United States Industrial Environmental Research EPA-60O/7-79-171
Environmental Protection Laboratory ju|y 1979
Agency Research Triangle Park NC 27711
Summary of Gas Stream
Control Technology
for Major Pollutants
in Raw Industrial Fuel Gas
Interagency
Energy/Environment
R&D Program Report
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RESEARCH REPORTING SERIES
Research reports of the Office of Research and Development, U.S. Environmental
Protection Agency, have been grouped into nine series. These nine broad cate-
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tems. The goal of the Program is to assure the rapid development of domestic
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essary environmental data and control technology. Investigations include analy-
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EPA-600/7-79-171
July 1979
Summary of Gas Stream Control Technology
for Major Pollutants in Raw
Industrial Fuel Gas
by
F. D. Hoffert, W. Y. Soung, and S. E. Stover
Hydrocarbon Research, Inc.
Lawrence Township, New Jersey 08648
Contract No. 68-02-2601
Program Element No. EHE623A
EPA Project Officer: William J. Rhodes
Industrial Environmental Research Laboratory
Office of Energy, Minerals, and Industry
Research Triangle Park, NC 27711
Prepared for
U.S. ENVIRONMENTAL PROTECTION AGENCY
Office of Research and Development
Washington, DC 20460
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ABSTRACT
This report is a summary of coal gasification and clean-up
technology with emphasis on methods for producing a clean industrial
fuel gas as defined by agreement for study purposes. The coal-derived
industrial fuel discussed is one which produces no more than 0.5 Ibs
of S00, O.lf Ibs of NO and 0.1 Ibs of particulates per million Btu
£. - . X
of fuel gas. In general, existing state-of-the-art control technology
will allow these emission guidelines to be met although the end use
for the fuel gas will strongly influence the choice of the pollution
control technology that is used.
Many but not all important factors pertinent to control
technology application were considered. Costs are an example of
an important factor which was not evaluated because the objective
was to first determine appropriate technology that could be applied.
Emissions other than the three major pollutants indicated were given
only a cursory treatment. Nevertheless, a general overall background
of control technology for industrial fuel gas has been covered.
-------
SUMMARY
Industrial fuel gas, classified as a fuel whose combustion
products exhaust directly to the atmosphere, logically fits into emission
guideline levels established by the Clean Air Act of 1971 even though
the S02 emission level was not specifically covered by that regulation.
For this reason, tentative guidelines for emissions of 0.5 Ibs S0?,
0.4 Ibs NO and 0.1 Ibs of particulates per million Btu of coal-
derived gaseous fuel were selected for study purposes. With worsen-
ing fuel shortages, the importance of industrial fuel gas is expected
to increase and its manufacture and control technology should be
examined.
Commercially available gasification and clean-up technology,
as individually described, has promised, even if not yet demonstrated,
the availability of systems to convert coal into fuel gas while con-
trolling at least sulfur, nitrogen and particulates to levels satis-
factory for pipeline gas. The less stringent standards for industrial
gas, as illustrated by the lower percent removal of major pollutants,
are then within reach with some imaginative design because new or
previously rejected technology might be appropriate under such circum-
stances. Since this lowered clean-up requirement makes it possible
to reconsider some old technology as well as modern processes, design
techniques and typical design for the use of iron-oxide processes
have been made and compared to the Stretford process.
End use of the fuel gas must also be considered. If the
fuel is used in a cement kiln, clean-up requirements might be minimal
because SO. would be removed in the cement-making process and parti-
culate control of process exhaust gas would preclude the need for
ill
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particulate removal from fuel gas. Carbon dioxide, which need not
be removed from industrial fuel, may lessen thermal NO formation
due to cooler combustion.
While sulfur removal is covered in some detail, the prob-
lems of nitrogen, tar and particulates also are considered. High
levels of nitrogen compounds may deactivate or cause losses in the
sulfur clean-up processes. Consequently, when nitrogen levels are
high, HCN might be controlled by an HCN guard or polysulfide scrub-
bing and basic nitrogen compounds with acid scrubbing. Use or dis-
posal of tars require more study because of the high sulfur and
nitrogen contents. Combustion of a typical tar from coal gasifica-
tion would probably yield a flue gas exceeding NO emission guide-
lines. By-product tar utilization is, therefore, a fertile field
for future research. Particulate control technology demonstrates
capability to achieve at least 0.0001 grains per SCF while the
industria] fuel guideline has been estimated to be 0.1 grains per
SCF. Technology is available and the proper control systems can
then be selected to meet the particulate level required for speci-
fic industrial fuel use.
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TABLE OF CONTENTS
ABSTRACT ii
SUMMARY iii
Introduction 1
Conclusions and Recommendations 6
Section A - Description of Gasification Systems 8
1. Discussion of Basic Technology 9
2. Koppers-Totzek Process 15
3- Lurgi Process 21
b. Riley-Morgan Process 27
5. WeiIman-Galusha Process 30
6. Wilputte Process 35
7. Winkler Process 39
8. Woodall-Duckham/Gas Integrale Process *»6
Section B - Description of Gas Clean-up Systems on
Operating Gasifier Installations 53
1 . Introduction 5^
2. Clean-up System for the Koppers-Totzek Gasifier 5$
3. Clean-up System for the Lurgi Gasifier 68
k. Wilputte Gas Clean-up System 77
5. Clean-up System on Woodall-Duckham/Gas
Integrale Gasifier Effluent 80
Section C - Comparison of Iron Based Clean-up Processes and
the Stretford Process 86
1. Introduction 87
2. Development of Gas Clean-up Processes 89
3. Basic Chemistry 92
4. Design Basis and Assumptions 97
5. Iron Oxide Sox Purifiers 99
6. Liquid-phase Iron Oxide Processes lO^t
7- Stretford Process 108
Section D - Operational Evaluation of Converter Output
Control Systems 112
1. Typical Clean-up Systems for Industrial Fuel 113
2. Dependency of Clean-up on End Use of Fuel Gas 116
3. Sulfur Emission Control with an Industrial Fuel 118
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k. Effect of Nitrogen Compounds on Chemical
Clean-up Systems 121
5. Tar and Oil By-products 123
6. Reduction of Particulates for Industrial Fuels 128
References 130
Append i ces
A. Generalized Formula for Sulfur Removal From a
Fuel to Meet a Specific Level of SO- Emissions 136
B. Sample Calculations of Clean-up Processes 138
C. Material Balances on Nitrogen and Sulfur
Components for Riley-Morgan Gasification
Systems T*9
VI
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LIST OF TABLES
A-2.1 Gasification Plants Using the K-T Process 19
A-3.1 List of Lurgi Gasifiers 25
A-7.1 Plant List - V/inkler Generators H
A-8.1 WD Two-Stage Coal Gasification Plants ^9
B-2.1 Koppers-Totzek Gasifier Gas Compositions 61
B-2.2 Koppers Coal Gasification - Water Analysis,
Kutahya, Turkey &k
vii
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LIST OF FIGURES
Page
A-l.l Generalized Performance of Gasifiers 12
A-1.2 Generalized Performance of Gasifiers 13
A-2.1 Koppers-Totzek Gasifier 16
A-3.1 Lurgi Pressure Gasifier 23
A-4.1 Riley-Morgan Gas Producer 29
A-5.1 WeiIman-Galusha Fuel Gas Generator 31
A-6.1 Wilputte Gasifier 37
A-7.1 A Winkler Gasifier *4l
A-8.1 Two-Stage Gasification A8
B-l.l Sulfur Removal Requirement for
Industrial Fuel 56
B-2.1 Koppers-Totzek Process - Gasification,
Cooling and Particulate Removal 59
B-2.2 Gas Preparation for Synthesis 60
B-3-1 Lurgi Process « Gasification and Ash
Handling 69
B-3-2 Lurgi Process - Gas Cooling, Shift
Conversion and Gas Liquor Processing 70
B-3-3 Rectisol Gas Clean-up and Methanation 71
B-5.1 WD/GI Process for Cold Desulfurized Fuel
Gas 31
C-5-1 Iron Oxide Design Factor Chart (American
Practice) 101
viii
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LIST OF FIGURES Continued
C-6.1 Flow Diagram of Liquid Iron Oxide Process
for H2S Removal 105
C-7-1 Flow Diagram of Stretford Process 109
D-3-1 Fuel Gas Desulfurization System Schematic
Diagrams 120
ix
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INTRODUCTION
Gas manufactured from coal (General Reference #6) was
first produced in the late 18th century by heating coal in the
absence of air. To supply the necessary heat, additional coal
was burned outside of the vessel. Combustion gases were segre-
gated from the air-deficient interior gas. By 1812, the first
coal-gas company was chartered in London to distribute a product
used for lighting. Four years later, the first U.S. company was
chartered in Baltimore.
Initially, gas with a heating value ranging from
to 560 Btu per cubic foot (depending on the type of coal and
process conditions) was produced by destructive distillation of
coal. Coke ovens, which manufacture coke mainly for steel indus-
try use, produce an off-gas similar in composition. This coke-
oven gas, where available, often supplemented the supply of coal
gas made from plants involving the destructive distillation of
coal. Unfortunately, about 70 percent of the feed coal remains
as a solid residue in these processes and disposal of the residue
was a problem except in the coking- type operation where coke was
the primary product. The solution to the disposal problem for
the carbon-rich residue lead a step beyond distillation to gasi-
fication of the residue.
Coal gasification, which depending on the process may in
the same vessel be preceded by distillation, involves the subsequent
reaction of the solid with air or oxygen and steam. The distillation
step gases which are first released have a high 3tu content because
methane and higher hydrocarbons contained in the coal are among the
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first components to emerge as the coal decomposes. The gasification
step makes a gas with a much lower heating value because the gas
produced is essentially a mixture of carbon monoxide, carbon dioxide
and hydrogen.
The gasification of coal follows two basic paths, using
either air or oxygen supported combustion to supply required heat.
This heat-producing step is necessary to maintain endothermic gasifi-
cation reactions. Gasification with air produces a clean gas of low
(100-250 Btu/cu.ft.) heating value due to a significant concentration
of nitrogen introduced in the air supply. To make a low inert content
gas suitable for synthesis, it is necessary to gasify with oxygen.
This second route produces clean gas of either medium (250-550 Btu/cu.ft.)
or high (950-1000 Btu/cu.ft.) heating value. The latter case requires
additional process steps to reach the higher heat content and, as such,
is not pertinent to the basic objective of this overview which is limited
to the study of industrial fuel.
Industrial gaseous fuels can generally be classified as low
or medium Btu fuels that are burned in equipment designed to exhaust
the products of combustion through a chimney directly to the atmos-
phere. On the other hand, pipeline or "towns-gas" quite often dis-
charge the products of combustion directly into a closed environment
such as a house or factory. In such an environment, the combustion
products are com ing led wi th air and can be breathed by humans. For
example, the typical household gas stove or oven is often poorly
vented or not vented at all. Hence, the products of combustion from
the gaseous fuel are breathed in a diluted form in the home. Conse-
quently, the sulfur content of pipeline gas is severly restricted to
accommodate this probability, and the maximum sulfur level permitted
in pipeline gas is about 4 ppm (1/4 grain per 100 SCF).
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Conversely, industrial fuel gases used to fire kilns,
boilers, heat-treating furnaces, etc., are combusted in equipment
fitted with appropriate stacks for discharge of the products of
combustion directly to the atmosphere. Therefore, it follows that
the pollution effects of the fuel-gas can be assessed on a more
global basis, using criteria more related to ambient concentrations
of pollutants that can be expected at ground levels.
These assessments already have been made and the Clean
Air Act of 1971 specifies that the large-scale burning of fuels
conform to these New Performance Standards for units that started
construction after August 17, 1971.
BOILER EMISSIONS STANDARDS
(HEAT INPUT GREATER THAN 250x1O6 BTU)
Reference: W.C. Wolfe, "Controlling Industrial
Boiler Emissions", PLANT ENGINEERING, 1/2AM
lb/10 Btu Approx. ppm
Emiss ion Fuel Input (dry)
Particulate All 0.1 (0.12 grains
per SCF)
550
520
165
227
525
so
£P
NO
X
Liquid
Solid
Gaseous
Liquid
Sol id
0.8
1.2
0.2
0.3
0.7
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However, the Clean Air Act failed to specify the allowable
emissions of SCL when burning gaseous fuels, presumably because the
gaseous fuels then currently burned were of pipeline quality and
the resultant fuel-gases were innocuous. The absence of a sulfur
specification for gaseous fuels should not be interpreted as re-
quiring zero pollutant emissions from such fuels. A rational analysis
of the fuel substitution problem would suggest that permissable sulfur
levels in gaseous industrial fuels should be close to, but less than,
levels for liquid and solid fuels. Lower sulfur levels in the gas
might be implied by analogy to the NO level for a gas as compared
J\
to liquid and solid fuels.
It is, therefore, possible to infer permissable emissions
levels that would conform in principle with the Clean Air Act criteria.
For the purpose of conducting the study from which the summary of
control technology would be developed and to determine if control
technology required would be significantly reduced from that necessary
for pipeline gas, discussions with the Fuels Process Branch of the
EPA led to the following guideline pollutant levels for coal-derived
industrial fuel gas. Although these emission specifications provide
a basis for study, no regulation by the EPA is implied or recommended
by their use.
Pollutant Maximum Emission Level
Sulfur 0.5 Ibs SO, per 106 Btu
6
Nitrogen 0.*» Ibs NO per 10 Btu
(measured as NO-)
Particulates 0.1 Ibs per 10^ Btu
Note that the suggested nitrogen level is twice the nitro-
gen level in the 1971 Clean Air Act. This increase in the allowable
-------
nitrogen level for industrial fuel gas recognizes the fact that
coal-derived fuel gases may contain nitrogen compounds not normally
found in natural gas. As a consequence, the NO that is formed during
combustion is derived from fuel-bound nitrogen and also from the
thermal fixation of atmospheric nitrogen itself. Thus, the anticipated
NO levels when burning coal-derived industrial fuel gas are higher
/\
than the levels expected when combusting natural gas (which has no
fuel-bound nitrogen).
Therefore, it can be said that industrial gas is a fuel that
can be expected to produce more pollutants than natural (pipeline) gas
when burned but less pollutants than liquid or solid fuels. As the
natural gas shortage worsens, the use of industrial gas to replace
natural gas can be anticipated especially in instances where liquid
or solid fuels are difficult or impossible to burn as substitutes.
Suitable control technology for fuel converters producing
industrial fuel gases must be developed in order to assure an
acceptable environment. This is not to say that acceptable control
technology is not available, but just that previous applications
either lacked control or applied them to more stringent standards
such as synthesis gas. A review of existing gasification processes
that have had a history of successful commercial operation suggests
that operable systems have already been developed to reduce and
control the particulate, sulfur and nitrogen content of coal-derived
fuel gases to meet process requirements. Since these process
requirements are more stringent than the emission levels discussed
previously for industrial fuels, the control technology that they
use should be satisfactory with a little adaptation to the guideline
levels. The purpose, then, of this overview is to discuss operable
gasification systems to cover clean-up methodology, and to suggest
improved methods for raw gas clean-up that will result in an
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industrial fuel gas that meets the proposed guidelines.
A review of many coal gasification systems that are capable
of producing an industrial fuel gas from coal is presented in the
following section of the report. Only those systems which have some
degree of proven operability have been considered. Economics were
not considered in choosing which processes should be included in
this review.
Processes selected for review are as listed below:
Koppers-Totzek
Lurgi
Riley-Morgan
Wellman-Galusha
Wilputte
Winkler
Woodal1-Duckham/Gas Integrals
Air-blown as well as oxygen-blown data are included where
available.
CONCLUSIONS AND RECOMMENDATIONS
° The purpose of this overview was to assess the capability
for producing an industrial fuel gas that will meet the following
environmental guidelines:
Maximum 0.5 lb S02/MM Btu fuel gas
Maximum Q.k lb NO /MM Btu fuel gas
/\
Maximum 0.1 lb particulates/MM Btu fuel gas
Technologies exist today that can meet these standards. However,
the best choice of pollutant control systems is dependent upon the
type of coal gasifier, the coal utilized and the end use of the fuel
gas.
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o Industrial fuel gas standards are less stringent than
those required to produce a pipeline gas and this allows a wider
choice of systems.
o The utilization of tars produced, the disposal of spent
liquors from acid gas removal, and the clean-up of waste water must
be considered when developing the appropriate processing sequence
for manufacture of environmentally acceptable industrial fuel gas.
o Coals containing high nitrogen and sulfur will produce
tars which are too high in nitrogen and sulfur to permit the tars
to be burned directly as fuel and still meet the standards set by the
Clean Air Act of 1971 for liquid fuels. A fertile field for future
development is to establish economic methods for removal of sulfur
and nitrogen from tar.
o Only recently has much consideration been given to environ-
mentally acceptable methods for disposal of spent acid gas removal
liquors. This is an area which requires additional development.
o A significant cost in producing an acceptable industrial
fuel is the clean-up of the waste water leaving the particulates
removal system. Therefore, it is important to use methods for
particulates removal which minimize the quantity of waste water.
-7-
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SECTION A - DESCRIPTION OF GASIFICATION SYSTEMS
1. Discussion of Basic Technology
2. Koppers-Totzek
3. Lurgi
A. R(ley-Morgan
5. WeiIman-Galusha
6. Wilputte
7. Winkler
8. Woodall-Duckham/Gas Integrale
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A. 1 Discussion of Basic Technology
Means for the production of gaseous products from coal
have been known for many years, and commercial processes for coal
gasification are available. Currently, considerable effort is
being expended to develop efficient means for producing gaseous
products from coal suitable for use as an energy source that meets
environmental regulations.
Gasification of coal transforms a cumbersome, inconvenient,
dirty solid fuel into a convenient, clean, gaseous fuel or into
synthetic gas. Some of the heating value of the coal is expended
to accomplish this transformation.
Primary gasification of coal entails the treatment of
coal with air or oxygen and steam to yield a combustible gaseous
product. The product of primary gasification obtained from devolati-
lization of coal and reaction of coal carbon with the gasifying agent
is usually a mixture of H2> CO, C0_, CH., inerts {such as NZ), and
minor amounts of higher hydrocarbons (such as tar, C.H,, C_H. , etc.)
and impurities (for example, H?S, NH, and dust). The product gas
is called a low Btu gas if an air-steam mixture is used directly to
gasify the coal and it contains nitrogen as a major component. Low
Btu gas is suitable for use as an energy source near its point of
generation. Intermediate Btu gas (synthesis gas), which contains
only a small amount of nitrogen, is obtained when an oxygen-steam
mixture is used to gasify the coal. Intermediate Btu gas can be
used either as an energy source or as a synthesis gas for the pro-
duction of chemicals and synthetic liquid and gaseous fuels.
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Three basic reactor concepts have been developed for the
gasification of coal. These are:
1. Fixed-bed or moving-bed (i.e., Lurgi reactor).
2. Fluidized-bed (i.e., Winkler reactor).
3- Suspension or entrained reactors (i.e., Koppers-Totzek)
In a moving-bed reactor, the gasifying medium is passed counter-
current to the coal with ash removal from the bottom and coal
addition at the top. If the velocity of the gasifying medium
and the size of the coal particles are such that the bed behaves
as a fluid, the system is called a fluidized bed. An entrained
system operates with pulverized coal particles carried by the
gasifying medium.
Generally, a gasification reactor can be divided into
three zones: the devolatization, gasification and combustion zones.
In the devolitization zone, if the gasifier has one, coal is dried
and carbonized by hot gas. Most of the volatile matter and moisture
of coal is distilled as CH., tar and aqueous compounds. In the
gasification zone, carbon in the coal reacts with steam and carbon
dioxide to produce carbon monoxide and hydrogen. Both the devola-
tizat ion and gasification zones are endothemnic. The third zone
is a combustion zone in which coal is oxidized to CO and CO..
The heat generated in the combustion zone sustains the chemical
reactions in the gasification and devolatization zones.
The principal components of the product gas are formed by
a combination of the following reactions:
10
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(A) Heat generation reactions
C + 02 = C02 + Heat
C + 1/2 02 = CO + Heat
(B) Heat consumption reactions
C + H20 + Heat = CO + H2
C + C02 + Heat = 2 CO
The compositions of numerous producer gas products have
been examined to gain a better understanding of the performance of
gas generating systems. Data from various types of autothermic
noncyclic gas generators have been plotted in Figures A-1.1 and
A-1.2. Yield patterns on air-blown as well as oxygen-blown systems
have been plotted. The bulk of the data is from coal-based gasi-
fiers but some data is included from oil-based systems for comparison.
As would be expected, the data from coal and oil-based units corre-
late separately because of different C/H ratios in the feed material.
Figure A-1.1 is a plot of the yield of CO + C02 vs. CO + H2- These
data points follow a well-defined yield pattern and correlate well
considering the variety of feed coals and the wide range of system
pressures involved in the correlation. It should be emphasized
that this correlation applies to noncyclic systems only. Cyclic
systems, due to the separation of the blow gas from the run product
gas, follow a different yield pattern. The ratio of CO + C0~ to
CO + H7 indicates the extent of coal combustion versus gasification
and is less than 1.0, as indicated by the slopes of the lines in
11
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ca
ID
ce.
«M
O
CJ
O
o
70 -i
60 -
50 _
40 H
30
20
LEGEND
A R!ley-Morgan
® Wilputte
20
O Winkler
O Woodall-Duckham
A Wellman-Galusha
• Koppers-Totzek
9 Lurgt
B HRI Fluid-Bed (Anthracite)
O Slagging Gasification
B Fixed Bed Generator (Anthracite)
® Flesch-Demag Generator
ffl Ruhrgas Cyclone Gasifier
O Morgantown, W. V. Gasifier
© yaw 0
Oxygen-Blown Gasifier
Air-Blown Gasifier
Q
Coal
HRI Kerosene Gasification
Texaco 051 Gasification
Shell Oil Gasification
—I 1 1 1 1 1
30 *»0 50 60 70 80
CO + H2, % in Raw Gas
FIGURE A-l.l GENERALIZED PERFORMANCE OF GASIFIERS
—r
90
100
Source of Data: Points calculated from compositions presented in many of the Appendix references,
-------
LEGEND
0.7 T
0.6 -
E
c
o
0.5 -
o
-4-
I
O
o o
o o
+ 8
C-J CD
8
o
o
0.3 -
0.2 _
0.1 -
A Ri ley-Morgan
® Wilputte
O Winkler
D Woodal 1-Duckham/Gas Integrale
Wei 1 man-Gal usha
El HRI Fluid-Bed (Anthracite)
G> HRI Kerosene Gasification
-------
Figure A-l.l. This ratio obviously becomes infinitely large when
combustion reactions are complete, forming no carbon monoxide and
hydrogen.
