Environmental Protection Technology Series
REDUCTANT GASES FOR FLUE GAS
DESULFURIZATION SYSTEMS
Industrial Environmental Research Laboratory
Office of Research and Development
U.S. Environmental Protection Agency
Research Triangle Park, North Carolina 27711
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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 five series. These five broad
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2. Environmental Protection Technology
3! Ecological Research
4. Environmental Monitoring
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This report has been assigned to the ENVIRONMENTAL PROTECTION
TECHNOLOGY series. This series describes research performed to develop and
demonstrate instrumentation, equipment, and methodology to repair or prevent
environmental degradation from point and non-point sources of pollution. This
work provides the new or improved technology required for the control and
treatment of pollution sources to meet environmental quality standards.
EPA REVIEW NOTICE
This report has been reviewed by the U.S. Environmental
Protection Agency, and approved for publication. Approval
does not signify that the contents necessarily reflect the
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This document is available to the public through the National Technical Informa-
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EPA-600/2-76-130
May 1976
REDUCTANT GASES
FOR FLUE GAS DESUL FURIZ ATION
SYSTEMS
by
D.W. Hissong, K.S. Murthy, andA.W. Lemmon, Jr.
Battelle-Columbus Laboratories
505 King Avenue
Columbus, Ohio 43201
Contract No. 68-02-1323. Tasks 21 and 36
ROAPNo. 21BJV-038
Program Element No. 1NB458
EPA Task Officer: Charles J. Chatlynne
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
The use of coal or residual oil gasification to produce a
hydrogen/carbon monoxide-rich gas for use as a reductant for regenerable
flue gas desulfurization (FGD) processes is considered in this study. Two
different reduction systems are considered - one for the type of FGD
process that produces a concentrated S02 stream and one for the type that
uses a liquid-phase Glaus reactor.
Detailed data on the composition of the raw gas from several
gasifiers are analyzed. To supplement the data on trace constituents
in the gas, thennodynamic calculations were made to determine the equili-
brium gas-phase concentrations for a typical coal and typical gasification
conditions. Mass transfer calculations were made to determine the extent
to which certain gaseous species could be removed by water washing of
the gas. The effects of the remaining trace constituents on the components
of the reduction systems are analyzed. Recommendations for additional
research on trace constituents and their effects are made.
The capital and operating costs for reduction systems based
on gasification of coal and residual oil are estimated and compared with
those for reduction systems based on natural gas.
iii
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TABLE OF CONTENTS
Page
1. INTRODUCTION 1
2. SUMMARY 3
3. FLUE GAS DESULFURIZATION PROCESSES 6
4. POSSIBLE ROUTES TO ELEMENTAL SULFUR 11
4.1 Type A FGD Processes 11
4.2 Type B FGD Processes . • 13
5. REDUCTANTS AND REDUCTION REACTIONS 15
5.1 Reduction of S02 to Sulfur 15
5.2 Reduction of SO. to H-S 24
6. USE OF PURCHASED HYDROGEN SULFIDE 28
7. GENERATION OF REDUCING GAS BY GASIFICATION 30
7.1 Coal Gasification Processes 31
7.2 Residual Oil Gasification Processes 33
8. TRACE CONSTITUENTS IN GASIFIER PRODUCTS 35
8.1 Experimental Data on Gasifier Product Trace Constituents . 35
8.2 Thermodynamically Predicted Trace Constituent Data .... 41
8.3 Contacts with Process Vendors 45
9. EFFECT OF.TRACE CONSTITUENTS ON FGD SYSTEMS 47
9.1 Overall FGD Plus Reduction Systems 47
9.2 Effects on FGD Processes 47
9.3 Effects on Sulfur Production Processes 49
10. GAS CLEANUP SYSTEMS 57
10.1 Scrubber Systems 57
10.2 Effectiveness of Scrubber Systems 58
iv
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TABLE OF CONTENTS
(Continued)
Page
11. COSTS OF GASIFIER-BASED REDUCTANT SYSTEMS FOR FGD PROCESSES . . 60
11.1 Type A FGD Processes 60
11.2 Type B FGD Processes 75
11.3 Other Considerations 81
11.4 Use of Oil Gasification 82
12. CONCLUSIONS 84
13. RECOMMENDATIONS 86.
14. ACKNOWLEDGMENTS 88
15. REFERENCES 89
APPENDIX A
FLUE GAS DESULFURIZATION PROCESSES A-l
APPENDIX B
USE OF PURCHASED HYDROGEN SULFIDE B-l
APPENDIX C
GASIFICATION PROCESSES C-l
APPENDIX D
CALCULATION OF TRACE CONSTITUENT DISTRIBUTIONS D-l
APPENDIX E
GAS CLEANUP SYSTEMS E-l
APPENDIX F
DETAILS ON ENGINEERING ANALYSIS F-l
APPENDIX G
SOURCES OF COST INFORMATION . . „ G-l
APPENDIX H
CONVERSION FACTORS H-l
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LIST OF TABLES
Table 3-1.
Table 3-2.
Table 5-1.
Table 7-1.
Table 7-2.
Table 8-1.
Table 8-2.
Table 8-3.
Table 8-4.
Table 8-5.
Table 11-1.
Table 11-2.
Table 11-3.
Table 11-4.
Table 11-5.
Table 11-6.
Table 11-7.
Table 11-8.
Table 11-9.
Flue Gas Desulfurization Processes
Typical Compositions of Type A FGD Product Gases ....
Data on Reduction of S02 with Hydrogen
Comparison of Selected Commercial Atmospheric Pressure
Coal Gasifiers
Typical Gas Compositions from Oil Gasification
Processes
Experimental Data on Koppers-Totzek Gas Trace
Constituents
Trace Constituent Data for Riley-Morgan Gasifier ....
Mean Analytical Values for Constituents in 101 Different
Comparison of Predicted and Experimental K-T Gas
Comparison of Thermodynamically Predicted Trace
Constituent Data
Reducing Gas Compositions Used in Material Balance
Investments for Reduction Options, Type A FGD Process,
1000-MW Power Plant
Investments for Reduction Options, Type A FGD Process,
500-MW Power Plant
Operating Costs for Reduction Options, Type A FGD
Process, 1000-MW Power Plant
Operating Costs for Reduction Options, Type A FGD
Process, 500-MW Power Plant
Unit Costs Used in Operating Cost Calculations
Effects of SO- Dilution on Sulfur Recovery in Two-
Stage Glaus Plant
Compositions and Flow Rates for H?S Production Based on
Coal Gasification
Investment for H-S Generation System for Type B FGD
Pag<
7
10
17
32
34
36
38
40
42
44
62
67
68
70
71
72
73
77
78
vi
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LIST OF TABLES
(Continued)
Page
Table 11-10. Operating Costs for H,2S Generation System for Type B
FGD Process ' 80
Table 11-11. Cost Comparison for Coal and Oil Gasification
Systems 83
LIST OF FIGURES
Figure 3-1. Schematic Flow Sheet of Westvaco FGD Process 8
Figure 4-1. Possible Routes to Elemental Sulfur with Type A Flue
Gas Desulfurization Process 12
Figure 4-2. Possible Routes to Elemental Sulfur with Type B Flue
Gas Desulfurization Process 14
Figure 5-1. Data on Reduction of SO with Carbon Monoxide 18
Figure 5-2. Flow Sheet of ASARCO-Phelps Dodge SO. Reduction Pilot
Plant 7 20
Figure 5-3. Allied Chemical SO. Reduction Process 22
Figure 5-4. Reduction of Sulfur Compounds as First Step in Shell
Glaus Off Gas Treating (SCOT) Process 26
Figure 5-5. Flow Sheet of Conceptual Reduction System Proposed for
Shell FGD Process 27
Figure 8-1. Battelle Experience with Time-Temperature Effects on
Predictive Power of Thermodynamic Calculations .... 43
Figure 9-1. Possible Routes to Sulfur Production Based on Coal
Gasification Reducing Gas in FGD Processes 48
Figure 9-2. Schematic for Coal Gasification, Cleanup, and H.S
Generation 54
Figure 9-3. Schematic of American Smelting & Refining Company
(ASARCO) Elemental Sulfur Production Plant Adapted to
Use of Coal Gas 55
Figure 10-1. Conceptual Design of a Spray Scrubber, Gas Cleanup
System 59
vii
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LIST OF FIGURES
(Continued)
Page
Figure 11-1. Material Balance Flow Sheet for Type A FGD, 85% SO
Stream, Oxygen-Blown Gasifier 63
Figure 11-2. Material Balance Flow Sheet for Type A FGD, 25% SO-
Stream, Oxygen-Blown Gasifier 64
Figure 11-3. Material Balance Flow Sheet for Type A FGD Process,
85% SO^ Stream, Air-Blown Gasifier 65
Figure 11-4. Material Balance Flow Sheet for Type A FGD Process,
25% SO. Stream, Air-Blown Gasifier 66
Figure 11-5. Material Balance Flow Sheet for H2S Production Based
on Coal Gasification 76
viii
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REDUCTANT GASES FOR FLUE GAS DESULFURIZATION SYSTEMS
by
D. W. Hissong, K. S. Murthy, and A. W. Lemmon, Jr.
1. INTRODUCTION
One of the major environmental challenges facing the United States
today is that of reducing the emissions of sulfur oxides from the flue gases
of fossil fuel-fired steam electric power plants. Extensive research and
development work is in progress ort flue gas desulfurization (FGD) processes
to accomplish this. Although the most advanced processes are generally of
the "throwaway" type which produce an unusable mixture of sulfur compounds,*
development work is also in progress on a number of "regenerable" processes
which can ultimately produce either elemental sulfur or sulfuric acid as a
by product. For many power plants which will have to install flue gas
desulfurization systems, elemental sulfur would be the preferred product.
Elemental sulfur is desirable because it is a noncorrosive solid which is
easily handled, stored, and shipped, and which can be used in a number of
industrial operations.
Since the sulfur in flue gas is in the form of sulfur oxides
(primarily SO.), it follows that any flue gas desulfurization system which
ultimately produces elemental sulfur must involve some form of reduction.
There are a number of ways in which this reduction can be effected, depending
on the nature of the flue gas desulfurization process involved. Theoretically,
any carbon source, such as coal, coke, petroleum fractions, or natural gas,
can be used as the ultimate reducing agent for the SO-. Conversion of these
materials into more reactive forms, may be necessary before the actual
reduction of the S0« takes place. Most of the alternatives currently
being proposed for use with regenerable FGD processes involve the use of
natural gas as the ultimate reductant. With the current and projected
severe shortage of natural gas, il: is imperative that alternate sources of
reductants be carefully examined. This is the overall objective of this study.
* The distinction between throwaway and regenerable processes depends upon
whether the sulfur specie prodxtced can be utilized profitably.
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The specific objectives of chis study are to:
(1) Summarize the reduction technology currently being
proposed for use with regenerable flue gas desulfuri-
zation process that yield elemental sulfur as a
product.
(2) Identify various routes by which materials other
than natural gas or light petroleum fractions
(liquefied petroleum gas or naphtha) can be used
as the ultimate reductants in connection with
regenerable flue gas desulfurization processes.
(3) Assess the technical and economic feasibility of
the most promising routes identified under
Objective 2. This will involve estimating the
costs of these routes and comparing with the
costs of the currently proposed options identified
under Objective 1.
(4) Define the needs for additional development studies
or demonstration plants to advance these new options
toward the point of commercial use.
This study is directed toward fossil-fuel-fired steam-electric
power plants in the United States. The objective at these power plants
is to remove about 90 percent of the SO- from the flue gas and to
recover elemental sulfur as the product from this operation. The material
used as the reductant for this operation must be coal, coke, petroleum
residual oil, or purchased hydrogen sulfide. Any conversions of these
reductants to other forms must be done on the premises of the power plant.
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2. SUMMARY
This report deals with the production of reductant gases for
regenerable flue gas desulfurization processes from materials other than
natural gas. The alternative considered in depth is that of gasifying
coal or residual oil to produce a hydrogen/carbon monoxide-rich gas which
can be used as the reductant. A number of suitable gasification processes
are commercially proven and available in this country. The complexity of
the reduction system depends largely on the type of flue gas desulfurization
process. For the FGD processes that produce a concentrated SO stream,
the reduction system will involve three reaction stages with the first
reactor containing either the American Smelting and Refining Company
(ASARCO) SO,- reduction catalyst or a cobalt-molybdenum catalyst. The last
two reactors will contain conventional Glaus catalysts. For the FGD
processes that use liquid-phase Glaus reactors, the required H.S can be
generated by reducing some of the elemnetal sulfur product with reducing
gas from a gasifier.
The raw gas from the gasifier will have to be washed with
water to remove ammonia, tar, and particulate matter. Although experimental
data in this area are lacking, thermodynamic and mass transfer calculations
indicate that water washing should effectively remove these substances and
the associated trace elements. The effects of the remaining trace species
on most of the FGD process itself will be insignificant because the amounts
are very small in comparison to those already entering the FGD process via
the flue gas. For a power plant using 3.5 percent sulfur coal, the amount
of coal fed to the gasifier to make the reducing gas will be only about 3.6
percent of the amount of coal fed to the power plant boilers. On the other
hand, there are portions of some FGD processes which are sufficiently free
of trace species that the introduction of such species via the reducing
gas could cause difficulties. The effects of trace species on the reduction
systems are not well defined. The only known adverse effects are related
to ammonia (for vapor-phase Glaus: catalyst) and COS (for liquid-phase Glaus
systems), and these species can be removed from the reducing gas. Experi-
mental studies related to trace species in the reducing gas and their
effects on the reduction systems are recommended.
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The extent to which the reducing gas must be processed to
increase its hydrogen content prior to the reduction is not well defined,
particularly for the case in which sulfur is to be reduced to produce H-S.
In this case, shifting of carbon monoxide to hydrogen and removal of
carbon dioxide may be required. Experimental studies in this area also
are recommended.
Reduction systems based on gasification of coal or residual
oil will be relatively expensive, at least in comparison with systems based
on natural gas. The following cost estimates apply to a 1000-megawatt
power plant fired with 3.5 percent sulfur coal. For the FGD processes
that produce a concentrated S0_ stream, the investment (1975) for the
reduction system is about $15 to $19/kW and the operating cost is 0.77
to 0.92 mills/kwh (for coal at $22/metric ton). The ranges given cover
the range of product gas SO,, concentrations from such FGD processes and
cover gasification with both air and oxygen. These costs (both investment
and operating) represent roughly 15 percent of the estimated costs of the
FGD process itself.
For the FGD processes that use liquid-phase Glaus reactors, the
investment for the reduction system is about $32/kW and the operating
cost is 1.13 mills/kW. in this case the gasification must be done with
oxygen. These costs are based on the assumption that the reducing gas
must be upgraded by CO shifting and C0« removal. If such upgrading can
be eliminated, the investment would be reduced by about 30 percent.
Including the upgrading, the costs of the reduction system represent
roughly 30 percent of the estimated costs of the FGD process itself.
A gasifier-based reduction system will increase significantly
the complexity of the overall FGD system. On the other hand, there are
other possible benefits of such a reduction system. The gasifier could
be used to supply some of the energy requirements of the FGD process,
such as for stack gas reheating. This energy could be supplied either
as excess by-product steam from the gasifier or as excess fuel gas produced.
Purchased H-S will play only a minor role as a reductant for
S02 from power plant FGD systems. The use of H-S will involve almost
entirely by-product streams obtained "across the fence" from a source
such as an oil refinery or a natural gas processing plant, and geographical
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considerations will restrict this usage. For a 1000-megawatt power plant
fired with 3.5 percent sulfur coal, H«S can be produced by reducing
sulfur with a coal-based reducing gas for $55-$64/metric ton H S (2.5-2.9C/
Ib) when coal costs $ll-$33/metr:lc ton. Therefore, purchased H.S would
be economical only if the price were lower than this level as applied to
the specific situation.
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3. FLUE GAS DESULFURIZATION PROCESSES
An analysis of regenerable flue gas desulfurization processes
was made in order to:
(1) Identify those processes which could be used in
combination with a reduction process to yield
elemental sulfur as a product
(2) Examine the compositions of their product gas
streams so that appropriate reduction systems
could be selected.
Emphasis was placed on processes which are fairly well advanced in their
development, although it should be reemphasized that none of these regener-
able processes is as advanced as limestone scrubbing, which has been
demonstrated in a number of full-scale installations.
A list of flue gas desulfurization processes which can produce
elemental sulfur is presented in Table 3-1. These processes have been classi-
fied into four groups on the basis of their reduction requirements. The first
two groups (Type A and B) will be referred to frequently in this report.
The Type A FGD processes are those which produce a concentrated SCL
stream which must be reduced. The Type B FGD processes require H.S as a
reductant because these processes include liquid-phase Glaus plants.
The Westvaco carbon adsorption process is in a category by
itself because, although it requires H.S internally (to reduce the H.SO,
adsorbed on the carbon), the H.S generation step is an integral part of
the process. A flow sheet of this process is shown in Figure 3-1. The
process requires a reductant such as hydrogen which can be used to
reduce sulfur internally to generate the H^S.
The Atomics International Aqueous Carbonate process is in
a separate category in that it inherently fulfills the desired goal of
using a "dirty" (sulfur-containing) ultimate reductant, namely, coke or
coal. Any Type A FGD process when coupled with Foster Wheeler's RESOX
reduction process would also be in this category. The Bergbau-Forschung
process is often coupled with the RESOX process.
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TABLE 3-1. FLUE GAS DESULFURIZATION PROCESSES
WHICH CM PRODUCE ELEMENTAL SULFUR
Type A -Processes Which Require Reduction of Concentrated S0? Streams
1. Very concentrated S02 streams containing only SO- and H-0
e Wellman-Lord - 85% S02
o Shell-UOP CuO ~ 90% SO- with absorber-stripper, 46% without
2. Less concentrated S0« streams containing other species
e Ammonia scrubbing - 27% SO- + HO, NH
e Magnesia scrubbing - 24% S02 + H20, K , CO,,, 02
• Bergbau-Forschung Carbon Adsorption - 23% SO + H90, N2, CO.
Type B - Processes Which Require H^S to Reduce SO^ Liquid Phase
i. 2.
e USBM Citrate
• Stauffer Powerclaus
• IFP-Catalytic NH3 Scrubbing
Type C - Processes Which Require Reduction of Sulfur Internally to
Generate H S
• Westvaco Carbon Adsorption - H.S needed to reduce H2SO, on
carb on
Type D - Processes Which Inherently Use a "Dirty" Reductant
e AI Aqueous Carbonate Process - coke
• Any Type A process when coupled with the Foster-Wheeler
RESOX reduction process
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1/2
Cleaned
Flue Gas
H20
65-150 C
(150-300 F)
Flue Gas
S02
removal
Acid-Laden Carbon
Spent Gas to Boiler
Carbon
Sulfur-Laden
Carbon
Sulfur stripper/
H2S generator
650 C (1200 F)
3H2 + AS
-*- 3H2S + S
Reductant
(H2/CO)
oo
Source: "Westvaco SO. Recovery Process" brochure, Westvaco Corporation, Charleston, South Carolina.
FIGURE 3-1. SCHEMATIC FLOW SHEET OF WESTVACO FGD PROCESS
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For the Type A processes an analysis was made to determine
typical compositions of the product gas streams. This involved analyzing
literature on the processes and, in some cases, making calculations based
on physical properties or reaction stoichiometry. The typical SO. con-
centrations are included in Table 3-1. More information on flue gas
desulfurization processes in general and particularly the product gas
compositions is presented in Appendix A. Based on the data available,
the compositions shown in Table 3-2 were selected as typical for the two
classes of Type A FGD processes. These compositions were used in the
material balance calculations for the FGD plus reduction systems to be
discussed.
There are two Type A processes listed in Table 3-1 for which
the product gas contains a species which could be troublesome in the
reduction system. In ammonia scrubbing, the residual ammonia in the
product gas could act as a poison toward the catalysts used in the
reduction system. The effects of ammonia will be further discussed later
•in this report. In magnesia scrubbing, the oxygen in the product gas
(about 4 mole percent) will increase the consumption of the reductant.
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10
TABLE 3-2. TYPICAL COMPOSITIONS OF TYPE A FGD PRODUCT GASES
Volume Percent in Product Gas from FGD Process
Class 1 - Very Concen- Class 2 - Less Concen-
Specie trated SO- Stream trated SO- Stream
so,
2
co2
N2
H.O
2
85
15
100
25
17
43
15
100
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11
4. POSSIBLE ROUTES TO ELEMENTAL SULFUR
The various possible routes by which a flue gas desulfurization
system can be integrated with a reduction system to yield an elemental sulfur
product are considered in this section. At this point no consideration will
be given to whether the technology for a given step has been proven or even
attempted. The status of this technology will be reviewed in Section 5.
The various options will be considered separately for the two types of flue
gas desulfurization processes of interest in this study.
4.1 Type A FGD Processes
(Those Producing Concentrated SO,.)
The possible routes to elemental sulfur which are within the scope
of this study and apply to Type A flue gas desulfurization processes are shown
in Figure 4-1. The problem here is one of reducing S02 in a relatively concen-
trated stream to elemental sulfur. The reduction schemes may be classified
as being either "direct" or "indirect." in direct reduction, the SO is
reduced catalytically to elemental sulfur in one step without going through
any other observable oxidation states. Hydrogen, carbon monoxide, and carbon
have been studied as reductants for direct reduction. Indirect reduction
involves more than one step, with the final step being the reaction of SO
with H2S in a Glaus plant to produce elemental sulfur. The H2S may be
obtained by hydrogenation (reduction) of either SO or process-produced
elemental sulfur. The H2S may also be purchased externally. Excluding the
external purchase of H2S, an indirect reduction scheme means that part of
the sulfur initially present in the 44 valence state (SO.) is reduced all
the way to the -2 valence state (H2S) before the desired zero valence state
is achieved.
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Coal
Coke
Petroleum
Residuum
C and H Source
Fluo Cue
Cleaned
Flue Cas
Type A Flue Gas
Desulfurization Process
Cone. SO,
Reductant
Conversion
Steam
0. or Air
Reductant
SO 2
Direct Reduction
Process
Elemental
Sulfur
Tail Gas Recycle
to FGD Process
Purchased
HS
Clous Sulfur Plant
Tail Gas Recycle
to FCD Process
Elemental
Sulfur
FIGURE 4-1. POSSIBLE ROUTES TO ELEMENTAL SULFUR WITH TYPE A FLUE GAS DESULFURIZATION PROCESS
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13
Within the scope of this study, the ultimate source of the
reductant is coal, coke, or petroleum residuum. The ultimate reductant
either can be used directly or can be converted into a more "convenient"
(reactive) form. The reductant conversion step normally would be some
type of gasification process which would convert the carbon and hydrogen
in the original reductant into primarily H_ and CO.
A. 2 Type B FGD Processes
(Those Requiring H^S)
The possible routes to elemental sulfur which are within the
scope of this study and apply to Type B flue gas desulfurization processes
are shown in Figure 4-2. The H2S required by this type of flue gas
desulfurization process either can be purchased externally or can be
generated by reducing part of the elemental sulfur product. For the
latter route, the ultimate reductant again can be used directly or con-
verted into another form.
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Coal
Coke
Petroleum
Residuum
Flue Gas
C and H Source
Purchased
H2S
Type B Flue Gas
Desulfurization Process
Reductant
Conversion
Cleaned
Flue Gas
Steam
Reductant
or Air
Elemental
Sulfur
FIGURE 4-2. POSSIBLE ROUTES TO ELEMENTAL SULFUR WITH TYPE B FLUE GAS DESULFURIZATION PROCESS
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15
5. REDUCTANTS AND REDUCTION REACTIONS
In this section the state of the art of the technology for a
number of reduction reactions is reviewed. These reactions and the types
of flue gas desulfurization processes for which they are of interest are
as follows:
Reduction Types of FGD Process
SO. to S . Type A with direct reduction
Type A with indirect reduction of SO
bU2 to H2b Type B with H^s made from SQ^
- Type A with indirect reduction of S
t0 2 Type B with H2S made from S
For each of these types of reductions a number of possible reductants will
be discussed.
5.1 Reduction of SO,, to Sulfur
Hydrogen as Reductant
Catalytic reduction of S09 by hydrogen has been studied by many
(1 2)
investigators. ' Recently, a kinetic model was proposed by a research
(2)
group at the University of Akron,, using an activated bauxite as the
catalyst. Sulfur dioxide is reduced according to the following chemical
reaction:
S02 + 2H2 -»• 2H20 + 1/2 S.
However, an undesired side reaction is always occurring:
H2 + 1/2 S2 -»• H2S.
That is, the reaction does not stop at elemental sulfur but goes on to
form H,,S. Thus, S0? is reduced only partially to condensable elemental
sulfur, and the remaining effluent gas is a mixture of H2S and unreacted
SO . Presumably, the mixed S02 and H S could be reacted in a separate
Glaus reactor to complete the conversion. However, the experimental
(2)
results indicate that total SO conversion is low and consequently
the H^S/SO. ratio in the effluent: is always far below 2.0, the ratio
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16
required for the Glaus- reaction. This is shown in the data presented in
Table 5-1. Thus, it appears that the concept of reducing S02 with hydrogen
but stopping at elemental sulfur has not been adequately demonstrated as
of this time.
Carbon Monoxide as Reductant
In the presence of suitable catalysts, carbon monoxide has
proven to be effective for reducing SO.. The chemical reaction is
2CO + S02-»- 1/2 S2 + 2C02.
However, a parallel reaction produces noxious carbonyl sulfide (COS) at
the expense of sulfur yield.
CO + 1/2 S2 -*- COS.
Exploratory studies have been done by Chevron, U.S. Bureau of
Mines/ ' ' and the University of Massachusetts/ ? All of these were
aimed at the removal of SO. from stack gas (typically 2000 ppm), with elemental
sulfur as the end product. Copper on alumina has been found to be a preferred
catalyst. Typical reaction time is about 0.2 seconds at temperatures above
425 C (800 F). The results of Quinlan, et al^ ' indicate that SO. conversion
increases with increasing temperature and CO/SO- feed ratio. Unfortunately,
the undesired COS formation behaves in the same manner. These data are shown
in Figure 5-1. Note that the formation of COS increases very rapidly as high
S02 conversions are approached. The total SO- plus COS in the effluent gas
goes through a minimum at about 90 percent SO. conversion, and at this point
the COS concentration is about 1.5 times the SO- concentration.
When the reducing gas contains water vapor as well as CO, the COS
formed can be hydrolyzed to H2S. That is,
COS + H20 -*- H2S + C02.
This reaction may require a catalyst different from that used for the reduction
reaction. The product of a reduction plus hydrolysis scheme would be H?S and
not the desired elemental sulfur. Thus, it appears that the concept of reducing
with CO to obtain elemental sulfur has not been adequately demonstrated as of
this time.
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TABLE 5-1. DATA ON REDUCTION OF SO WITH HYDROGEN
Initial Concentrations: S0_ = 2.67 mole percent
H_ = 5.33 mole percent
Temperature
°C °F
345
360
375
390
653
680
707
734
Equilibrium Values
SO- Conversion S/H-S Residence Time
00 Ratio (sec.)
93.9 6.9 0.83
92.9 5.7 0.81
*•» { l:ll
89.6 3.4 6.78
Experimental
S00 Conversion
L (%)
16.9
24.3
28.1
44.9
36.8
Data
S/H S
Ratio
22.4
15.9
10.0
7.5
7.7
H S/SO
Ratio
0.009
0.019
0.036
0.096
0.067
Source: Murdock, D. L. and Atwood, G. A., "Kinetics of Catalytic Reduction of Sulfur Dioxide
with Hydrogen", Ind. Eng. Chem. Process Des. Develop. 13_, No. 3, p. 254, 1974.
-------
18
too
S»0
90
40
30
CONTACT TIME . 23 SECS
INLIT SO, • 2000 pern (NOMINAL)
INUT CO, • 14%
• • CO RATIO •
• - CO RATIO •
1.36 - 1 44
.»! • I.OS
TOO 723
790 773 «00
TEMPERATURE, -r
S23
Dependence of SOj conversion on temperature
too
LI 1.2 I.)
CO RATIO
Dependence of SOj conversion on carbon monox-
ide concentration
,50 -
• OUCMIOO'I DATA (1*711
• cumcNT worn • cxrr LINC AT rj«r
SYSTEM VMIAM.C3
CATALYST' HAMSNAW C« 0803
Mimae sizes- 20/30 MESH ANO i/««.
TCMtMATUMC' 700 • 9BCTf
CONTACT TIME • .210 • .«IO SCCS
CO KAPO- .SO • l.«
203O4OSOtOTOM*OKM
SO, REDUCED. KRCENT OF INLET SO,
Dependence of COS production on SO; conversion
s
I*
40-
30-
20t-
• OUEHIOO'S DATA (1*711
• DATA FROM CURRENT WORK
NOTE-
TOTAL SULFUR COMPOUNDS
DEFINED AS SO. AND COS
IN EFFLUENT.
o 20 40 to to no
SO, MCOUCEO, PERCENT OF INLET SO,
Dependence on total remaining effluent SO: and
COS upon SO] conversion
CO Ratio
CO cone, ppm
(2)(S02 cone, ppm)
Source: Quinlan, C. W., et al, "Kinetics and Yields for Sulfur Dioxide Reduction
by Carbon Monoxide", Ind. Eng. Chem. Process Des. Develop., !£ (1), p 107,
1973.
FIGURE 5-1. DATA ON REDUCTION OF S02 WITH CARBON MONOXIDE
-------
19
Hydrogen/Carbon Monoxide Mixtures as Reductant
The S09 reduction process which is being developed by the American
Smelting and Refining Company (ASARCO) and Phelps Dodge Corporation '
involves the reduction of S0? to sulfur using H^/CO mixtures as the reductant.
A pilot plant of this process is in operation at ASARCO's El Paso, Texas,
copper-lead smelter. A flowsheet of this pilot plant is shown in Figure 5-2.
Actually, ASARCO has developed the SO reduction process and Phelps Dodge has
developed the process for reforming natural gas to provide the reducing gas.
The ASARCO process uses a two-stage reactor system. In the first
stage, the SO^-rich stream and the H-/CO-rich stream are reacted over a
proprietary catalyst. The conversion of S0? to sulfur in this reactor is
about 69 percent when reducing an SO- stream containing 12 percent SO- and
is about 80 percent when reducing a pure SO- stream, v ' The gases are
introduced into this reactor at a temperature of about 340 C (650 F) „ The
reactor is a vertical she 11 -and -tube heat exchanger with catalyst in the
tubes and a heat transfer fluid circulating through the shell side to control
the reactor temperature.
The second stage is a Glaus reactor. Part of the SO- is reduced to H
in the first stage, and the reaction conditions are adjusted to obtain the
desired 2:1 H^S/SO- ratio in the product gas.
The H-/CO mixture used as the reductant in this process is generated
by the partial oxidation of natural gas. The principal reaction is
CH4 + 1/2 02-*- 2 H2 + CO.
Thus
, the reducing gas contains HL and CO in a molar ratio of 2:1. The
overall SO^ reduction reaction (both stages), therefore, can be written as
3/2 S02 + 2 H2 + CO -*- 3/4 S2 + C02 + 2 H20.
The ASARCO reduction technology is very close to that required when a
coal-based reducing gas is used (as will be seen later in the report), but
one difference is that in the coal gas the H^/CO ratio is not 2 but rather
-------
Molten Sulfur
Air
Sulfur
burner
Steam
S02
Tail Gas Recycle
Boiler
Preheated Air
I
Natural Gas
Reformed
Gas
Reformer
ALTERNATE SO2 SOURCE
(
S02 storage
tank
Liquid
S02
S02
Primary
reactor
Steam
Vaporizer
Primary
condenser
i
L1(luid
Sulfur
Secondary
reactor
Secondary
condenser
Li
O
Incinerator
Natural
FIGURE 5-2. FLOWSHEET OF ASARCO-PHELPS DODGE S02 REDUCTION PILOT PLANT
-------
21
about 0.5-0.6. The water gas shift reaction
CO + H20 ->• H2 + C02
can be used to increase the H2/CO ratio, but this adds another step to the
process. An important question is whether a gas with an H /CO ratio as low
as 0.5 can be used in the ASARCO reduction system. ASARCO is now studying
the use of coal gas as a reductant in their system.
If the H«/CO molar ratio in the reducing gas is designated as R,
then the general overall equation for the desired SO- reduction is
on 2 9 OP
S02 + ST H2 + m co -*• 1/2 S2 + SH C02 + m H2°«
Two moles of H2 + CO are required per mole of SO..
Natural Gas as Reductant
The SO- reduction process which has been developed by the Allied
Chemical Company accomplishes a reduction of SO- to sulfur using
natural gas as the reductant. A flowsheet of this process is shown in Figure
5-3. The process is similar to the ASARCO process except that the natural
gas is used directly instead of being reformed. The process uses a two-stage
reactor system. In the first stage, the SO -rich stream and natural gas are
fed to a reactor containing a special, proprietary catalyst. About 40 percent
of the inlet S0_ is converted to sulfur in this reactor, and this fact makes
this process appropriate for discussion in this section. The overall
reaction involved is:
CH^ + 2 S02 -»~ S2 + C02 + 2 H20.
H.S is clearly an intermediate in this reaction; Allied states that one of
the accomplishments in the primary reduction stage is to generate suffi-
cient H-S to provide a 2:1 H-S/S02 ratio for the ensuing Glaus reaction.
The important fact is that Allied's catalyst does result in some reduction
of SO- to sulfur in a single step.
Since only about 40 percent of the SO is reduced to sulfur
directly in the first stage, a Glaus reactor is used as a second stage to
recover the rest of the sulfur. Some S02 is reduced to-H-S in the first
-------
Steam Reducing
• Gas i
S02 Gas - , . > = i == £o) 2
Preheoter ) Main
blower
z>9=cr
ticeam
i
/•-». 1
^••n/ ""^~-^"~^ i *
^LJilFeli^J^ ^>'=d
Culf nr^^S^lS / .-
DUITUT ^5*-,_/
condenser | Converters 8
•• ' •> x
Culfni-
o u 1 1 u r ' •
condenser
Sulfur-holdinc
pit
Feed gas
heater
I -It
1
^ Ciilfur
DUITUr 1 1 •
condenser |
Reduction Steam
reactor . «
system • *V^)^==..^
Reactor e xi t^j^S--/""1
gas cooler '
Fuel *
Gas *
-------
23
stage, and the reaction conditions are adjusted to get the desired 2:1
H-S/SO ratio in the product gases. The Allied process reportedly
produces very little COS and CS .
The feasibility of this process has been demonstrated commercially
by treatment of the S0« effluent stream from smelters at the Falconbridge
Nickel Mines, Ltd. facility near Sudbury, Ontario, Canada. Since this process
requires natural gas, it is not considered "acceptable" within the scope of
this study but rather is used as a baseline for comparison with other options.
It should be noted that Allied is now working on the extension of this process
to higher molecular weight hydrocarbon feedstocks.
Hydrogen Sulfide as Reductant
The reduction of SO- with H~S is the basis of the commercially
proven Glaus process. This process is shown as part of the Allied process
in Figure 5-3. In this figure, converters 7 and 9 and the following equipment
constitute the Glaus plant, A two-stage catalytic converter with intercooling
and condensing is used. A bauxite catalyst is used. The reaction temperatures
are 230-400 C (450-750 F), with lower temperatures maintained in the
(13)
second stage than in the first stage. This is done because the
reaction is exothermic and the equilibria are less favorable at higher
temperatures. The conversion is 70-80 percent in the first stage and
about 95 percent overall. A third stage can be used to increase the conversion
somewhat, but this is not justified when the Glaus plant follows an S0«
recovery facility to which the tail gas can be recycled. The tail gas must
be oxidized before being recycled in order to convert the H S into SO..
Coal as Reductant
The RESOX process ' developed by Foster-Wheeler Corporation
employs anthracite as a noncakinga low-volatile, solid carbon reducing agent
for SO-, with elemental sulfur as the end product. The overall reaction is
C + S02-*~ C02 + 1/2 S2.
In the RESOX process, rice-size anthracite enters the top of a vertical
tubular reactor by gravity through a rotary airlock valve. The concentrated
-------
24
SO. stream flows countercurrently upward with air injected at several ports
around the lower vessel circumference. The reactor is operated under atmos-
pheric pressure in the temperature range of 590-820 C (1100-1500 F). Typical
gas residence time is 3-4 seconds. Foster-Wheeler reports that conversion of
90-95 percent of the inlet SO to elemental sulfur can be realized.
This process was scheduled for commercial demonstration at Gulf
Power Company's Scholz Plant at Chattahoochee, Florida, during 1975.
However, these tests, at least on the reduction process, have not been
successful as of this writing.
5.2 Reduction of SO to H S_
Hydrogen/Carbon Monoxide Mixtures as Reductant
SO- can be reduced all the way to H«S with hydrogen and/or carbon
monoxide. A high activity hydrogenation catalyst can be used for this purpose,
and the cobalt-molybdenum catalysts used extensively for hydrogenation
reactions would appear to be a logical choice. However, the use of such a
catalyst for reducing a concentrated (>25%) SO- stream has not yet been demon-
strated. Reduction of dilute (<5%) SO- streams has been done as part of two
Glaus tail gas treatment processes, and reduction of a concentrated stream
has been proposed by Shell as part of their FGD process. These items will be
discussed in this section.
The ability of CO to serve as a reducing agent arises through inter-
mediate reactions involving water vapor, such as
CO + H20 -*- H2 + C02 (water-gas shift)
COS + H20 -*• H2S + C02 (COS hydrolysis).
The water vapor can be either initially present in the gas or generated by
the reactions involving hydrogen. If the H /CO molar ratio in the reducing
gas is designated as R, then the general overall equation for the reduction
reaction is
To 7 -5 7R-1
S°2 + FT H2 + ST C° - V +^T C02 + §T H2°'
-------
25
The two Glaus tail gas treatment processes in which H /CO mixtures
are used to reduce SO- are the Shell Glaus Offgas Treating (SCOT)
process ' and the Beavon sulfur removal process. ' In both
processes the reduction step converts all the sulfur compounds in the
tail gas to H2S, which is then recovered. The H-S recovery is done by
amine scrubbing in the SCOT process and by the Stretford system for the
Beavon process.
In the SCOT process the reduction is conducted over a bed of cobalt-
molybdenum on alumina catalyst at a temperature of about 300 C (570 F). With
an excess of reductant, virtually complete conversion to H S is obtained. The
water-gas shift reaction is very rapid at the conditions of the SCOT process.
The presence of CO in the reducing gas results in an increased reaction rate
as compared with reduction with H. alone. This is shown in Figure 5-4. At
the SCOT conditions the concentrations of COS and CS_ in the effluent gas
from the reducer approach the thermodynamic equilibrium values. Methanation
reactions are negligible.
The Beavon process also uses a cobalt-molybdate catalyst and uses
"moderate conditions of temperature and pressure, resembling those in the
Glaus unit.' The catalyst is "very effective in reacting water vapor
present with carbon monoxide to form hydrogen, or reacting the water vapor
with COS and CS to give H0S."^ ' The space velocity of tail gas plus
(10)
reductant in the reduction reactor is about 2000 vol/hr/vol of catalyst.
The catalyst is reasonably rugged but evidently can be damaged by some trace
species. Reference 17 states "Because there are no heavy metals or halogens
to poison the catalyst, or tars to clog the system, the catalyst should last
for several years."
It should be re~emphasized that, as tail gas treatment processes,
the SCOT and Beavon processes are designed for BO streams containing less
than 5 percent S0_.
The S0» reduction system proposed by Shell for use with their FGD
/i a\
proce^o is shown in Figure 5-5.. The reduction occurs in three stages
with intercooling, the last two stages being Glaus reactors. The first
stage, labeled "partial reduction reactor", involves a reduction of SO to
H-S with a HL/CO mixture. Neither the intended catalyst nor the H.-/CO ratio
are specified. This reduction system has been proposed by Shell but has not
been demonstrated in conjunction with the basic Shell FGD process.
-------
26
percen
c
0)
o>
<
CT
c
'o
3
•a
a>
01
•«—
o
c
.o
'35
v
c
o
o
o
o
u>
o
o>
o
->i
o
-------
27
H2/CO-Rich Stream
S02-Rich Stream
RFRC -
Tail Gas Recycle 170 °C
Partial reduction
reactor
First Claus
reactor
Second Claus
reactor
Steam
FIGURE 5-5. FLOWSHEET OF CONCEPTUAL REDUCTION SYSTEM
PROPOSED FOR SHELL FGD PROCESS
-------
28
6. USE OF PURCHASED HYDROGEN SULFIDE
The possibility of using purchased hydrogen sulfide as a reductant
for flue gas desulfurization processes has been mentioned in previous sections.
An analysis was made of the feasibility of this option. This analysis is
summarized in this section, and additional details are presented in Appendix B.
Problems associated with the use of hydrogen sulfide as a reductant
include: extreme toxicity, corrosion properties, explosive hazard, transpor-
tation, storage, adequate supply and unfamiliarity with the product by operating
personnel at electricity generating plants.
