EPA-600/2-75-078
November 1975
Environmental Protection Technology Series
FUEL GAS ENVIRONMENTAL IMPACT:
Phase Report
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 categories were established to
facilitate further development and application of environmental
technology. Elimination of traditional grouping was consciously
planned to foster technology transfer and a maximum interface in
related fields. The five series are:
1. Environmental Health Effects Research
2. Environmental Protection Technology
3. Ecological Research
4. Environmental Monitoring
5. Socioeconomic Environmental Studies
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 views and policies of the Agency, nor
does mention of trade names or commercial products constitute endorse-
ment or recommendation for use.
This document is available to the public through the National
Technical Information Service, Springfield, Virginia 22161.
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EPA-600/2-75-078
FUEL GAS
ENVIRONMENTAL IMPACT:
PHASE REPORT
by
F.L. Robson, A.J. Giramonti, andW.A. Blecher
United Technologies Research Center
Gerald Mazzella
Foster Wheeler Energy Corporation
United Technologies Research Center
400 Main Street
East Hartford, Connecticut 06108
Contract No. 68-02-1099
ROAP No. 21ADD-027 and -104
Program Element No. 1AB013
EPA Project Officer: William J. Rhodes
Industrial Environmental Research Laboratory
Office of Energy, Minerals, and Industry
Research Triangle Park, NC 27711
Prepared for
U.S. ENVIRONMENTAL PROTECTION AGENCY
Office of Research and Development
Washington, DC 20460
November 1975
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ABSTRACT
A program was carried out to evaluate the technical and economic
feasibility of: (l) Lurgi-type fixed-bed gasifiers and BCR-type entrained-
flow gasifiers in combination with low- and high-temperature fuel gas
cleanup systems, (2) advanced technology combined-cycle power systems, (3)
integrated gasification systems, cleanup processes and power systems. Pro-
cesses and systems considered were those using technology currently avail-
able for power station configurations which the Contractor judged could
appear in commercial applications in the 1975-1978 time frame (first-gener-
ration systems) and those using technology potentially applicable in the
1980-decade time period (second-generation systems). The objective of this
analytical study of fuel gas environmental impact is the definition of
combinations of: (l) fossil fuel gasification systems, (2) low- and high-
temperature fuel gas cleanup processes, and (3) advanced-cycle power
systems for central power stations that appear to result in the lowest
practicable emissions of air, water, and solid pollutants consistent with
the environmental constraints, while producing low-cost electrical power.
The method of analysis is based upon the systems approach in which the
technical and economic characteristics of the overall integrated gasifica-
tion, cleanup and power system are evaluated as a whole. A Contractor-owned
digital computer program was utilized to define the performance of the
system from coal in to kilowatts out. The modular approach to analysis by
this unique analytical tool permits wide flexibility in fuel process config-
urations and power cycle arrangement. However, lack of substantial data on
gasifier operation limited the approach to design point calculations.
The analyses indicate that high-temperature cleanup systems have the
potential of improving the efficiency and reducing the capital costs of in-
tegrated gasification systems. However, unacceptable emission levels for
NOX could result with some gasifier types. The use of commercially avail-
able low-temperature cleanup systems results in lower efficiencies and
higher costs, but would still allow generation of electrical power at costs
competitive with conventional steam stations with stack gas cleanup while
having sulfur emissions which are below regulations for conventional stations,
iii
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This report was submitted in partial fulfillment of Project Nos. 21
ADD 27 and 10U, Contract No. 68-02-1099, by the Research Center of the United
Technologies Corporation, under the sponsorship of the Environmental Protec-
tion Agency. This report covers work performed between July 1, 1973 and
November 1,
iv
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TABLE OF CONTENTS
CONCLUSIONS
RECOMMENDATIONS ................ . ............... 3
INTRODUCTION .................................. ^
SECTION 1
GASIFIER & CLEANUP • SYSTEM PROCESS REVIEW .................... 7
COAL GASIFICATION SYSTEMS .............. . . . . . ..... . 7
Gasifier Operating Data ........................ IQ
Bureau of Mines Stirred-Bed Gasifier ............... 10
BCR Entrained Flow Gasifier .................... 15
FUEL GAS CLEANUP PROCESSES ........................ 17-
Low-Temperature Cleanup Systems .................... 17
Chemical Solvant Processes .................... "19
Physical Solvent Processes .................... 26
Direct Conversion Processes .................... 27
Dry Bed Processes ......................... 28
Selection Considerations. . '. ............... ; ..... 29
High-Temperature Cleanup Systems .................... 59
Bureau of Mines Process ...................... 20
Consolidation Coal Process .................... 22
Air Products Process ....................... J^
IGT Meissner Process ........ • ............. . . . [^
Battelle Northwest Process .................... I|_2
Selection Consideration ........................ [
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TABLE OF CONTENTS (Cont'd)
SECTION 2
PARAMETRIC ANALYSIS OF POWER SYSTEMS 1+7
DIRECT COAL-FIRED STEAM CYCLES l±Q
Design Considerations 51
Unit Size 51
Steam Conditions 51
Exhaust Conditions 52
Environmental Considerations 52
Auxiliary Power 5!).
GASIFIED COAL FIRED STEAM SYSTEMS 5!+
System Configuration. 56
System Performance 56
ADVANCED COGAS SYSTEMS 60
Gas Turbine Cycle 62
Turbine Inlet Temperature. 62
Steam Systems 69
COGAS Performance 71
Integrated System Performance Evaluation ............ 78
Integrated System Studies -82
Boiler Firing Studies 83
Steam Conditions • 83
Fuel-Gas Temperature Effects .................. 87
Gasifier Heat Balance 91
Effect of Pressure Rato on Power Cost ......... 91
Boost Compressor Location 96
vi
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TABLE OF CONTENTS (Cont'd)
Parametric Analysis of Nitrogen Oxide Emissions From Gas Turbine Power
Systems Burning Low-Btu Gas • 96
Status of NOV Pollution Modeling 98
J\i
Nitric Oxide Emissions From Gas Turbine Combustors . . 99
Fuel Nitrogen 103
Discussion of Results of NOX Modeling 105
SECTION 3
COMPARISON AND'SELECTION OF CLEANUP SYSTEMS 10?
COMPARISON OF CLEANUP' SYSTEMS. . . . '. 107-
Low-Temperature Systems 107
High-Temperature Systems 109
Preliminary Performance of Integrated Cleanup Systems 113
Integrated Low-Temperature Cleanup Systems 113
Integrated High-Temperature Cleanup Systems. . .' 12 U
SELECTION OF STANDARD INTEGRATED SYSTEMS ,
First-Generation Standard Systems
Bureau of Mines/Selexol System-Process Description
Bureau of Mines/Iron Oxide System Process Description 157
Second-Generation Standard Systems. 158
BCR/Selexol System-Process Description • '158
BCR/Consol System-Process Description 176
Utilities and Energy Balances 178
Investment Cost 178
Comparison of Standard Integrated Systems , 186
VI1
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TABLE OF CONTENTS (Cont'd)
SECTION U
INTEGRATION OF GASIFICATION CLEANUP AND POWER SYSTEMS 188
STEAM SYSTEMS ." 190
Performance Calculations 192
Gasifier Performance 1914.
Cleanup System Performance 195
Steam Cycle Performance 196
COMBINED CYCLE SYSTEMS 197
System Descriptions 202
Bureau of Mines/Selexol 202
Bureau of Mines/Sintered Iron Oxide
BCR/Selexol System
BCR/Consol System 20?
Performance. 207
COST COMPARISON 217
' Gasified Coal-COGAS Systems 223
Coal-Fired Steam Station 225
Gasified Coal-Fired Steam Plant 232
SECTION 5
ENVIRONMENTAL IMPACT AND GOALS 23^
IDENTIFICATION OF ENVIRONMENTAL IMPACT 23^
Air Emissions from Gasifier/Cleanup Systems 23^
Water Emissions from Gasifier/Cleanup Systems 237
Solid Wastes from'Gasifier/Cleanup Systems 21*2
VI11
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TABLE OF CONTENTS (Cont'd)
Air Emissions From Power Systems 242
Water Emissions From Power Systems 245
.DEFINITION OF CLEANUP SYSTEM GOALS 246
EPA Standards ' 2^
Gas Turbine Fuel Specifications 246
Particulate Loading 248
Suggested Low-Btu Fuel Gas Specification 251
COMPARISON OF CLEANUP SYSTEM CHARACTERISTICS 251-
Low-Temperature Systems 253
High-Temperature Systems 255
s
Comments on Cleanup System Selection 25o
TECHNOLOGY REQUIRED TO OBTAIN FUEL SPECIFICATION GOALS 257
Technology for Sulfur Removal 257
Technology for Fuel Nitrogen Removal • 257
Technology for Particulate Removal • . . 258
RECOMMENDATIONS 26l
Gasification 26l
Low-Temperature Cleanup Systems 263
High-Temperature Cleanup Systems 263
Advanced Power Systems 264
Environmental Impact 264
REFERENCES 266
APPENDICIES 273
A - Performance Evaluation Program 273
B - Gas Turbine Combustion Pollution Model 276
ix
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TABUE OF CONTENTS (Cont'd)
C - Power Plant Cost Analysis. 284'
D - Description of Gas Turbine Cost Model 290
E - COGAS Performance Evaluation 292
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LIST OF FIGURES
1. Bureau of Mines Stirred-Bed Gasifier ^
2. BCR Entrained-Flow Gasifier l6
3. Typical Low-Temperature Acid Gas Removal Unit 25
U. Equilibrium Constant For HoS Absorption by Iron Oxide 33
5. Equilibrium Constant for I^S Absorption by Iron Oxide 3^
6. Equilibrium Constant for Water Gas Shift 35
7. Equilibrium Constant for COS Absorption by Iron Oxide 36
8. Equilibrium Constant for H S Absorption by Half-Calcined Dolomite. 38
9. Dissociation Pressure For Calcium Carbonate ' 39
10. Equilibrium Constant For COS Absorption by Half-Calcined'Dolomite .^0
11. Equilibrium Constant For HoS Absorption by Calcined Dolomite ... ^2
12. Equilibrium Constant for COS Absorption by Calcined Dolomite ... W-
13. Basic Steam Cycle Flow Diagram 50
lU. Effect of Reheating 53
15. Gasified Coal-Steam Cycle Systems 57
16. Waste-Heat Combined Gas and Steam Turbine System 6l
17. Estimated Turbine Inlet Temperature Progression 63
18. Effect of Cycle Parameters on Turbine Cooling Air Flow Requirements 65
19. COGAS Station Performance 66
XI
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LIST OF FIGURES (Cont'd)
20. Gas Turbine Performance MAP 67
21. Performance of 2600 F Gas Turbine 68
22. Temperature Distribution in Waste Heat Recovery Boiler For Oil-Fired
COGAS Station 70'
23. COGAS Station Performance With Fired Boiler 72
2k. COGAS Station Performance With Conventional Air Cooling 73
25.. COGAS Station Performance With Precooled Air To Vanes 7^.
