freeboard
heating surface)
Coal
250,000
600
750
87.7
337

195.6
68.5
24.9
289.0

1530
672
350
275,000
10.0
298,000
40.0
.0
6.2
46.2

45
100
50

3,200
1,'TOO
7,710
860
12,870
16/40/17

Single unit (with
sloped tubes in bed
and freeboard
heating surface)
Coal
250,000
600
750
87.7
337

244.0
20.1
24.9
289.0

925
672
350
275,000
10.0
298,000
40.0
4.6
7.1
51.7

67
90
50

3,200
4,970
3,510
970a
12,650
16/40/13

             aDoes not include 620 ft* in the bed.
V-3-6

-------
         FLUIDIZED
           BED
         CHAMBER
^-\  "
 FAN
                    CAR-
                    BON
                    BURN-
                     UP
                    CELL
                                       ECONOMIZER
.__!__.)
                             CONVECTION
                                ZONE

                                      CARBON-RICH
                                      ~~"ASH "'""
                     I        I
                    MECHANICAL
                    Li  DUST  T-*-
                     COLLECTOR

                    .W
                                                              ELECTRO
                                                               STATIC
                                                              PRECIPI-
                                                               TATOR
                                                             W
                                                                         STACK >
                                                                  ASH TO
                                                                 DISPOSAL
                Figure 1.  FBC system components - arrangement No. 1.
                                                                         XXX
  FAN
                                                           I       I
                                                         ELECTROSTATIC
                                                         PRECIPITATOR

                                                                      'STAClO
                                                            1	J	
                                                                ASH TO
                                                               DISPOSAL
                 Figure 2. FBC system components - arrangement No. 2.
SATURATED STEAM
FROM STEAM DRUM
              	SUPERHEATER
              	ELEMENTS ABOVE	
              	FLUIDIZED BED _ -
               •SUPERHEATER ELEMENTS
                  SUBMERGED IN
                  FLUIDIZED BED
                                            FLOW-BIASING
                                               VALVE
                                                                 SUPERHEATED
                                                                 •  STEAM TO
                                                                   PROCESS
                                                    SUPERHEATER OUTLET
                                                         HEADERS
                        Figure 3.  Superheater arrangement.
                                                                           V-3-7

-------
                                              4.  DESIGN FEATURES
                                                     OF PRESSURIZED
                                        FLUIDIZED-BED BOILERS
                               R. W. BRYERS
                       Foster  Wheeler  Corporation
INTRODUCTION

  As part of its effort on a NAPCA-funded
study* to evaluate the fluidized-bed combus-
tion  process,  Foster  Wheeler is studying
conceptual designs of fluidized-bed boilers.
Although  atmospheric-pressure fluidized-bed
boilers for a 600-MW plant  are also  to  be
studied, work  thus far has concentrated  on
pressurized boilers for a 300-MW plant. These
pressurized 300-MW designs are the subject of
this presentation.
ADVANTAGES OF PRESSURIZATION

  Application of fluidized-bed combustion to
a utility steam generator requires selection of
a cycle and  conceptual design which will
make best use of the improved heat transfer
offered by fluidization.  Considering the rela-
tively low  average  temperature at  which
combustion takes place  in the bed, the com-
bined steam-turbine/gas-turbine  power plant
with a pressurized boiler (see Figure 1) must
be considered an immediate candidate. The
higher heat transfer coefficients and high gas
densities of pressurized operation should offer
an excellent opportunity to reduce the first
cost of the steam generator.

  Pressurizing a conventionally fired boiler in
this manner  reduces the amount of heating
surface required, which reduces its size and
' Westinghouse is prime contractor on this program.
 weight. Ultimately these reductions are re-
 flected in a reduction in first  cost as comp-
 ared to a conventional non-pressurized boiler
 designed  for the same steam conditions and
 net cycle output. The reduction is achieved as
 a result of an increase in emissivity of the
 non-luminous gases, higher gas densities and
 available pressure drop, higher permissible gas
 mass  flows and  convection  heat  transfer
 coefficients, and  the reduced boiler duty
 requirement which  results  from  splitting
 power generation between the gas turbine and
 steam turbine.  Fluidized-bed vcombustion
 should offer further improvements in  first
 cost by  making use of high heat transfer
 coefficients  throughout virtually  the entire
 exchange of energy from the combustion gas
 to the steam, which is true when gas-turbine
 inlet  temperature closely  approaches  the
 fluid-bed temperature.
 CYCLE SELECTION

   Cycle  selection is  ultimately dictated by
 requirements for gas side pressure and steam
 conditions.
Gas Side Pressure

  Selection of the optimum gas side pressure
is, of course, determined by studying the
effect of pressure on overall plant economy.
An examination by Westinghouse of the influ-
ence of pressure on plant heat rate and output
                                      V-4-1

-------
 indicates  that there  is  little advantage  in
 operating  the gas  turbine  at  boiler inlet
 pressures  above 10 atmospheres. For a given
 boiler capacity, the cost of the pressure shell
 is most strongly related to gas pressure. Sur-
 face  requirements  are a function of  steam
 cycle conditions and, except for some second-
 ary  effects,  remain  unchanged  by gas side
 pressure. Effects of pressure on cost of auxil-
 iaries such as coal handling equipment cannot
 be ignored. However, since they are beyond
 the  scope  of  this  presentation  they will
 temporarily be overlooked.

   The cost of the shell is affected by pressure
 level, gas  velocity,  capacity, and  diameter.
 The latter two variables are related in that the
 designer can resort to modular design  at an
 appropriate plant capacity  to minimize the
 effects of a further increase in shell size and
 to make  use  of shop fabrication techniques
 rather than field erection. Below a 12-1/2 ft
 diameter,  a substantial reduction in cost can
 be realized by using shop fabrication. Modular
 construction, however, is not without bounds.
 When more than four or five  vessels are con-
 sidered, the effective reduction in vessel size
 diminishes rapidly, as shown in Figure 2. Gas
 velocity is more or  less fixed by the combus-
 tion process. Any effect  it may have on shell
 size is secondary  compared to the other varia-
 bles.  As  pressure level  increases,  shell size
 decreases  rapidly due to the increase in gas
 density. Above 10 atmospheres, this effect on
 shell  size  diminishes,  as shown in  Figure  3.
 This diminished effect is compounded by the
 increase in shell thickness required to  with-
 stand the  higher pressure.  Considering that
 the incremental cost  of  auxiliary equipment,
 such  as  the coal handling system,  also in-
 creases with an increase in pressure,  10 atmos-
 pheres appears  to  be an optimum pressure
 level.
Steam Conditions

  The economics can be enhanced by apply-
ing tne pressurized boiler to either the super-
V-4=2
critical or subcritical pressure steam cycle and
using the once-through  principle.  Doing so
eliminates the large and costly steam drum
required by  a  natural circulation  boiler, as
well as  the  risers and  downcomers  which
would  otherwise require numerous penetra-
tions of the relatively small pressure contain-
ment shell. Once-through  circuitry permits
greater  freedom in surface arrangement and
location.

  In factors influencing  boiler cost, the sub-
critical  cycle  offers advantages  (over the
supercritical) of thinner wall tubes,  lower
tube metal temperatures, and some improve-
ment in mean  temperature between the gas
and the steam.

  Westinghouse thermodynamic   and  cost
analyses of plants using subcritical and super-
critical steam cycles combined with 10-atmos-
phere gas turbines also show that using super-
critical steam conditions is  not advantageous.

  Based on the  foregoing considerations, the
design conditions of Table  1 were  selected as
a basis for the conceptual design of fluidized-
bed boilers for a 300-MW plant. The approxi-
mate steam conditions in the boiler are shown
superimposed  on  the temperature-enthalpy
diagram in Figure 4.
CONCEPTUAL DESIGNS

Basic Criteria

  Conceptual arrangements of the pressurized
fluid-bed boilers are based on several funda-
mental criteria.

  1. Where possible, modular construction is
     used to take the greatest possible advan-
     tage  of shop fabrication techniques and
     (thereby) of lower cost.
  2. A  minimum of  four  completely  self-
     contained modules  per boiler facilitate
     the highest possible turndown ratio  in
     the shortest permissible time.

-------
  3. The  fluid  beds are further  divided  by
     heating functions (i.e.,  pre-evaporative
     heating, evaporation, superheating,  and
     reheating)  to  simplify  startup proce-
     dures,  reduce startup time, and improve
     control of the various heating functions.
  4. Pre-evaporative heating and evaporation
     may be combined in a single bed, where
     vertical tubes  are  being considered to
     minimize tube  bending and manifolding
     and  to avoid possible tube overheating
     due  to stratified flow  of  steam  and
     water.  (Under  these circumstances a
     minimum of 12 beds are required of two
     different sizes.1)

  Pressurized  operation offers the advantage
of higher permissible pressure drops through
the bed without  appreciably affecting plant
economy. This means it is possible to go to
higher gas velocity and much deeper beds. For
example, at 1 atmosphere and gas velocities of
4 ft per second, the bed may be about 2 or
2-1/2 ft deep. A comparable bed at 10 atmos-
pheres and  15- ft  per second could be 30 ft
deep. Going to higher pressures changes  the
ratio of height to plant area considerably.