It is estimated that for 1974, the total hydrogen sulfide potentially
available from refineries and natural gas plants in the United States was
3
approximately 15.2 x 10 metric tons per day, increasing to approximately
3
21.4 x 10 metric tons per day by 1980. This is the total hydrogen sulfide
from all sources and includes the large percentage converted to elemental
sulfur in Glaus plants or burned to SO- and vented to the atmosphere.
It is expected that plants in which coal or oil is gasified or
desulfurized will provide additional potential sources of hydrogen sulfide in
the future; however, their contribution to national production will be insig-
nificant through 1985.
Sulfur dioxide emissions from fossil-fired plants will exceed the
potential supply of hydrogen sulfide for use as a reductant in all power plants
by a factor of at least 3 in 1980. The potential demand for hydrogen sulfide
is much greater than any anticipated supply. Nevertheless, there is potentially
enough hydrogen sulfide available to reduce a significant fraction of the
sulfur oxides produced at power plants.
Hydrogen sulfide presently is available at prices ranging from
$143 to $221 per metric ton (6.5-10 cents per pound),* f.o.b., and most
hydrogen sulfide probably is sold at the lower figure. The current prices
have little relevance to the price of large volumes of hydrogen sulfide
that might be required by an electric utility as a reductant (approximately
* Because of the dominance of English units of measurement in the indus-
tries to which this report is addressed, some English units are used in
this report. Both units are often given together. A table of conversion
factors between English and metric units is presented in Appendix H.
-------
29
693,000 liters [183,000 gallons] per day for a 1000-MW power plant), and
the price would probably be equal to or slightly less than the price of
the contained sulfur. Results contained in Section 11 of this report
indicate that the cost of producing ELS via gasification of coal or
residual oil would be between $55 and $71 per metric ton (2.5-3.2 cents
per pound) of H.S, depending on the feedstock cost and plant size. Thus,
a power plant having access to an "across the fence" source of H S would
probably need to purchase H-S at these price levels or lower for this
approach to prove economically attractive. Another factor to consider is
that if H-S is purchased by an electric utility to reduce the SC>2 it
generates, the utility is tripling the amount of sulfur that it has to
market.
At the present time, liquefied hydrogen sulfide is transported
only by truck or railroad car at a cost of approximately 3 to 4 cents per
metric ton-mile„ Barge transportation probably would cost less but would
be of limited applicability. Pipeline transportation is impractical except
for short distances. It may be feasible if the power plant is located in the
vicinity of a refinery or other hydrogen sulfide source.
In summary, the potential supply of hydrogen sulfide is sufficient
to handle only a portion of the power plants which will be installing FGD
systems. Geographical considerations will further restrict the use of hydrogen
sulfide at power plants. Thuss it: is concluded that purchased hydrogen sulfide
will play only a minor role as a reductant for FGD systems.
-------
30
7. GENERATION OF REDUCING GAS BY GASIFICATION
Before quantifying the trace constituents in gasifier-produced
reductant gases and defining their effects on FGD systems, it is necessary
to examine the characteristics of the gasification systems themselves.
Coal and petroleum residuum can be gasified with steam and oxygen (or air)
to produce a gas rich in hydrogen and carbon monoxide. This gas can be
used as a fuel, either directly or after upgrading to increase its
heating value. However, the intention in this study is to use this gas
not for its fuel value but for its reducing power.
The consideration in this study is limited to gasification
systems which are commercially proven and available. For the gasification
of coal, the consideration is also limited to gasifiers operating near
atmospheric pressure. This is so because the reducing gas is required
only at low pressure, since the FGD processes with which it will be used
operate at essentially atmospheric pressure. Thus, there is no reason to
incur the additional complication and expense of a lock-hoppering system
for feeding the coal into a pressurized gasifier. The following atmos-
pheric pressure gasifiers are accepted as commercially proven and available
for commercial usage in the U.S.:
(1) Koppers-Totzek gasifier
(2) Wellman-Galusha gasifier
(3) Winkler gasifier
(4) Riley-Morgan gasifier.
For the gasification of petroleum residuum, the following
gasification processes are commercially available:
(1) Shell Gasification Process
(2) Texaco Synthesis Gas Generation Process.
These processes operate on a variety of petroleum feedstocks including
atmospheric and vacuum residua. Also, coal tar can be used as a feedstock
for the Texaco process.
Additional information on the gasification processes listed
above is presented in Appendix C. The following sections highlight the
features of these processes and the compositions of the gases that they
produce.
-------
31
7.1 Coal Gasification Processes
The Koppers-Totzek gasifier must be blown with oxygen (or at
least an oxygen/air mixture) to reach the high operating temperature
used. The other three gasifiers can be blown with either oxygen or air.
When oxygen is used, these gasifiers generate a gas which contains primarily
H« and CO and only very little N? (usually 1-2 percent). When air is used,
the gas is diluted by the N« in the feed air, the N- concentration being
about 50 percent. As compared with air-blown systems, the oxygen-blown
systems will result in:
• Smaller gasification vessels, gas cleanup systems,
reduction reactors, and gas transfer ducts
o Higher reaction rates in the reduction reactors
because of higher partial pressures of the reacting
species.
The decreased investment for the smaller equipment tends to be offset by
the cost of the oxygen plant. Differences in operating cost items are
also important, such as the fact that oxygen plants require considerable
quantities of electrical power. The choice between air- and oxygen-blown
systems requires a detailed analysis of the entire gas generation and
utilization system for the specific situation involved.
For a 1000-MW power plant fired with 3.5 percent sulfur coal,
the gasification system to supply reductant for the FGD process will
be equivalent to about 270 metric tons/day of coal feed and to about 1.4 x
9 9
10 kcal/day (5.6 x 10 Btu/day) of gas output. This corresponds to
several of the smaller gasifiers or one of the larger ones (Koppers-Totzek
two-headed). For comparison, the feed rate of coal to the power plant
itself is about 8,840 metric tons/day. If the gasifiers are oxygen-blown,
the capacity of the oxygen plant will be about 200 metric tons 0 /day.
The operating characteristics of the four coal gasifiers are
compared in Table 7-1. Due to very high temperature operation, the K-T
does not produce ammonia or methane in any significant quantities, but
the gas contains up to 50 percent of the ash in the coal as fly ash. The
K-T has the lowest overall thermal efficiency (67 percent) of the four
gasifiers, whereas the Wellman-Galusha has the highest efficiency (82
percent).
-------
TABLE 7-1. COMPARISON OF SELECTED COMMERCIAL ATMOSPHERIC PRESSURE COAL GASIFIERS
Oxidizing medium
Gaslfler
Supplier
Casifler Bed Type
Commercial Units Built
Gas Discharge Conditions
Temperature, C (F)
Pressure, Inches water (gauge)
Composition, mole percent
CO
H2
co2
CU4
CH4+
N2 + Ar
H2S, etc.
HHV, Btu/scf
Gas Produced, scf/lb coal (approx.)
Overall Thermal Efficiency, percent
Turndown, percent of full load
Feed Coal Size
Pure Oxygen Feed, lb/10 Btu cold gas
Steam Feed, lb/10 Btu cold gas
Export Steam Produced, lb/10 Btu
cold gas
Coal feed, lb/10 Btu cold gas
Oxygen/Coal Weight Ratio
Gasification Rate, Ib coal/hr-ft
Oxygen-Steam
Koppers-Totzck
Koppers Company
Inc., Pitts-
burgh, Pa.
Entrained
53
1480 (2700)
70
56
35
7
0
0
2
(b)
290
31
67
35 (2-burner)
65 (4-burner)
90Z(-200)mesh
80
30
173
114
0.7
16
Rl ley-Morgan
Riley-Stoker
Corp. , Wor-
cester, Mass
Fixed-bed
1
650 (1200)
35
41
39
16
3
0.7
0.5
(b)
305
33
70
ND
ND
86
0
110
0.46
ND
Wellman-
Galusha
McDowell-
Wcllman
. Engg. Co.,
Cleveland,
Ohio
Fixed-bed
0
680 (1250)
70
52
33
12
1
0
2
(b)
282
ND
82
90
1/2 - 2 inches
90
50
62
94
ND
99/ft2
Winkler
Davy Powergas,
Lakeland ,
Fla.
Fluldized
20
980 (1800)
30
40
37
18
3
—
2
(b)
280
26
76
50-150
<3/8 In.
70
20
ND
140
0.5
2-4
Rl ley-Morgan
Fixed-bed
0
35
26
19
4
1.6
0.2
50
(b)
163
ND
NA
NU
NA
ND
Air-Steam
Wellman-
Galusha
Fixed-bed
3
70
29
15
3.4
2.7
0
50.3
(b)
170
65-68
NA
ND
NA
ND
Winkler
Fluldized
16
30
34
31
12
0.5
—
22.5
(b)
120
ND
NA
ND
NA
ND
(a) Before waste heat boilers at gasifier exit.
(b) Depends on coal sulfur content. Usually about 0.6 percent for 2 percent sulfur coal, oxygen-blown gasifier.
(c) For 2-burner gasifier; 28 for 4-burner gasifier.
NA - Not applicable; ND - No data.
Conversion factors: kg/m2 - 25.40 (inches water), kcal/std o3 - 8.935 (Btu/scf), std m3/kg coal - 0.06243 (acf/lb coal), kg/106 kcal - 1.793
(lb/10 Btu).
Co
ro
-------
33
All the gasifiers generate CO and K- with the ratio CO/H greater
than 1. For some applications, a lower CO/H_ ratio (i.e., more H~) may
be desired. This ratio can be affected, but only slightly so, by changing
the ratios of steam and 0- (or air) to coal fed to the gasifier. Shifting
the CO to make more H is also possible, since sulfur-resistant shift
catalysts are available..
7.2 Residual Oil Gasification Processes
Typical gas compositions for the two heavy oil gasification
processes are given in Table 7-2.,
-------
TABLE 7-2. TYPICAL GAS COMPOSITIONS FROM OIL GASIFICATION PROCESSES
(Dry Basis)
Shell Process
Feeds tock
Oxidizing Medium
Composition, volume percent
H2
CO
co2
CH.
4
N + argon
(a)
H2S and COSV '
Heavy Fuel
Oil
Oxygen
46.7
46.2
4.3
0.6
1.4
0.8
Propane
Asphalt
Oxygen
44.2
47.8
4.5
0.6
1.3
1.6
Texaco Process
Vacuum
Residuum
Oxygen
47.1
50.8
1.2
0.6
1.2
0.04
Air
14.5
23.6
1.0
0.2
60.7
0.02
Straight
Residuum
Oxygen
47.0
47.0
5.6
0.1
0.2
0.1
(a) Depends on feed sulfur content.
-------
35
8. TRACE CONSTITUENTS IN GASIFIER PRODUCTS
Available experimental data on trace constituents in products
from various gasifiers are presented in this section. These data are scanty
and do not provide a sufficient basis for evaluating the effects of trace
substances on user processes. Therefore, thermodynamically calculated
values of the equilibrium composition in gases at various temperatures are
presented based on a modified equilibrium constant approach. Because of
the need to compare the relative contribution of trace constituents by
both the flue gas being scrubbed and the coal gas used for sulfur production,
the trace constituent analyses of coal-fired boiler flue gas also are
presented.
8.1 Experimental Data on Gasifier
Product Trace Constituents
Koppers-Totzek Gasifier
Experimental trace constituent data on the K-T gasifier are
(19 20 21)
limited to the concentrations of gaseous species in product gases. ' '
All operating K-T gasifiers are located outside the USA in countries where
public concern for the environment: is less intense than in this country.
However, available trace data are tabulated (Table 8-1) as a function of
coal sulfur content and heating value. It can be noted that the total
sulfur compounds formed are about 0,3 mole percent when the feed coal has
1 percent sulfur. Acid gas treatment steps can be employed to reduce the
sulfur content.
Wellman-Galusha Gasifier
Data on W-G gasifier trace constituents are not available.
Thus, the computer calculated values will be the only source'of data.
Also, since the Riley-Morgan gasifier operating conditions are very similar
-------
TABLE 8-1* EXPERIMENTAL DATA ON KOPPERS-TOTZEK GAS TRACE CONSTITUENTS
Reference Number 20
Gas Location Gasifler Outlet
Coal
Ash, weight percent
S, weight percent
HHV, Btu/lb
Gas Composition, (volume
percent dry basis)
CO
H2
CO,
N2
H2S
COS
so2
HCN
NO
Ar
Particulate matter,
grains/scf
52.76
35.55
10.07
0.42
0.11
0.32
0.025
31 vppm
0.24
0.041
10 vppm
0.45
11.57
11.57
20 20 21 22
To Compression and Product Gas After Acid Gas Gasifler Outlet
Acid Gas Removal Removal
52.77
35.55
10.04
0.43
0.12
0.32
0.025
16 vppm
0.23
0.032
7 vppm
0.45
0.002
0.002
17.3
0.95
8,830
53.17 60.88
36.52 32.60
9.44 5.23
0.44 1.16
0.12 0.10
3 vppm 0.02
1.5 vppm 0.01
0.5 vppm
1 vppm
1 vppm
3 vppm
0.46
0.002
0.002
22.14 19.12
0.67 4.88
9,888 11.388
58.68 55.38
32.86 34.62
7.04 7.04
1.12 1.01
0.28 1.83
.02 0.12
CO
ON
Conversion factors: kcal/kg - 0.5578 (Btu/lb), grams/std m3 - 2.288 (graino/scf)
-------
37
to the W-G operating conditions, the experimental data presented for the
R-M gasifier can be adapted if necessary to the W-G unit. Further, there
are no operating oxygen-blown W-G gasifiers, so that obtaining experimental
trace constituents data on W-G gasifiers will not be possible.
Riley-Morgan Gasifier
For an oxygen-blown R-M gasifier the available experimental
trace constituent data for raw and clean gases are presented in Table
8-2. It can be seen that the data are totally inadequate with respect
to trace elements. However, ammonia is present in the clean gas, even after
liquid scrubbing, to the extent of 1800 ppm (0.18 volume percent). The
effect of ammonia on the Glaus catalyst should be watched.
Winkler Gasifier
Although commercial Winkler gasifiers have operated for over
30 years, the product gas has not: been analyzed for trace constituents.
A U.S. team which visited commercial gasification plants during 1974 has
reported that the staff of Azot Sanayii (Kutahya, Turkey) was engaged in
(21)
analyzing various streams for heavy metals. This did not include the
actual product gas analyses. Thus, experimental data on trace constituents
are not available for the Winkler gasifier.
Synthane Process
During 1973-75, a series of gasification tests was run in
the Synthane process development unit at Bruceton, Pennsylvania, to
(22)
define the trace elements' distribution in gasifier effluents.
The gasifier operates at 40 atmospheres pressure and has a capacity of
20-40 Ib/hr coal. Trace elements in the gas were determined by neutron
activation "but results were incomplete and not conclusive" due to the
extremely small amounts of trace metals present. For example, Hg, the
only element measured, was on the order of 0.0001 ppm by weight.
-------
38
TABLE 8-2. TRACE CONSTITUENT DATA FOR RILEY-MORGAN GASIFIER
(Oxygen-Blown System)
Raw Gas
Clean Gas
Participate matter, grains/scf 1-75 0
Tar, grains, scf 15.00 0.003
(submicron aerosol)
Ammonia, ppmv
HCN, ppmv
2 '
COS, ppmv
cs2
Tar Level
Btu/scf
2000 1800
<100 traces
3000 1200 (a)
300 250 (a)
traces
300
Coal Data
Ash, weight percent
Sulfur, weight percent
HHV, Btu/lb
8.0
1.0
14,000
(a) These levels can be reduced to any desired lower value by
further scrubbing.
Source: Extracted from letter correspondence between BCL and
Riley Stoker Corporation, Worcester, Mass.
Conversion factors: grams/std m = 2.288 (grains/scf), kcal/std
nT = 8.935 (Btu/scf), kcal/kg = 0.5578 (Btu/
Ib).
-------
39
Hygas Process
An experimental program, has been conducted at the Institute of
Gas Technology using samples from their pilot plant of the Hygas process.
This process is fundamentally different from the other gasification
processes discussed in this report in that it involves a hydrogasification
of coal with a hydrogen-rich gas stream instead of using oxygen or air.
This increases the production of methane in the gasifier. The hydrogasi-
(23)
fication is conducted at about 650 C (1200 F) and 74 atmospheres pressure.
(23)
The first phase of this program involved analyses of the
coal and of solid materials at various stages of processing. No gas
samples were analyzed. Assessment of these data indicate that significant
fractions of the arsenic, cadmium, lead, mercury, selenium, and tellurium
in the coal went into the gas. However, the report stated that "since
these results are based on a limited number of samples, further analysis
of a larger number of samples is required before any firm conclusions can
be drawn".
The second phase of this program did involve some analyses of
gas samples, but the results have not yet been published. Preliminary
(24)
results indicate the presence of some arsenic, selenium, and germanium
in the gas.
The above discussions illustrate the extent of data unavailability
as well as the difficulty involved in chemically analyzing the gas for
trace constituents when they are present in parts per billion levels. A
partial solution to circumvent these difficulties will be to estimate
trace constituents in the gas by employing thermodynamic equilibrium
calculations based on reactivity and volatility of the trace constituents.
Data from this approach are presented as follows.
-------
40
TABLE 8-3. MEAN ANALYTICAL VALUES FOR CONSTITUENTS IN 101
DIFFERENT COALS ON WHOLE-COAL BASIS
Constituent
As
B
Be
Br
Cd
Co
Cr
Cu
F
Ga
Ge
Hg
Mn
Mo
Ni
P
Pb
Sb
Se
Sn
V
Zn
Zr
Al
Ca
Cl
Fe
K
Mg
Na
Si
Ti
Org. S(a)
Pyr. S
Sul. S
Tot. S
SXRF
Ash
Btu
C
H
N
O
Mean
14.02 ppm
102.21 ppm
1.61 ppm
15. 42 ppm
2. 52 ppm
. 57 ppm
13. 75 ppm
15.16 ppm
, 60.94 ppm
3. 12 ppm
6. 59 ppm
0.20 ppm
49.40 ppm
7. 54 ppm
21.07 ppm
71. 10 ppm
34.78 ppm
1 . 26 ppm
2.08 ppm
4.79 ppm
32. 71 ppm
272.29 ppm
72.46 ppm
1.29%
0. 77 %
0.14%
1.92%
0.16%
0.05%
0.05%
2.49%
0.07%
1.41 %
1.76%
" 0. 10 %
3.27%
2.91%
11.44%
12,748.91
70.28%
4. 95 %
1.30%
• 8.68%
Standard Deviation
17.70
54.65
0.82
5.92
7.60
7.26
7.26
8.12
20.99
1.06
6.71
0.20
40.15
5.96
12.35
72.81
43.69
1.32
1.10
6.15
12.03
'' 694. 23
57.78
0.45
0.55
0.14
0.79
0.06
0.04
0.04
0.80
0.02
0.65
0.86
0.19
1.35
1.24
2.89
464.50
3.87
0.31
0.22
2.44
Minimum
0.50
5.00
0.20
4.00
0.10
1.00
4.00
5.00
25.00 '
1.10
1.00
0.02
6.00
1.00
3.00
5.00
4.00
0.20
0.45
1.00
11.00
6.00
8.00 .
0.43
0.05
0.01
0.34
0.02
0.01
0.00
0.58
0.02
0.31
0.06
0.01
0.42
0.54
2.20
11, 562.00
55.23
4.03
0.78
4.15
Maximum
93.00
224.00
4.00
52.00
65.00
43.00
54.00
61.00
143.00
7.50
43.00
1.60
181.00
30.00
80.00
400.00
218.00
8.90
7.70
51.00
78.00
5,350.00
133.00
3.04
2.67
0.54
4.32
0.43
0.25
0.20
6.09
0.15
3.09
3.78
1.06
6.47
5.40
25.80
14,362.00
80.14
5.79
1.84
16.03
(a) Abbreviations other than standard chemical symbols: organic sulfur (Org. S), pyriric sulfur (Pyr. S). sulfate
sulfur (Sul. S), total sulfur (Tot. S). sulfur by X-ray fluorescence (SXRF). air-dry loss (ADL). moisture (Mots.).
volatile matter (Vol.). fixed carbon (Fix. C). high-temperature (HTA). low-temperature ash (LTA).
Source: Reference 25.
-------
41
8.2 Thermodynamically Predicted Trace Constituent Data
Thermodynamic predictions were made of the distribution of
species between the gaseous and condensed phases for a number of sets
of operating conditions. These calculations were based on the average
(25)
analysis for 101 coals determined by the Illinois Geological Survey,
which is shown in Table 8-3. The operating conditions for the gasifiers,
i.e., the temperature, steam/coal ratio, and oxygen/coal ratio, were taken
from Table 7-1.
The technique employed to compute thermodynamic equilibrium
(26 27^
values was based on the EQUICA Program developed by Battelle. '
Detailed results of computed values of trace substances for the various
gasifiers using the EQUICA Program are presented in Appendix D. Salient
data are summarized in this section.
A comparison of experimental and calculated gas compositions for
the Koppers-Totzek gasifier is shown in Table 8-4. The data indicate close
agreement on the major gaseous components, namely, CO and H^. On the
other hand, the agreement is poor for CO-, CH. , and NH,,. Since experimental
values for trace elements are not available, the validity of the computer
calculated trace element concentrations cannot be checked. The approach
to equilibrium, and hence the validity of the calculated values, is probably
reasonably good when the temperature of the equilibrium mixture is above
1000 C (1800 F) . Below this temperature, the assumptions of thermodynamic
equilibrium and stability may not be valid. Thus, when predicted gaseous
compositions for 315 C (600 F) were compared with experimental values
for major components (CO, H~), the deviations were extremely large.
This confirms earlier observations (see Figure 8-1) on the predictive
(27)
power of thermodynamic calculations at temperatures below 1000 C.
Insufficient time was available for investigating the causes of these
deviations and, when possible, for modifying the program to avoid these
difficulties.
The calculated gas compositions are summarized in Table 8-5.
The detailed data are presented in Appendix D. It should be noted that these
compositions apply to the gas phase only and do not include any fly ash
suspended in the gas. Removal of particulate matter is considered
-------
42
TABLE 8-4. COMPARISON OF PREDICTED AND EXPERIMENTAL K-T GAS COMPOSITION
Composition
at 2700 F, Gasifier Exit
Concentrations, mole percent
CO
H2
co2
N2
V
COS
Concentrations, volume ppm
cs2
H2°
CH4
NH3
Concentrations, weight ppm
Fe
K
Na
so2
Mg
Cu
Pb
Se
C12
Hg
Experimental
Values
56.0
35.0
7.0
1.0
1.0
0.1
ND
1.0
800
1700
ND
ND
ND
ND
ND
ND
ND
ND
ND
ND
Predicted Values
60.9
38.7
0.3
0.5
0.8
0.08
20
0.73
0.1
1.5
2768
878
270
3
211
8
19
1
~0
0.112
ND = Not determined.
-------
43
Mo*
Days
Time
Hrs
Sec
\ ,.,—.,. -_
\ 9 AEC-NUCLfcAH REACTOR
v AEC-SNS
\ PuO,
\
NAVAL TURBINE COATINGS E3B\
••VViS^^V
BaddBBB
SEMI-
\ QUANTITATIVE QUANTITATIVE
GLASS MELT * \
HEATER ELEMENT \
SEMI-QUANTITATIVE v
NATURAL GAS BURNER • \
\
V
V \
\, \
X>. \
^S . N. ROCKET PROPULSION
^^ v
r\\ I A! IT ATlWC/Ti* ^ AND COATINGS
•v^ COAL COMBUSTION ,
_ „ ,_. \ fcJBi SANDIA PLASMA O ^
ftASTM-CHETA..I,l ^Xj POWDER PRODUCTION , \J
1000
2000
3000
: FIGURE 8-1. BATTELLE EXPERIENCE WITH TIME-TEMPERATURE EFFECTS
ON PREDICTIVE POWER OF THERMODYNAMIC CALCULATIONS
Note that Coal Combustion and Gasification Processes Fall into the
Semiquantitative Estimation Category
Source: Reference 27.
-------
44
TABLE 8-5. COMPARISON OP THEKIDDYNAMICALI.Y PREDICTED TRACE CONSTITUENT DATA
System
Temperature, C (F)
ppm bv weight in gas phase
H2S
COS
so.
Fe
K
Na
cs2
Mg
Cu
Pb
Se
C12
Hg
HgO
HgCl2
P4°10
As
ZrCl4
T1C14
Zn
P
Cd
Coal-Fired
Boiler
650 (1200)
0
0
9,755
0
0
0.408
0
0
0
0
0
0
0.098
0.004
0.012
0.001
0.231
12.85
0
0
0
0
0
. 0
0
K-T
Gasifier
1815 (3300)
14,292
1,358
3
0.1
2,768
878
270
81
212
8.3
9.5
1.14
0.008
0.112
0
0
0
0
7.7
0
5.1
0
150
39
1.38
K-T
Gasifier
650 (1200)
14,811
2,197
0
5.65
0
898
0.02
8.3
0
0
0.004
1.17
0.002
0.115
0
0
0
0
7.9
11.0
5.3
303
0
0
0
Winkler
Gasifier
980 (1800)
19,445
720
0
7.0
0
1,095
336
24
0
0
5.82
1.43
0
0.14
0
0
0
0.611
9.7
0.005
2.5
6.06
181
48
1.64
Wellnan-
Galusha
Gasifier
650 (1200)
12,352
587
11.5
0
697
0.01
0.3
0
0
8.6
0.91
0
0.07
0
0
0
71.3
6.1
0
0
0
0
0
0
Ri ley-
Morgan
Gasifier
680 (1250)
17,042
1,720
11.2
0
990
0.08
9.4
0
0
12.2
1.3
0
0.10
0
0
0
101
8.7
0
0
0
0
0
0
Total species present
in gas phase
19
14
17
11
11
-------
45
necessary before the gas can be used as a reductant. Data on the composition
of the ash also are included in Appendix D. The data in Table 8-5 support
the following conclusions:
(1) The flue gas from the coal-fired boiler, after
cooling to 650 C (1200 F) contains eight gaseous
trace species. This number is lower than that
for any of the gasifier products.
(2) Among the gasifier product gases, the general level
of trace species is determined primarily by the
gasification temperature. The Koppers-Totzek
gasifier operating at 1815 C showed 19 trace consti-
tuents, the Winkler gasifier operating at 980 C
showed 17, and the two lower temperature gasifiers
(Wellman-Galusha and Riley-Morgan) operating at
650-680 C showed 11 species.
(3) Cooling the Koppers-Totzek product gas to 650 C
(1200 F) and reestablishing equilibrium reduces
the number of gaseous trace species to 14. Note
that this is still higher than the value for the
two gasifiers which operate at this temperature
level.
Actually, of course, the gasifier product gas will be cooled
below 650 F during the cleanup steps, and hence the gaseous trace constituent
concentrations will be less than the values shown'in Table 8-5. If no
gaseous species have to be removed but only fly ash, a high-temperature,
high-efficiency electrostatic precipitator can be used. This can be done
at a temperature of about 430 C (800 F), thus eliminating the need for
reheating the gas prior to using it as a reductant. In the more likely case
that gaseous species such as NH have to be removed, the gas will have to be
scrubbed with water. This will require reheating the gas but will result in
even lower trace constituent concentrations in the gas, since equilibrium
will then be established (or at least approached) at a much lower temperature
(about 50 C) and some species will be at least partially soluble in the water.
8.3 Contacts with Process Vendors
The developers and/or vendors of several sulfur-producing FGD
and related processes were contacted to obtain their opinions regarding
the effects of trace components in coal-based reducing gases upon their
systems. The processes considered included the Peabody Citrex process,
-------
46
IFF Catalytic NH scrubbing process, and the American Smelting and Refining
Company (ASARCO) reduction process. The individuals contacted felt that
the trace components should not present any problems since: (1) the
proportion of the trace substances introduced by the coal gas is only
about 3.6 percent of the main boiler flue gas constituents and (2) the
small volume of the gas generated will be thoroughly scrubbed to give
near zero levels of particulate matter and most gas phase contaminants. One
purpose of the thermodynamic calculations was to examine if there was a
theoretical basis for the assumption of near complete removal of trace
elements in the coal gas by liquid scrubbing to near ambient temperatures.
This purpose was not completely achieved because of the difficulties
encountered with the lower temperature calculations.
-------
47
9. EFFECT OF TRACE CONSTITUENTS ON FGD SYSTEMS
In Section 7, several coal gasification processes were discussed.
In Section 8, the trace contaminants present in the producer gas were
examined. In this section the effect of the coal gas trace contaminants
on the FGD processes proper and on the subsequent SO™ reduction catalysts
will be evaluated.
9.1 Overall FGD Plus Reduction Systems
To understand the effect of trace elements on the FGD chemistry *
a knowledge of both the overall sulfur-producing conceptual systems and
the chemistry of the important steps which utilize the coal-based reducing
gas will be useful. The possible routes to sulfur production employing
various FGD processes and subsequent reduction processes are described in
Figure 9-1. The reduction processes include: (1) the Glaus sulfur plant,
(2) the ASARCO-Phelps Dodge direct sulfur process, and (3) the liquid
Glaus unit.
9.2 Effects on FGD Processes
\
An analysis of the data shown in Figure 9-1 indicates that:
(1) 3.6 pounds of coal will be fed to the coal gasifier
for each 100 pounds of coal fed to boiler for
generating steam. Thus, even if the trace species
volatilization ra.te is assumed equal for the gasifier
and boiler, the gasifier will introduce only 3.6
percent of the trace species which will be introduced
into the FGD system via the boiler flue gases.
(2) Prescrubbing of boiler flue gas with water is a
desirable step to eliminate all condensible and
soluble trace species at a temperature of 30-50 C
(90-120 F). Whether water prescrubbing will reduce
also the trace species in gaseous state, and if so,
to what degree has not been quantified experimentally.
(3) The FGD scrubber handles 144 volumes of flue gas from
boiler for every volume of reducing gas entering the
process from the gasifier.
-------
H2S +S02(traces)
Coal,100 Ib
CLAUS OR ASARCO SYSTEM
Toothers
Sulfur \ Uses
To
Process
Coal,3.6 Ib.
Ozor Air
Steam
Gasifier
/.
Tar. char,
and dust remover
(quench, etc.)
Gas Conditions and Predicted Trace Species
At CD At @
A* - Steps in Type A FGD Processes
B* = Steps in Type B FGD Processes
-Coal pile
FIGURE 9-1. POSSIBLE ROUTES TO SULFUR PRODUCTION BASED ON COAL
GASIFICATION REDUCING GAS IN FGD PROCESSES
(Schematic of the Overall FGD Process System)
Temperature, F
Pressure , •.
Trace elements, Ib/min
Fe
K
Na
CS2
Pb
Se
Cl-
Hg
HgO
SeO
P,0^
As
200
-1.0
.09
02
.0009
0026
.051
-83
200
-1.0
.436
0
0.004
0.005
-0
0
-0
0
0
°-°4
0.004
(a) For a 1000-MW plant burning 406 ton/hr
(13,540 Ib/min) coal.
-------
49
Thus, it appears reasonable to estimate that the effect of
trace species in the reducing gas on the FGD scrubbing system itself will
be inconsequential. However, special consideration must be given to FGD
processes which require a reductant for some purpose in addition to that
needed for reducing the SCL. This is the case for the Shell FGD process
and the Catalytic-IFP NH» scrubbing process. The Shell process requires
a reductant to regenerate the CuSO, and thereby liberate the SO-. Shell
recommends hydrogen for this purpose. The Catalytic-IFP process requires
a reductant for its sulfate reduction reactor. For such purposes it would
be convenient to obtain both reductants from the coal. If this is to be
done, consideration must be given to the effects of trace contaminants
from the coal upon the portion of the FGD process where the reductant is
introduced. An important consideration is the extent to which trace
constituents are already present at the location in question by transfer
from the flue gas. One would not: expect problems in the Shell process
because the copper-containing sorbent is periodically contacted with the
flue gas. In the Catalytic-IFP process, the extent to which impurities
in the flue gas reach the sulfate reduction reactor is not well defined.
9.3 Effects on Sulfur Production Processes
As shown in Figure 9-1, there are three processes in which the
sulfur dioxide stream concentrated by the Type A FGD processes can be
converted to elemental sulfur. Two of these, commercially proven and in
use for at least a decade, are
(1) The vapor-phase Glaus sulfur plant
(2) The liquid-phase Glaus reactor.
The basis of both these processes is the reaction:
S02 + 2H2S + 3/2 S2 + 2H20,
which indicates gaseous H^S to be a requirement. The third process does
not need gaseous HLS and can use the coal gas directly according to the
following vapor-phase reaction:
1.5 S02(g) + CO + 2H2 -» 0.75 S2(g) + C02(g) + 21^0.
-------
50
This reaction is representative of the ASARCO-Phelps Dodge process. A
pilot plant with a sulfur production capacity of 20 tons/day has been in
operation at ASARCO's El Paso, Texas, smelter since August, 1971. '
However, the process employs reformed natural gas (containing CO and H«)
rather than coal gas. The reformed natural gas is generated by the Phelps
(28)
Dodge Reforming process. In these processes, the common unknown is
the technical feasibility of using coal gas in place of reformed natural
gas. Conversations with ASARCO and others have indicated that experimental
studies in this area are under way; no data are available yet. Also, there
is a reluctance to sharing of available information since proprietary
interests are involved. In this section, a critical analysis of the effect
of coal gas trace constituents on the three processes listed above will be
attempted.
Vapor-Phase Glaus Reaction
In this process a stoichiometric mixture of SO. and H~S is
reacted over a bauxite catalyst at 230-400 C (450-750 F) and atmospheric
pressure as follows:
S02 + 2H2S * 3/2 S£ + 2H20.
The reactor inlet temperature is restricted to about 230 C (450 F) because
the reaction is exothermic and the equilibria are less favorable at higher
temperatures. The catalyst performance is unaffected by the presence of
C02« However, any ammonia and hydrocarbons must be removed from the feed,
since these species are detrimental to the catalyst. The sulfur vapor is
condensed at about 150 C (300 F).
According to a RJ5 production process vendor coal gas can be
converted to H_S to provide the following product composition as predicted
(29)
by thermodynamics:
Coal Gasifier Product,
mole percent dry
CO 50
H2 34
C0 14
-------
51
H2S Generator Product,
mole percent dry
H2S 55.7
S02 0.3
C02 40.4
COS 1.8
CS2 0.3
CO 0.4
N2 0.7
H2 0.4
Thus, the product gas exiting at 430-540 C (800-1000 F) can be fed directly
to the Glaus converter in admixture with the S02 from the FGD processes.
The trace species in the coal gasifier product entering the H2S generator
can be assumed to be too small to cause problems of catalyst poisoning;
the gas cleanup (scrubbing and high efficiency particulate matter removal)
will have removed over 98 percent of gaseous and condensed trace species
as discussed in Section 8. However, the CO and traces of methane can
result in the following undesirable reactions in the Glaus reactor:
CH4 + S02 -* COS + H20 + H2
CO + S ->• COS
C0 + HS f COS + H0
CO + S2 ->• CS2 + 1/2 02.
The CS2 and COS thus produced can be burned in the tail gas incinerator
to S02 and recycled to the FGD scrubber. Any trace species in the inciner
ator gas with tendency to augment oxidation of the scrubbing liquor in the
FGD absorber will be minor. The oxidation products are handled as a purge
stream which already exists in then parent FGD process. These discussions
indicate that a coal gasifier product can be employed to replace other
sources of reducing gas for a Glaus sulfur plant.
-------
52
Liquid-Phase Glaus Processes
In the Citrate, Phosphate, and Catalytic-IFP processes, the
S02~laden absorber liquor is reacted with an H2S-rich gas stream in a
two- or three-stage liquid reactor system.* Available data indicate
that catalysts are not employed in these processes. Therefore, potential
catalyst poisoning problems due to trace species in coal gas do not arise.
However, catalysts are utilized in the H2S generation and SO- reduction
steps.
For these processes, there is no experimental evidence regarding
the effects of using a coal gasifier-based H^S stream. The following
information provided by Peabody Engineered Systems (CITPJSX process vendor)
and Catalytic, Inc., (Catalytic-IFP process vendor) may be helpful.
• Water washing of the coal gas should eliminate nearly
all trace species having a potential to cause adverse
effects on sulfur production steps.
e The C02, CS2, and NH3 in the H S gas do not adversely
affect the catalyst or the liquid Glaus process
• A high COS concentration in the H2S feed gas will
result in an undisclosed adverse reaction in the liquid
Glaus step. The degree of adversity is not quantified.
• Apparently, for the citrate process removal of CO-
from the shift converter gas with at least a single
stage CO. removal step (e.g., the hot carbonate process)
is desirable. This will provide to the H2S generator
a feed rich in H2 and low in CO and COS. The net
result is an H2S stream with >80 mole percent H S
content, which insures high sulfur conversion efficiency
and high-quality sulfur product. Reportedly, the
Catalytic-IFP process seems to have the advantage of
utilizing the coal gasifier product directly without
a shift converter.
• According to Catalytic, Inc., the impurities contained
in the coal gas stream may build up cm the H_S
generator catalyst but may not poison it. Tnus,
periodic washing of the catalyst bed is provided for
by installing a spare catalyst chamber in the Catalytic-
IFP process design.
The Catalytic-IFP liquid Glaus unit is a single-stage system which
accomplishes essentially 100 percent conversion to bright commercial
grade elemental sulfur in a single pass.
-------
53
Discussions with the various vendors suggest that a system for
coal gasification, cleanup, and H«S generation shown in Figure 9-2 would
be necessary. Since the gas necessarily is cooled both before and after
the shift reactor, it would probably be advisable to use a low temperature
shift catalyst to minimize the extent of gas reheating. Shift catalysts
are now available which can be used at temperatures of 200-250 C (390-480 F)
and can tolerate reasonable concentrations of H^S without poisoning. The
choice is not obvious, however, because the high temperature shift catalysts
have the advantage that they effect nearly complete hydrolysis of COS to
form H.S, thereby eliminating an undesirable specie in the product gas.
Based on the available information and best judgment, the effect
of trace species on the liquid Glaus reactions is expected to be a minor
problem in comparison with the other problems of successfully operating a
regenerative FGD system. The magnitude of the problems of utilizing coal-
based reducing gas can only be judged accurately by conducting experimental
tests with various coals and FGD processes.
ASARCO Direct Sulfur Process
The ASARCO process for reducing SO to sulfur with H./CO was
discussed in Section 5. A flow sheet of this process as adapted for use
with coal gas is shown in Figure 9-3. Reportedly, the technical feasi-
bility of the process has been, proved in a 20 ton/day pilot plant
operating since 1971 on smelter gas and using reformed natural gas as
the reductant. The requirement of the process is that the feed have
at least 12 mole percent SO.. This is easily met by all the Type A FGD
processes, since these processes produce streams containing 20 to 85
percent SO^.
As shown in Figure 9-3,.the process employs two catalytic
reactors, the second of which is a Claus reactor. The type of catalyst
in primary reactor is proprietary information. Probably, the secondary
reactor contains bauxite or alumina as the Claus catalyst. We have
already noted that trace species in coal gas, if properly scrubbed, will
not harm the Claus catalyst. However, since no information on the primary
-------
Water
Coal
Steam
I
Shift
Converter
Sulfur from
FGD process
H2S
Generator
H2S (80%)
01
To waste
FIGURE 9-2. SCHEMATIC FOR COAL GASIFICATION, CLEANUP, AND H^S GENERATION
-------
Tail Gas Recycle (S02>
so2
FGD
process
module
Coal
gasi tier and
gas cleanup
1 •
i
sT
«h
8
i '
ASARCO ••
primary
reactor
M350 C)l
w
i
Shift
converter
(possible)
Steam
o
cs
W
M
S^\. c^t
( ) t V......I . . re
cs
/*>>. °
*( ^r W
\y y
X_X o
\ 0
I o
ater — ' U r
r
\N
Steam
n
Sulfur
Condenser I
(205 C)
i
f — ^^\
Secondary
reactor
(240 C)
^—- ^
' — 1
1 ' Steam
1 ... .^ Air
1 *^ "••
5\J , }
L-X ' 1
1-i.elicaLet bultur InrmPrntnrU
incineraior i
Condenser IE • •
i •
t I (140 C) Liquid
1 4 Sulfur
Liquid
Sulfur
Ul
FIGURE 9-3.
SCHEMATIC OF AMERICAN SMELTING & REFINING COMPANY (ASARCO)
ELEMENTAL SULFUR PRODUCTION PLANT ADAPTED TO USE OF COAL GAS
-------
56
reactor catalyst is available, the effect of trace species in coal gas
on it cannot be evaluated. ASARCO is studying the feasibility of substi-
tuting coal gas for reformed natural gas. If the initial work in this
area is promising, more specific studies of the effects of trace species
on the process could be conducted.
-------
57
10. GAS CLEANUP SYSTEMS
Previous sections of this report (and Appendix C) have presented
information on gasification systems and their products. It is clear from
the discussion of Section 9 that the raw gas from a gasification system
will have to be washed with water before it can be used as a reducing gas
in the manner intended in this report. This water washing will remove
principally particulate matter, ammonia, and hydrogen cyanide. The purpose
of this section is to discuss possible water washing systems and to examine
the effectiveness of such systems! in removing ammonia and hydrogen cyanide
from gases. Additional information on these subjects is presented in
Appendix E.