26. COGAS Station Performance With Ceramic Vanes and Conventional Cooled
Blades 75
27. COGAS Station Performance With Ceramic Vane and Blades 76
28. COGAS Station Performance With Zero Bleeds and Leakage 77
29, Integrated COGAS/Coal Gasification Power Station .... 79
30. Integrated COGAS/Coal Gasification Power Station With Low Temperature
Cleanup Systems 80
31. Integrated BCR/Consol Gasification Power Station .... 8l
32. Effect of Steam Cycle on Station Performance 8k
33. Integrated COGAS/BCR Gasification Station Performance 86
3k. Performance Improvement Due to Elevated Fuel Supply Temperature. . 88
35. Effect of Fuel Temperature 89
36. Capital Cost Equivalent of Plant Efficiency ..... 92
37. Effect of Pressure Ratio on Power Cost 95
38. Integrated COGAS/Oil Gasification Station Performance 97
XI1
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LIST OF FIGURES (Cont'd)
39. NC) in Gas Turbine Combustor . . . 101
X
UO. WOX vs. Axial Distance 102
Ul. NO Production vs. T^,.-, at1 Various Pressures for Coal Gas From Four
U.C_L .
Gasifier Cleanup Systems - 10^
k2. Integrated Low-Temperature Cleanup System 110
1+3. Integrated Consol High-Temperature Cleanup System 125
kk. Integrated Calcined Dolomite High-Temperature Cleanup System . . . 126
^5. Integrated Bureau of Mines High-Temperature Cleanup System .... 127
14-6. Integrated Battelle Molten Salt High-Temperature Cleanup System. . 128
V7. Sulfur Absorption by Half Calcined Dolomite 13U
U8. Sulfur Absorption by Calcined Dolomite 136
149. Process Flow Diagram-Bureau of Mines/Selexol System
50. Process Flow Diagram-Bureau of Mines/Sintered Iron Oxide System. .
51. Process Flow Diagram Sulfur Recovery Unit Glaus Process 156
52. Process Flow Diagram BCR/Selexol System 159
53. Process Flow Diagram BCR/Consol System .160
5^. BCR Consol/Steam System 198
55. Gasified Coal-Steam System 199
56. Bu Mines/Selexol System. 203
57. Bu Mines/Sintered Iron Oxide ; 205
58. BCR/Selexol System 206
Xlll
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LIST OF FIGURES (Cont'd)
59. BCR/Consol System 208
60. BCR Gasifier-Selexol Cleanup Performance Characteristics 209
6l. Performance Estimates for Second-Generation Turbine Systems. . . . 210
62. Performance Estimates for First-Generation Turbine Systems .... 211
63. Combined Cycle Representation 213
6k. Cost Summary 218
65. Effect of Output Power on Gas Turbine Price 22k
66. S02 Emissions as a Function of HpS in Fuel Gas 2kb
67. Effect of Particle Size on Engine Lifetime 250
68. Operation of SOAPP System 27^
69. Swirl-Stabilized Can Flow Field Model. ... 278
70. Primary-Zone Flow Pattern Observed in a Can-Type Combustor .... 279
xiv
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LIST OF TABLES
1. Comparison of Gasification Systems ................ 9
2. Bureau of Mines Gasifier-Gasification of Kentucky Coal ...... 13
3. Bu Mines Gasifier-Gasification of Kentucky Coal with Tar Recycle . Ik
k. BCR Entrained-Flow Gasifier-Gasification of Illinois No. 6 Coal. . 18
5. Low-Temperature Desulfurization Processes ............. 20
6. High-Temperature Desulfurization Processes ............ 31
7. Characteristics of Basic Coal-Fired Steam Power Plant ...... '. ^9
8. Auxiliary Power Requirements ..... .............. 55
9. Steam Power Station Efficiency Comparison ............. 58
10. Design Characteristics for Reference Integrated COGAS Power Station °5
11. BCR-Selexol Fuel Gas Regenerator Costs .............. 90
12. Effect of Steam Extraction for Gasifier on Station Performance . . 93
13. Characteristics of Oil Gasification System ............ 9^
Ik. Low-Btu Gas Characteristics at Design Points- for Four Gasifier-Cleanup
Systems .............................. 10°
15. Comparison of Selected Low-Temperature Cleanup Processes .....
16. Comparison of High-Temperature Cleanup Processes ....... . . H2
17. Material Balance for BCR Gasifier ...... ........... H1*
18. Material Balance for Selexol Cleanup System ............ 117
19. Material Balance for Benfield Cleanup System ........... 118
xv
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20. Material Balance for Rectisol Cleanup System 120
21. Summary of Low-Temperature Integrated Systems 123
22. Material Balance for Consol Cleanup System 129
23. Material Balance for Air Products Cleanup System 130
2k. Material Balance for Bureau of Mines Cleanup System 132
25. Material Balance for Batelle Molten Salt Cleanup System 133
26. Summary of High-Temperature Integrated Systems .......... 139
27. Bureau of Mines/Selexol System Equipment List
28. Material Balance for Bureau of Mines/Selexol System
29. Bureau of Mines/Iron Oxide Equipment List 152
30. Material Balance for Bureau of Mines/Iron Oxide System 153
31. BCR/Selexol System Equipment List l6l
32. Material Balance for BCR/Selexol System 16U
33. BCR/Consol Equipment List 168
3^. Material Balance for BCR/Consol System 172
35. Summary of Bureau of Mines/Selexol Desulfurization Utilities
Consumption 179
36. Summary of Bureau of Mines/Sintered Iron Oxide Utilities
Consumption 180
37. Summary of BCR Gasification/Selexol Desulfurization Utilities
Consumption l8l
38. Summary of BCR Gasification/Consol Desulfurization Utilities
Consumption. . 182
xvi
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39« Gasification/Cleanup Systems Energy Balance. ... 183
kO. Summary of Integrated Gasification Cleanup System Investment Costs 185
1+1. Comparison of Gas if ier/Cleanup System Performance 18?
b2. Gasified Coal-Steam Cycle Performance Summary 191
it-3. Gasified Coal-Steam Cycle Component Performance Characteristics. . 193
MK BCR/Consol/Steam Cycle Performance ...... •. 200
if5. Integrated System Performance Summary. . . 201
h6. Combined-Cycle Energy Utilization Characteristics . 2lU
k7. Effect of Combined Cycle on Gasifier and Cleanup System. ...... 215
14-8. Integrated System Cost Summary ' 220
it-9. Cost Summary for Coal-Fired Steam With Stack Gas Cleanup . . . . . 221
50. Cost Summary for BCR Consol Steam-Cycle Power Plant 222
51. Power System Cost Detail 226
52. Effect of Location on Power System Cost 228
53. COGAS Power System Cost Summary 229
5^. Gasifier and Cleanup System Capital Cost Breakdown 230
55- Cost Summary - Coal-Fired Steam Plant 231
56. BCR/Consol/Steam Plant Cost Summary 233
57. Estimated Environmental Impact of Integrated Power System 235
58. Gasifier/Cleanup System Effluent Summary 236
59- Air Emissions From Gasification/Cleanup Systems 238
xvi i
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60. Waste-Water Stream Analysis 239
6l. Trace Element Analysis of Illinois Coal 2hO
62. Potential Waste-Water Stream-Analysis for Power Gasifier 2*4-1
63. Power System Emissions 2^3
6U. Comparison of Emissions for Gas Turbine Systems 2^7
65. Gas Turbine Fuel Specification 2^9
66. Suggested Low-Btu Fuel Gas Cleanup System 252
67. Comparison of Cleanup System Characteristics ... 25^
68. Size Distribution of Flyash From Boilers '........ 260
69. Nitric Oxide Formation Kinetics Relationships 283
70. Specific Components Whose Costs Are Calculated in Power Station
Analysis 286
71. Illustrative Example of Electric Power Cost Estimating 289
xvi 11
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ACKNOWLEDGEMENTS
The support work of Mr. W. R. Davison and Dr. E. B. Smith of United
Technologies Research Center is acknowledged with sincere thanks.
The efforts of the Foster Wheeler Energy Corporation carried out under
the direction of Mr. Gerald Mazzella provided both excellent technical
information and considerable insight into the•interface problems associated
with mating power systems with complex chemical plants.
The contributions and technical direction offered by Messrs. W. J.
Rhodes and T. K. Janes of the Industrial Environmental Laboratory, EPA, are
acknowledged with deep appreciation.
xix
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LIST OF CONVERSION FACTORS
Btu x 0.252 = Kcal
ft x 0.30U8 = m
in. x 25.^ = mm
F subtract 32 x 0.555 = C
Ib x O.U53 = Kg .
Ib x 0.U53 = Kg .
scf (@ 60 F & 30 in. Hg) x 0.028U = m3 (@ 15.5 C & 762 mm Hg)
Btu/scf x 8.88 = Kcal/m3 (@ 15-5 C & 762 mm Hg)
lb/106 Btu x 1.798 = Kg/10 Kcal
ton x 1.10^ = metric ton
xx
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CONCLUSIONS
1. First-generation coal gasifiers and commercially available low-temperature
cleanup processes used in conjunction with combined-cycle power systems
have the potential of producing electrical power at costs competitive with
conventional coal-fired steam plants with stack gas cleanup. Thes.e,inte-
grated gasification/power systems would also have lower emissions of SOo,
NOV, and particulates, reduced thermal pollution, and would provide recovered
.A.
sulfur in the elemental form.
2. High-temperature cleanup processes used in conjunction with combined-cycle
power systems have the potential of providing significant advances in
efficiency and lower capital costs which could result in the production of
• electrical power at costs 20% or more lower than conventional coal-fired
steam plants with stack gas cleanup.
3. Emissions of N02 from power systems using high-temperature cleanup processes
could be unacceptable. These higher NOV emissions result from: (a) higher
J\.
combustion temperatures due to the higher temperature fuel; and (b) fuel-
bound nitrogen compounds formed in some types of gasifiers which would
pass directly through the high-temperature cleanup process. It may be possi-
ble to mitigate the NO formation due to temperature effects by careful
J\.
combustor design or modifications. At this time, there appears no accept-
able method of removing fuel-bound nitrogen at high temperature.
h. Gasifiers operating at temperatures beyond those considered in this study
(l800 F) would tend to break down fuel-bound nitrogen to N2 and IL, and thus
could reduce the fuel-bound nitrogen carryover and subsequently the NO
JC
emissions from the power system.
5. Overall system performance improvements resulting from advancing gas turbine
inlet temperatures HOO F (second-generation power systems) are comparable
to those improvements associated with the change from low-temperature to
high-temperature cleanup processes.