Figure A-1.2 shows a plot of the fraction of carbon that
undergoes combustion to supply heat for the endothermic gasification
reactions, in terms of CO«/(CO, + CO + CH. + C H ) vs. a ratio of
/ L 4 n m
H- and CO. The curves also show the general trends for bituminous
coal, anthracite and oil gasification processes. As would be ex-
pected, the fraction of combusted carbon increases when the H./CO
ratio increases. When gasification reactions proceed to a greater
extent to produce more H. via the steam-carbon reactions, the system
must supply more heat by burning more carbon to sustain the endo-
thermic gasification reactions.
Figures A-l.l and A-1.2 are, therefore, useful for check-
ing and predicting the performance of a gasifier. This understanding
is a good prerequisite for sound environmental study. If the data
from a specific gasifier do not fall in line with the general trends
of the curves shown in the figures, it would indicate that either
the data are not accurate (perhaps due to bad sampling or analysis)
or that the heat leak on the gasifier system is excessive.
-------
A.2 Koppers-Totzek Process
The Koppers-Totzek gasifier is an entrained-bed producer,
operating at high temperature (3300 to 3500°F at the burner discharge
with an exit gas temperature of about 2700°F) and at near atmospheric
(5 to 7 psig) pressure. Operation at high temperature results in
complete gasification of carbon and organic sulfur in the feed and
produces a nonreactive ash in the form of a molten slag. The two-
headed gasifier is a refractory-lined, horizontal cylinder with
opposing coaxial burner heads at each end. Capacity of each gasi-
fier unit is about 400 tons of coal per day for the two-headed gasi-
fier and over 800 tons per day with a four-headed gasifier. Figure
A-2.1 shows a Koppers-Totzek unit.
The coal is pulverized to 70% through 200 mesh using roller
or ball type wind-swept mill and dried to between 2% and 8% moisture
content. Up to a capacity of 150 tons per hour roller mills are
used. Ball mills are applied in the 150 to 250 tons per hour range.
Pulverizers are designed to use combustion gases tempered to 800-900 F
as a drying medium. Gas at this temperature level keeps the coal
particle temperature at 180°F where there is no devolitization or
chemical reaction of the coal particles. As a result the evaporated
coal moisture, after particulate removal, can be discharged as a
vapor to the atmosphere reportedly without detrimental affects to
ai r quality.
The dried, pulverized coal is conveyed with nitrogen to gasifier
service bins and continuously discharged from twin variable-speed screw
feeders into a mixing nozzle where it is entrained in a mixture of oxy-
gen and low pressure steam. The feed mixture is then delivered through
15
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FIGURE A-2.1
KOPPERS TOTZEK GASIFIER
FEED
BIN
FEED WATER
SCREW FEEDER
PIV GEAR
WATER
*H.P. STEAM
\
GAS OUTLET
FEED
BIN
L.P. STEAM
ASH
Reference: toppers Company Brochure
16
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a transfer pipe to the burner head of the gasifier at a speed
higher than flame velocity. Moderate temperature and h-igh burner
velocity prevent the reaction of the coal and oxygen until entry
into the gasification zone. The gasifier designs either have two
or four gasifier heads with opposed inlets. Alignment of in-flow
lines is offset to create a cyclone for effective mixing. The
steam shrouds the high temperature reaction zone and protects burners
and refractories from excessive temperature in addition to its
principal role of supporting the gasification reactions. Product
gas passes through a central, water-cooled vertical gas outlet.
Some ash (30% to 50%), in the form of a molten slag, leaves the
gasifier with the product gas; the balance (50% to 70%) is removed
from the bottom of the gasifier.
The raw gas from the gasifier passes through the waste
heat boiler where high pressure steam (up to 1500 psig) is pro-
duced. After leaving the waste heat boiler, the gas at 350 F is
cleaned and cooled in a high energy scrubbing system. Particulate-
laden water from the gas cleaning and cooling system is piped to a
clarifier. The cool gas leaving the gas cleaning system may contain
sulfur compounds which must be removed to meet gas specifications.
A list of Koppers-Totzek units that have been built is
given in Table A-2.1.
Process Characteristics:
1. Scale of Operation: Commercial
2. Heat Supply: Autothermic
3- Flow: Concurrent
*». Gasifying Media: Steam, Oxygen
17
-------
5. Ash Removal: Molten Slag, Continuous, 30 to 50$
Bottom Ash Removal
6. Pressure: Slight Positive Pressure
7. Temperature: 2700 to 3500°F
8. Product Gas: 300 Btu/SCF
Raw Gas Analysis of Major Components (dry basis) - Volume %:
Western Coal Illinois Coal Eastern Coal
CO
CO
H2
N2
COS
Dried Feed Coal
C
H
N
S
0
Ash
Moisture
58.68
7.04
32.86
1.12
0.28
0.02
Analysis, Weight %:
56.76
4.24
1.01
0.67
13-18
22.14
2.00
55.38
7.04
34.62
1.01
1.83
0.12
61.94
4.36
0.97
4.88
6.73
19.12
2.00
55.90
7.18
35.39
1.14
0.35
0.04
69.88
4.90
1.37
1.08
7.05
13-72
2.00
Koppers-Totzek Reference #2: Farnsworth, J.F.
18
-------
TABLE A-2.1
GASIFICATION PLANTS USING THE K-T PROCESS
Location
Charbonnages de France,
Paris, France
Typpi Oy, Oulu
Finland
Nihon Suiso Kogyo Kaisha, Ltd.
Tokyo, Japan
Nitrogen Works in Puentes de
Garcia Rodriquez, Coruria, Spain
Typpi Oy, Oulu
Finland
S.A. Union Chimique Beige,
Brussels, Belfium
Amoniaco Portuguts S.A.R.L.,
Lisbon, Portugal
Government of the Kingdom of Greece
Ptolemais, Greece
Nitrogen Works in Puentes de
Garcia Rodriquez, CoruTia, Spain
Nitrogen Works of Societe' el Nasr d1
Number of
Gasif ier
Units
1
3
3
3
2
2
2
k
1
1 Engrai s 3
Capacity:
CO + H2
in 24 Hours
75,000-
150,000 Nm3
2,790,000-
5,580,000 SCF
HO, 000 Nm3
5,210,000 SCF
210,000 Nm3
7,820,000 SCF
242,000 Nm3
9,000,000 SCF
140, 000 Nm3
5,210,000 SCF
176,000 Nm3
6,550,000 SCF
169,000 Nm3
6,300,000 SCF
629,000 Nm3
23,450,000 SCF
175,000 Nm3
6,500,000 SCF
778,000 Nm3
Year
of
Order
19^9
1950
195^
195^
1955
1955
1956
1959
1961
1963
et d1Industries Chimiques, Attaka, Suez
United Arabian Republique
Chemical Fertilizer Company Ltd.,
Mae Moh, Lampang, Thailand
Azot Sanayii T.A.S., Ankara,
Kutahya Works, Turkey
28,950,000 SCF
217,000 Nm3 1963
8,070,000 SCF
775,000 Nm3 1966
28,850,000 SCF
19
-------
TABLE A-2.1 (Cont'd)
Number of Capacity: Year
Gasiffer CO + H2 of
Location Units in 24 Hours Order
Chemieanlagen Export-Import G.m.b.H., 2 360,000 Nm3 1966
Berlin fur VEB Germania, Chemieanlagen und 13,400,000 SCF
Apparatebau, Karl-Marx-Stadt, Germany
Kobe Steel Ltd., Zambia, Africa 1 214,320 Nm3 1967
7,980,000 SCF
Nitrogenous Fertilizers 1 165,000 Nm3 1969
Industry S.A., Athens, 6,150,000 SCF
Ptolemais, Greece
The Fertilizer Corporation 4 2,000,000 Nm3 1969
of India Ltd., New Delhi, 0 of them 74,450,000 SCF
Ramagundam Plant, India as stand-by)
The Fertilizer Corporation 4 2,000,000 Nm3 1970
of India Ltd., New Delhi, 0 of them 74,450,000 SCF
Talcher Plant, India as stand-by)
Nitrogenous Fertilizers 1 242,000 Nm3 1970
Industry S.A., Athens 9,009,000 SCF
Nitrogenous Fertilizers Plant
Ptolemais, Greece
The Fertilizer Corporation of India Ltd., 4 2,000,000 Nm3 1972
New Delhi, Korba Plant, India 0 of them 74,450,000 SCF
as stand-by)
AE & Cl Ltd., Johannesburg, 6 2,150,000 Nm3 1972
Modderfonteln Plant, South Africa 80,025,000 SCF
Indeco Chemicals Ltd., 1 220,800 Nm3 1974
Lusaka, Kafue Works, Zambia 8,220,000 SCF
Indeco Chemicals Ltd., 2 441,600 Nm3 1975
Lasaka, Kafue Works, Zambia 16,440,000 SCF
20
-------
A.3 Lurgi Process
The Lurgi producer is a moving-bed, intermediate pressure
gasifier with counter-current flow of coal and gas. Dry ash is
removed from the bottom and gas exits from near the top.
Coal of 1/4 to 1-3A inch is fed to the top of the gasi-
fier through a pressurized lock hopper. Fine coal particles are
not acceptable in the feed; large particles are needed to permit
sufficient gas passage. The coal, ideally nonagglomerating, flows
downward through the moving-bed gasifier. A rabble arm levels the
coal on the surface of the bed.
Steam and air or oxygen are introduced at the bottom of
the reactor through a revolving grate; as these gases rise through
the reaction zone, they react with residual carbon to produce heat
necessary to carry out the gasification reactions. To insure com-
plete gasification, a sufficiently high temperature must be main-
tained by control of the steam/oxygen ratio. This temperature
must be below the ash fusion point but high enough to provide the
heat necessary in the endothermic gasification zone.
The moving bed has a layered temperature profile, which
increases as the coal proceeds downward until reaching approximately
2000°F in the bottom combustion zone. At the top of the bed, feed
coal is preheated, dried, and heated up to start devolati1ization.
Carbonization products, such as tar, oils, naphtha, light hydrocarbons,
phenols and ammonia, together with a mixture of hydrogen and carbon
monoxide are produced starting at about 900°F to 950°F. From a
temperature of 1150°F to 1400°F onward, devolati1ization is accom-
panied by gasification of the resulting char. The kinetics of the
21
-------
devolatiIization and gasification reactions is related to the re-
activity of the coal and, thus, the minimum temperature at which
the reaction will proceed. The approximate final reaction tempera-
tures at which the reaction rates of various coals approach zero
are:
Lignite 1200°F
Sub-bituminous 1350 F
Semi-anthracite 1450°F
Coke 1550°F
To achieve a good performance of the reactions, the minimum required
residence time of a coal grain at the desired temperature level of
H»00°F to 1600°F is about one hour.
Gas leaving the bed contains, in addition to the carboni-
zation products, coal and ash particulates plus some hydrogen sul-
fide and carbony 1 sulfides. The exiting gas temperature is at 700 F
to 1100 F depending on the type of coal.
The relatively uncomplicated Lurgi gasifier design shown
in Figure A-3.1 utilizes a double-walled vessel with hot water or
boiling water between the walls. Gasifiers of recent design are not
internally insulated. The Lurgi design incorporates a mechanically-
driven grate at the bottom of the gasifier. The grate supports the
fuel bed, removes the ash into a lock hopper, and allows introduction
of steam and air to the gasifier. The ash> a granular, low-carbon
material, is discharged from the pressure reactor through an ash
lock chamber. The literature indicates the physical limit to the
diameter of Lurgi gasifiers has been established by the need to
22
-------
o
FEED COAL
RECYCLE TAR
DRIVE
GRATE
DRIVE &
STEAM
OXYGEN
\ :
SCRUBBING
COOLER
GAS
WATER JACKET
Reference:
Paul F. H. Rudolph, "The Lurgi
Process The Route to SNG from
Coal", Ath Synthetic Pipeline Gas
Symposium, Chicago, Oct. 30 & 31 ,
1972.
LURGI PRESSURE GASIFIER
23
FIGURE A-3.1
-------
maintain good distribution of gas through the downward moving bed
of coal. When the size of the reactor becomes too great, perhaps
15 feet or more, it would be difficult to maintain good distribu-
tion of gas across that large a cross-sectional area. Capacities
up to 570 x 10 Btu of coal per hour are reported for each Lurgi
gasifier uni t.
A list of Lurgi gasifiers that have been built is given in
Table A-3-1- Some of these units are no longer operating.
Process Characteristics:
1. Scale of Operation: Commercial
2. Heat Supply: Autothermic with externally heated
steam
3. Flow: Counter-current
4. Gasifying Media: Steam, air or oxygen
5- Ash Removal: Dry-ash, continuous
6. Pressure: 10 to 30 atm.
7. Temperature: 1150-1600 F in gasification zone
8. Product Gas: 310 Btu/SCF (oxygen-blown) - MO-180
Btu/SCF (air blown)
9. Heat Recovery: Side wall
-------
A.A Riley-Morgan Process
The Riley-Morgan gas producer is a low-pressure (up to
kO inches of water gauge), stirred moving-bed gasifier. The new
design is an improved version of the old Morgan gas producers of
which over 9,000 were sold through the 19^0's. Improvements included
replacing castings with weldments, an outer gas and dust tight casing
for personnel protection, water seals for higher operating pressure,
increased volume to handle the swelling tendency of certain coals,
a water-cooled agitator to handle caking coals, and automated con-
trols to reduce the number of operating personnel.
Coal of I/A to 1-1/2 inch is fed to the top of the gasi-
fier. Both noncaking and caking coals can be gasified because the
deep bed agitator can break up agglomerated masses.
The fuel bed is supported on an ash bed which, in turn,
is supported on a rotating ash pan. The pan, barrel and charge
all rotate together. The fuel bed is smoothed out by leveler arms
and moves downward as the gasification of fuel proceeds. As the
ash level becomes too high, ash is moved by a helical plow from
the bottom of the gasifier.
Steam and air or oxygen are fed to the bottom of the
gasifier. Air is brought in without a fan by an injection system
which mixes and inducts air by steam flow in Venturis.
Gases leave the top of the reactor, pass through a cyclone
to remove dust, and are cooled first in an air condenser and then in
a water condenser. Cooler condensate is separated from the gas which
27
-------
proceeds on to further processing as required by downstream gas
utili zation.
A Ri ley-Morgan unit is shown in Figure A-**.!.
A sulfur and a nitrogen balance for the Riley-Morgan
system is given in Appendix C.
Process Characteristics:
1. Scale of Operation: Two years full scale laboratory
test
2. Heat Supply: Autothermic
3. Flow: Counter-current
4. Gasifying Media: Steam, air or oxygen (proposed)
5. Ash Removal: Dry ash, intermittent
6. Pressure Atmospheric
7. Product Gas: 305 Btu/SCF (oxygen-blown) - 160 Btu/SCF
(air-blown)
Raw Gas Analyses of Major Components (dry basis and sulfur-free basis)
Volume %:
Ai r-Blown Oxygen-Blown
CO
H2
Ch\
C02
cnHm
N2
26.0
18.0
1.4
3.8
0.2
50.0
41.2
38.9
2.8
15.9
0.7
0.5
Approximate Analysis of Feed Coal (U.S. Eastern Bituminous Coal)
Moisture 1.2%
Volatile Matter 34.4%
Fixed Carbon 42.7*
Ash 15.7*
Riley-Morgan Reference #3: Rawdon, A.M., et al.
28
-------
FIGURE A-A.l
RiLEY-MORGAN CAS PRODUCER
GAS OUTLET
FIXED COVER
LEVELER BOX
COAL INLET
WATER COOLED LEVELERS
WATER COOLED AGITATOR
STATIONARY OUTER CASING
R6fRACTORY LINING
WATER SEAL
BLAST HOOD
ROTATING BARREL
BARBEL SUPPORT SPIDER
ROTATING ASH PAN
ASH PLOW
ASH HOPPER
EXTERNAL SUPPORT WHEELS
WATER SEAL
DRIVE GEAR
STRUCTURAL SUPPORTS
AIR AND STEAM INLET
Reference: Riley-Morgan Brochure
29
-------
A. 5 WeiIman-Galusha Process
The WeiIman-Galusha producer is a single-stage, low-temperature,
low-pressure gasifier with a fixed bed through which coal moves slowly
downward toward a revolving eccentric grate. Gas flows upward counter-
current to the flow of coal. Sizes of the water jacketed automatic
producers with revolving grates range from k2 inches to 10 feet in
inner diameter. Capacities of these units are, respectively, 192 to
1600 Ibs per hour (*»0 million Btu/hr) using 3/16 to 5/16 inch anthra-
cite. With an agitator, the capacity of the 10 feet diameter unit
is 2000 Ibs per hour for anthracite and 7000 Ibs per hour on 1-1/4
to 2 inch bituminous coal. A high pressure unit (300 psi) has been
designed and built at the Bureau of Mines Research Center in Morgan-
town, West Virginia. Figure A-5.1 illustrates a typical unit.
A two-compartment fuel bin, at the top of the producer,
feeds crushed and dried coal to the downward moving bed. A slowly
revolving horizontal arm, which spirals vertically below the surface
of the fuel bed, retards channeling and maintains a uniform bed.
The producer can gasify anthracite and coke but, when bituminous
coal is used, an agitator is required to avoid channeling and caking
(agglomerating) of coal particles. The bed is supported by a slowly
revolving eccentric grate through which dry ash is continuously
ejected into an ash hopper.
Gas flow is upward through the gasifier. Steam, generated
in the water jacket surrounding the gasifier, and air or oxygen are
introduced through the revolving grate. Depending on the type of
gas required, CO^ can be substituted for the steam. This procedure
would avoid formation of hydrogen from the steam in order to make
30
-------
FIGURE A-5.1
WATER JACKET
AGITATOR
COMBUSTION
ZONE
TYPICAL RUILDirir,
AND FUEL CLF.VATOR
OUTLINE
FUEL DIM
VALVES CLOSED
LOCK HOPPER
.HATER SEAL
AMD DUST
COLLECTOR
•GASIFICATION
ZONE
WELLMAN-GALUSHA FUEL, GAS GENERATOR.
Reference: WeiIman-Galusha Brochure
-------
a gas with improved burning characteristics. Raw gas containing
particulates, tars and oils leaves the top of the producer at a
temperature between 1000 F and 1250 F depending on the type of
coal.
After the reactor, the gas may pass through a waste heat
recovery section. Ash, carried over by the gas, and tar are removed
by scrubbing. The cooled gas is then compressed and further pro-
cessed if pipeline gas is to be made.
Process Characteristics:
1. Scale of Operation: Commercial
2. Heat Supply: Autothermic with externally heated
steam
3- Flow: Counter-current
4. Gasifying Media: Steam, oxygen or air
5. Ash Removal: Dry-ash, continuous
6. Pressure: Atmospheric
7. Temperature: 1000° to 1250°F (normal off-take
temperature of the gas)
8. Product Gas: 260 Btu/SCF (oxygen-blown) - 160 Btu/SCF
(ai r-blown)
9. Heat Recovery: Sidewall
32
-------
Raw Gas Analysis of Major Components (Unspecified Coal)
Mol
Component High-Btu Low-Btu
CO 29-6 26.0
COz 12.3 3-0
H2 30.3 13.9
H20 25.3 8.3
CHJL 0.7 2.5
N2 1.1 1*5.6
02 0.1
H2S/COS 0.6 0.7
Total 100.0 100.0
Higher Heating Value
(dry basis) , Btu/SCF 263
General Reference #1
33
-------
Raw Gas Analysis of Major Components (dry and sulfur-free basis)
Vo1ume %:
Bituminous A
Coke W. Va.
31.0 15.0
0.6
54.1 28.6
11.3 3.4
0.4 2.7
2.6 50.3
Anthracite
Feed
H2
°2
CO
C02
CH/,
N2
Coal :
PA
41.0
0.1
40.0
16.5
0.9
1.5
Proximate Analysis, W%
Moisture
Ash
Volatile Matter
Fixed Carbon
6.4
10.2
4.5
78.9
Ultimate Analysis, W&
C
H
N
0
S
Ash
78.7
2.4
0.7
7.4
0.6
10.2
5.0 3-7
11.9 3-5
36.3
83.1 56.5
84.1
1.8
0.8
1.8
0.6
10.9
WeiIman-Galusha References #3 and #7-
34
-------
A.6 Wilputte Process
The Wilputte producer is a moving-bed gasifier, operating
at atmospheric pressure. The producer is equipped with a rotating
grate and ash pan assembly which rests on roller supports. A sta-
tionary ash plow removes the ash from the ash pan and discharges the
ash into an ash trough. The producer is supported above the rotating
grate and ash pan assembly by columns from the floor and is sealed
by water in the ash pan. The producer can gasify either bituminous
or lignite coal with air or oxygen. A rotating rabble arm (agitator)
is used to mix and break up the agglomerates when a caking coal is
gas i f ied.
The crushed coal of less than k inches is fed through the
top of the gasifier. Moist air or oxygen flows upward through a
rotating grate. The fresh coal is pyrolyzed to expel volatile hy-
drocarbon products at the upper portion of the fuel bed, and the
carbon residue from pyrolyzed coal is burned and gasified with air
and steam to carbon dioxide, carbon monoxide and hydrogen. The
product gas leaves the producer at about 1100-1200 F.