  Going to high gas velocities runs the risk of
erosion of  tube  surface, elutriation of large
quantities of fly ash and carbon, and increas-
ing the carbon  loss. Instead, conceptual  de-
signs are based on stacked beds with gas veloc-
ities of about 4 ft per second; the volume
occupied  by the bed remains the same as if
gas  velocities  of  15  ft  per  second  were
selected. This same volume  is accomplished
by splitting the air flow and putting four beds
in parallel on  the air/gas side rather than pass-
ing all the air through a single bed.
Preliminary Layouts

  Figures 5 and 6 show two possible arrange-
ments of the boiler. In  both these arrange-
ments, air  enters the shell at the bottom of
the vertical boiler and flows upward through
the space between the shell and a finned tube
water wall which encloses the boiler proper.
The required portion of the air is drawn off at
each  bed level, and passed  into  a plenum
chamber housing   the  tube  headers, and
through a grid plate into the bed. Gases leav-
ing the fill space above each bed pass through
an opening in the inner water-cooled wall into
a central channel and thence downward out
of the vessel. The gases are returned to the gas
turbine in an air-cooled double-lined  pipe in
which the air from the  compressor  is the
coolant.

   This  arrangement  not  only  provides  air-
cooling of pressure parts, headers, and coal

   Table 1. DESIGN CONDITIONS FOR COAL-
    FIRED, PRESSURIZED, FLUIDIZED-BED
         BOILER FOR 300-MW PLANT

Design Characteristic                   Criteria

Primary steam flow, Ib/hr              1,727,000

Superheater outlet pressure, psig         2500

Superheater outlet temperature," F       1000

Reheat steam flow, Ib/hr               1,644,000

Reheat inlet temperature, °F            650

Reheat outlet temperature, °F           1000

Reheat inlet pressure, psig              600

Reheat outlet pressure, psig             580

Feedwater temperature, °F             578

Fluidized-bed pressure, atm             10

Air flow to boiler, Ib/hr                2,340,000

Air temperature to bed," F              575

Fluid-bed temperature, ° F              1700

Gas temperature leaving boi ler," F        1600

Bed-to-tube heat transfer coefficient,      50
   Btu/hr/ft2/°F	

                                      V-4-3

-------
 conduits, but also ensures that in-leakage of
 air (rather  than out-leakage of hot gas) will
 occur in the event of any breach in the finned
 tube  enclosure  wall. It also reduces the total
 external piping required  by  carrying, the hot
 gases within the pressure vessel a part of the
 way to the gas turbine.

   Annular  beds  are used  in the  Figure  5
 arrangement, which has the advantage of
 making maximum use of the space available
 within  the  shell. It  also separates zones of
 high  differential pressure on  the  air/gas side
 with   cylindrical  surfaces which  minimizes
 bracing and reinforcement.

   Coal is introduced tangentially at a number
 of points on the outer periphery of the beds.
 Three injection  points per bed should  give
 adequate distribution to  permit blowing the
 coal  in  through  tubes located between the
 water tubes of the enclosure wall.

   Tube  bundles  in the  superheater and re-
 heater sections consist of vertical, multi-start,
 helical coils that fill the annular beds.  The
 vertical evaporating tubes are arranged  radi-
 ally in the annular bed and are connected to
 ring-type headers. The enclosure wall  tubes
 serve as part of the pre-evaporator surface; the
 balance is within the evaporator bed.

   Helical coils are advantageous for this appli-
 cation because of the relatively small number
 of tubes leading into and  out  of the bed. The
 same  is true of flat spiral coils, which might
 also  be  employed. However, coils have the
 disadvantage of high fabrication cost. Also, it
 is impossible  to  replace  individual  tubes
 within a coil bundle; the only recourse is to
 plug the ends of leaking tubes.

   The arrangement of Figure 6 uses rectan-
 gular  beds and  the more conventional  hori-
 zontal return-bend tube elements in the super-
 heater and  reheater. This arrangement  pro-
 vides  for easier  maintenance but has  the
 disadvantage of large flat  surfaces which have
 to be reinforced to withstand gas differential
V-4-4
pressures. Space utilization within the shell is
not as efficient as with annular beds.

  Cyclones and  dip  legs (not  shown  in
Figures 5 and 6) are in  the free space above
beds to collect elutriated material and return
it to the beds.

  Note that the just-discussed  concepts are
preliminary. They are two  of several bed-and-
surface  arrangements  that will be  critically
evaluated prior to developing a single detailed
design. Other  concepts include a single deep
bed  in each  vessel, and  horizontal vessels
containing relatively shallow beds. Much more
detailed  study of the  manufacturing, assem-
bly,  and  maintenance problems involved in
these arrangements is required.
Surface Requirements

   Heating surface area and bed volume re-
quirements of the boiler have been calculated
by assuming that 2-in. O.D. tubes are  used
throughout and that the tubes in the bed are
placed on an average spacing equivalent to a
4-in. square pitch. Although 2-in.  O.D. tubes
throughout may not be optimum, they are a
reasonable choice, considering  problems  of
support, possible  vibration,  and the number
of circuits and tube-to-heater connections. It
has also been conservatively assumed that the
bed-to-tube heat transfer coefficient  is con-
stant throughout the bed at  50 Btu/hr/ft2/°F,
regardless of tube orientation. Table  2 gives
the results of calculations for the annular bed
arrangement  shown in  Figure 5.  For  the
rectangular bed arrangement of Figure 6, the
amount of pre-evaporator surface in the bed
would  be reduced  because of the  greater
amount of enclosure wall surface.

   Using smaller than 2-in O.D. tubes of the
same pattern  and pitch-to-diameter ratio re-
duces bed volume considerably. For example,
for 1-1/2 in. tubes on 3-in. centers and 1-in.
tubes on 2-in. centers, the bed volume will be
about  three-quarters and one-third as much,

-------
respectively, as  for  2-in.  tubes  on  4-in.
centers.

COMPARISON WITH CONVENTIONAL
BOILER

  Table  3 compares a pressurized fluidized-
bed  boiler to a conventional boiler for the
same steam conditions and plant output. The
maximum  mean tube wall  temperatures are
nearly the same  in both units. The cost of
material  in  the superheater and  reheater of
the fluidized-bed  boiler is only  about one-
fifth of that in the conventional boiler. Sur-
prisingly, no austenitic steel is required in the
reheater. All this, however, is based  on the
relatively conservative bed-to-tube heat trans-
fer coefficient of 50  Btu/hr/ft2/°F. Any  in-
crease in this coefficient changes the material
picture considerably.
                    Table 2. CALCULATED SURFACES AND BED VOLUMES-
                     PRESSURIZED, FLUIDIZED-BED BOILER FOR 300-MW
                                        PLANT3
Component
"Pre-evaporator
cEnclosure walls
Immersed in bed
Evaporator
Superheater
Reheater
Heating
surface,
ft2
,
10,500
3,300
9,780
15,820
8,350
Bed
volume,
ft3

-
700
2,070
3,360
1,770
Overall
H.T. Coeff,
Btu/hr/ft2/°F

20d
47
47
43
39
                   aAII values are approximate.
                   ^Enclosure wall length based on assumed 15-ftfree height above
                    each separate bed.
                   cPortions of the enclosure walls are immersed in the beds.
                   ^Weighted average.
                                                                                    V-4-5

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                    Table 3. COMPARISON OF CONVENTIONALLY FIRED AND
                      PRESSURIZED FLUIDIZED-BED STEAM GENERATORS
                                 FOR 320-MW PLANT OUTPUT
                Characteristic
Conventional
Fluid-bed
Primary steam flow, Ib/hr
Superheater outlet pressure, psig
Superheater outlet temperature, ° F
Reheat steam flow, Ib/hr
Reheat inlet temperature, "F
Reheat outlet temperature, °F
Reheat inlet pressure, psig
Reheat outlet pressure, psig
Feedwater entering steam generator, "F
2,189,000
2,500
1,000
1,936,000
650
1,000
600
580
486
1,727,000
2,500
1,000
1,644,000
650
1,000
600
580
578
                Superheater
                   Surface area, ft2                        61,660         15,820
                   Max. mean tube wall temperature," F        1,095          1,078
                   Material weights, Ib
                     Carbon steel                        108,000         36,000
                     C-1/2%Mo                                        24,000
                     112% Cr -172% Mo                   62,000         17,000
                     1%Cr-1/2%Mo                                    19,000
                     1 -174% Cr -1 /2% Mo                 73,000         22,000
                     2-1 /4% Cr -1 % Mo                  430,000         22,000
                     18Cr-8Ni                         20,000          7,000
                   Total material weight, Ib                693,000        147,000
                   Relative material cost, %                    100             19