6. Steam power systems using gasification/cleanup processes appear competitive
with conventional coal-fired steam systems with stack gas cleanup only when
second-generation gasifiers and high-temperature cleanup processes are used
in conjunction with a let-down turbine to recover compressor power. These
systems would be subject to NOX emission problems similar to those of the
combined-cycle system.
T. Particulate carryover from the gasifier may pass through the high-temperature
cleanup process with little or no reduction. Thus, both the environment and
the turbine may suffer. However, some high-temperature particulate removal
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devices, such as the Aerodyne-type high energy cyclone, appear to be nearer
commercialization than do the cleanup processes.
8. The high-temperature iron oxide process, which is nearer commercialization
than other high-temperature processes, requires a larger amount of clean
fuel to convert the SC>2 to elemental sulfur. Other forms of ultimate
sulfur recovery would not be as energy intensive.
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RECOMMENDATIONS
The overall attractiveness of the integrated gasification/cleanup/combined-
cycle systems even when using essentially present-day technologies lead to
the recommendation that this system should be demonstrated in a large enough
scale (*~100 Mw or larger) to obtain experience on systems which could be
commercialized at an early date. Prior to, or parallel with, this expensive
effort (about $300 million for two programs) several other less ambitious but
essential programs are recommended: :
1. The potential benefits of the high-temperature cleanup system are great enough
to warrant extensive development effort to bring these systems to commerciali-
zation-at an early date. -Programs -to verify physical configuration, materials
of construction, and sulfur capture ratios at pressures and temperatures typical
of combined-cycle operations (e.g. 300-^00 psi, 1500-1700 F) should be carried
out in a fuel gas stream, preferably from a gasifier or from .a synthesized gas
source. .
2. Methods of removing nitrogen compounds, mainly ammonia, from high-temperature
fuel gas streams should be investigated. In addition, studies should be initiated
on the mechanism of ammonia formation during gasification with the objective of
determining operating conditions consistent with minimum ammonia production.
Alternatively, gasifiers operating at temperatures at,which ammonia breaks
down to nitrogen and hydrogen should be considered.
3. In the off gas from a gasifier, the actual carryover of materials potentially
harmful to the gas turbine; e.g., particulates, alkiline metals, vanadium, etc.
should be measured. Only when these data are available can performance and
cost estimates be made with relative confidence.
^. Programs leading to the commercialization of high-temperature particulate
removal devices should be continued or initiated at the easrliest possible time.
Removal efficiencies of 99% down to 2y are estimated to be required for adequate
turbine life and to meet environmental constraints.
5. The present study was"constrained to specific gasifier operating points due
to a lack of adequate data. The development of a simple model of the gasifier
would allow preliminary estimates to be made of variations in overall system
performance due to changes in gasifier operation.
5. The potential emission of thermal NOX, (NO resulting from the combustion process,
not from fuel nitrogen) from high-temperature, high-pressure ratio gas turbines
should receive further consideration. The preliminary model of NOX formation
presented in this study should be further developed, if possible, to take into
consideration more extensive current and future data on the combustion of low-
Btu fuel in gas turbine combustors.
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INTRODUCTION
There are many possible coal gasification/advanced-cycle power system
combinations that may be integrated into future power generating installations.
Before committing large resources to any specific combination, an assessment
must be made of the technical and economic characteristics and the environ-
mental intrusion from the coal conversion process, the fuel gas cleanup system,
and the power cycle. Towards this objective an analytical program was under-
taken by UTRC and the Foster Wheeler Energy Corporation to define combinations
of coal gasification systems, low- and high-temperature fuel gas cleanup pro-
cesses, and advanced power cycles that appear to result in the lowest practi-
cable emissions of air, water and solid pollutants while producing low-cost
electric power from central power stations.
When initiated, this study was a part of the EPA Control Systems Labora-
tory plan to evaluate methods of reducing or eliminating emissions from coal-
fired central generating stations. Additional impetus was given to the EPA
program by the 1973 October War and the resulting embargo on mid-east petroleum
imports. Subsequent examinations of this nation's energy supplies and identi-
fication of the goals of Project Independence have spotlighted coal as the
"fuel of the future." Thus, the prior work done by the EPA and its predeces-
sor, the National Air Pollution Control Administration (NAPCA) on pre- and post-
combustion cleanup of coal has attained even greater importance.
The energy crunch with its accompanying conservation measures has reduced
somewhat the projected growth of the electric utility industry. Nonetheless,
even conservative estimates(1) of the need for electrical energy in 1980 indi-
cate that about 625 GW will be required, of which ^25 GW will be fossil fueled.
By 1995, it is estimated that 1350 GW will be installed, 610 GW of which would
be fossil fueled.
With all the projected additional installed capacity of fossil-fueled
power systems, the already difficult problem of utility-caused sulfur dioxide
(302), particulate, and nitrogen oxides (NO) pollution could become intolerable.
There are three alternative solutions open for utility consideration: (1) use
of naturally occurring clean fuels, (2) treatment of the powerplant exhausts
to remove pollutants, and (3) treatment of dirty fuels before combustion to
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remove harmful pollutants. Because of the very high costs and demand for the
naturally occurring clean fuels by other segments of the economy having little
or no alternative, widespread use of such fuels will not be attractive to a
major portion of the utility industry. The second alternative, stack gas
cleanup, shows promise as a method of meeting the near-term (1975) EPA stan-
dards for SOo and these systems, some with the help of special particulate
removal devices, will meet the standards for particulate emission levels for
1975. Currently, however, no commercial-scale stack gas cleanup system has
been demonstrated that also removes significant amounts of NO , nor have re-
moval efficiencies much beyond those required for 1975 standards been demon-
strated conclusively. The third alternative, fuel pretreatment, can in several
of its forms such as hydrodesulfurization of residual oils or solution of coal,
meet near-term goals for sulfur and particulates. However, only gasification
followed by some type of sulfur cleanup process has the promise of attaining
sulfur levels significantly below those required by 1975 standards.
In order to offset some of the inefficiencies in converting coal to a
clean fuel, an advanced-cycle power system of high efficiency could be utilized.
The use of gasified coal.for fueling advanced-cycle power systems is not a new
idea. As far back as the mid-1960's Steinkohlen-Elektrizitat AG (STEAG) and
Lurgi Gesellschaft fur Warme und Chemotechnik mbH proposed^) a combined-cycle
plant using Lurgi gasifiers to fuel a pressurized boiler which would exhaust
into a gas turbine. The original concept did not involve a fuel gas desul-
furization step although one has subsequently been added. Widespread atten-
tion to the concept of the integrated gasifier/cleanup/power system was brought
to focus by the publication(3) in December 1970 of the United Aircraft Research
Laboratories/Burns and Roe, Inc./FMC report "Technological and Economic Feasi-
bility of Advanced Power Cycles and Methods of Producing Non-Polluting Fuels
for Utility Power Stations." The major conclusion of this report was that ad-
vanced technology available in the aircraft gas turbine industry could be ap-
plied to advanced-cycle combined Gas And ESteam (COGAS) power systems which
would allow the use of costlier gasified and desulfurized fuel in power sta-
tions which would generate essentially pollution-free power at costs competi-
tive with or less than conventional steam stations with no pollution control.
This study was based upon NAPCA-prescribed power systems of 1000-Mw nom-
inal size and technology judged best available, i.e., technology, which could
appear if given immediate and necessary research and development funding. In
the years following the publication of this report, it is apparent that the
gasified coal-fired COGAS system, while recognized as the most attractive
future method of generating clean power, has not received funding in the amounts
and within the time required to match the study projections. Additionally,
subsequent effortCM at UTRC has indicated that performance benefits might be
obtained through the use of high-temperature fuel gas cleanup systems, but the
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use of such systems could result in increased NO production over systems using
low-temperature fuel gas cleanup processes. The technology level used in these
later studies has been that judged to be attainable based upon currently avail-
able funding and also upon marketing considerations; i.e., projected acceptance
of COGAS systems and its effect on the development of large, advanced-technology,
industrial turbines.
In order to put into perspective the technical and economic characteristics
of these advanced power systems based upon the current assessment of technologi-
cal development and also to identify the potential environmental intrusion of
the systems, a study was conducted to evaluate combinations of gasification,
cleanup processes, and power cycle configurations capable of producing low-cost",
pollution free electric power from coal. The results of that study are pre-
sented in this report. Section 1 contains the definition of the operating
characteristics of a fixed-bed gasifier (U.S. Bureau of Mines) and a two-stage
entrained-flow gasifier (Bituminous Coal Research) and a review of low- and
high-temperature fuel gas cleanup processes. In Section 2, the operating
characteristics of advanced power systems that could be used with the gasifi-
cation and cleanup systems are identified. Preliminary integration efforts
leading to the selection of systems for detailed evaluation are also given in
Section 2. The results of the foregoing analysis were used to select appli-
cable cleanup processes for use in integrated gasification/cleanup/power sys-
tems. Low- and high-temperature cleanup processes for each gasifier type
were selected for detailed technical and economic evaluation. Detailed sche-
matics and flow sheets were prepared. This work is presented in Section 3«
In Section h, the performance and cost of the integrated gasification clean-
up and power systems is presented. As a basis for comparison, performance
and cost of a direct coal fired steam system is also given. Capital costs of
the selected systems are tabulated and resultant power costs compared. Lastly,
environmental goals were selected and comparison of the emissions from the
various systems with these goals were carried out. In addition, the status of
the technology required to meet the various environmental goals was identified
and recommendations made for further effort in these areas. This work is re-
ported in Section 5.
The report contained herein covers the first three phases of a four phase
study. The results presented are thus preliminary in nature, subject to re-
evaluation during the fourth phase. Many of the areas of analysis identified
in the recommendations will be carried out during the fourth phase.
-6-
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SECTION I
GASIFIER AND CLEANUP SYSTEM PROCESS REVIEW
This section presents a survey of low- and high-temperature fuel gas
cleanup processes which could provide essentially sulfur-free fuel to advanced-
cycle electric power generating stations. These cleanup systems would be used
in conjunction with two generations* of coal gasifiers: the first-generation
gasifier would have off gases containing condensible tars as typified by the
fixed-bed type (e.g., Lurgi or McDowell-Wellman); the second-generation gasi-
fier would be of the entrained flow type having off gases without condensibles
(e.g., BCR-two-stage).
COAL GASIFICATION SYSTEMS
Production of low-calorific fuel gas from coal is achieved by gasifica-
tion with air and steam at elevated temperatures. The overall gasification
process is endothermic primarily due to the steam-carbon reaction which re-
quires about 5000 Btu/lb of carbon:
C + H20 - CO + Hg (1)
This heat requirement is usually satisfied by burning a portion of the coal
feed with air.
Gasification systems are commonly classified into three categories accord-
ing to the characteristics of the gasifier coal bed; i.e., fixed, fluidized,
or entrained. Generally speaking, fixed-bed systems operate with counter-
current flow of coal and gas at temperatures below the ash fusion point.
.