The raw product gas contains coal tar, light oil, hydrogen
sulfide and particulates. The gas is passed through a cyclonic dust
collector to a waste heat boiler that lowers the temperautre of the
gas stream to about 800 F. The gas is then quenched with an aqueous
liquor and further scrubbed with the liquor in a packed tower to
remove light oil, tar and particulates. When operating on a low-
sulfur coal, the scrubbed gas can be used at this point as a fuel
without additional processing. If high-sulfur coal is used, or
complete sulfur removal is required, the gas stream is further
35
-------
purified by a secondary scrubber and by an electrostatic precipitator
before entering the sulfur removal system. The Holmes-Stretford
process may be used for the removal of hydrogen sulfide. A system
to remove HCN from the gas stream prior to entry into the Holmes-
Stretford unit is not required because the HCN content of the gas is
so low that fouling of the absorbent solution with thiocyanates is
not a serious problem.
The heating value of the producer gas is affected by the
volatile content of the feed material. Bituminous coal and lignites
will produce a fuel gas with a heating value of 160-170 Btu/SCF.
The HHV of gas from anthracite or coke will be somewhat lower.
Process Characteristics:
1. Scale of Operation - Commercial
2. Heat Supply - Autothermic
3. Flow - Counter-current
k. Gasifying Media - Moist air or oxygen
5. Ash Removal - Dry ash, continuous
6. Pressure - Atmospheric
7. Product Gas - 165 Btu/SCF (air-blown) - 300 Btu/SCF
(oxygen-blown)
8. Heat Recovery - Sidewall
A Wilputte gasifier is illustrated in Figure A-6.1.
-------
FIGURE A-6.1
WILPUTTE GASIFIER
COAL /-etc
WATCH COOLtO
AOITATO* 4
COVA/£C7M/6 MM
SHELL
SU/fOKT-
CQLUUH
Reference: Wilputte Co. Brochure
37
-------
Purified Gas
Composition, V % (l) Air-blown
H2
CHI,
N2
0
CO
co0
16.6
3.6
51.0
0.2
22.7
5.9
Feed Coal: U.S. Eastern Bituminous Coal.
NOTE: 1) After Purification by Stretford Process
Wilputte Reference #1
38
-------
A.7 Winkler Process
The Winkler producer is a f1uidized-bed gasifier, operating
at low (up to 3 psig) pressure. The Winkler generator was initially
designed to make producer gas from coal with air and steam gasification.
The gas was used as an industrial fuel for both direct combustion and
use in stationary internal combustion engines. A proposed recent ver-
sion of the Winkler gasifier has been designed to operate at a pressure
of about 125 psi. Operation at this proposed pressure level is said
to improve the economics of both the oxygen-blown and the air-blown
models of this system.
Coal is normally crushed in the range of 0 to 3/8-inch and
dried. Past operating experience has indicated that the coal feed
need be dried only if surface moisture is present and the coal can-
not be handled without plugging screens, conveyors, etc. Generally,
coals with moisture contents of up to 13% can be handled and gasi-
fied without drying.
Coal enters the producer through a variable-speed screw
feeder. The screws, in addition to providing control on the coal
feed rate, serve to seal the producer preventing steam from wetting
the coal feed and blocking the feed line.
The primary supply of steam and oxygen (or air) enters the
bottom of the gasifier. Coal reacts with oxygen and steam to pro-
duce a gas rich in carbon monoxide and hydrogen. A secondary blast
of steam and oxygen or air above the fluidized bed converts unreacted
carbon in the gas-entrained particles and raises the bed to maximum
temperature.
39
-------
Gasification reactions in the Winkler gasifier are primarily
a combination of combustion and water-gas reaction at temperatures of
1500°F to 2000°F depending on the type of coal. The high temperature
reacts all tars and heavy hydrocarbons. Before leaving the gasifier,
the gas is cooled by a radiant heat boiler section to prevent ash
particles from melting and forming deposits in the exit duct.
As a result of high fluidization gas velocity, the ash
particles are segregated according to size and specific gravity.
The larger, heavier ash particles fall down through the fluidized
bed and pass into the ash discharge unit at the bottom of the gasi-
fier while the lighter particles are carried up and out of the fuel
bed by the product gas. Experience has shown that approximately
30% of the ash leaves at the bottom while 70% is carried overhead.
An auxiliary unit provides a hot bed of coal for initial start-up.
Raw gas leaving the gasifier is passed through a further
waste heat recovery section. Fly ash is removed by cyclones followed
by a wet scrubber and finally, if appropriate for the application, an
electrostatic precipitator. The gas may then be compressed and
shifted if continuing on to pipeline gas. The gas coming from the
shift converter is purified, methanated, dehydrated, and compressed
further if necessary to produce pipeline quality.
The ash removed in the dry state is conveyed pneumatically
to an ash bunker. That which is removed wet is recovered as a slurry
in a settler and then mixed with the warm dry ash where the contained
water cools the ash and wets it to prevent dusting problems during
ultimate disposal.
A Winkler unit is shown in Figure A-7.1.
-------
FIGURE A-/.1
A WINKLER GASIFIER
PURGE £ INERT GAS LINES
WITUT
FUEL BURNER
TO STACK
WATER COOLED
SHAFT
/M/VVWVW
RATCHET DRIVE
STEAM
GASIFIER
GAS TO DUST
V-^-COLLECTOR £
WASTE HEAT BOILER
o o
STEEL SHELL
REFRACTORY LINING
WATER JACKETED
SCREW CONVEYOR
SCRAPER FOR ASH
REMOVAL
OXYGEN OR
ENRICHED
AIR
RATCHET DRIVE
WATER COOLED
SHAFT
ASH
RECEIVER
WATER
JACKETED
SCREW CONVEYOR
Reference: Davy Powergas, Inc.
-------
A list of Winkler gasifiers that have been built is given
in Table A-7.1- Not all of the units are still operating.
Process Characteristics:
1. Scale of Operation: Commercial
2. Heat Supply: Autothermic with externally heated steam
3- Flow: Counter-current
4. Gasifying Media: Steam, oxygen or air
5. Ash Removal: Dry-ash, continuous
6. Pressure: Atmospheric (up to 8 psig)
7. Temperature: 1500-2000°F
8. Product Gas: 270 Btu/SCF (oxygen-blown, German Coal)
105 Btu/SCF (air-blown, German Coal)
9. Heat Recovery: External
Raw Gas Analysis of Major and Sulfur Components (Coal Unspecified):
MoU
Component
H2
CO
C02
H20
CHl,
N2
H2S
COS
Total 100.0 100.0
Higher Heating Value
(dry basis), Btu/SCF 275 118
General Reference #1
High-Btu
32.2
25.7
15.8
23.1
2.4
0.8
2500 ppm
400 ppm
Low-Btu
11.7
19.0
6.2
11.5
0.5
51.1
1300 ppm
200 ppm
-------
Raw Gas Analysis of Major Components (dry basis) - Volume %;
Oxygen-Blown
H2 38.5
CO 35.3
C02 21.8
CHij 1.8
N2 1.1
H2S 1.5
Ultimate Analysis of Feed Coal (German Brown Coal):
W %
Moisture 8.0
Ash 15.4
C 54.6
H 4.1
N 0.6
S 3-3
0 14.0
Winkler Reference #4.
43
-------
TABLE A-7-1
PLANT LIST - WINKLER GENERATORS
Capacity Per Generator
Normal Maximum
Plant
No.
1
2
3
k
5
6
7
8
9
Plant
Leuna-Werk
Leuna, Germany
Braunkohle-Benzin AG
Boh 1 en, Germany
Braunkohle-Benzin AG
Magdeburg, Germany
Yahagi
Japan
Braunkohle-Benzin AG
Zeitz, Germany
Da i -N i hony 1 nzo-H i ryo
Japan
Nippon Tar
Japan
Toyo-Koatsu
Japan
Sudetenlandische
Year
1926-
1930
1936
1936
1937
1938
1938
1938
1939
Product
Fuel Gas
Water Gas
Water Gas
Water Gas
Water Gas
Water Gas
Synthesis Gas
Water Gas
Synthesis Gas
1000
NM3/hr
60
30
27.6
27.6
8.75
22.5
14
14
15
1000 1000 1000
SCFH NM3/hr SCFH
2240 100 3730)
1120 50 1870)
1030 30 1120
1030 33 1230
330
840
520
520
560 20 750
No.
Ger
3
3
1
3
2
2
2
Treibstoffwerke
Brux, Czechoslovakia
*10 Fabrika Azotnih
Jendinjenja
Gorazde, Yugoslavia
11 Calvo Sotelo
Puertollano, Spain
12 Union Rheinische
Braunkohlen
Wesseling, Germany
1943 Water Gas
1953 Synthesis Gas
1954 Water Gas
27.6 1030
5 190
9-5 350
1956 Synthesis Gas 12
30 1120 5
j
1
450 17 630 1
(Continued)
44
-------
TABLE A-7-1 (Cont'd)
Plant
No.
13
*15
*16
Plant
Calvo Sotelo
Puertollano, Spain
Azot Sanyyi i TAS
Kutahya, Turkey
Neyveli Li gni te
Corporation
Madras, India
Union Reinische
Braunkohlen
Wesseling, Germany
Capacity Per Generator
Normal Maximum
1000 1000 1000 1000 No.
Year Product NM3/hr SCFH NM3/hr SCFH Gen.
I960 Synthesis Gas 12
1957 Synthesis Gas 9-5 350
1959 Synthesis Gas 12
1959 Synthesis Gas k].6 1550
^450
1
18 670 2
17 630 1
* Presently operating
-------
A.8 Woodall-Duckham/Gas Integrale Process
The Woodall-Duckham/Gas Integrale process is a two-stage,
moving-bed gas producer. The producer consists of a distillation
retort surmounting a gasification shell. The producer is top fed,
operating at atmospheric pressure with ash continuously withdrawn
through a grate at the bottom of the producer. The lower portion
of the producer is water jacketed and the upper portion is made of
hollow refractory tile.
Coal of 3A to 2-3/4 inch is fed to the top of the producer
and descends in the distillation retort (first stage) through gradually
Increasing temperature zones. All volatile matter is expelled until
only carbon and ash (distillation coke) remain to enter the gasifica-
tion zone (second stage).
In the gasification zone air and steam enter through a
rotating grate at the base of the producer and the carbon in the
distillation coke is gasified. A part of the hot producer gas
(called clear gas) travels upward through the fresh coal in the
upper portion of the producer while the sensible heat of the gas is
used to distill off the volatile matter of the fresh coal. The
producer gas mixes with distillation vapors to leave the retort as
mixed gas at temperatures of 210 F to 300°F. The mixed gas consists
of 60-75% water gas (CO & Hj,), with the balance being coal distilla-
tion gas, tar, light oils and unreacted steam. The remaining part
of the clear gas (a mixture of CO, H_, CO-, CH. and N_ and completely
free from oil and tar) leaves the gasifier through a blast gas out-
let at a temperature of 800 F to 1000 F. The producer cannot handle
a highly caking coal.
-------
Refer to Figure A-8.1 for a WD/GI illustration.
A list of WD/GI gasifiers is given in Table A-8.1.
Process Characteristics:
1. Scale of Operation: Commercial
2. Heat Supply: Autothermic
3. Flow: Counter-current
4. Gasifying Media: Steam, oxygen or air
5. Ash Removal: Dry-ash, continuous
6. Pressure: Atmospheric
7. Product Gas: 340 Btu/SCF (oxygen-block) - 175 Btu/SCF
(ai r-blown)
8. Heat Recovery: Sidewall
Raw Gas Composition of Major Components (dry and sulfur-free)
- Vo 1 ume %:
Air-Blown
H2
02
CO
CO,
CHjJ
C H
M11 ^
o
Cont i nuous
17.0
Nil
28.3
4.5
3-0
47.2
Cycl ic
50-56
0.1-0.3
26-31
6-10
5-8
2-6
Oxygen-Bl
38.4
0.0
37.5
18.0
3.5
0.4
2.2
Proximate Analysis of Coal, W%
Moisture 11.0
Volatile Matter 35.4
Fixed Carbon 44.9
Ash 8.7
Woodall Duckham/Gas Integrale References # 2 £ # 3
47
-------
drying zone
low temperature
distillation zone
gasification zone
ash
FIGURE A-8.1
Two Stage Gasification
Reference: Wooda!1-Duckham Brochure
-------
TABLE A-8.1
WD TWO-STAGE COAL GASIFICATION PLANTS
The WD Two-Stage process has been in commercial use as a cyclic medium-Btu gas
generator since the 1920's and as a continuous air-blown fuel gas generator since
19^2. Over 100 units have been built in Europe alone. The following lists are
confined to plants built since 1946.
A. Industrial Fuel Gas Plants
These units all employ a continuous air/steam blast as gasffication agent.
Different numbers of standard sTze gasffiers (there are several standard
sizes) are used to obtain the required output. A variety of coal types
are used, as indicated. Mixtures of different coals have also been
successfully employed. Gas purffication is included on some plants.
Client and Location
Austrian-American Magnesia Co.,
Radenthein, Austria
VOEST Steelworks, Linz, Austria
Cellulose & Paperworks,
Frantschach, Austria
Reforming Plant, Wels, Austria
Gas Utility Co.,Graz, Austria
Coke Plant, Strasbourg, France
Coke Plant, Drocourt, France
Steelworks, Audincourt, France
Steelworks, Firminy, France
English Steel Corp.,
Sheffield, England
Weldless Steel Tube Co.,
Wednesfield, England
Coal Type
Bituminous
Lignite/Bituminous
Ligni te
Ligni te
Ligni te
B iturn!nous
B i t urn i nous /Coke
Bi turn!nous
Bi turn!nous
Bituminous
Bituminous
No. of Units
3
1
1
1
3
3
1
1
7
(Cont inued)
-------
TABLE A-8.1 (Cont'd)
Client and Location
Ziar Aluminum Works,
Czechoslovakia
Chomutov Tube Works
Czechoslovakia
Istanbul Gas Utility Co.,
Tu rkey
Australian Consolidate
Industries Ltd., Sydney,
Australia
Melbourne Gas Works
Melbourne, Australia
Elgin Fireclay Ltd.,
Springs, South Africa
Vaal Potteries Ltd.,
Meyerton, South Africa
Union Steel Corp.,
Johannesburg, South Africa
Stewards & Lloyds Steelworks
South Africa
Masonite, Escault,
South Africa
SAAPI, Mandini, South Africa
Rand Water Board,
Vereeniging, South Africa
Drlefontein, Carltonville,
South Africa
Vereeniging Refractories
South Africa
Coal Type
Bituminous
Lignite
Lignite
Bituminous
Bituminous/Brown Coal
Bituminous
Bituminous
Bituminous
Bituminous
Bituminous
Bituminous
Bituminous
Bituminous
Bituminous
No. of Units
3
3
2
1
2
2
(Continued)
50
-------
TABLE A-8.1 (Cont'd)
B. Publ ic Utility Gas Plants
The following plants employ the same type of gasifiers, Including coal
and ash handling systems, as the Industrial Fuel Gas Plants, but operate
in a cyclic mode so as to produce a gas with a very low nitrogen content.
The calorific value of the product gas is from 330 to 500 Btu/cu. ft.,
depending on the extent of enrichment, e.g. by carburation with distillate
or residual oil, or by enrichment with LPG.
Location of Utility No. of Units
St. Poelten, Austria 2
Naples, Italy 2
Rome, Italy 5
Trieste, Italy 2
Milan, Italy 2
LaSpezia, Italy 1
Como, Italy 1
Genoa, Italy *»
Vierzon, France 2
Dijon, France 2
Kensal Green, England 1
Gloucester, England 1
Ulm, West Germany 2
Freiburg, West Germany 2
Zagabria, Yogoslavia '
Prague, Czechoslovakia 6
Warsaw, Poland 3
Thorn, Poland 2
Tokyo, Japan 5
Posen, Poland 3
51 (Continued)
-------
TABLE A-8.1 (Cont'd)
C. Synthesis Gas and Water Gas Plants
The manufacture of ammonia or methanol requires a low level of methane in
the synthesis gas. Synthesis gas from coal is produced in WD plants by
cyclic operation, including autothermic reforming of hydrocarbons, or by
continuous gasification with oxygen or an oxygen/air mixture. If coke is
specified as the feed, the upper (distillation) section of the gasifier is
omitted.
Client and Location
OSW Fertilizer Plant
Linz, Austria
Vetrocoke, Porto Marghera,
Italy
Montecatini, Crotone, Italy
Montecatini, St. Giuseppe
di Cairo, Italy
I.M.A.D., Naples, Italy
State Works, Semtin,
Czechoslovakia
D. Swarovski Co.,
Wattens, Austria
Edison S.p.A., Milan, Italy
Marconi S.p.A., Aquila, Italy
Public Utility, Paris, France
Public Utility,
Fuerth, West Germany
Mode
Oxygen Blown
Bituminous Coal ,
Cyclic
Oxygen Blown
Oxygen Blown
Bituminous Coal,
Cyclic
Coke, Cyclic
Coke, Cyclic
Coke, Cyclic
Coke, Cyclic
Coke, Cyclic
Coke, Cyclic
No. of Units
2
2
2
1
1
3
1
52
-------
SECTION B - DESCRIPTION OF GAS CLFAN-UP SYSTEMS
ON OPERATING GAS IF IER INSTALLATIONS
1 . Introduction
2. Koppers-Totzek
3. Lurg!
k. Wilputte
5. Woodal1-Duckham/Gas Integrals
53
-------
B.I Introduction
Particulates, sulfur and nitrogen are the three pollutants
in industrial gases that have received the most attention due to
their potential environmental and health effects. Their removal has
been standard practice for decades in the manufactured and natural
gas industry. Perhaps more attention has been paid to the removal
of particulates and sulfur than to nitrogen. It is only recently
that health effects of nitrogen oxides have begun to be given seri-
ous consideration. Much of the need to remove particulates is associ-
ated with potential problems in compressing and/or burning dirty
gases.
Raw coal-derived gases also contain tars along with sulfur
compounds, and good removal of tars and particulates was a practical
necessity in the early days of manufacturing and distributing coal-
derived fuel gases. Hence, many years ago, means were developed to
remove tars and particulates from raw gas streams. These methods
may not be sophisticated in today's world but they served the purpose
in their day. Many of these methods were dual-purpose schemes. For
example, the dry box hydrated iron-oxide process removes essentially
all hydrogen sulfide from the gas stream and will also remove traces
of tar. Furthermore, in the event of an upset in the tar scrubbing
system, the boxes will capture a slug of material. Although not
specifically intended as such, the iron-oxide boxes were effective
in preventing tars from entering the gas distribution system.
Because the industrial gas is burned in vented equipment
(as compared to household gas which is frequently burned in nonvented
equipment), the pollutant levels in industrial fuel can be somewhat
-------
higher than permitted in pipeline gas. It would appear reasonable
to assume that industrial fuel gas should have about the same, or
perhaps, as previously suggested, lower levels of contaminants as
any alternative liquid fuel would have that complies with the
Clean Air Act of 1971. Consequently, a basis can be established
to evaluate and develop pollution control systems for the generation
and use of industrial fuel gas from coal.
If a sulfur emission criterion of 0.5 Ibs of SC^ per 10
Btu is used as a guide, application of this specification to a typi-
cal air-blown gasifier product shows that the degree of sulfur removal
required for industrial fuel gas is substantially different from that
required for pipeline gas. This difference can be clearly seen in
Figure B-l.l which has been prepared to show, in a generalized way,
the effect of coal sulfur content on the degree of sulfur removal
required to meet the proposed level. The chart has been based on
calculations for a low heating value feed (lignite), a typical
bituminous coal with a heating value of 12,500 Btu/lb, and the
following assumptions:
a. 75% gasification efficiency
b. 80% of the sulfur in the coal appears in the raw
product gas.
For other coals, a generalized formula for sulfur removal is given
in the Appendix. The sulfur removal curve for pipeline gas looks
as if it were a vertical straight line because all the values are
so close to 100%, but some of the points as calculated follow:
55
-------
-------
% Sulfur in Coal % Removal for Pipeline Gas
0.2 99-79
0.5 99-92
1.0 99-96
2.0 99.98
5.0 99-99
10.0 99.996
While pipeline gas requires just about complete sulfur removal, the
chart shows that a 1.^5% sulfur coal with a 12,500 Btu/lb heating
value would require only 30% removal from the raw product gas to
meet the 0.5 lb S02 per 10^ Btu level. Even when processing coals of
a high sulfur content, there is a great difference in the required
removal efficiency for the two gases. These generalized curves are
of importance because they show the great difference in sulfur re-
moval requirements to prepare an industrial fuel as compared to pipe-
line gas. This difference man ifests itself in sulfur removal process
selection and the design criteria for these processes.
The technology of the past two decades has been developed
and improved from the point of view of essentially total sulfur
removal from gases as would be required for the pipeline industry
and the chemical industry. This currently practiced technology
of almost total sulfur removal is examined in this section in con-
sideration of its proven capability, but a new look at partial re-
moval of sulfur is recommended. The ability to achieve a level of
removal exceeding industrial fuel requirements suggests that partial
removal might be easier and more economical.
57
-------
B.2 Clean-up System for the Koppers-Totzek Gas ifier
Gasification:
The Koppers-Totzek gasifier and a proposed clean-up system
for producing clean, desulfurized utility gas or synthesis gas
is shown schematically in Figures B-2.1 and B-2.2. Oxygen, steam
and coal react in the gasifier, converting the coal volatile matter
and carbon into gas and the coal ash into molten slag. About 50
to 70% of this slag leaves the gasifier through the bottom and
solidifies upon contact with water in the quench tank situated
beneath the gasifier. The remainder of the slag, along with any
ungasified carbon, is entrained with the gas leaving the top of
the gasifier. If necessary, water sprays freeze any slag droplets
prior to entry into the waste heat boiler to prevent solidification
on the tubes. In the waste heat boiler, steam up to 1500 psig is
generated.
A typical composition for the gasifier outlet gas before
any clean-up is given in Table B-2.1. As shown, the gas may have
about 12 grains/SCF of particulates, 0.2-0.3% HgS and COS, and about
0.2% nitrogen compounds including ammonia, cyanides and oxides of
nitrogen. Due to the high reaction temperature, phenols, pyridenes,
tar, oil or other condensable hydrocarbons are not contained in the
gas.