                Reheater
                   Surface area, ft2                        67,800          8,350
                   Max. mean tube wall temperature, "F        1,115          1,136
                   Material weights, Ib
                     Carbon steel                        140,000
                     1 /2% Cr - 1 /2% Mo                  133,000
                     1 -1 /4% Cr -112% Mo                 59,000         24,000
                     2-174% Cr -1 % Mo                   64,000         14,000
                     9%Cr-1%Mo                                     21,000
                   Total material weight, Ib                396,000         59,000
                   Relative material cost, %                    100             18
V-4-6

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STEAM TO

STEAM TURBINE

CYCLE
                                                                            GENERATOR
                      FEEDWATER FROM

                      STEAM TURBINE CYCLE
         Figure 1.  Combined steam-turbine/gas-turbine cycle with pressurized boiler.
      1.0
 _
 uj co
 CO UJ
 to >
 UJ *^
      0.8
 5 2 °'7
 cr t

 "- § 0.6

 ui 2
 E <
 D^ 0.5
     0.4
 1- OC
      0.3
      0.2
 25  o.l
f
T
I     I     I     I     I     I
I     I
                                                                                    I
             I     I	I	I	I	L
                                 5    6    7     8     9   10   11   12   13   14   15   16

                                     NUMBER OF VESSELS SELECTED
           Figure 2.  Effect of the number of vessels selected on the vessel size.
                                                                                     V-4-7

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                                              345
                                          GAS FLOW, 106 Ib/hr
                Figure 3.  Effect of gas flow and pressure  on vessel  diameter.
     LU
     DC
     i
1100

1000

 900

 800

 700

 600

 500

 400
                                  PRE-
                              EVAPORATOR
               EVAPORATOR
                        I
I
I
I
            300  400   500   600   700
                                800   900   1000   1100   1200   1300  1400  1500
                                   ENTHALPY, Btu/lb
                 Figure 4.  Temperature-enthalpy diagram for steam and water.
V-4-8

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                             HELICAL TUBES IN
                             FLUIDIZED BED
                             FEEDWATER
                             PREHEAT
                             TUBES
                            VERTICAL TUBES IN
                            FLUIDIZED BE
EVAPORATOR
SECTION
                                                     SUPERHEAT
                                                     SECTION
                                                     TWO
                                                     SECTIONS
                                                     (ONE NOT
                                                     SHOWN)
                                 VERTICAL TUBES
                                 IN FLUIDIZED BED
                                                                                FEEDWATER
                                                                                PREHEAT
                                                                                TUBES
                              HORIZONTAL TUBE BANKS
                              IN FLUIDIZED BED
   Figure 5.  Pressurized fluidized-
   bed boiler — annular beds.
Figure 6.  Pressurized fluidized-
bed boiler - rectangular beds.
                                                                                     V4-9

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                                5.  SOME  ECONOMIC ASPECTS
                                       OF HIGH-TEMPERATURE
                                                        STEAM  CYCLES


                 D.  E. ELLIOTT AND E. M. HEALEY
                      University of Aston, England
INTRODUCTION

  One of the  most important aspects of
fluidised-bed  combustion  is  the  very  low
corrosion rates which occur when metals are
immersed in the solids; hitherto, steam cycles
have been limited to a temperature of 1050°F
because of fouling and corrosion of super-
heater tubes. Thus the heat rate of steam
plant, which showed a downward trend until
the mid  1950s, has  now flattened out at
about  2.5 kW heat/kW electricity. As  flui-
dised-bed  combustion may  allow  us to
continue the downward trend of the heat rate
curve,  we  have undertaken a  brief examina-
tion of the cost of achieving better steaming
conditions.
FOULING AND CORROSION

  Dainton and Elliott1  immersed a typical
stainless steel probe for  120 hours in a fiui-
dised bed burning Thoresby coal at a tempera-
ture  of  1300°F. This coal,  notoriously
viscous, has  a high chlorine/sulphur content
and would cause a normal power station to
shut down within a day because  of  savage
fouling. The  probe was cooled internally  by
air so that the exterior in contact with the ash
varied in temperature from 900 to 1200°F.
The metal surface which was at a temperature
above  1100°F  was clear of any  deposit  or
corrosion; at temperatures below  1100°F,
fouling  became progressively worse as the
temperature  decreased. The deposit, mainly
                                      V-5-1
sodium chloride, was accompanied by a small
amount of surface  pitting. By contrast, the
fouling of superheater tubes in a test rig at
B.C.U.R.A., simulating pulverised fuel  firing
and using the same coal, gave massive deposits
which almost  bridged across the tube bundle
in a 24-hour experiment.

   As the absence of fouling and corrosion is
one  of the cornerstones of our present pro-
posals,  it  is worth looking at some of the
fundamentals  involved. Figure 1  shows the
variation with temperature  of the  vapour
pressure of some  of the  alkali  metal salts
present in most coal ashes.  It will be seen that
the vapour pressure levels at normal flame
temperatures  of about 2500°F  are several
orders of magnitude higher than those in a
fluidised bed operating at 1500°F. As fouling
is  a  process of deposition from  the vapour
phase, it is reasonable to predict much lower
fouling  rates at  the lower temperatures.
Figure  1 also shows that the vapour pressure
of sodium chloride at 1300°F is significantly
higher than that of most of the  other  com-
pounds and would therefore be the one most
likely to be deposited; this agrees with experi-
mental  results. It is interesting to note that
the vapour pressures of the sulphates, which
are a common cause of fouling  in  normal
boilers, is negligible under fluidised-bed condi-
tions.

  The absence of fouling  on  the metal sur-
faces which are above 1100°F can be likened
to the  phenomenon of deposition of dew

-------
 from moist  air.  Unless  there is a sufficient
 temperature  differential,  little  deposition
 occurs; if there  is any, it could  well be re-
 moved  by the mildly abrasive action of the
 solids. Only  when there is a substantial  tem-
 perature difference does the deposit adhere. It
 should  be pointed out that  the experiments
 were conducted  with very  fine coal  at low
 fluidising  velocities;  we could expect  less
 deposits when using larger  coal,  unless the
 nature of the fluidising conditions gives rise to
 particle  temperatures well in excess of bed
 temperature. When similar trials were under-
 taken with coals which are considered to be
 bad (but tolerable) to use in  normal systems,
 no fouling whatsoever occurred at any metal
 surface temperature.

   These preliminary encouraging results  have
 been supported by experience in the larger
 combustors now running; although  we need
 more evidence from much  longer  duration
 experiments, it is  useful to enquire if the
 potential to go to higher temperatures would
 be worthwhile economically.
 STEAM CYCLES

   Four steam cycles have been investigated:
 the first two were based on the design condi-
 tions of an existing supercritical p.f.-fired
 power  station;  the other  two were  cycles
 considered by Downs.2 Table 1 gives working
 conditions and the efficiency of these cycles,
 together with those of a standard 500-MW set
 which has been taken as datum. It is not sug-
 gested that these cycles would be adopted in
 an actual design; however, they represent (at
 the lower end) plant which could be achieved
 without  departing from current steam turbine
 technology  and  (at the  upper end)  plant
 which could possibly be developed in the next
 decade.
FLUIDISED-BED CONDITIONS

   Tube  cost  calculations  are  based on the
fluidised-bed operating conditions  shown  in
V-5-2
Table  2. These  conditions were deliberately
chosen on  the pessimistic side; e.g., we have
evidence that heat  transfer coefficients in
excess of 70 Btu/ft2/hr/°F can be achieved,
and it is likely that  operating the bed above
1500°F would produce  a more  economical
tube bundle.

  In each case it is assumed that all the super-
heater and  reheater tubes are in the fluidised
bed to  take  full advantage of reduced heat
transfer surface.  Cooling of the gases after the
fluidised bed is  by  normal  convective  sur-
faces. The assumed fluidising velocity is rather
low (2 ft/sec) to correspond to conditions in a
pressurised  system; however, the same argu-
ment would hold for higher velocities.
TUBING DESIGN

  A wide  range of high-temperature  alloys
have been  developed for use in gas turbines
which may prove suitable for boiler applica-
tion; the final  choice would most likely be
made because of weldability and resistance to
internal waterside corrosion, rather than to
optimise the  cost of the tubing. Nevertheless
the cost of tubing to perform a given duty can
be estimated as follows:

  1. For each  metal considered, calculate
     values  of tube  thickness  for metal
     temperatures in the range 572-1472°F
     (300-800°C), using the formula:

         t_dip +40(2f + P)
            2f-P  P(2f-P)
      where,


         t   =  wall  thickness  (in.)  but
                never less than 0.092 in.
         dj  =  internal diameter (in.)
         P   =  working pressure (psig)
         f   =  working stress of metal at
                the  appropriate tempera-
                ture in psi.