10.1 Scrubber Systems
A number of water scrubbing systems have been proposed for use
with coal or oil gasification processes. These are discussed in Appendix
E. A frequent omission in these conceptual systems is that of providing
no means for removing the absorbed gaseous species, primarily ammonia,
from the scrubbing liquor. .Without such removal, the ammonia concentration
would be expected to build up to some steady-state concentration in the
scrubbing liquor loop and, from that time on, there would be little
change in ammonia concentration in the gas phase through the scrubber.
Of course, the particulate matter must also be removed from the scrubbing
liquor.
A conceptual design for a spray scrubber which would remove gaseous
impurities as well as particulate matter and tars is shown in Figure 10-1.
The water loop would be provided with a separator, settler, and/or filter
to remove tars, tar oils, naphtha, and particulate matter and then a
rectification-stripping column to drive off the absorbed materials and
soluble materials more volatile than water. Subsequent to the stripping
column, a water cooler would reduce the water temperature to a level
suitable for use in the spray scrubber. A water purge and water makeup
also are indicated.
-------
58
A key feature in this conceptual design is the configuration of
the rectification-stripping column. By using closed steam heating of the
reboiler, additional water is not added to the water loop unnecessarily.
Also, provision for a rectification section permits the recovery of materials,
such as ammonia, which may have some value and prevents environmental con-
tamination through indiscriminate discharge of these materials to the atmos-
phere.
10.2 Effectiveness of Scrubber Systems
The conceptual gas cleanup system shown in Figure 10-1 includes
a spray scrubber, because this type of equipment can handle a solids-laden
gas while still providing reasonably good gas-liquid contacting. The
mass transfer characteristics of a spray tower are not so favorable as
those of a packed tower or tray tower, but the latter type of equipment
would readily be plugged by the particulate matter contained in the raw
gasifier gas. The maximum number of mass transfer units achievable in a
spray tower is about three. Therefore, some calculations were made to
determine the removals of gaseous species that could be obtained with three
mass transfer units. The results indicate that up to 93 percent of the
ammonia (NH,) and 85 percent of the hydrogen cyanide (HCN) can be removed
from a typical raw coal gas by water scrubbing at 30 C. A substantial
fraction of the hydrogen sulfide (H.S) in the gas will also be removed.
A few exploratory calculations were performed on the stripping
of absorbed gases back out of the water. Ammonia can easily be stripped
from the water; less than two equivalent plates would be required to
reduce the ammonia concentration in the water to 1 ppm.
-------
59
I
Clean
Gas Out
Spray
Scrubber
Raw
Gas
In
A AA A
Water
Out
Water
Makeup
Water
In
Cooler
Separator
and/or
Filter
Noncondensibles
Rectification and
Stripping Column
Reboiler
Purge
Solids
or Slurry
FIGURE 10-1. CONCEPTUAL DESIGN OF A SPRAY
SCRUBBER, GAS CLEANUP SYSTEM
-------
60
11. COSTS OF GASIFIER-BASED REDUCTANT
SYSTEMS FOR FGD PROCESSES
In this section the costs of gasifier-based reductant systems
for FGD processes are considered. It should be recognized at the outset
that these costs are necessarily rather approximate because they apply to
systems that are still only conjectural. Detailed system designs have
not been made and could not be accurately made at this time because some
questions on the technology involved have not yet been answered.
In order to make the cost estimates as realistic as possible,
segments of the systems such as oxygen plants and Glaus plants were
costed as complete units based on current cost data from vendors. Addi-
tional detail on the sources of cost data and how these data were used
is presented in Appendix G. Additional details on the engineering
calculations involved in the system designs (i.e., material and energy
balances) are included in Appendix F.
In the following, the application of reductant systems to the
two types of FGD processes will be considered separately.
11.1 Type A FGD Processes
The Type A FGD processes can use a H./CO mixture for reduction
of the concentrated S02 stream produced. The most desirable way to
accomplish this would be with a system similar to the ASARCO process.
This involves a two-stage reaction system with some sulfur produced in
both reactors. The first reduction reactor contains the ASARCO catalyst
and the second reactor a Glaus catalyst. It is possible that a third
reactor containing Glaus catalyst might be necessary if 95 percent sulfur
recovery is to be achieved. In this case, the reduction system would
actually contain a. two-stage Glaus plant and would, in effect, be very
similar to a three-stage Glaus plant overall. The only difference is that
a different catalyst would be used in the first stage so that a H2/CO
mixture could be fed instead of H_S.
On the other hand, a reductant system for Type A FGD processes
is not entirely dependent upon the ASARCO technology because a standard
-------
61
cobalt-molybdenum hydrogenation catalyst could be used in the first stage.
The only disadvantage of this is that the first reactor would then
produce only H«S and no sulfur would be removed at this point. As far
as the system cost is concerned, it should make little difference which
type of catalyst is used in the first reactor. In either case the
reduction system should be very similar in cost to a three-stage Glaus
plant. Any differences in reactor size (space velocity) or catalyst cost
would be unknown at this time anyway.
The cases considered in this section include both oxygen- and
air-blown gasifiers and both classes of Type A FGD processes, i.e., those
which produce an 85 percent S02 stream and those which produce a 25 percent
SO- stream. Thus, four different levels of gas dilution are included.
Material Balances
The reducing gas compositions used in the material balance calculations
are shown in Table 11-1. The oxygen-blown case is based on a Koppers-Totzek
gasifier and the air-blown case on a Wellman-Galusha, although other gasi-
fiers would give similar compositions.
The material balance flow sheets for the four cases are presented
in Figures 11-1 through 11-4. The first two figures are for oxygen-blown
gasifiers and the second two are for air-blown gasifiers. In these
figures the S0« reduction system is shown as consisting of two stages
with the second being a Claus reactor. As mentioned previously, there
will probably be in fact two separate Claus reactors, and the costs have
been developed to reflect this. It is assumed that no shifting of CO to H£
is required before the first reduction step.
Investments
The estimated investments for reduction systems for 1000-MW and
500--MW power plants are presented in Tables 11-2 and 11-3, respectively. In
addition to the four .cases based on coal gasification, two cases are
shown for the Allied SO,, reduction process, which uses natural gas as
the reductant. The investments for the coal gasifier-based systems are
consistently greater than those for the Allied process.
-------
TABLE 11-1. REDUCING GAS COMPOSITIONS USED IN MATERIAL BALANCE CALCULATIONS
Specie
Concentration (mole % except as otherwise noted)
Oxygen-Blown Gasifier (Koppers-Totzek)
Leaving Heat.
Recovery System
Air Blown Gasifier (Wellman Galusha)
Leaving Heat,. »
Recovery System
Tar, grains/scf
Particulate matter, grains/
scf
100.00
11.57
11.57
100.00
100.00
3.82
0.1-0.5
0.1-0.5
H2
CO
co2
CH.
4
COS
NH3
HCN
N2
Ar
H.O
24.43
39.08
4.97
0.08
0.92
0.08
0.17
0.03
0.71
0.32
29.19
30.65
49.03
6.24
0.10
1.06
0.10
0.01
—
0.89
0.40
11.52
13.32
25.39
3.02
2.40
0.54
0.02
0.50
0.01
44.65
0.15
10.00
13.18
25.11
2.99
2.37
0.49
0.02
0.01
—
44.16
0.15
11.52
100.00
(a)
(b)
(c)
From data on Illinois coal (Table C-l) after adjustment to 3.5 percent S coal and addition of HO,
NH_, Ar, CH,, and HCN from Reference 19.
Based on reduction of NH» to 100 ppm, high removal of HCN, 8 percent removal of H»S, and saturation
with water at 120 F and I atm.
From data on bituminous coal (Table C-2) after addition of HO (assumed cone), H S, COS, Ar, and HCN
(adjusted from oxygen-blown data), and NH-.
3
Note: grams/std m = 2.288 (grains/scf).
-------
Air
i
Dry Coal,
-287.8 T/D
Steam
(2)
Oxygen
plant
©
Oxy
Nitrogen
(3)
SO.-Rlch Stream
gen, 220.5 T/D Watcr from FGD Air
" I V if
Gasifier
i
r
Gas Cleaned Tall Gas Recycle
Raw Gas Cleanup Gas 1st Stage 2nd Stage Toil Gas to FGD
($\ and (7) " Reducer *~ Reducer /gV^ Incinerator /7j\
Reheat (Clous)
1
V '
Bottom Ash
Fly ash, Tar,
Water
Sulfur, 299.8 T/D
Ib mole/hr for 1000-MW Power Plant
[2J,
©
2643.6
702.7
360.3
2632.1 10.6 0.9
138.5 518.4 45.8
706.2
1129.6
143.8
2.4
26.6
20.5
843.7
674.4
1078.8
3346.3 360.3
2770.6 529.0
Temperature, C
Temperature, F
27
80
121
250
46.7
27
80
798.3
140.9
939.2
27
80
97.8
156.5
1117.0
2.3
3.4
40.0
21.4
1034.9
4.1
.11.8
2489.2
288
550
550.7
146,4
697.1
27
80
3057.4
LO
FIGURE 11-1. MATERIAL BALANCE FLOW SHEET FOR TYPE A FGD, 85% SO. STREAM, OXYGEN-BLOWN GASIFIER
-------
Mr
S
Dry Coal,
-287.8 T/D ' ^
Steam
Oxygen
plant
©,
Oxy
Nitrogen
i
SO.-Rlch Stream
gen, 220.5 T/D Water from FGD Mr
•® . (_ ® • ®
Gasifier
i
Gas Cleaned Tall Gas Recycle
Raw Gas Cleanup Gas 1st Stage 2nd Stage Tail Gas to FGD
(e) and (7) " Reducer " Reducer /§\' Incinerator (fj\
Reheat (Clous)
i
* ! " '
Bottom Ash
Fly ash. Tar,
Water
Sulfur, 299.8 T/D
Ib 'mole/hr for 1000-MW Power Plant
Total
Temperature, C
Temperature, F
360.3
3346.3 360.3
2632.1 10.6 0.9
138.5 518.4 45.8
2770.6 529.0
27
80
121
250
46.7
27
80
706.2
1129.6
143.8
2.4
26.6
20.5
843.7
17.3
2890.1
177
350
4743.2
288
550
550.7
550.7
146.4
697.1 5311.4
27
80
FIGURE 11-2.
MATERIAL BALANCE FLOW SHEET FOR TYPE A FGD, 25% SO- STREAM, OXYGEN-BLOWN GASIFIER
-------
Water Air
SOj-Rlch Stream
from FGD
Air
Dry Coal- , _
296.3 T/D
Air (?) ' Raw Gaa Cleanup Gas
Gosifier fi\ and ($\
Steam © _ ^ Reheat ^
1
Bottom Ash Fly ash, Tar,
Water
11
©
1
st Stage 2nd
®
i
Tall Gas Recycle
Stage Tail Gas to FGD
Reducer * Reducer ^\" Incinerator (§\
(Glaus) W W
1
i
Sulfur, :
b mole/hr for 1000-MW P
s
£99.9 T/D
wer Plant
Specie' 0 © © © © © © © ©
H, 661.6 604.4 . 87.6 0.
CO 1260.5 1151.5 167.0 • 0.
CO 150.2 269.6 1254.1 1529.5
CH2 118.7 108.4 108.4 0.
H0S 26.8 22.6 3.3 0.
SO 2.0 798.3 40.0 43.3
N 2 2205.2 2205.2 471.8 2677.0 2677.0 1401.7 4078.7
H20 548.3 496.5 656.1 140.9 1333.1 1640.8
O2 586.2 125.4 11.4 11.4 372.6 35.0
Other 33.8 9.4 9.0 9.0
Total 2791.4 548.3 4953.3 597.2 5512.0 939.2 5690.9 1774.3 7336.3
Temperature, C 27 121 177 27 343 27 288 27
Temperature, F 80 250 350 80 650 80 550 80
ON
l/i
FIGURE 11-3. MATERIAL BALANCE FLOW SHEET FOR TYPE A FGD PROCESS, 85% S00 STREAM, AIR-BLOWN GASIFIER
-------
Water Air
SO.-Rich Stream
from FGD
Air
Dry Coal, ,— _•_ __.", ... ,-
296.3 T/D
Air (?) Raw Gas Cleanup Gaa 1
* Gasitier /Jx * and /g\ *~
Steam © ^ Reheat ^
' i
1 1
Bottom Ash Fly ash, Tar,
Water
»
11
©
<
st Stage 2nd Stage
Reducer * Reducer (?
(Clous) -^-
i
Sulfur. 299.9 T/D
> mole/hr for 1000-MH Power Plant
,®
Tail Gas Recycle
Toil Gas to FGD
\* Incinerator /g\ *"
Specie' ©'© © © © © © © ©
H, - 661.6 604.4 . 87.6 0.
CO 1260.5 1151.5 167.0 • 0.
CO. 150.2 269.6 542.8 1796.9 2072.3
CH7 118.7 108.4 108.4 0.
H,5 26.8 22.6 3.3 0.
SO 2.0 798.3 40.0 43.3
N,2 2205.2 2205.2 471.8 2677.0 1373.1 4050.1 1401.7 5451.8
H*0 548.3 496.5 656.1 479.0 1671.2 1978.9
0* 586.2 . 125.4 11.4 11.4 372.6 35.0
Other 33.8 9.4 ?-0 9.0
Total 2791.4 548.3 4953.3 597.2 5512.0 3193.2 7944.9 1774.3 9590.3
Temperature, C 27 121 177 27 343 27 288 27
Temperature, F 80 250 350 80 650 80 550 80
CTv
0\
FIGURE 11-4.
MATERIAL BALANCE FLOW SHEET FOR TYPE A FGD PROCESS, 25% SO- STREAM, AIR-BLOWN GASIFIER
-------
TABLE 11-2. INVESTMENTS FOR REDUCTION OPTIONS, TYPE A FGD PROCESS, 1000-MW POWER PLANT
Allied Process (Natural Gas)
Inlet S0_ Concentration (vol %) 85 25
Gasification Agent
(a.} 3
1975 Investment^ ' (10 $)
Gasification system
Oxygen plant
Glaus plant
Ductxvork
Blowers
Total 8,258 11,837
Total Investment in $/kW 8.3 11.8
Mole % H2 + CO + H2S + S02 at
first reduction stage
85
°2
8,110
5,180
3,967
29
52
17,338
17.3
79.3
Coal Gasification
85
Air
8,372
5,983
42
113
14,510
14.5
40.0
25
°2
8,110
5,180
5,444
36
66
18,836
18.8
46.8
25
Air
8,372
7,168
47
128
15,715
15.7
29.6
(a) Chemical Engineering Plant Cost Index = 190, Marshall & Stevens Equipment Cost Index = 460. Invest
ments are based on flow rates 10 percent greater than those shown on material balance flow sheets.
-------
TABLE 11-3. INVESTMENTS FOR REDUCTION OPTIONS, TYPE A FGD PROCESS, 500-MW POWER PLANT
Allied Process (Natural Gas)
Inlet S0_ Concentration (vol %) 85 25
Gasification Agent
1975 Investment (a) (10 3$)
Gasification system
Oxygen plant
Claus plant
Ductwork
Blowers
Total 6,062 8,587
Total Investment in $/kW 12.1 17.2
Mole % H2 + CO + H2S + S02 at
first reduction stage
85
°2
4,992
3,188
2,617
21
32
10,850
21.7
79.3
Coal Gasification
85
Air
5,153
3,947
30
41
9,171
18.3
40.0
25
°2
4,992
3,188
3,592
26
70
11,868
23.7
46.8
25
Air
5,153
4,729
34
79
9,995
20.0
29.6
(a) Chemical Engineering Plant Cost Index = 190, Marshall & Stevens Equipment Cost Index = 460.
Investments are based on flow rates 10 percent greater than those shown on material balance
flow sheets.
oo
-------
69
Systems based on oxygen-blown gasifiers require greater invest-
ments than those based on air-blown gasifiers. The cost of an air-blown
gasification system is only slightly greater than that of an oxygen-
blown system (excluding the oxygen plant), and this difference is due to
a small increase in capacity because of a greater reheat requirement.*
The greater gas volumes for an air-blown system increase the cost of the
reduction system and other equipment, but this increase is not as great
as the cost of the oxygen plant for the oxygen-blown system.
Operating Costs
The estimated operating costs for reduction systems for 1000-MW and
500-MW power plants are presented in Tables 11-4 and 11-5, respectively. The
unit costs used in calculating these operating costs are shown in Table
11-5. For the coal-based cases, two coal costs are considered, $10 and $30
per ton, since there is currently a wide range of coal prices. Even with
coal at $10/ton the operating cost for the coal-based systems are consis-
tently greater than those for the Allied process. Again, the air-blown
systems are less costly than the oxygen-blown systems.
Comments
Although the costs for systems based on air-blown gasifiers
are less than those for oxygen-blown gasifiers, the choice between these
two is not obvious. In some cases, particularly for the more dilute S0_
streams, the additional dilution caused by the use of air-blown gasifiers
may decrease the reaction rates in the reduction system to a point where
satisfactory sulfur recovery cannot be obtained. As the composite H«S +
SO- stream entering a Claus plant becomes more dilute, the efficiency of
sulfur recovery decreases. This effect for a two-stage Claus plant is illus-
trated in Table 11-7. For a three-stage Claus plant the recoveries will be
higher but the trend will be the same. Thus, for reduction of relatively
dilute S0? streams, the gasification may have to be done with oxygen.
* Reheating the reducing gas to 340 C (650 F) after the water wash increases
the coal fed to the gasifier by about 8.7 percent for the air-blown
case versus 4.5 percent for the oxygen-blown case. (See Appendix F.)
-------
TABLE 11-4. OPERATING COSTS FOR REDUCTION OPTIONS, TYPE A FGD PROCESS, 1000-MW POWER PLANT
75 Percent Load Factor
Inlet SO- Concentration (vol Z)
Gasification Agent
Operating Cost, 10 $/yr
Natural gas
Coal, $10/T ($30/T)
Electricity
Gasification system
02 plant
Reduction system
Blowers, etc.
Steam
Gasification system
Reduction system
Water
Gasification system
02 plant
Reduction system
Catalysts & chemicals
Gasification system
0. plant
Reduction system
Total Feed and Utilities
Operating labor
Investment related (24Z/yr)
Total Gross
Sulfur credit ($10/LT)
Total Net
Net Operating Cost in mills/kWh
Net Operating Cost In $/LT Sulfur
Allied Process
85
1,016
44
14
-24
29
4
1,083
237
1,982
3,302
-713
2,589
0.39
36
(Natural Gas)
25
1,016
44
19
-18
29
4
1,094
237
1^841
4,172
-713
3,459
0.53
49
85
°2
788 (2,364)
19
249
44
18
-129
-24
37
6
29
6
59
4
1,106 (2,682)
377
4,161
5,644 (7,220)
-733
4,911 (6,487)
0.75 (0.99)
67 (89)
Coal Gasification
85 25
Air 02
811
19
44
27
-24
39
6
29
6
4
961
259
3.482
4,702
-733
3,969
0.60
54
(2,433) 788 (2,364)
19
249
44
20
-129
-24
37
6
29
6
59
4
(2,583) 1,108 (2,684)
377
4,521
(6,324) 6,006 (7,582)
-733
(5,591) 5,273 (6,849)
(0.85) 0.80 (1.04)
(76) 72 (93)
25
Air
811 (2,433)
19
44
20
-24
39
6
29
6
4
964 (2,586)
259
3.772
4,995 (6,617)
-733
4,262 (5,884)
0.65 (0.90)
58 (80)
Note: $/metric ton » 1.103 ($/T) - 0.9844 (S/LT).
-------
TABLE 11-5. OPERATING COSTS FOR REDUCTION OPTIONS, TYPE A FGD PROCESS, 500-MW POWER PLANT
75 Percent Load Factor
Allied Process (Natural Gas)
Inlet SO. Concentration (vol X)
Gasification Agent
3
Operating Cost, 10 $/yr
Total feed and utilities^
Operating labor
Investment related (24%/yr)
Total Gross
Sulfur credit ($10/LT)
Total Net
Net Operating Cost in mills/kWh
Net Operating Cost in $/LT Sulfur
85
542
193
i»455
2,190
-357
1,833
0.56
51
25
547
193
IjOei
2,801
-357
' 2,444
0.74
69
85
°2
553 (1,341)
306
1,604.
3,463 (4,251)
-366
3,097 (3,885)
0.94 (1.18)
85 (106)
Coal Gasification
85
Air
481 (1,292)
210
1*201
2,892 (3,703)
-366
2,526 (3,337)
0.77 (1.02)
69 (91)
25
°2
554 (1,342)
306
2 ,848
3,708 (4,496)
-366
3,342 (4,130)
1.02 (1.26)
91 (113)
25
Air
482 (1,293)
210
2, 399.
3,091 (3,902)
-366
2,725 (3,536)
0.83 (1.08)
74 (96)
(a) One-half of value for 1000-MW plant.
(b) Scaled down from 1000-MH case using 0.3 power of capacity.
Note: $/metric ton - 1.103 ($/T) - 0.9844 ($/LT).
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72
TABLE 11-6. UNIT COSTS USED IN OPERATING COST CALCULATIONS
Item Cost
Natural gas $0.75/10 scf
Coal $10/ton ($30/ton)
Electricity $0.012/kWh
Steam $0.50/106 Btu(a)
3
Process water $0.25/10 gal
3
Boiler feed water $0.50/10 gal
Operating labor $9/hour total cost
(a) Heat relative to saturated liquid water
at 49 C (120 F).
Conversion factors: $/std m - 35.31 ($/scf)
$/metric ton = 1.103 ($/ton)
$/10° kcal = 3.953 ($/106 Btu)
$/10J liter = 0.2642 ($/103 gal)
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73
TABLE 11-7. EFFECTS OF SO DILUTION ON SULFUR RECOVERY
IN TWO-STAGE CLAUS PLANT
Inlet H S + SO Concentration, Sulfur Recovery,
volume percent dry basis percent
20 92.7
40 93.5
60 94.4
80 95.0
Source: Kim, B. C., et al (Reference 30).
-------
74
The operating cost calculations for the oxygen-blown systems
include a rather substantial credit for steam produced in the gasification
system. Since the oxygen-blown cases are based on the Koppers-Totzek
gasifier, data on the steam production by K-T gasifiers were used. Using
the data for Illinois coal (Table C-l) plus two material balances obtained
from Koppers, ' the net steam productions and heat contents thereof
were calculated. The average net steam production for these cases was
1,436 kg/metric ton dry coal (2,990 Ib/ton) or 0.912 x 106 kcal/metric
ton dry coal (3.27 x 10 Btu/ton) of energy relative to liquid water at
49 C (120 F). This is primarily high temperature (e.g., 260 C or 500 F)
steam which can be used in the FGD process or the boiler itself. No steam
credit was taken for the air-blown cases, since the available data on
Wellman-Galusha systems showed neither a steam credit nor a cost. Air-
blown gasifiers generally operate at lower temperatures than oxygen-blown
gasifiers and hence generate less steam and lower temperature steam. It
is probably reasonable to assume that the steam generated by such a
gasifier will just meet its own steam requirement.
Although comparisons of the coal-based reduction systems with
the Allied process are instructive, it should be kept in mind that these
two options accomplish a rather different objective because of the
difference in the starting materials. The use of a "dirty" (sulfur-
containing) reductant would be expected to cost more than the use of a
clean reductant (natural gas), just as the use of dirty fuels in an
environmentally sound manner generally costs more than the use of clean
fuels. The present and projected shortage of natural gas will require
factors other than cost to be considered.
The general level of the reduction costs are worth noting.
For a 1000-MW power plant fired with 3.5 percent sulfur coal, any type
of SO- reduction system will add 8-19 $/kW to the investment for the
power plant and 0.4-1.0 mills/kWh to the operating cost. This represents
a significant fraction, roughly 15 percent, of the estimated coat of the
FGD process itself. The costs reflect the fact that the reduction step
further decreases the overall energy conversion efficiency of the power
plant. The increase in the complexity of the power plant is apparent
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75
from the flow sheets presented in this section. This is not meant to imply
that the reduction step is undesirable, since it is a necessary part of
a sulfur-producing FGD system. It is important to recognize all the
costs and other disadvantages which are concomitant with the advantages of
an FGD system which is environmentally acceptable in all respects.
>
11.2 Type B FGD Processes
The Type B FGD processes require H_S as a reductant because
these processes include liquid-phase Glaus reactors. The reduction of
some of the elemental sulfur to generate the H0S can be done with commercially
available processes such as the Thio-Pet and Girdler processes. It is
possible that a coal-based H«/CO mixture could be used directly in a process
of this type in place of the EL-rich gases which have been used to date.
Thio-Pet personnel are optimistic about this on the basis of some thermo-
dynamic calculations they have made (these were discussed in Section 9).
However, they are quick to point out that this has not been tested
experimentally.
Lacking a demonstration of this more direct route, it seemed
advisable to estimate the costs associated with the conventional technology,
i.e., using a H.-rich stream for the H_S generation. Hydrogen sulfide
generation plants operating today use a reductant gas containing at least
80 percent H,.. Producing such a gas from coal gas requires an oxygen-
blown gasifier plus two stages of CO shifting and acid gas removal. A
flow sheet for a coal-based H,,S generation system is shown in Figure 11-5,
and the associated stream compositions and flow rates are given in Table
11-8.
Investments
The investments for coal-based H-S generation systems for use
with Type B FGD processes are shown in Table 11-9. The investment for
the H9S generation facility itself is a quotation supplied by C & I/
(33)
Girdler, Inc., for this project. The investment for the reductant
gas upgrading equipment (shift converters, acid gas removal facilities,
-------
C.W.
Coal Gas From,
Cleanup Syste
Water
21 gpm
Steam
I'KjQ. drum_
To FGD
j _* Process
C.W.
JK.O. drum I
18,000 Ib/hr
t
Water
18 gpm
l_
0>
JO
L.
o
in
Q
<
r
•x
On
1X2
o
•*-
o
Q)
C
Q>
O>
0>
CE
Molten Sulfur
Ib/hr
Water
12 gpnCT
• Steam
115,200
Ib/hr
6,500 Ib/hr
43,700 Ib/hr
FIGURE 11-5.
MATERIAL BALAMCE FLOW SHEET FOR H2S PRODUCTION BASED ON COAL GASIFICATION
-------
TABLE 11-8. COMPOSITIONS AND FLOW RATES FOR US PRODUCTION BASED ON COAL GASIFICATION
Stream Nuober
Temperature, C
Temperature, F
Pressure, psia
Composition, mole percent
H2
CO
co2
CH4
H2S
cos
N2
Ar
cs2
s
H20
Total
Total Flow Rate
103 Ib-molea/hr
106 ft3/hr
1
52
125
15
30.65
49.06
6.24
0.10
1.04
0.10
0.89
0.40
—
—
11.52
100.00
2.319
0.969
2
316
600
120
30.65
49.06
6.24
0.10
1.04
0.10
0.89
0.40
-- -
—
11.52
100.00
2.319
0.219
3
482
900
120
39.11
8.15
24.64
0.06
0.62
0.06
0.53
0.24
~
—
26.61
100.02
3.912
0.475
4
107
225
120 .
44.89
9.35
28.28
0.07
0.71
0.07
0.61
0.27
—
—
15.76
100.01
3.408
0.208
5
110
230
120
72.76
15.15
4.59
0.11
—
—
0.98
0.44
—
—
5.97
100.00
2.103
0.130
6
38
100
120
76.76
15.99
4.84
0.12
—
—
1.04
0.47
--
—
0.79
100.01
1.993
0.100
7
204
400
120
76.76
15.99
4.84
0.12
—
—
1.04
0.47
—
—
0.79
100.01
1.993
0.153
8
260
500
120
67.59
1.43
14.07
0.09
—
—
0.77
0.35
—
—
15.71
100.01
2.678
0.230
9
107
225
120
67.59
1.43
14.07
0.09
—
—
0.77
0.35
—
—
15.71
100.01
2.678
0.164
10
110
230
120
90.50
1.91
1.89
0.12
—
—
1.03
0.46
—
/
4.08
99.99
2.000
0.123
11
454
850
120
1.56
0.33
1.63
0.04
76.71
1.32
0.89
0.40
0.06
13.51
3.53
99.98
2.313
0.270
12
54
130
15
1.81
0.38
1.81
0.05
88.70
1.53
1.03
0.46
0.07
—
4.08
100.00
2.000
0.843
(a) Refer to Figure 13-5.
9 33
Conversion factors: kg/m abs. - 703.1 (psia), g-moles/hr - 453.6 (Ib-moles/hr), m /hr - 0.02832 (ft /hr).
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78
TABLE 11-9. INVESTMENT FOR H S GENERATION SYSTEM FOR TYPE B FGD PROCESS
Power Plant Capacity
1975 Investment (103$)
Gasification system
Oxygen plant
Shift converter 1
Waste heat recovery system
Acid gas removal 1
Shift converter 2
Acid gas removal 2
H2S generation facility '
Total
Total Investment in $/kW
1000 MW
8,141
4,840
295
1,885
5,047
466
1,773
9,700
32,147
32.1
500 MW
5,011
2,979
182
1,160
2,524
287
1,091
4,850
18,084
36.1
(a) Two trains for 1000 MW, one train for 500 MW.
(b) Four trains for 1000 MW, two trains for 500 MW.
-------
79
f f\ i \
etc.) are based on cost data obtained from the Bureau of Mines on
making hydrogen from coal gas. Adjustments were made for capacity and
operating pressure.
The investments shown in Table 11-9 are 1.5-2.2 times greater than
the investments for coal-based reductant systems for Type A FGD processes
(Tables 11-2 and 11-3). The investment would be reduced by about 7
percent if the second stage of shifting and acid gas removal could be
eliminated and by about 30 percent if both stages could be eliminated.
Operating Costs
The operating costs for the coal-based H S generation systems
are shown in Table 11-10. The requirements of the H.S generation facility
itself were supplied by Girdler. The steam requirement for the reductant
upgrading steps, which is quite large, was determined by material and
energy balances.
The total operating costs shown in Table 11-10 are 1.5-2.5 times
greater than the operating costs for coal-based reductant systems for Type
A FGD processes (Tables 11-4 and 11-5). The total cost of H2S made in this
manner is $55-$71/metric ton (2.5-3.2c/lb), depending on the coal price
and plant size. An "across the fence" purchase of H-S for use as a reduc-
tant (the concept discussed in Section 6) would be economical only if the
prices were lower than these cost levels as applied to the specific
situation. Only by-product H-S streams could meet this requirement; high
purity merchant H^S would be expensive.
Comment
It should be noted that a portion of the costs shown here
would be incurred even if the H-S were to be made from natural gas instead
of coal. In this case, the cost of the H«S generation facility itself
would be the same as shown here, the cost of the gas upgrading step would
be somewhat lower (less CO shifting and C02 removal), and the cost of a
steam reforming plant for the natural gas would be less than that of the
coal gasification system.
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80
TABLE 11-10. OPERATING COSTS FOR H2S GENERATION SYSTEM FOR TYPE B FGD PROCESS
Power Plant Capacity 1000 MW 500 MW
3
Operating Cost (10 $/yr)
Coal, $10/T ($30/T) 792 (2,377)
Electricity
Gasification system 19
02 plant 231
H.S generation facility 36
Other 22
Steam
Gasification system -130
Shift converters, regenerators 648
H2S generation facility 76
Water
Gasification system 37
0- plant 6
H S generation facility 14
Other 10
Catalysts & chemicals
Gasification system 6
02 plant 54
Other 10
Total Feed and Utilities 1,831 (3,416) 916 (1,708)
Operating labor 377 306
Investment related (25%/yr) 7.715 4.340
Total 9,923 (11,508) 5,562 (6,354)
Operating Cost in mills/kW 1.51 (1.75) 1.69 (1.93)
Operating Cost_in_$/metric
" tonH2S " " 55~.T (63.9)" " " "61.7 (70.6)
Note: Elemental sulfur for H2S generation facility is taken as haying
no cost.
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81
11.3 Other Considerations
There is one aspect of the integration of a coal gasification-
based reduction system with an FGD process to which an allusion was made
earlier in this section which warrants further discussion. This involves the
fact that the gasification system, in addition to supplying the reductant,
can supply at least part of the energy requirements of the FGD process.
For the wet FGD processes, the major use of energy is for reheating the
cleaned flue gas. The energy requirement for reheating the flue gas is
about 0.15 x 10 kcal/metric ton (0.52 x 10 Btu/ton) of coal fired to
the boiler (see Appendix F). As mentioned previously, a Koppers-Totzek
gasifier produces about 0.912 x 10 kcal/metric ton of coal gasified
in usable heat as high temperature steam. Since the quantity of coal
fed to the gasifier is about 3.6 percent of that fed to the boiler,
the heat available from the by-product steam is
(0.036)(0.912 x 10 ) = 0.033 x 10 kcal/metric ton of coal fired to boiler.
Thus, with a Koppers-Totzek gasifier about one-fourth of the reheat
requirement could be met with the by-product steam. For other gasifiers
operating at lower temperatures and hence producing less steam, a lower
fraction of the reheat requirement could be met in this way. Even
though the by-product steam can only partially satisfy the reheat require-
ment, the quantity of heat is large and the economic credit for it is
substantial.
Additional energy, beyond that available from by-product steam,
can. be supplied to the FGD process for some uses by increasing the size
of the gasification system and using the low- or intermediate-energy fuel
gas directly. The flue gas can be reheated by burning some fuel gas and
mixing the combustion products with the flue gas. Admittedly this adds
some S0_ to the cleaned flue gas, but this may be permissible if the S0_
reir.oval by the FGD process itself is high enough. The effect is similar
to the option of reheating cleaned flue gas by bypassing some of the flue
gas untreated.
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82
11.4 Use of Oil Gasification
In order to investigate the effects of using residual oil as the
gasification feedstock instead of coal, the Shell gasification process (SGP)
for oil was compared with the coal gasification systems. The SGP is a
partial oxidation process and includes facilities for removing the soot
from the gas and recycling it to the gasifier. Therefore, it is comparable
to a coal gasification system including the gas washing step.
This cost comparison is shown in Table 11-11. The coal gasification
costs are from Tables 11-2 and 11-4. For the Shell process the investments
and operating requirements (except labor) were taken from a paper by J. B.
Plummer, et al, of Shell. The operating labor was taken as the same as
for coal gasification. It should be pointed out that the investments for
the Shell process are subject to considerable error since they had to be
scaled down from data for plants about 10 times as large. This scaling
was done using a 0.7 exponent.
Compared to coal gasification, the investments for the Shell
residual oil process are lower if oxygen is used for the gasification but
slightly higher if air is used. One reason for the large difference in
gasification system cost between the air- and oxygen-blown versions of
the Shell process is that this is a high pressure process and hence the
air-blown version requires extensive air compression facilities.
The price of residual oil is, like other energy costs, quite
variable at this time. Therefore, two cost levels differing by a factor
of three were used, just as was done for the coal. These costs were $5
and $15 per barrel. At these feedstock costs the total operating costs
of the two processes are quite comparable if the gasification is done
with oxygen. If the gasification is done with air, the Shell process is
somewhat more expensive. The Shell process does deliver the reducing gas
at an elevated pressure which could be an advantage for some particular
applications. Of course, there are also high pressure coal gasification
processes.
The choice between coal and oil as a feedstock for making
reducing gas will depend largely on relative availability and convenience.
In most cases this will mean that the power plant should produce reducing gas
from the same fuel that it is using to fire its boilers.
-------
TABLE 11-11. COST COMPARISON FOR COAL AND OIL GASIFICATION SYSTEMS
1000—MW Power Plant 75 Percent Load Factor T^^s A FGD Process
Gasifier Feed
Gasification Agent
3
1975 Investment (10 $)
Gasification system
Oxygen plant
Total
3
Operating Cost (10 $/yr)
Item Unit Cost
Coal $10/T ($30/T)
Residual oil $5/B ($15/B)
Naphtha $12/B
Electricity $0.012/kUh
Steam $0.50/105 Btu
Water & chemicals
Operating labor $9/hr
Subtotal
Investment related (24% /yr)
Total
Operating Cost in mills /kWh
Coal
°2
8,110
5,180
13,290
788 (2,364)
19
-129
43
377
1,098 (2,674)
1,190
4,288 (5,864)
0.65 (0.89)
Air
8,372
8,372
811 (2,433)
19
45
259
1,134 (2,756)
2,009
3,143 (4,765)
0.48 (0.73)
Residual Oil
°2
3,933
5,180UJ
9,113
1,321 (3,964)
32
44
55
26, .
377(a)
1,855 (4,498)
2^187
4,042 (6,685)
0.62 (1.02)
Air
8,820
8,820
1,388 (4
32
33
-39
42, ^
259 (a)
1,715 (4
2J..17
3,832 (6
0.58 (1
,164)
,491)
,608)
.01)
oo
(a) Value taken to be same as for coal gasification.
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84
12. CONCLUSIONS
The following conclusions, arranged by topics, may be drawn
from the work discussed in this report.
Trace Constituents in Coal Based Reducing Gases
Experimental data from the direct analysis of trace
species in coal gases are not available, at least so
far as public information is concerned. In the
absence of such data, one can use either thermodynam-
ically calculated concentrations or concentrations
obtained by difference using coal and ash analyses.
The general level of trace species in raw coal gas
depends primarily on the composition of the specific
coal and on the gasification temperature, with more
species being present at higher temperatures.
Removing particulate matter from the coal gas and
the cooling necessarily associated with this operation
will remove the great majority of the trace species
from the gas. Water scrubbing will reduce the gas
temperature to such an extent that only very low
concentrations of trace species will remain in the
gas. Scrubbing decreases the thermal efficiency of
the system, thereby increasing the size of the
gasifier.
Effects of Trace Constituents
The effects of trace species in the coal-based
reductant gas upon the FGD SO--removal step should
be inconsequential, particularly if the gas is water
scrubbed. Even without scrubbing, the quantities of
coal involved indicate that the reductant gas would
add less than 4 percent to the quantity of trace
species already entering the FGD process via the flue
gas.
For the FGD processes which produce a concentrated
SO. stream, the process will be followed by an SO.
reduction reactor (e.g., the ASARCO type) and a
vapor-phase Glaus reactor. The only known adverse
effect of trace species on these systems is that
NH, will poison the Claus catalyst. Hence, the
reducing gas must be water scrubbed.
-------
85
For the FGD processes which use a liquid-phase
Glaus reactor, H2$ must be generated from the reducing
gas and sulfur. No catalyst is used in the liquid-
phase Glaus reactors, so no poisoning problems are
possible. The only known adverse effect of trace
species in this system is a side reaction caused by
high COS concentrations. The COS can be hydrolyzed
to H-S, and in fact this may occur in the shift
reactor which will precede the H S generator in at
least some possible versions of this system.
Costs of Reductant Systems
Reduction systems based on gasification of coal
or residual oil will be relatively expensive,
at least in comparison to systems based on natural
gas. For an FGD process that produces a concen-
trated SO- stream, the cost of a gasification-
based reduction system will be at least 15 percent
of the estimated cost of the FGD process itself.
For an FGD process that requires t^S, this figure
will be close to 30 percent unless technology can
be developed for using the reducing gas without
shifting or acid gas removal.
Complexity and Side Benefits
A gasifier-based reduction system will signifi-
cantly increase the complexity of the overall FGD
system. On the other hand, there are other possible
benefits of such a reduction system. The gasifier
could be used to supply some of the energy require-
ments of the FGD process, such as for stack gas
reheating. This energy could be supplied either
as excess by-product steam from the gasifier or as
excess fuel gas produced.
Use of Purchased EL,S as Reductant
e Purchased hydrogen sulfide will play only a minor
role as a reductant for SO- from power plant FGD
systems. The use of purchased H S will involve
almost entirely by-product streams obtained "across
the fence" from a source such as an oil refinery
or a natural gas processing plant, and hence geograph-
ical considerations will restrict this usage.
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86
13. RECOMMENDATIONS
Based on the work reported here, recommendations are made for
experimental studies in three areas, which are described below.
1. Trace Species in Coal Gases
Experimental measurements of the concentrations of trace species
in raw and water washed gases from leading coal gasification processes
should be made.
Probably the most meaningful way to do this would be to sample
streams from one of the commercial Koppers-Totzek gasifiers operating
overseas. Since this is the highest temperature gasifier now available,
it has the greatest potential for transferring trace species from the
coal into the gas.
2. Use of Coal Gas as SO^ Reductant
The use of coal gas as a reductant in the ASARCO SO. reduction
system should be studied. The objectives would be to determine whether
shifting of CO to H2 is required and whether trace species in the water
washed gas have any detrimental effects on the catalysts. Depending on
these results, it may also be desirable to study the effects on a cobalt-
molybdenum catalyst (the alternative for the first reduction stage).
It is our understanding that ASARCO is now conducting some
studies in this area using their catalyst and presumably their pilot plant.
Although this should be a valuable start, small-scale experiments are
limited in that they cannot normally use real coal gas but only blends
which simulate it. Therefore, tests on a demonstration-scale (>100-
megawatt) FGD system should be conducted as soon as this is feasible.
The NIPSCO unit of the Wellman-Lord process, which will be started up
soon, would be a good facility for such testing. The NIPSCO unit includes
the Allied Chemical SO. reduction process (using natural gas).