58
-------
FIGURE B-2.1
KOPPERS-TOTZEK PROCESS
GAS IFI CAT I ON, COOLING & PARTICULATE REMOVAL
Coal BFW
LJ
Gasi f ier
1
Steam
Slag
<
> i
Quench
Tank
N Spray
Washe
Coole
Wash
Water
>
\
l\ ^
1
99 99
Granular Slac
to Mine
i
r
r
Gas
»
The! sen
Di si ntegra-
tors
Overflow
Water
*
El
Cool ing
J^ater
-<-J
Slurry
Gas
i
Mist
imi nator
l
Gas
»
Electrostat i c
Preci pi tator
— ^-
-^-
-^-
<
Cool ing
Tower
-1
Blow DC
Treatm<
Clarif
Wate
Ian f ier
>lu Ige Fi
1 '
Fi ter
1
Gas to Chemical
Clean-up
Fi 1ter Cake
to Mine
59
-------
FIGURE B-2.2
GAS PREPARATION FOR SYNTHESIS
Gas from
Koppers-Totzek Process
I
Compressor
Rectisol
Sulfur
Remova1
1
Compressor
Shift
Conversion
I
Rectisol
C02
Remova1
Ni trogen
Wash
To Synthesis
Process
60
-------
TABLE B-2.1
KOPPERS-TOTZEK GASIFIER GAS COMPOSITIONS
(Coal Unspecified)
Volume Percent
To Compression &
Component Gasifier Outlet Acid Gas Removal
CO 37-36 49.50
C02 7-13 3-k2
CHij 0.08 0.11
H2 25.17 33-35
N2 0.30(1) 0.40
H2S 0.23 0.30
COS 178 ppmv 235 ppmv
HCN 288 ppmv 300 ppmv
NH, 0.17 0.22
H20 29.19 6.20
Ar 0.32 0.42
S02 ' 22 ppmv 15 ppmv
NO 7 ppmv 7 ppmv
Particulates (grs/SCF) 11.57 <0.002
NOTE: (1) Possible Sources of Nitrogen With Oxygen-Blown Gasifica-
tion includes Impurity in Feed Oxygen and Conversion of
Fuel-Bound Nitrogen
Reference: Farnsworth, J.F., Mitsak, D.M., Kamody, J.F.,
"Clean Environment with K-T Process", presented
at EPA Meeting: Environmental Aspects of Fuel
Conversion Technology, May, 1974.
61
-------
Gas Cooling with Participate Removal:
The next two or three process steps reduce the participates
in the gas to very low levels. The effluent stream from the waste
heat boiler enters a washer-cooler where sprays of water cool the
gas from 350°F to about 100°F while simultaneously removing 90$ of
the entrained particles. Then two Theisen-type irrigated disinte-
grators in series reduce the dust loading to about 0.002 grains per
SCF. If the gas is to be compressed to high pressures for chemical
synthesis or for the production of high Btu gas, wet-type electro-
static precipitators are used to reduce particulates to 0.0001 grains
per SCF. For compression up to 175 psi, precipitators may not be
necessary and the gas, after passing through a mist eliminator and
fan, would be manifolded into the suction of the gas compressors.
The composition of the gas at this point may also be seen in Table
B-2.1.
When Koppers (U.S.A.) first undertook design of Koppers-
Totzek gasification units, they introduced the venturi scrubbers
as a substitute for the washer-cooler and The!sen disintegrator
system. Since no units had been built that way, the German Koppers
Company took a conservative position and suggested a return to
the original design proposal. While the venturi system does consume
power, use of the disintegrator units requires even more power con-
sumption. Each of the two Theisen units would require 700 HP for
gas from one ^-headed or two 2-headed gasifiers. With the venturi
system, the gasifier is operated at slightly increased pressure
to provide the 50 inches of water pressure loss in the venturi
scrubbers. The first unit that might be installed in the U.S.A.
is likely to have Theisen units because the designers will look
62
-------
favorably on the proven approach. However, in the future, venturi
scrubbers might be a good alternate choice. For blast furnace appli-
cations, venturi scrubbers are now being installed rather than Theisen
disintegrators.
Cooling Water Clean-up:
Particulate-laden waters from gas cleaning and cooling plus
overflow from the quench tank are piped to a clarifier. Thickened
clarifier sludge is filtered, and the filtrate is returned to the
clarifier. Filter cake and granular solids removed from the quench
tank by means of a scraper-conveyor assembly are loaded into rail-
road cars for disposal at the mine site. Clarified water is pumped
to a cooling tower and recirculated to the gas cooling/cleaning and
the solids quench systems.
Water recirculation permits the build-up of many chemicals
in the water. Analyses at various steps in the process have been
reported from a plant in Kutahya, Turkey. The data given in Table
B-2.2 show the order of magnitude of the chemical concentrations and
identify some possible contaminants. Although not reported in the
Kutahya washer-cooler analysis, dissolved hLS might be expected and
therefore stripped out by air in the cooling tower. If all the
dissolved H_S were stripped into the air, the discharge air concen-
tration would be 1-2 ppm by volume. While this is far above the odor
threshold, Koppers Company experience shows that there is no odor
problem. A previous Environmental Protection Agency study verifies
this finding (K-T Reference #6). Appreciable bioxidation, common
in the cooling water circuit, may account for the indicated analytical
results and Koppers' observations.
63
-------
TABLE B-2.2
KOPPERS COAL GASIFICATION
WATER ANALYSES. KUTAHYA, TURKEY(l)
Sample Location
pH Value
Conductivity
Total Hardness
CaO
MgO
Ma
K
Zn
Fe
NH/,
N02
N03
P0|, Total
Cl
S0{,
KMnO/t Consumed
COD
Si02
Suspended Solids
Hot Residue, 800°C
Stripped Residue
Hot Residue, 800°C
Cu
mho/cm
0 dH
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
8.8
7.6
10-*
20.8
78
97
17-5
5.6
0.01
0.05
0.32
0.02
58.2
1.89
18
42
0.26
8
14
14.8
14
4
568
268
0.01
8
1
10
33
101
161
17
8
0
0
157
0
3
0
85
216
0
9
18
16
4612
3918
812
550
0
.8
.8
-3
.5
.5
.8
.03
.22
.13
• 32
.81
.52
.0
.01
8
2
10
36
78
194
17
10
0
1
184
4
13
1
96
155
12
400
128
14
5084
4356
940
588
0
• 9
.0
-3
.8
.5
.0
.02
• 95
.47
.7
.21
.5
.8
.01
7
9
22
85
102
17
6
0
0
25
5
34
1
53
147
7
r»z»
Uc
14
18
30
278
134
606
366
0
.5
.7
-4
.8
.5
.8
.03
.26
.34
.0
.69
.0
+ ar~i
teci
.6
.06
8
1
10
34
135
145
17
8
0
0
137
0
24
0
57
255
1
•orl
11
16
19
3072
2690
706
526
0
.8
.8
-3
.0
.5
.0
.02
.20
.24
.7
.81
.4
.8
.01
8.
1.
10-
34.
179
113
17-
8.
0.
0.
122
4.
22.
2.
46
109
14.
145
63
42.
50
46
724
512
0.
9
8
3
8
5
0
02
64
37
9
70
0
6
06
8.9
1.7
10-3
35.2
179
129
17-5
7-9
0.02
0.24
72
23.7
42.0
2.41
36
153
0.7
60
30.6
58
42
828
534
0.27
1) Cooling water to gasifier seal pot.
2) Water from the gasifier seal pot.
3) Wash water after washer-cooler.
4) Wash water after Theisen washer.
5) Water into clarifier.
6) Water out of clarifier.
7) Water out of cooling tower,
NOTE: I) Process Water Streams Circulated within Process Unit. Any purge
Stream from the System Would Require Treatment before Discharge.
Reference: Farnsworth, J.F., Mitsak, D.M., Kamody, J.F., "Clean Environment
with K-T Process", presented at EPA Meeting: Environmental Aspects
of Fuel Conversion Technology, May, 1974.
64
-------
An additional cooling tower effluent is the drift loss of
mist from the cooling tower. The mist will contain dissolved and
suspended solids, which will result in deposits on the ground and on
nearby equipment. To minimize solids build-up in the cooling tower
circulation and the intensity of its attendant problems, blowdown
is necessary. Acceptable stream standards must be met by treatment
of the blowdown, principally to destroy HCN and NH^ through chemical
chlorine oxidation and pH adjustment. Evaporation, windage and blow-
down water losses at the cooling tower, plus moisture in the filter
cake from clarifier sludge and in slag, necessitate the addition of a
small quantity of make-up water to this system. If water is at a
premium, air cooling may be used for cooling down to 1^0 F in certain
applications and the cooling tower can be reduced in size to provide
only the final trim in water temperature.
Gas Clean-up: .
The cool, clean gas leaving the gas cleaning system con-
tains sulfur compounds which must be removed to meet gas specifica-
tions. The type of system chosen depends upon the end uses and
pressure of the product gas. For low pressures (up to 150 psig) and
low Btu gas application, there are the chemical reaction processes,
such as amine and carbonate systems. At higher pressures, the physi-
cal absorption processes, such as Rectisol, Purisol and Selexol,
are recommended. The choice of the process is also dependent upon
the desired purity of the product gas and the desired selectivity
with respect to the concentrations of carbon dioxide and sulfides.
There are several acid gas removal processes with the
capability of reducing the sulfur content in the gas to 5 ppm by
65
-------
volume. The processes are based on absorption in solution and subse-
quent stripping of the acid gases, H2S and CQ-2, from the solution.
The physical absorption processes, which operate at pressures of
300-400 psig, exhibit the greatest selectivity with respect to
hydrogen sulfide and carbon dioxide removal. Since no chemical
reactions occur, these processes do not form stable compounds, such
as thiosulfates and thiocyanates. Some chemical reaction processes,
such as carbonate and atnine, which form the aforementioned stable
compounds, can be used but will require periodic dumping of the
solution in order to maintain removal efficiency. Dumped solution
will require treatment to meet permissible discharge limitations.
The choice of process is dependent upon economics, environmental
control, purity of product gas, and desired acid gas selectivity.
A plant can be designed to reduce sulfur in product gas to 5 ppm
by volume, control the H2S level in carbon dioxide to 10 ppm by
volume, and control the liquid effluent to zero pollutants.
The tappers Company, it is understood, is proposing the
use of MOEA (methyl dlethanolamtne) for selective removal of H2S.
This chemical has sufficient selectivity to provide about a 22%
H2$ concentration in the acid gas to the Claus plant. The acid
gas removal facilities consist of a COS hydrolysis column followed
by an absorber. Within the COS hydrolysis column, the gas is con-
tacted with hot circulating MDEA solution to promote hydrolysis of
COS to H2S in order to facilitate a high degree of sulfur removal
withtn the absorber. The acid gases are stripped from the MDEA
absorbent and sent to Claus units. Gas leaving each H?S absorber
contains approximately 115 ppm of HjS, plus COS, or 0,076 pounds
of S02 equivalent per million Btu of gross heat content of the fuel
gas.
66
-------
At the African Explosives and Chemical Industries Ltd.
ammonia plant in South Africa, where coal is gasified using the
Koppers-Totzek process, gas is cleaned up with the Rectisol process.
First a methanol wash at -36°F and 30 atm. desulfurizes the gas.
Then carbon dioxide is removed, following the CO shift conversion,
by washing with methanol at -72 F and 51 atm. Finally, residual
CO, argon and methane are removed by washing with liquid nitrogen
at -310°F (K-T Reference #10).
The acid gas stream, containing a minimum of I1* volume
percent HjS, is catalytically converted to elemental molten sulfur
in a Claus unit. The tail gases exiting the Claus unit contain S02
and can be treated to catalytically reduce the S02 to H2$• Scrubbing
with an amine solution absorbs the H.S, and subsequent stripping
yields an HjS stream which is recycled to the Claus unit. This
combination results in overall sulfur recovery of 99+%.
67
-------
B.3 Clean-up System for the Lurgi Gasifier
Gas ificat ion:
The Lurgi gasifier and its clean-up system for producing
SNG has been carefully studied and, therefore, is illustrated in Figures
B-3.1, B-3.2 and B-3.3 The proposed El Paso Burnham Coal Gasification
Complex exemplifies such an application for the Lurgi process when used
for pipeline gas. Industrial gas, as explained in the Introduction,
would require less stringent control of pollutants and thus, would pro-
vide more flexibility in the control methods for pollutants.
Coal entering from lock hoppers reacts with a mixture of
oxygen and process steam introduced into the bottom of the gasifier.
About 86% of the coal fed to the gasifier is gasified and the remaining
lk%, which is mostly carbon, is burned in the combustion zone. Ash,
bearing only a very small amount of unburned carbon, passes down through
the grate, out of the gasifier through an ash lock hopper, and into an
•
ash bin where ash is cooled by water quenching. The ash is separated
from quench water in a clarifier and sent to the mine site for disposal.
It is estimated that 1.k% of the DAF coal is not consumed and leaves
with the ash.
Gas Quenching:
Raw gas leaves the gasifier at about 850°F containing
carbonization products such as tar, oil, naphtha, phenols, ammonia,
and traces of coal and ash dust. On an oil-free and dry-gas basis,
the gas leaving the gasifier will have the following approximate
composi tion:
68
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FIGURE B-3.1
LURGI PROCESS
GASIFICATION AND ASH HANDLING
Coal
Coal
Lock Hopper
Vent Gas to
Gas Cool ing
Steam
Gasi fier
i
Ash
Lock Hopper
Ash
Quench
Clari fier
Gasifier Outlet Gas
to Gas Cool ing
Vent Gas to
Gas Cool ing
Water
Ci rculation
Ash to Mine
-------
FIGURE B-3.2
LURGI PROCESS
GAS COOLING, SHIFT CONVERSION AND GAS LIQUOR PROCESSING
Gasi fier
Outlet Gas
Spray
Wash
Coo 1e r
100 psig Steam
Waste Heat
BoiIer
Raw
Gas
i
Shift
Convers ion
Gas
Cool ing
Steam
I
Waste Heat
Boiler
Gas
Cool ing
Gas to Acid Gas Removal
70
H
Compressor
Gas Liquor
Flash
Separation
Tar
Separat ion
Gas Liquor
Treatment
Treated Effluent Liquor
to Biological Treatment
Tar to
Storage
Phenol
Ammon i a
-------
FIGURE B-3.3
RECTISOL GAS CLEAN-UP AND METHANATION
Synthesis Gas
Refr.
Naphtha
Water to
Gas Liquor
Treatment
t
Prewash
Tower and
Flash Tank
Methanol
i
Naphtha
Methan'oJ | Extractor &
Azedtrope
Tower
Methanol-
Water
Column
Gas
Acid Gas to
Claus Unit
H2S
Absorber
Methano Flash
Regenerator
Recovered
Methanol
Hot
Regenerator
Compressor
1st Stage
Compressor
2nd Stage
SNG
71
-------
C02 28.6 MoU
CO 20.2
H2 37.9
Chi, ll.it
C2H6 .62
H2S + COS .1*9
N2 + A .33
This crude gas is then cooled rapidly to AGO F by quenching
in the spray washer with gas liquor, an aqueous phase condensed from
the gas. Further cooling in the waste heat boilers drops the temperature
down to about 360°F to 370°F while generating 100 psig steam. As higher
boiling tar fractions are condensed, coal and ash dust are bonded to the
tar. Some of the liquid condensed in the waste heat boilers is recycled
to the wash coolers, and the excess is drawn off to gas liquor separation.
Gas liquor from the spray washer and the gas cooling area is
flashed to atmospheric pressure in an expansion vessel to remove dis-
solved gases. Heavy tar is separated out in another vessel and sent to
storage. The mixture of tar and dust is returned to the gasifier for
cracking and gasification. The detarred liquor is sent to the gas liquor
treatment area to remove dissolved phenol and ammonia.
Shift Conversion and Gas Cooling:
Raw gas leaving the gasifier section is divided into two
streams; one is sent to shift conversion where the hydrogen content
is increased and the other goes directly to gas cooling. Crude gas
vented from the cyclic operation of the coal lock hoppers, the expansion
gas from gasification, and small quantities of recycle gas from other
areas are compressed and injected into the stream which goes directly
to the gas cooling area.
The crude gas bypassing shift conversion is cooled in a
series of units comprised of the following: a waste heat boiler
72
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generating 60 psig steam, a low pressure steam generator, an air
cooler and possibly a water trim cooler. Gas liquor and tar from the
first two units are transferred to the primary gas liquor separator.
The remaining condensate streams, comprised of gas liquor and a tar
oil naphtha mixture, are separated in a second gas liquor separator.
Converted gas is cooled by the following series: exchange
with boiler feedwater, generating low pressure steam, boiler feedwater
deaeration, and finally air and water trim cooling. The two gas streams
are combined into a single stream before acid gas removal. Gas liquor
and condensate from the first three steps are cooled and combined with
condensate streams from subsequent cooling. The total stream is then
sent to the gas liquor separator where separation of the tar oil naphtha
mixture from gas liquor will occur. Gas liquor will be pumped to the
gas liquor treatment area and the tar oil mixture will be transported
to storage.
Gas Liquor Treatment:
Gas liquor condensed in coal gasification and gas processing
contains phenols, ammonia, carbon dioxide, hydrogen sulfide and caustic
effluent. The Lurgi Phenosolvan Process can be used to recover the
phenols. Ammonia is recovered in aqueous solution. The other con-
taminants are removed from the liquor by heating and stripping.
Incoming gas liquor entering the Phenosolvan Process is
filtered in gravel filters to remove suspended matter. The filtered
liquor is then mixed with an organic solvent (isopropyl ether) in
the extractors where phenols are dissolved in the solvent. The
phenol-rich solvent extract is collected for feed to the solvent
distillation column, where crude phenol is recovered as the bottoms
product and solvent as the overhead product. Recovered solvent is
73
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separated from the water by settling and then, with some make up of
fresh solvent, recycled to the extractors.
Before being heated and steam stripped, the lean liquor
(rafflnate) from the extractors is stripped with fuel gas to remove
traces of solvent which are picked up in the extraction step. The
resulting solvent-laden fuel gas is scrubbed with crude phenol to
recover the solvent. The phenol-solvent mixture is then fractionated
in the solvent recovery stripper to produce the crude phenol product
and collect the solvent for recycle to the extraction step.
Solvent-free liquor is heated and steam stripped to remove
carbon dioxide, hydrogen sulfide and ammonia.- The carbon dioxide
and hydrogen sulfide are removed separately from the ammonia and
returned to the process for sulfur recovery.
Ammonia is stripped from the liquor and condensed as an
aqueous solution of about 25 weight percent NHj. In the event that
a market does not develop for this product, the wet ammonia vapor
can be consumed as plant fuel provided NOX emissions are within
limits.
Treated effluent liquor from the ammonia stripper is cooled
and delivered to the biological treatment plant for further reduction
of contaminants to render it suitable for use as cooling tower make-up.
Before biological treatment, the effluent contains less than 20 ppm
phenols and less than 60 ppm free ammonia. Afterward, the effluent
meets the regulations for disposal of waste liquor.
-------
Gas Clean-up and Methanation:
For SNG manufacture, the gas purification plant is designed
to remove h^S and COS to a total sulfur concentration of 0.1 ppm by
volume using the Lurgi Rectisol Process. After methanation and
first-stage compression, the gas is washed to reduce C02 content.
Because of the low operating temperatures (down to -50°F), all
hydrocarbons heavier than G£ are removed, leaving a very clean gas
stream for the methanation section.
The mixed gas is chilled before entering the prewash tower
where water and naphtha are removed by cold methanol wash. Naphtha
is recovered from methanol and water by means of the naphtha extractor.
Naphtha recovery is maximized by recycling the naphtha-methanol mix-
ture through the azeotrope column. The methanol is recovered by
distillation in the methanol-water column.
The naphtha-free gas enters the h^S absorber where H2$ and
COS are removed down to 0.1 ppm by volume total sulfur by cold methanol
wash first used for C02 absorption. Heat of absorption is removed
by refrigeration. Some of the absorbed acid gases are removed from
the methanol wash by multi-flash in the flash generator. The remain-
ing acid gases are completely stripped in a second regenerator operating
at a higher temperature. All the acid gas streams are combined and
delivered to the sulfur recovery plant.
The sulfur-free synthesis gas leaves the Rectisol Unit
absorber, exchanges heat with returning methanated gas to save on
refrigeration, and moves on to the methanation unit. The returned
methanated gas enters the CO. absorber. The C02 content of the gas
is reduced by regenerated cold methanol wash. The heat of absorption
75
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is removed by a refrigerant. The high Btu purified dry gas is warmed
and sent to the second-stage compression unit.
The mechanical compression refrigeration unit provides
refrigeration at two temperature levels. The high temperature
level refrigeration (32°F) is used to condense most of the water
out of the mixed gas and the methanation product gas. The remaining
water vapor in the gases is prevented from freezing by methanol in-
jection. The low temperature level refrigeration (about -50°F) is
used to achieve the low temperature required for effective methanol
wash.
-------
B.4 WMputte Gas Clean-up System
The Wilputte gasifier is a moving-bed gasifier, operating
at atmospheric pressure. The operating principles are described in
Wilputte Bulletin No. 7662. as quoted below:
"The gasification process is counter current.
Thus, the coal is fed downward and the gas flow
is upward. Moist air flows upward through an ash
zone to the cumbustion zone, in which the carbon
residue from pyrolyzed coal is burned to carbon
dioxide (the water in the air does not enter
this reaction). Various reactions occur in the
bottom of the coal bed which lays directly above
the combustion zone. The carbon reacts with the
carbon dioxide to form carbon monoxide or with
water to form hydrogen and carbon monoxide or
carbon dioxide. In the middle of the coal bed,
carbon reacts with carbon dioxide to from carbon
monoxide. In the top of the coal bed, coal is
pyrolyzed to form a carbon residue and volatile
hydrocarbon products. The coal in the top of the
charge passes through a plastic state when it be-
comes heated to about 850°F if the coal is a coking
coal. A rotating rabble arm is used to mix and
break up this plastic layer so as to get a uniform
distribution for the upward flow of gas."