-------
                                             Table 1. ASSUMED STEAM CYCLE CONDITIONS

Cycle
No.



Datum

1

2


3

4

>s. Point of
^vcomparison
>v
Cycle3 ^-v
X^
Fluidised-bed 'a'
500-MW cycle
(b)
(a)
Drakelow C
(b)
Logical single-
reheat
(b)
Optimum-pressure(a)
single-reheat ...
(b)
Optimum-pressure(a)
double-reheat (b)
(c)

Generated
power.
MW


500

375

464


460

528


Steam
pressure.
psi

2300
585
3500

735
3500

770
6500
1430
6500
1430
286

Boiler working fluid conditions
Temperature
Inlet,
°F
493
689
505

683
500

700
500
700
500
700
700
Outlet,
°F
1050
1050
1100

1050
1300

1250
1300
1250
1300
1250
1250
Enthalpy
Inlet,
Btu/lb
476
1344
498

1324
488

1341
489
1292
489
1292
1369
Outlet,
Btu/lb
1488
1545
1495

1540
1629

1651
1572
1638
1572
1638
1661

Steam
flow
rate.
Mlb/h
3.4
2.66
2.405

1.794
2.405

1.804
2.405
1.804
2.405
1.804
1.353

Steam
cycle
efficiency.
%

42.8

45.9

47.9


48.6

49.4

Assumed
sent out
peak
efficiency.
%

39.5

42.4

44.2


44.8

45.6


Reduction
in fuel
consumption.
%

0

7.35

11.9


13.4

15.4

aEach cycle shows separate conditions for: (a) main steam; (b) reheater No. 1; and (c) reheater No. 2.

-------
        The second term in the formula is
        a   corrosion  allowance;  vary  it
        according to experience.
 2.  Using t, and knowing the fluidised-bed
     temperature, 0b, calculate the overall
     bed-to-steam   heat transfer coefficient,
     h, assuming:
        (a) a  bed-to-tube  heat  transfer
            coefficient  of 70 Btu/ft2/h/
            °F,
        (b) a tube-to-steam heat transfer
            coefficient of 500 Btu/ft2/h/
            °F,and
        (c) values of thermal conductivity
            for  the  material  under  con-
            sideration.
 3.  Knowing  0fc  and  h,  calculate the
     working fluid temperature 0f.
 4.  Knowing 0^, t, h,  0f, and the metal
     density, determine the relative weight
 Table 2. ASSUMED FLUIDISED-BED OPERATING
                CONDITIONS
 Characteristic
Condition
Combustion bed temperature, °F           1472

Gas temperature at inlet to air
   heater, °F                           662

Waste gas temperature at input
   to stack, °F                          248

Fluidising velocity, ft/sec                    2

Excess air, %                            20

Base plate pressure drop, in. wg               5

Bed depth, ft                             2

Density of fluidised bed, Ib/ft3              50

Bed-to-tube heat transfer coefficient,
   Btu/ft2/hr/°F                          70

Tube-to-working-fluid heat transfer
   coefficient, Btu/ft2/hr/°F               500

Tube internal diameter (I.P.), in.	1

V-5-4
       of  material  to  transfer one unit of
       heat at the  appropriate temperature.
       If the cost of the metal and its cost of
       fabrication is estimated,  the cost to
       transfer a unit of heat can  be calcu-
       lated. This allows the cheapest metal
       to be chosen for any given steam con-
       dition. Normally it would not be de-
       sirable to  use more than four different
       materials  in  one  boiler in  order to
       minimise  the  number of  dissimilar
       welds. Figure 2 shows how the relative
       costs to transfer a unit of heat would
       appear  using 2-1/4  percent chrome
       steel, Austenitic  316, and  Nimonic
       PE16 with  costs per ton  of  £750,
       £3000,  and  £7500,  respectively, for
       steam  pressures  of  3500  psig,  and
       using 1-in. I.D. tubes.
    5.  Finally, calculate the individual costs
       of metal  for the different sections of
       the  boiler  tubing  for  the various
       cycles. Table 3 gives  the results, using
       the above materials and mild steel.

  For comparison,  the cost of the tubes in
the  datum  standard 500-MW, 2300  psi  x
1050°F  x  1050°F  set,  using  the same
fluidized-bed conditions would be £ 0.6/kW.
(We have chosen to compare the  advanced
cycle conditions with those of a fluidised-bed
boiler; we believe the latter will prove more
economic  than  normal   pulverised-fuel
systems.)

  In addition to the increased boiler tubing
cost when going to the higher steaming condi-
tions, there is also an increase in the  cost of
the  economiser  tubing,  due to the  higher
operating pressure.  This increase  has been
allowed for in the overall comparison shown
later.
            FUEL COST SAVINGS

              The present-day worth of the savings in
            fuel has been  calculated assuming an equiva-
            lent  life time load factor of 50 percent,  a

-------
                            Table 3. CAPITAL COSTS3 OF 1-IN. BOILER TUBES
                      Tubing
Cycle 1   Cycle 2    Cycle 3   Cycle 4
                  Main and superheater

                    NimonicPE16         0.24      0.74      2.28      1.92
                    Austenitic316         0.20      0.17
                    2-1/4% chrome steel    0.18      0.11

                  Reheater

                    Nimonic PE16
                    Austenitic316
                    2-1/4% chrome steel
                    Mild steel

                  Total                  0.69      1.29      2.58      2.35

                  aCosts given in £/kW.
                  ^Figures include both reheaters.
	
0.02
0.03
0.02
0.17
0.06
0.03
0.01
0.18
0.08
0.03
0.01
0.26b
0.1 1b
0.05b
0.01b
station  life of  30 years, and  a return on
capital  of 7-1/2 percent  for  fuel  costs  of
3d/therm and 5d/therm. Table 4 shows these
savings together with tubing and other costs.
OTHER COST SAVINGS

  The importance of overall efficiency to the
overall  cost of a power  station is not always
appreciated.  The cost of access roads,  coal
storage, conveyor and mixing systems,  coal
bunkers,  mills, fans,  gas  cleaning plant,
chimnies, foundations, services, cooling water
supplies, condensers, and water  purification
plant, all  costs come down  pro-rata with the
lowering of the heat rate. Since these items
constitute  more  than half  the  total station
cost,  "other savings" are indeed  most  im-
portant. The converse (decrease of efficiency
leads to increased capital cost) is also true, as
has been found in numerous studies aimed at
producing a cheap peak-load power station by
cutting  out  fuel-saving  systems;  e.g.,  feed
heaters and high-temperature superheaters.
        DISCUSSION

           Table 4 shows that the value of the fuel
        cost savings far outweighs the increased cost
        of the high-temperaiuic boiler  tubing and,
        together with the other savings, should allow
        more than enough margin to pay for the in-
        creased cost of the steam turbine. Note that,
        although we could realize a moderate increase
        in turbine inlet temperature, say to 1100°F,
        without altering turbine design, an increase to
        1200°F would  almost certainly  require  a
        change in construction of the steam turbine
        rotor.  This would be a more difficult problem
        than the design of the turbine blading which
        could  rely  heavily  on gas turbine practice.
        However, there would  not appear  to  be any
        insuperable problems in ultimately  producing
        1300°F machines.

           Although the estimates made in this prelim-
        inary investigation are not precise and until a
        detailed design study is made, we  could not
        indicate the optimum  conditions; there seems
        to be a good case for serious consideration of
        high-temperature systems.  In addition  to
                                              V-5-5

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             Table 4. SAVINGS3 IN CAPITAL AND CAPITALISED REVENUE RELATIVE
                                    TO DATUM CYCLE
Characteristic
Boiler tubing costs
Economiser tubing costs
Other station costs'3
Fuel costs
3d/therm
5d/therm
Total savings
3d/therm
5d/therm
Datum
500-MW unit
0.6
0.6
23
72.9
110.1
-
Savings
Cycle 1
-0.09
-0.1
+1.8
+3.8
+6.4
+5.4
+8.0
Cycle 2
-0.69
-0.1
+2.7
+5.9
+9.9
+7.8
+11.8
Cycle3
-1.98
-0.5
+3.0
+6.6
+11.0
+7.1
+11.5
Cycle 4
-1.75
-0.5
+3.6
+7.5
+12.5
+8.8
+13.8
         aSavings given in £/kW.
         bMost likely an underestimate; the full costs influenced by efficiency have not been com-
          pletely examined.
design studies, we need further experimental
evidence of both the internal and  external
corrosion  resistance  of high-temperature
materials; accordingly, it is recommended that
in the next round of pilot-scale plant, oppor-
tunity is taken to incorporate at least some
tubes  running  under  advanced  conditions.
Only if we can show evidence of good possi-
bility of achieving high temperatures will
turbine manufacturers be justified in under-
taking  the  expensive  development  of ad-
vanced machines.