-------
87
3. Use of Coal Gas in H,,S Generation
The use of coal gas to provide a reductant for elemental sulfur
in. an H.S generation process, such as the Girdler or Thio-Pet processes,
should be studied. The objective would be to define the minimum amount
of processing (shift conversion, acid gas removal, etc.) of the coal gas
required for satisfactory operation of the H_S generation facility. The
use of H?S generated from coal gas in an FGD process such as the Citrate
process which contains a liquid-phase Glaus reactor should be studied.
The limitations of small-scale tests previously mentioned apply
as well to this research area. However, there are as yet no demonstration-
scale units for FGD processes of the type involved here. Such a unit for
the Citrate (or Citrex) process may be built in the near future. In the
meantime, some work in this area could be started using one of the two
existing Citrate pilot plants. The USBM pilot plant in Kellogg, Idaho,
has the advantages that it includes the Thio-Pet H S generation process
(using natural gas) and is currently in operation. However, this plant
treats smelter gas rather than power plant flue gas. The Peabody Citrex
pilot plant in Terre Haute, Indiana, has the advantage that it treats
flue gas from a coal-fired boiler. This plant is not presently in
operation but could be restarted. It does not include an H2$ generation
step (cylinder H?S was used).
Work in the above areas will focus on the major technical
uncertainties associated with the alternative principally addressed in
this work, i.e., the gasification of a "dirty" (sulfur-containing) fuel
to provide a reducing gas for use with FGD systems. There are alternatives
which accomplish the same objective and which were discussed only briefly
in this report. One is Foster-Wheeler's RESOX process, which reduces SO-
directly with coal. Another is the Atomics International Aqueous Carbonate
FGD process, which accomplishes its regeneration directly with coke or
coal. Research on these alternatives should also be continued until the
relative advantages of the various options are clarified and the field
can be narrowed.
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88
14. ACKNOWLEDGMENTS
A number of individuals and organizations contributed their
advice and assistance to this study. In this regard, the Project Officers
for these tasks, Dr. Charles Chatlynne and Mr. Roger C. Christman,* as well
as Mr. Richard D. Stern of EPA deserve mention. The Allied Chemical
Company and the Electric Power Research Institute provided a number of
helpful suggestions after reviewing draft reports on this project.
Several staff members of Battelle-Columbus also contributed to this study,
including G. R. Smithson, Jr., E. H. Hall, P. S. K. Choi, and D. A. Ball.
T. W. Zegers and D. M. Jenkins contributed to the study of the use of
purchased hydrogen sulfide. D. M. Treweek, a consultant to Battelle-
Columbus, contributed to the thermodynamic predictions of trace consti-
tuent distributions.
* Mr. Christman served as Project Officer for most of the work done on these
tasks.
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89
15. REFERENCES
(1) Doumani, T. F., Deery, R. F., Bradley, W. E., "Recovery of Sulfur
from Sulfur Dioxide in Waste Gases", Ind. Eng. Chem., 36, p 329,
1944.
(2) Murdock, D. L. and Atwood, G. A., "Kinetics of Catalytic Reduction
of Sulfur Dioxide with Hydrogen", Ind. Eng. Chem. Process Des.
Develop., 13 (3), p 254, 1974.
(3) Ryason, P. R. and Harkins, J., "Studies on a New Method of Simul-
taneously Removing Sulfur Dioxide and Oxides of Nitrogen from
Combustion Gases", J. Air Pollution Control Assoc., YJ_ (12), p 796,
December, 1967.
(4) Haas, L. A., McCormick, T. H., and Khalafalla, S. E., "Activity
Patterns in the Catalytic Reduction of SO. on Some Transition
Elements in Alumina", Bureau of Mines, Report of Investigation 7647,
1972.
(5) Haas, L. A. and Khalafalla, S. E., "Removing Sulfur Dioxide by
Carbon Monoxide Reduction", Instruments and Control Systems, 45,
p 101, March, 1972.
(6) Goetz, V. N., Sood, A., and Kittrell, J. R., "Catalyst Evaluation for
the Simultaneous Reduction of Sulfur Dioxide and Nitric Oxide by
Carbon Monoxide", Ind. Eng. Chem. Prod. Res. Develop., j.3 (2) ,
p 110, 1974.
(7) Quinlan, C. W., Okay, V. C., and Kittrell, J. R., "Kinetics and Yields
for Sulfur Dioxide Reduction by Carbon Monoxide", Ind. Eng. Chem.
Process Des. Develop., JL2_ (1), p 107, 1973.
(8) Querido, R. and Short, W. L., "Removal of Sulfur Dioxide from Stack
Gases by Catalytic Reduction to Elemental Sulfur with Carbon
Monoxide", Ind. Eng. Chem. Process Des. Develop., _12_ (1), p 10, 1973.
(9) Henderson, J. M. , "Reduction of S02 to Sulfur", Mining Congress
Journal, p 59, March, 1973.
(10) Pfeiffer, J. B. (Editor), Sulfur Removal and Recovery from Industrial
Processes, Advances in Chemistry Series 139, 1975.
(11) Hunter, W. D., Jr., and Michener, W., "New Elemental Sulfur Recovery
System Establishes Ability to Handle Roaster Gases", Engineering and
Mining Journal, p 119, June, 1973.
(12) Hunter, W. D., Jr., "Reducing SO. in Stack Gas to Elemental Sulfur",
paper presented at Flue Gas Desulfurization Symposium sponsored by
EPA, New Orleans, May, 1973.
-------
90 •
. REFERENCES
(Continued)
(13) Dryden, C. E. and Corrigan, T. E., "Chemical Technology Outline
Series", The Ohio State University, 1961.
(14) Steiner, P., Juntgen, H., and Kinoblauch, K., "Foster Wheeler's
RESOX Process", Heat Engineering, p 113, May-August, 1974.
(15) Bischoff, W. F., Jr., and Steiner, P., "Coal Converts S02 to S",
Chemical Engineering, p 74, January 6, 1975.
(16) Naber, J. E., et al, "New Shell Process Treats Glaus Off Gas",
Chemical Engineering Progress, J9_ (12), p 29, December, 1973.
(17) Beavon, D. K., "Add-On Process Slashes Glaus Tail Gas Pollution",
Chemical Engineering, p 71, December 13, 1971.
(18) Dautzenberg, F. M., Naber, J. E., and van Ginnekin, A.J.J., "Shell's
Flue Gas Desulfurization Process", Chem. Eng. Progress, 67 (8),
pp 86-91, August, 1971.
(19) Mitsak, D. M., "Koppers-Totzek: Take a Long Hard Look", 2nd Annual
Symposium on Coal Gasification, Liquefaction, and Utilization; Best
Prospects for Commercialization, University of Pittsburgh,
August 5-7, 1975.
(20) Magee, E. M., et al, "Evaluation of Pollution Control in Fossil
Fuel Conversion Processes K-T Gasifier", EPA 650/2-74-009a, U.S.
EPA, January, 1974.
(21) Telephone conversation with W. J. Rhodes, EPA, RTP, North Carolina,
regarding a visit to four commercial gasification plants in Westfield,
Scotland, Kutahya, Turkey, and Sasolburg, South Africa.
(22) Forney, A. J., et al, "Trace Element and Major Component Balances
Around the Synthane PDU Gasifier", ERDA Report No. PERC/TPR-75/1,
Pittsburgh Energy Research Center, Pittsburgh, Pa., August, 1975.
(23) Attari, A., "Fate of Trace Constituents of Coal During Gasification",
EPA 650/2-73-004, August, 1973.
(24) Personal communication from W. J. Rhodes, EPA, RTP, North Carolina,
August 28, 1975.
(25) Ruch, R. R., Cluskoter, A. J., and Shimp, N. F., "Distribution of
Trace Elements in Coal", presented at EPA Symposium on Environ-
mental Aspects of Fuel Conversion Technology, St. Louis, Missouri,
May 13-16, 1974.
(26) Cruise, D. R., J. Phys. Chem., 68., pp 3797-3802 (1964).
-------
91 and 92
15. REFERENCES
(Continued)
(27) Alexander, C. A., Hoyland, J. R. , and Treweek, D. N., "Evaluation of
Computerized Techniques for Predicting Chemical Reactivity and
Stability", Final Report from Battelle's Columbus Laboratories to
U.S. Coast Guard, DOT, April 11, 1975.
(28) U.S. Patent No. 3,071,454, January 1, 1963, referred to in
Reference 9.
(29) Telephone and letter communication with K. A. Charters, Senior Staff
Engineer, Thio-Pet Chemicals, Ltd., 304 Sixth Avenue West, Calgary,
Alberta, October 27, 1975.
(30) Kim, B. C., et al, "Development of Information for Standards of
Performance for the Fossil Fuel Conversion Industry", BCL report on
EPA Contract No. 68-02-0611, Task 7, October 11, 1974.
(31) Letter to Dr. Douglas W* Hissong (BCL) from Mr. James W. Bumbaugh
(Koppers Company, Inc.),, March 6, 1975.
(32) Letter to Dr. Benjamin C. Hsieh (BCL) from Mr. James W. Bumbaugh
(Koppers Company, Inc.), May 30, 1975.
(33) Letter to Mr. Paul Choi (BCL) from Mr. Donald M. Hess (C & I/Girdler,
Inc.), November 20, 1975.
(34) Personal communication between BCL and the U.S. Bureau of Mines,
Morgantown, West Virginia, 1975.
(35) Plummer, J. B., et al, "The Generation of Clean Gaseous Fuels from
Petroleum Residues", presented by Shell Development Company, Houston,
Texas, at the AIChE meeting, Tulsa, Oklahoma, March 11-13, 1974.
-------
APPENDIX A
FLUE GAS DESULFURIZATION PROCESSES
-------
A-2
APPENDIX A
FLUE GAS DESULFURIZATION PROCESSES
This appendix presents some additional information on flue gas
desulfurization processes in general and on the analysis which was made
of the reduction requirements of some of these processes.
Classification of Processes by Type of Operation
There are two general types of flue gas desulfurization processes -
the "throwaway" processes which produce an unusable mixture of sulfur com-
pounds and the "regenerable" processes which produce saleable by products.
The throwaway processes, particularly lime and limestone scrubbing, are
generally at a more advanced stage of development than the regenerable
processes. Only the regenerable processes are of interest in this study.
Figure A-l shows a functional classification of the regenerable
flue gas desulfurization processes. The three types of processes of interest
in this study are those using absorption by liquids, sorption by solids,
and catalytic reduction. Typical processes currently under development in
these three groups are shown in Figures A-2, A-3, and A-4.
Detail on Reduction Requirements of Sulfur-Producing Processes
A classification of sulfur-producing FGD processes according to
their reduction requirements was presented in Section 3 of the text. For
the Type A processes, which produce concentrated SO- streams, typical
S0_ concentrations in the product streams were given (Table 3-1) and
average compositions were developed for each of the two classes of Type
A processes. As backup information, brief process descriptions are given
in this section with emphasis on where the stream requiring reduction
comes from and what it contains.
-------
A-3
RECOVERY
PROCESSES
CATALYTIC
REDUCTION
CATALYTIC
ADSORPTION
SORPTION
BY SOLIDS
ABSORPTION
BY LIQUIDS
"•* ~-
— —
REDUCTION
BY CARBON
REDUCTION
BY CO
REDUCTION
BY H2S
WET
SYSTEMS
DRY
SYSTEMS
REGENERATION
BY REDUCTION
REGENERATION
BY WASHING
REGENERATION
BY HEATING
ORGANIC
SOLIDS
METAL
OXIDES
SODIUM
COMPOUNDS
ALKALINE EARTH
ABSORBENTS
A 1 \S A 1
LIQUIDS
I—
., ..
HIGH
LOW
CALCIUM
COMPOUNDS
MOLTEN
CARBONATES
POTASSIUM
COMPOUNDS
COMPOUNDS
AMMONIA
SCRUBBING
FIGURE A-l. SELECTED TECHNOLOGIES FOR THE REMOVAL OF S02 FROM STACK GAS
-------
S02 REMOVAL
ABSORPTION' BY LIQUIDS
ALKALI
ABSORBENTS
Wellman-Lord
(Davy Power gas)1''
N'a2S03
USBM
Citrate*
Sodium Citrate
S'tauf fcr
Powerclaus*
ALKALINE EARTH
ABSOKBANTS
^Example only--other similar processes exist
FIGURE A-2. SELECTED TECHNOLOGIES FOR TOE REMOVAL OF
SO FROM STACK GAS BY LIQUID ABSORPTION
-------
A-5
METAL OXIDES
Shell-UOP*
CuO
S02 REMOVAL
SORPTION BY SOLIDS
CARBON
ADSORPTION
Thermal
Regeneration
Foster Wheeler-BF*
Activated Carbon
Reduct ive
Regeneration
Westvaco"
Activated Carbon
^Example only--other similar processes exist.
FIGURE A-3. SELECTED TECHNOLOGIES FOR THE REMOVAL OF
SO FROM STACK GAS BY SOLIDS SORPTION
-------
A-6
S02 REMOVAL
CATALYTIC REDUCTION
HYDROGEN SULFIDE
CARBON MONOXIDE
CARBON
PETER SPENCE
CHEVRON
RESEARCH
PENN STATE
FIGURE A-4. SELECTED TECHNOLOGIES FOR THE REMOVAL OF
SO, FROM STACK GAS BY CATALYTIC REDUCTION
-------
A-7
Wellman-Lord (Davy Powergas)
Shell-UOP CuO
General Description. The SCL is absorbed by a solution of
sodium sulfite, thereby forming sodium bisulfite. The
solution is thermally regenerated in a crystallizer, liber-
ating so2.
Product Stream. The gas from the crystallizer contains
about 85 percent SCL with the balance being mostly water
vapor.
General Description. The SCi is picked up by reaction with
a solid sorbent composed of copper on alumina, thereby
forming CuSO,. The sorbent beds are periodically regenerated
with a hydrogen-steam mixture, liberating SO . The regen-
eration off-gas is concentrated in a quench tower and an
absorber/stripper system.
Product Stream. -The SCL concentration in the product stream
is about 6 percent leaving the reactor, about 46 percent
after the quench tower, and about 90 percent after the
absorber/stripper. After the absorber/stripper, the stream
contains only SO and HO, but before this it also contains
impurities fr.ocn the hydrogen used for regeneration.
'Special Reduction Requirements. A reductant is needed for
two steps in this process: (1) for regenerating the sorbent
and (2) for reducing the SO , For the former, Shell now
feels that hydrogen must be used, although they have done
some work on other reductants .
-------
A-8
Ammonia Scrubbing
General Description. The SO is absorbed by an aqueous
ammonia liquor containing sulfite, bisulfite, and sulfate.
The sulfites are thermally decomposed and the sulfates are
reduced with recycled sulfur product, thereby producing an
SCL-rich gas. In one particular process, the IFP-Catalytic
process, H?S is generated by reducing SO with an H -CO
mixture and this H S is used to reduce the SO to S in a
liquid phase catalytic Glaus reaction.
Product Stream. The SO -rich gas produced in the regener-
ation step contains about 27 percent SO , 24 percent HO,
and 49 percent NH . After the SO has been removed by
reduction, the ammonia in this stream is recycled to the
absorber.
Magnesia Scrubbing
General Description. The SO is absorbed by an aqueous
slurry of MgO, thereby forming hydra ted MgSO and MgSO .
These salts are then dried to remove surface water and water of
hydration and then decomposed in a calciner to form MgO and SO . The
.- •=- / *
decomposition is thermal for the MgSO , but coke is added '
to reduce the MgSO .
Product Stream. Because the calciner is direct-fired
(normally with oil), the product gas stream contains com-
bustion gases. When the desired product is H.SO excess
air is fed to the calciner to avoid contamination of the
acid with sulfur. In this case, the gas contains about
24 percent SO , 56 percent N , 9 percent CO , 7 percent HO,
and 4 percent 0 . When the desired product is sulfur, the
excess air will evidently be reduced or eliminated. When
this is done, carbon from the oil may serve to reduce some
of the MgSO , thereby decreasing the need for coke.
-------
A-9
Foster Wheeler-BF Carbon Adsorption
General Description. The SO is adsorbed onto activated
carbon in the form of H2S04. A thermal regeneration at
about 650 C (1200 F) decomposes the ^SO^, thereby
i. T
liberating the SC^.
Product Stream. Because the regeneration temperature is so
high, oxygen in the system results in combustion of some
of the carbon to form C02. The oxygen comes from two
sources—air initially present (adsorbed on the carbon)
and 0 evolved from decomposition of the H^SO^. The
product gas also contains N^ from the initial air and H^O
from the decomposition of the H2S°4' The gas comPosition
is about 23 percent SO , 33 percent N2, 21 percent CO^
and 23 percent HO.
U.S.B.M. Citrate
General Description. -The flue gas is cooled and washed in
a spray tower and then scrubbed with an aqueous solution
of sodium citrate, citric acid, and sodium thiosulfate.
These compounds buffer the solution, thereby increasing the
absorption capacity, and inhibit oxidation of the sulfite
to sulfate. The resulting solution, which contains about
10 g SCL/liter, is contacted with H2S in a stirred reactor,
thereby reducing the SCL to sulfur. The sulfur is separated
by oil or froth flotation and melting, and the citrate solution
is recycled to the absorber. The H2S is normally generated
by reducing sulfur, in which case 2/3 of the total sulfur
produced is recycled in this manner.
-------
A-10
Stauffer Powerclaus
General Description.-The flue gas is cooled and washed in
a spray tower and then scrubbed with an aqueous solution
of Na^PO^. The resulting solution is contacted with H S in
a stirred reactor, thereby reducing the SO to sulfur. The
sulfur is recovered by one of several available methods and
is purified by autoclaving.
Westvaco Carbon Adsorption
General Description. The SO is adsorbed onto activated
carbon in the form of H.SO.. The H SO, is reduced to sulfur
24 24
with H S, and the sulfur is stripped from the carbon with
steam. Both the adsorption and the reduction steps are
conducted in fluidized beds. The H S is normally generated
by reducing sulfur with H in a third fluidized bed. When
this is done, only 1/4 of the total sulfur produced is
stripped from the carbon, and the remainder is converted to
H0S while still on the carbon.
Aqueous Carbonate Process
General Description. -The flue gas is scrubbed with a
sodium carbonate solution in a modified spray dryer, the
reaction product being a dry powder containing sodium
sulfite and sulfate. This product is reduced with coke
(or other carbon source) in a molten salt reactor. The
molten product from this reactor is quenched and dissolved
in water. The resulting solution is cooled, filtered, and
reacted with C0? (from the reduction reactor) to liberate
H S. The H2S-rich gases are sent to a Glaus plant for
recovery of elemental sulfur. The carbonation generates
sodium carbonate for recycle to the scrubber.
-------
A-11
Status of Regenerable FGD Processes
Since this report deals with regenerable FGD processes, a brief
review of the development status of these processes is appropriate at this
point. Such a review has recently been provided by Murthy, et al.
This review will be highlighted in this section, with emphasis on the
processes most closely related to the subject of this report.
A brief summary of the status of the regenerable processes is
presented in Table A-l. Additional details on selected processes follow.
MgO Scrubbing*
The magnesia scrubbing process has been shown to work with no
major problems on a 1QO-MW coal-fired power plant boiler. The operating
cm-stream availability of the units was 60 percent compared to boiler on-
(2)
stream time of 90 percent. The Potomac Electric Power Company estimated
the installed cost at $100 per kW.* Both MgO recycle possibilities and
material balance (loss of MgO) were undetermined as of 1974. The unit
was shut down due to the closing down of the spent magnesia regenerating
plant at Rumford, Rhode Island. A 120-MW unit at the Eddystone station of
Philadelphia Electric Company should be ready for operation in 1976. Economic
operation of the MgO process may require a central regenerating plant for spent
magnesia.
Wellman-Lord Process
The Wellman-Lord system has been successfully operated on
tail gas from Claus and H?SO, plants, and on oil-fired boiler flue gas
but not on coal-fired boiler flue gas. Demonstration on coal-fired
boilers is scheduled to start soon on a 115-MW boiler at the Dean
Mitchell station of Northern Indiana Public Service Company (NIPSCO) .
The 85 percent concentrated S02 stream from the Wellman-Lord regenerator
will feed an Allied Chemical SO reduction process to produce >99.5
percent pure sulfur. The NIPSCO estimate for capital cost is $117/kW
and operating cost is 6.3 mills/kWh or 65 cents per million Btu input.
* Costs reported in this section should be considered approximate to ±40
percent as they are highly dependent on many site and situation specific
conditions.
-------
TABLE A-l. BRIEF STATUS SUMMARY OF REGENERATIVE PROCESSES
(a)
Process Name
Year, Installation Site,
Vendor, Size, and Type of Boiler
Status
MgO Scrubbing
Wellman-Lord
Cat-Ox
Chiyoda CT-101
Citrate^
1972, Mystic, Boston Edison, Chemico, 150 MW, oil
1974, Dickerson, Potomac Electric, Chemico, 100 MW,
coal
1976?, Eddystone, Philadelphia Electric, UEC, 120 MW,
coal
1976?, D. H. Mitchell, NIPSCO, Davy Powergas/
Allied Chemical, 115 MW, coal
1972, Wood River, Illinois Power Co., Monsanto,
110 MW, coal
1975, Scholz power plant. Gulf Power Co., Chiyoda,
23 MW, coal
1973, Pfizer's Vigo Chemical complex, McK.ee/
Peabody, 1-MW slipstream, coal
Shut down since June, 1974
Shut down since July, 1975
Shut down for modification
Under construction
Shut down since 1974
Operating since June, 1975
Shut down September, 1974
after data collection
-------
TABLE A-l. (Continued)
Process Name
Year, Installation Site,
Vendor, Size, and Type of Boiler
Status
Phosphate
Catalytic-IFP
Consol-Potassium Salts
AI-ACP
FW-BF
Westvaco
SFGD
1974, Norwalk Harbor Station, Connecticut Power &
Light Co., Stauffer, 0.1 MW, oil
No data in open literature
1972, Cromby, Philadelphia Electric Co., Consol/
Bechtel, 10 MW, coal
1971, Mohave, So. California Edison, Al, 0.5 MW, oil
1975, Scholz power plant. Gulf Power Co., Foster
Wheeler, 20 MW, coal
1970, Westvaco Research Center, Westvaco, 0.2 MW, oil
1974, Big Bend Station, Tampa Electric, Shell, 0.6-MW
slipstream, coal
Shut down June, 1974, after
data collection
No data in open literature
Shut down since 1972 after
data collection (a smaller
plant, 1000 acfm was operated
till 1975)
Shut down 1972 after data
collection on open-loop system
Started commissioning
January, 1975, many problems
Shut down 1974 after data
collection
Tests are in progress
(a) Installations in the United States only are discussed in this table.
(b) Extensive pilot-plant studies on a lead smelter gas are being conducted by USBM Salt Lake City
Metallurgy Research Center.
-------
A-14
Citrate Process
Many versions of the Citrate process are being offered now.
The Peabody Citrex process and the McKee Citrate process are based largely
on the results and expertise gained from a Citrate process pilot plant
operated during 1973-74 by McKee, Peabody, and Pfizer at Pfizer's Vigo
(4)
Chemical complex in Terre Haute, Indiana. Flue gas to the pilot plant
was fed from a coal-fired boiler stack breaching. The Citrate process
was pioneered by the USBM Salt Lake City Metallurgy Research Center.
The USBM pilot plant at Bunker Hill lead smelter, Kellogg, Idaho, is the
only currently operating pilot plant. This plant employs an H S generator
built by the Thiopet Chemical Company, Alberta, Canada, using natural gas
as the feedstock. The estimated cost of full-scale Citrate process ranges
from $50-60 per kW.
Phosphate (Powerclaus) Process
Stauffer's phosphate process is similar to the Citrate process
in that the phosphate radical replaces the citrate radical as the hydrogen
acceptor and buffering agent used to increase the solubility of S02 in
water. The various steps in the process were tested separately at the
10 to 30 scfm level at Stauffer's Delaware City CS plant. A 100 scfm pilot
plant has been operated on an oil-fired boiler in Connecticut, demonstrating
all features of the process except the H,S generation. Stauffer has
( 6)
licensed the process to Chemico.
Catalytic-IFP Process
The Catalytic-IFP process also is called the Fumeless Ammonia
Scrubbing process. The process promises to produce sulfur by liquid-
Glaus regeneration step. Details of the process status are still
proprietary.
Other ammonia scrubbing processes are available, and some of
these have been used for SO absorption in a number of foreign plants,
-------
A-15
but in these applications the products are primarily ammonium compounds
for fertilizer.
Consol Potassium Salts Process
The Consol potassium salts process is a new regenerative
process which has been demonstrated on both a 1000 acfm and 30,000 acfm
pilot plant. The latter pilot plant operated with flue gas drawn from
a coal-fired power plant of Philadelphia Electric Company. The exact
status of the process is not clear. Economics estimated by the Rust
Engineering Company for 1977 startup indicate investment costs of $70/kW
(8)
arid operating cost of 4.1 mills/kWh.
Aqueous Carbonate Process
Atomics International's aqueous carbonate process has been
under development since 1970 as an alternative to AI's molten carbonate
process tried earlier at the Arthur Kill Station of Consolidated Edison
Company. The molten carbonate process was a victim of severe materials
corrosion problems. The new process (AGP) was developed with in-house
funding and shows better promise of success. The process can be used
as open loop if large quantities of cheap sodium carbonate are available.
If the economics dictate regeneration and closed-loop operation, AI is
(9)
prepared to offer a completely closed-loop system.
The entire process has not yet been tested in a single
installation, but both the scrubbing system and the reduction reactor
have been tested separately. The size of the reduction reactor used
was equivalent to 3-5 MW of power plant capacity.
Carbon Adsorption Processes
The two carbon adsorption processes plus the Shell CuO process
are the only dry processes listed in Table A-l. The Foster Wheeler-BF
carbon process is the only dry process undergoing large-scale demonstration
-------
A-16
on a power plant boiler. The Westvaco process, although shown to be
technically feasible from pilot-plant tests at Westvaco Research Center,
is still undergoing evaluation for large-scale applicability.
Shell CuO Process
Demonstration units of the Shell process have been operated on
an oil-fired process heater at a Shell refinery in The Netherlands (600
scfm) and an oil-fired boiler plus a Glaus plant in Japan (79,000 scfm).
Small units have been operated on slipstreams from coal-fired boilers in
The Netherlands and in Florida.
Other Processes
The Cat-Ox and Chiyoda processes are not considered further
here because they do not use reductants. The Cat-Ox process produces
sulfuric acid and the Chiyoda process produces gypsum (CaSO.). Whether
the Chiyoda process should be considered a regenerable process is
debatable.
-------
A-17 and A-18
REFERENCES
(1) Murthy, K. S., Rosenberg, H. S., and Engdahl, R. B., "Status and
Problems of Regenerable Flue Gas Desulfurization Processes", 68th
Annual AIChE Meeting, Los Angeles, California, November 16-20, 1975.
(2) Erdman, D. A., "Mag-Ox Scrubbing Experience at the Coal-Fired
Dickerson Station of the Potomac Electric Power Company", Symposium
on Flue Gas Desulfurization, Atlanta, Georgia, November 4-7, 1974.
(3) Mann, E. L. and Bailey, E. P., "Power Plant Flue Gas Desulfurization
by the Wellman-Lord SCL Process", Symposium on Flue Gas Desulfurization,
Atlanta, Georgia, November 4-7, 1974.
(4) Vasan, S., "The Citrex Process for S0? Removal", Chemical Engineering
Progress, May, 1975.
(!)) McKinney, W. A., et al, "Pilot Plant Testing of the Citrate Process
for S0_ Emission Control", Symposium on Flue Gas Desulfurization,
Atlanta, Georgia, November 4-7, 1974.
(6) Saleem, A. and Sheehan, J. D., "Application of the Phosphate Process
to a Large Utility Boiler", Paper No. 75-087, 68th Meeting of the
Air Pollution Control Association, Boston, Massachusetts,
June 15-18, 1975.
(7) "Efficient, Economical, Reliable Industrial SO Removal", promotional
literature on Fumeless Ammonia Scrubbing available from Catalytic,
Inc., Philadelphia, Pennsylvania.
(8) Struck, R. T., Gorin, E., and Clark, W. E., "Consol Stack Gas Process
for S0? Removal", 68th Annual AIChE Meeting, Los Angeles, California,
November 16-20, 1975.
(9) Botts, W. V. and Gary, D. C., "Regenerative Aqueous Carbonate Process
for Utility and Industrial SO Removal Applications", 167th ACS
Meeting, Los Angeles, California, April 4, 1974.
-------
APPENDIX B
USE OF PURCHASED HYDROGEN SULFIDE
-------
B-2
APPENDIX B
USE OF PURCHASED HYDROGEN SULFIDE
This appendix contains the details of the study which was made
of the feasibility of using purchased hydrogen sulfide as a reductant in
connection with flue gas desulfurization processes. The following sections
are included:
c Technical Considerations
c The Hydrogen Sulfide Market
e Geographic Considerations
• Shipping and Handling Considerations
• Conclusions.
Technical Considerations
Process Stoichiometry
The reduction of sulfur dioxide with hydrogen sulfide is the
well known Glaus reaction. The Stoichiometry of this reaction is
2H2S + S02 = 3S + 2H20.
On a weight basis, about one unit of hydrogen sulfide is required per unit
of sulfur dioxide. On a sulfur-equivalent basis, two units of sulfur in
the form of hydrogen sulfide are required to reduce one unit of sulfur in
the form of sulfur dioxide. In other words, purchase of hydrogen sulfide
to reduce the sulfur dioxide triples the amount of sulfur which the utility
must dispose of by sale.
Properties of Hydrogen Sulfide
Physical Properties. The physical properties of hydrogen sulfide
are summarized in Table B-l. The vapor pressure and density are shown as
functions of temperature in Figures B-l and B-2, respectively. Hydrogen
sulfide is a gas at ambient temperatures, but it may be liquefied under
-------
B-3
TABLE B-l. PHYSICAL PROPERTIES OF HYDROGEN SULFIDE
Color
Corrosivity
Molecular Weight
Melting Point, C/F
Boiling Point at 14.7 psia, C/F
Vapor Pressure at 25.5 C
Specific Gravity Liquid 15.5 C/15.5 C
(60 F/60 F)
Density (pounds/gallon at 15.5 C)
Critical Temperature, C/F
Critical Pressure, psia
Flammability Limits in Air, vol percent
Cp Vapor (Btu/lb) at 15.5 C and 14.7 psia
Cv Vapor (Btu/lb) at 15.5 C and 14.7 psia
Cp Liquid (Btu/lb) at -46 to -18 C
(-50 to 0 F)
At 15.5 C and 14.7 psia
Auto Ignition Temperature, C/F
Heat of Vaporization (Btu/lb) at -60.2 C
(-76.4 F) and 14.7 psia
15.5 C (60 F) and 229 psia
Colorless
Corrosive to eyes and
respiratory tissues
34.08
-82.9/-117.2
-60.2/-76.4
20 atm
0.790
6.58
100.4/212.7
1306
4.3-46 percent
0.254
0.192
0.454
0.532
260/500
236.3
191.1
Source: Stauffer Chemical Company, "Safe Handling of Liquid Hydrogen
Sulflde".
-------
Temperature, C
2,000
-100
3
O
!-j
0)
E
Q)
M
3
CO
0)
a,
1,000 -
600
400
200 _
100 _
60
40
O
a.
20
10
-75
i
-50
-25
L_
25
50
t_
75
t
-200
mm of mercury
100
L_
2,000
1,000
600
400
200
100
i
•150
-100 -50 0 50 100
Temperature, F
150
200
250
cfl
CO
O.
„,
M
c/3
60 g.
cO
40 >
20
10
FIGURE B-l. H S VAPOR PRESSURE VERSUS TEMPERATURE
Source: Stauffer Chemical Co., "Safe Handling of
Liquid Hydrogen Sulfide".
-------
B-5
-------
B-6
pressure. The vapor pressure of hydrogen sulfide is higher than that of
ammonia and propane but lower than carbon dioxide. This means that
stronger tanks would be required to ship and store hydrogen sulfide than
are used for ammonia to liquefied petroleum gases.
Toxicity. Hydrogen sulfide is a highly toxic gas, extremely
dangerous in very low concentrations. The threshold limit value for
hydrogen sulfide is 10 ppm as defined by the American Conference of
Governmental Industrial Hygienists. This represents the condition under
which it is believed that nearly all workers may be repeatedly exposed,
day after day, without adverse effects. It is both an irritant and an
asphyxiant and may be characterized by its repugnant odor of "rotten eggs".
The toxic hazard is aggravated because unprotected exposure to nontoxic
levels of hydrogen sulfide will result in temporary olfactory fatigue
and can give a false sense of security in dangerous concentrations.
As the concentration increases (200-500 ppm), effects on the
nervous system become pronounced. Headaches, dizziness, excitement,
staggering gait, and diarrhea are some of the results that occur.
Bronchitis or bronchial pneumonia may also follow.
The nervous system response reverses as hydrogen sulfide
concentration increases. While at lower concentrations it may be a
depressant, at higher values, it stimulates. Still higher values will
paralyze the respiratory center.
The acute effects of hydrogen sulfide inhalation represent the
most serious threat in handling. Although it is not a cumulative poison,
severe exposure may result in brain damage from oxygen deficiency if
breathing is not rapidly regained. Permanent olfactory damage (loss of
sense of smell) may result from extended exposure at less than lethal
levels. Exposures in the 700-1000 ppm range may be fatal in 30 minutes.
Very high concentrations will be instantly fatal.
The effects of hydrogen sulfide on humans are summarized in
Table B-2. Liquid hydrogen sulfide poses another hazard in that contact
with the skin can result in severe frost bite.
Trained operating personnel at chemical plants and petroleum
refineries can handle large volumes of toxic materials like hydrogen
-------
B-7
TABLE B-2, ACUTE EFFECTS OF HYDROGEN SULFIDE GAS
Hydrogen Sulfide, H2S
Concentration, ppm
0.13
10.0
50-100
200-300
500-700
900
1000
Time
Sniff
8 hours
1 hour
1 hour
1/2: hour
Minutes
Minutes
Effect
Odor detectable
Threshold limit
Mucous membrane
Mucous membrane
(severe)
Coma
May be fatal
Fatal
irritation
irritation
Source: Stauffer Chemical Co., "Safe Handling of Liquid Hydrogen Sulfide".
-------
B-8
sulfide without undue hazard. Operating and other personnel at electricity
generating plants are not trained to handle such hazardous materials. The
magnitude of the potential safety problems cannot be discounted. The
toxicity of hydrogen sulfide may cause problems at power plants.
Corrosion Properties. Severe corrosion and cracking problems
in piping and storage vessels are experienced if water in any significant
amount is associated with the hydrogen sulfide. The hydrogen sulfide in
combination with water can attack metals and their oxidation products in
storage vessels or piping to form sulfide, which is hazardous due to its
pyrophoric nature. Stress cracking is also a direct result of these
conditions.
All metallic portions anticipated for use in a liquid hydrogen
sulfide system should be carefully reviewed in the light of the publica-
tion of the National Association of Corrosion Engineers. This publi-
cation is considered to be one of the most comprehensive studies of the
sulfide stress cracking phenomenon and should be considered basic to the
reference material for any plant using liquid hydrogen sulfide.
Hydrogen sulfide piping generally is constructed of 316 Stainless
Steel to minimize the corrosion problems. The corrosion properties of
hydrogen sulfide are severe enough to cause great concern for any poten-
tial power plant use. For these reasons, the installation cost of
equipment to handle hydrogen sulfide and the cost of the special materials,
monitors, and precautions required because of the corrosive nature of
hydrogen sulfide are an important consideration in evaluating the potential
use of this reductant.
Fire and Explosive Properties. The flammable or explosive
range for hydrogen sulfide gas in air is 4.3-46 percent by volume. The
vapor is slightly heavier than air, which permits the gas to "pocket" in
"still" conditions. The auto-ignition temperature in air of hydrogen
sulfide is 500 F. For this reason, all sources of ignition such as open
flames, sparks, glowing surfaces, and all electrical equipment should be
protected to prevent contact with the gas. Only non-sparking tools should
be used in areas where hydrogen sulfide. is handled.
-------
B-9
The Hydrogen Sulfide Market
Supply
Hydrogen sulfide currently is a specialty chemical with a
limited market. The current market and distribution system would not
suffice for large scale uses like flue gas desulfurization. Therefore,
historic data are of little use in analyzing supply.
Potential Supply
Although there are some specific processes for the manufacture
of hydrogen sulfide, most hydrogen sulfide is recovered from either natural
gas or from refinery gas streams. The hydrogen sulfide is removed from
these gas streams by absorption in an alkaline solution, generally an
ethanol amine. Concentrated hydrogen sulfide is then recovered by-
stripping the amine solution. The hydrogen sulfide generally is converted
into sulfur, although at some sms.ll gas processing plants it may be flared.
The largest potential sources of hydrogen sulfide are petroleum
refineries and natural gas processing plants. In the future, hydrogen
sulfide also may be. available at coal gasification plants and coal desul-
furization plants using hydrodesulfurization (for example, the solvent-
refined coal process). In 1973, refineries converted hydrogen sulfide
recovered from process streams into an estimated 5,330 metric tons*
(2)
sulfur per day. The hydrogen sulfide removed from refinery gases is
generally converted to sulfur in the refinery. The two major sources
of hydrogen sulfide in the refinery are fuel gases and recycle streams
from hydrogen processing operations. There are also a few minor sources
of hydrogen sulfide in refineries like sour water strippers. In the
past, much of the hydrogen sulfide in fuel gas has not been recovered.
The quantity of hydrogen sulfide removed from process streams
is expected to increase significantly due to environmental regulations
requiring removal of hydrogen sulfide from refinery fuel gases. It is
estimated that the total hydrogen sulfide generated at refineries in 1973
* A metric ton equal to 2205 pounds. A long ton, which is the standard
unit of measure in the sulfur industry, is 2240 pounds. A ton in this
report refers to a short ton of 2000 pounds.
-------
B-10
was 11,000 metric tons sulfur equivalent per day and 11,500 in 1974. The
unrecovered 6000 metric tons per day will be recovered as soon as regulations
limiting sulfur content of refinery fuel gas become effective. Further-
more, the higher sulfur content of crude oil processed in U.S. refineries,
increased refining capacity, and increased hydrodesulfurization will
result in increased hydrogen sulfide recovery. It is estimated that by
1980 the recovery of hydrogen sulfide at refineries will increase to the
equivalent of 17,700 metric tons per day (see Table B-3).
The second largest source of hydrogen sulfide is natural gas
processing plants. In 1973, it is estimated that the average sulfur
(3)
production from these sulfur plants was 2,440 metric tons per day. It
may be noted that this is significantly higher than the 1,740 metric
tons per day reported in the literature. The higher estimate was based
on industry contacts. It is believed that reported production omits
some significant plants.
In addition to the sulfur recovered at the natural gas process-
ing plants, there is an estimated 1,230 metric tons per day hydrogen sulfide
removed at small gas plants which is not converted to sulfur. This
hydrogen sulfide currently is burned and vented to the atmosphere. It
is uneconomic to recover as sulfur, but it might be liquefied and sold
if a suitable market were developed. If future regulations prohibit
flaring of hydrogen sulfide, sale of hydrogen sulfide would probably be
more attractive than making sulfur at these small plants.
Natural gas production in the lower 48 states is expected to
decline in the future. The sulfur content of new gas fields is not known.
There will probably be environmental pressure to flare less hydrogen
sulfide at gas plants. In view of these uncertainties, it has been assumed
that the production and availability of hydrogen sulfide from gas plants
will remain constant through 1980 (see Table B-3).
It should be noted that there is some discrepancy between
statistics on hydrogen sulfide availability. The U.S. Bureau of Mines
reports recovered elemental sulfur, which includes both petroleum refiner-
ies and natural gas processing plants, in 1973 was 2.4 million metric tons.
This is equivalent to 6,619 metric tons per day, somewhat below the estimated
7,770 metric tons per day (5,330 from refineries plus 2,440 from natural gas
-------
B-ll
TABLE B-3. ESTIMATED HS AVAILABLE FROM OIL AND GAS PROCESSING
(a)
Average sulfur content of crude oil, wt. 1*
Crude oil processing rate, 10 bbl/D.
Natural gas processing rate, 10 SCF/D
3
Total H S produced, 10 Met tons/D
(d)
Oil refineries , N
\e )
Natural gas plants
Total
o
H S sent to Glaus plants, 10 Met tons/D
(£)
Oil refineries v ' , ,.
( e )
Natural gas plants
Total
1974
0.88
12'9M
55.6<-c;
11.5
3.7
15.2
5.6
2.4
8.0
1980 (est.)
1.07
16.6
55.6
17.7
3.7
21.4
17.7
2.4
20.1
(a) NPRA Survey, Letter from D. C. O'Hara to R. E. Train (EPA), August 30,
1974.
(b) Oil and Gas Journal, 72, April 1, 1974, plus crude run/capacity ratios
from Oil and Gas Journal, 1970-1974.
(c) Oil and Gas Journal, 72, July 8, 1974.
(d) Based on average crude oil density of 301 Ib/B and 70 percent S in
crude going to HLS.