First-generation systems are those using current technology which the Con-
tractor judges could appear in the 1975-1978 time frame. Second-generation
systems are those using technology judged by the Contractor to be potentially
applicable for commercial configurations in the 1980-decade.
-7-
-------
Consequently, these systems are characterized by relatively low gasification
rates and substantial tar formation. High-temperature gasification under ash
slagging conditions affords higher gasification rates together with reduced
reactor volume and essentially a tar-free producer gas. Co-current entrained
flow gasifiers typically operate under the latter conditions. However, at the
higher temperatures, the fraction of the coal heating value represented by
sensible heat of the product gas is substantially greater than for fixed-bed
gasifiers. To achieve maximum thermal efficiency, this.sensible heat must be
used effectively. It is this problem of effective utilization .that forms a
major portion of this study.
For use in advanced-cycle power systems the producer gas is required at
elevated pressure. Since the volume of the producer gas is roughly twice
that of the air required for gasification, it is normally advantageous to em-
ploy pressure gasification rather than product gas compression to obtain a given
delivery pressure. Moreover, the specific gasification rate, gas produced per
unit of reactor volume, is greater at higher pressures.
First-generation technology for'pressure gasification of coal is limited
to the fixed bed-type gasifier. Second-generation technology will probably
evolve from current developmental efforts on entrained-flow and fluid-bed
gasifiers. For the purpose of this study, the two-stage entrained flow-type
gasifier was selected, as representing advanced coal gasification technology.
Typical operating characteristics of fixed-bed and entrained-flow gasifiers
are given in Table 1.
The two prominent differences in the performance of these gasifiers are
directly related to their operating temperature:
1. The fixed-bed system operates at 900 F to 1100 F and has rela-
tively low carbon conversion and a large amount of by-product
tar is formed.
2. The entrained-flow gasifier operates at 1800 F and above and
has a ratio of sensible heat of product gas to coal heating
value twice as great as the fixed-bed system. No condensibles
are formed.
The presence of tar in the off gas of the fixed-bed gasifier complicates
downstream heat recovery and subsequent gas purification. For low-temperature
purification of the raw gas, direct quenching is the preferred method, thereby
avoiding heat exchanger fouling, but the level for sensible heat recovery is
degraded. The recovered tar product could be recycled to the gasifier to
increase the overall carbon gasification. If a high-temperature purification
-8-
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Table 1
COMPARISON OF GASIFICATION SYSTEMS
Gasifier type:
Temperature, F
Pressure, psia
Coal
Air/Coal, Ib/lb
Steam/Coal, Ib/lb
% Carbon Gasified*
Producer Gas
Fixed-Bed
1000
500
Western Kentucky
2.69
0.3^9
Two -Stage
Entrain ed-Flow
1800
500
Illinois No. 6
0.56T
99
SCF/lb Coal
Vol. %
N2
' CO
C02
H2
CH^
H c
2
COS
NHo
H20
, Ib/lb coal
', Btu/ scf (tar free)
.sible Heat/Coal HV , %
58.8
U7.61
20.55
5.88 -
13.83
2.76
0.60
0.10
0.25
8.1*2
100.00
0.11
139
12.2
71*. 6
1*7.70
16.7!*
. 8.81*
11.98
3.11*
0.1*6
0.10
0.38
10.66
100.00
0.0
125
21*. 6
*No tar recycle
-------
system is integrated with the fixed-bed gasifier such that no tar is allowed
to condense, then the tar represents a component in the delivered producer
gas. This may present a problem from a pollution standpoint in that the tarry
material typically contains 1.5-3-0 percent sulfur and about 1.0 percent nitro-
gen which, during combustion, will convert to S09 and NO .
^ X
The large sensible heat content of the gas from an entrained-flow gasifier
makes it imperative, from an efficiency standpoint, to use heat recovery tech-
niques in conjunction with low-temperature purification. Fortunately, the ab-
sence of condensible tars facilitates indirect heat exchange and affords the
possibility for regenerative heating of the fuel gas sent to the power produc-
tion system. Because of the high effluent temperature, the entrained flow gasi-
fier is particularly amenable for integration with high temperature purification
systems operating in the range of 1^00-1700 F.
Gasifier Operating Data
Descriptions and operating data on the fixed-bed (Lurgi) and entrained-
flow (BCR) gasifiers were to be Government-furnished information. For the ana-
lytical program carried out during this study, the following gasifier data were
required: coal, steam, and air feed rates and conditions, operating tempera-
tures and pressures, fuel gas output temperature, pressure, and composition,
size limitations on the gasifiers, capital and operating costs, potential emis-
sions of pollutants such as ash, coal, tar derivatives, or thermal discharge
to the environment. Fuel gas constituents such as sulfur compounds, alkaline
metals, chlorine, and particulate matter which could be harmful to the environ-
ment or to the operation of downstream system components were to be quantita-
tively identified.
Unfortunately, the Government-furnished information was not available in
a timely fashion. The study program was therefore based on gasifier operat-
ing data collected by Foster Wheeler Corporation and UTRC from other sources.
The information so obtained, although the best available, was single point
data which necessitated extrapolation to conditions applicable to the subject
study. Furthermore, there was a definite lack of information on the gasifi-
cation system costs as well as on the pollutants discharged from the gasifier.
Bureau of Mines Stirred-Bed Gasifier - The Bureau of Mines (BOM) stirred-bed
coal gasifier was selected as the gasification system for the study of first-
generation powerplants. This selection was made on the basis of the availabil-
ity of operating data and the similarity in operational mode to the Lurgi sys-
tem which is the only present-day high-pressure gasifier in commercial opera-
tion.'
-10-
-------
A schematic diagram of the BOM gasifier is shown in Figure 1. Although
this system is not commercially available, a one ton/hour pilot unit is in
operation in the Morgantown Energy Research Center and a commercial-sized
unit is in the design stage. Because of the rabble arm agitator, the gasi-
fier is capable of processing both caking and noncaking coals.
In operation^), Coal screened to 1 x 2 inches is fed intermittently to
the top of the bed through lock hoppers pressurized with inert gas. Air and
steam are admitted below the rotating grate and the coal bed moves downward
countercurrent to the ascending stream of hot gases produced from partial
combustion of coal. The steam rate is adjusted to control the bed tempera-
ture below the ash fusion point to avoid clinker formation. Ash is discharged
by eccentric rotation of the grate and is removed through a pressurized lock
hopper. Producer gas exiting from the top of the converter is partially cleaned
of entrained dust in a cyclone separator.
The Bureau of Mines supplied a summary of raw data for 35 runs made during
the 1973 operation of the pilot gasifier. These runs used four Eastern caking
coals as feedstock: Illinois Wo. 6 (1.3 percent sulfur), Loveridge (2.9 per-
cent sulfur), Upper Freeport (3.8 percent sulfur), and Western Kentucky No. 9
(3-9 percent sulfur). These data were evaluated with the object of providing
a basis for the study of the first-generation gasifier. Eight runs showed an
acceptable material balance (±5 percent) and, arbitrarily, the highest pres-
sure acceptable run made on the high-sulfur Western Kentucky coal (Run 55-3)
was selected. The data from this run was adjusted to close on elemental bal-
ances, particularly sulfur, which showed only 65 percent recovery. In addi-
tion, the heat balance was adjusted to reduce the large heat rejection to
jacket cooling water which represented 16 percent of the feed coal heating
value. Considering that a commercial-scale gasifier would have a much lower
surface to volume ratio, a 3 percent heat rejection to jacket cooling was
taken as a reasonable value. The adjusted heat and material balance is given
in Table 2 for a Western Kentucky coal. The cold gas efficiency of this pro-
cess, defined as the higher heating value of the tar-free producer gas divided
by the higher heating value of the coal feed, is 72.5 percent. In a commercial
operation where the tar components are condensed out, it would be advantageous
from an overall fuel utilization viewpoint to recycle the tar to the gasifier.
The heat and material balances for the BOM gasifier were therefore revised to
take into consideration recycle of by-product tar. Since the pilot gasifier
has not been operating in this fashion, the resulting gasifier yields
shown in Table 3 represent a judicious extrapolation to commercial scale
-11-
-------
BUREAU OF MINES STIRRED- BED GASIFIER
FIG. 1
COAL
STEAM
ASH
Rl-19-1
-------
Wt. %
C
6U.U?
H
5.U8
S
3.90
0
IK 18
N
1.58
ASH
15.^
R75-951726-100
Table 2
BUREAU OF MINES GASIFIER
GASIFICATION OF WESTERN KENTUCKY COAL
Feed Coal Analysis:
H20
1.58 15.^
H.V. = 11,^50 Btu/lb
Gasifier Operation:
Temperature, F 1000
Pressiire, psia 135
Air (150 F, 185 psia), lb/lb Coal 2.733
Steam (1165 F, 1^5 psia), lb/lb Coal 0.3^9
Gasifier Products:
*Gas, scf/lb Coal 59-21
Tar, lb/lb Coal 0.108 .
Dust, lb/lb Coal 0.016
Ash, lb/lb Coal 0.166
Gas Analysis:
N2 CO COp H2 C% H2S H20
Vol. % U8.26 20.16 6.06 13.U9 2.7^ 0.71' 8.58
M.W. = 25.0^9 HHV = lUO.3 Btu/scf
Tar Analysis:
C H S N+0
Wt. % 82.1 7.6 1.5 8.8
H.V. = 16,000 Btu/lb
*Dust and Tar Free
-13-
-------
Table 3
BUREAU OF MIKES GASIFIER
GASIFICATION OF WESTERN KENTUCKY COAL
WITH TAR RECYCLE
Feed Coal Analysis;
C H S 0 N ASH H20
Wt. $ 6U.U? 5.^8 3.90 U.18 1.58 15.1+U U.95
H.Vo = 11,^50 Btu/lb
Gasifier Operation;
Temperature, F 1000
Pressure, psia 135
Air (150 F, 185 psia) Ib/lb Coal 3.18?
Steam (1165 F, lij-5 psia) Ib/lb Coal 0.^05
Gasifier Products:
*Gas, scf/lb Coal 68.92
Dust, Ib/lb Coal 0.016
Ash, Ib/lb Coal 0.166
Gas Analysis:
N2 CO C02 H2 CI% HpS H20
Vol. % kB.Zk 20.15 6.06 13.52 2.?if O.blf 8.65
a
M.W. =2^.926 HHV = 139-9 Btu/scf
*Dust and Tar Free
-Ik-
-------
y
operations. This extrapolation was made by assuming the same distribution of
gasified carbon as the base case data, using the same air to steam ratio, and
maintaining the same approach to the water gas shift equilibrium. The cold
gas efficiency of the gasification process with tar recycle is thereby increased
to 8^.2 percent.