An example of the gas clean-up system used on Wilputte
gasifiers is the installation at the Holston Defense Corporation in
Tennessee. This plant consists of 12 gasifiers each having a diameter
of 9'-2". Each unit is rated at 2k T/D of coal feed. The plant was
built in 19^5. Two gasifiers at any one time are in operation using
a metallurgical grade of bituminous coal having a sulfur content of
less than 1%. These gasifiers have had an excellent maintenance re-
cord. The original brick lining in the gasifier is still in use.
77
-------
A very simple and straightforward gas clean-up system is
used on the Holston gasifier installation. Raw gas leaves the gasifier
at a temperature of about MOO F and enters a refractory lined cyclonic-
type separator that removes larger sized entrained particles. The gas
then flows to a "primary gas cooler" through a collector main. Hot
liquor is sprayed into the main reducing the gas temperature to about
200 F and removing the bulk of the tar prior to entry into the cooler.
No waste heat is recovered from the gasifier effluent stream.
The "primary gas cooler" is packed tower that is irrigated
with cooled liquor. Condensed tars and wash liquor from the collector
main and the gas cooler flow by gravity to a sump where separation of
liquor and tar occurs. The tar has a specific gravity of about 1.15 and
therefore settles to the bottom of the sump. This material is pumped
to storage for subsequent use as a boiler fuel. Solids are periodically
raked from the sump.
Liquor from the sump is pumped and recirculated in a
split-flow arrangement. Part of the liquor is returned directly to
the collector main as"hot" liquor while the remaining part is cooled
via water exchange to enter the top of the cooler. Excess liquor is
treated by sand filters and carbon absorption to produce an acceptable
effluent stream.
Overhead has from the "primary gas cooler", which actually
is a counter-current gas scrubber, flows to an exhauster that boosts
the scrubbed gas to a fuel gas header pressure of about 10 psi. The
gas then flows to the individual burners for combustion without a
sulfur removal step.
78
-------
As can be seen, the clean-up system is extremely simple,
comprising no moving parts other than the tar and liquor pumps and the
exhauster. Over 30 years of operation of this gas clean-up system is
ample proof of its basic operability and simplicity. Unfortunately,
this simplicity cannot be maintained if the plant were tooperateon
a high-sulfur coal and would, therefore, require sulfur clean-up
on the gas product. Nevertheless, the existing gas clean-up system
used at the Holston installation is a working example of a simple,
pratical approach to removing tar, dust and some ammonia from a
coal gasifier effluent so that the product gas can be properly burned
as an industri al fuel.
79
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B.5 Clean-up System on Woodal1-Duckham/Gas Integrals Gasffier
Effluent
The principle of Woodal1-Duckham/Gas Integrale (WD/Gl) two-
stage coal gasification is to separate the volatile matter of the
coal before subjecting the remainder of the coal to the high tempera-
tures of the gasification reactions. As the coal descends through
the fixed-bed gasifier, the coal is dried, then evolves gas, light
hydrocarbons and tar, and reaches the bottom of the upper part of
the gasifier as semi-coke, or char if the coal is a nonagglomerating
coal such as sub-bituminous or lignite.
The semi-coke or char now passes into the gasification
zone, where temperatures rise from about 1200°F at the side gas
off-take to typically 2200°F at the final combustion stages.
The gasification agent converts the semi-coke into carbon
monoxide, carbon dioxide, hydrogen, some undecomposed steam and nitrogen.
This product Is termed clear gas. About half the clear gas is with*
drawn at 1200°F via the side gas off-take, the remainder passing
overhead to dry and devolatilize the coal.
This split-flow system, together with a very low gas off*-
take velocity, keeps fines carryover to a very low level. Very low
fines carryover is one of the features of two-stage gasification
which distinguish it from single-stage gasifiers. Figure B-5.1
shows how two-stage gasification lends itself to a simple and effec-
tive gas clean-up. No direct contact scrubbing Is used. The top
gas at 250°F, containing all the coal volatile matter, is electro-
statically detarred, indirectly cooled, and passed through a second
80
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•GASIFICATION SECTION
•DESULFURIZATION SECTION-
oo
vCool
^Steam Tar Oil
Precipitator „ Precipitator H2S
Absorber
/F0TW
r CW| r
&,
K
x
E
,, Gasifier
Ash
Vacuum
Filter
n
Air Blower Tar& Oil Aqueous
(or O2 supply) Tank Effluent
Pumping Air Oxidiser Sulfur
Tank Blower foam tank
Fuel
Gas
e
JSulfur
Cake
Fuel
Oil
e
wryci PROCESS FOR COLD DESULFURIZED FUEL GAS
FIGURE B-5.1
-------
precipitator to ensure full demisting.
The clear gas at 1200°F, containing no volatile matter,
ts dedusted and cooled in a combination of waste heat boiler and
tubular cooler.
The combined gas can be desulfurized by several well-known
processes, such as the Stretford process, A Stretford unit in simpli-
fied schematic form is shown in Figure C-7-1, and this process has
been used to desulfurize gas from WD/GI gasifiers to produce a sul-
fur cake.
Host of the sulfur contained in the coal feed will be in
the product gas, mainly as H2S with 5 to 10? of the sulfur appear-
ing as COS and CS_, The H.S is easily removed in a Stretford unit
to give a final product gas which can comply with SO. emission re-
quirements for industrial fuel gas. For low-sulfur coals, HgS
removal may not be necessary. Due to the low temperature pre-
distillation of the coal obtained in the two-stage process, the
product contains only traces of hydrogen cyanide (a few ppm), unlike
many gases obtained from coal. This simplifies and reduces the cost
of the Stretford unit.
The cooling of the sulfur-bearing gases prior to entry
into the sulfur removal system produces a contaminated aqueous
stream. An important feature of the process is that by using
indirect cooling the volume of contaminated aqueous effluent is
sharply reduced compared to direct contact cooling. This minimizes
the cost of effluent disposal and, indeed, often makes simple
incineration as attractive as biological treatment of this aqueous
stream. Recycle of this stream to the gasifier has been practiced,
82
-------
but further development work is required before WD/G1 feels it is
appropriate to offer this process variant on future plants.
It should be emphasized that each feedstock coal presents
differing operating problems. Therefore, the gas clean-up system
would normally be tailored to the specific type of feedstock and
may not be the same for each plant. The flow sheet shown in Figure
B-5.1 does represent a design that has been reduced to practice (in
South Africa). Specific gas analyses are not available, but the pro-
duct gas should meet all proposed EPA guidelines for the following reasons.
Part iculates:
Solids removal is effected by cyclones, electrostatic precipi-
tators and finally by scrubbing with the acid-gas removal liquid (Stretford
solution). It is highly unlikely that the fuel gas, after these successive
processing steps, would exceed the particulates emission limit of 0.1 Ibs per
10 Btu of gas, and it most probably contains a small fraction of this amount,
Particulates formed during the combustion step for low Btu gaseous fuels
are generally known to be far less than 0.1 Ibs per 10 Btu; hence, the
sum of the ash particles plus the soot particles should be less than 0.1
Ibs per 10 Btu.
Sulfur Dioxide:
About 90% of the sulfur-containing gas in the raw fuel
gas is H.S, the remaining 10% being COS and CS_ with traces of
mercaptans. The Stretford unit removes essentially all of the
H_S but does not remove COS, CS_ and the mercaptans. If it is
assumed that 85% of the sulfur in a feed coal containing 3% sulfur
appears in the gas product, then this 10% of "organic" sulfur that
is not removed will produce nearly 0.5 Ibs of S02 per 10 Btu when
this fuel is burned. Therefore} given the suggested EPA guideline
33
-------
of a maximum of 0.5 Ibs of S0_ per 10 Btu for industrial gas, it
appears that coal of up to 3% sulfur can be converted to fuel gas
of acceptable quality using the equipment as shown In Figure B-5.1.
Use of coals higher than 3% sulfur may require the partial removal
of organic sulfur compounds (in addition to removal of H-S) in order
to meet the emissions requirement guideline. Some processes are
available to meet this requirement.
The NOX content of combusted gas is determined mostly by
flame temperature, excess air and the chemically-bound nitrogen in
the fuel itself. Because the theoretical flame temperature of the
WD/GI fuel gas is some 200°F colder than high-Btu (natural) gas,
the thermal fixation of N0x should be reduced. Fuel-bound nitrogen,
if present, should exist in the converter effluent gas as HCN and
NH, with only traces in the HCN form. NH^ will probably be found
in varying amounts, depending on the quality of the coal, time-
temperature history of the coal particle, and hydrogen partial pres-
sure. Not all of this ammonia will end up as NO because some
J\
ammonia is removed by the gas clean-up as an aqueous condensate and
various studies show limited combustion of the remainder to NOX.
Thus, ?t is believed that NO formation from this fuel gas will be
substantially below the guideline emission level of 0,4 Ibs of NO
6 x
per 10 Btu.
The presence of water vapor in a gas is also known to lower
the formation of NO during the combustion process due to a cooler
y»
flame. WD/GI gas normally contains an appreciable amount of water
vapor because it is saturated with water at about 100 F or higher.
Therefore, the beneficial effect of decreased NO formation is generally
^
realized when this gas is burned.
84
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Commercial Operations:
The commercial operations of several WD/GI units to produce
synthesis gas for ammonia or methanol synthesis are further proof
that gas clean-up methods have been developed to adequately purify
coal-derived gases to a degree of purity higher than required by
industrial fuel gas. The sulfur and particulate contents of
synthesis gas must be essentially zero in order to avoid the poison-
ing of catalysts; hence, gas clean-up to less stringent criteria
would appear to be demonstrated. This is not to say that significant
improvements cannot be made, but coal-derived gases can be purified.
85
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SECTION C - COMPARISON OF IRON-BASED CLEAN-UP PROCESS AND
THE STRETFORD PROCESS
1. Introduction
2. Development of Gas Clean-up Processes
2.1 Iron Oxide Box Purifiers
2.2 Liquid-phase Iron Oxide Processes
2.3 Stretford Process
3, Basic Chemistry
3J Iron Oxide Box Process
3.2 Liquid-phase Iron Oxide Processes
3-3 Stretford Process
A. Design Basis and Assumptions
5. Iron Oxide Box Purifiers
5.1 Design Based on American Practice
5.2 Design Based on European Practice
5.3 Discussion
5.4 Advantages and Disadvantages
6. Liquid-phase Iron Oxide Processes
6.1 Process Description
6.2 Process Requirements
6.3 Advantages and Disadvantages
7. Stretford Process
7,1 Process Description
7.2 Process Requirements
7,3 Advantages and Disadvantages
86
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C. I Introduction
Recent interest in gas clean-up has generally been directed
toward the use of processes which almost completely remove pollutants
because of the stringent standards necessary for pipeline gas or
synthesis gas. With revived interest in industrial gas generation,
a new prospect to be considered is a reexam5nation of old technology.
This type of process may not be so complex or costly and may be applied
if a lower percentage of sulfur removal is satisfactory as is the
case of industrial fuel compared to pipeline gas. Processes possibly
falling into this category are those based on iron oxide and a great
deal of information on both the iron oxide box and liquid versions
are given in Kohl and Riesenfeld's text (iron Oxide Process Reference
#3). This section, therefore, will consider the old and new techno-
logies that might be applied to make industrial gas by comparing
designs of older iron oxide processes with the Stretford process
as the representative of modern day development. Either way, the
sulfur level can be brought within the guideline level of 0.5 Ibs
of S0_ per 10 Btu for industrial fuels while making elemental sulfur
directly without a Claus unit.
Iron oxide boxes are one of the oldest methods for gas
purification still in industrial use. The engineering design of
units for industrial gas clean-up is possible from publicly available
sources of information. American and European practices differ and
the results of both are presented.
Improvements of the iron oxide process to save on space
requirements and labor and to recover sulfur resulted in the develop-
ment of a continuous liquid-phase sulfur clean-up system. The
economic factors that promoted this development are still valid
87
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today when cleaning industrial gas. Thus, a generalized combination
of the Ferrox, Gluud and Manchester iron-oxide-based liquid-phase
processes was used for a process design comparison case.
-------
C.2 Development of Gas Clean-up Processes
2.1 Iron Oxide Box Purifiers
The iron oxide process was introduced in England around
the middle of the 19th century. Before that time, a wet purification
process utilizing calcium hydroxide as the active agent was used.
Although iron oxide processes are still used on a large scale for
treatment of coal gases, recently developed wet purification pro-
cesses have gradually been replacing oxide box purifiers.
A simple form of the dry-box process utilized for the first
batch-type installations completely removed hydrogen sulfide with
hydrated ferric oxide. At the completion of a cycle, ferric sulfide
formed in the reaction is oxidized to elemental sulfur and ferric
oxide by exposure to air. Hydrogen sulfide is oxidized to elemental
sulfur and water in the overall chemical reaction. The cycle is
repeated until sulfur fills most of the pores and coats most surfaces
with tar and sulfur. Then the bed is less active, pressure drop
increases and the bed must be removed from the box for cleaning.
The oxide was often reused after the sulfur had been removed. Any
current applications of this process must dispose of the tar and
sulfur laden spent oxide not suitable for regeneration in a way com-
patible with present environmental standards.
More economical revivification methods were discovered
later. In one, the iron is revivified continuously by addition
of small amounts of air or oxygen to the purification plant inlet
gas. The other method involves a cyclic in situ revivification
by circulation of oxygen-containing gas after the bed has been
fouled. Eventually, the iron oxide bed must be removed to avoid
89
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excessive back pressure and to maintain good gas contact. Benefits
of the revivification improvements were a savings due to reduced
frequency of loading and unloading and the achievement of a higher
level of sulfur content before it was necessary to dispose of the
iron oxide batch,
2,2 Liquid-phase Iron Oxide Processes
Since the main disadvantages of dry iron oxide purification
are large ground space requirement, high labor costs of the purifica*-
tion plants and the disposal of a large quantity of spent iron oxide,
a search for more efficient methods for hydrogen sulfide removal
from the gases was undertaken. This search-resulted in the develop-
ment of the liquid-phase iron oxide processes. A logical step employed
liquids in regenerative cycles and utilized reaction between iron
oxide and hydrogen sulfide followed by conversion of iron sulfide
to iron oxide and elemental sulfur. Starting with the work of
Burkheiser shortly before the first World War, several processes
which were developed in Europe and the United States used iron oxide
suspended in alkaline aqueous solutions. The Koppers Company of
Pittsburgh introduced the Ferrox process in the 1920's, Gluud intro-
duced a similar process in Germany. In England, a more recent modi-
fication of the Ferrox process, known as the Manchester process,
was introduced.
Use of these processes has been diminishing to some extent.
At present the Burkheiser process is not in commercial use although
a new proposal for a novel coal-gas purification scheme has been
made. While a few Ferrox plants are still operating in the United
States, most have been replaced by more modern systems. The Gluud
process still finds some use. In Great Britain, where the Manchester
90
-------
process was popular, the Stretford process covered later in this
section is replacing it.
2.3 Stretford Process
The Stretford process was developed in the early 1960's
by the Western Gas Board and the Clayton Aniline Company for removal
of hydrogen sulfide from coal gas. Initially, an alkaline solution
containing the sodium salt of 2-6 and 2-7 anthraquinone disulfonic
acid (ADA) absorbed H2S and converted it to sulfur. Due to the
excessively long reaction time between f-LS and ADA, equipment sizes
were large and the solution had an excess dissolved salt accumulation.
The discovery of adding vanadate salts to the alkaline ADA solution
reduced the reaction time. The vanadium salt participates as an
oxidizing agent which ADA later restores to the oxidized form.
91
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C.3 Basic Chemistry
3.1 Iron Oxide Box Process
The following equations illustrate the chemistry involved
in iron oxide purification:
Absorption: Fe2°3^H2°^x + 3 H2S = Fe2S3 + ^x+3^ H2° + heat
Revivifying: 2 Fe2S3 + 3 QZ + x H20 = 2 Fe,/) (H20) + 6 S + heat
Overall Reaction: 2 H^ + 02 - 2 H20 + 2 S + heat
Depending on operating conditions, a large number of
other reactions may occur. The reaction mechanism is principally
influenced by temperature, moisture content, and pH of the purifying
material.
Both, mixed and unmixed, iron oxides are used for gas-
purifying materials. Both types contain iron oxide that may be
prepared by the air oxidation of iron borings in the presence of
water and lime. The make-up of each class follows:
Unmixed oxide - pure hydrated ferric oxide and sometimes
fibrous materials such as those occurring
in natural iron oxides ores.
Mixed oxide - artifically prepared by supporting finely
divided iron oxide on materials of large
surface and loose texture such as wood
shavings and granulated or crushed slags.
Mixed oxides have the following advantages:
Control of bulk density, iron oxide content, moisture
and pH.
-------
Reduced tendency to cake.
Free passage of gas.
Higher final sulfur loading
Capacity and activity are two important aspects in the
selection of purifying materials. In theory, 0.61* Ib of hydrogen
sulfide will react with 1 Ib of anhydrous ferric oxide, but only
up to about 0.56 Ib sulfur/lb ferric oxide has been achieved in
operations. Capacity decreases progressively with every cycle
after the first.
3-2 Liquid-phase Iron Oxide Processes
The chemistry involved in the liquid-phase processes is
based on the reaction of H S with either sodium carbonate or ammonia
and the subsequent reaction of the hydrosulfide formed with iron oxide.
Regeneration follows by.aerating the iron sulfide and converting it to
iron oxide and elemental sulfur. The reaction mechanism for the
process using iron oxide suspended in an alkaline aqueous solution
is shown by the following equations:
H2S + Na2C03 = NaHS + NaHCO
Fe20-3 . 3 H20 + 3 NaHS + 3 NaHCO = Fe2$ ' 3 H,0 +
3KJ A ^ O -i- O LJ ft
I"i a w \j i j n u
2 Fe2S3 ' 3 H20 + 3 02 = 2 Fe20 ' 3 H20 + 6 S
Several, mostly undesirable, side reactions occur depending
on the operating conditions and the composition of the gas to be treated.
Some thiosulfate formation is inevitable and it has been reported that
sometimes it may even be desirable to completely convert to thiosulfate
as follows:
93
-------
2 NaHS + 02 = Na2S2°3 * H2°
Na2S + 1-* 02 + S - Na2S203
Another undesirable side reaction occurs when hydrogen
cyanide absorbed in the alkaline material is converted to thiocyanate.
The mechanism first involves formation of sodium cyanide which is
then oxidized by elemental sulfur as follows:
HCN + Na2CO = NaCM + MaHCO
NaCN + S » NaSCN
Not all of the HCN is converted because a majority Is stripped from
solution by regeneration air.
Significant concentrations of hydrogen cyanide in the gas
can possibly lead to a mechanism in which the iron oxide reaction
with hydrogen sulfide is Inhibited. Noticeable color changes have been
observed in the solution when treating a gas in which hydrogen cyanide
approaches 10% of the hydrogen sulfide concentatfon. Oxidized solution
has a blue coloration due to the presence of ferric-ferrocyanide com-
plexes which become pale yellow in the fouled condition. While reaction
between iron oxide and hydrogen sulfide is quite slow, the blue
complexes support rapid conversion of hydrogen sulfide to sulfur. It
Is hypothesized that the reactions involve oxidation of H_S by con-
version of the ferric-ferrocyanide complex to ferrous ferrocyanfde.
Ferric-ferrocyanide Is reestablished in the regeneration step. The
reactions can be represented by the following equations:
2 H2S + Fe(t(Fe(CN)6) + 2 Na2CO -
2 Fe.Fe(CN), + Na, Fe(CN), + 2 H,0 + 2 CO, + 2 S
Z 0 *t o Z i
2 Fe2Fe(CN)6 -f Na^FefCN)^ 02 + 2 H2C03 -
Fe^(Fe(CN)6) + 2 Na^Oj + 2 HjO
-------
Iron, lost with the sulfur as iron cyanide compounds, is replenished
with a dissolved solution of iron sulfate.
3.3 Stretford Process
Hydrogen sulfide is absorbed into a solution consisting
mainly of sodium metavanadate, sodium anthraquinone disulfonate (ADA),
sodium carbonate and sodium bicarbonate in water. Sodium carbonate
reacts with the H_S to produce sodium hydrosulfide as follows:
H2S + Na2CO. *• NaHS + NaHCO-3
Free sulfur is formed by oxidation with sodium metavanadate. Vanadium
in this reaction is reduced as shown:
HS" + 2 V5+ +• 2V + S + H+
The full reaction equation is:
2 NaHS + *» NaVO- + H20 +• Na2V^°9 + 2 S + ** NaOH
Vanadium is reoxidized by reacting with ADA in the oxidizer as follows:
2 \l^+ ADA (oxidized) *• 2V5+ + ADA (reduced)
or:
Na^Og + 2 NaOH + H20 + 2 ADA *k NaVO + 2 ADA (reduced)
Unused caustic (k moles formed in the vanadium reduction equation and
only 2 moles are consumed in the previous equation) reacts with sodium
b icarbonate:
2(NaOH + NaHCO. »• Na2C03 + H20)
The reduced form of ADA is oxidized by air blowing:
2 ADA (reduced) + 02 *• 2 ADA + 2 H20
Overall, the above combination of reactions reduces to:
2 H2S + 02 *• 2 S + 2 H20
Side reactions include oxidation of sodium hydrosulfide to thiosulfate,
conversion of hydrogen cyanide to sodium thiocyanate, and formation of
sodium sulfite and sulfate from SO..
-------
2 NaHS + 2
2 HCN + 2 NaHS + QZ *• 2 NaCNS + 2
S02 + 2 Na2C03 + H20 ^ Na2$03 + 2
2 Na2S03 -i- 02 ^2 Na^O^
-------
C.A Design Basis and Assumptions
As a comparison of requirements to remove hydrogen sulfide
from product gas of a typical gasifier, the following calculation
basis and assumptions have been made:
q
1. Ax 10 Btu/day capacity of coal gasification plant.
2. Heating value of gasification product gas is 175 Btu/SCF.
3- Assume particulate and tars in the product gas have
been removed before the gas enters into a hydrogen
sulfide removal system.