(e) Murthy, K. S., BCL report on EPA Contract 68-02-0611, August 28, 1974.
(See Reference 3.)
(f) Based on 70 percent utilization of capacity; capacity based on Beers,
W. D., "Characterization of Glaus Plant Emissions", Process Research,
Inc., Report No. EPA R2-73-188, April, 1973.
-------
B-12
plants) discussed above. A small portion of this discrepancy may be due
to a portion of the hydrogen sulfide being converted directly to sulfuric
acid or sold as hydrogen sulfide. These probably account for less than
500 metric tons per day. The higher estimate is believed to be more
accurate.
As mentioned earlier, it is expected that coal desulfurization
and coal gasification plants may provide additional potential sources of
hydrogen sulfide. In particular, in the event of coal gasification
located at an electric utility, enough hydrogen sulfide would be recovered
to satisfy the needs of a coal-burning generation facility of twice the
capacity o-f the gas burning facility. Although these coal-processing
plants may be a significant local source of hydrogen sulfide for a
specific power plant, their contribution to the national production will
be insignificant through 1985.
Historic Supply
The historic supply of hydrogen sulfide has little bearing on
the potential supply for pollution abatement uses. Hydrogen sulfide is
used mainly for the manufacture of alkali sulfides, which are used in the
dye and leather industries. There are also other minor uses. The Bureau
of Mines reports that between 1969 and 1973 the total production of sulfur
dioxide and hydrogen sulfide varied between 88,000 and 149,000 metric tons
per year sulfur equivalent. Industry sources indicate that these data
may be in error, but in any event the total sales of hydrogen sulfide are
small.
The major merchandiser of hydrogen sulfide is the Stauffer
Chemical Company, which shipped approximately 29,000 metric tons of hydrogen
sulfide in 1973. The second producer is the Montana Sulphur and Chemical
Company, which shipped between 5,440 and 6,350 metric tons of hydrogen
sulfide in 1973. PPG Industries and a few other conpanies sell small
quantities of hydrogen sulfide in cylinders. Most of the hydrogen sulfide
sold in the United States originates at petroleum refineries.
-------
B-13
Thio-Pet Chemicals Ltd, operates a pe'trochemical plant at Fort
Saskatchewan, Alberta for manufacture of carbon disulfide and hydrogen
sulfide by-product by the methane-sulfur process.1 The Thio-Pet plant
commenced operations in spring 1965 with an initial daily hydrogen sulfide
capacity of 8 tons and was expanded in 1967 to a hydrogen sulfide capacity
of 13 tons per day. A portion of: the hydrogen sulfide by-product is sold
by pipeline to an adjacent chemical plant as an impure vapor at 150 psig.
Facilities for production of purified hydrogen sulfide liquid product were
put in service in December 1965. Pure liquid hydrogen sulfide is marketed
in pressure cylinders and in bulk by railroad tank cars. The first tank
car shipment was made in June 1966.
Hydrogen sulfide produced by Thio-Pet is a by-product of the
methane-sulfur carbon disulfide process represented by the following
equation:
CH. + 4S > CS- + 2H.S.
4 22
This reaction is carried out over a catalyst at high temperature and low
pressure. It is customary for the hydrogen sulfide by-product to be recon-
verted to sulfur by the modified Glaus process and to recycle the sulfur
recovered from the by-product hydrogen sulfide to the carbon disulfide
reactor. Reconversion of hydrogen sulfide to sulfur w,as not, however,
economically justified for the Thio-Pet operation because of the relatively
small capacity of the plant. Instead, the hydrogen sulfide is purified to
varying degrees and sold as both vapor and liquid.
Demand
The demand for hydrogen sulfide to reduce all the SO- which could
be recovered from power plant flue gas far exceeds the potential supply of
hydrogen sulfide. The total emissions of SO,, from power plants in the
(7)
United States in 1970 was about 19.4 million tons. Between 1960 and
1970 these emissions had increased at an average rate of about 6,7 percent
per year. If this rate were to continue, the potential (uncontrolled)
S0? emissions in 1980 would be about 37 million tons. To reduce emissions
some power plants have recently switched to lower sulfur fuels. For
example, the average sulfur content of coal burned at power plants
-------
B-14
(8)
decreased from about 2.6 percent in 1969 to 2.3 percent in 1972. On
the other hand, the shortage of natural gas will almost certainly result
in a decrease in the consumption of this fuel at power plants, and con-
version to other fossil fuels will increase the SO,, emissions. In spite
of the uncertainty in the potential S0« emissions, the following calcu-
lation illustrates the point. A total SO emission of only 25 million
tons per year would require about 66,100 metric tons of H S per day for
reduction. This is about 3.3 times the estimated potential H S supply in
1980 (see Table B-3). Looking at it in the other direction, there is
potentially enough hydrogen sulfide available to reduce about one third
of the SO- produced at power plants.
Requirements for 10QO-MW Power Plant
A 1000-megawatt power plant operating at full load on 3.5
percent sulfur coal and recovering 90 percent of the SO from the flue
gas will require about 10,600 scfm of reductant (based on 2 moles of
reductant per mole of S0?). If H^S is used, this equivalent to about 658
metric tons per day. If the H S was received in liquefied form, the
requirement would be about 805,000 liters (213,000 gallons) per day.
Using present shipping practices (see section on Shipping), this corresponds
to about 10 railroad tank cars per day.
Prices
There are currently only two significant marketers of hydrogen
sulfide, Stauffer Chemical Company and Montana Sulfur and Chemical
Company. Stauffer's current list price is $0.10 per pound (220 dollars
per metric ton). Montana Sulfur quotes the price of $0.065 per pound (14.3
dollars per metric ton), and most hydrogen sulfide is probably sold at the
lower figure.
-------
B-15
The current prices have little relevance to the price of the
large volumes of hydrogen sulf'ide that might be required by an electric
utility. Such purchases will probably be made on a long-term contract,
witih the price of hydrogen sulfide, f.o.b. refinery, equal to or slightly
less than the price of contained sulfur.
The market value of sulfur is quite volatile. In recent years,
the. price of recovered sulfur, f.o.b. plants, has ranged from $29.15 per
metric ton in 1969 to $15.45 in 1.973. ' In 1974, prices increased
rapidly. For example, the average export value for the first 7 months of
1974 was $32.63 per metric ton, compared with $19.38 in 1973. For the
(9)
month of July, the value of export shipments was $41.61 per metric ton.
Because a refinery or large natural gas processing plant con-
verts recovered hydrogen sulfide into sulfur, they should be indifferent
to selling the sulfur in elemental form or as hydrogen sulfide, provided they
can obtain an equivalent price. In most cases, by the time a utility
decided to purchase hydrogen sulfide from a refinery, the refinery would
have already invested in the sulfur plant. Therefore, the price of
hydrogen sulfide would have to cover depreciation and return on invest-
ment on the sulfur plant and perhaps on a sulfur-plant tail-gas treating
unit. In order to sell hydrogen sulfide, the refiner will have to install
liquefaction, pumping, storage, and shipping facilities. The price of
hydrogen sulfide will have to include depreciation and return on invest-
ment on these new investments as well as on the existing sulfur plant.
Because hydrogen sulfide 'and sulfur are generally considered a nuisance
by-product by refiners, in times of low-sulfur prices the refiner may be
willing to accept less than an "adequate" return on investment. Another
factor to consider is the additional tail gas cleanup which will be
required under EPA regulations. A refiner may be willing to sell H S at
a loss instead of investing in cleanup facilities for Glaus plants.
The existing refiner will have the option of either selling
hydrogen sulfide or manufacturing and selling sulfur. Thus, the minimum
price of hydrogen sulfide, on a contained sulfur basis, would be the price
of sulfur less direct operating expenses for sulfur manufacture plus the
cost: of handling and shipping hydrogen sulfide. The direct manufacturing
expenses are estimated to be between $6 and $18 per metric ton.
-------
B-16
Since high-purity H S is not required for reduction purposes,
gasification of coal or residual oil may prove feasible as a source of
H2S. In Section 11 of this report it is estimated that the cost of
producing H,S via gasification of coal or residual oil would be between
2.5 and 3.2 cents per pound ($55-71/metric ton) of H2S, depending on the feed-
stock cost and plant size. Thus, a power plant having access to an
"across the fence" source of H-S would probably need ti> purchase H S at
these price levels or lower to prove economically attractive.
Geographic Considerations
The major sources of hydrogen sulfide supply are not necessarily
located near electric utilities. The U.S. petroleum refineries ^nd their
sulfur plant capacities are summarized in Table B-4 and the natural gas
processing plants and associated sulfur production are summarized in
Table B-5. From these tables it can be seen that the greatest potential for
production of hydrogen sulfide exists in California, Texas, Louisiana, New
Jersey, Pennsylvania, and Delaware. In general, petroleum refineries are
apt to be located near power generation plants while gas processing
plants are more likely to be in isolated areas.
Electric power generation plants are generally located near
population centers. The location of the major coal and oil-fired power
plants is illustrated in Figure B-l. The power plants located near Chicago,
Toledo, Philadelphia, Los Angeles, San Francisco, and in New Jersey and
Delaware and a few in the Ohio River Valley are apt to be located very
near petroleum refineries. Rather short pipelines (10 miles or less)
would probably be the most practical means to deliver hydrogen sulfide to
power plants in these areas. Power plants in other areas would have to
obtain hydrogen sulfide either by rail or barge. For power plants located
farther from refineries or natural gas processing plants, transportation
costs could significantly affect the overall economics of flue gas
desulfurization using hydrogen sulfide as a reductant.
Another important geographic factor concerns the market for
sulfur. The largest market is in the Gulf Coast states, especially
Florida where the phosphate fertilizer industry is concentrated. In 1971,
-------
B-17
TABLE B-4. REFINERY GLAUS PLANT CAPACITIES
State
Crude Oil
Capacity(a),
1000 Met tons/D
Glaus Plant
CapacityCb),
Met_tons Sulfur/D
A leib ama
Alaska
Arkansas
California
Colorado
Delaware
Florida
Georgia
Hawaii
Illinois
Indiana
Kansas
Kentucky
Louisiana
Maryland
Michigan
Minnesota
Mississippi
Missouri
Montana
Nebraska
New Jersey
New Mexico
New York
North Dakota
Ohio
Oklahoma
Oreigon
Pennsylvania
Rhode Is land
Tennessee
Texas
Utah
Virginia
Washington
West Virginia
Wisconsin
Wyoming
TOTAL
5.31
7.33
6.55
234.60
6.76
19.15
0.68
1.68
8.66
142.48
73.54
52.94 •
21.68
212.21
3.27
18.11
23.71
41.90
14.09
18.95
0.68
80.98
6.46
14.04
7.22
79.89
63.12
2.26
88.63
1.03
3.97
477.10
16.59
6.57
46.39
2.67
4.86
19.53
1,835.59
9
25
2556
18
782
--
--
-~
572
425
44
--
585
_-
91
174
30
80
115
-.-
622
30
50
207
90
—
--
345
—
—
1207
12
50
--
..
15
__
8,134
(a) Based on crude oil density of 301 lb/bbl (1 metric ton = 7.31 bbl).
January 1, 1974 data.
(b) Beers, W. D., "Characterization of Claus Plant Emissions",
Process Research, Incorporated, Report No. EPA-R2-73-188, April,
1973. Added 5 percent for "greater than" figures.
-------
B-18
TABLE B-5. SALIENT DATA ON SULFUR RECOVERY
IN NATURAL GAS PROCESSING^3)
Arkansas
Florida
Mississippi
New Mexico
North Dakota
Texas
Wyoming
Total
Associated (Sour)
Gas Throughput,
106 scf/D
31.9
130
20
54.3
74.8
783.0
68.2
1,162.2
Total Sulfur
Production,
Met tons/D
5.3
650
221.4
30
116
1,320.6
100
2,443.3
(a) Source: Murthy, K. S. (Reference 3.)
-------
•--T-S.J *""*>• f-
oooo
.
jow- *ooi- root- ' loo*- »ooi- to.ooi-
cooo rooo 1000 »ooo 10.000 is. ooo
(3d
I
V£l
FIGURE B-l. LOCATION OF MAJOR COAL- AND OIL-FIRED POWER UNITS - 1971
Source: McGlamery, G. G., et al.
-------
B-20
10,630 metric tons sulfur per day were consumed in Florida, Louisiana,
and Texas. Transportation costs to ship sulfur' from the power plant to
customers can greatly affect the economics of a given project (3
metric tons sulfur are produced per metric ton sulfur removed from flue
gas).
Shipping and Handling Considerations
Shipping
In the section on Technical Considerations, the extremely toxic,
corrosive, and explosive nature of hydrogen sulfide was discussed.
In principle, liquefied hydrogen sulfide can be shipped by
tank truck, railroad car, barge, or pipeline. The toxic and corrosive
nature of hydrogen sulfide place severe limitations on the use of these
transportation modes. At the present time, liquefied hydrogen sulfide
is only transported by tank truck or railroad car at a cost of approxi-
mately 3 to 4 cents per metric ton-mile. Barge transportation would probably
cost somewhat less but would be of quite limited applicability. Pipeline
transportation probably is impractical for all except very short distances
because special lined or stainless steel pipelines would be required to
avoid corrosion.
The Code of Federal Regulations classifies hydrogen sulfide as
a flammable compressed gas (49 CFR 170.1, P 172.5) and shippers should
conform with applicable DOT regulations. For tank car shipments the
Code requires a pressure vessel classified as DOT-106A800X. These are
multiunit tank car tanks which are designed to be removed from the car
structure for filling and emptying. The maximum allowable water capacity
of each tank is 2600 pounds. For hydrogen sulfide the maximum permitted
filling density is 68 percent; this would give each tank a maximum capacity
of 1768 pounds hydrogen sulfide.
Compliance with the above Code would be impractical for bulk
shipments. Both U.S. companies that ship hydrogen sulfide in bulk have
a special permit from the Department of Transportation to use DOT-105A-600W
tank cars with a maximum loading capacity of approximately 70 tons, or
-------
B-21
tank trailers with a capacity of approximately 14 tons. The DOT permit
is valid for 2 years. Renewal is contingent upon providing the Department
of Transportation with satisfactory reports on safety inspections, total
shipments, and any accidents that may have occurred.
Storage
Whether the hydrogen sulfide is received by the power plant as
a liquid from shipping vessels or as a gas piped from a nearby source,
sufficient storage capacity must be provided to insure continuous opera-
tion. It is believed that a minimum supply for 15 days of operation
should be on hand. For a 1000-megawatt plant burning coal containing 3.5
percent sulfur, it is estimated that a total storage capacity for close
to 10,000 metric tons of hydrogen sulfide to be necessary. If the hydrogen
sulfide were to be shipped in railroad tank cars 'of present size, approxi-
mately 10 cars would be needed every day.
The volume occupied by 10,000 metric tons of liquefied hydrogen
sulfide is approximately 11,436 cubic meters (413,800 cubic feet) at room
temperature. The storage and handling of liquefied hydrogen sulfide is
complicated by the fact that the material must be kept under pressure.
The vapor pressure is 18 atmospheres at 21 C (70 F) and 31 atmospheres at
43 C (110 F). Also, special precautions must be taken to avoid problems
due to corrosion, particularly sulfide stress cracking. This is treated
rather extensively in Reference 1.
Even if a power plant would receive hydrogen sulfide by pipe-
line:, similar though somewhat less storage would be required. Hydrogen
sulfide would be produced in refineries at a relatively steady rate. The
dema.nd by the power plant could fluctuate by daily, weekly, and seasonal
cycles.
The H«S storage requirements could be reduced, albeit at the
expense of more H_S consumption, by providing an incinerator in which
some H S could be burned and then sent through the FGD system at times
when, the H S supply got ahead of the power plant demand.
-------
B-22
Conclusions
Hydrogen sulfide presently is marketed on a very modest scale,
and the market is not expected to expand significantly through 1985.
However, a large-scale use (such as flue gas desulfurization) would
require a vastly expanded supply which would involve oil refineries and
some natural gas processing plants. The potential supply of hydrogen
sulfide from these sources is sufficient to handle only a portion of the
power plants which will be installing flue gas desulfurization systems.
Geographical considerations will further restrict the use of
hydrogen sulfide at power plants. The present and potential sources of
hydrogen sulfide are concentrated in the South and the West. In contrast,
the demand for electricity is concentrated in the population density
centers of the East and Midwest. Since the costs of shipping hydrogen
sulfide are relatively high (approximately 3-4 cents per metric ton-mile), the
use of hydrogen sulfide as a reductant gas is not expected to be feasible
when the transportation distances exceed a few hundred miles. Some use
of hydrogen sulfide may be practical for power plants located near
refineries or sour gas wells.
It is concluded that purchased or merchant hydrogen sulfide
will play only a minor role as a reductant for sulfur dioxide from FGD
systems.
-------
B-23 and B-24
REFERENCES
(1) National Association of Corrosion Engineers Publication 1F166,
M. R. Chance, Chairman. "Sulfide Cracking Resistant Metallic
Materials for Values for Production and Pipeline Service",
September, 1966.
(2) Genco, J. M. and Tarn, S. S., "Characteristics of Sulfur from
Refinery Fuel Gas", Battelle's Columbus Laboratories, Final Report
to EPA, Contract No. 68-02-0611, March 29, 1974.
(3) Murthy, K. S., Final Report on "Characterization of Sulfur Recovery
in Oil and Natural Gas Production", EPA Contract No. 68-02-0611,
Industrial Studies Branch, EPA, Durham, North Carolina, p 52,
August 28, 1974.
(4) Oil and Gas Journal, 72. 80, July 8, 1974.
(!>) U.S. Bureau of Mines, Sulfur in 1973, Mineral Industry Surveys,
June 11, 1974.
(6) Geddes, T. H., "Manufacture and Transportation of Liquid Hydrogen
Sulphide", Quarterly Bulletin of Alberta Sulphur Research Ltd.,
_V, No. 11, 5-16, January-March, 1969.
(7) National Academy of Sciences, Office of the President, "Air Quality
and Stationary Source Emission Control", report to Senate Committee
on Public Works, Serial No. 94-4, 94th Congress, 1st Session, p 239,
March, 1975.
(8) Jimeson, R. M., Oral statement upon behalf of staff of the Federal
Power Commission before Environmental Protection Agency Stationary
Source Enforcement Proceeding, p 5, October 29, 1973.
(9) U.S. Bureau of Mines, Frasch and Recovered Sulfur in July, 1974,
Mineral Industry Surveys, September 23, 1974.
(10) McGlamery, G. G., et al, Conceptual Design and Cost Study, "Sulfur
Oxide Removal from Power Plant Stack Gas", TVA Bulletin Y-61,
EPA R2-73-244, EPA, Office of Research and Monitoring, Washington,
D.C., p 104, May, 1973.
-------
APPENDIX C
GASIFICATION PROCESSES
-------
C-2
APPENDIX C
GASIFICATION PROCESSES
This appendix contains additional information on the coal and
oil gasification processes discussed in the text. The following processes
are included:
Coal Gasification
(1) Koppers-Totzek gasifier
(2) Wellnian-Galusha gasifier
(3) Winkler gasifier
(4) Riley-Morgan gasifier
Oil Gasification
(1) Shell Gasification Process
(2) Texaco Synthesis Gas Generation Process.
Coal Gasification Processes
(1 2)
Koppers-Totzek Gasifier ' '
The Koppers-Totzek gasifier, shown schematically in Figure C-l,
is an oxygen blown, atmospheric pressure, entrained system. Over 50
gasifiers of this type have been built. In this system, pulverized
coal, 90 percent under 200 mesh, is screw fed from hoppers into pairs of
burners situated opposite each other on the same axis so that their jet
discharges converge. A gasifier may contain two or four such burner
heads. Oxygen conveys the coal from the screw feeders to the burner
nozzles at velocities sufficient to prevent flame propagation into the
feeder tubes (300 ft/sec). Coal is dried to about 1 weight percent
moisture, but moisture contents of 4 to 8 weight percent can be tolerated.
The gasifier is refractory-lined with a water jacket where process.steam
is raised and added to the burner nozzles in amounts sufficient to main-
tain the flame temperature at about 1930 C (3500 F). The gas temperature at
the reactor exit is about 1480 C (2700 F). About 50 percent of the coal ash
-------
C-3
WASTE HEAT
BOILER
HIGH PRESSURE
STEAM
-—GAS OUTLET
COAL
FEED WATER
COAL
wvwj
FIGURE C-l. THE KOPPERS-TOTZEK GASIFIER
-------
C-4
leaves as slag which is quenched and granulated in water at the bottom of
the reactor, and 50 percent leaves as fly ash with the product gas and is
removed in gas cleanup. Waste heat boilers are used to recover the sensible
heat in the high temperature product gas. The gas is further cooled and
cleaned by conventional, commercially proven processes.
The gasifier thermal efficiency (70 to 80 percent) will be
lower if the sensible heat in the product gas is not recovered. Essentially
no methane is produced in the gasifier. Gasifier outputs and typical raw
gas composition for different coals are shown in Table C-l. The K-T
gasifier is highly sophisticated in its design and operational aspects and
hence can be more reliable in operation but at a higher cost than other
... (3)
gasifiers.
' Cl 4}
The Wellman-Galusha (W-G) Gasifier '
The\ W-G gasifier (Figure C-2) employs a fixed bed and operates
similar to the Lurgi gasifier but at atmospheric pressure with air or
oxygen. The product gas heating value is 290 Btu/scf when oxygen blown.
Gas cleaning systems offered include the Holmes-Stretford system offered
by the Peabody Engineered Systems, Stamford, Connecticut.
Complete system design and construction services are offered
by A. G. McKee, Cleveland, Ohio, and Mason & Hanger-Silas Mason Co., Inc.
The McDowell-Wellman Company, Cleveland, Ohio, offers only the gasifier.
Presently, seven air-blown commercial units are operating in the U.S.
Typical composition of Wellman-Galusha gas from air- and oxygen-blown
systems is presented in Table C-2.
Winkler Gasifier(1>5)
The Winkler process is a commercially proven, fluidized-bed,
atmospheric pressure system for the production of synthesis gas from
coal. An air-blown version of the Winkler gasifier is available from
Davy Powergas of Lakeland, Flordia. A typical gasifier is shown in
Figure C-3.
The generator is a cylindrical, refractory-lined vessel, 12
to 18 feet in diameter and 65 feet high. Coal usually is crushed to
-------
C-5
TABLE C-l. K-T GASIFIER DATA FOR U.S. COALS
Gasifier Feed
Dried coal to gasifier analysis,
weight percent
C
H
N
S
0
Ash
Mois ture
Gross heating value, Btu/lb
Oxygen, ton/ton dried coal
Purity, percent
Process steam, Ib/ton dried coal
Gasifier Products
t
Jacket steam, Ib/ton dried coal
High pressure steam, Ib/ton dried
coal
Raw Gas
Analysis (dry basis) , volume
percent
CO
co2
H2
N2
COS
Western Coal
56.76
4.24
1.01
0.67
13.18
22.14
2.00
9,888.0
0.649
98.0
272.9
347.8
2,147.1
58.68
7.04
32.86
1.12
0.28
0.02
100.00
Illinois Coal
61.94
4.36
0.97
4.88
6.73
19.12
2.00
11,388.0
0.704
98.0
541.3
404.9
2,292.2
55.38
7.04
34.62
1.01
1.83
0.12
100.00
Eastern Coal
69.88
4.90
1.37
1.08
7.05
13.72
2.00
12,696.0
0.817
98.0
587.4
464.9
3,023.6
55.90
7.18
35.39
1.14
0.35
0.04
100.00
Source: Reference 1.
-------
C-6
TYPICAL BUILDING
AND FUEL ELEVATOR .—,
OUTLINE |
WATER JACKE
DISTRIBUTOR
COMBUSTION
ZONE
FUEL BIN
VALVES CLOSED
LOCK HOPPER
WATER SEAL
AMD OUST
COLLECTOR
GASIFICATION
ZONE
FIGURE C-2. WELLMAN-GALUSHA FUEL GAS GENERATOR
-------
C-7
TABLE C-2. TYPICAL WELLMAN-GALUSHA PRODUCER GAS COMPOSITIONS
Mr-Steam Mode
Composition (dry basis) ,
volume percent
CO
H2
co2
CH4
°2
N2
Bituminous
Coal
28.6
15.0
3.4
2.7
0
50.3
Anthracite
Coal
27.1
16.6
5.0
0.5
0
50.8
Coke
29.0
10.0
3.5
0.7
0
56.8
Oyxgen-Steam Mode
Anthracite
Coal
42.0
39.0
17.0
0.7
0.2
1.1
Coke
54.0
31.0
11.3
0.4
0.6
2.7
Source: Telephone conversation with Mr. David D. Woodruff, General Sales Manager,
McDowell-Wellman Engineering Company, Cleveland, Ohio, for oxygen-blown
systems data. Air-blown data are from Reference 22.
-------
C-8
PURGE i INERT GAS I IS-S
\\n_njT
GAS TO OUST
COLLECTOR
WASTE HEAT
WATER COOLED
SHAFT
RATCHET DRIVE
STEAM
STEEL SHELL
REFRACTORY LINING
SC3APER FOS A3.H
REMOVAL
OXYGEN OR
ENRICHED AIR
RATCHET ORIVE
'JATES COOL SHAF7
ASH
RECEIVER
WATER JACKETED
SCREW CONVEYOR
FIGURE C-3. WINKLER GASIFIER
-------
C-9
0-3/8 in. and dried, although moisture contents of up to 18 percent can be
tolerated. Coal is fed to the generator through variable speed screws
which serve as seals to prevent back flow of gases. The gasifying
medium fluidizes the coal bed and gasifies the coal at a uniform
temperature. Oxygen (or air)-steata mixtures are added at the bottom and
top of the bed. The highest temperatures in the system occur at the top
of the fluid bed. Typical gas compositions for oxygen-blown operation
are. shown in Table C-3.
Ash leaves the generator with the product gas (70 percent)
and at the bottom of the fluid bed (30 percent). Particles that leave
as fly ash are sometimes molten since the maximum temperature occurs at
the top of the bed. A radiant waste heat boiler is used at the top of
th«i generator to cool the gas and thus prevent blockage of exit ports by
buildup of molten ash particles.
The Winkler generator can gasify coal tars, heavy fuel oils, and a
variety of coalss although some are more desirable than others, Caking
coals, as a rule, cannot be gasified in fluid-bed systems without
pretreatment to render them noncaking. Processing of coals of low
reactivity requires higher operating temperatures to assure high char
utilization. Coals with high ash content can be processed in the Winkler
generator. In respect to turndown and capacity, a generator with a
nominal capacity of 2 x 10 scfh can be operated in the range of 500,000
to 3 x 10 scfh. Total shutdown can be achieved in minutes simply by
shutdown of all feed streams. The heat content in the generator fuel
bed, is' sufficient to allow startup by reintroduction of oxygen and steam
even after several days. Heat loss during shutdown can be compensated
by injection of air into the fuel bed once or twice daily. On-stream
availability of 91 percent is claimed.
Of over 30 Winkler generators built since 1926, six are
presently operating: two in Kutahya, Turkey, three in Madras, India, and
one in Gorazde, Yugoslavia.
-------
TABLE C-3. TYPICAL WINKLER PRODUCER GAS COMPOSITIONS
Plant Location
Gorazde,
Yugoslavia
Kutahya,
Turkey
Wesseling,
Germany
Toyo-Koatsu,
Japan
Plant Number
Year Built
Product
Current Status
Feed Coal Type
Ash, percent weight
Water, percent weight
Volatiles, percent weight
Gas Composition, volume percent
CO
H2
cd
HHV, dry
Btu/scf
10
1953
Synthesis gas
Operating
Subbituminous
14
1959
Synthesis gas
Operating
Lignite
16
1960
Synthesis gas
Unknown
Lignite
8
1939
Synthesis gas
Unknown
Moderately caking
bituminous
24
3
39
37
37
20
3
3
27
3
37
35
40
20
3
2
4
9
47
48.2
35.3
13.8
1.8
0.9
34
31
12
0.5
22.5
o
i
269
272
287
215
-------
The Riley-Morgan Gasifler
C-ll
(3)
The Riley-Morgan gasifier is a modified version of the Morgan
gasifier. About 3000 units of this gasifier were used in the U.S. before
1940. With the advent of natural gas these were abandoned. Thus, the
Riley-Morgan gasifier can be considered commercially demonstrated.
The Riley gasifier can be either air blown or oxygen blown.
The gas at the exit of the gasifier is at 650 C (1200 F) and a pressure
of +35 in water gauge.
The gasifier is of conventional fixed-bed design with a 10-1/2-
foot diameter chamber. The outer shell of the unit rotates while the
grate on the inside of the unit remains stationary. This rotating action
helps to break up the fuel bed and prevent clinkering tendency of caking
coals. Riley Stoker Company has been operating a 10-1/2 foot commercial
size unit at their Worcester, Massachusetts, facilities since June, 1975;
prior to which a pilot,plant was in operation. Trace elements data
appear to have been collected at their facilities in response to requests
from clients who are vendors of sulfur-producing FGD processes, but Riley
is not willing to publish such data for proprietary reasons.
A detailed flow diagram of the Riley coal gasification system
for medium-Btu gas generation is shown in Figure C-4. Gas composition
for air and oxygen operations with Ohio bituminous coal are presented in
Table C-4.
Reported oxygen consumption in the Riley system is 0.93 Ib per
pound of fixed carbon in the coal. On this basis, for the coal in Table
C-4, the oxygen consumption is 0.4 Ib per Ib of coal. Since this coal
can be considered typical, the oxygen/coal ratio of 0.4 is much lower
than the average value of 0.7 for the other gasifiers discussed. Because
the Riley oxygen requirement is not a published value, but was obtained
during telephone communication with Riley personnel, a closer evaluation
of the oxygen/coal ratio may be necessary at a later stage.
-------
Coal
13,000 Ib/hr
750°F
Gas 22,350 Ib/hr
Tar 920 Ib/hr
Steam 4720 Ib/hr
n
Ash
1070 Ib/hr
Oxygen
6020 Ib/hr'
1
n
Gasifier
fr
130 Ib/l
/
Cyclone
Throttle Valve
65 psia
Sat. steam
10,550 Ib/hr
167 psia
Steam
10,170 Ib/hr
2700 Ib/hr
Cooler
Liquor @ 100°F
30,850 Ib/hr
T = 165°F
Gas 22,350 Ib/hr
Tar 550 Ib/hr
Steam 9430 Ib/hr
Cooling Tower
80°F
Make Up
Water 171,880 Ib/hr
T = 130°F
Electrostatic
Precipitator
Gas
Condenser
3 Pass
T= 110°F
Gas 22,350 Ib/hr
Tar 510 Ib/hr
Steam 1830 Ib/hr
Sulfur
Removal
i
Cleaned
Fuel Gas
33,740 Ib/hr
Water 26,140 Ib/hr
Tar 370 Ib/hr
T = 165°F
Water 171,880 Ib/hr
T = 80°F
Water 7600 Ib/hr
Tar 40 Ib/hr
-Tar Separation
Tank
Tar 410 Ib/hr
Water 50 Ib/hr
Tar Storage
& Mixing Tank
(Heated)
460 Ib/hr
70 Ib/hr H2S
„ Ib S02
0.8 * . Limit
106 Btu
o
i
Tar 507 Ib/hr
Feed
Water
\
Tar By Product
47 Ib/hr
Tar Boiler
50 psig
FIGURE C-4. FLOW DIAGRAM OF THE RILEY-MORGAN INTERMEDIATE-BTU COAL GASIFICATION SYSTEM
-------
C-13
TABLE C-4. TYPICAL ANALYSES OF FUEL GAS FROM RILEY-MORGAN GASIFIER
Air-Steam Oxygen-Steam
Operation Operation
Coal Type Ohio bituminous Ohio bituminous
Coal Assay, percent
Moisture 7.2 7.2
Volatiles 34.4 34.4
Fixed carbon 42.7 42.7
Ash 15.7 15.7
Heating value, Btu/lb 11,315 11,315
Gas Assay
CO
H
CO.
CE2.
CH,+
N 4+ Ar
HHV
Sp. gravity
26.0
18.4
3.8
1.6
0.2
50.0
163.0
0.82
41.2
38.9
15.9
2.8
0.7
0.5
305
0.70
Source: Telephone communications with Thomas F. Walsh, Sales
Manager, Riley Stoker Corp., Worcester, Mass.
-------
C-14
Summary of Coal Gasifier Characteristics
While the K-T gasifier can be operated only on oxygen, the others
can use either air or oxygen. Due to very high temperature operation,
the K-T does not produce ammonia or methane in any significant quantities,
but the gas contains up to 50 percent of the ash in the coal as fly ash.
The K-T has the lowest overall thermal efficiency (67 percent) of the four
gasifiers, whereas the Wellman-Galusha has the highest efficiency (82
percent). The turndown capability of the K-T is lower than that of the
other three gasifiers.
The K-T gasifier, being a suspension type, has the highest rate
of gasification (pounds of coal gasified per hour per unit volume of
reactor). Thus, K-T will offer a smaller sized unit.
Residual Oil Gasification Processes
Two commercially proven heavy oil gasifiers are: (1) the Shell
Gasification process and (2) the Texaco partial oxidation process. More
than 100 Shell gasification reactors with >95 percent on-stream
availability and about 60 Texaco gasification plants are in commercial
use.
Shell Gasification Process ' '
In the Shell process, any hydrocarbon feedstock can be converted
into gas by controlled partial oxidation in a special noncatalytic reactor.
A generalized flow diagram of the process is presented in Figure C-5. The
heavy oil feed, mixed with recycled carbon in a homogenizer, is heated
and fed to the reactor along with oxygen and steam. Air can be used instead
of oxygen. In the reactor, partial combustion of the feed provides the
energy for cracking and gasification at temperatures of 1000 to 1600 C
(1800 to 2900 F). The waste heat boiler producing high pressure steam is
said to be extremely compact due to prevailing high heat transfer rates.
-------
C-15
Reactor
Waste
heat
/^N boiler
High Pressure Sceam
to Turbines
"Clean Gas
Carbon
slurry
separator
Reactor
bn
r
Cooler/scrubber
?Carbon Slurry
1 1>
Pelletizer
Fresh
Water
Boiler
Feed Water
Cond.
f
Steam From Oxygen
Turbines
" Carbon-Free
Circulation
Water
Waste-
water
Homogenizer
To
*"Boiler
Oil
FIGURE C-5. GENERALIZED SCHEMATIC DIAGRAM OF THE SHELL GASIFICATION PROCESS
Source: Reference 7.
-------
C-16
Most of the carbon for recycle to the process is recovered in
the gas purification section. The gas is virtually carbon-free (less
than 4 ppm). Typical gas compositions for air-steam and oxygen-steam
modes of reactor operation are provided in Table C-6.
(8}
Texaco Partial Oxidation Process
The Texaco process, also called the Texaco Synthesis Gas
Generation Process (TSGGP), is used in over 20 countries on a variety
of feedstocks, primarily heavy oils. One of these installations at the
Texaco Refinery in Los Angeles has met all the stringent air and water
pollution control regulations. Eventually, the process is expected to
use coal also as a feedstock.
The TSGGP employs a noncatalytic flame reaction in an oxygen
deficient atmosphere to produce hydrogen and carbon monoxide. Reactor
temperatures range from 1100 to 1600 C (2000 to 2800 F) and pressures
from nearly 1 to over 136 atmospheres. A typical gas composition from a
straight residue feedstock is shown in Table C-7. A general process flow
diagram is provided in Figure C-6.
-------
C-17
TABLE C-6. TYPICAL GAS COMPOSITION FROM SHELL GASIFICATION PROCESS
(Dry Basis)
Feedstock
Oxidizing Medium
Heavy Fuel
Oil
Oxygen
Propane
Asphalt
Oyxgen
Vacuum
Residuum
Oxygen
Air
Composition, volume percent
H2
CO
CO 2
CH,
N« + argon
H2S and COS
46.7
46.2
4.3
0.6
1.4
0.8
44.2
47.8
4.5
0.6
1.3
1.6
47.1
50.8
1.2
0.6
1.2
0.04
14.5
23.6
1.0
0.2
60.7
0.02
(a) Depends on feed sulfur content.
Sources: (1) Hydrocarbon Processing, April, 1975, p 117.
(2) Reference 6.
-------
C-18
TABLE C-7. TYPICAL SYNTHESIS GAS COMPOSITION FROM TEXACO GASIFIER
Feed
Oxidizing Medium
Product Gas Composition.
Straight petroleum residue
Oxygen/steam
volume percent
CO
H2
co2
N2 + A
CH
H S(a)
H2S
47.0
47.0
5.6
0.2
0.1
0.1
(a) Proportional to feed sulfur content.
Source: Telephone communications with Dr. Edward T. Childs,'
Texaco Development Corporation, New York, New York,
(212) 953-6000.
-------
C-19
Feed
O)
a.
a.
Steam
Cleaned gas
Oil and soot
-*/Separa»orj
f
Naphtha and soot
^ Steam
Water and soot
Water
FIGURE C-6. SCHEMATIC DIAGRAM OF THE TEXACO GASIFICATION PROCESS
Source: Hydrocarbon Processing, April, 1975, p 133.
-------
C-20
REFERENCES
(1) Mudge, L. K., et al, "The Gasification of Coal", A Battelle Energy
Program Report, Battelle Memorial Institute, Columbus, Ohio,
July, 1974.
(2) Mitsak, D. M., "Koppers-Totzek: Take a Long Hard Look", 2nd Annual
Symposium on Coal Gasification, Liquefaction, and Utilization: Best
Prospects for Commercialization, University of Pittsburgh,
August 5-7, 1975.
(3) Cobb, R. W., et al, "Analysis of a Coal Gasification Facility and
Potential Gas Using Industries for Pike County, Kentucky", BCL
report to Appalachian Regional Commission, Commonwealth of Kentucky,
and Pike County Fiscal Court, September 1, 1975.
(4) McDowell Wellman Engineering Co., "Wellman-Galusha Gas Producers",
brochure (Form No. 576).
(5) Banchik, I. N., "The Winkler Process for the Production of Low-Btu
Gas from Coal", Day Powergas, Inc.
(6) Plummer, J. B., et al, "The Generation of Clean Gaseous Fuels from
Petroleum Residues", presented by Shell Development Company, Houston,
Texas, at the AIChE meeting, Tulsa, Oklahoma, March 11-13, 1974.
(7) Shell Development Company, "The Shell Gasification Process" brochure.
(8) Child, E. T., "Texaco: Heavy Oil Gasification", Symposium on Coal
Gasification and Liquefaction: Best Prospects for Commercialization,
University of Pittsburgh, August 6-8, 1974.
-------
APPENDIX D
CALCULATION OF TRACE CONSTITUENT DISTRIBUTIONS
-------
D-2
APPENDIX D
CALCULATION OF TEACE CONSTITUENT DISTRIBUTIONS
This appendix presents additional details related to the computer
calculation of trace constituent distributions for gasification processes.
The first section includes a comparison of the computer program used in
this study with other similar programs. The second section presents the
detailed computer output data, i.e., the predicted equilibrium composi-
tions of the gaseous and solid phases.
Comparison of Equilibrium Calculation Computer Programs
The technique employed to compute thermodynamic equilibrium
values in this study was based on the EQUICA program developed by Battelle
based on the work of Cruise. Other programs available for computing
equilibrium calculations include the NASA program, which utilizes the
free energy minimization method, and the Naval Ordinance Test Station
C2)
(NOTS) program. In a recent Battelle study for the U.S. Coast Guard,
these programs were evaluated for relative accuracy. A brief summary
of the results of the study is presented in Tables D-l and D-2. Accordingly,
the difference between estimating accuracy of NASA and EQUICA is negligible.
However, the NASA program uses more automatic computing processes and
therefore can be less expensive in computing time cost when very large
calculations are involved.
Computer Output Data on Trace Constituents
This section contains the computer printouts giving the detailed
trace constituent distributions. The following tables are included:
-------
D-3
TABLE D-l. COMPARISON OF EQUILIBRIUM COMPOSITIONS CALCULATED
(Thermal Decomposition of Acrylonitrite at
1300 K and 1 atm Pressure)
Species
HCN
NH3
CCS)
H
2
2
CH
•T
H
C2H2
A» 4-
CH.
•(
J
EQUICA
-5
3.459 x 10
-5
1.746 x 10
6.004 x 10"1
_i
2.970 x 10
1.004 x 10"1
2.121 x 10~3
_7
4.503 x 10
1.789 x 10~7
-8
4.831 x 10
2.699 x 10~10
NASA
-5
3.5258 x 10
-5
1.6047 x 10
6.0034 x 10"1
-1
2.9756 x 10
1.0032 x 10"1
1.7257 x 10~3
-7
3.7969 x 10
2.0284 x 10~7
-8
5.1507 x 10
3.0380 x 10~10
Source: Reference 3.