The BOM data did not indicate the presence of any ammonia or carbonyl sul-
fide in the raw product gas, although these compounds would be expected to be
formed at the gasification conditions. In the subsequent use of these data,
the material balances were adjusted to reflect the presence of KH-^ and COS by
assuming that all of the nitrogen in the coal feed was converted to NHo (con-
servative) and that HpS and COS were present in the off-gas in a volumetric
ratio of 10 to 1.0.-
BCR Entrained-Flow Gasifier - The two-stage entrained flow gasifier developed
by Bituminous Coal Research, Inc. (BCR) was selected as the second-generation
gasification system for this study. An oxygen-blown, 120 ton/day pilot plant
to produce 2.5 MMSCFD of SNG is under construction at Homer City, Pennsylvania.
Design of an air-blown version is currently underway at the Foster Wheeler
Corporation.
A schematic diagram of the BCR gasifier with its auxiliaries is shown in
Fig. 2.'°' Run-of-the-mine coal is crushed, dried, and pulverized until 70
percent passes through a 200 mesh screen. The pulverized coal is metered from
the feed hopper into hot transport gas recycled from the gas purification sec-
tion, and then fed into the upper stage of the gasifier. In this stage, the
coal reacts with synthesis gas from the lower stage together with steam pro-
ducing methane, carbon monoxide, hydrogen, and unreacted char. The gases leave
the upper stage at around l800 F.
Entrained residual char is removed from the gas by cyclone separators and
recycled via superheated steam to the lower stage of the gasifier. The char
then reacts with steam and air at about 2800 F to form synthesis gas and molten
slag. The hot synthesis gas flows to the upper stage for reaction with coal as
described above. Molten slag collects and drains from the bottom of the lower
stage into the slag pot where it is water quenched.
Overall,'the gasifier reactions are endothermic and the process heat re-
quirement is supplied by combustion of char with air. The air rate is regulated
^Commercial systems using Lurgi gasifiers typically recycle the tar or utilize
the tars as a binder for briquetting coal fines into sizes required for use
in the gasifier.
-15-
-------
FIG. 2
BCR ENTRAINED- FLOW GASIFIER
COAL
GAS
TRANSPORT
GAS
SLAG
STEAM
SLAG
R1-19-3
-16-
-------
to maintain the operating temperature in the upper stage while the lower stage
temperature is controlled by steam addition. Temperature in the lower stage
is fairly critical since too high a temperature will damage the refractory and
too low a temperature will cause the ash slag to freeze and accumulate.
Operating data for the two-stage BCR gasifier was supplied by Gilbert
Associates, Inc.(7) who were acting as project monitor on an OCR contract for
the design and construction of a 1200 ton/day developmental gasifier system.
A summary of the data made available on the production of a low-Btu fuel gas
from Illinois No. 6 coal is given in Table k. This information is judged to
represent the expected performance of the BCR gasifier although it remains to
be confirmed via actual operating experience.
The calculated cold gas efficiency for this gasification system is 83.3
percent. Content of particulates, chlorine, alkaline metal vapor, and trace
components in the product gas was not available. It is expected that the high
gasifier operating temperature with sufficient residence time would preclude
the formation of coal tars, naphthas, and phenols resulting in a relatively
clean product gas compared to the fixed-bed gasifier. Coal nitrogen was
assumed to form ammonia.
FUEL GAS CLEANUP PROCESSES
Fuel gas cleanup systems have been divided into two categories; those
requiring substantial cooling of the dirty fuel gas to 250 F or below (low-
temperature systems) and those requiring little or no cooling of the dirty
gas prior to cleaning (high-temperature systems).
Low-Temperature Cleanup Systems
Low-temperature processes for desulfurizing raw producer gas are commer-
cially available and have been widely used for natural gas sweetening and
treating synthesis gas in the chemical process industry; e.g., the manufacture
of ammonia, methanol, oxo alcohols. These systems normally operate below 250
F and are commonly classified into the following categories:
Chemical Solvent Processes'
Physical Solvent Processes
Direct Conversion Processes
Dry Bed Processes
-17-
-------
Table
BCR ENTRAINED-FLOW GASIFIER
GASIFICATION OF ILLINOIS NO. 6 COAL
Feed Coal Analysis:
Wt.
C
67.1*
H
5.1
S
3.8
0
9.6
N
1.2
HHV = 12,200 Btu/lb
Gasifier Operation:
Temperature , F
Pressure, psia . .
Transport Gas (300 F, 555 psia), scf/lb Coal
Air (800 F, 515 psia), Ib/lb Coal
Steam (950 F, 915 psia), Ib/lb Coal
Gasifier Products:
Gas, scf/lb Coal
Slag, Ib/lb Coal .
Product Gas Analysis:
Vol. %
N2
1*5.76
CO
77.27
0.087
C0
18.13 8.31 13.07 3.60
M.W. = 2U.519 HHV, Btu/scf = lUl.9
Transport Gas Analysis:
Vol. %
N2
M.W. =
. CO
19.58
00-
n2
lU.10
3.91
HHV, Btu/scf = 1U7.6
ASH
8.T
l800
6.560
3.088
0.567
COS
0.10
H20
U.2
NH3 H20
0.39 10.15
NH-
H20
0.08 6.69
-18-
-------
In Table 5 are listed the gas treating processes available for removing
sulfur compounds, particularly hydrogen sulfide. from raw producer gas. This
list was compiled from a literature survey(°-11) and, while it is believed to
represent the major technology in this area, it is not necessarily all inclu-
sive. Figure 3 illustrates the process scheme for a typical absorption-strip-
ping process for low-temperature acid gas removal.
Chemical Solvent Processes - Chemical solvent processes employ aqueous solu-
tions of organic and/or inorganic agents which are capable of forming "com-
plexes" with the acid gas components, notably HpS and C02, present in the
raw gas stream. The absorption solution is regenerated by decomposing the
"complex" at elevated temperature thereby releasing the acid gases for subse-
quent recovery and recycling the solution for further absorption. These pro-
cesses are essentially insensitive to the partial pressure of acid gases in
the feed and generally exhibit little or no selective absorption of H2S over
carbon dioxide. The chemical processes may be sub-divided into those using
amine scrubbing solutions and those based on alkali scrubbing solutions.
The principal reactions involved in gas sweetening with amine solutions
(10-30 percent weight) may be represented as:
RNH2 + H2S - RKH3 HS • (2)
+ C02 + HO - RNH3 HC03 (3 )
Monoethanolamine (MEA) will easily reduce the HpS content below k ppm; however,
it is not considered selective, even though the rate of C0~ absorption is less
than for HgS. The principal disadvantage of MEA is .that it will react with COS
and CS2 forming nonregenerable compounds. Diethanolamine (DEA) will not react
with these contaminants and is favored for service where COS and CS2 are likely
to be present. Like MEA, DEA solutions are not selective for H2S, and, will
seldom reduce the H2S content below 100 ppm. Tertiary amines, such as trieth-
anolamine and methyl-diethanolamine , while not as reactive as the other amines,
have the advantage of being selective towards H2S removal. The tertiary amines
are two to four times more costly and find little application in industrial gas
sweetening.
The alkali scrubbing system may be represented by the following chemical
reactions :
C03 + H2S - MHS + MH C03
C0 + C02 + H20 - 2 MH C0 (5)
-19-
-------
Table 5
LOW-TEMPERATURE DESULFURIZATION PROCESSES
PROCESS
Chemical Solvent
MEA
SNPA:DEA
TEA
i MDEA
ro
o
Econamine
ADIP
Alkazid
Hot Potash
SORBENT
Processes
Monoethanolamine
Diethanolamine
Triethanolamine
Methyl-diethanolamine
D i gly co lamine
Di-Isopropanolamine
Potassium Dimethyl
Amino Acetate
Potass ium Carbonate
TEMPERATURE
P
80-120
100-130
100-150
100-150
80-100
80-100
70-120
200-250
SELECTIVE
HPS REMOVAL
No
Wo
Yes
Yes
No
No
Yes
Partial
STATUS
Commercial
Commercial
- Not Commercially
Important
Not Commercially
Important
Commercial
Commercial
Commercial
Obsolete
Catacarb
Solution
Activated Potassium
Carbonate Solution
150-250
Partial
Commercial
-------
Table 5 - Continued
ro
PROCESS
Benfield
Seaboard
Vacuum Carbonate
TPP
Sodium Phenolate
Aqua Ammonia
SORBENT
Activated Potassium
Carbonate Solution
Sodium Carbonate
Solution
Sodium Carbonate
Solution
Tri-Potassium
Phosphate Solution
Sodium Hydroxide
+ Phenol Solution
Aqueous Ammonia
TEMPERATURE
F
150-250
80-100
80-100
80-120
80-100
60-120
SELECTIVE
HpS REMOVAL
Partial
Partial
Partial
Partial
Partial
Partial
STATUS
Commercial
Obsolete
Obsolete
Commercial
Obsolete
Commercial
Solution
-------
Table 5 - Continued
i
ro
ro
i
PROCESS
Physical Solvent Processes
Sulfinol
SORBENT
Selexol
Fluor Solvent
Purison
Rectisol
Estasolvan
Water Wash
Sulfolane +
Di-Is opropanolamino
Polyethylene
Glycol Ether
Water
TEMPERATURE
F
80-120
20-80
70-100
SELECTIVE
H2S REMOVAL
Possibly
Yes
Yes
STATUS
Commercial
Commercial
Propylene Carbonate
N-Methyl
Pyrrolidone
Methanol
Tri-N- Butyl
Phosphate
1*0-80
70-100
< Zero
80-120
Yes
• Yes
Yes
Yes
Commercial
Commercial
Commercial
Not Commer<
Import an1
Not Commercially
Important
-------
ro
OJ
i
PROCESS
Direct Conversion Processes
SORBEWT
Table 5 - Continued
TEMPERATURE
F
Stretford
Takahax
Giammarco-
Vetrocoke
Thylox
Ferrox
Manchester
Perox
Sodium Carbonate' +
Anthraquinone
Sulfonic Acid
Sodium Carbonate +
Wathoquinone
Sulfonic Acid
Arsenic Activated
Potassium Carbonate
Arsenic Trioxide
+ Sodium Carbonate
Iron'Oxide +
Sodium Carbonate
Iron Oxide +
Sodium Carbonate
Aqueous Ammonia
+ Hydroquinone
100-300
SELECTIVE
HpS REMOVAL
Yes
Yes
Yes
Yes
Yes
Yes
Yes
STATUS
Commercial
Commercial
(outside U.S.)
Commercial
(outside U.S.)
Commercial
(outside U.S.)
Commercial
(coutside U.S.)
Commercial
(outside U.S.)
Commercial
(outside U.S.)