A. Use 3.0 W % sulfur coal .
5- Assume all sulfur in coal appears mole for mole in
the product gas (conservative assumption).
6. Assume thermal efficiency of a gasification process
is 80%.
7. Assume heating value of coal is 12,500 Btu/lb.
Calculation of sulfur content of product gas:
12,500 Btu/lb of coal x 80% = 10,000 Btu in product gas/lb coal
k,000,000,000/10,000 = AOO.OOO Ibs of coal/day
= 200.0 tons of coal/day
10,000/175 = 57-1*» SCF product gas/lb of coal
0.03 lb sulfur in coal/32 Ib/mole = 0.0009375 moles sulfur
compounds in gas (assuming
100% recovery of sulfur in
coal)
= 0.3562 SCF of sulfur com-
pounds in gas per lb of
feed coal
= 0.619 V % concentration of
sulfur compounds in product
gas
= 6190 ppm by volume
97
-------
k,000,000,000/175 - 22,857,000 SCFD of product gas
- 952,380 SCFH
0.03 1b sulfur x 7000 grains/lb « 210 grains sulfur compound
in gas per 57 SCF product gas
» 368 grains/100 SCF of unpuri-
fied product gas
98
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C.5 Iron Oxide Box Purifiers
5.1 Design Based on American Practice
Bed Size Estimation -
Empirical rules, for the most part, provide the process
design method for iron oxide bed purifiers. The Steere Engineering
Co. has proposed the most commonly used method in the United States,
The formula is expressed by the following equation:
GS
3,000 ( D + C )
where:
A = cross-sectional area of gas flow through any one box,
in series, of a set.
G = maximum gas rate, SCF/hr.
S = correction factor determined by the hydrogen sulfide
content of the inlet gas.
D = total depth of oxide in feet through which the gas will
pass in the purifier set. In split flow designs where
half the gas volume passes through each layer, the gas
flow area is twice the cross-sectional area of the box,
while D is the depth of one layer of oxide.
C = factor determined by the number of boxes as follows:
A for 2 boxes; 8 for 3 boxes; and 10 for k boxes.
The values of the S correction factor are:
Grains H2S/100 SCF
of Unpurified Gas Factor
1000 or more 720
900 700
800 675
700 6*»0
600 600
500 560
kOO 525
300 500
200 or less
99
-------
The influence of H2S loading may be observed from the graphical pre-
sentation of this data in Figure C-5.1.
Estimated Box Purifier Design -
Detailed calculations are given in the Appendix.
No. of iron oxide boxes 10
Unit Size: Square 37 ft x 37 ft x 10 ft high
Circular k\ ft I.D. x 10 ft high
Gas Volume (Inlet) 22.86 million cu ft/day
(4 x 109 Btu/day)
1*2$ Content (inlet) 368 grains/100 SCF gas
(Outlet) 29 grains/100 SCF gas
5.2 Design Based on European Practice
Bed Size Estimation -
The space velocity through one box, known at the R ratio,
governs European oxide-purifier design. Selected values of R may
be applied in the space velocity equation given below to design
box size.
_ cubic feet of gas per hour
R
cubic feet of oxide in one box
R values can be varied from 15 for a conservative design to as high
as 100 for some installations. Good design practice for systems
operating at essentially atmospheric pressure and with revivification
in situ use an R ratio of 20 to 50.
Minimum standards of the oxide-purifier design require an
oxide bed to be at least 10 feet deep to produce sufficient pressure
drop for proper gas distribution over the entire cross-sectional
area. Also, vessel diameter should limit deposition to a maximum
of 15 grains/square foot of cross-sectional area of bed per minute.
100
-------
K-C '0 * '0 TO THE CENTIMETER 16 x Jb CM
*C KEUFFEL & ESSER CO. M»OE IN U 5 »
46 1513
L...L- -
-------
Estimated Box Purifier Design -
Detailed calculations are given in the Appendix.
No. of iron oxide purifiers k
Unit Size: Square 35 ft x 35 ft x 10 ft high
Circular 39 ft I.D. x 10 ft high
R Factor 20
Gas Volume (inlet) 22.86 million cu ft/day
(*» x 109 Btu/day)
H2S Content (inlet) 368 grains/100 SCF gas
(Outlet) 29 grains/100 SCF gas
5-3 Discussion: Dry Iron Oxide Process
The iron oxide box purification process provides reliable
and effective hydrogen sulfide removal. Its space requirement, however,
may be excessive and the operation produces a considerable quantity of
spent iron oxide (when revivification is not effective) which must be
disposed. The limiting factor in the iron oxide process is that it
requires a large ground space. This may rule out many large existing
plants.
Hydrogen sulfide removal in iron oxide box purification is
a surface reaction. Dust, light oils, naphthalene, and tars should
be removed before purification since these materials coat the iron
oxide and render it unreactive. For optimum operation of the process,
the gas should not contain more than 0.4 grain per 100 cubic feet of
tars, oils and dust. This low tar loading requirement may limit the
application of the iron oxide purification process to some gasification
processes since they produce significant amounts (about 10 W %) of
tars and light oils per pound of feed coal.
102
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5.** Advantages and Disadvantages of the Dry Iron Box
Process
Advantages -
1. Completely removes small to medium concentrations of
hydrogen sulfide without removing carbon dioxide.
2. Equally effective at any operating pressure.
3- Removes mercaptans or converts them to disulfides.
k. Produces a final product in elemental sulfur form.
5. Well proven commerical process.
Disadvantages -
1. A batch process, requiring duplicate installations or
flow interruption of process gas.
2. Requires large ground space for large gas plants.
3. Hydrogen cyanide reacts irreversibly with iron oxide
causing a loss of purifying material.
k. The iron oxide sponge bed may become coated with en-
trained oil, tar or distillates and require more
frequent changing. Thus, it may be necessary to
wash the gas with oil (e.g., benzene scrubber) before
iron oxide box purification.
5- High labor cost.
6. Disposal of a large quantity of iron oxide sponge
from spent beds is necessary.
103
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C.6 Liquid-phase Iron Oxide Processes
6.1 Process Description
The schematic flow diagram of a liquid-phase iron oxide
process is shown in Figure C-6.1. The system consists of a packed
column (absorber) and a regenerator. Scrubbing solution, normally
containing 3.0 W % Na.CO. and 0.5 W % ferric hydroxide, is pumped
to the top of the absorber where it is counter-currently contacted
with the raw product gas fed into the bottom of the vessel. The
liquid is circulated at such a rate that a two to threefold excess
of ferric hydroxide over the stoichiometric quantity necessary for
the complete reaction with hydrogen sulfide is present.
The hydrogen sulfide-contain ing solution flows from the
bottom of the absorber to the regenerator. Elemental sulfur, formed
in the regenerator by contact of the solution with air, accumulates
as a froth on the liquid surface, flows to the slurry tank, and is
pumped from there to a filter where excess liquid is removed. A
part or all of this liquid may be discarded, thereby purging undesir-
able salts from the system. Oxidized solution is pumped from the
regenerator back to the absorber to complete the cycle. Foul gas
from the oxidizer is vented to either a boiler or a washer.
6.2 Process Requirements
The estimated sizes of the equipment for treating k x 10^ Btu
per day of low-Btu gas (175 Btu/SCF) are:
Packed Tower -
Packed Column: 9 ft 3 in I.D. x 50 ft overall height
Packing Height: 26 ft
Packing Material: 2" Intalox saddles with Fp -
-------
FIGURE C-6.1
FLOW DIAGRAM OF LIQUID IRON OXIDE PROCESS FOR H9S REMOVAL
Absorber
Vent
Purified Product Gas
Mist
Eliminator
Liquid
Distributor
Liquid
Distributor
packing
Raw _
Product"1*
Gas
From
Gasifier
Regenerator
f////W/S
A/NAA
\
/
X
/
X
u u u u
\
x
\
/
/
— —
\
/
\
\
-~-y
H20
Na
-£
^
iCQi
ZUU3
1
Fe(
r
Makeup
Chemical s
Storage
Tank
•^
i
i
— (•
i
CD
T3
0)
I/I
V)
O
C
o
'
")
•**.
^ /
\J W
/
y
(
_ ^
Air
Filtrate
Storage
Tank
Recr red
Su r
-{X—
-------
Regenerator -
Total Volume Required: 777 ft3, i.e., 7-0 ft x 20 ft
high (for example)
Pressure Drop Through Packed Tower -
0.5 in H20 per ft of packed bed
Flow Rates -
Product Gas: 952,380 SCFH or 22.86 x 106 SCFD
Solution Circulation Rate: 165 Ib/sec or 1161 gpm
Air Requirement: 815 SCFM (2x stoichiometric amount)
Detailed calculations are shown in the Appendix.
6.3 Advantages and Disadvantages of the Liquid-phase Iron
Oxide Process
Advantages -
1. Continuous gas clean-up process, using inexpensive
chemicals.
2. Small ground space requirement.
3- Flexible to accommodate the variations of hydrogen
sulfide content of the raw gas. Easy to control
the process operation.
k. Selective removal of H2$ from C02-
5. Operates over wide pressure ranges.
6. Low labor cost; little supervision required.
7. Elemental sulfur is produced as a final product.
8. Well proven commercial process.
Disadvantages -
1. Scrubbing chemicals react with HCN irreversibly to pro-
duce thiocyanates and to form thiosulfates by side
reactions, causing a loss of active purifying material.
2. Requires pumping a large amount of recycle liquid to
the absorber and air to the regenerator.
3. H£S is not always completely removed; however, the I^S
removal efficiency of the process is adequate to meet
projected EPA air pollution control guidelines for
industrial gas users.
106
-------
k. Like any other liquid-phase h^S removal process,
the corrosion problems must be surmounted.
5. Bleed-off liquid streams must be treated before
discharge. The streams contain sodium thiocyanate,
thiosulfates, sodium carbonates and ferric hydroxide.
6. Does not remove most organic sulfur compounds.
7. Operates at low temperatures, ambient to 100°F.
107
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C.7 The Stretford Process
7-1 Process Description
As shown in Figure C-7-1, feed gas to be purified is scrubbed
in an absorber by a counter-current flow of an alkaline solution at
about 80°F to 100°F. The solution, containing a vanadium salt along
with an anthraquinone derivative, oxidizes H2$ into elemental sulfur
while the vanadium is reduced. To insure complete precipitation of
sulfur, a delay tank beneath the absorber retains the Stretford solu-
tion for 10 to 20 minutes. The solution reaches an equilibrium with
respect to the carbon dioxide in the gas and only relatively small
amounts of CO- are removed by the process.
Liquor from the absorber is fed to oxidizers to restore
vanadium to the oxidized form through a mechanism involving oxygen
transfer via the anthraquinone derivative. Air blown through the
oxidizer also separates sulfur by froth flotation. The scum pro-
duced is either filtered or centrifuged, washed, and melted into
high-quality sulfur.
7-2 Process Requirements
H2S removed from raw product gas to meet the study
basis guideline (952,380 SCF/hr gas, 3-68 gr H2S/SCF,
S2% removal):
0.128 Ib H2S/sec
Stoichiometric Na.CO. requirement to convert H.S
CO- + H2S - NaHS + NaHCO.):
3.U8 Ib Na2CO-/lb H2$
or
0.399 Ib Na2C03/sec
108
-------
Absorber
.*~ Purified Gas
Liquid
Distributor
packing
Raw
Product
Gas
Vent
Reaction
Tank
Oxidfzer
(Regenerator) Sulfur Foul
Troth Air
Regenerated
Solution
Makeup
Chemical
tl
Makeup
.Water
Pump
Surge
Tank
I
v v v/
p
I iItrate
ank
Air
Discharge
Sulfur
Cake
FIGURE C-7.1
FLOW DIAGRAM OF STRETFORD PROCESS
-------
Practical Na_CO_ requirement (usually two to three times
theoretical; using 3):
1.197 lb Na2CO-/sec
Amount of scrubbing solution required (sodium carbonate:
0.1 N; sodium bicarbonate: 0.3 N, or equivalent Na~CO_
solution of 0.505% by weight):
237.1 lb solution/sec
or
1625 gpm @ 65.5 lb/ft3
Process Unit Summary -
Raw Gas Rate: 22.86 x 10^ cu ft/day
H2S in, grains/100 SCF: 368
H-S out, grains/100 SCF: 29 (to meet EPA guidelines
of 0.5 Ibs S02 per 10^ Btu)
Absorber
Overall Height
Diameter
Packing Type
Packing Height
Reaction (Holding) Tank
Circulation Pump
Oxidizer
Air Blower
Filter
85 ft (about)
10 ft I.D.
2" saddles with Ff
2 sections each H
ko
.5 ft
Bottom section of absorber
27 ft high
2 Units (one stand-by),
1625 gpm
14 ft 6 in I.D. x 20 ft high
(25 ft overall height)
60,000 ct ft/hr
Vacuum Rotary
Detailed calculation estimates are given in the Appendix.
110
-------
7-3 Advantages and Disadvantages of the Stretford Process
Advantages -
1. Well proven commercial process.
2. H^S removal to below 1 ppm possible.
3. Can recover H.-S as pure saleable sulfur.
^4. Insensitive to h^S/CO,, ratio.
5. Operates over wide pressure ranges (0 psig - 1000 psig)
6. Accepts process fluctuations.
7. Primarily mild steel construction.
8. Little supervision and maintenance required.
Disadvantages -
1. Does not remove most organic compounds.
2. Requires preprocessing for feeds which contain large
quantititles of SO™, HCN or heavy hydrocarbons.
3. Produces a large purge stream containing a vanadium
compound, ADA, thiocyanates and thiosulfates.
!*. Operates at low temperatures (ambient to 120 F) .
5. Probably less economic for treating streams with an
acid gas concentration greater than 15% H«S than some
other processes.
Ill
-------
SECTION D - OPERATIONAL EVALUATION OF CONVERTER OUTPUT
CONTROL SYSTEMS
1. Typical Clean-up Systems Applied to Industrial
Fuel
2. Dependency of Clean-up on End Use of the Fuel
Gas
3. Sulfur Emission Control with an Industrial
Fuel
k. Effect of Nitrogen Compounds on Chemical
Clean-up Systems
5. Tar and Oil By-products
6. Reduction of Particulates for Industrial
Fuels
112
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D.1 Typical Clean-up Systems Applied to Industrial Fuel
The clean-up systems covered in Section B illustrate
representative applications for each of the particular gasifier
systems examined. Some utilize simple clean-up techniques and
others are complex, but basically they all have a means to quench
or cool the gas and knock out the particulate matter. Further
steps required depend on the type of coal being gasified, the
process conditions, and the end use of the gas.
The same considerations are, for the most part, necessary
for industrial gas as for synthesis gas. Particulates and tars will
have to be removed. Gas cooling, along with waste heat recovery
for energy efficiency, would be appropriate. Sulfur and nitrogen
removal may be necessary, although the reduction in sulfur and
nitrogen levels need not be complete to reach the proposed EPA
guidelines on industrial gas. These levels as suggested by the
EPA are:
Sulfur - 0.5 Ibs of SO,, per 10 Btu.
Nitrogen - 0.1* Ibs of NO per 106 Btu (as NOJ.
x 6
Particulates - 0.1 Ibs per 10 Btu.
The sulfur limit may be reached through the use of the processes
already discussed. Nitrogen in the form of gaseous compounds is,
more or less, controlled by sulfur removal processes as discussed
later. Organically bound nitrogen, in the tars and oils knocked
out in the particulate and tar removal portion of the gasification
systems must be given consideration to prevent pollution and this
aspect is also considered later.
113
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Carbon dioxide removal, which is required in the manufacture
of pipeline gas, is not necessary for industrial fuel gas. The lower
heating value of the fuel as a result of leaving carbon dioxide in
the industrial gas is not likely to adversely affect most applications.
In fact, a beneficial effect due to the lower flame temperature would
result from a decrease in the thermal fixation of nitrogen. Of course,
if a high concentration of nitrogen compounds such as hydrogen cyanide
or ammonia were contained in the gas, they might have to be controlled
to prevent the formation of NO .
X
Considering the point that the subject of this outline is
control technology for industrial fuel, the ability of the illustrated
systems to clean-up gas for the high standards necessary to make pipe-
line or synthesis gas suggests the possibility that the more tolerant
requirements for industrial gas can be more easily met. The simpler
systems appear to be directly applicable for industrial gas clean-up
but, where inadequate, the complex systems would be able to provide
the necessary performance with a relaxation of design specifications
for the clean-up system. Examples of this decreased demand on acid
gas clean-up requirements for industrial gas manufacture may be con-
veniently illustrated by calculating the effect of changed operating
conditions on H2$ removal with a cold methanol wash. The following
three cases are estimates of an H2$-bearing gas flowing in contact
and at equilibrium with a cold methanol wash.
Case
1
3
2
Pressure
psia
500
500
14.7
Temperature
°F
-80
-60
-25
Approximate
Methanol Circulation
moles/mole H2S
3200
3200
130
Gas H2S
Level
ppm
0.1
0.25
520
-------
The first case shows the methanol circulation ratio that is necessary
per mole of absorbed H.S in order to establish equilibrium with a
gas containing 0.1 ppm of hLS. To clean-up a gas for methanation,
acid gas removal down to 0.1 ppm level would be appropriate. The
effect of a refrigeration loss, as might occur due to a process upset
from the loss of a compressor, would be to raise the gas hLS content
above the desired level. Case 2 illustrates the equilibrium H-S
level attained in the gas for an arbitrary was temperature rise to
-60°F. While this might be an undesirable level for a methanation
feed gas, it would still be more than adequate for an industrial
fuel gas. Case 3 is, then, an example of how acid gas removal design
criteria could be relaxed to deliver a product satisfactory for indust-
rial fuel gas as compared to methanation feed gas. A 175 Btu per
SCF higher heating value industrial fuel gas with a sulfur level of
0.5 Ibs of S02 per million Btu would have about 520 ppm of HLS. A
methanol wash at atmospheric pressure and -25°F would be in equilibrium
with this gas using a circulation of 130 moles per mole of absorbed
H2S.
Adapting high performance systems, capable of removing
pollutants to low levels, for industrial gas clean-up may be accomp-
lished in two ways. Naturally, one could clean up a fraction of
the gas stream while bypassing the remainder in an amount such that
the recombined stream would meet guidelines for an industrial fuel
gas. Although such a procedure would require a somewhat smaller
capacity clean-up unit (compared to treating the whole stream for
synthesis), there is a good possiblity that a process design for
treating the entire flow to meet requirements for an industrial gas
might be advantageous. Such a consideration should be given a thorough
evaluation by the designer contemplating a unit for industrial fuel
manufacture. Not only is there a whole new set of design parameters for
clean-up processes, but many older processes might once again be con-
s idered.
15
-------
D.2 Dependency of Clean-up Systems on the End Use of the Fuel Gas
In the final analysis, the clean-up methods used to condition
the raw-gas from a gasifier are more a function of the end use for the
fuel gas than it is of a particular gasification system. For example,
the firing of a hot, dusty, sulfur-laden fuel gas into a cement kiln
should be quite satisfactory because the cement-making process itself
requires a dust control system and the limestone in the kiln feed will
remove substantially all of the sulfur in the fuel. In a sense, the
process provides the gas cleaning step on the raw fuel-gas and a gas
of high purity is not required.
An example of an end use that is intolerant of dust is the
combustion of coal-derived gas in a gas-turbine. For this application,
the fuel must be extremely clean, in the order of one ppm by weight
of solids. However, complete removal of sulfur is not required and
the degree of sulfur removal would be dependent on local regulations
governing S0» emissions.
An example of a gas requiring an intermediate degree of
cleanliness with respect to dust is the mixture of coke-oven gas and
blast-furnace gas that is normally used in steel mills to fire soaking
pits, reheating furnaces, coke-ovens, etc. This gas is normally cleaned
to a solids level of 15-30 ppm (by weight). Experience has shown that
this level of cleanliness is satisfactory for distribution of gas
throughout the steel mill.
Any gas containing hydrogen and carbon oxides that is to
be methanated (perhaps for Btu adjustments) using a nickel catalyst
must be essentially sulfur-free or contain no more than 0.1 ppm
(volume) of sulfur-bearing compounds. As mentioned earlier, some
116
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industrial gas is burned without sulfur removal; hence, sulfur removal
requirements can vary from no removal to complete removal in the
event that catalytic processing is involved. Such a wide variance
demonstrates quite clearly that the end use of a fuel-gas determines
the degree of clean-up that is required provided, of course, that
the final emissions meet the required EPA guidelines. This example,
plus the three cases cited previously on differences in dust levels
in fuel gases, confirms the importance of end use of the gas vis-a-vis
the clean-up system(s) used on a given gasifier effluent.
The systems studied in Section B show this end use dependency
For the Wilputte unit, a very simple system is described because low-
sulfur coal is used and the product gas is then immediately applied
as a fuel. The end use in the WD/GI example, while also intended
for direct use as a fuel, would tolerate medium-sulfur coals because
the Stretford process was employed to remove sulfur. The Koppers-
Totzek and Lurgi cases make use of the Rectisol process because fur-
ther processing into ammonia and pipeline gas required a high degree
of clean-up.
117
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D.3 Sulfur Emission Control with an Industrial Fuel
The easiest way to limit the amount of sulfur dioxide
formed when combusting a coal -derived fuel gas is, perhaps, to
gasify a coal sufficiently low in sulfur that the product gas,
when burned, will conform to the guideline limit of 0.5 Ibs of
sulfur per 10 Btu. Unfortunately, this method would require a
feed coal containing approximately 0.25% sulfur and such coals are
too rare to be of real significance. Nevertheless, the gasification
of a low-sulfur fuel such as wood or wood refuse does produce a fuel
gas low in sulfur, and such plants have been built and operated
successfully for commercial ventures in Africa and elsewhere.