TABLE D-2. COMPARISON OF DECOMPOSITION TEMPERATURES CALCULATED
BY PROGRAMS NOTS AND NASA
Nitroethane
Epichlorohydrin
Acrylonitrile
Nitrobenzene
Acetic acid
Butyraldehyde
Propylene oxide
Acetylene + 0
Decomposition
NOTS
1161
1042
1896
1540
652
800
951
3341 (To)
Temperature, K
NASA
1161
1042
1891
1538
655
• 801
951
3341 (To)
-------
D-4
Table Temperature,
Number System C (F)
D-3 Power plant boiler 650 (1200)
I>-4 Koppers-Totzek gasifier 1815 (3300)
D-5 Koppers-Totzek gasifier 650 (1200)
D-6 Winkler gasifier 980 (1800)
D-7 Wellman-Galusha gasifier 650 (1200)
D-8 Riley-Morgan gasifier 680 (1250)
-------
TABLE D-3". TRACE EJLEMENTS DISTRIBUTION IN BOILER AT 1200 F
t 96 PERCENT tJY WEIGHT OF THE FOLLOWING CONDENSED SPEOIES ARE
SPECIES
SI03
CAO
FE304
M&O
K20
CACL2
TI02
NA20J3
A 1333
0203
ZNO
UNO
ZR02
NIO
V204
CUO
POO
TIN
TI
MN
C
SI
7. KM
ZR
K
HI
BEO
CAS
CA
F£
AL
ALN
V/N
V
CR203
NAOH
CU
an
B
CR
OE
CRCL2
CDCL2
PB
BE3M2
NA
PBCL2
SB203
IN
MOL. MT.
60.1
56.1
231.6
1*0.3
9-».2
111. 0
70.9
10-i.O
101.9
6-J. fa
ai . t*
70.9
123.2
7,. 7
165.9
79.5
233.2
61.9
47.9
5-».9
12.0
2b.l
105.2
91.2
39.1
5 i . 7
25. 0
72.2
40.1
55.9
27.0
41. 0
65.0
51.0
152.0
4J.O
6?. 5
24.8
10.8
52.0
9.0
122.9
183.3
207.2
55.1
23.0
278.1
291.5
THE PRODUCT GAS STREAM AS FLY
MOLES
i^jiiT'jMj/
.75779E-02
,14730£-02
.97950E-03
.17578E-03
.17391E-03
.16797E-03 .
.12299£-03
.925B1E-04
.43264E-04
.39751E-04
. 355tt'JE-04
.672b3t:-05
.50938E-05
.2.101JE-05
. 20584t-05
.14025E-05
. 11473t-05
.97&10t-0b
.976G9S-06
. y6095E-06
.g^t+J-OE-Ob
.92515E-0&
. H489iE-06
.8t693E-06
.7930dE-0&
.7670d£-06
.761+10E-0&
.76080E-06
.7o080E-06
.7l7';8il-06
.b3672£-06
.68672S-06
.63616c:-06
.6d606£-06
.67172£-06
.65258E-06
.63750E-06
,63097£-06
.b3097£-06
.56W33E-06
. 38195E-06
.35193E-06
,189'*7t-06
.17939E-0&
.127H2E-06
.12195£-06
.10869£-06
.332Qiȣ-07
GRAMS
(ASnTGAS)
.(«55<«3E+00
.62633E-01
.22685E+00
.73839E-02
.1&382E-01
. IBb'iSE-Ol
.98269E-02
.98136E-02
,'4'»086E-02
.27667E-02
.28969E-02
. «.7693£-03
.b2756E-03
. 1 7 1 9 1 E - 0 3
. 3'*1'<9E-03
.11150E-03
.25608E-03
. 60U21E-04
.H6755E-Qi»
. t2756E-0<«
. 11331.E-01*
.25997E-Qi«
.agsiSE-O'*
. 77>t22E-0<*
. 310 lOE-O't
. '.5Q28E-0'*
. 19102E-OV
,5i«930E-Oi«
. 30508E-0'*
,<*0113E-0'*
. ISS'^IE-O'*
.2815bE-0'*
. 'iHbOOE-Oi*
. 3'»989E-0't
. 10210E-03
. 26103E-0<*
. 'tOit81E-0'*
. ISb'tBE-O'*
.681i«5E-05
.293<*5E-0(»
.3U376E-05
. i*3252E-0<*
.3^5i*7E-0't
.37170E-0<*
.70208E-05
.280i«9E-05
.30227E-0"*
.56790E-05
ASH? THE BALANCE IS
-------
TABLE D-3. (Continued)
SB
ZN
COO
CK3C2
CAC2
SIC
vc
TIC
NA20
H&
MG
CD
MGCL2
SPECIES
H2
co2
H20
0?
S02
p<*oi3
K
AS203
SE02
CL2
KO
HGO
HG
HGCL2
sto
SE
CO
PB
H2
t!02
Ft (OH)2
. NA
3ECL2
AS
ALCL3
Ticm
NH3
00
ZRCLi.
S
H2S
COS
CA(OH>2
F£
ALCH
l?1.8
65. <•
12 1.<*
183.1
6-..1
(»0.1
292£-07
% 17603=:-07
.3<*2ll£-08
.i*0000£-09
.20000E-09
.20000E-09
,2)000£-09
.200 JOE-09
. 10000E-09
.10000E-09
.39990E-10
.623616-11
.0
MOLES
(ASH+GAS)'
.273XaE*01
,50022ii + 00
.22759Z+00
. 117«i7t*00
n&^99r-oi
.08
. 18183C-10
.6998UE-10
. 2H&38C-10
. 2378HE-11
.E1250E-11
.59337E-12
. 69921E-16
. 1!»-«17E-16
,177«t9E-18
. 19723E-21
.1863lli-21
.70024ȣ-24
. V+922E-2i»
.38125E-25
"""" .53356E-26
.I.3191E-27
.21183E-31
.0000021
".0000011
.000000'*
.0000001
.0000000
.0000000 .
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
SPECIES
. PERCENT
IN" GAS)'
.701,',05'j
.302^392
.77885J7"
.«.7029tt9
';9755510
.001281*7
.00001*08
.00011*60
.0000231
.0000098
.0000060
.0000012
.OOQOOQi*
.0000001
.0000000
.0000000
."0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000003"
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
. OTTOTODir-
'.021
.011
.OOi*
.001
.000
"" .000
.000
.00-3
.000
.000
.000
.000
0.003
WERE PREDICTED)
PPM ay WT.
(IN GAS)
7070'»H.O»5
20302^.392
37786.537
31,702.939
9755. btfl
12.8i«7
.<*08
1 .i«60
.231
.093
.060
.012
.OOi*
.001
.000
.000
.000
.003
.000
.000
.000
.003
.000
.000
.000
.000
.000
/.ooo
.003
.000
.000
.000
.000
.000
.000
.022
.en
. 00«»
.001
.000
.000
.000
.GOO
. 000
. 000
.000
.000
0.000
TOTAL
PPM BY WT.
' (ASH+GAS)
706989.285
203008.665
37765.610
31*700.300
9751*. 75^
12.6<«6
' .'.oa " "
l.^bO
.231
. 098
.060
.012
. 00<*
.001
.000
. 000
.000
.000
.000
.000
.000
. 000
.000
. 000
.000
.000
.000
.coo
.000
.000
.000
.000
.000
.000
.000
2.798
1.371
.523
.087
.015
.010 """ "" ""
.015
. Oil*
.007
.021*
. 001
.001
0.000
1000000. 00
7
cr-
-------
TABLE. Dr3. .(Qontinued)
.11.0 .658U9£-i*0 .20<»13E-38 .0000000 .000 .000
CH»» lo.O .SgoaSc-ftS . 158'*'*E-'«3 .0000000 .000
CS2 7o.l .26998£-5't «205'«fcE-52 .0000000 .000
SI 21.1 .60100E-57 .1&888E-55 .0000000 .000
""""""*" " *»••
.000
.000
.000
TOTALS """ --.--- .lOS^lE+03 100.000000 1000000.00 1000000.00
-------
TABLE D-4. TRACE ELEMENT DISTRIBUTION IN K-T COAL GASIFIER AT 3300 F
SPECIES
sio2
FE
CAS
TI02
MNO
ZR02
NI
V20i»
TI
AL
NA2C03
CAO
SIC
ZRN
ZR
B203
ALN
CAC2
TIC
CACL2
C
MGO
B
VC
VN
SI
F£30i*
TIN
3N
NIO
AL2G3
MN
V
3EO
CR203
CR3C2
BE3N2
K20
NA
AS203
ZNO
NA20
CUO
K
HG
MOL. HT.
60.1
55.9
72.2
79.9
70.9
123.2
5d.7
165.9
i»7.9
27.0
106. 0
56.1
<*0.1
105.2
91.2
69. 6
1*1. 0
6i*. 1
59.9
111.0
12. 0
t»0.3
10.8
63.0
65.0
2d. 1
231.6
61.9
2^.8
7<*.7
101.9
5 + .9
51.0
25.0
152.0
183. 1
55.1
9^.2
23.0
197 . 8
81. it
62 . 0
79.5
39.1
200. 6
IN THE PRODUCT GAS
MOLES
(ftSHtGAS)
.7575^E-02
.21657E-02
.16397E-02
.12268E-03
.71i*72E-05
.50938E-05
.2i»931E-05
.18761E-05
.97609E-06
.95118E-05
,9'«562E-06
.87V37E-06
.86732E-06
.8t897£-06
.84893E-06
.7927i»E-06
.77777E-06
.76671E-06
.7Q157E-06
.67502E-06
.65987E-06
.6<+719£-06
.603i*5E-0&
.60075E-06
.600^0E-06
.59990E-06
.59158E-06
.58718E-06
.57773E-06
.57533E-06
. 5i*802£-06
,ii»05:*£-06
.53602E-06
.36297E-06
.77625E-07
.^7193E-07
.31935E-07
. 10000E-09
. 100QOE-09
. 10000E-09
.10000E-09
.10000E-09
.10000E-09
. 10000E-09
. 10QOOE-09
STREAM AS FLY
3RAHS
(ASH+GAS)
.i*553i»E+00
.12106E+00
. 11839E+OQ
.9U018E-02
.50673E-03
.62755E-03
. li+63t,E-03
. 3112^^-03
,'*o755E-0'*
.25682E-0'*
.1002'«E-03
.if9063E-0'*
.3i«780E-0'»
.69311E-0'*
.77i»22E-Qi»
.55175E-0^
. 3ia88E-0
. 7i»927E-0't
.79185E-05
. 2S082E-0'*
.65173E-05
. 378i<7E-Q<+
. 39026E-01*
. 16857E-01*
. 13701E-03
. 35347E-0'*
. U329E-0'*
. "42977E-04
. 55S'i3E-0'*
. 29676E-01*
.27337£-0't
. 907'«2E-05
. U799E-0'*
.65i«19E-05
. 17596E-05
.9U200E-03
.230 OOE- 08
. 19780E-07
. 81i*OOE-08
.62000E-08
. 795 OOE-Od
.3910CE-08
.20060E-07
ASH? THE BALANCE IS ^EfOi/E
HT. PERCENT
(IN GAS)
i.ii&16<»73
.38861iiO
.3800256
. 031^*635
.0016266
,00201<«i»
.0001*698
.0009991
.0001501
.0000821*
.0003218
.0001575
.0001116
.0002867
.00021*85
.0001771
.0001021*
. 0001578
.000131*9
.00021*05
.0000251*
.0000837
.0000209
.0001215
.0001253
.000051*1
.0001*398
.0001167
.0000'-* SO
.0001330
.0001793
.0000953
.0000873
.0000291
.0000379
.0000277
.0000056
.0000000
.OOQGOQO
.0000001
.0300000
.0000000
.0000000
.0000000
.0000001
PPM BY HT.
(IN GAS)
1<»&16.!*73
3386.11*0
3800.255
31*. 635
16 .255
20.11*!*
•*.693
9.991
1 .501
.82i»
3.218
1.575
1.116
2.857
2.1*85
1.771
1.021*
1.578
1.31*9
2. <*05
.251*
.837
.239
1.215
1.253
.51*1
4.393
1.157
.1*50
1.380
1.793
.953
.878
.291
.379
.277
.055
.000
.00)
.001
.000
.003
.003
.000
.031
D AS ASH)
PPM BY WT.
(ASH+GAS)
28583.901
7599.716
71*31.761
615.299
31.810
39.391*
9.187
19.538
2.935
1.612
6.292
3. 080
2.183
5.606
i».860
3.'*6'*
2.002
3. 085
2.638
1».703
.<*97
1.637
.1*09
2.376
2.<»50
1.058
8.601
2.282
.899
2.698
3.505
1.863
1.716
.570
.7<«1
.5<*2
.110
.001
. 000
.001
.001
.000
. 000
.000
.001
TOTAL
PPH BY WT.
(IN ASH)
61*3708.682
17111* 5. 4,06
167363.057
13856.520
716.356
887. 158
206. 882
1*39.990
66. 096
36.306
li»l. 700
69.359
1*9. 167
126.257
109.1*1*9
77.999
1*5.080
69.«*77
5 9. i« 0 9
105.922 |
11.191* oo
36. 871
9.213
53.501*
55. 170
23.831
193.687
51.382
20.256
60. 756
78. 9<*i*
1*1.952
38. 6i*6
12.828
16.680
12.217
2.1*88
.013
.003
. 028
.012
. 009
.011
.006
.028
1000000.00
-------
T^ABLE D-4. (Continued)
(THE FOLLOWING GASEOUS SPECIES WERE PREDICT-0)
SPECIES
CO
H2
H2S
H20
N2
C02
FE
COS
K
NA
S
MG
ALCL3
BO
MGCL2
ZN
P
CS2 •
CU
CR
B02
AS
NAOH
BECL2
SI
PB
PB
S02
CA
SE
CD
FE(OH)2
NH3
CHi*
SB
OE
ALOH
HG
CRCL2
CL2
KO
PBO
02
PBCL2
SEO
ZRCL<*
TICLt,
CA(OH»2
Pi*010
DQCL2
HGO
MOL. HT.
23.0
2.0
3i*. 1
18.0
2B.O
1*1*. 0
55.9
60. 1
39.1
23.0
32.1
24.3
133.3
26.6
95. 2
65.1*
31.0
76. 1
63.5
52.0
1*2.8
7i+. 9
(.0.0
79. 9
28.1
207.2
207,2
6^.1
40.1
79.0
112.1,
33.9
17. 0
16.0
121.8
9.0
i*<*. 0
200.6
122.9
70.9
55.1
223.2
32.0
273.1
95.0
233.1
189.7
7*. 1
283.9
183.3
216.6
MOLES
(ASH+GAS)
,i*9770E + 00
.316<«i*E + 00
.65285E-02
.60358E-02
,396i*7E-02
.22367E-02
.77157E-03
.35205E-03
,3i»986E-03
.13301S-03
.17711E-03
.13563E-03
.8505iE-0<»
.76150E-Q1*
.39506E-OI*
.35607E-01*
.19622E-Q1*
.16613E-01*
.2Q399E-05
.19566E-05
. ,18i»82E-05
.16001E-05
.10430E-05
.997Q4E-06
.87363E-06
.71755E-06
.71755E-06
.71689E-06
.<»519QE-06
.22536E-06
.19190E-06
.1353!»E-06
,127i,5£-06
,861<»6E-07
.857QOE-07
.72<*8^£-07
.19577E-07
.86997E-08
.7451.3E-08
.16773E-08
.55<*15E-09
.22779E-09
.98931E-10
.73661E-10
.'«'»lfa6E-10
,206'43E-10
. 206U8E-10
.712i«3E-ll
.<.1061£-11
.17706E-11
.31585E-12
GRAMS
(ASHtGASl
.13936E+02
,63289E*00
.22262E+00
.103B4E+00
.11101E+00
.98413E-01
,i,3131E-01
.21158E-01
,13b80E-01
.953E-0'«
. 18121E-0'*
. 17803E-0"*
.21569E-0'*
. 12167E-0'*
,21o66E-05
. 13783E-05
,10<«3BE-0<»
. 65236E-Oo
.85137E-06
.17<»52E-05
.91S13E-06
.11892E-06
.3053*tE-07
.508(»3E-07
.31658E-08
,20'»85E-07
.^1977E-08
. 1.3129E-08
.391b8E-08
.5Z791E-09
.llb57E-08
,32i,5i*E-09
.6B<«1<»E-10
HT. PERCENT
(IN GAS)
&9.<*662801
<«. 0631276
l.*»292163
.697<*968
.7126922
.6318097
.2768987
.13583<«3
.0878220
.0270225
.036H981
.0211585
.0727893
.0131019
.02
9.5^5
2.950
1.163
1.11*3
1.385
.731
.133
.083
.670
.01*2
.055
.112
.059
.003
.002
.003
.003
.001
.000
.000
.003
.003
.000
.000
.000
PPM BY HT.
(ASH+GAS)
871*799.038
39729.160
13971*. 81,1
6820.106
6968.686
6177.819
2707.509
1328.185
858.721
261*. 226
356.878
206.887
711.732
128.110
236.092
1U&.181*
38.185
79.359
8.131
6.387
<». 966
7.523
2.619 T
5.001 «>
1.550
9.333
9.333
2.885
1.138
1.118
1.35V
.76<*
.136
.067
.655
.Gi*l
.051*
.110
.058
.00?
.002
.003
.000
.001
.000
.000
.000
.000
.000
.000
.000
-------
TABLE D-4. (Continued)
COO
HGCC2
SE02
SB203
128.4
271.5
111. 0
291.5
.16873E-13
.15277E-15
.i6<*87E-i&
.52085E-25
.21665E-11
,<»i<«7eE-13
. 18301E-14
. 15183E-22
.0000000
.0000000
.0000000
.0000000
.000
.000 •
.000
.000
.000
.000
.000
.000
TOTALS
0.817305
. 15576E. + 02
100.000000
1000000.00
1000000.00
Y
M
— 0
-------
TABLE D-5. TRACE-ELEMENT DISTRIBUTION IN K-T COAL GASIFIER AT 1200 F
SPECIES HOL.
NA2C03
2ND
PBCL2
NA20
NA
NAOH
K
ZN
K20
COCL2
AL20J
T102
ZR02
TIC
TIN
QEO
PO
FE
MGCL2
CR203
VC
FE30i+
SI02
NI
MNO
CACL2
B203
ON
CAS
SIC
V
ALN
CU
C
TI
CA
PBO
SB
CR
CRCL2
MGO
ZR
ZRN
AL
VN
S6203
BE3N2
dE
(
IN
wr.
106.0
81 .1*
275.1
62.0
83.0
i»0.0
39.1
65 . U
9<*.2
1M.3
10 1.9
7'J.9
123.2
59.9
61.9
25.0
207.2
55.9
95.2
152.0
63.0
231.6
•JO.l
5fl.7
70.9
111.0
69.6
2k. 6
72.2
1*0.1
51.0
i*1.0
63.5
12.0
<«7.9
<«0.1
223.2
121.8
52. 0
122.9
UJ.3
91.2
105.2
?7.0
65.0
291.5
55.1
9.0
0 PERCENT UY WEIGHT OF THE FOLLOWING COUOENSEO SPECIES ARE
THE PRODUCT GAS STREAM A3 FLY ASH? TH£ BALANCE IS REMDVEO
MOLES
IASH+GAS)
.90279E-Ot»
. 33051£-Qi»
.10757E-05
.716n9£-06
.716<«ti£-06
.70161--06 •
.68332E-06
. 55636S-06
.512A9E-06
.19162E-06
.32928E-08
.ai^E-oa
.16222E-08
. d73tj'3E-09
.t»36BOE-D9
.i*36BOE-09
.4UOOOE-09
,258b3£-09
.20000E-09
.17500E-09
.17500t-09
. 166D4E-09
. 1'5000£:-09
.1500QE-09
.150005-09
.15000E-09
. i^oae-og
.12500S-09
. 12500E-09
.125UOE-09
.10000E-09
,10000£-09
.10000E-09
.10300E-09
.10000E-39
.10000E-09
.lOOOOt-09
.10000t-09
. 10000£-09
.10000£-09
.10000Z-09
.10000£-09
.10000£-09
.10003E-09
.10000E-09
.10000£-09
.10000E-09
. 10000E-09
SRAMS WT. PERCENT PPM 3Y HT.
(A3M+&AS)
-------
TABLE D-5. (Continued)
cuo
CACZ
HG
CO
NIO
MN
CR3C2
SI
CAO
Q
CDO
MG
SPECIES
CO
H2
CH<*
C02
H2S
N2
COS
K - - -
ALCL3
TICLt*
NH3
P"*013
CS2
AS
3ECL2
SE
MA
HG
" CL2 """ "
PB
HGCL2
FE(OH)2
S02
P
S
FE
KO
SEO
80
•302
HGO
ALOH
AS203
79.5
6<*.l
20). 6
165.9
7t.7
5 -.9
23. 1
56. 1
10.3
12S.i*
2\. 3
MOL. WT.
2<1.0
2.0
""16.0 '
i»<» . 0
li . 0
3«. 1
23.0
50.1
------ 39.1
133.3
139.7
lf.0
235.9
76. 1
71*. 9
79.9
Z3J. 1
79.0
23.0
203.6
" ' 70.9
207.2
271.5
89.9
6«*. 1
31. 0
32.1
55.9
55. 1
95.0
2b. 8
1*2.8
216.6
Hi*. 0
197. 8
. 10000E-09
.10000E-09
. 10000E-09
.8-* 199E-10
.62533£-10
.-;oooo£-io
.50300£-10
" ' .53303E-10
.25000£-10
.25000£-10
.0
1
MOL£S
(ASH*-GAS5
. 16HU9E-I-00
* .70756t-01
.b2309£-01
.16217E-01
.65822E-02
. 39620E-02
.55i*10E-03
. 3 ^+ 8 15 £~0 3
. 369 27 C — 0 1*
.21+215C-Q1*
. 50360E-05
.16539E-05
"" .1&002£-05~
*.225i.OE-!!l
.«68i*6£-08
" .i*'-30biE-09
. .31582E-09
iiw^lS
. 1JS5COE-11
.62307E-12
.11571E-12
.27021E-15
. 12621* E>16
.71353E-18
.62356E-18
.ii»326E-22
.10332E-22
.79500E-08
. 6m OOE-08
.200 t-OE-07
. 9'toi»0t>0fl
.10369E-07
. 3/350E-08
. 91550E-08
.70250E-09
. lt*025E-08
. 51963E-10
.52012E-11
. 0
ITHS FOLLOWING (
3RAMS
(ASH*GAS)
.32937E*00
. 11321E+01
.29190E+00
" ~ . 11093E*00
. 33301E-01
.13613E-01
. l*5936E-02
.85612E-01*
. 13927E-02
• .12586E.-03
. 11985E-03
. 79637L-01*
. lof 6/C.-U3
. 17807E-01*
. 286i*7E-06
. 17i*21E-05
" ~" . 3i*78i*E-07
.12509E-08
. 8S793E-10
. 19315E-10
.37142E-11
. 151G«»E-13
.790 Odc-li*
. 11992E-1H
. 19122E-16
. 26688E-16
. 63036E-21
. 20 <* 3 YE~ £Q
0.0000000
0.0000000
0.0000000
0.0000000
0.0000000
0.0000000
0.0000000
0.0000000
0.0000000
0.0000000
0 .0000000
0.0000000
0.0000000
JASEOUS SPECIES
WT. PERCENT
JIN GAS)
67.7381*223
2.1735592
7.1*7081*61
18.0920852
1.9262936
1.1*81191'*
.7320725
.2197607
.0898322
. 032^830
.030313'*
,0005650
.0091903
.0008306
.0007909
.0005255
.0011065
.0001175
~ " .0000019 '
.0000115
.0000002
.OOOQOO<*
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
0.000
0.000
0.000
0.000
0 .000
0.000
0.000
0 .030
0.000
0.000
0.000
0.000
3.030
HERE PREDICTED)
PPM OY WT.
(IN CJAS)
677381*. 223
21735.592
7i*70(J.i*61
180920 .852
19262.936
11*811 .911*
" 7320 .725
2197.607
898.322
32:*. 830
303.131*
5 .650
91.903
3 .30i
7.909
5.255
11.065
1.175
.019
.115
" ".002" "
.OQi*
.000
.000
.000
.003
'.000
/. 000
.033
.003
.000
.000
.000
.000
.000
.001
.000
.001
.001
. 001
.000
.000
.001
.000
.000
.000
.000
0.000
TOTAL
PPM OY WT.
JASh+GAS)
676805.1,38
21717.020
71,61,'*. 627
180766.265
' 192'*6.<*77
11*799.258
7311*. i*70
2195.729
897.551* "
321*. 552
302.875
5.61*5
91.825
8.299
7.903
5. 251
11.055
1.171*
.019 """
.115
.002
.00<*
.000
.000
.000
.000
.000
. 000
.000
.000
.000
.000
.000
.000
.000
.613
" .1.95
.730
. 800
.288
.212
.706
.108
.001*
.000
0.000
1000000.00
O
ro
-------
02
-$TS2
SI
CA
-------
TABLE D-6. TRACE ELEMENT DISTRIBUTION IN WINKLER GASIFIER AT 1800 F
( 09 PEHCENr HY WEIGHT OF THE FOLLOWING CONQENSEn SPEClEi AHE INCLUDED
1<\ THE PHOIlUCT GAS STHEAM AS FLY A$HJ THE BALANCr IS HtMUVtO AS ASH)
sprcits " "0|_. tai." MOLES «HAMS wr. PERCENT PPM HY wi. PPM BY WT.
(ASH+GAS) (ASH»«AS) (IN GAS) (IN GAS) (ASH»GAS)
c
Ft
TI"2
AL'03
7^2
vc
M
CO
. . /NO . _.
TIN
TI
SI
N A "? 0
NAOH
Cun _.
C*
fiL
r
CAO
PN
e
V204
MC
CH7U3
CRrL2
K20U*
HE
PEO
PHO
PH
RE1N2
SB?03
12.0
S-^
72.2
40.3
111.0
101 .9
6S.6
70.9
A3.0
«1 .4
Al.9
47.9
91,2
*t .0
40 .0
40.1
41.0
ll\(\
M!O
?H,8
10.8
1*5.9
7<..7
1^2.0
^2.0
211 .6
s.o
?b.O
2?3.2
217.2
S5.1
2^1.5
, 17144E + 00
.75779E-02
.29375K-02
. 17S09K-03
. 1661 7t-03
.43223E-04
.397SIE-U4
.672ftHt-05
.50935h-i)5
,29l(S Jh-05
. 12 75DE-U5
.97610t-06
.9609SE-06
.92S15K-06
,H4li'jnt-06
. .H4M93E-06 ...
.77306E-06
.760POL-06
-Ocjtiljt""'Jb
• Ocf^nht.**0b
.66078E-06
.63097E-06
.63D97E-06
.575T3E-06
.5650 JE-06
.56493E-U6
.5606U-06
.S40?7t-06
,3Bl9bE-06
.17939E-06
,12742t-U6
.16081E-07
,20573E*01
!u*?u!oo
:5!!sMtiJS
(JTQlZAt^QP
^fifO/fAK^O?
.27667E-0?
.6^75?t-03
.183 73E-03
.8n962tl-04
.81 06SL-04
l^bStlSt
. 77422E-04
.50H27E-04
.8S44«St-04
.30922E-04
,60818^-04
,3(>50et-04
. l668st-04
.349H9E-04
.3707ot:-04
. 1S648E-04
.429 /7E-04
.85907E-04
.293/6E-04
.1?984E-03
. J4376t-05
,37l69t£-04
.6S209E-05
.46877E-05
11.5678064
2.5607651
.9232980
.5987585
.0396756
,1037087
.0550745
.0247650
.0155564
.003s?fl4
.0010331
.U008229
.0004552
.0003397
,000?629
.0002966
.0001462
,0005022
.0004353
,0002«58
.0004804
.0001739
,0003420
.0001715
,0001583
,0001043
.0000938
.0002506
.0001967
,0002084
.0000880
.0000383
.OOOS600
.0002417
.0004830
.0003905
.0001652
.0007300
.0002862
.0000193
.0000533
.0002252
.0002090
.0000395
.0000367
.0000264
1 15678.064
25607. 6bl
9232.908
5987.505
396. 7b6
1037.007
550.745
247.650
155.564
26.816
10.3J1
8.22
5 «0
-------
TABLE'
'(Continued)
*" " N
Sl<"
Tir
K
no
MGTU2
SECIES
~2
cc
K2
cc*
K
CC«:
p
Nh-J
cs?
^ s
p p ^ 1_ 2
P8
SE
CC
ALTL3
P4D10
CCrL2
HG
S
FE
8C
SC?
PC?
CDO
KC
ALOM
SEO
HGTU2
(•GO
C2
A4.1
4U .1
?3.0
C '" i» * 6
95.2
"°L* W"
2.0
?Jl?
?8.0
18.0
16.0
44.0
?l!o
£b!i
11 .0
17.0
76.1
27S.1
IPS. 7
2 0 / . f.
79.0
112.4
133,3
2"3.9
1P3.3
2^3,1
fl-i.9
76. a
70.9
42. fl
'•S.I
44.0
95.0
271.5
2U.6
32.0
.200nOE-09
.20000E-09
.10000t-09
.IOOOOE-09
. 1 0 U ii 0 c - u '•»
.0
" (THE
MOLES
(ASH+GAS)
,41947E*00
:?5*JbtUS
,396?OE~02
..)2402E""2
ifl^Elol
.34878E-03
.14899E-03
,346126-04
.19blbE-04
.4S44.3t.-05
.1600 JE-05
.72692E-06
,397fllE-06
,3Hb]OE-n6
.22540E-06
.18U6t-06
,80 72^f->07
.26786E-07
.10236E-07
.870nOE-D8
.1402SE-08
iJIs^E-lO
.22940E-10
.11418E-10
.68B36E-11
.136R9E-12
.2639JE-13
.45HSHE-14
.264S5E-15
.48433E-19
.1^^20E-07
.3662ot"07
.Hi'icfont-'Ofl
. 1 1 98flt»07
. 3^1 OOt-08
• 2.*r)OOE~0 A
• 2fi060t-07
.0 0
FOLiOwlNii GASEOUS
GHflMS WT
(ASH*GAS) (
.83893E+00 6
,9(1939E»01 73
.11094E+00
,5u323E-01
;^SO«E!O!
.13637E-01
.41H4U-02
. 604^7E-03
. U4053E-04
.30201E-03
.20216E-03
,3076qt-04
,l7807t-04
.2041 QE-04
.l()761t-04
.18762E-OS
,1745?E-05
.450l9t-07
.68122E-07
.11328E-07
.6J47HE-09
.44124E-09
.17834E-10
.14543E-11
.20177E-12
.2V718E-12
.71826E-13
.3S235E-14
.15499E-17
,0000001
.0000002
,0000000
.0000001
,0000000
.0000000
5 0 0 u i) 0 0 i
.0000000
SPECIES
. PEHCERTT
IN (>AS)
.7367362
.0470787
.9445963
.4604821
,3601944
,3976707
,1095414
,0336084
,0719252
,0181823
.0048594
.0006752
,0024259
,0009628
,0016238
.0006062
,0002472
.0005820
,0001430
.0001640
.OOOOB64
.0000611
.0000151
,0000140
,0000004
jOOOQOOS
.0000001
,0000000
,0000000
,0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.oouoooo
.0000000
.0000000
.0000000
,001
.002
,000
,001
.000
.000
. () U 1
0.000
WE«E PREDICTED)
pPM BY wT.
(IN GAS)
67387.362
730470,767
19445.903
8910. 8V6
4684,821
3601 .944
3976.707
1095,414
336. Ob4
719.232
181,823
46.5*4
6.7D2
24,2S9
16.238
6.062
2.4 12
5.820
1.4JO
1 .640
,864
.611
.140
.004
.005
.OUl
.000
.000
.000
.QUO
.000
.000
.000
.QUO
.000
.000
.000
• QUO
.001
.003
.001
.001
.000
.000
.002
0,000
TOTAL
PPM BY #T.
(ASM*GAS>
63086,715
683852.299
18204.926
8342,205
. 4385, U36 ...
3372,069
3722.914
1025.505
314.636
673.350
_ . 170,219
&!321
22,710
9.013
15,202
5,675
2.314
5.448
1,339
1.535
.809
.572
' !l31
,003
.005
.001
.000
.000
.000
,000
.000
.000
.000
.000
.000
.000
.000
.000
,005
.013
,003
.004
,001
.001
.007
0.000
1000000.00
0
u.
- -
•
-------
TABLE D-6. (Continued)
SI ?".! .SJ4T9E-20
CMOH)? 74.1 ,l>jS)?6t-23
4S'03 1J2t-^2 ,0000000
.i^'t^gE+oa loo.ouoooo
.000
.000
.000
,000
1000000.00
.000
.000
.000
.000
1000000.00
I
M
-------
TABLE D-7. TRACE ELEMENT DISTRIBUTION IN WELLMAN-GALUSHA GASIFIER AT 1200 F
SPECIES
SI02
Ft
CAS
HGO
CACL2
TI02
NA2C03
AL203
B203
ZHO
MNO
ZR02
NI
V204
CU
TIN
TI
MN
C
ALN
SI
ZRN
ZR
NA20
CUO
bEO
CAO
NA
CR203
AL
MG
VN
V
K
K20
• CAC2
9N
B
NIO
FE304
ZN
MAOH
BE
P8
CR
PBCL2
CR3C2
CRCL2
t 10 PtKC£NT BY WEIGHT OF THE
IN THE PRODUCT GAS STRtAM AS
HOL. WT. MOLES
(ASHtGAS)
60.1
55.9
72.2
t*0.3
111. 0
79.9
lOo.O
101.9
69.6
HI. 4
70.9
123.2
58.X
165.9
63.5
61.9
i»7.9
54.9
12.0
41.0
23.1
105.2
91.2
62.0
79.5
25.0
56.1
23.0
152.0
27.0
24.3
65.0
51.0
39.1
9*.2
6t. 1
2i*. 8
10.8
74.7
231.6
65. 4
40.0
9.0
207.2
52.0
278.1
183.1
122.9
.75779E-02
.29.375t£-02
.14727E-02
.17509E-03
.16840E-03
. 122995-03 '
.'31519E-04
.43143E-04
.39751E-04
.3SO'jl£-04
.672b8il-05
.50407E-05
.24931E-05
.2058411-05
.12750E-05
.•37610E-06
.'37609£-06
.96095E-06
.93452c>06
.92857E-06
.92515£-06
.90203E-06
.84893E-06
.81712E-06
.7b50lE-06
.76409E-06
.75993E-06
.72633E-06
. 70637E-06
.68G72E-06
.6366'«£-06
.68616E-06
.68606E-06
.68332t-06~ "
.6.3018E-06
.6i»33b£-06
.63097E-06
.53097E-06
.57533E-06
.56061E-06
.55636E-06
.SZZ^Qf.-Of>
.38195E-06
.35883E-06
.28251E-06
.179U9E-06
.I'tl't6£-06
,l'»128c-06
GRAMS
(ASH+GAS)
,t*55'*3L*00
, 161.21E+00
,10633E*00
•70563E-02
. 18693E-01
.902&7E-02
.97010E-02
.W3962E-02
.27667E-02
.28532E-02
. <,7693E-03
. 621 02t-03
.li«63<«E-03
. 3"tl'»9E-03
,6a962E-0'»
.&a'*20E-0'»
.'.6755E-0'*
. 52756E-0'«
,112li4E-0'»
. 38072E-01*
,25997E-Oi»
. 9<*89'«E-0'»
. 77i»22t-Qit
.50&61E-0'*
.60618E-0'*
. 19102E-0;*
,'*2b32E-0'»
.1&706E-0
WT. PERCENT PPM BY WT.
(IN GAS) UN GAS)
.2334300
,08«»l&
-------
TABLE D-7. (Continued)
OE3N2
CD
PUO
20
COCL2
CDO
SB203
HG
VC
TIC
SIC
CA
MGCL2
SPECIES
HZ
CO
C02~ ' " ~~
H20
CHi*
H2S
K
COS
NHJ
P«,OiO
AS
SE
CS2
NA
HG
FE(OH>2
S02 "
3ECL2
"S ' ' "" """
p
TICLi*
ALCL3
KO
FE
CL2
SEO
HGCL2
B02
no
HGO
ZRCL<*
AS203
ALOH
55.1
112. i.
32J.2
121.8
113.3
29 1 . 5
6 J.O
5 J.9
1*0.1
1,0.1
95. 2
MOL. WT.
2.0
28. 0
i»i». 0
11.0
16. 0
2(5. 0
39.1
60.1
l/.O
243.9
7"+. 9
20 7.2
79. 0
""" " "" 76.1
23.0
2fl'J.6
flj.9
79.9
""32.1
31. 0
1*3.7
133.3
55. 1
55.9
70.9
95.0
?71.5
1,2.8
26.8
216.6
233.1
197.8
4!+. C
.127i*2£-06
.95950c-07
.a-J70ot-07
.53533E-07
,i»8025E-07
.i*7925E-07
. 16081E-07
. 20725E-03"
.20000E-09
.20000E-09
.20000E-09
".'0
.0
MOLtS
(ASHvGAS)
.2J390E+00
.1-3730E + 00
.70676E-02
. 39578E-02
.3H782E-03
.1908QE-03
."<,9056E-OS"
. 16003E-05
.77679E-07
,<»2659E-08
.37567E-11
. 6<,3i,8E-ll
.75i,bOE-13
."25 2 00" """-"""U
.<*5113E-15
.U1Q33E-15
'.18627E-15
.39755E-16
. 16804E-17
.61050E-18
.2307i*E-18
.212>*dE-21
.15730S-22
.70207E-05
. 10785C-01, "
.20022L-01*
.65209E-05
.88030E-05
.61536E-05
. 1.6877E-05
. 12600E-07
. 11980E-07
. 80200E-08
. 0 0
.0 0
UH£ FOLLOWING GASEOUS
r.RAMS WT
.0000036
.0000055
.0000103
.0000033
.000001,5
.0000032
.0000021,
.0000002
.0000000
.0000000
.0000000
.0000000
.0000000
SPECIES WERE
. PERCENT
(ASHtGAS) (IN GAS)
. B2'J<,1E*00 " 1,
.82293E+01 1,2
.22979E+01 11
.24101E+00 1
. 11082E+00
. 13600E-01
. lli,67E-01
. 13927E-02""
. 11986E-03
. 16729E-03
. 17307E-01,
. 19012E-06
.31009E-07
. 56131E-09 "
.2i,223f-ll
. 15017E-11
.i*9702E-12 "
. 1&723E-12
. 22937E-13
. 13206E-13
. 37767E-11,
.61397E-15
.71920E-16
. 16361E-16
. i*9979E-16
6i,229E-l7 ~ " ~ "
.i,2029E-19
. 69211E-21
.2203959
.179231,2
.1,71*9379
.7780368
.0037322
.2352776
.5679927
.069701*8
.0537735
.00 111,16
.0071382
.000611,1+
.0008571,
.0000913
.0000303
.0000010
.0000069
.0000002
.0000000
.0000000
.00000 00
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.036
.055 •
.103
.033
.032
.021,
.002
.000
.000
.000
0.000
0.000
PREDICTED)
PPM Bf WT.
(IN GAS)
1*2203.959
1*21792.31*2
35<«7<,J.379"~~
1 17780 .3&8
',0037.322
12352.776
5679.927
697.01*8
587.735
11 .'*!&
71.382
6.11*1*
8.57<*
.913
.303
.010
.059
.002
.000
.000
~ .000 "
.000
.000
.000
.000
.000
'.000
/.GOO
. .000
.000
.000
.000
.000
.000
.000
.31*7
.533
.990
.323
.1*35
.301*
.232
.021
.001
.001
.000
0.000
0.000
TOTAL
PPM (1Y WT.
(ASH+GAS)
1*0730.218
1*07063.565
31,2361.708
113667.537
38639.23d
11921.1,23
51*81.587
672.707
567.212
11.017
68.890
5.929
8.275
.881
.292
. 009
.067
.002
.000
.000
.000
.000
.000
.000
.000
. 000
.000
.000
. 000
.000
.000
.000
.000
.000
. 000
8.951
13.7<*9
25.526
8.313
11. 223
7.81.5
5.976
.530
.016
.015
.010
0.000
0.000
1000000.00
O
oo
-------
TABLE D-7. ' (Continued)
02 T?.Q ^526655-23 L16853E-21 .0000000 .000 .000
"T£U2 ~ llT. 0 .55231£-2'T .bl306£"-22 .0000000 .ti~OT8 ."000"
Cfl(OH)2 7^.1 .117y5E-30 ,67i»28E-29 .0000000 .000 .000
SI 29.1 .ll<«38E-35 . 321<*l£-3<* ,0000000 ' .000 .000
TOTALS .19510E+02 100.000000 1000000.00 1000000.00
7
I-1
vo
-------
TABLE D-8. TRACE ELEMENT DISTRIBUTION IN RILEY-MORGAN GASIFIER AT 1250 F
SPECIi.3
SI02
FE
CAS
MGO
CACL2
TI02
NA2C03
AL20J
B203
ZNO
•1NO
ZR02
VC
HI
CU
TIN
TI
MN
ALN
SI
ZRN
ZR
NA20
CUO
OEO
CAO
NA
CR203
AL
MG
VN
V
K
K20
CAC2
8N
0
C
NIO
FE301*
ZN
NAOH
V20i*
BE
Pb
CRCL2
PDCL2
CR3C2
MOL. WT.