   One interesting aspect of such development
would be the result of an increased demand
for high-temperature metals: would it create a
shortage and a higher price, or lower material
costs? Our calculations suggest that, even with
the very advanced conditions, about 600 tons
of Nimonic  (or  other equivalent  high-
temperature) material would be used  in  a
2000-MW station. Comparing this with the 20
tons of such alloys which are needed for the
engines of a Concorde, it would appear that,
although the demand would  be large enough

V-5-6
to justify quantity production, it is unlikely
to produce a scarcity market.
CONCLUSIONS

   1. The incremental  cost  of  the  tubing
needed  to  achieve  high-pressure  high-
temperature steam in a fluidised-bed boiler is
relatively low: a 6500 psi x 1300°F x 1250°F
x 1250°F boiler should not cost as much as a
conventional  pulverised-fuel-fired 2300 psi x
1050°Fxl050°F boiler.

   2. The fuel  cost savings  which  would
accrue from these higher efficiency cycles are
substantial; they would more than pay/for^the
extra  cost of the  high-temperature  steam
turbine.

   3. In addition to the reduced fuel  costs,
there should  also be a saving in capital costs
associated with the higher efficiency, because
all plant items related to fuel throughput have
their costs spread over a greater output.

-------
  4. Modest  increases in steam temperature
(within  the  existing capabilities  of steam
turbines)  could  be the prime  objective of
pilot-scale plant  because they should present
no  extra development  problem  for the
fluidised-bed  boiler. Opportunity should also
be taken to incorporate much higher tempera-
ture trial  sections in such plant in order to
assess the capabilities of such systems.
BIBLIOGRAPHY

1. Dainton, A.D. and D.E. Elliott. Researches
  into Combustion of Coal. Seventh World
  Power Conference. Moscow, 1968.
2. Downs, J.E. ASME Paper No. 55-SA-76.
ACKNOWLEDGMENT

  The authors wish to thank Mr. D. H. Stock-
well of the Coal Research Establishment for
carrying out detailed analyses and for general
assistance in the  preparation of this paper.
They also wish to  thank  the National Coal
Board for  permission to publish the work.
The ideas and conclusions expressed are those
of the authors, not necessarily of the Board.
                                                                                  V-5-7

-------
                                         TEMPERATURE, 104/T«K
                                                  9
                                                                10
                                     11
      1400
900        800
  TEMPERATURE, °C
600
           Figure 1.  Variation (with temperature) of vapour pressure of alkali  salts.
     10
o
H
Z
z
I
8
uj
I
    0.5
             NIMONIC PEI6
                     FLUID PRESSURE 3500 psig
                          TUBE I.D. 1 'n.
                                            I	I
      400      500      600      700      .800.     900     1000      1100     1200     1300    1400
                                    WORKING FLUID TEMPERATURE °F

  Figure 2.  Relative cost of tubing for the high-pressure conditions of cycles No.  1 and 2.

V-5-8

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                                                        6.   COMBUSTION
                                                         OF OIL OR  GAS
                                                 IN FLUIDIZED BEDS

                                     L. REH
       Lurgi Gesellschaft fuer Chemie und Huettenwesen GmbH
ABSTRACT

  Over  the last 10 years,  Lurgi  has  built
numerous plants for combustion  of liquid
refinery wastes with additional heating by
refinery waste gas.  Bad experiences during
startup  of  the  first  plants  led to  a
fundamental  study  of the combustion  of
liquid and  gaseous  fuels in  a  fluidized bed.
The  principal  results of this  work are dis-
cussed in this paper.
INTRODUCTION

  About 10 years  ago, the Lurgi Company
had to solve a problem  for several German
refineries:  burning  liquid  wastes from  their
waste water treatment plants. Because of the
ever-changing net  heating value  of  these
viscous waste sludges, the fluidized-bed  prin-
ciple was chosen.1 >2 >3
COMBUSTION OF LIQUID REFINERY
WASTES

  The  plants,  built as compact units, had
fluidized-bed furnaces  of approximately 4.5
ID. The furnaces burned, for instance, 1400
Ib/hr oil sludge with oil content of 12 per-
cent, water content of 70 percent and inor-
ganic  matter  content  of  18  percent  The
composition often changed rapidly, so  that
nearly  all oil-contaminated water had to be
"burned" and the heat requirement of com-
bustion had  to be  supplied  by  injected
gaseous fuel.
  Figure 1 shows the principal flow scheme
of such a plant. The incoming air was first
preheated  by the flue gases of the fluidized-
bed  furnace to 900°F  and then  split into
primary and secondary air  streams. The first
was sent through the grate as fluidizing air to
a sand  bed  into  which  the oil sludge was
pumped; the second was blown at a very low
angle of inclination over  the fluidized-bed
surface  level. When necessary, gaseous fuel
was injected over nozzles in the side wall of
the fluidized-bed furnace at a distance not too
high  above the grate. The temperature of the
fluidized bed was held in the'  1500°F range;
however, even with means for distribution of
fuel into different fluidized bed sections, the
temperature  of the furnace  off-gas ranged up
to 1650°F  under  normal  operating condi-
tions, due to afterburning above the sand bed.
The flue gases, completely freed from organic
matter,  passed through a tube-and-shell heat
exchanger. The entrained ash was precipitated
in a following cyclone before the gases left for
final dust cleaning and disposal via the stack.

  Figure 2 shows that, between  net heating
values of 2000 and 2500 Btu/lb of sludge, the
optimal combustion temperature  can be
reached without any addition of gaseous fuel
(which  is  necessary for lower  net heating
values) and without any cooling  (needed for
higher heating values).

  When, for instance,  an   oil sludge  with
nearly  no  heating value was burned, heavy
afterburning  (caused  by gaseous  fuel)  oc-
curred in the upper part of the furnace. The
bed temperature went down and  the flue gas
                                       V-6-1

-------
 temperature up, sometimes leading to damage
 in  the heat exchanger. The same  occurred
 when light volatilizing fuels were burned.4

   This problem  led  us to the study of the
 influence of the radial mixing of gases in  a
 fluidized bed on combustion.
fluidized-bed furnace. The samples were taken
symmetrically to the injection points of fuel
in two vertical planes rectangularly arranged
to each other. By means of balance calcula-
tions, the mean content of unburned carbon
in the fluidizing gas was found in every hori-
zontal layer of the furnace.
 INFLUENCE OF RADIAL GAS MIXING
 ON COMBUSTION

   In cooperation  with the Technical Univer-
 sity,  Karlsruhe,  Germany, first  tests were
 made  with cold models under similar fluidi-
 zation conditions as in an actual fluidized-bed
 furnace. One of the cold models was the same
 size as a  32-in.  diameter fluidized  furnace
 which could be fired with gaseous fuels, diesel
 oil, or heavy oil.5

   For the  conditions existing in  a fluidized
 bed with   grate   tuyeres  with horizontally
 blowing openings, we studied,  for instance,
 the injection length of the jet into a nearly
 two-dimensional fixed bed of sand  of approxi-
 mately  1/20-in.  particle diameter.  Even  at
 high velocities (in the region  of several hun-
 dred feet per second) the jet reached only the
 middle of  the model at a width of approxi-
 mately 1 ft; in the experiment,  the influence
 of the jet  affected  the whole  width of the
 sand bed, 2 ft above the injection point.

   Following the  model  tests,  a fluidized
 furnace was constructed, having a grate from
 seven  horizontally blowing  tuyeres  into
 which gaseous fuel could be introduced either
 by the  central  tuyere  in the  furnace  axis
 through the grate or by injecting nozzles by
 the side; the side nozzles could be used also
 for liquid fuels. Because premixed air/fuel gas
 mixtures,  entering a fluidized  bed by  ade-
 quate grates to avoid back-ignition, give com-
 plete  combustion,  only  those  cases were
 studied in which mixing and  combustion  of
 fuel gases occur inside the bed.

   Concentration   and  temperature  profiles
 were  measured  at different heights of the
   Figure  3 is the temperature  profile for a
test, where town gas,  injected  through the
central tuyere of the grate, was burned  in a
sand bed (1/20-in. mean grain diameter) with
50 percent excess air at a fluidizing velocity
of approximately 5 ft/sec. The profile, typical
for other measurements,  shows clearly  that
the temperature  in the bed adjacent to the
wall (where, preferably,  the air  should be
passing)  is considerably lower  than in  the
centre  of the bed. Some combustion is  still
taking place above the bed level, regulated to
60 in. above the grate.