-------
Table 5 - Continued
ro
PROCESS
Lacy-Keller
Townsend
IFF
Deal Process
Nalco Process
Dry Bed Processes
Haines
Molecular Sieve
Iron Sponge
Activated Carbon
SORBENT
Proprietary
Triethylene Glycol
Proprietary
I
Aqueous Sulfolane
Proprietary
TEMPERATURE
F
—
100-250
200-300
—
__
SELECTIVE
HpS REMOVAL
Yes
Yes
Yes
Yes
Yes
STATUS
Not Commercially
Important
Not Commercial
Not Commercial
Not Commercial
Not Commercial
Zeolites (+S02)
Molecular Sieve
(adsorption)
Hydrated Ferric Oxide
Carbon
70-100
70-100
70-100
70-100
Yes
Yes
Yes
Yes
Not Commercially
Important
Commercial
Commercial
Commercial
-------
TYPICAL LOW-TEMPERATURE ACID GAS REMOVAL UNIT
ro
VJl
i
TREATED GAS
ABSORBER
RAW GAS
FEED
RICH
SOLVENT
ACID GAS
LEAN SOLVENT
REGENERATOR
-STEAM
P
w
-------
The earlier processes, such as Seaboard and Vacuum Carbonate, were based on
dilute solutions of sodium carbonate (3-^ percent weight) and were capable
of removing 80-90 percent of the HpS. Regeneration in the Seaboard process
was by air resulting in a dilute acid gas stream while the Vacuum Carbonate
system used vacuum regeneration with steam. These processes were superseded
by the hot potassium carbonate system. In the "hot pot" processes, an aqueous
solution of 25-35 weight percent Y^ C0_ is used to absorb acid gases at tempera-
tures in the range of 200-250 F. With low J^S/CO ratios, the process is capable
of sweetening the gas to 5 ppm. A degree of selective E^S absorption over C02
can be achieved by taking advantage of the relatively slow rate for COp absorp-
tion. In addition to removing I^S and C02, the process can remove COS and CS2
by hydrolysis of these components to C02 and HpS. The Catacarb and Benfield
processes are improved versions of the Bureau of Mines "hot pot" systems insofar
as they employ activators to increase the rate of absorption thereby decreasing
the required solution circulation rate. Disadvantages of the hot potassium
carbonate systems are the relatively high steam consumption for regeneration,
required operating pressure above 300 psi, and the inability to remove mercap-
tans.
The tripotassium phosphate process was developed by the Shell Oil Company
specifically for HpS removal via the reaction:
K_ PO^ + HgS - Kg HPO^ + KHS (6)
The nonvolatility of the agent, its nonreactivity with COS and CS25 and partial
selectivity toward ^S in the presence of C02 gives it certain advantages over
the amine systems. However, when operated for high ^S selectivity, the process
only gives about 90 percent removal efficiencies, and conversely, with high H2S
removal the steam consumption becomes excessive due to C02 absorption.
Physical Solvent Processes - Physical solvent processes all use organic solvents
to remove acid gases by physical absorption, rather than chemical reaction,
which is directly proportional to the partial pressure of the acid gas components.
These processes are most applicable to high-pressure gas treating where appre-
ciable quantities of sour gases are present. After absorption, the "loaded"
solvent is regenerated by heat and/or pressure reduction giving a concentrated
stream of ^S plus CO and a recyclable lean solvent. Due to the higher solu-
bility to H2S in these organic solvents, selective absorption of ^S over COg
can be achieved. In general, these processes have two major disadvantages; the
solvents have a great affinity for absorbing heavy hydrocarbons (C,-+) which con-
taminate the gas stream fed to sulfur recovery units, and the solvents are quite
expensive so that large solvent losses cannot be tolerated.
-26-
-------
As a group, these processes were developed for bulk removal of acid gases
but, for low H2S concentration, they are capable of giving a sweetened gas
having less than 5 ppm ffeSo In order to maximimize the solubility of acid gases
and minimize solvent loss through vaporization, the processes are generally
operated at or below ambient temperature. In addition to removing H2S and CC>2,
the solvent processes are all capable of removing COS, 082 and mercaptans without
solvent degradation as well as dehydrating the gas to a low dew point. The low
heats of solution for acid gases result in appreciably lower -steam requirements
for solvent regeneration compared to the chemical solvents.
The Sulfinol process is unique in that it combines the characteristics of
a solvent process and an amine process. The physical absorbent, Sulfolane, gives
high acid gas loadings at high acid gas partial pressures, giving it bulk removal
capacity and the chemical absorbent, DIPA* reduces residual acid gases to very
low values. However, the presence of the chemical solvent reduces the H>>S selec-
tive absorption for this system compared to the straight solvent processes.
Direct Conversion Processes - Two types of processes fall into this category:
a. those based on oxidation reduction reactions, and
b. those based on the stoichiometric reaction of HgS with SOg
in the-presence of a solvent.
The first type involves the absorption of H2S in alkaline solutions containing
oxygen carriers. The K^S is subsequently oxidized to elemental sulfur by air
fed to the regenerator where the sulfur product is fHotated and collected at the
regenerated solution interface as a froth. Processes of this type are in common
use in Europe for removal of H^S and sulfur recovery from manufactured gases and
coke-oven gas. The Ferrox and Manchester processes employ a suspension of iron
oxide in an aqueous solution of sodium carbonate to absorb t^S. With multi-
stage absorption, essentially complete removal of H2S is obtainable; however, the
product is of low quality due to salt contamination and chemical replacement costs
are high. Both the Thylox and Giammarco Vetrocoke'^"^^) processes use alkaline
solutions of arsenates and are capable of reducing the I^S to less than 1 ppm.
Partial removal of COS, CS2, and mercaptans is also possible. Again the sulfur
product is contaminated and the use of arsenates makes these processes poten-
tially hazardous. The Stretford and Takahax processes are similar in that
alkaline solutions of quinone sulfonic acids are employed. The addition of
vanadium salts increase the rate of oxidation of hydrosulfide to sulfur resulting
in higher solution loadings. Close to 100$ H2S removal is claimed for these
processes along with high purity (99$) sulfur product. However, substantial
amounts of thiosulfates are formed resulting in sludge deposition and corres-
ponding chemical makeup.
*di-isopropanolamine
-27-
-------
Generally speaking, the low solution loadings exhibited by this group of
processes make them uneconomical for treating large very sour gas streams.
They are best suited for sour gases containing less than 1.0$ H2S with sulfur
production under 20 tons/day. These processes, as well as those in the follow-
ing group are almost totally selective for H2S removal.
In the second group of direct conversion processes are those in which I^S
is absorbed in a solvent and converted to elemental sulfur by the Glaus type
reaction with S02.
2 H2S + S02 -* 3 S + H20 (?)
The Townsend process uses an aqueous solution of an organic solvent, such as
triethylene glycol, to sweeten the gas, dehydrate the gas, and convert H2S to
elemental sulfur. A portion of the product sulfur is burned to S02 which is
absorbed by fresh solvent and the S02 rich solvent is used to contact the sour
ga's. The IFF, Nalco, and Deal processes operate in a similar manner employing
other solvents. While a high purity (99-7$) sulfur product is claimed, none of
these processes have been commercialized.
Dry Bed Processes - These sweetening processes are based on absorption of acid
gases by a fixed bed of solid absorbent. Due to their low absorbent loading,
they are applicable to gases containing low concentration of I^S and mercaptans,
perhaps less than 500 ppm. These processes can be subdivided into the iron oxide
processes and the various molecular sieve processes.
The iron oxide or dry box process is one of the oldest processes known for
removing sulfur compounds from industrial gases. In the iron sponge system,
wood shavings impregnated with hydrated ferric oxide are used to. absorb H2S:
2 Fe2 03 + 6 H S •* 2 Fe2 S3 + 6 HgO ,g<
Regeneration of the absorbent is carried out with air:
2 Fe2 S + 302 -*• 2 Fe2 03 + 6s (9)
This process is best suited for small to medium gas volumes with low sulfur
contents, otherwise the sponge bed life would be too short to be economical.
The process is selective towards H2S and mercaptans and will partially remove
COS and CS2- Sweetened gas of less than 5 ppm H2S is easily obtained. However,
sulfur recovery would not be economical when using the iron sponge system.
-28-
-------
Molecular sieves can be tailor made to have pore sizes which will permit
selective absorption of H2S over CC^. These processes are characterized by the
various regeneration schemes employed; i.e., via hot combustion gases, hot SC>2
gas as in the Haines process, or hot air. In the latter two modes, elemental
sulfur is produced via the oxidation of the absorbed H2S. The sieve processes
also appear to be economically attractive for small to medium gas volumes having
low HjS content. Additionally, for efficient I^S removal, the raw sour gas should
have a water content below 20 Ib/MMSCF since water will also be absorbed by the
molecular sieve structure.
Selection Considerations
Low-temperature desulfurization systems for application in low-Btu fuel gas
plants will have to treat large volumes of sour gas, 500-1000 MMSCFD, having total
sulfur content in the range of O.U to 1.0$ (weight). In addition to I^S, the
raw gas will contain C02, COS, CSp, probably mercaptans, cyanides and heavy hydro-
carbons. Of the types of processes described above, it is evident that the liquid
scrubbing processes, physical solvents and some chemical solvents, are the best
suited. These processes are currently available and can easily reduce the sulfur
content to 100 ppm, which, when combusted, would result in S02 emissions well
within present EPA regulations for conventional steam stations. As such, these
processes are capable of serving both first generation and second generation coal
gasification plants.
High-Temperature Cleanup Systems
High-temperature systems for sulfur removal are not presently available in
commercial scale although there are several in various stages of development.
Most of the active work presently involves use of limestone and dolomites which
have potential in the range of 1500-2000 F. Other systems receiving attention
employ iron oxides, molten salts, and liquid metals. These systems operate by
chemical reaction of the absorbent with sulfur compounds in the gas, forming the
corresponding metal sulfides. The degree of desulfurization attainable depends
on the chemical equilibria for the particular system at operating conditions. As
with low-temperature processes, economics dictate that the sulfided absorbent be
regenerated for reuse.
The only commercial experience with high-temperature desulfurization reported
in the literature is that of the Frodingham Desulfurization Process.(1^~15) This
process employed fluidized beds of pulverized iron oxide operating at 650 F to
800 F. In the early 1960's, a commercial plant treating 32 MMCFD of crude coke
oven gas containing 1.0$ H2S was operated at the Exeter Works of the South Western
Gas Board. Essentially complete (99-9$) removal of HgS was achieved with
removal of all organic sulfur compounds other than thiophene. The sulfided
-29-
-------
absorbent was regenerated with air at 1000-1100 F resulting in a S02 stream
which was subsequently converted to sulfuric acid. Major difficulties were
experienced in the solids handling system which produced fine oxide dust
resulting in excessive losses of the absorbent. In addition, operation of
the sulfuric acid plant was erratic due to low S02 concentrations in the regen-
erator off -gas.
Several processes currently under development which may prove commercially
viable for use with second generation gasification systems are listed in Table 6.