The use of low sulfur fuels such as wood is recognized
by the EPA. Recently an amendment to the New Source Performance
Standards was made which will affect large steam generators by
allowing wood residue as a fuel supplement. The heat content of
the wood residue (defined as bark, sawdust, slabs, chips, shavings,
mill trim and other wood products) would be used for determining
compliance with the standards so long as there is no increase in
sulfur or nitrogen oxide emissions as a result. The impact of
the amendment on particulate emissions has not yet been defined,
and information on this subject is currently being gathered. The
amendment appeared in the Federal Register on November 22, 1976
FR 51397).
In Section C, some old clean-up systems used primarily for
the removal of sulfur are examined and compared to a modern process.
While these old systems may not compete when synthesis gas standards
are required, a design is presented for comparison purposes in the
manufacture of industrial gas. Depending on the particular situation
118
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under consideration, it appears from equipment requirements that
the old processes might be economically attractive for some indus-
trial gas applications.
The most practical way to remove sulfur (mostly H«S) from
coal-derived fuel gas is, of course, to use a one-step process that
directly converts the hLS to elemental sulfur. Section C gives
examples of some one-step processes. Two-step processes that
selectively remove a stream rich in H?S for subsequent processing
to form sulfur in a Claus-type unit are also commonly used. These
two-step operations are more complex, and the tail gas from a Claus
unit introduces an additional pollution stream to be cleaned-up.
An idea of the difference in the complexity of the one-step versus
the two-step systems can be noted in schematic flowsheets shown in
Figure D-3-1. Nevertheless, the economics of two-step processing
might be suitable for a large installation.
The removal of sulfur from coke-oven gas (which can be
considered a special type of industrial fuel) has historically
favored the one-step systems. The older liquid-phase sulfur removal
systems, such as Thylox, Ferrox and Manchester, and the solid-phase
iron oxide boxes are one-step processes that were so used. More
recently developed one-step processes that are in use removing sulfur
from coke-oven gases are:
Stretford
Holmes-Stretford
G i amma rco-Ve t rocoke
Takahax.
The Stretford process is also being applied with modern
processes. In Japan, H,>S is being removed from large volumes of
low-Btu gas produced in the operation of a Flexicoker.
119
-------
FIGURE D-3.1
FUEL GAS DESULFURIZATION SYSTEM SCHEMATIC DIAGRAMS
A. ONE-STEP SYSTEM
Purified Gas
Impure
Vent
Gas
Rich Solution
B. TWO-STEP SYSTEM
Purified Gas
Impure
Sour Gas
Gas
Claus Unit
Reducing Gas
Reactor
\
I
Heater
Off-Gas
To Incinerator
Cool ing
Tower.
Lean
Am me
Absorber
I
Sulfur
Stripper
Rich Amine
Condensate to
Sour Water Stripper
120
-------
D. k Effect of Nitrogen Compounds on Chemical Clean-up Systems
The nitrogen content of coal, which generally runs in the
range of 0.6 to 2.0 percent, can lead to problems when coal is gasi-
fied. Not only is there a risk of excessive environmental contamin-
ation, but an impact on sulfur clean-up processes may also be found.
During gasification nitrogen compounds such as ammonia, cyanides,
oxides of nitrogen and pyridenes may, depending on conditions, be
included in the raw product gas. High temperature, for example,
favors HCN but no pyridene formation.
Unfortunately, hydrogen cyanide reacts in most chemical
clean-up processes to form undesirable chemicals which precipitate
out of the treatment solution or are removed by purging a part of the
circulating stream. In the Stretford process, almost all of the HCN
will be converted to sodium thiocyanate as shown in the following
equations:
HCN + Na-CO > NaCN + NaHCO
NaCN + NaHS + 1/2 0^ > NaCNS + NaOH
NaHC03 + NaOH—yNa^Oj + H20
or overal1:
HCN + NaHS + 1/2 0^ >NaCNS + H20
The effluent stream purged from the process, containing thiocyanate,
thiosulfate, vanadium and anthraquinone disulfonic acid, has in the
past been considered an innocuous stream for disposal, but now, to
meet current requirements, other methods such as recovery or biodegrad-
at ion must be evaluated. In the MEA treating process, HCN can react
and partially deactivate the solution. Activity of the solution can
be reclaimed by treatment with sodium carbonate or caustic soda.
Activated hot potassium carbonate almost completely absorbs HCN.
However, the hot potassium carbonate solution is not degraded by HCN
121
-------
since it is removed from the solution along with C0_ and KLS.
When high levels of HCN are encountered, removal from the
process stream may be necessary by scrubbing with a polysulfide
solution. Scrubber efficiencies of over 37% can be achieved, and
the HCN content of the gas can be reduced to under 30 ppm by volume.
Disposal of the sodium thiocyanate, formed in the reaction between
HCN and polysulfide, is necessary. As indicated for the purge solu-
tion in the previous paragraph, alternates to discarding this chemi-
cal stream must now be considered.
Ammonia dissolves in Stretford solution and can be stripped
from solution during regeneration. When feed gas concentrations are
above approximately 300 ppm, special precautions are necessary to
prevent ammonia from being released to the atmosphere by the stripping
action of regeneration air. The basic nature of ammonia makes it
amenable to removal from gasifier gas in those instances of excessive
ammonia concentrations. Removing ammonia from coke-oven gas by water
or acid scrubbing has been the practice historically. Pyridene bases
may also be scrubbed out of the gas with an acid absorbent.
122
-------
D.5 Tar and Oil By-products
The tar/oil yields on coal gasification systems are
determined, to a large extent, by the intrinsic nature of the gas-
ification system. Moving-bed counter-flow systems, such as the Lurgi,
WD/Gl and Wilputte gasifiers, all produce an appreciable amount of by-
product liquids. On the other hand, high temperature entrained co-
current gasifiers, such as Koppers-Totzek, produce no tars. The yield
of these tars from the counter-flow systems, including the light-oils,
represents about 4-12 wt % of the coal feed. Because the heating
value of the tar/oils is higher than the heating value of the coal,
the percentage of the heating value of the feed coal appearing in
the tar/oil is even greater.
If the tar/oil is to be burned as a fuel, there may be
problems in meeting sulfur, nitrogen and particulate specifications.
Tars produced by gasification processes generally have a reduced
sulfur content compared to the feed coal but have about the same
nitrogen contents as the coal. A typical tar from a gasification
system having a heating value of 16,500 Btu/lb requires a maximum
sulfur content of 0.66 wt % to meet a specification of 0.8 Ibs of
S07 per 10 Btu. Such sulfur contents are typical of tars produced
from low to medium sulfur coals, but tars produced from high sulfur
coals may have trouble meeting an emission level of 0.8 Ibs of S00
g t-
per 10 Btu when burned.
Information on the properties of tars produced by various
gasfication systems is limited. Therefore, it is appropriate to review
data from systems or equipment that simulate the conditions in a moving-
bed counter-flow gasifier. Such a simulation is found in assay tests.
123
-------
Assay tests for coal that distill (carbonize) coal give a clue to the
sulfur content of tars as related to the sulfur content of the coal.
Assay tests on a number of different coals indicate that
the sulfur content of the tar is invariably less than for the coal.
These data are listed below.
DISTILLATION/CARBONIZATION ASSAY
Coal Source
Pennsylvania
Kentucky
V i rg i n i a
Maryland
Alabama
Pennsylvania
111inois
British Columbia
Pennsylvania
Alabama
West Virginia
West Virginia
West Virginia
Alabama
Utah
Pennsylvania
Some data has been found on actual gasification systems
and the data tabulated below is derived from the Sythane gasification
of coal (fluid-bed pressurized gasification).
(Bureau of Mines - Monograph 5)
1 S in
Tar Product
0.65
0.5
0.5
0.75
0.7
0.7
0.5
a 0.55
0.6
0.6
0.5
0.55
0.85
0.4
0.6
0.6
% S in
Coal Feed
1.1
0.6
0.6
1.5
0.8
1.0
0.8
0.6
1.3
1 .1
0.6
0.9
1.8
0.7
1.0
1.3
S in Tar
as % of
S in Coal
59
83
83
50
87
70
62
92
46
55
83
61
47
57
60
46
124
-------
SYNTHANE GASIFICATION
Coal Source
Pittsburgh Seam
111 inofs No. 6
Montana Sub-bit.
N. Dakota Lignite
% S in
Tar Product
0.8
2.7
0.5
1.0
% S in
Coal Feed
1.5
3.5
0.6
1.1
S in Tar
as % of
S in Coal
53
77
83
91
As can be seen from all data, the sulfur content of the
tar is appreciably lower than for the coal. In view of this in-
formation, it is reasonable to conclude that tars produced by any
gasification process generally would have a reduced sulfur content
compared to the feed coal.
Nitrogen poses a somewhat more difficult problem. The
Clean Air Act of 1971 specifies that new, large boilers have an
emission level for NO of less than 0.3 Ibs. per 10 Btu when firing
X
liquid fuels. This level may be difficult tomeetwhen firing tars
produced from gasifiers. Most American coals contain 0.6-2.03; nitrogen,
and the tars produced from gasifiers will have about the same amount
of nitrogen.
The nitrogen contents of tars produced from most coal
processing systems appear to be similar, regardless of the exact
processing mechanism. This can be see from the comparison noted
below.
125
-------
Type of Processing % N In Coal
Pressurized Fluid-Bed
(Synthane)
1.1
% N in
By-Product Tar
1.1
In Situ Underground
Gasification with Air
(Hanna, Wyoming)
0.6-0.7
0.7*1-0.79
Carbonization of
111inois Coal @
1000°C (U.S. Bureau
of Mines
\.k
1.1
When the tars are combusted, a portion of this fuel-bound
nitrogen is converted to NO and, from experiments, the general
X
relationship between NO and the fuel nitrogen has been determined
as shown in the graph below.
TJ
x c
o 3
•Z. O
CQ
cn i
— «
ut 3
m LL.
-------
One source of data, based on the combustion of fuel oil,
indicates that a fuel nitrogen content in the order of 0.35 wt %
would be required to meet an emission level of 0.3 Ibs of MO per
6 x
10 Btu. This is about one-third of the nitrogen content of typical
tars produced from gasification systems. If this data is assumed to
apply to the combustion of tar (a reasonable assumption), then tars
produced in any large scale gasification project may require a hydro-
genation pretreatment for the partial removal of nitrogen before the
tars are burned. Such a pretreatment would also remove part of the
sulfur.
An alternate disposal method would involve blending the
tar into a low-nitrogen, low-sulfur fuel-oil so that the mixture, when
burned, would produce pollutants at acceptable levels.
The sale of by-product tar to recover valuable chemicals
obviates, of course, all of the problems associated with the combustion
of tar. Such an ideal route for the disposal of tar may not always be
avai1able.
The combustion of by-product tars has been practiced by
industry for many years. Special attention must be paid to the burner
design because the tars will frequently contain erosive ash particles
and other solids that can plug mechanical atomization devices. Stream-
atomized and rotary-cup burners are generally believed to be the most
trouble-free type of burner when combusting by-product tars.
127
-------
D.6
Reduction of Participates for Industrial Fuels
The previous sections have shown a great variety of devices
employed in the removal of particulates from gasifier output. In
most cases, a combination of such units is used depending on the gas
end use. Examples of such equipment are cyclones, spray washers,
coolers and waste heat boilers (which trap tars and/or particulates
in condensate), packed scrubbers, and sometimes (as in the case of
iron oxide boxes) the sulfur purification process. Usually, where very
low particulate levels must be achieved for applications such as
synthesis gas, additional high energy removal units are also neces-
sary. These applications might use disintegrators, venturi scrub-
bers or electrostatic precipitators.
Generally, data on the levels of particulate reduction
accomplished through each system are not covered in the source
literature. Koppers-Totzek information was an exception and parti-
culate reduction through the various units is reported as follows:
IJnit Operation
Gasifier
Wash-Cooler
Theisen-Irrigated
Disintegrators
Venturi Scrubber
Wet-type Electrostatic
Precipitator
Particulate Level
(grains/SCF)
12
1.2
0.002
0.002-0.003
0,0001
Remarks
50-70% slag leaves bottom
of gasifier; remainder is
entrained.
90$ removal of entrained
particles.
As currently recommended
by Koppers Co,
Possible alternate to
disintegrators.
For production of high-Btu
gas.
128
-------
In the examples presented for the other units, it may be inferred
that the participate removal systems utilized are adequate for the
discussed applications.
To put particulate removal necessary to meet EPA guidelines
for an industrial fuel gas in perspective, the 0.1 Ibs per 10 Btu
may be converted to 0.12 grains/SCF for a 175 Btu fuel gas. While
a definition of particulate removal operations necessary to meet
the guideline is not specifically covered in the literature, some
generalizations can be observed. Gasifier, waste heat boiler or
wash-cooler outlet particulate levels are higher than the limit,
but high energy systems are excessive for the purpose of making
industrial fuel gas. Possibly some combination of a cyclone, spray
washer and/or packed scrubber would, in some cases, provide the
necessary clean-up of particulates. In other cases, a venturi scrub-
ber design with decreased performance requirements compared to syn-
thesis gas might be appropriate. Depending on the Btu content of
the gas made and assuming that combustion of the fuel gas adds only
a negligible amount of particulates from soot formation, the clean-up
system that is selected will have to remove particulates down to
approximately 0.1 grain per SCF.
129
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REFERENCES
GENERAL
1. Bodle, W. W., and K. C. Vyas. Clean Fuel from Coal. Oil and Gas
Journal, August 26, 1974.
2. Boyd, N. F. Coal Conversion Processes Loom Big as a Source of
Hydrocarbon Fuels. Mining Engineering, September 197**.
3. C6EN Staff. Coal Gasification Development is Languishing. Chemical
6 Engineering News, November 1, 1976.
4. Fieldner, A. C., and J. D. Davis. Gas, Coke and Byproduct-Making
Properties of American Coals and Their Determination. American
Gas Association - 1934.
5. Musick, James K., and Fred W. Williams. Hopcalite Catalyst for
Catalytic Oxidation of Gases and Aerosols. Naval Research Laboratory,
Washington, DC 20375-
6. Perry, H. The Gasification of Coal. Scientific American, March 1974,
Volume 230, No. 3-
7. U. S. Steel. The Making, Shaping and Treating of Steel. Eighth Edition
1964.
KOPPERS-TOTZEK PROCESS
1. Farnsworth, J. F., D. M. Mitsak and J. F. Kamody. Clean Environment
with K-T Process. Presented at the EPA meeting on Environmental
Aspects of Fuel Conversion Technology, St. Louis, May 1974.
2. Farnsworth, J. F. Coal Gasification System Could Ease Energy Supply
Punch. 33 Magazine, The Magazine of Metal Producing, August 1973.
3. Farnsworth, J. F., et al. Production of Gas from Coal by the
Koppers-Totzek Process. Presented at Clean Fuels from Coal Symposium
at Institute of Gas Technology, Chicago, September 10-14, 1973.
4. Gas Manufactured. Kirk-Othmer Encyclopedia of Chemical Technology.
Second Edition. Volume 10. pp. 353-442.
5. Loeding, J. W., and J. G. Patel. Coal Gasification Review. Presented
at ASME Joint Power Generation Conference, Portland, Oregon, September
1975.
130
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KOPPERS-TOTZEK PROCESS (Continued)
6. Magee, E.M., C.E. Jahnig and H. Shaw. Evaluation of Pollution
Control in Fossil Fuel Conversion Processes; Gasification; Section 1:
Koppers-Totzek Process. PB-231-&75, EPA-SSO/Z-y^-OOSa, January 1971*.
7. Mitsak, D.M., J.F., Farnsworth and R. Wintrell. Economics of the
K-T Process. Koppers Company, Inc., Pittsburgh, August 6,
8. Mitsak, D.M., H.J. Michaels and J.F. Kamody. Koppers-Totzek
Economics and Inflation. Presented at the Third Annual International
Conference on Coal Gasification and Liquefaction, Pittsburg, August
1976.
9. Mitsak, D.M. Phone conversation with authors, March 17, 1977-
10. Partridge, L.J., South Africa Firm Gets Operating Exeprience with
1100-ton/day Coal-based Ammonia Plant. Oil and Gas Journal, November
2k, 1975.
11. Wintrell, R. The K-T Process; Koppers Commercially Proven Coal and
Multi-fuel Gasifier for Synthetic Gas Production in the Chemical and
Fertilizer Industries. Presented at AlChE National Meeting, Salt
Lake City, Utah, August 1971*.
LURGI PROCESS
1. Application of El Paso Natural Gas Co. before the Federal Power
Commission, Dated November 7, 1972.
2. Gallagher, J.T. The Lurgi Process State of the Art. Presented at
the Coal Gasification and Liquefaction Symposium, Pittsburgh, August
197**.
3. Hatten, J.L. Plant to Get Pipeline-Quality Gas from Coal. Oil and
Gas Journal, January 20, 1975.
k. Hendrickson, T.A. (Compiler). Synthetic Fuels Data Handbook.
Cameron Engineers, Inc. 1975-
5- Huebler, J. Coal Gasification: State of the Art. Heating/Piping/
Air Conditioning, T*9-155 (1973) January.
6. Loeding, J.W., and J.G. Patel. Coal Gasification Review. Presented
at ASME Joint Power Generation Conference, Portland, Oregon, September
1975
131
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LURGI PROCESS (Continued)
7. Lowry, H.H. (Editor). Chemistry of Coal Utilization, Supplementary
Volume, Chapter 20, John Wiley 6 Sons, Inc., New York, 1963.
8. Moe, J.M. SNG from Coal Via the Lurgi Gasification Process. Fluor
Engineers and Constructors, Inc., Los Angeles.
9. Robson, F.L., et at. Technology and Economic Feasibility of Advanced
Power Cycles and Methods of Producing Nonpolluting Fuels for Utility
Power Stations. Prepared by United Aircraft Corp. for National Air
Pollution Control Administration. Final report, December 1970.
10. Rudolph, P.F.H. Coal Gasification - A Key process for Coal Conversion.
Presented at Conference on Synthetic Hydrocarbons at the AI ME 1973
Annual Meeting, Dallas.
II. Rudolph, P.F.H. Gas from Coal. Chemical Economy 6 Engineering
Review, October 1973, Volume 5, No. 10 (No. 66).
12. Rudolph, P.F.H. The Lurgi Process The Route to SNG from Coal.
Presented at the Fourth Synthetic Pipeline Gas Symposium, Chicago,
October 1972.
13- Rudolph, P.F.H. The Lurgi Process Route Makes SNG from Coal. 90-92
(1973) January 22.
14. Shaw, H., and E.M. Magee. Evaluation of Pollution Control in Fossil
fuel Conversion Processes; Gasification; Section 1: Lurgi Process.
PB-237-634, EPA-650/2-74-009-C, July 1974.
RILEY MORGAN PROCESS
1. Kohl, A. and Riesenfeld, F., Gas Purification, 2nd Ed., 1974, Gulf
Publishing Co., Houston, Texas.
2. Private communication with Riley Stoker Co.
3- Rawdon, A.H., Lisauskas, R.A. and Johnson, S.A., "NO Formation in
Low and Intermediate Btu Coal Gas Turbulent-Diffusioft Flames",
The Proceedings of the NO Control Technology Seminar, San Francisco,
Ca., EPRI, February, 1976?
4. Walsh, T.F., The Riley-Morgan Gasifier. Presented at Third Annual
International Conference on Coal Gasification and Liquefaction,
Pittsburgh, PA, August 1976.
WELLMAN-GALUSHA PROCESS
1. Gas Generator Research and Development. Survey and Evaluation Phase
One, Volume 2, R & D Report No. 20. Prepared for Office of Coal Research,
Department of the Interior, March 1965.
132
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WELLMAN-GALUSHA PROCESS (Continued)
2. Hamilton, G.M. Gasification of Solid Fuels. Cost Eng. 8 (3),
4-11, 1963-
3. Hendrickson, T.A. (Compiler). Synthetic Fuels Data Handbook.
Cameron Engineers, Inc., 1975.
4. Mudge, L.K., et al. The Gasification of Coal. A Battelle Energy
Program Report, July 1974.
5. Robson, F.L., et al. Technological and Economic Feasibility of
Advance Power Cycles and Methods of Producing Nonpolluting Fuels
for Utility Power Stations. Prepared for U.S. Department of
Commerce, December 1970.
6. Wright, C.C., K.M. Borchay and R.F. Mitchell. Production of
Hydrogen and Synthesis Gas. Ind. Eng. Chem. 40 (4), 591-600 (1948).
7- Lowry, H.H. (Editor), Chemistry of Coal Utilization Supplementary Volume,
Chapter 20, John Wiley & Sons, Inc., New York, 1963.
WILPUTTE PROCESS
1. Wilputte Low-Btu and High-Btu Fuel Gas Process. Wilputte Co.,
Murray Hill, NJ, Bulletin No. 7762, 7763, June 1, 1976.
2. Private communication with Wilputte Co.
WINKLER PROCESS
1. Banchik, I.N. Power Gas from Coal Via the Winkler Process. Symposium
on Coal Gasification and Liquefaction, Pittsburgh, August 1974.
2. Banchik, I.N., T.K. Subramaniam and J.H. Marten. Pressure Reaction
Cuts Gasification Costs. Hydrocarbon Processing, March 1977.
3. Banchik, I.N. The Winkler Process for the Production of Low-Btu Gas
from Coal. Presented at Clean Fuels from Coal Symposium at Institute
of Gas Technology, Chicago, September 10-14, 1973-
4. Gasification of Solid Fuels in Germany by the Lurgi , Winkler and Leuna
Slagging-Type Gas-Producer Processes. Bureau of Mines Information
Circular No. 7415.
5- Gas Manufactured. Kirk-Othmer Encyclopedia of Chemical Technology,
Second Edition 10, pp. 353-442.
133
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WINKLER PROCESS (Continued)
6. Henderson, T. A. (Compiler). Synthetic Fuels Data Handbook. Cameron
Engineers, Inc., 1975.
7- Pressure Reactor Cuts Gasification Costs. Hydrocarbon Processing,
March 1977.
8. White, J. W., et al (Editor). Clean Fuels from Coal Symposium II
Papers. Sponsored by IGT, Chicago, Illinois, June 1975.