   Combustion inside  the  bed is strongly de-
pendent on the fluidization velocity, as shown
in Figure 4, in which the content of unburned
carbon for the  aforementioned combustion
conditions is plotted versus the  axial coordi-
nate of the sample point above the grate. At
higher  fluidization velocities,  higher  bed
heights are needed to reach complete combus-
tion because  of  a limited radial mixing rate,
which also means longer mixing ways in the
direction  of  gas flow.  Especially at higher
fluidization velocities,  the content of  un-
burned carbon leaving the fluidized  bed  at a
bed level  60-70  in. above the grate  is rather
high; it reached values of 15-20 percent.

   The concentration profiles, which will be
published5 later, indicate an excess  01" air in
the bed adjacent to the wall; this excess is due
to insufficient radial mixing, and nearly corre-
sponds with the temperature profile.

   When  designing fluidized-bed combustion
systems for gaseous or liquid fuels, the possi-
ble advantage of premixing fuel  and  combus-
tion air should always be kept in mind.
V-6-2

-------
  Different  constructions,  fulfilling the    2. Reh, L. Gaswaerme, 15: No. 8, p. 265-270,
above-mentioned requirements, are already in       1966.
use.
                                            3. Reh, L. Chemie-Ingenieur-Technik. 39: No.
                                               4, p. 165-171, 1967.

                                            4. Reh, L. Chemie-Ingenieur-Technik. 40: No.
BIBLIOGRAPHY                                11, p. 509-515, 1968.

l.Giese,  W.  and  P.  Schwarz.  Brennstoff.    5. Boehm, E. Dissertation, TH Karlsruhe, not
  Waerme, Kraft. 18: p. 227-230, 1966.            yet published, 1970.
                                                                                V-6-3

-------
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                                  FUEL (GAS)
                                900 OF
                             .PRIMARY AIR
                                               TUBE HEAT EXCHANGER
                                                                  AIR
                                                                BLOWER
                                                                 1000 OF
                                                                    CYCLONE
                                                                          ASH
              Figure 1. Combustion of refinery waste sludge in a fluidized bed.
   2500
    2000
 Q
 UJ
 m
 Q
 m
 N
 5  1500
 5
         OPTIMAL COMBUSTION TEMPERATURE
                                    frMMl'X*'
 EC


 1
 LU
 CL
1000
     500


OIL CONTENTS   NET H.V.

     5              0
    10            910
    15           1820
    20           2720
   ASH CONTENTS 10%
                  500
                          1000         1500        2000

                                  NET HEATING VALUE, Btu/lb
   2500
                                                                             3000
                         3500
   Figure 2-  Combustion temperature for aqueous oil sludge in a fluidized bed without air
   preheating.
V-6-4

-------
   90,
C/5
o
Q.
X
<
                     1560°F

                FLUIDIZEI>BED LEVEL
   10 —
              808
               RADIAL POSITION, in.
                                                Figure 3. Temperature profiles, town
                                                gas combustion in sand bed, fluidized-
                                                bed furnace.
                               30        40       50        60

                                     HEIGHT ABOVE GRATE, in.
        Figure 4.  Carbon content leaving fluidized bed at various sampling heights.
                                                                                   V-6-5

-------
SESSION VI:
     Discussion Panel and Summary
PANEL MEMBERS:
     Mr. R. P. Hangebrauck, NAPCA, Chairman
     Dr. D. H. Archer, Westinghouse
     Mr. Shelton Ehrlich, Pope, Evans and Robbins
     Professor D. E. Elliott, University of Aston, England
     Mr. D. B. Henschel, NAPCA
     Dr. C. Y. Wen, University of West Virginia

-------
             1. MINUTES  OF THE PANEL  DISCUSSION
                                        AND SUMMARY SESSION
                              D. B. HENSCHEL

              Division  of Process Control  Engineering

          National Air Pollution Control  Administration
  The  final session  (Session  VI)  of  the
Second  International  Conference   on
Fluidized-Bed Combustion  was  a summary
discussion, led by a panel, covering the major
topics and questions  arising during the Con-
ference.

  Panel members were Dr. D.H. Archer, Mr.
Shelton Ehrlich, Prof. D.E.  Elliott, Mr. R.P.
Hangebrauck (chairman), Mr. D.B. Henschel,
and Prof. C.Y. Wen.

  The subjects discussed were:

    1. Atmospheric versus  pressurized flui-
      dized-bed combustion.
    2. Gas/solids distribution problems.     i
    3. SO2 sorbent regeneration versus non-
      regeneration.
    4. Fluidized-bed boiler design factors.
    5. Methods of carbon burnup for com-
      bustion efficiency.
    6. Corrosion/erosion/deposition  in
      fluidized-bed boilers.
    7. Fluidized-bed  combustion for indus-
      trial-size boilers.
    8. Potential for NOX control.
    9. Gasification versus combustion in the
      fluidized bed.
    10. Fluidized-bed   combustion  for  new
      boilers, versus  an add-on to existing
      units.
  Summaries of the discussions appear below
and are number-keyed to correspond with the
list of subjects, above.

   1. Atmospheric versus pressurized flui-
      dized-bed combustion.

  It was generally agreed that the first genera-
tion  of fluidized-bed  boilers will be atmos-
pheric units. However, many felt that pressur-
ized  operation  will  be  important  in  the
future. It was argued that some of the items
requiring investigation for the development of
pressurized  units can  and should be studied
soon, even  before the technology for neces-
sary  high-inlet-temperature turbines is avail-
able.

   Dr. Archer (Westinghouse) felt that pres-
surization may be desirable from the stand-
point of both economics and air pollution
control. Both Dr. Archer  and Mr. Nicole (Na-
tional Coal  Board) said that  work should be
started as soon as possible to develop the tur-
bines needed for such an application. Profes-
sor  Elliott  (University of Aston) indicated
that much development  work on pressurized
combustors  can be conducted even before
turbine technology is available. Mr. Broadbent
(National Coal Board) stated that gas turbine
development is the major hurdle to be over-
come  for pressurized combustors, and that
                                     VI-1-1

-------
 this development can  proceed concurrently
 with  the  development  of  atmospheric
 fluidized combustors, since turbine work can
 be  conducted without  a fluidized-bed boiler.

  Mr. Walker (Babcock and Wilcox) indicated
 that  1500-2000  MW "conventional"  boilers
 are currently  on the drawing board,  adding
 that if 660 MW fluidized-bed units will not be
 available until 1980, then fluidized-bed com-
 bustion will be  "too  late with too  little."
 Professor  Elliott argued  that  once  the 660
 MW fluidized-bed unit  is developed, scale-up
 to  larger  units  can  proceed  rapidly.  Mr.
 Ehrlich (Pope, Evans and Robbins) and  Mr.
 Broadbent pointed out that the development
 of  fluidized boilers could  be accelerated if
 more money were available.


  2.  Gas/solids distribution problems.

  Professor Elliott stated that gas distribution
 is not a problem, adding that, to study solids
 distribution, a sufficiently large unit must be
 built  and   operated.  Dr. Archer  felt  that
 greater solids  feeding problems (possibly in-
 cluding hot spots and localized reducing con-
 ditions) would be encountered when  coal is
 finely ground. Mr.  Hammons (Esso Research
 and Engineering Co.) indicated that, although
Esso  is currently  studying pulverized  coal
 combustion in a fluidized bed  of lime, he
would not defend the use of coal as  fine as
 that presently being studied (average particle
 size 200 jum)  in  later stages of the  develop-
ment. Mr. Ehrlich argued that possibly Esso
need not use such fine coal in order to achieve
 their  goal of  complete ash elutriation from
 the bed. He indicated that, with one  partic-
 ular  coal  that   Pope,  Evans and  Robbins
 tested, the ashing of 1-inch lumps of this coal
yielded ash that was,  for the most part,
 smaller than 20  mesh. Such  fine ash would
elutriate from  the bed as  desired by Esso.
Possibly not all coals would behave this way.


  3. SC>2 sorbent  regeneration  versus non-
     regeneration.

VI-1-2
   Dr. Gorin (Consolidation Coal Co.) stated
that the purpose of regenerating the sorbent is
not  to recover  the  value of the  resulting
byproduct, but rather to alleviate the need for
the vast quantities of sorbent which would be
necessary  if it is not regenerated. Mr. Ehrlich
also  voiced this feeling: he said that regenera-
tion  and recovery possibly would apply more
to utility boilers than to industrial boilers,
adding that  industrial  units  might well  be
fired with a clean fuel such as the char result-
ing from coal gasification processes.