60.1
53.9
72.2
i*0.3
111.0
79.9
106. 0
101.9
69.6
81.**
70.9
123. 2
'i . 1
23.0
152.0
27.0
2<*.3
65.0
51.0
39. 1
9<*.2
f>'+. i
2i*. 8
10 .8
12. 0
7-*. 7
231.6
65.1*
i»0.0
165.9
• 9.0
207.2
122.9
278.1
183.1
( 10 Pl~k3ENT JY WEIGHT OF THE FOLLOWING CONDENSED SPECIES 4Ri INCLUDED
IN THE PRODUCT GAS STREAM AS FLY ASH? THE BALANCE IS *EH3VEO AS ASH)
riOLi'S GKAI1S HT. PEKCENT PPM DY WT. PPM BY HT .
(ASH+GA3) (ASH+GAS) (IN GAS) (IN GAS) (ASH+GAS)
./5779E-02
.29375E-02
. li,728£-02
. 17509E-03
.16826E-03
.12299E-03
.91500£-Qi»
."+31<*3£-0<*
. 39751E-Q1+
.35 051 £-0
-------
TABLE D-'8. (Continued)
CR
HE JNZ
CO
"" PBO
su
CDCL2
COO
50201 '
HG
" TIC
SIC
MGCL2
CA
SPECIES
CO
H2
CH-+
C02
H20
H2S
N2
COS
K
NH3
P4010
CS2
AS
PB
SE
NA
HG
•JECL2
FE(OH)2
P
S02
S
ALCL3
FE
CL2
KO
SEO
HGCL2
ZRCL1+
BO
802
HGO
ALOH
52.0
•55.1
223. *2
121.8
18J.3
12(1. i*
291.5
200.6
59.9
95.2
1*0.1
MOL. WT.
2.0
16,0
t.i». 0
1 lo.o
34.1
60. 1
39.1
17. 0
281.9
7o.l
" 7 !, . 9
207.2
79.0
2J.O
^90 . 6
79.9
89.9
31.0
1*9.'7
32. 1
133.3
53.9
70.9
55.1
95.0
271. 5
233.1
26.8
1+2.8
216.6
,li»125E-06
,1271*2£-06
.95960E-07
.89706E-07
.19170E-08
.20000E-09
.20000E-09
.0
.0
(THE
MOLES
(ASH+-GAS)
'.aJ^ztio!
. 313ii6E-Oi
,13'*29E-01
,68b21E-02
.39593E-02
.39295E-03
.31+782E-03
.9096511-05
.17050E-05
.16003L-05
.«0738£-06
,i+57&0£-07
.63830E>U8
.82816E-10
.12066E-10 "
.56068E-11
.72621E-12
. 35932E-12
.60371E-13
,82&2<+E-15
.60753E-15
.18761E-15
.-91172E-17
.52855E-17
.17170E-17
.89201E-19""
.10i*20i£-2l
73i»52E-05
10785E-OI+
20022E-Q1*
88030E-05
615 36E-05
<«6S77E-05
38i*56E-06
11980E-07
80200E-08
a
0
.0000053
.0000051
.0000079
.000011*6
.00000<*7
.0000061+
. UU U UU HP
.0000031*
.0000003
.0000000
.0000000
0.0000000
0.0000000
FOLLOWING GASEOUS SPECIES WERE
GRAMS
(A3H + GAS)
. 8&378E+ 01
.23591E+-01
. 13810E+01
. 2'*173C + 00
. 2 J1* 0 OE+ 00
. 11087E+00
.23616E-01
.13600E-01
.ll¥27r-0"2~
.12975E-03
. 11980 E- 03
.16729E-03
"". 17807E-0'*"
.10525E-05
"'.13807E-05
.b6170E-Od
. 108£,8E-03
. 17381E-09
.1.6550E-10
. 66163E.-10
.e&27lE-ll
. 80i»7itE-ll
,i»&187£-13
.10337E-13
. 13602E-11*
".2'»753E-1'."
. 12320E-H*
.92261E-16
.73i*87E-16
;i9321E-16"
WT. PERCENT
(IN GAS)
62.90931*1*1
3.2660208
18.63/7915
10 .0578818
1.7605332
1.70i»2i27
.8071*950
.1719995
.09901*68
.0011263
' .01011*30
.00091*50
.0000730
.0012181*
.0001297
.0000077
.UUU0101
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.0000000
.053
.051
.079
.11*6
.01*7
.06<*
rt 1 f
»Uf i»
.03i*
.003
.000
.000
0.000
3.000
PREDICTED)
PPM BY WT.
(IN GAS)
629093. !»!*!
32663.203
186377.915
100573.013
17605.332
1701*2.127
807"+. 950
1719.995
990.1*58
l 1.263
101. 1*30
9.1*50
8 .730
12.18"+
1.297
.077
.101
.000
.000
.000
.000
.000
.000
.000
.000
.000
.000
xOOO
.000
.000
.000
.000
.000
.000
.509
.u&
.7«*7
1.387
.610
.325
.027
.001
.001
0.000
0.000
TOTAL
PPM BY WT.
(ASH+&AS)
598333.170
31063.21*8
177261*. 7»»6
95660. 897
16208 .832
7680.115
1635.891*
91*2.038
10.712
96.1*71
8.987
8.303
11.588
1.233
.073
.096
.000
.000
.000
.000
.000
.000
.000
.000
.000
.000
.000
.000
.000
.000
.000
.000
.000
9.365
8.951
13.751
25.528
8.311+
11.22** ~" "
5.977
.1*90
.015
.010
0.000
0.000
1000000.00
a
M
AS 2 03
.79»39E-21
,0000000
.000
,000
-------
TABLE D-8. (Continued)
02
SE02
CA(OH)2
SI
32. 0
111.0
71.. 1
23. 1
.63ftfa6£-2H
.63058E-25
. 11082E-31
,26<»21£-33
.20123E-22
.69995E-23
.82117E-30
.7<»2't2t-32
.0000000
.0000000
.0000000
.0000000
'.000
.000
.000
.000
.000
.000
.000
.000
TOTALS
.13X3J.E + 02
100.000000
1000000.00
1000000.00
ho
N3
-------
D-23 and D-24
REFERENCES
(1) Cruise, D. R., J. Phys. Chem., 6_8, pp 3797-3802 (1964).
(2) Gordon, Sanford and McBride,. Bonny J. , "Computer Program for Calcu-
lation of Complex Chemical Equilibrium Compositions", "Rocket
Performance, Incident, and Reflected Shocks", and "Chapman-Jouquet
Detonations", NASA SP-273, 1971.
(3) Alexander, C. A., Hoyland, J. R., and Treweek, D. N., "Evaluation of
Computerized Techniques for Predicting Chemical Reactivity and
Stability", Final Report from Battelle's Columbus Laboratories
to U.S. Coast Guard, DOT, April 11, 1975.
-------
APPENDIX E
GAS CLEANUP SYSTEMS
-------
E-2
APPENDIX E
GAS CLEANUP SYSTEMS
Scrubber Systems
A number of water scrubbing systems have been proposed for use
with coal or oil gasification processes. These were shown in Appendix C.
Such systems have been depicted as part of gasification systems in
Figure C-4 for the Pviley-Morgan gasifier system, in Figure C-5 for the
Shell gasification process, and in Figure C-6 for the Texaco gasification
process. Each of these scrubbing-systems has different characteristics
which are dependent upon the characteristics of the raw gas and the
desired purity of the clean gas. Scrubber systems for other processes
which have not been documented undoubtedly have differences in details
of design but must be generically similar.
For the Riley-Morgan scrubber system (Figure C-4), a spray
tower is followed in sequence by (1) a quench chamber, (2) a three-pass
gas condenser, and (3) an electrostatic precipitator. The gas leaves
the quench chamber at 76 C (169 F) and the gas condenser at 43 C (110 F).
As will be discussed later, a temperature as high as 76 C (169 F) will
not be conducive to the removal of ammonia. Also, there is no provision
in the scrubbing liquor loop for the removal of the ammonia values. Thus,
the ammonia concentration would be expected to build up to some steady-
state concentration in the scrubbing liquor loop and, from that time on,
there would be little change in ammonia concentration in the gas phase
through the scrubber.
The scrubber shown associated with the Shell gasification process
in Figure C-5 also has no provision for removal of contaminants such as
ammonia in the scrubbing liquor loop. However, since this is a heavy oil
gasifier, the ammonia content in the raw gas should be very low. A purge
and a fresh water makeup stream do provide some capability for removal
of contaminants.
The scrubber separator and water stripper shown as part of the
Texaco gasification process in Figure C-6 is probably the most effective
-------
E-3
of th.e gas cleaning systems depicted from the standpoint of removal of
impurities like ammonia and hydrogen cyanide. The water stripper should,
as will be discussed later, drive off the gaseous impurities before the
water is recycled to the scrubber. One deficiency appears to be that no
cooling is performed on the stripped water which goes from the bottom of
the stripper to the top of the scrubber. To provide effective srubbing
(and cooling) there must be some provision in the water loop for cooling,
although it is possible, for simplicity in the drawing, that this provision
was omitted.
Raw Gas Compositions
Representative compositions of the raw gas leaving the heat
recovery section of coal gasifiers are shown in Table E-l. The oxygen-
blown case is based on a Koppers-Totzek gasifier. The concentrations of
H£ j CO, C02J H^S, anc* ^^ were based on the data for Illinois coal from
Table C-l. The concentrations of fUS and COS were adjusted to a coal sul-
fur content of 3.5 weight percent, based on the data in Table C-l. The
concentrations of the five species listed above were then reduced proportion-
ally to permit the addition of t^O, NH , Ar, CH,, and HCN, with these con-
centrations being directly obtained from Reference 19 of the text. The
concentration of water vapor is fairly high (29 volume percent) because the
Koppers-Totzek heat recovery system sprays water into the gas.
The particulate matter concentration in the Koppers-Totzek gas also
is from Reference 19 and applies for coal containing about 10 percent ash.
For this gasifier about half the ash in the coal goes into the gas. The
operating temperature of the K~T gasifier is so high that very little tar
is formed.
The air-blown case shown in Table E-l is based on a Wellman-Galusha
gasifier. The concentrations of '3.2, CO, C0~, CH,, and N£ were based on
the data for bituminous coal from Table C-2. The concentrations of I^S,
COS, Ar, and HCN were estimated by adjusting the oxygen-blown data based
on the fact that the gas production is 66 scf/lb coal for Wellman-Galusha
air-blown and 31 scf/lb coal for Xoppers-Totzek oxygen-blown. The NH,
concentration was taken as 0.5 volume percent, reflecting the fact that
-------
E-4
TABLE E-l. REPRESENTATIVE RAW GAS COMPOSITIONS
Concentration (mole % except as otherwise noted)
Oxygen-Blown Gasifier
Specie (Koppers-Totzek) (a)
H2
CO
C02
CH4
H2S
COS
NH3
HCN
N2
Ar
v
Tar -(grains /scf)
Particulate matter (grains/
scf)
24.43
39.08
4.97
0.08
0.92
0.08
0.17
0.03
0.71
0.32
29.19
100.00
11.57
Air-Blown Gasifier
(Wellman-Galusha) (b)
13.32
25.39
3.02
2.40
0.54
0.02
0.50
0.01
44.65
0.15
10.00
100.00
3.82
0.1-0.5
(a) From data on Illinois coal (Table C-l) after adjustment to 3.5 percent
S coal and addition of H20, NH , Ar, CH , and HCN from Reference 19 of
the text. J 4
(b) From data on bituminous coal (Table C-2) after addition of HJD (assumed
concentration), H2S, COS, Ar, and HCN (adjusted from oxygen-blown data),
and NH ..
-------
E-5
lower temperature gasifiers such as the Wellman-Galusha produce more NH-
than the higher temperature Koppers-Totzek gasifier. Admittedly, the
NH_ concentration for the air-blown case (actually the lower temperature
case) is not well defined.
For the Wellman-Galusha gas, the particulate matter and tar
concentrations were based on data stating that the particulate matter in
the gas comprise 1 to 5 percent of the ash in the coal and the tar in the
gas about 3.6 weight percent of the moisture free coal. The particulate
matter concentration was calculated for a coal containing 10 percent ash.
Performance of Spray Towers
The performance of spray towers in the removal of ammonia from
air have been explored by a number of investigators. Discussions of these
investigations are contained in Perry and in Sherwood and Pigford.
The most extensive investigation appears to be the one of Pigford and
Pyle; the information in this latter paper was used almost exclusively
as the practical basis for the analysis performed to predict the removal
of NH and HCN from reducing gases produced from coal.
The performance of spray towers is dependent upon a number of
variables. The characteristics of the spray, i.e., the droplet size and
velocity, are very important. This means that the characteristics of the
spray nozzle, the pressure drop through the nozzle, and the flow rate of
the liquid will strongly affect the droplet characteristics and, therefore,
the absorption characteristics of the column. Also of importance are the
column diameter, the number of spray nozzles used, and their placement.
If there is substantial interaction among the sprays from the various
nozzles and if the sprays impact the walls so that a significant fraction
of the liquid flows down the walls;, then absorptive performance will be
degraded. Thus, care must be taken in the design and operation of a spray
absorption column that desired absorptive performance is achieved.
'In the investigation reported by Pigford and Pyle, ' ammonia
absorption from air was measured in a 31.5-inch diameter spray tower. Tower
heights were 52 inches and 26 inches in two distinct series of experiments.
Gas inlet rates were varied from 218 to '817 Ib/hr-sq ft and water inlet
-------
E-6
rates were varied from 285 to 955 Ib/hr-sq ft. In the majority of experi-
ments six Sprayco 5-B nozzles were used. Nozzle pressure drops varied from
2 to 35 Ib/sq in. and were dependent upon the water inlet flow rates.
The number of overall gas-film transfer units, NQGJ was evaluated
for each experiment run at constant conditions. It was found that N
(for a fixed tower height) increased approximately in proportion to the
liquid rate and decreased as about the square root of the gas velocity.
For the range of gas and liquid flow-rate conditions stated, experimental
values for NQG ranged from 0.30 to 2.43 for the 52-inch high column and
from 0.29 to 1.31 for the 26-inch high column.
Further pertinent results and conclusions stated by Pigford and
Pyle relate to transfer unit height and maximum number of transfer
units to be expected. They state that fewer transfer units per unit of
height can be obtained in a tall chamber than in a short one. This is
because longitudinal gas mixing is prevalent in spray towers and the gas
may be so thoroughly mixed that it has nearly the same composition through-
out. Thus, spray towers are limited to applications in which only 1 to 3
transfer units are required. If true countercurrent action were to be re-
quired from a spray-type unit, it would be necessary to collect the liquid
at intervals of a few feet of tower height and respray it.
With this background, it becomes rather apparent that absorption
by a spray tower would not be tremendously effective unless the liquid would
be collected and resprayed several times; this technique, however, would
appear to be precluded by the presence of the particulate matter which would
almost certainly clog most known nozzle designs. With a single group of
spray nozzles at the top of the scrubber, a common, practical and workable
arrangement, the maximum number of transfer units achievable would appear
to be about 3. This value has been selected as a basis for the subsequent
analysis which is described.
-------
E-7
Method of Computation
Mathematical Relations
Computational methods for determining the number of transfer
units are succinctly described by Geankoplis. He indicates [Equation
(8.3-58).] that
(1-y) *M
N
OG
1-y
ave y2
_v*
y-y
where y = mole fraction of gas being absorbed in the bulk gas
y* = mole fraction in the gas in equilibrium with the
bulk liquid phase.
He also indicates that a simplification can be obtained if the
solution is dilute and the equilibrium line is straight over the range of
concentrations under consideration. Both of these conditions are true
for the ammonia and hydrogen cyanide absorptions being evaluated here.
Therefore, since
y* = mx
where m = slope of the equilibrium line at a constant
temperature
x = mole fraction of absorbed gas in the bulk liquid
then „ „
dy '2 '1
y-y*
'(y-y*)1 - (y-y*)2'
(y-y*).
In
ln (y-y*) 2
Furthermore,
(l-y)*M
1-y
ave
(1-y*) - (1-y,,)
In^Z?
O1^
(1-y*) -
In '^P
-------
E-8
One further relationship is required before N can be computed.
U(j
This is the so-called "operating line". In reality, this is merely the
equation which results from applying a material balance between the gas
phase and the liquid phase in the absorption column. It can be stated as
x2 - X]_ + (yry2)G
(y1-y2)G + L
G = gas flow rate, Ib mole/hr-sq ft
L = liquid flow rate, Ib mole/hr-sq ft.
Since, for the conditions of interest, the equations for
computing N can be expressed analytically, a direct computational procedure
Utr
can be established. This was done and translated into a Fortran IV code
for machine computation. A printout of this code is included at the end
of this appendix.
Experimental Data
To permit computation, experimental values were needed for the
slopes of the equilibrium lines for ammonia (NH«) and hydrogen cyanide
(HCN). These were obtained from literature sources. Data for NH, were
obtained from Perry and from Geankoplis. Values used for computation
were as follows:
Temperature, Slope of
C Equilibrium Line (for NHj , m
j
30 1.198
40 1.936
50 2.791.
Data for HCN were obtained from Linke and from International Critical
Tables. Values used were
Temperature, Slope of
C Equilibrium Line (for HCN), m
30 (30.2) 7.229
40 (40.2) 9.947
50 14.129.
-------
E-9
As will be noted for the HCN values, actual temperatures for the reported
data were 30.2 and 40.2 C. In all results, however, these are labelled
20 and 40 C, respectively.
Machine Computations
Computations utilizing the Fortran IV program previously
mentioned were performed on Battelle's CDC 6400 computer. Specific values
for G, L, y., y_, and X1, were selected and the number of transfer units
(N ) computed. Nominal values for G and L were those representing
(4)
maximum performance in the Pigford and Pyle experiments. Variations in
the L values were selected as twice and one-half the nominal value. All
values used for computation are expressed with the results which are
summarized in the following section.
Removal of NH_ and HCN by Spray Towers
Computations were made for the range of conditions shown in
Table E-2; the results obtained form the basis for the assessment of
performance which is contained In. this section.
A plot of results for removing ammonia from the raw gas is
shown in Figure E-2. The number of transfer units, Nn_, for various
(j\j
inlet water ammonia concentrations is plotted versus the ammonia concen-
tration in the outlet gas. Other factors are T = 30 C, G = 7.556 Ib
moles/hr-sq ft, L = 53.01 Ib moles/hr-sq ft, and y = 1700 ppmv. As will
be observed, a greater number of transfer units is required to achieve a
given exit gas concentration when the concentration of ammonia in the
inlet water is higher. This is riot unexpected behavior.
A series of these plots was constructed and then interpreted
to obtain further performance estimates. For example, two additional plots
-------
E-10
TABLE E-2. MAXIMUM EXPECTED REMOVALS OF NH BY WATER SPRAY ABSORPTION TOWERS
(For Inlet Concentration of NH in Water = 0)
Water
Temperature,
C
30
I
30
1
40
40
I
i'
50
1
T
50
1
i
Liquid
Rate, L,
Ib moles/hr-sq ft
26.505
53.01
106.02
26.505
53.01
106.02
26.505
53.01
106.02
26.505
53.01
106.02
26.505
• 53.01
106.02
26.505
53.01
106.02
Inlet Gas
Concentration ,
ppmv
1700
I
5000
1
1700
1
1
5000
1
1
1700
1
\
5000
1
if
Outlet Gas
Concentration ,
ppmv
163
120
101
479
352
296
232
143
112
681
426
330
331
180
126
1071
529
421
-------
E-ll
0.0001 0.0002 0.0003 0.0004 0.0005
Outlet Gas Concentration , y2 , mole fraction
0.0006
OJ0007
FIGURE E-2. NUMBER OF TRANSFER UNITS FOR AMMONIA REMOVAL
-------
E-12
similar to Figure E-2 were constructed at similar conditions except for
the inlet water temperatures which were 40 C and 50 C, respectively. Then,
recognizing, as previously discussed, that the probable maximum number of
transfer units which could be achieved in a spray scrubber would be 3,
cross plots were constructed. The first of these is shown in Figure E-3.
In Figure E-3, the ammonia concentration in the outlet gas is
shown as a function of the ammonia concentration in the inlet scrubbing
water. As stated, this is for an N = 3 and an inlet ammonia concentration
00
in the gas of 1700 ppmv. Smoothed curves (straight lines) through the
points indicate the manner in which outlet gas concentration of ammonia
will increase with inlet water concentration. Limiting (maximum) values
for removal occur when the inlet water concentration of ammonia is zero.
At 30 C, the minimum ammonia concentration in the outlet gas would be
115 ppmv. At 40 C it would be 140 ppmv and at 50 C, 180 ppmv.
Figure E-4 is a plot similar to Figure E-3 but for an inlet
concentration of ammonia in the gas of 5000 ppmv. For this case, the
minimum concentrations of ammonia in the outlet gas would be 350 ppmv
at 30 C, 420 ppmv at 40 C, and 525 ppmv at 50 C.
Similar plots could be prepared for hydrogen cyanide removal.
The behavior of the system would be similar except that the concentrations
of HCN in the water streams entering and leaving would be much lower than
those for NH .
In Table E-2, the maximum expected removals of ammonia for
the variety of conditions which were computed are tabulated. A similar
tabulation for hydrogen cyanide is shown in Table E-3.
In Tables E-2 and E-3, the liquid rates shown can be compared
with those used experimentally by Pigford and Pyle. ^ ' The liquid rate
of 53.01 Ib moles/hr-sq ft corresponds to their L at their maximum L/G.
The other two liquid rates used, 26.505 and 106.02 Ib moles/hr-sq ft,
were selected for these computational purposes as one-half and twice,
respectively, the nominal value. The lower circulation rate would be
less costly and the higher rate more costly for operation than the nominal
one.
Inspection of the results Table E-2 indicates the penalty
in removal efficiency resulting from higher operating temperatures. The
-------
E-13
0.0006
0.0005
c
o
o
o
o
e
c
o
0.0004
•= 0.0003
a
c
o
O
at
O
0.0002
0.000 I
L = 53.01
y,=O.OOI7
NH3
T=50C
= 40C
0.00002 0.00004 0.00006 0.00008 0.00010
Inlet Water Concentration , X, mole fraction
0.00012
FIGURE E-3. OUTLET GAS AMMONIA CONCENTRATION AS A FUNCTION INLET WATER
AMMONIA CONCENTRATION FOR AN INLET GAS AMMONIA CONCENTRATION
OF 0.17 MOLE PERCENT (1700 ppmv)
-------
E-14
0.0008
0.0007 —
c
o
0.0006 —
o
£
.1 0.0005
o
c.
o
O
0.0004
O
O
0.0003 —
0.0002
0.00002
0.00004
0.00006
0.00008
0.00010
Inlet Water Concentration, x(,mole fraction
FIGURE E-4.
OUTLET GAS AMMONIA CONCENTRATION AS A FUNCTION INLET WATER
AMMONIA CONCENTRATION FOR AN INLET GAS AMMONIA CONCENTRATION
FOR AN 0.50 MOLE PERCENT (5000 ppmv)'
-------
E-15
TABLE E-3. MAXIMUM EXPECTED REMOVALS OF HCN BY WATER SPRAY ABSORPTION TOWERS
(For Inlet Concentrations of HCN in Water = 0)
Water
Temperature,
C
30
30
*
40
f
40
*
50
50
Liquid
Rate, L,
Ib moles /hr-sq ft
53.01
106.02
53.01
106.02
53.01
106.02
53.01
106.02
106.02
106.02
Inlet Gas
Concentration,
ppmv
100
*
300
*
100
1
300
1
100
300
Outlet Gas
Concentration,
ppmv
26
13
78
45
37
17
111
52
25
76
-------
E-16
results also indicate that, with inlet ammonia concentrations of 1700 ppmv,
outlet concentrations less than about 120 ppmv cannot be achieved in a
single-stage scrubber. For an inlet concentration of 5000 ppmv the corres-
ponding outlet concentration is about 352 ppmv.
In the computations for hydrogen cyanide removal, a number of
the cases selected for computation did not result in operable conditions.-
This is indicated by the data contained in Table E-3, in which only the
operable conditions are listed. The minimum outlet gas concentrations of
HCN are about 26 ppmv with 100 ppmv at inlet and 78 ppmv with 300 ppmv at inlet.
Influence of H,,S on NH- Removal
by Spray Towers
Subsequent to the performance of the above reported computations,
a limited analysis was made to estimate the extent to which hydrogen sulfice
(H.S) would affect the absorption of ammonia (NH.). This analysis also
yielded information on the expected removal of H-S from gas in a spray
scrubber system.
The computer program previously described was modified to permit
the simultaneous consideration of NH, and H-S absorption. This involved a
trial and error solution wherein a match of the number of transfer units
for NH- and H S adsorption was sought by modifying the assumed extent of
adsorption of H S.
(8)
Data for the computation were those derived by Beychok from
(9)
the work of Van Krevelen, et al. The data used were those describing
the simultaneous absorption of NH- and H.S. No attempt was made to consider
the influence of the presence of CO in this system even though data are
(9)
available from the work of Van Krevelen, et al. Consideration of CO
absorption along with that for NH- and H.S could easily have introduced
an order of magnitude of complexity into the computational procedure. This
did not appear warranted.
Results of the computations are shown in Tables E-4 and E-5.
In both tables the absorption of NH- when influenced by simultaneous
absorption of H_S is compared with the absorption of NH- when it was
assumed that there was no influence of H_S. These latter results are those
-------
E-17
TABLE E-4. ABSORPTION COMPARISONS AT 1.5 TRANSFER UNITS
FOR y-L(NH3) = 0.0017 and y-^S) = 0.0092
x1(NH3)
0.00002
0.00004
0.00006
0.00008
0.00010
...
0.00002
0.00004
0.00006
0.00008
0.00010
0.00002
0.00004
0.00006
0.00008
0.00010
0.00002
0.00004
0.00006
0.00008
y2(NH3)Ca)
L
0.000393
0.000409
0.000413
0. 0004 20 ^
0.000436
— . . — ...
L =
0.000391
0.000402
0.000414
0.000425
0.000463
L
0.000411
0.000417
0.000438
0.000442(c)
0.000449(c)
LJ ~"
0.000400
0.000416
0.000433
0.000452
Previous, ,
y2(NV
= 53.01, T =
0,, 000445
0., 000463
0,000481
0,000499
0.000517
-----
106.02, T =
0.000421
0.000439
0.000458
0.000476
0.000495
= 53.01, T =
0.000486
0.000515
0.000543
0.000571
0.000600
106.02, T =
0.000447
0.000477
0.000505
0.000535
x1(H2S)
30 C
0.000004
0.000008
0.000012
0.000016
0.000020
30 C
0.000004
0.000008
0.000012
0.000016
0.000020
40 C
0.000004
0.000008
0.000012
0.000016
0.000020
40 C
0.000004
0.000008
0.000012
0.000016
y2(H2s)(a)
0.001642
0.001645
0.001583
0.001573(c)
0.001569
' a
0.001838
0.001789
0.001787
0.001801
0.001924
0.001384
0.001291
0.001293
0.001188(c)
0.001131(c)
0.001647
0.001619
0.001603
0.001590
(a) Values obtained considering the simultaneous absorption of
NH3 and H2S.
(b) Values computed ignoring influence of H2S absorption.
(c) Underline indicates linearly extrapolated values; all other
values were linearly interpolated.
-------
E-18
TABLE E-5. ABSORPTION COMPARISON AT 1.5 TRANSFER UNITS
FOR y1(NH3) = 0.0050 and y1(H2S) = 0.0054
XjCSH)
0.00006
0.00008
0.00010
y2(NH3)(a)
L
0.001160
0.001117(c)
0.001175
Previous, N
y2(NH3)tb) Xl(H2S)
= 106.02, T = 30 C
0.001308 0.000012
0.001327 0.000016
0.001345 0.000020
y2(H2s)(a)
0.001758
0.001656(c)
0.001632
(a) Values obtained considering the simultaneous absorption of
NH3 and H S.
(b) Values computed ignoring influence of H S absorption.
(c) Underline indicates linearly extrapolated values; all other
values were linearly interpolated.
-------
E-19
previously discussed. Table E-4 covers inlet gas concentrations of 1700
ppmv for NH_ and 9200 ppmv for H»S; Table E-5 is for concentrations of
5000 ppmv for NH_ and 5400 ppmv for H«S. Comparisons were made at 1.5
transfer units because valid solutions could not be obtained at higher
values (i.e., 3.0 as used previously).
Inspection of the results indicates that the additional absorption
of NE_ obtained when considering the simultaneous absorption of RJ* is not
significant. Only about 3 percent more NH is absorbed. However, there
are significant amounts of H^S removed. However, the amounts of H^S
absorbed would probably be reduced to some extent in any actual situation
by the simultaneous absorption of CCK.
It is concluded that the results obtained for NH- absorption in
the calculations where the simultaneous absorption of H S was ignored are
reasonably valid.
There may also be an interaction between NH_ and HCN, since
these compounds are known to form NH CN in the liquid phase. However, the
concentrations of HCN are so low that an investigation of this interaction
did not appear warranted.
Stripping Column Performance
A major question in attempting to predict the extent of removal
of ammonia and hydrogen cyanide in a spray scrubber is that pertaining to
the concentrations in the inlet scrubbing water of the species being
removed. In an attempt to answer this question, some exploratory compu-
tations were performed to evaluate the ease with which ammonia can be
stripped from water. Computational techniques presented by Robinson and
Gilliland . and data from Perry were utilized. Because of the
exploratory nature of the computations, the techniques and data will not
be described here.
Results of the computations indicate that ammonia can be easily
stripped from water. Less than two equivalent plates would be needed in
the stripping section of the column to achieve an ammonia concentration
in the recycle water stream of 1 ppm (mole fraction of 0.000001). This
is substantially zero and should permit achieving within about 5 ppm of
-------
E-20
the minimum gas phase ammonia concentrations shown in Table E-3. With
more equivalent plates in the stripping section, even lower ammonia
concentrations could be obtained, which would result in an even closer
approach to the minimum values shown in Table E-3.
Similar computations were not made for hydrogen cyanide; this
could be accomplished, however, if there were sufficient interest.
Conclusions
Computations performed have indicated that substantial amounts
of ammonia and hydrogen cyanide are removed by spray scrubbers which are
installed primarily to achieve tar, tar oil, naphtha, and particulate
removal. Despite substantial removal of NH- and HCN by water scrubbers,
however, there are still significant amounts remaining in the cleaned
gas. These amounts are shown in Table E-2 for ammonia and Table E-3
for hydrogen cyanide.
There are several ways, not explored computationally, in which
modifications to the scrubber design, operation, or conditions might
result in greater removals of ammonia and hydrogen cyanide. Three of
these ways worthy of mention are:
e Increasing water circulation rates
e Increasing the number of transfer units by collecting
and respraying water
• Using lower water temperatures.
Each of these ways would result in a more costly operation but substantial
benefit might be achieved.
Of the three improvement methods listed, increasing the number
of transfer units by collecting and respraying the water would probably
be the most effective. Some difficulty would probably be encountered,
however, with particulate matter clogging the spray nozzles. Such problems
could probably be overcome by including filters in the respray loop and
by using nonclogging spray heads.
Additional computations to determine more extensively the
behavior of scrubbing systems could be performed and would probably be
worthwhile.
-------
E-21
Computer Program for Number of Transfer Units
This section contains the FORTRAN computer program used for
calculating the number of transfer units (N ) for gas washing equipment.
-------
POQGRAH Lf- »
DATA LARR/53.01,106 .C2.26.505/
flflTA Mfl(?P/1 . 1 QAk1 •?. 1 . qT^A-t 0 . 7 . 7 Q fl 5 A U .7 _ ??Qn 7 « . Q . Qi.71 Kf. . 1 /. 1 ?Q/,A?C«
OAT A M(~
1
-------
c
r
C
r.
r*
w
r.
C
c
C
r.
*
M = HARR(XTMO£X)
LNCNT-c;e;
V^PY LT3UIP PATC
nn 1UOJ TTMT*y-it7
L=LARR( IIMO&X)
VfiPY I Tallin TNI FT r.nMr.FNTRATTnN
IF NM7, tfARY BY INCREMENTS OF .00002
IF wr-j, uflpv t»v IMCPEMEKITS nr ,u03P02
XI =-o. 0 0 0 f J
XII NC = 0.j;J02
TFtVTND.-y -I". ^) AOTn 111
XI = -0.03J032
y-i TK>r = ri . i,in nr ?
11Q
n
C
r.
C
/*
C
CONTINUE
nn 1^00 jTNin.-Y = 1 , f,
X1=X1*X1INC
VARY INUtT GAS CONCENTRATIONS
t
LMIN=1 W
JP H^M KlK-npy =tt THP" 6)t "SE A niFppBr^T SET "F T"LET »?*S 'M
CONCiNTPATIONS ' ; ^
TF ( KTNOP X- ^T . 7 1 IM'TNsI
LMAX=LMIN*1
HO 1R11Q I I^4^i~y = ! MTM ,1 MAX
Y1=Y1ARR(LINOEX)
VARY ODTl FT GAS CONCENTRAT TON5
w
n
YINC=0.0
no irnn MT^n^xai ,?•;
YINC = YINCO.02
Y? = YTwr « vi
C
r
C
C
2 JO
r
C
r
TH^r.K TF Y? TS niiT OF pflur:F
TFfY? .T.F. M » Y1< KOTO ?.!(!
OtHUGGING CODE CAN GO HERE
•KOTO in. in
CONTINUE
X2 = U (Y1-Y2) » G) / ( ( IY1-Y2) » G> * L) > * XI
YF1 = M » Y?
-------
r.»rr.<< TH- \ffli THTTY OF YI
IFtYl . r,E. VE3 1 GO.IO 3d I
QEBUn-.ING CODE CAN GO HERE
-£0X0 10-3-0
303 CONTINUE
WE ARE AT THE LOW POINT OF THE LOOP.
_ACiUJ
YE2 = M * XI
nilANTI = Cl Or. f f 1 . - YF?I / (1 .-Y? 1
OUANT2 = ALOGt (1.-YE1)/(1,-Yl) )
_5LLt) YF1= Yt - Y p 1
Y2MYE3 = Y2 - Yf2
/wniinr = t f Y?HYF? /
1 ( Y1MYF1 / QHftNT2 ) ) / 2.
TNY2Y1 = (/1-Y?) / f fY1MYF1 - Y?MY-?> /
1 ( ALOG( Y1MYF1 / Y2MYE2 ) ) )
S—ky Q MOT *
OUT THE ANSWERS
IF(LNCNT.LE.51) GOTO 850 L|
-LNCtJTsj . NJ
PRINT S2C ^
1 11HLIOUIO RATE,i*X,12HLIQUIO INLET,
_t—4.X,13HLiCUIJ—OU-IL£J-,.5j( ^
2 3X.9HINLET GAS.5X, IjHOUTLfT GAS, 2X, 8HTRANSFER /
16HLJ-MOLFS/HR-SQFT, IX, 13HCONCENTRATION , IX,
L3.HCJ1K C c N T ?. fl T IJJtL, _ IX, _ 1 ft Hpnil-TJ T T?.I MM LINE, jgy, i3nr.nNCENT P.fl T ION
5, IX, 15HCOMCEHTRATION, IX, 7H UNITS / 1H 16(1H-),1X,
_6 ib MB^.L. I X «J-3-U.H=J-t-ljC , J_3J_LH-_K,:
7 IX , 13(1H-) ,IX, 13(1H-),1X, 13(1H-),1X, 8(lH-),/>
flSO r.ONTTNIIF
PRINT y>,j,';,L,Xl,X2,M,(MC(II,KINDEX) ,11 = 1,3>,Y1,Y2,NOG
FnQMflTt1y.?y.F1r.q.ax.F1n.q. qY.Fi?.7.iy. FI?.7.7Y.FIr.q ./Y.
1 Ai*,A<*, A2 ,2X ,Flf .5 ,4X,F12. 7, F11.5
10JO CONTINUE
C
_C CflHPI.irATinN-
c
END
-------
E--25 and E-26
REFERENCES
(1) Mason & Hangar - Silas Mason Co., Inc., and McDowell Wellman Engineering
Co., "Introduction of a Modernized Concept for Utilization of a Low-Btu
Coal Gasification Unit" brochure, June, 1974.
(2) Perry, John H. (Ed.), Chemical Engineers' Handbook, Third Edition, McGraw-
Hill, New York, 1950, pp 171-172, 674, 1685.
(3) Sherwood, Thomas K. and Pigford, Robert L., Absorption and Extraction,
Second Edition, McGraw-Hill, New York, 1952, pp 268-277.
(4) Pigford, Robert L. and Pyle, Cyrus, "Performance Characteristics of
Spray-Type Absorption Equipment", Ind. Eng. Chem., 43, pp 1649-1662 (1951).
(5) Geankoplis, Christie J., Mass Transport Phenomena, Holt, Rinehart, and
Winston, New York, 1972, pp 385-394, 478.
(6) Linke, William F., Solubilities - Inorganic and Metal-Organic Compounds,
Fourth Edition, American Chemical Society, Washington, B.C., 1958, Volume
1, p 1106.
(7) Washburn, Edward W. (Ed.), International Critical Tables, McGraw-Hill.
New York, 1928, Volume III, p 261.
(8) Beychok, Milton R., Aqueous Wastes from Petroleum and Petrochemical
Plants, John Wiley, New York, pp 161-175, 1967.
(9) Van Krevelen, D. W., Hoftijzer, P. J., andHuntjens, F. J., "Composition
and Vapor Pressures of Aqueous Solutions of Ammonia, Carbon Dioxide, and
Hydrogen Sulfide", Rec. Trav. Chim. Pays-Bas, 68, pp 191-216, 1949.
(10) Robinson, Clark S. and Gilliland, Edwin R., The Elements of Fractional
Distillation, Third Edition, McGraw Hill, New York, pp 114-118, 1939.
-------
APPENDIX F
DETAILS ON ENGINEERING ANALYSIS
-------
F-2
APPENDIX F
DETAILS ON ENGINEERING ANALYSIS
This appendix contains additional details related to the engineer-
ing analyses made as part of this project.
Quantity of Reductant Required
Basis: 1000-MW power plant
Coal heating value = 12,000 Btu/lb
Coal sulfur content =3.5 weight percent
Power plant heat rate = 9,750 Btu/kWh (35% efficiency).
Hourly coal burning rate to generate electricity
= 10 kW x 9,750 Btu/kWh x Ib coal/12,000 Btu x
ton/2000 Ib = 406.25 ton/hr.
Sulfur burned per hour
= 406.25 x 0.035 = 14.22 ton/hr.
generated per hour
64.06 SO
QO n*
J/ . (JO
x 14-22 = 28.41 ton/hr or 887.0 Ib-mole/hr.
Reduction of one mole of SO,, to sulfur requires 2 moles of
reductant (H. + CO + H S). The FGD process will remove about 90 percent
of the SO- from the flue gas. Some reductant in excess of the stoichio-
metric amount will be required to insure a high SO conversion. Using
about 11 percent excess reductant, the moles of reductant required can be
taken as twice the moles of S02 generated by the power plant. That is,
Reductant (H + CO + H S) requirement = 2 x 887.0 = 1,774.0 Ib-mole/hr.
or 1,774 x 359.2 scf/lb-mole x hr/60 min = 10,620 scfm.
-------
F-3
This does not include the gas which must be burned to reheat the reducing
gas. This is discussed in the following section.
Quantity of Reducing Gas Burned to Reheat Gas
When the water wash step is followed by a high temperature step,
as it is for the application to Type A FGD processes, the reducing gas
must be reheated. To reach a temperature around 340 C (650 F), the most
efficient method will normally be to burn a portion of the gas. The amount
of gas which must be burned and the effects on the final reducing gas
composition are determined by material and energy balances around the
burner system, as follows:
— Incoming gas ,_*
r
1
> 1 SCF^ ,
(1) '
1
1
l_
(l-x)SCF
x S(JKJ
.•"j But lit! 1.' • ' "
•f
Air or CL
1
1
i
1
1
i
i
!
Outgoing gas
x = fraction of incoming gas which is burned.
The basis for the calculations presented here is as follows:
(1) The incoming gas is at 30 C (86 F) and the outgoing
gas at 343 C (650 F).
(2) For the gas produced by an oxygen-blown gasifier,
the burner is also fired with oxygen. For the air-
blown gasifier the burner is fired with air.
(3) 10 percent excess oxygen or air is supplied to the
burner.
-------
F-4
Tables F-l and F-2 show the material and energy balances for
the oxygen-blown and air-blown cases, respectively. The quantity of gas
which must be burned is 4.50 percent of the incoming gas for the oxygen-
blown case and 8.65 percent for the air-blown case.
Flue Gas Quantity and Reheat Requirement
The applicable combustion reactions are
C + 02 -> C02
H2 + 1/2 0£ + H20
S + 02 -> S02
N + 1/2 02 •*• NO.
On the basis of 100 pounds of coal, the quantities involved are
C 70.5 Ib or 5.88 Ib atoms
H 5.0 Ib 2.50 Ib-moles
S 3.5 Ib 0.11 Ib atoms
N 1.5 Ib 0.05 Ib-moles
0 8.8 Ib 0.28 Ib-moles.
The theoretical amount of reaction products after combustion is
5.88 + 2.50 + 0.11 + 0.05 = 8.54 lb-noles/100 Ib coal.