WOODALL-DUCKHAM/GAS INTEGRALE PROCESS
1. Brochure of Woodal1-Duckham Gasification Process. WD Engineering
and Construction Co.
2. Grant, A. J. Application of the Woodal1-Duckham Two-Stage Coal
Gasification. Presented at the Third Annual International Conference,
Pittsburgh, PA, August 1976-
3. Hendrickson, T. A. (Compiler). Synthetic Fuels Data Handbook. Cameron
Engineers, Inc., 1975-
4. Private communication with Woodal1-Duckham Co.
IRON OXIDE PROCESSES
1. Duckworth, G. L., and J. H. Geddes. Oil and Gas Journal September 13
1965, P. 94. '
2. Gas Engineers Handbook. McGraw-Hill, 1934.
3. Kohl, A. L., and F. C. Riesenfeld. Gas Purification, Chapter 8,
2nd Edition. McGraw-Hill Book Co., Inc., New York, 1974.
4. Perry, R. H., and C. H. Chilton. Perry's Chemical Engineers' Handbook,
5th Edition. McGraw-Hill Book Co., Inc., New York, 1973, pp. 18-21 to
18-23.
5. Peters, M. S., and K. D. Timmerhaus. Plant Design and Economics for
Chemical Engineers, 2nd Edition. McGraw-Hill Book Co., Inc., 1968,
p. 647.
6. Zapffe, F. Oil and Gas Journal, September 10, 1962, p. 136.
134
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STRETFORD PROCESS
1. British Gas Corp. Hydrocarbon Processing, April 1975.
2. Moyes, A. J., and S. Vasan. Holmes-Stretford h^S-Removal Process
Proved in Use. Oil and Gas Journal, September 2, 1974.
3. Riesenfeld, F. C., and A. L. Kohl. Gas Purification, 2nd Edition.
Gulf Publishing Co., Houston, TX, 197^-
k. The Stretford Process. A Report for the EPA, Research Triangle
Park, NC prepared by Catalytic, Inc., Philadelphia, PA, December 1976.
135
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APPENDIX A
GENERALIZED FORMULA FOR SULFUR REMOVAL FROM A FUEL TOMEETA SPECIFIC
LEVEL OF S02 EMISSIONS
A formula has been derived to calculate the required percent removal
of sulfur from a fuel to meet specific sulfur dioxide emissions levels in
the flue gas. This equation is as follows:
R = 100 -,
where
E = percent efficiency of Btu recovery when processing fuel
(i.e. gasification). If fuel is used directly, set E equal
to 100.
G = percent of sulfur in fuel ending up in processed product
which is to be burned (i.e. fuel gas from a gasifier). If
fuel is burned directly, set G equal to 100.
H » high heating value of fuel, Btu/lb.
L * level of SO. contained in flue gas, Ibs S02 per 10 Btu.
R « percent removal of sulfur to meet desired emission level, L.
S = percent sulfur in fuel (i.e. coal to a gasifier).
An example in applying the formula to the gasification of coal and
burning the product gas is as follows:
136
-------
E = 75% gasifier efficiency.
G = 80$ of the sulfur in coal ending up in the gasifier product,
H - 8500 Btu/lb coal heating value.
L = 0.5 Ibs S02 per 10 Btu (desired emission level).
S = 2% sulfur in the coal
100 - 97>n I 75 Xfln5°V °'5l « 90.0« sulfur removal required.
ZUU \ OU X / I
137
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APPENDIX B
SAMPLE CALCULATIONS OF CLEAN-UP PROCESSES
NOTE: Basis, assumptions and basic calculations are stated in
Section C.
1. Dry Iron Oxide Purifiers
A. Design Based on American Practices
G = 952,380 SCFH
S = 515 (factor for 368 grains per 100 SCF gas)
EPA air pollution control guideline for gaseous fuel is 0.5 pound
SO- per 10 Btu. It is equivalent to 520 ppm by volume of sulfur
compounds per SCF of gas with a heating value of 175 Btu/SCF.
The overall efficiency required to meet EPA guidelines is
100 - 520 * 100/6190 = 91.6%.
The hydrogen sulfide removal efficiency of the commercial iron
oxide boxes ranges from 85 to 95%. Assuming one dry box will give
a 91.6% or more efficiency:
C = 2 (for one box)
It has been suggested in the literature that the oxide bed should
be at least 10 ft thick for good gas distribution.
D = 10 ft.
A- (952'380) (515) =13,630 ft2
(3000) (10+2)
Considering the limitation of physical size and the requirement of
uniform gas distribution inside the iron oxide boxes, ten trains of
138
-------
gas clean-up units, in parallel, are suggested. The size of each
box is
13,630/10 = 1363 ft2
i.e., 41.7 ft I.D. x 10 ft high (cylindrical container)
37 ft x 37 ft x 10 ft high (rectangular container)
Total amount of Fe20o required is
1363 x 10 ft3/box x 10 boxes x 9 lb Fe203/ft3 = 1,226,700 Ibs of Fe203.
S umma ry
Ten boxes of 37'x37'xlO' (or 41.7' I.D. x 10') are required for
treatment of 22.86 million cubic feet per day of producer gas with
a heating value of 175 Btu per SCF.
B. Design Based on European Practices
1 . For R = 20
952,380/20 = 1*7,600 ft3 of oxide in box.
Assume the depth of bed height is 10 ft.
2
Cross-sectional area of box is 4760 ft .
Linear gas velocity = 952,380/4768/60 = 3-3 ft/min.
Mean gas residence time - 10/3-3 = 3-0 min.
Sulfur deposition rate is
952,380 x 3-68 x 0.92/60/4760 =11.3 grain/ft2/min
2
Sulfur deposition is below maximum rate of 15 grains/ft /min.
139
-------
Considering the limitation of physical size and the requirement
of uniform gas distribution inside the boxes, four boxes in parallel
are suggested.
4760/4 - 1190 ft2
i.e., 39 ft I.D. x 10 ft high
(4 columns are required)
Total amount of Fe£0o required is 428,000 pounds.
2. For R - 50
952,380/50 = 19,050 ft3 of oxide in box.
Cross-sectional area is
19,050/10 = 1905 ft2
Linear gas velocity = 952,380/1905/60 - 8.3 ft/min.
Mean gas residence time = 10/8.3 ~ 1-2 min.
Sulfur deposition rate is
952,380 x 3.68 x 0.92/60/1905 = 27-6 grains/ft2/min.
Since sulfur deposition rate is larger than acceptable value, 15,
R equal to 50 cannot be used for design.
Table I lists designs at varying values of R.
Summary
According to European design practices for iron oxide box, the values
of R can be 20 or 25 for this case. The dimension of the boxes is:
39 ft I.D. x 10 ft high, 4 boxes required; for R = 20.
35 ft I.D. x 10 ft high, 4 boxes required; for R = 25.
140
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TABLE I: IRON OXIDE BOX DESIGN BASED ON EUROPEAN PRACTICES
Total Cross
R Volume of Bed Sectional Linear Mean Gas Rate of S No. of
CFH Gas Iron Oxide Depth Area Needed No.of Box Gas Velocity Residence Deposition Box Boxes
per CF Oxide Ft^ Ft Ft^ Suggested Ft/min Time, min Grains/ft2 min Dimension Required
20 47,600 10 4,760 4 3-3 3-0 11.3 39' I.D. 4
X
10' High
25 38,100 10 3,810 4 4.1 2.4 14.1 35' I.D. 4
x
10' High
30 31,700 10 3,170 5.0 2.0 16.0 *
50 19,050 10 1,905 8.3 1.2 27-6 *
n
Rate of sulfur deposition is greater than 15 grains/ft /min. limit.
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11. Liquid-phase Iron Oxide Process (Ferrox)
H2S must be removed from product gas to meet EPA guidelines:
3.68 x 952,380 x 0.92/7000 = 460.6 Ibs H2S/hr = 0.128 Ibs H2S/sec
2 Fe (OH)3 + 3 H2S —*• Fe^ + 6 H20
3 H2S (3k) ... Ib H7S .
2 Fe (OH)3 " '-5 x (58.5 + 51) " 0.466 Ib Fe (OH)3 (theoretical)
Theoretical amount of Fe (OH), required to remove 0.128 Ibs per
sec of H2S is
0.128/0.466 - 0.275 Ibs of Fe (OH)3/sec (theoretical)
Actual amount of Fe (OH) 3 used is usually about 3 times that of the
theoretical one,
0.275 x 3 = 0.825 Ibs Fe (OH)3/sec (actual use)
This amount corresponds to 0.5 W % in scrubbing solution, the
amount of solution is
0.825 x 122-= 165 Ibs solution/sec
0.5
= 1161 gpm % 63.6 lb/ft3
0.83 Ib/sec : Fe (OH), }
7 fl65 Ib/sec
Scrubbing Solution* 4.94 Ib/sec : Na2C03 > Total \
( (J2.59 ft3/sec)
J58.93 Ib/sec : H20 )
Physical properties of gas and solution:
JU.= 1.0 cp
|°g = 20/380 = 0.053 lb/ft3
f1 - 63.6 lb/ft3
VjJ =PH20//>1 " 64.2/63.6 = 0.981
FD ** packing factor
• 40 for 2" Intalox Saddles
142
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Calculate Flooding Point for a Packed Column
L' = 165 Ib/sec
G1 = 952,380 x 0.053/3600 = 14.02 Ib/sec
The limiting vapor velocity for practical operation of a packed
twoer is set by the flooding point. The flooding point can be
predicted as presented in Perry's Chemical Engineer's Handbook or
in the generalized pressure drop correlation of an aritcle by
J. S. Eckert (Chemical Engineering Progress, Vol. 66, No. 3,
March 1970, p. 40).
L/G (Pg/P,)1/2 = (Pg/P,)172 = (053} '/2 „ 0.339
From the correlation, it is found that the value of ordinate is
_ = 0.063
*\ A —
G2 = 0.063 -
G2 = 0.063 (0.053) (63.6) (32.2) =0>,7A
(40) (0.981) (l.O)0'2
G = 0.418 lb/ft2 sec
Cross-sectional area of packed column at flooding point is
14.02/0.418 = 33.6 ft2
The vapor velocity at flooding point is, V f,
Vgf = 952,380/3600/33.6 = 7-9 ft/sec
Calculate Packed Tower Diameter
Because the packed column operation may become unstable as the
flooding point is approached, the design value for allowable vapor
143
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velocity is usually estimated to be 50 to 70% of the maximum
allowable velocity, and this allowable velocity is used to
establish the column diameter. Say at 50%,
Vg = Vgf x 50* = 7-9 x 0.5 = 3.9 ft/sec
A - 952,380/3600/3.9 - 67.2 ft2
i.e., 9 ft 3 inch I.D. column.
Packed Bed Height
The height of packing material required can be estimated by
NOG
- (mGM/LM)
where NQQ » number of overall transfer units
m * slope of equilibrium curve dy /dx
X£ - mole fraction solute in liquid fed to top of column
yj • mole fraction solute in gas fed to bottom of column
Y2 ** mole fraction in gas leaving top of column
Gu - superficial molar mass velocity of gas stream,
Ib moles/(hr)(sq ft)
LM - superficial molar mass velocity of liquid stream,
Ib moles/(hr)(sq ft)
The number of overall transfer units is estimated to be 15 from
Perry's Handbook.
(15)
26 3 ft
* *
20 (o. i.o)
plus about 10 ft of freeboard for gas-liquid disengagement, 15 ft for
bottom head, the overall height of the packed column is about 50 ft.
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Size of Regenerator
It has been indicated in literature that a minimum 5 minutes of
liquid mean residence time is required for regenerating ferric
sulfide into ferric oxide.
V - 2.59 ftVsec x 60 x 5 = 777 ft3
i.e., 7-0 ft I.D. x 20 ft high
Summary
Packed Tower
Packed column : 9 ft 3 inches I.D. x 50 ft overall height
Packing Height : 26 ft
Packing Material: 2" Intalox Saddles with F = AO
Regenerator
Total Volume Required: 777 ft3
i.e., 7-0' I.D. x 20' High (for example)
Flow Rates
Product Gas: 952,380 SCFH or 22.86 x 10 SCFD
Solution Circulation Rate: 165 Ib/sec or 1161 gpm
III. Stretford Process
The amount of HnS to be removed from raw product gas to meet EPA
guidelines is
3.68 gr/SCF x 952,380 SCF/hr x 0.92/7000 gr/lb =
460.6 Ib H2S/hr - 0.128 Ib H2S/sec.
-------
From process basic chemistry, it is shown that one mole of sodium
carbonate is required to react with each mole of hydrogen sulfide.
i.e., Na2C03 + H2S = NaHS + NaHCOj
Na2C03 _ 106 _ . Ib Na2C03
H2S 3k ' Ib H2S
Hence, theoretical amount of Na^CO, required to absorb 0.128 pound
per second of H2S is
0.128 x 3-118 = 0.399 Ib Na2C03/sec.
Actual amount of Na2C03 used in practical operation is usually about
two to three times that of the theoretical one. Say three times,
0.399 x 3 = 1.197 Ibs Na2C03/sec.
The scrubbing solution is an aqueous solution containing sodium
carbonate and bicarbonate in the proportion of about 1:3 and 2,7
anthraquinone disulfonic acid (ADA). The concentration of Na2CO? is
about 0.1 N and NaHCO. is 0.3 N in order to keep the pH range of 8.5
to 9.5- This is equivalent to a solution with 0.505 W % of Na2COj.
Thus the amount of solution required is
1.197 x 100/0.505 = 237-1 Ib solution/sec
= 1625 gpm £ 65.5 lb/ft3
The following physical properties of gas and solution have been
estimated:
jU = 1.0 cp
pg - 0.053 lb/ft3
f, = 65.5 lb/ft3
= 0.952
-------
Calculate Flooding Point of a Packed Column
L1 = 237-1 Ib/sec
G1 = 952,380 x 0.053/3600 = 14.02 Ib/sec
From Perry's Chemical Engineers' Handbook, it is found that
G F tyjlH^
pP o] = o-0**1
Assuming 2" saddles are used as a packing material with a packing
factor of 40, i.e. F =40
1/2
r
G =
,, PgPl >c 1/2 [(0.041) (0.053) (65.5) (32.2)]
'" FpVxiO-2} =[ (40) (0.952) (1.0)0-2 J
= 0.347 lb/ft2 sec
Cross-sectional area of packed column at flooding point is
14.02/0.347 - 40.4 ft2
The vapor velocity at flooding point is
952,380/3600/40.4 = 6.5 ft/sec
Calculate Packed Column Diameter
Assume the design value for allowable vapor velocity is 50% that of
flooding point.
V = 6.5 x 0.5 = 3-3 ft/sec
Thus, the cross-sectional area is
A = 952,380/3600/3-3 = 80.2 ft2
i.e., 10 ft I.D.
147
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Packing Height
A mean residence time of 10 seconds is assumed.
3-3 ft/sec x 10 sec = 33 ft.
Reaction Tank (Delay Tank)
The reaction tank is located at the bottom of the absorber tower.
A minimum of 10 minutes holding time is required to ensure a
complete sulfur deposition. The height of the reaction tank is
h - (2.371) (60) (10) m ,
n (65.5) (80.2)Z/ tt-
Oxidizer (Regenerator)
The oxidation of reduced ADA to normal ADA will take 20 to 60 minutes
for completion. However, 10 to 15 minutes will be practically enough
for reefrculation of the solution. Say 15 minutes for this calculation.
Then the size of the oxidizer is
2?7 *6° x 15 = 3260 ft*
65.5
i.e., }k ft 6 inches I.D. x 20 ft high
(Twenty feet high is typical for oxidizer)
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APPENDIX C
MATERIAL BALANCES ON NITROGEN AND SULFUR COMPONENTS fOR
RILEY-MQRGAN GASIFICATION SYSTEM
1. DATA
The following data are extracted from the paper presented by Rawdon
et al.
A. Dry Gas Composition V %
CO 2k.6
H2 13.1
CHi, 4.4
C02 5.9
N2 + A 51.7
CnHm Q.I
NH3 0,07
H2S 0.12
B. Average gas flow Rate is 40 scfm
C. Ammonia concentration in gas is 666 ppm
D. Carbon, nitrogen and sulfur contents of feed coal are 81.0
1.54 and 0.7 W %, respectively.
E. Molar conversion of coal nitrogen to ammonia is 9*
F. Tar loading in the gas is 0.8 grams/scf
G. Nitrogen and sulfur contents of tars in the gas are 1.2 and
0.5 W %, respectively.
2. CALCULATION
A. Back-calculation of coal feed rate
1) Based on nitrogen balance in gas
666 ppm NH3 x 40 x 60 - 1.598 scfh of NH3
1.598/380 x 17 = 0.0715 Ibs of NH3
- 0.0588 Ibs of Nitrogen
0.0588 - 0.09 x 0.0154 x W
W - 42.4 Ibs/hr of coal feed
149
-------
2) Based on carbon balance and an assumed value of carbon
conversion.
Gas Component Carbon Balance
CO 2k.6% x 40 x 60 = 590.4 scfh
CH^ 4.4% x 40 x 60 = 105.6 scfh
C02 5.3% x 40 x 60 - 141.6 scfh
C H 0.1% x 40 x 60 x 3 = 7.2 scfh
n m
844.8 scfh
2.22 Moles/Hr
26.7 Ibs/Hr
If 80% of carbon in coal is in the gas, the the coal feed rate
is 26.7/0.8/0.81 =41.2 Ibs/Hr.
Since both figures come out very close to each other, the coal
feed rate is taken to be 42 Ibs/Hr in subsequent calculations.
B. Sulfur-component balance
In: 42 x 0.7% = 0.294 Ibs of sulfur per hour
Out: 0.12% x 40 x 60 x 32/380 = 0.243 Ibs/Hr in HgS
0.8 x 2400/453.6 x 0.5% = 0.021 Ibs/Hr in Coal Tars
0.294 - 0.243 - 0.021 = 0.030 Ibs/Hr in (Coke + Char)
C. Nitrogen-component balance
Accurate calculation of nitrogen balance is impossible because no
data on the ratio of air feed rate to coal feed rate have been given.
The nitrogen-component balance, however, is estimated as follows:
In: 42 x 1.54% = 0.647 Ibs of N in Coal/Hr
Out: 0.07% x 40 x 60/380 x 14 = 0.062 Ibs of N in MH3
0.8 x 2400/543.6 x 1.2% = 0.051 Ibs of N in Coal Tar
150
-------
The balance 0.53** (0.6^7 - 0.062 - 0.050 appears as free nitrogen in
the gas and as bonded nitrogen in the coke and char. According to data
of Kohl, the split of the remaining nitrogen is estimated as 3**% in coke
and char and 66% in the gas.
Hence 0.53^ x lk% = 0.182 Ibs/Hr in Coke and Char
0.531* x 66% = 0.352 Ibs/Hr in Gas as N0
3- SUMMARY
A. Sulfur Balance
In, Ib/Hr
Coal
0.29**
(100%)
Out, Ib/Hr
H S Tars Coke & Char
0.253 0.021 0.030
(82.65%) (7.14*) (10.21%)
B. Nitrogen Balance
In. Ib/Hr
Coal
0.6^7
(100%)
NH
0.062
(9.58%)
Out, Ib/Hr
Tars N Coke + Char
0.051 0.352 0.182
(7.88%) ($k.50%) (28.
151
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TECHNICAL REPORT DATA
(Please read Instructions on the reverse before completing)
1. REPORT NO.
EPA-600/7-79-171
2.
3. RECIPIENT'S ACCESSION-NO.
4. TITLE AND SUBTITLE
Summary of Gas Stream Control Technology for
Major Pollutants in Raw Industrial Fuel Gas
5. REPORT DATE
July 1979
6. PERFORMING ORGANIZATION CODE
7. AUTHOR(S)
F.D.Hoffert, W.Y.Soung, and S.E.Stover
8. PERFORMING ORGANIZATION REPORT NO.
9. PERFORMING ORGANIZATION NAME AND ADDRESS
Hydrocarbon Research, Inc.
Lawrence Township, New Jersey 08648
10. PROGRAM ELEMENT NO.
E HE 62 3 A
11. CONTRACT/GRANT NO.
68-02-2601
12. SPONSORING AGENCY NAME AND ADDRESS
EPA, Office of Research and Development
Industrial Environmental Research Laboratory
Research Triangle Park, NC 27711
13. TYPE OF REPORT AND PI
Final; 7/77 - 3/79
PERIOD COVERED
14. SPONSORING AGENCY CODE
EPA/600/13
15. SUPPLEMENTARY NOTES
IERL-RTP project officer is William J. Rhodes, Mail Drop 61, 919/541-2851.
is. ABSTRACT
repor|. summarizes coal gasification and clean-up technology with em-
phasis on methods of producing a clean industrial fuel gas as defined by agreement
for study purposes. The coal-derived industrial fuel discussed produces no more
than 0. 5 Ib of SO2 , 0. 4 Ib of NOx, and 0. 1 Ib of particulates per million Btu of fuel
gas. In general, existing state-of-the-art control technology will allow these emission
guidelines to be met, although the end use for the fuel gas will strongly influence the
choice of the pollution control technology that is used. Many but not all important
factors pertinent to control technology application were considered. Costs are an
example of important factors which were not evaluated because the objective was to
first determine appropriate technology that could be applied. Emissions other than
the three major pollutants indicated were given only cursory treatment. Neverthe-
less , a general overall background of control technology for industrial fuel gas has
been covered.
17.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b.IDENTIFIERS/OPEN ENDED TERMS
c. COSATI Field/Group
Pollution
Goal Gasification
Coal
Gases
Fuels
Sulfur Dioxide
Nitrogen Oxides
Dust
Pollution Control
Stationary Sources
Industrial Fuel Gas
Particulate
13 B
13H
2 ID
07D
07B
11G
18. DISTRIBUTION STATEMENT
Release to Public
19. SECURITY CLASS (ThisReport)
Unclassified
21. NO. OF PAGES
20. SECURITY CLASS (Thispage)
Unclassified
22. PRICE
EPA Form 2220-1 (9-73)
152
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