   4. Fluidized-bed boiler design factors.

   Mr. Seibel (Erie City Energy Division, Zurn
Industries) indicated  that users  of conven-
tional boilers have had difficulties with steam
tubes of the relatively small diameter that is
being considered in  some current  fluidized-
bed boiler designs. He said that, without going
to such  small tube diameters, it would  be
difficult to  fit the required amount of tube
surface  in the bed. Professor Elliott stated
that  nuclear power plants incorporate tubes as
small as 1/2 to 1 inch in diameter—as small as,
or smaller than,  the  diameters proposed for
tubes in fluidized boilers.  Mr.  Seibel agreed,
but  indicated that he felt that  the nuclear
industry has a different situation.

   Dr. Wright (National Coal Board)  said that
the British  have  observed high heat transfer
coefficients  in the bed (up to 100 Btu/hr/ft2/
°F) at high  tube  metal temperatures. He also
indicated  that beds composed of dense fine
particles  gave  the best  results.  Mr. Bishop
(Pope, Evans and Robbins) added that coeffi-
cients  immediately  above  the bed  can  be
fairly high,  much higher than would be ex-
pected in the absence of a bed.


   Mr. Walker questioned the means by which
heat pickup could be controlled in the various
sections of  a  fluidized-bed unit during turn-
down. In  particular, he wondered what would
happen to the water side  of the tubes in one
cell of a multiple-bed boiler if that cell had to
be turned off during turndown. Would  the

-------
water/steam  continue  to be circulated
through the cold tubes in that cell? Or would
each  of  the many  cells have an individual
water/steam valve to shut off the flow to that
cell? The superheater tubes were of particular
interest. Mr. Bishop indicated that, according
to the Pope, Evans and Robbins design, most
of the superheater sections would be  in the
carbon  burnup cell, which would not be
turned off when the boiler is operating; the
limited superheater tubes in any cells which
were  turned off would be run cold.  Mr.
Walker was concerned about the effect that
cold  operation  of some of the superheater
tubes might have on  steam conditions and the
steam turbine.

  Mr. Walker and Mr. Demmy (UGI Corpora-
tion)  were skeptical of the ability of cyclones
to collect the -325 mesh limestone that some
investigators are proposing to inject. Professor
Elliott and Mr. Ehrlich  stated  that their ex-
perience  has  shown their  cyclones  to be
effective in collecting the fine lime.

  5. Methods of carbon burnup for combus-
      tion efficiency.

  The National  Coal Board representatives
indicated that,  operating at low superficial gas
velocities, they have been able to achieve 99
percent combustion efficiency in their 3-foot
square unit  by recycling primary fines. High
dust loadings result. Mr. Bishop reported that,
at the high gas velocities employed in the PER
combustors, the  best combustion efficiency
attainable by primary recycle is 94 percent.
By  utilizing the carbon  burnup cell concept
developed by PER, combustion efficiencies of
98  percent  have been achieved.  The carbon
burnup cell does not  create an increase in dust
loading, since recycling  is not  involved. The
British felt that the high dust loading resulting
from  flyash  recycling  should  not  create
serious  erosion  problems  in  the  overbed
convection passes because the  relatively low
operating  temperature  in  the  fluidized bed
results in softer flyash that is not fused and
sintered as is the ash emitted from higher-
temperature conventional coal-fired units.
  6.  Corrosion /erosion /deposition  in  flui-
      dized-bed boilers.


  The lower  operating  temperature of the
fluidized bed should result in a less abrasive
flyash and, at the lower temperatures that
have  been  studied, less  volatilization of the
alkali components and less tendency for the
ash to become sticky. Under  these circum-
stances,  corrosion, erosion,  and  deposition
should be less serious.

  Mr. Demmy, pointing out  that  lime is a
high-temperature  fluxing agent  for the ash,
stated that glass  could be formed on boiler
interiors  during additive injection. Mr. Ehrlich
agreed that fluxing and clinkering might occur
in a  marginal situation,  when the fluidized
boiler is  being operated at a temperature near
which the  ash in the particular  coal being
burned will be affected in this way. However,
for the low temperatures at which fluidized-
bed boilers can be operated, such effects can
be  avoided with  most  coals.  The National
Coal Board, operating at even  lower tempera-
tures  than are the Americans,  should be even
safer in this regard.
   7. Fluidized-bed combustion for industrial-
     size boilers.


   Mr.  Walker  stated  that there  will be no
market for coal-fired  industrial units unless
the boilers are extremely simple,  fully auto-
mated, and foolproof. Mr. Broadbent agreed,
stating that the market for coal-fired boilers
smaller than 100,000 Ib steam/hr would be so
small that even the development of a control
system for such  small units would not be
justified. Mr. Walker felt that even 100,000
Ib/hr would be too small. One disadvantage of
coal-fired industrial units is the need for fuel
storage space.

   Mr.   Demmy  added  that  environmental
considerations also enter into the  small user's
choice  of  fuel. A small operator does  not
want to be  bothered by local air pollution
                                     VI-1-3

-------
control officials. Therefore, he will opt for gas
or even electricity.


   Professor  Elliott  indicated  that  he  has
studied  gas-fired  fluidized-bed  combustors
using a 100,000 Btu/hr unit.
   8.  Potential for NOx control.
   Dr. Ulmer (Combustion Engineering) stated
that NOX emissions appear to be affected by
the nitrogen content of the fuel, not just by
the reactions of atmospheric N2 and 62- Mr.
Jonke  (Argonne National Laboratory) indi-
cated that when natural gas, containing no
organically  bound  nitrogen, is burned in a
fluidized bed, NOX emissions  are  low, near
thermodynamic  equilibrium levels for the N2
  O2 reaction. When Argonne burned coal in
their fluidized bed, NOX emissions were such
that, if the NOX were entirely from nitrogen
in the  coal, then only about one-third of the
fuel nitrogen would have  been  converted to
NOX. Mr. Ehrlich felt that no one has fully
studied how to reduce NOX emissions from
fluidized-bed combustion.  He suggested  that
one possibility might be to operate the bed at
stoichiometric air and add  secondary air over
the bed;  work  at  PER indicates that NOX
emissions might be cut in half by  this tech-
nique.


   The  question was raised concerning the fate
of the  nitrogen in the fuel when the unit was
operated as a gasifier. One suggestion was that
NH3 might be formed.
   9. Gasification  versus  combustion in the
     fluidized bed.

   It was pointed out  that gasification and
combustion are  different.  Dr. Archer sug-
gested  that both  should be developed. Mr.
Jonke  added that  char  combustion  in  a
separate fluidized-bed unit will probably have
to be tied in with a  coal gasification process.

   Dr. Gorin stated  that  the technology for
gasifying a  weakly caking coal,  with sulfur
removal, is  available today. However, it was
suggested that the current technology might
not be economic.
10.  Fluidized-bed combustion for  new
     boilers, versus an add-on to existing units.

   Mr.  Walker  stated  that existing  conven-
tional units have  been designed for a  given
heat release  rate. He added that modifying the
heat transfer surface in the boiler, to accom-
modate an  add-on fluidized unit,  would be
very expensive. He  suggested  that  a new
fluidized-bed unit  be built, rather than modi-
fying an existing unit by adding on a fluidized
bed. The Conference seemed in general agree-
ment. Mr.  Ehrlich added  that ductwork re-
quired  for the add-on would be prohibitive if
the fluid bed had to be sited far from the
existing boiler due to space limitations.

   Professor Elliott felt that, in  the case of
gasification, an add-on process might be more
feasible,  specifically  the  process being de-
veloped for oil by Esso Research in England.
Minutes submitted by:
   D.B. Henschel
October 22, 1970
VI-1-4
                                          U. B. GOVERNMENT PRINTING OFFICEl IBM	74S4el/41O3

-------
APPENDIX A




    Attendance List

-------
                                                           A.  APPENDIX
                                                  ATTENDANCE LIST

(NOTE: To facilitate their identification, attendees are listed alphabetically together with the
 name of the organization they represent. The complete address of each organization represented
 at the Conference appears at the end of the list of attendees.)
 LIST OF ATTENDEES