Bureau of Mines Process ( 16-17)- This process, under development at the Morgan-
town Energy Research Center, is based on a sintered absorbent consisting of a
mixture of iron oxide (Fe203) and fly ash. This sorbent satisfied the primary
requirements for high-temperature sulfur removal in that it is readily available
and inexpensive, it has reasonable absorption capacity for sulfur, can be regen-
erated for repeated use, and is resistant to fusion and disintegration over the
operating temperature range of 1000-1500 F. The absorbent is prepared by mixing
iron oxide with "as received" fly ash to a total iron oxide content of about 35$.
Iron oxide contents above Uo$ were unsatisfactory because the fusion temperature
was lowered within the operating range. The mixture is extruded into 1/U" x 3/8"
pellets and then sintered to develop the required hardness.
Absorption studies over the range of 1000-1500 F show sulfur capacities of
10 to 25$ by weight, respectively, for dry simulated producer gas. The presence
of water vapor reduces the capacity to 6-10$ by weight but there is no evidence
of loss of absorption effectiveness over 150 cycles of regenerations.
The reaction mechanism is chemisorption wherein hydrogen sulfide diffuses
into the sorbent particle and reacts with iron oxide forming iron sulfide.
Fe0 + 3HS + FeS + FeS + 3H0 (10)
FeO + H2S ->• FeS + HgO (ll)
Sulfided absorbent is regenerated with air at temperatures of 1000-1500 F produc
ing an SOg containing off -gas and reusable Fe203. Since sulfur recovery in the
elemental form is preferable from a pollution standpoint, the formation of S02
is a disadvantage to this process.
In accordance with the above absorption reactions, the residual sulfur con-
tent in the treated gas is governed by the chemical equilibrium relationships:
[H2S] = [H20]
12S] = [H2
2 2
[H2S] = [HO] / K2
-30-
-------
Table 6
HIGH TEMPERATURE DESULFURIZATION PROCESSES
Process
Bureau of Mines
Babcock & Wilcox
Consolidation Coal
(CONSOL)*
Air Products
Battelle Northwest
IGT-Meissner
Absorbent
Sintered Iron
Oxide + Fly Ash
Iron Oxide
Half-Calcined
Dolomite
Calcined Dolomite
Molten Carbonates
Proprietary
Temperature, F
1000-1500
700-1200
1500-1800
1600-2000
1100-1700
800
Status
Pilot
Experimental
Pilot
Pilot
Pilot
Unknown
*Now Conoco Coal Development Company
-31-
-------
where [ ] = the gas phase mol fraction of the component and K]_, K2 = equilibrium
constants for reactions (10) and (11 ), respectively. KI and K2 are given in
Figs, k and 5. Removal of H2S by this process is independent of operating pressure,
and for a given gas composition, will decrease with increasing temperature. Iron
oxide catalyzes the water gas shift reaction.
+ CO - H2 + C02 (lU)
so that the gas composition approaches equilibrium with respect to this reaction
as it passes through the absorption bed. Equilibrium constants for the shift
reaction are shown in Figure 6.
(1Q1
Data on the absorption of IS from actual producer . gas v •" indicate that
the equilibrium absorption closely corresponds to equation (13) above.
Although the absorption of other sulfur compounds on this sorbent have not
been studied, thermodynamically, the removal of COS appears to be practical:
COS + FeO - FeS + C02 (15)
where, [COS] = [C02]/Ko, KO = equilibrium constant.
Values for K^ are given in Figure 7-
To date, this process has been operated on a pilot scale absorbing H2S in
the raw producer gas from a 1.0 ton/hour stirred-bed coal gasifier. The H2S
content was reduced from 0.6$ to 150 ppm at 1100-1200 F and 120 psi . Tar present
in the raw gas was not removed by the absorbent. Further work on the regeneration
cycle is currently in progress. Commercial- scale plants would be based on the mul
tiple fixed-bed principle alternating between absorption and regeneration cycles.
Fluidized bed operation would not be considered due to potential attrition of the
absorbent.
Consolidation Coal Process - This process*1 ®' evolved as an adaption of the C02
Acceptor Process(21) for producing low-Btu fuel gas from coal and incorporates the
use of a half-calcined dolomite acceptor for sulfur capture as studied by Squires
and coworker s' 2^-25 ). Basically the process chemistry involves the following
reaction:
[CaC03 • MgO] + I^S - [CaS - MgO] + HgO + C02 (16)
for which the equilibrium H2S concentration in the gas is given by
[H2S] = [H20] [C02] P/K (1?)
-32-
-------
FIG. 4
EQUILIBRIUM CONSTANT FOR hS ABSORPTION BY IRON OXIDE
CO
z
o
CJ
cc
CO
D
a
01
106
io5
104
103
102
I I
Fe203 + 3 H2S—»-FeS + FeS2 + 3 H20
K = (H20)3/(H2S)3
I
800 1000
1200 1400 1600
TEMPERATURE, F
1800 2000
R 1-19-9
-33-
-------
FIG. 5
EQUILIBRIUM CONSTANT FOR H2S ABSORPTION BY IRON OXIDE
to
o
o
cc
CD
O
LU
10*
103
102
10
FeO + H2S-*-f eS + H2O
K = (H20)/(H2S)
I _]_
800
1000 1200 1400
TEMPERATURE, F
1600 1800
R1_19_10
-------
FIG. 6
EQUILIBRIUM CONSTANT FOR WATER GAS SHIFT
10
V)
z
o
CJ
DC
CD
O
UJ
1.0
CO + H,0.
K = (H2) (C02)/(H20) (CO)
I
1000 1200 1400 1600
TEMPERATURE, F
1800 2000
R1-19-24
-35-
-------
FIG. 7
EQUILIBRIUM CONSTANT FOR COS ABSORPTION BY IRON OXIDE
106
105
o
o
a:
CD
O
UJ
104
103
I
FeO + COS-
• FeS
K = (CO2)/(COS)
1
600 800
1000 1200 1400
TEMPERATURE, F
1600 1800
R1-19-23
-36-
-------
where [ ] = mol fraction of component
P = total pressure, atm
K = equilibrium constant
From Fig. 8, showing the variation of K with temperature, it is obvious that the
removal of H2S increases with temperature for a given gas composition. Also, for
a given gas composition and temperature, the removal efficiency is inversely pro-
portional to the operating pressure.
A maximum operating temperature for this process is imposed by the partial
pressure of carbon dioxide in the gas phase; i.e. the temperature cannot exceed
that at which the C02 partial pressure is equal to the decomposition pressure for
CaCOo via the following reaction:
Ca C03 - CaO + C02 (18)
The equilibrium decomposition pressures for calcium carbonate are given in Fig. 9-
Although no data have been reported for COS adsorption by half-calcined
dolomite, high COS removal efficiencies are predicted thermodynamically according
to the reaction:
[CaC03 • MgO] + COS - [CaS • MgO] + 2 C02 (19)
where the residual COS concentration is
[COS] = [C02]2 P/K ' (20)
The equilibrium constant for this reaction is shown in Fig. 10.
The process, as described by Consol involves desulfurizing the raw gas in
a fluidized bed of half-calcined dolomite acceptor at 1600-1700 F, according to
reaction (16). The sulfided acceptor is regenerated by the addition of steam and
C02 at reduced temperature, thereby reversing the absorption reaction. Regenera-
tion is conducted in a fluidized bed at around 1300 F giving a dilute H2S off-gas,
less than 10$ (volume). Because the low H2S off-gas content prohibits the direct
use of a vapor-phase Glaus unit for sulfur recovery, Consol is proposing the use
of a liquid phase sulfur recovery system based on the Wachenroeder reaction
2H2S + H2 SO - 3S + 3H20 (21)
-37-
-------
FIG. 8
EQUILIBRIUM CONSTANT FOR H2S ABSORPTION BY HALF-CALCINED DOLOMITE
1000
K
z
CO
I
oc
CO
o
100
10
1.0
0.1
K =
-------
FIG. 9
DISSOCIATION PRESSURE FOR CALCIUM CARBONATE
100
Ca C03-
CO
10
ui
oc
D
(ft
to
LU
OC
CL
•z.
CJ
o
to
00
1.0
0.1
0.01
1200 1400 1600 1800
TEMPERATURE, F
2000
2200
R1_19_19
-39-
-------
FIG. 10
EQUILIBRIUM CONSTANT FOR COS ABSORPTION BY HALF-CALCINED DOLOMITE
104
in
a
o
00
O
LLJ
102
10
K = (C02)2 P/ICOS)
(Ca CO3- MgO) + COS
I I
1000 1200 1400
TEMPERATURE, F
1600
1800 2000
R1_19_16
-40-
-------
At the present time, the Consol desulfurization process is under development in
the pilot plant stage. An apparent drawback to the process is the low degree of
regeneration obtainable for the sulfided acceptor, around 10-13$, which results
in a large recirculation of sulfided dolomite to the absorber. For operation
at 200 psia, Consol projects a reduction of H2S from 0.6 percent to 200 ppm in
the treated gas.
Air Products Process - As with the Consol process, this system employs dolomite
as the sulfur acceptor. However, the acceptor is in the fully calcined form; i.e.
[CaO • MgO], and the process therefore consists of three steps: absorption,
regeneration, and calcination.
Absorption of hydrogen sulfide takes place at around 1600-1700 F via the
reaction with calcined dolomite:
TCaO • MgO] + H2S -' [CaS ' MgO] + H20 (22)
The equilibrium HpS concentration may be calculated from the equilibrium
constant given in Figure 11 according to the following relationship:
[H2S] = [H20] /K (23)
where,
[ ] = mol fraction of component
K = equilibrium constant
For this system, the H2S removal efficiency is independent of pressure but, for
a given gas composition, it decreases with increasing temperature. From a
practical viewpoint, there is a minimum temperature at which this process should
operate; that being the temperature at which the decomposition pressure of
equals the partial pressure of C02 in the gas. Below this temperature carbon
dioxide will also be absorbed according to the reaction.
Ca 0 + C02 - Ca C03 (2k]
This is detrimental for two reasons:
a. Additional acceptor capacity is required, and
b. the large exothermic heat of C02 absorption, 75»000 Btu/mol, will
necessitate some means for heat removal from the absorption bed.
-------
FIG. 11
EQUILIBRIUM CONSTANT FOR H2S ABSORPTION BY CALCINED DOLOMITE
106
800 1000 1200 1400
TEMPERATURE, F
1600
1800
2000
R1-19-22
-1(2-
-------
Like the Bureau of Mines and Consol processes, COS removal appears to be
thermodynamically attractive. Residual COS content may be estimated from the
chemical equilibrium associated with the absorption reaction:
[CaO • MgO] + COS - [CaS • MgO] + C02 (25)
[COS] = [C02] /K (26)
where equilibrium constant is given in Figure 12.
Regeneration of the sulfided acceptor is conducted similarly to the Consol
process. Steam and carbon dioxide are reacted with the sulfided dolomite at
1100-1200 F resulting in a dilute H2S off -gas and the carbonated dolomite form.