The theoretical amount of oxygen required (neglecting that in the coal)
is
5.88 + (1/2)(2.50) + 0.11 + (1/2)(0.05) = 7.27 lb-moles/100 Ib coal.
At 20 percent excess air, the total air required is
7.27 x 1/0.21 x 1.2 = 41.51 lb-moles/100 Ib coal.
On combustion, the 7.27 Ib-moles of oxygen which reacts yields 8.54 Ib-
moles of reaction products. Thus, the total moles of reaction products
(including inerts in the air) is
41.51 + (8.54 - 7.27) = 42.78 lb-moles/100 Ib coal
or 855.6 Ib-moles/ton of coal.
-------
TASLE F-l. MATERIAL AND ENERGY BALANCES AKOUND REHEAT SYSTEM FOR OXYGEN-BLOWN GASIFIER
Basis: 1 scf Incoming Gas, x - Fraction of Incoming Gas Burned
Specie
H2
CO
co2
CH4
H20
Other
°2
Total
Incoming
scf
0.3065
0.4903
0.0624
0.0010
0.1152
0.0246
1.000
Gas at 30
Btu/scf
0.474
0.484
0.605
0.615
0.552
0.478
C (86 F)
Btu
0.1453
0.2373
0.0378
0.0006
0.0636
0.0118
0.4964
Oxygen for Combustion
at 21 C (70 F) Outgoing
scf Btu/scf Btu scf
0.3065 (1-x)
0.4903 (1-x)
0.0624 + 0.4913x
0.0010 (1-x)
0.1152 + 0.3085x
0.0090X 0.282 0.0025x 0.0246 + 0.0090x
0.4415x 0.282 0.1245x 0.0401x
0.4505x 0.1270x 1.0000 + O.OSllx
Gas at 343
Btu/scf
10.830
11.045
15.945
17=520
12.960
10.975
11.415
C (650 F)
Btu
3.3194 (1-x)
5.4154 (1-x)
0.9950 + 7.8338x
0.0175 (1-x)
1.4930 + 3.9982x
0.2700 + 0.0988x
0.4577x
11.5103 + 3.6362x
Energy Balance: (Incoming Gas) + (Oxygen for Combustion) + (Heat of Combustion)
0.4964 + 0.1270x + 248.14x - 11.5103 + 3.6362x
11.0139
(Outgoing Gas)
244.63
- 0.0450.
Amount of 98 percent oxygen required - (0.4505)(0.0450) - 0.0203 acf.
Amount of outgoing gas - 1 + (0.0511) (0.0450) ~ 1.0023 scf.
I
m
-------
TABLE F-2. MATERIAL AND ENERGY BALANCES AROUND REHEAT SYSTEM FOR AIR-BLOWN GASIFIER
Basis: 1 acf Incoming Gas, x •» Fraction of Incoming Gas Burned
Specie
H2
CO
co2
^4
H20
Other
Air
Total
Incoming
scf
0.1318
0.2511
0.0299
0.0237
0.1152
0.4483
1.000
Gas at 30
Btu/scf
0.474
0.484
0.605
0.615
0.552
0.478
C (86 F)
Btu
0.0625
0.1215
0.0181
0.0146
0.0636
0.2143
0.4946
Air for Conbustion
at 21 C (70 F) Outgoing
scf Btu/scf Btu scf
0.1318 (1-x)
0.2511 (1-x)
0.0299 + 0.2748x
0.0237 (1-x)
0.1152 + 0.1792x
0.4483 + 0.9615x
1.3753x 0.291 0.4002x 0.1250x
1.3753x 0.4002x 1.0000 + 1.1339x
Gas at 343
Btu/scf
10.830
11.045
15.945
17.520
12.960
10.975
11.035
C (650 F)
Btu
1.4274 (1-x)
2.7734 (1-x)
0.4768 + 4.38l7x
0.4152 (1-x)
1.4930 + 2.3224x
4.9201 + 10.5525x
1.3796x
11.5059 + 14.0202x
Energy Balance: (Incoming Gas) + (Air for Combustion) + (Heat of Combustion) - (Outgoing Gas)
0.4946 + 0.4002x + 140.94x - 11.5059 + 14.0202x
11.0113
127.32
- 0.0865.
Amount of air required - (1.3753)(0.0865) - 0.1190 acf.
Amount of outgoing gas - 1 + (1.1339)(0.0865) - 1.0981 scf.
>Tl
I
-------
F-7 and F^8
Quantity of Heat for Reheating Flue Gas After Scrubbing
The quantity of heat required to reheat the flue gas after
scrubbing can be estimated as follows. The flue gas usually must be
reheated to about 93 C (200 F) in order to provide adequate plume buoyancy
and to avoid condensation in the stack. Using a psychrometric chart for
air, the difference in enthalpy between 93 C (200 F) and 52 C (125 F) at
constant absolute humidity is about 21 Btu/lb. Since the molecular weight
of the flue gas is about 29, the heat required to reheat the flue gas from
52 C (125 F) to 93 C (200 F) is
855.6 Ib-moles 29 Ib 21 Btu _ ,01 nnn , ,
: rr :; rr = 521,000 Btu/ton coal.
ton coal Ib-moles Ib
If the coal has a heating value of 12,000 Btu/lb, the flue gas reheat
requirement represents about 2.2 percent of the energy in the coal input.
-------
APPENDIX G
SOURCES OF COST INFORMATION
-------
G-2
APPENDIX G
SOURCES OF COST INFORMATION
Coal Gasification Systems
This section includes cost information on a number of coal
gasification systems. A consensus of these data was used in developing
the costs used in this report. The consensus will be discussed after
each source of cost data is discussed individually. The investment data
are compared in Figure G-l, where they are plotted against the production
rate of reducing gases, i.e., H?, CO, and H^S (CH, is not considered
sufficiently reactive).
Wellman-Galusha, Air Blown
The basic cost data for Wellman-Galusha gasification systems were
taken from the brochure "Introduction of a Modernized Concept for Utili-
zation of a Lcw-Btu Coal Gasification Unit", which was prepared by Mason &
Hangar and McDowell-Wellman in June, 1974. This brochure gave the invest-
ment for a gasification system including one gasifier, a cyclone for parti-
culate matter removal, coal preparation facilities, and a small control
building. These data are shown in Table G-l. Also shown in this table is
the extension of these data to a system involving five gasifiers, this
being based on information obtained from McDowell-Wellman Engineering
Company in May, 1975. This information was that the additional gasifiers
would cost $400,000 apiece and that the five-fold capacity increase would
increase the coal preparation facilities cost by $240,000. These costs
do not include water washing of the gas, since the data did not permit
separating this from the H?S removal cost.
Wellman-Galusha, Oxygen Blown
A quotation was obtained from Mason & Hangar in 1975 on a gasi-
fication system involving three 10-foot diameter oxygen-blown Wellman-
-------
100.0
G-3
x
-------
G-4
TABLE G-l. COST DATA FOR WELLMAN-GALUSHA AIR-ELOWN GASIFICATION SYSTEMS
Coal Input Rate, Ib/hr
(a)
Dry Gas Production Rate, scfm
Production of H + CO + H S, scfnT )
(c)
Investment, escalated to 1/1/75 , $
Equipment
Erection (equipment)
Control building (erected)
Site work (producer plant only)
Total Construction Cost
Engineering
Interest During Construction
Startup (60 days)
Total Appropriation
Operating Cost Components, $/year
Electricity ($0.015/kwh)
Water ($0.25/1000 gal)
Chemicals (allowance)
Supplies (allowance)
Operating labor ($6/hr + fringes, G&A)
Maintenance (2 percent of equipment)
(a) Based on dry gas production of 65 scf/lb
(b) Based on gas composition given with cost
H2 + CO + H S.
One Gasifier
7,000
7,580
2,980
541,559
25,860
168,905
1,395
737,719
73,772
33,197
20,000
864,688
6,620
10,640
1,000
1,000
100,800
11,348
coal as given
data, i.e., 39
Five Gasifiers
35,000
37,900
14,900
2,704,688
with cost data.
.3 mole percent
(c) Chemical Engineering Plant Cost Index = 178.6.
(d) Operating time = 7000 hours/year.
-------
G-5
Galusha gasifiers. The production rate was 11,000 scfm of clean gas
having a low heating value of 290 Btu/scf. Adjusting the gas composition
from Table 9-2 to an oxygen-blown situation and adding 1.0 mole percent
H(?S, the gas should contain about 86.9 percent H + CO + H-S. Thus, the
production of H, + CO + ELS is 9,560 scfm. The investment quoted for
6
this system was about $2.67 x 10 but did not include engineering and
startup costs. The breakdown of this investment is shown in Table G-2.
Based on the air-blown Wellman-Galusha information, the engineering cost
was taken as 10 percent and the startup cost as $20,000. This made the
total investment about $2.96 x 10 .
Koppers-Totzek, Oxygen Blown
Two cost quotations were obtained from the Koppers Company in
March, 1975, on systems involving a single two-headed Koppers-Totzek
(2)
gasifier. The first quotation involved a gasifier feeding 283.4 tons/
day of Illinois No. 6 coal and producing 9,142 scfm of H + CO + H S.
However, this feed rate is less than the capacity of such a gasifier,
which is about 440 tons/day of coal. For this coal, 440 tons/day would
correspond to 14,190 scfm of H + CO + H0S. The total investment for
6
this system was quoted as $13 x 10 . The level of accuracy of this
estimate was stated as ±25 percent. This investment includes the water
wash step but not the oxygen plant.
(1 3)
The second quotation * involved a gasifier feeding 289.9
tons/day of Elkhorn No. 3 coal and producing 12,243 scfm of H- + CO + H S,
This feed rate was stated to be 66 percent of the capacity of the gasifier
(approximately 440 tons/day coal). Thus, for this coal the capacity of
the gasifier corresponds to 18,550 scfm of H2 + CO + H_S. The total
investment for this system was as follows:
Component Cost, 10 $
Coal handling, pulverizing, drying, and feeding 2.5
Gasification, cooling, cleaning (water wash) 8.2
General items 3.4
14.1
-------
G-6
TABLE G-2. CAPITAL COSTS FOR WELLMAN-GALUSHA OXYGEN-BLOWN GASIFICATION SYSTEM
Capital Cost,
Component $
Site (allowance) 3,000
Coal handling 27,000
Gas producer building erected 410,000
Three 10' diameter gas producers 960,000
Heat recovery system 152,000
Tar removal system 29,000
Ductwork for clean up system 38,000
Compressor 144,000
Electrical 120,000
Painting (allowance) 36,000
Miscellaneous producer plant piping, valves,
insulation (allowance) 120,000
Gas distribution lines (1000 lin. ft, 24" diam) 576,000
Fire protection 57,000
Total ex engineering and startup 2,672,000
Engineering (10 percent) 267,000
Startup 20,000
Total 2,959,000
-------
G-7
In both the above quotations, the gas demand was less than the
output of the normal-size Koppers-Totzek two-headed gasifier and yet one
such gasifier was specified and costed as such. It may not be feasible
to scale down this gasifier far below the normal size. In this case, any
gas not needed for the reduction could be used to supply the energy needs
of the FGD process and/or could be fed to the boilers. This extra energy
is, of course, obtained at some cost.
(4)
A paper published by Koppers gives investments for systems
involving one or more four-headed Koppers-Totzek gasifiers. A four-headed
gasifier can process about 590 tons/day of coal. These costs are not
directly useful here because they include the H^S removal facilities, but
it. is instructive to note that the size exponent for the investments is
about 0.59. This is for scaling by using multiple gasifiers (based on 1
to 5 gasifiers) and, as mentioned above, does include the H_S removal
facilities.
Koppers-Totzek, Oxygen + Air Blown
A cost estimate was obtained on a Koppers-Totzek gasification
system in which the gasifiers were to be blown with a 60/40 mixture of
oxygen/air. The system was to feed 1,853 tons/day of coal and to
produce about 61,100 scfm of total gas, including about 48,400 scfm of
H,, + CO + H S. The June, 1974, investment estimate was $21.5 x 10 ,
excluding the oxygen plant.
Riley-Morgan, Oxygen Blown
A quotation was obtained from the Riley-Stoker Corporation on
a system involving three 10'6" diameter Riley-Morgan gasifiers. * The
system was designed to feed 39,000 Ib/hr of coal and to produce 21,250
sc.fm of dry gas (67,050 Ib/hr). Using the clean, dry gas composition
from Table 9-4 but adding 1.0 mole percent H S, the gas should contain
80.3 percent H2 + CO + H S. Thus, the production of H_ + CO + H.S is
' 17,060 scfm. The investment for the gasifiers and associated equipment
was quoted as $2.9 x 10 , but this does not include water washing of the gas.
-------
r Q
(j-O
Applied Technology Corporation, Air Blown
The Applied Technology Corporation (ATC) is marketing a two-stage
gasifier essentially equivalent to the units which Wellman-Incandescent,
Ltd., had previously offered. Many of the Wellman-Incandescent units are
in use today and many have been operating for over 20 years. A quotation
was obtained on a system involving four 12-foot diameter gasifiers and
having a total production capacity of 388 x 10 Btu/hr of hot, raw gas or
322 x 10 Btu/hr of cold, clean gas. The gas composition from the ATC
gasifier should be similar to that from the Woodall-Duckham gasifier, since
the latter is also a two-stage gasifier. A typical Woodall-Duckham
. . (8) .
composition is:
Mole Percent
Component (Clean, dry)
H2
CO
N2
High heating value =
176 Btu/scf.
The gas from the Wellman two-stage gasifier is stated to have a high
(9)
heating value of about 175 Btu/scf. Using this gas composition but
adding 0.6 mole percent H S, the production of H + CO + H S for this
plant is calculated to be 13,900 scfm. The investment quoted was $5.0 x
10 , as shown in Table G-3.
Wilputte, Air Blown
Quotations were obtained from the Wilputte Corporation on
systems involving 7 and 20 air-blown Wilputte gasifiers. ' Each
gasifier has a 10'4" inside diameter, feeds 30 tons/day of coal, and
produces 2500 scfm of gas. A typical clean, dry gas composition is:
-------
G-9
TABLE G-3. CAPITAL COSTS FOR APPLIED TECHNOLOGY CORPORATION GASIFICATION SYSTEM
Estimated Capital
Section Cost, 106$
Coal handling and storage (on ground) 1.2
(a)
Gasifiers and auxiliariesv ' 2.7
Gas cooling and clean.up 1.1
Total 5.0
(a) Includes small package steam boiler. Does not include land,
site preparation, and civil work for producer gas plants.
(b) Includes cooling towers and effluent water clean up.
Note: Above estimates include equipment, installation, buildings,
and engineering.
-------
G-10
Component Mole Percent
H2
CO
CO 5.9 High heating value
170 Btu/scf.
i. j. u
°2
N2
After adding 0.6 mole percent H S, this gas contains 39.7 mole percent H2 +
CO 4- H S. The investments for the two systems are shown in Table G-4.
Consensus
Investment. As shown in Figure G-l, there is a fair amount of
scatter in the investment data. If one were to draw a best fit line
through all the data the slope would be considerably higher than the range
of 0.6-0.7 which fits the size exponents of most chemical processing equip-
ment. For the two sources for which two sizes were quoted, the calculated
exponents are 0.71 (Wellman-Galusha) and 0.89 (Wilputte). As mentioned
previously, a Koppers publication indicates an exponent of 0.59 for
multiple gasifier systems. For correlating the data here, a size exponent
of 0.7 was selected in order to stay within the bulk of chemical equipment
data.
The two points for air-blown Wellman-Galusha systems and the
point for the Riley-Morgan system were ignored because these costs did not
include the gas-washing facilities. For the other seven points shown on
Figure G-l, the cost was calculated for 10,000 scfm of H2 + CO + H2S using
the 0.7 exponent. The average of these seven costs was $7.03 x 10 . Thus,
the investment for a gasification system can be represented by the equation
Investment, 10 $ = 7.03(10 4 scfm H + CO 4- H S)°'7.
It should be noted that 10,000 scfm of H- + CO + H2S is close to the
reductant requirement for a 1000-MW power plant under the assumptions of
this study.
-------
G-ll
TABLE G-4. CAPITAL COSTS FOR WILPUTTE GASIFICATION SYSTEMS
Coal Input Rate, tons /day
Dry Gas Production Rate, scfm
Production of H. + CO + H.S, scfm
Investment, 10 $
Coal and ash handling
Gasifiers and waste heat boilers
Water treatment facilities
Seven Gasifiers
210
17,500
6,950
1.3
2.8
1.0
Twenty Gasifiers
600
50,000
19,850
3.0
8.0
2.0
Total plant cost 5.1 13.0
-------
G-12
This analysis did not disclose any general difference between
the investments for air-blown and oxygen-blown systems, not considering
the oxygen plant. One might expect the investment for air-blown systems
to be higher because of the larger vessel sizes required, a difference
which the cost of the oxygen plant would tend to offset, but if this is
true it could not be detected in view of the scatter of these data. All
that could be justified by the data was a single equation.
Operating Cost. The only source of operating cost data is
that for the air-blown Wellman-Galusha system (Table G-l). These data do
not include the water washing step, so this will have to be treated
separately. Some of the other data sources gave operating cost data
but included the H?S removal operation in the data. The values from
Table G-l, exclusive of maintenance, were used after being adjusted to
the basis of this study. That is, the annual operating costs were adjusted
proportionally to the annual operating time and the production rate of
TT i /"*("} ITT o
Residual Oil Gasification
Cost data on the Shell Gasification Process (SGP) for residual
oil were taken from a paper by J. B. Plummer, et al, of the Shell
Development Company. The January, 1974, capital cost of an SGP
unit processing 10,000 bbl/stream day of residual oil was given as $31
million for gasification with air and $16 million for gasification with
oxygen. This includes the soot removal and recycle facilities but not
H-S removal. For gasification with air, the air compressor train is
included. For gasification with oxygen, the oxygen plant is not included.
These are installed process equipment costs including engineering and
contingency. Updating these costs to 1975 (Chemical Engineering Plant
Cost Index = 190) gives $39.1 x 10 for gasification with air and $20.2 x
10 for gasification with oxygen. The investment was taken as proportional
to the 0.7 power of the oil feed rate.
-------
G-13
Typical gas compositions and process requirements for the Shell
process are given in Table G-5.
Oxygen Plants
The costs for oxygen plants were taken from a paper by J. T.
(i:
•
given.
(12)
Hugill. The following investments and electricity consumptions were
Oxygen Production Rate (T/D)
400 600 800
1974 Investment (103$)
Basic equipment 4.3 5.6 6.9
Erection 2.1 2.9 3.5
Total 6.4 8.5 10.4
Electricity Consumption
(kwh/T 02) 350 343 330
These investments indicate a size exponent of 0.70. After updating to
1975 (Chemical Engineering Plant Cost Index = 190), these investments can
be represented by the equation
Investment, 106$ = 0.1109(T/D 02)0'7.
In Figure G-2 these investments are compared with quotations obtained
from the Koppers Company ' and Mason and Hangar as adjuncts to the
gasification system costs previously discussed.
By plotting the electricity consumptions above, a value of
351 kwh/ton of 0_ was obtained for the range of oxygen plant sizes of
interest in this study (150-300 T/D). These values apply for the produc-
tion of 98 percent pure oxygen at slightly above atmospheric pressure.
According to Hugill, the 1974 cost of operating labor for 400-
800 T/D 0 plants is $220,000 per year, this corresponds to about three
operators per shift.
(13)
According to Shreve the total cooling water requirement is
8,640 gal/ton of 0^. In a recirculating system with cooling towers only
about 5 percent of this quantity would be required as makeup. According to
Shreve, oxygen plants require 12.5 pounds of caustic soda (NaOH) per ton of
O.,., The current cost of NaOH is about $159/ton ($150/ton of 73 percent Na20)
-------
G-14
TABLE G-5. DATA ON SHELL RESIDUAL OIL GASIFICATION PROCESS
Gasification Agent
Gas Composition, mole % dry
H2
CO
co2
CH,
4
N2
Ar
Lower Heating Value of Gas,
6
Requirements Per 10 Btu of
Item
Residual oil
Naphtha
Steam
1350 psia
1275 psia, 925 F
65 psia
Electricity
Boiler feed water / N
( 3 )
Total cooling water
scfm H + CO for 10,000 B/D
Oil L
basis
Btu/scf
Gas LHV
Unit
bbl
bbl
Ib
kwh
gal
gal
Residual
Oxygen
47.1
50.8
1.2
0.6
0.1
0.2
100.0
298.6
0.200
0.002
-81
122
5.1
2.80
0
1,590
113,800
Air
14.5
23.6
1.0
0.2
60.0
0.7
100.0
116.6
0.235
0.002
-335
224
2.5
1.96
13.9
1,860
96,600
(a) In a recirculating system only about 5 percent of this
quantity will be required as makeup.
Source: Plummer, J. B., et al (Reference 11).
-------
100
o
21
II
x
•o
c
G-15
0.
UJ
O
01
o
13
U3
O
c
0)
cu
10
Investment, 10s dollars
= 0.1109 (tons/day 02)a7
O Hugill
E3 Koppers Co.
" Mason & Hanger
J I L
I 1 I
100
1000
Oxygen, tons/day
FIGURE G-2. INVESTMENT FOR OXYGEN PLANTS
-------
G-16
Glaus Sulfur Plants
Data on the investment for vapor-phase Glaus sulfur plants are
shown in Figure G-3. Probably the best source is a pair of quotations
recently obtained from Ford, Bacon, and Davis, a leading vendor of Glaus
plants. The solid line on Figure G—3 is based on these data and is
expressed by the equation
Investment (1Q6$) = 0.1354 (LT/D Glaus sulfur)0'6.
This is for a three-stage Glaus plant with an inlet concentration of H.S
SO. of about 75 mole percent. The sulfur recovery in such a plant is
95.5 to 96 percent.
Ford, Bacon, and Davis indicated that the investment for a
two-stage plant is 8 to 10 percent less than that of a three-stage plant.
The data of Nelson indicate that the investment for a one-stage plant is
about 36 percent less than that of a three-stage plant.
The effect of the inlet gas concentration on the investment
for Glaus plants does not appear to be available in the open literature.
(14)
However, information obtained from the Allied Chemical Corporation
states that "Sizing of equipment in a Glaus unit is primarily a function
.of gas volume. Glaus converter sizing (packed bed reactor) is based on
contact time and, thus, gas velocity and volume. The major portion of
the heat load in the sulfur condensers is sensible heat of the gas and
thus a volume flow function. Incinerator design is based on retention
time and thus a volume flow function." It would, therefore, appear to
be reasonable to assume that the investment varies with the 0.6 power of
the inlet gas volume.
Based on information from vendors, the operating requirements
for Glaus plants include the following:
Item Amount
Electricity 50 kwh/LT sulfur
Boiler feed water 800 gal/LT sulfur
Activated alumina 0.4 Ib/LT sulfur (Cost = 14c/lb)
Steam produced 600 Ib/LT sulfur.
-------
G-17
10.0
C)
II
X
c:
•*-
CO
O
(»
in
en
0.1
Base point
Investment, 10 dollars
= O.I354(LT/D)a6
One stage
75
'Data are for 3-stage plant except as noted.
Numbers by points are inlet mole percent H.S + S0_.
Sources:
O Ford, Bacon, and Davis quotations (1975).
Q] Quotation from another vendor (1975).
Nelson, W. L., The Oil and Gas Journal, p 120,
March 18, 1974.
I I i
100
1000
LT/D Clous Sulfur
FIGURE G-3. INVESTMENT FOR GLAUS SULFUR PLANTS
-------
G-13
Allied Chemical Reduction Process
For the reduction systems for Type A FGD processes, the Allied
Chemical S0~ Reduction process was used as a base case for economic
comparisons.
The capital cost for the Allied process is primarily a function
of the gas volume processed and secondarily of the inlet S0? concentration
(i.e., the amount of sulfur processed). Based on a plot contained in a
paper by W. D. Hunter of Allied, the following equation was derived
for the on site (battery limits) investment:
6 R
Investment, (10 $) = A (megawatts capacity)
where the constants A and B are shown as functions of the inlet S0? concen-
tration in Figure G-4. In spite of the variation of the constants A and
B with inlet S0« concentration, the above equation should be considered
valid only for coal containing about 3.5 weight percent sulfur, since this
is the basis for the data on which it is based. The constant A has been
updated to 1975 using the Chemical Engineering Plant Cost Index.
An important component of the operating cost for the process is
the natural gas used as the reductant. The natural gas consumption varies
directly with the quantity of S0_ reduced. Based on a figure presented by
Hunter, the natural gas requirement for a 1000-MW power plant burning 3.5
weight percent S coal was calculated to be 4.95 x 10 scf/day.
The other components of the operating cost were estimated from
published information ' and process calculations.
Ductwork Costs
Much of the ductwork involved in the reduction system for Type
A FGD processes is included in the cost estimate for the gasification
system and the Glaus plant. However, there is some additional ductwork
connecting these major processing steps whicn must be considered separately.
It is very difficult to estimate these additional ductwork costs because
they are in reality very situation-specific. However, even a rough estimate
is better than none at all.
-------
G-19
0.7r-
0.6
<
c
O
O
0.5
0.4
0.3
20
1975 onsite investment (10s dollars)
= A (megawatts capacity)8
C.E. Plant Cost Index = 190
Coal sulfur content = 3.5 wt %
40 60 80 100
nlet Mole Percent S02 (Dry Basis)
0.481-
CD
^»
c.
a
.*-
«n
c
O
O
0.46
0.44
0.42
20 40 60 80 100
Inlet- Mole Percent S02 (Dry Basis)
FIGURE G-4. CONSTANTS IN INVESTMENT FUNCTION, ALLIED S02 REDUCTION PROCESS
-------
G-20
The materials and labor costs for ductwork were taken from
( T Q\
Mills. These costs are expressed in terms of the equivalent diameter
of the duct, even though such ducts are often rectangular in cross section.
After updating to 1975 using the Marshall and Stevens Index, the materials
cost for carbon steel ducts is
0 812
Materials cost, $/ft = 14.49 (diameter, feet)
At a total labor cost of $10/hour, the labor cost is
Labor cost, $/connection = 41.29 (diameter, feet)
For a linear gas velocity of 20 ft/sec at atmospheric pressure, the equiva-
lent diameter is
Diameter, feet = 0.003495 \J(Flow rate, Ib moles/hr) (Temp, R) .
The calculations made using these equations are summarized in
Tables G-6 and G-7. Note that only the streams which are not part of the
gasification system or the Glaus plant are included here. These estimates
must be regarded as order-of-magnitude costs only. Their magnitude is such
that they have little effect on the total system cost anyway.
Blower Costs
The reduction system for Type A FGD processes requires several
blowers which are not included in the cost estimates for the gasification
system and the Glaus plant. The estimation of the capital cost of these
blowers is shown in Table G~8. The equipment cost, including the blower,
(18)
motor, and drive, was taken from Mills. To obtain the total capital
cost, the equipment cost was multiplied by a "module factor" from Guthrie,
which reflects the cost of installation, foundations, electrical and piping
hookups, and other facilities associated with the basic piece of equipment.
-------
G-21
TABLE G-6. ESTIMATION OF INCREMENTAL DUCTWORK COSTS FOR AN OXYGEN-BLOWN GASIFIER
Unit costs ($/ft and $/connection) are for uninsulated carbon steel ducts
Values are for 1QOO--MW plant except as indicated
Stream
Number(a'
5
7
8
9
10
11
Total Cost,
1000 MW
500 MW
Temp,
R
540
1110
540
1010
540
760
ioVb)
Diameter,
ft
0.56
5.60
2.49
5.54
2.14
5.33
85% SO, Case
Material,
$/ft
9.1
58.7
30.4
58.2
26.9
56.4
16.1
12.2
Labor,
$/conn
23
250
107
247
91
237
13.0
9.1
Diameter ,
ft
0.56
5.60
4.59
7.65
2.14
7.02
25% S02 Case
Material,
$/ft
9.1
58.7
49.9
75.6
26.9
70.5
19.6
14.8
Labor,
$/conn
23
250
203
346
91
316
16.6
11.6
(a) Refers to Figure 13-1 of the text.
(b) Includes 10 ft and 2 connections for stream 10, 50 ft and 8 connections for all
other streams.
NOTE: In total costs, values for high-temperature streams (>540 R) reflect 60 percent
increase in material cost for alloy steel and insulation, 100 percent increase
in labor cost for insulation.
-------
G-22
TABLE G-7. ESTIMATION OF INCREMENTAL DUCTWORK COSTS FOR AN AIR-BLOWN GASIFIER
Unit costs ($/ft and $/connection) are for uninsulated carbon steel ducts
Values are for 1000-MW plant except as indicated
Stream
Number^)
4
5
6
7
8
9
Total Cost,
1000 MW
500 MW
Temp,
R
540
1110
540
1010
540
760
ioVb)
Diameter,
ft
1.98
8.64
2.49
8.38
3.42
8.25
85% S09 Case
Material,
$/ft
25.2
83.5
30.4
81.4
39.3
80.4
21.8
16.5
Labor,
$/conn
84
393
107
381
149
375
19.7
13.7
Diameter,
ft
1.98
8.64
4.59
9.90
3.42
9.44
25% S00 Case
Material,
$/ft
25.2
83.5
49.9
93.2
39.3
89.7
24.5
18.5
Labor,
$/conn
84
393
203
453
149
431
22.5
15.7
(a) Refers to Figure 13-3 of the text.
(b) Includes 10 ft and 2 connections for streams 4 and 8, 50 ft and 8 connections for
all other streams.
NOTE: In total costs, values for high-temperature streams (>540 R) reflect 60 percent
increase in material cost for alloy steel and insulation, 100 percent increase
in labor cost for insulation.
-------
TABLE G-8. INVESTMENT AND POWER REQUIREMENTS FOR BLOWERS
1000-MW Power Plant
3 (a)
Equipment Cost, 10 $
Air to reheat burner
SO.-rich stream
Cleaned reducing gas
Air to tail gas incin.
Total
Total Installed Cost of
Blower & Associated
Facilities, 103$(c>
„ (d)
Horsepower
Air to reheat burner
SO -rich stream
Cleaned reducing gas
Air to tail gas incin
Total
Oxygen-Blown
85% S02 Case
2.84
11.40
2.39
16.63
51.7
18
93
14
125
Gasifier Case
25% S02 Case
7.47
11.40
2.39
21.26
66.1
63
93
14
170
Air-Blown
85% S02 Case
2.18
2.84
26.81
4.58
36.41
•
113.2
12
18
223
35
288
Gasifier Case
25% SO Case
2.18
7.47
26.81
4.58
41.04
127.6
12
63
223
35
333
(a) From Mills (Reference 18).
(b) Two blowers for oxygen-blown case, four blowers for air-blown case.
(c) Based on module factor of 3.108 for compressors from Guthrie (Reference 19).
(d) Based on a pressure drop of 0.4 psi.
?
Ni
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G-24
H?S Generation Facilities
The costs for the H S generation facility required for the Type
B FGD processes were based on a quotation obtained from C & I/Girdler, Inc.,
for that purpose. This information is presented in Table G-9. For a
1000-MW power plant fired with 3.5 percent sulfur coal, a Type B FGD process
will require 725 tons/day of H_S. Girdler pointed out that they do not
have accurate cost data for such a large plant, since the largest ELS
plant they have built to date has a capacity of 60 tons/day.
The Girdler process is designed to utilize "bright" sulfur
and not "dark" sulfur. About 10 percent of the sulfur fed to the plant
must be blown down from the system. This sulfur is dark but is suitable
for use in the manufacture of sulfuric acid.
Reductant Gas Upgrading Facilities
Prior to its use for H«S generation, the coal gas must be
upgraded, through CO shifting and acid gas (CO,.,) removal, to increase
the concentration of hydrogen. The investments for these facilities were
(21)
based on data obtained from the U.S. Bureau of Mines on making hydrogen
from coal gas. These data were adjusted for differences in the capacity
and operating pressure. Table G-10 shows the resulting investments for
these facilities for a 1000-MW power plant. Each segment of the
processing sequence includes all the equipment, such as pumps and heat
exchangers, associated with the operation involved.
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G-25
TABLE G-9. DATA ON GIRDLER H S GENERATION PROCESS
Case A Case B
Plant Capacity 725 T/D 725 T/D
No. of Trains 3 4
Capacity/Train 242 T/D 181 T/D
Approximate Installed Cost $9,000,000 $9,700,000
Product Delivery Conditions
Pressure 80 psig
Temperature 54 C (130 F)
HO Same as H2 feed
Feed Gas Conditions
Pressure 140 psig
Temperature 38 C (100 F)
Feed Gas Analysis, mole percent
H 90.5
CO 1.91
CO 1.89
OT 0.12
"N, 1.03
AZ 0.46
H0 4.09
100.00
Approximate Product Analysis, mole percent
H, 2.0
ICS 88.5
CO 0.23
COS 1. 7
CO 1.9
on o.os
CS^ 0.04
N Z 1.0
AT 0.45
H2° 4.1
100.00
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G-26
TABLE G-9. (Continued)
Estimated Utility Requirements Quantity/Ton H S
Feed gas 25,000 scf
Bright sulfur (liquid at slight
positive head) 2,160 Ib
Electric power (440 V-3 PH-60 Cy) 15 kwh
Steam for jackets (40 psig) 700 Ib
Cooling water (32 C max temp) 5,500 gal
-------
G-27
TABLE G-10. INVESTMENT DATA ON REDUCING GAS UPGRADING FACILITIES
Facility
1975 Investment,
106$
Items Included
Shift converter 1
Waste hes.t recovery
Acid gas removal 1
(a)
Shift converter 2
Acid gas removal 2
0.295
1.885
5.047
0.466
1.773
Converter, feed gas heater, feed water pump
Waste heat boiler, boiler feed pump, cooler,
knockout drum
Absorber, regenerator, reboiler, solution
pumps, solution cooler, acid gas cooler
and knockout drum, clean gas cooler and
knockout drum, iron oxide tower, char
tower, solution storage and mixing tanks,
centrifuge, makeup pump
Converter, feed gas heater, feed water pump,
aftercooler, knockout drum
Absorber, regenerator, reboiler, solution
pumps, solution cooler, acid gas cooler
and knockout drum, clean gas cooler and
knockout drum, solution storage and
mixing tanks, centrifuge, makeup pump
Total
9.466
(a) Two trains.
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G-28
REFERENCES
(1) Cobb, R. W., et al., "Analysis of a Coal Gasification Facility and
Potential Gas Using Industries for Pike County, Kentucky", BCL report
to Appalachian Regional Commission, Commonwealth of Kentucky, and
Pike County Fiscal Court, September 1, 1975.
(2) Letter to Dr. Douglas W. Hissong (BCL) from Mr. James W. Bumbaugh
(Koppers Company, Inc.), March 6, 1975.
(3) Letter to Dr. Benjamin C. Hsieh (BCL) from Mr. James W. Bumbaugh
(Koppers Company, Inc.), May 30, 1975.
(4). Mitsak, D. M. and Kamody, J. F., "Koppers-Totzek: Take a Long Hard
Look", 2d Annual Symposium on Coal Gasification, Liquefaction, and
Utilization: Best Prospects for Commercialization, University of
Pittsburgh, August 5-7, 1975.
(5) Private communication.
(6) Letter to Mr. David A. Ball (BCL) from Mr. Thomas F. Walsh (Riley-
Stoker Corporation), May 20, 1975.
(7) Letter to Mr. David A."'Ball (BCL) from Mr. Sidney G. Nelson (Applied
Technology Corporation), June 6, 1975.
(8) Woodall-Duckhain (USA), "WD/GI Coal Gasification; Industrial Fuel Gas
from Coal", brochure.
(9) Gale, K., "Going Back to Gas with Modern Plant Ends Fuel Hangups",
The Engineer, July 11, 1974.
(10) Letter to Mr. David A. Ball (BCL) from Mr. George R. Cooper (Wilputte
Corporation), July 8, 1975.
(11) Plummer, J. B., et al, "The Generation of Clean Gaseous Fuels from
Petroleum Residues", AIChE Meeting, Tulsa, Oklahoma, March 11-13, 1974.
(12) Hugill, J. T., "Cost Factors in Oxygen Production", Efficient Use of
Fuels Conference, Chicago, Illinois, December 9-13, 1974.
(13) Shreve, R. N., Chemical Process Industries, 3d Edition, McGraw-Hill
Book Company, 1967.
(14) Letter from W. D. Hunter (Allied Chemical Company) to R. C. Christman
(EPA), June 2, 1975.
-------
G-29 and G-30
REFERENCES
(Continued)
(15) Hunter, W. D., "Application of SO- Reduction in Stack Gas Desulfuri-
zation Systems", Flue Gas Desulfurization Symposium, New Orleans,
May 14-17, 1973.
(16) Mann, E. L., "Power Plant Flue Gas Desulfurization by the Wellman-
Lord Process", Symposium on Flue Gas Desulfurization, Atlanta,
November 4-7, 1974.
(17) McGlamery, G. G. and Torstrick, R. L., "Cost Comparisons of Flue Gas
Desulfurization Systems", Symposium on Flue Gas Desulfurization,
Atlanta, November 4-17, 1974.
(18) Mills, H. E., "Costs of Process Equipment", Chemical Engineering,
p 133, March 16, 1964.
(19) Guthrie, K. M., "Capital Cost Estimating", Chemical Engineering, p 114,
March 24, 1969.
(20) Letter to Mr. Paul Choi (BCL) from Mr. Donald M. Hess (C & I/Girdler,
Inc.), November 20, 1975.
(21) Personal communication between BCL and the U.S. Bureau of Mines,
Morgantown, West Virginia, 1975.
-------
APPENDIX H
CONVERSION FACTORS
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H-2
APPENDIX H
CONVERSION FACTORS
To Convert (English)
atmospheres
barrels (42 gal)
Btu
Btu/ft3
Btu/lb
ft
3
ft (cubic feet)
gallons
grains
1C 3
grains/ft
in.
in. H20
long ton (2240 Ib)
pounds
pound-moles
2
pounds/in. (psi)
ton, short (2000 Ib)
To (Metric)
kilograms/m
liters
kilocalories
kilocalories/m
calories/gram
meters
3
m (cubic meters)
liters
grams
grams/m
cm
kilograms/m
metric ton (2205 Ib)
grams
gram-moles
kilograms/m
metric ton (2205 Ib)
Multiply By
10,333
159.0
0.2530
8.935
0.5578
0.3048
0.02832
3.785
0.06480
2.288
2.540
25.40
1.016
453.6
453.6
703.1
0.9070
Temperature in C =
- (Temperature in F) - 32
1.8
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TECHNICAL REPORT DATA
(Please read Instructions on the reverse before completing)
1. REPORT NO.
EPA- 600/2-76-130
2.
3. RECIPIENT'S ACCESSION1 NO.
4. TITLE AND SUBTITLE
Reductant Gases for Flue Gas Desulfurization
Systems
5. REPORT DATE
May 1976
6. PERFORMING ORGANIZATION CODE
7.AUTHOR Hissong, K.S. Murthy, and
A.W. Lemmon, Jr.
8. PERFORMING ORGANIZATION REPORT NO
9. PERFORMING ORGANIZATION NAME AND ADDRESS
Battelle-Columbus Laboratories
505 King Avenue
Columbus , Ohi o 43201
10. PROGRAM ELEMENT NO.
1NB458; ROAP 21BJV-038
11. CONTRACT/GRANT NO.
68-02-1323, Tasks 21 and 36
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 PERIOD COVERED
Task Final; 5/74-3/76
14. SPONSORING AGENCY CODE
EPA-ORD
15. SUPPLEMENTARY NOTES
Ext 2915.
Off i C6r fOI t MS
is C. J. Chatlynne, Mail Drop 61,
16. ABSTRACT
rep0rj- gj ves results of SL study of the use of coal or residual-oil
gasification to produce a hydrogen/carbon monoxide-rich gas for use as a reduc-
tant for regenerable flue gas desulf urization (FGD) processes. Two different
reduction systems are considered: one for the type of FGD process that produces
a concentrated S02 stream; the other, for the type that uses a liquid-phase Glaus
reactor. Detailed data on the composition of the raw gas from several gasifiers are
analyzed. To supplement the data on trace constituents in the gas, thermodynamic
calculations were made to determine the equilibrium gas -phase concentrations for
a typical coal and typical gasification conditions. Mass transfer was calculated to
determine the extent to which certain gaseous species could be removed by water-
washing of the gas. The effects of the remaining trace constituents on the compo-
nents of the reduction systems are analyzed. Additional research on trace constit-
uents and their effects are recommended. The capital and operating costs for
reductant systems based on gasification of coal and residual oil are estimated and
compared with those for reduction systems based on natural gas.
7.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b.lDENTIFIERS/OPEN ENDED TERMS
c. COSATI Field/Group
Air Pollution
Flue Gases
Desulf urization
Reduction (Chemistry)
Coal Gasification
Gasification
Residual Oils
Natural Gas
Air Pollution Control
Stationary Sources
Reductant Gases
Regenerable Processes
13B
2 IB
07A,07D
07B,07C
13H
2 ID
8. DISTRIBUTION STATEMENT
Unlimited
19. SECURITY CLASS.(This ReportJ _
Unclassified
21_._NO. OF PAGES
253
20. SECURITY CLASS (This page)
Unclassified
22. PRICE
EPA Form 2:!20-1 <9-73)
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