 Name
 Archer, Dr. David H.
 Bailie, Dr. Richard C.
 Bishop, Mr. John W.
 Broadbent, Mr. D.H.
 Bryers, Mr. R.W.
 Coates, Mr. N.H.
 Curran, Mr. G.P.
 Demmy, Mr. R.H.
 Diehl, Mr. E.K.
 Eckerd, Mr. J.W.
 Ehrlich, Mr. Shelton
 Elliott, Prof. Douglas E.
 Feldkirchner, Mr. Harlan L.
 Forney, Mr. A.J.
 Glenn, Dr. R.A.
 Glenn, Mr. Roland D.
 Godel, Mr. Albert A.
 Gorin, Dr. Everett
 Gwynn.Mr. J.D.
 Hammons, Mr. Gene A.
 Hangebrauck, Mr. R.P.
 Hanway, Mr. John E. Jr.
 Helfenstine, Mr. Roy J.
 Henschel, Mr. D.B.
 Jarry, Mr. R.L.
 Jonke, Mr. A.A.
 Keairns, Dr. Dale L.
 Lawroski, Dr. Stephen
 Lindquist, Mr. W.E.
 Lundberg, Mr. R.M.
 MacDonald, Mr. B.I.
 Mansfield, Dr. Vaughan
 Marshall, Mr. Keith
 Meyers, Mr. Sheldon
 Moss, Dr. Gerry
 Nicole, Mr. T.C.L.
 Rakes, Mr. S.L.
 Rawdon, Mr. A.H.
Representing
Westinghouse
West Virginia University
Pope, Evans and Robbins
National Coal Board (England)
Foster-Wheeler
Bureau of Mines (Morgantown)
Consolidation Coal
U.G.I.
Bituminous Coal Research
Bureau of Mines (Morgantown)
Pope, Evans and Robbins
University of Aston (England)
Institute of Gas Technology
Bureau of Mines (Pittsburgh)
Bituminous Coal Research
Pope, Evans and Robbins
Societe Anonyme Activit (France)
Consolidation Coal
Balfour-Beatty Power Consultants (England)
Esso Research and Engineering (Linden)
NAPCA (Cincinnati)
Chicago Bridge and Iron
Illinois State Geological Survey
NAPCA (Durham)
Argonne National Laboratory
Argonne National Laboratory
Westinghouse
Argonne National Laboratory
Fuller
Commonwealth Edison
Kennecott Copper
Peabody Coal
Balfour-Beatty Power Consultants (England)
NAPCA (Cincinnati)
Esso Petroleum (England)
National Coal Board (England)
NAPCA (Durham)
Riley Stoker
                                        A-l

-------
 List of Attendees (Cont)
Name

Reh, Dr. Lothar
Riley, Mr. Boyd T. Jr.
Seibel, Mr. R.V.
Shannon, Mr. Larry
Shultz, Mr. F.G.
Skopp, Mr. Alvin
Slikas, Mr. Charles A.
Smith, Mr. J.B.
Spector, Mr. Marshall L.
Squires, Dr. Arthur M.
Svoboda, Mr. Jean J.
Turner,  Mr. P.P.
Ulmer, Dr. Richard C.
Vogel, Mr. G.J.
Walker, Mr. James B. Jr.
Wen. Dr. C.Y.
Wheeler. Mr. Cecil M.
Williams, Dr. D.F.
Wright,  Dr. S.J.
Zoschak, Mr. Robert J.
Representing

Lurgi
Bureau of Solid Waste Management
Erie City Energy Division
Midwest Research Institute
Bureau of Mines (Morgantown)
Esso Research and Engineering (Linden)
American Petroleum Institute
Tennessee Valley  Authority
Air Products and Chemicals
CCNY
Babcock-Atlantique (France)
NAPCA (Durham)
Combustion Engineering
Argonne National Laboratory
Babcock and Wilcox
West Virginia University
Copeland Systems
National  Coal Board (England)
National  Coal Board (England)
Foster-Wheeler
LIST OF ORGANIZATIONS REPRESENTED

Name  (Represented by)

Air Products and Chemicals, Inc.
   (Mr. Spector)

American Petroleum Institute
   (Mr. Slikas)

Argonne National Laboratory
   (Mr. Jarry, Mr. Jonke,
   Dr. Lawroski, Mr. Vogel)

Babcock-Atlantique
   (Mr. Svoboda)

Babcock and Wilcox Co.
   (Mr. Walker)
Address

P.O. Box 538
Allentown, Pa. 18105

1271 Avenue of the Americas
New York, N.Y.  10020

9700 South Cass Ave.
Argonne, 111. 60439
48 Rue La Boetie
Paris, 8e, France

20 South Buren Ave.
Barberton, Ohio 44203
A-2

-------
List of Organizations Represented (Cont)

Name (Represented by)

Balfour-Beatty Power Consultant Group
Engineering and Power Development
Consultants (Mr. Gwynn,
  Mr. Marshall)

Bituminous Coal Research, Inc.
  (Mr. Diehl, Dr. Glenn)

Bureau of Mines, U.S.D.I.
  (Mr. Coates, Mr. Eckerd,
  Mr. Shultz)

  (Mr. Forney)
Bureau of Solid Waste Management
  (Mr. Riley)
Chicago Bridge and Iron Co.
  (Mr. Hanway)

The City College of the City
  University of New York (CCNY)
  (Dr. Squires)

Combustion Engineering, Inc.
  (Dr. Ulmer)

Commonwealth Edison Co.
  (Mr. Lundberg)

Consolidation Coal Co., Inc.
  (Mr. Curran, Dr. Gorin)

Copeland Systems, Inc.
  (Mr. Wheeler)
Erie City Energy Division
  Zurn Industries
  (Mr. Seibel)

Esso Petroleum Co., Ltd.
  (Dr. Moss)
Address

Marlow House
109 Station Road
Sidcup, Kent, England
 350 Hochberg Road
 Monroeville, Pa.  15146

 P.O. Box 880
 Collins Ferry Road
 Morgantown, W.Va. 26505

 4800 Forbes Ave.
 Pittsburgh, Pa. 15213

 Twinbrook Building
 12720Twinbrook Pkwy
 Rockville, Md. 20852

 Route 59
 Plainfield, 111. 60544

 245 West 104th St.
 New York, N.Y.  10025
1000 Prospect Hill Road
Windsor, Conn.  06095

72 West Adams St.
Chicago, 111.  60690

Library, Pa.  15129
120OakbrookMall
Suite 220
Oakbrook, 111. 60521

Erie, Pa. 16503
Esso Research Centre
Abingdon, Berkshire, England
                                                                                 A-3

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List of Organizations Represented (Cont)

Name (Represented by)

Esso Research and Engineering Co.
  (Mr. Hammons, Mr. Skopp)
Foster-Wheeler Corp.
   (Mr. Bryers, Mr. Zoschak)

Fuller Co.
   (Mr. Lindquist)

Illinois State Geological Survey
   (Mr. Helfenstine)

Institute of Gas Technology
   (Mr. Feldkirchner)

Kennecott Copper Corp.
   (Mr. MacDonald)

Lurgi Gesellschaft Fur Chemie und
   Huttenwesen mbH
   (Dr. Reh)

Midwest Research Institute
   (Mr. Shannon)

National Air Pollution Control
   Administration (NAPCA),
   U.S. DHEW, PHS (Mr. Hangebrauck,
     Mr. Meyers)
     (Mr. Henschel, Mr. Rakes,
     Mr. Turner)

National Coal Board
   (Mr. Broadbent, Mr. Nicole,
   Dr. Williams, Dr. Wright)

Peabody Coal Co.
   (Dr. Mansfield)

Pope, Evans and Robbins
   (Mr. Bishop, Mr. Ehrlich,
   Mr. Glenn)
Address

Government Research Laboratory
P.O. Box 8
Linden, N.J. 07036

12 Peach Tree Hill Road
Livingston, N.J. 07039

Research and Development Dept.
Catasauqua. Pa. 18032

Natural Resources Bldg.
Urbana,Ill.  61801

3424 South State St.
Chicago, 111.  60616

161 East 42nd St.
New York, N.Y.  10017

Lurgihaus, Gervinusstrasse 17/19
PostFach 9181
6000 Frankfurt (Main), Germany

425 Volker Blvd.
Kansas City, Mo. 64110

5710WoosterPike
Durham, N.C. 27701

411 West Chapel Hill St.
Cincinnati, Ohio 45227
Hobart House, Grosvenor Place
London S.W. 1, England
301 North Memorial Drive
St. Louis, Mo. 63102

515WytheSt.
Alexandria, Va.  22314
A-4

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list of Organizations Represented (Cont)

Name (Represented by)

 Riley Stoker Corp.
   (Mr. Rawdon)

 Societe Anonyme Activit
   (Mr. Godel)

 Tennessee Valley Authority
   (Mr. Smith)

U.G.I. Corp.
   (Mr. Demmy)

University of Aston at Birmingham
   (Prof. Elliott)

Westinghouse Electric Corp.
   (Dr. Archer, Dr. Keairns)
 West Virginia University
   (Dr. Bailie, Dr. Wen)
Address

90 Neponset St.
Worcester, Mass.  01606

66 Rue d'Auteuil
Paris, XVIe, France

720 Chattanooga Bank Bldg.
Chattanooga, Tenn. 3 7401

247 Wyoming Ave.
Kingston, Pa. 18704

Gosta Green
Birmingham, B4 7ET, England

Research and Development Center
Beulah Road, Churchill Borough
Pittsburgh, Pa.  15235

 Morgantown, W.Va. 26506
                                                                                  A-5

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