[CaS • MgO] + H20 + C02 - [CaC0 • MgO] + H2S (2?)
Hydrogen sulfide in the regenerated off -gas may be converted to elemental sulfur
via the liquid-phase Wachenroeder process or first concentrated and then fed to
vapor-phase Glaus unit .
The regenerated dolomite must be calcined before recycle to the subsequent
absorption cycle. Calcination is effected at 1900 F with air to drive off carbon
dioxide:
[Ca C03 • MgO] - [CaO , MgO] + C02 (28)
This reaction is endothermic and requires the use of fuel to preheat the
air. Excessive temperature, above 2000 F, during calcination can result in
deactivation of the acceptor .
The development of this process by Air Products is currently in the pilot
plant stage.' ®> The absorption cycle has been successfully tested on actual
producer gas. Operating conditions for regeneration and calcination are still '
under investigation to determine the maximum acceptor life cycles .
IGT-Meissner Process - This process is under development by the Institute of Gas
Technology in association with its "U-Gas" process. (2?) Details of this technol-
ogy have not been divulged except that the process operates at 800 F and selectively
removes H2S and COS relative to CO-.
Battelle Northwest Process - The Battelle process' ' utilizes calcium carbonate,
CaCOo, dissoDved in a tertiary mixture of alkali metal carbonates to remove H2S
at high temperature. The tertiary carbonate system consisting of potassium
carbonate, lithium carbonate, and sodium carbonate, has a eutectic melting point
-1*3-
-------
FIG. 12
EQUILIBRIUM CONSTANT FOR COS ABSORPTION BY CALCINED DOLOMITE
106
105
t/3
z
o
0
cc
00
D
o
104
103
(CaO • MgO) + COS-»-(CaS • MgO) + CO2
K = (C02)/(COS)
I
1000 1200 1400 1600
TEMPERATURE, F
1800
2000
R1-19-11
-hk-
-------
around 750 F and under operating conditions contains about 15 mol$ CaCOo. Besides
removing sulfur compounds from the gas stream, this solvent will also scrub out
the fly ash constituents from the raw gas.
The system under.study at Battelle contacts the molten salt and raw gas in
a co-current flow venturi scrubber at temperatures from 1100 to 1700 F. Hydrogen
sulfide is removed from the gas by chemical reaction with the carbonate solvent:
Ca C03. + H2S -* CaS + C02 + E^O (29)
The equilibrium H2S content in the treated gas may be expressed as:
(CaS) ' [H?0] [C02]
x P
(30)
(Ca C03) K
where,
( ) = activity of component in solution
[ ] = mol fraction of component in gas
P = total pressure in atm
K = equilibrium constant
Unfortunately, due to the chemical complexity of the molten salt system, the
equilibrium constant cannot be accurately predicted. Observed K values have
been a factor of ten below the calculated values. Qualitatively, the H2S
removal efficiency improves with temperature and is inversely proportional to
pressure. Experimental data at atmospheric pressure have indicated high H2S
removal, > 9^%> at salt loadings under 50$ of capacity. Regeneration of the
salt is conducted with steam and C02 at 1000-1100 F giving an off-gas having
H2S concentrations suitable as feed to a Glaus unit for recovery of elemental
sulfur.
This process is presently in the pilot plant stage where it will be demon-
strated on an actual gasifier raw gas at low pressure, 0-10 psig. It is doubtful
that this system can achieve high H2S removal efficiencies at high pressure and
its ability to handle other sulfur contaminants is yet to be demonstrated.
Materials of construction for a commercial unit will also present a formidable
problem.
-------
Selection Considerations
In general, commercialization of high-temperature systems appears to be
5-10 years away. The processes under development are all capable of selectively
removing HoS down to acceptable levels, less than 500 ppm. Areas which need
further definition include ability to contend with other sulfur compounds and
contaminants present in raw producer gas, absorbent life expectancy, actual
performance data to demonstrate long term reliability, and recovery of elemental
sulfur from regenerator off-gases. It would appear that the Bureau of Mines
process shows promise of being the first to mature. This process has been demon-
strated on actual producer gas, involves relatively simple operation without
complex solids handling, and has proven cyclic operation without loss in acceptor
activity. The problem of sulfur recovery from the regeneration step represents
the major area for refinement.
-------
SECTION 2
PARAMETRIC ANALYSIS OF POWER SYSTEMS
Preliminary analysis, of the power systems considered for use with a fuel
gasification and cleanup system are presented in this 'Section. They fall
into two general categories, steam systems and gas turbine systems. It is
increasingly apparent that steam cycles have reached a stage.of maturity where
little -or no improvement can be expected in the next decade. This is best
demonstrated by the decline in the percentage of new steam plants using super-
( PQ ^
critical pressures in the 3500 psi range.\ ?' The marginal operating cost
savings have apparently proven insufficient to justify the added capital costs
associated with the higher pressures. Therefore, current steam plant perfor-
mance will not change significantly with present and future gasifier technol-
ogy.
Gas turbine technology, on the other hand, has not yet matured and
further improvements in the performance of simple-cycle machines can be
expected. These improvements generally involve increased turbine inlet tem-
peratures which, when coupled with increased pressure ratios, not only result
in better simple-cycle performance but enhance the attributes of the combined
gas and steam (COGAS) systems by permitting higher steam temperatures. Both
simple- and combined-cycle performance characteristics have been examined to
mate with first- and second-generation gasifier technology.
As a point of departure, the performance of a coal-fired steam system
with stack gas cleanup has been included in this analysis and in the subse-
quent economic evaluation as a source of relative comparison.
Since one of the objectives of the program is to identify the differ-
ences between and advantages of low- and high-temperature cleanup systems,
it was decided to fix the fuel input at a given level and allow the power
output to vary. In this way, any scale effects for the gasifier/cleanup
systems would be negated. This gasifier plant size has been selected to
handle a coal input of 700,000 Ib/hr, the amount nominally required for a
1000-Mw coal-fired station. The resultant electrical output is a function
-------
of system conversion efficiency and was found to vary from about 800 to over
1,100 Mw. In the combined-cycle generating systems, an even number of gas
turbines has been assumed of the largest size practicable using a 3600-rpm
single-end power turbine. For these systems, steam turbine size is maintained
at or greater than 100 Mw. When the gasified fuel is used directly in con-
junction with a steam cycle, performance is based on the use of two 500-Mw
units. These factors are discussed in more detail in the following paragraphs
of this section.
DIRECT COAL-FIRED STEAM CYCLES
Because of the commanding position currently maintained by steam power
systems, they are the logical yardstick against which the feasibility of alter-
nate generation systems must be compared. This section first presents esti-
mates of the performance that can be expected from a typical late 1970-decade
steam system without any type of emission controls. The effect of sulfur
emission control by stack gas cleanup on that basic plant is then estimated,
providing a basis against which the various combinations of gasifiers and
cleanup systems can be measured when used in conjunction with either a steam
cycle or in advanced power cycles.
The reference steam power plant consists of a turbine-generator and steam
cycle mated to a coal-fired boiler. Overall performance for that system is
presented in Table 7- The steam cycle used is typical of systems currently
under construction with 1000-F, 2^00-psi throttle conditions and a single reheat
to 1000 F. This same steam cycle is used in evaluating gasified coal-fired
steam systems. The significant steam cycle parameters are shown in Figure 13.
Without environmental restrictions, a coal-fired system can be expected to
have a heat rate of about 9235 Btu/kwh based on the higher heating value (HHV)
of the fuel. Of course this increases with condenser temperature and as other
auxiliary power requirements increase. With stack gas cleanup, the power
required for scrubber operation and the necessary reheat of the stack gas
would reduce output power by 5$ and give a net heat rate of 9721 Btu/kwh. This
varies with the amount of sulfur in the fuel and with the removal process
employed. Throwaway systems will have a slightly lower power consumption than
those using a regenerable sorbent, but are faced with a severe disposal problem.
The cleanup penalty assumed herein, is intended to be representative of systems
that will be operational in the 1970's.
The system performance assumes that the yearly average condenser pressure
is 2.0 in. Hg when rejecting heat to a wet cooling tower. During the months
of June through September the average condenser pressure will-be about 3-5 in.
Hg. This will result in an increase in the net steam cycle heat rate from
7,836' to 8,058 Btu/kwh or 2.83 percent. This same percent increase applied
-U8-
-------
TABLE 7
CHARACTERISTICS OF BASIC COAL-FIRED STEAM POWER PLANT
500-Mw Nominal Size Unit
Single Unit Characteristics
Steam Conditions
Throttle Pressure
Throttle Temperature
Reheat Temperature
Condenser Pressure
Final Feedwater Temperature
Steam Cycle
Net Steam Cycle Output Power
Auxiliary Turbine Driven Feed Pump Power
Gross Steam Cycle Power
Gross Heat Rate
Net Steam Cycle Heat Rate
Station Efficiency
Boiler Efficiency (Illinois #6 Coal) (HHV)
Plant Auxiliary Power
Net Unit Output
Plant Heat Rate (HHV)
Efficiency (HHV)
Effect of Stack Gas Cleanup
Reduction in Net Plant Output
Plant Rate (HHV)
Efficiency (HHV)
2,HOO psia
1,000 F ,
1,000 F
2.0 in. Hg.
U80 F
525,000 kw
11,779 kw
536,779 kw
7,661> Btu/kwh
7,836 Btu/kwh
88.5 %
21,630 kw
503,370 kw
9,235 Btu/kwh
36.96 %
9,721 Btu/kwh
35.1 %
-------
BASIC STEAM CYCLE FLOW DIAGRAM
VJl
O
3)
O
THROTTLE FLOW
3,497,193 LB/HR
T= 1000°F
P = 2400 PSIG
h = 1461.2BTU/L8
STEAM FLOW
TO REHEATER
h = 1312.9 BTU/LB
REHEAT FLOW
3,099,565 LB/HR
T = 1000°F
h = 1519.8 BTU/LB
P = 478 PSIG
XV
ljT0 © j>
© ©
FEED WATER
TO BOILER 3,497,193 LB/HR
T = 480°F
h = 464.9 BTU/LB
NOTE:( A }THRU CG) INDICATE EXTRACTION
POINTS FOR FEEDWATER HEATING
GENERATOR OUTPUT
513,221 KW
2,251,329 LB/HR
127,348 LB/HR
BOILER FEED
PUMP
11,779 KW
CONDENSER
2.0 IN. HG.
7 STAGE FEEDWATER HEATING
TOTAL STEAM EXTRACTED:
1,118,516 LB/HR
CONDENSATE
PUMP
P
CO
-------
to the overall plant heat rate during those months, raises it from 9721
to 9996 Btu/kwh.
Heat rates presented here and used throughout the study were based on
performance at rated plant output. Typical operation at 70 to 80 percent of
capacity would produce a five percent increase in these values for conven-
tional steam plants. No data are available for the corresponding factor in
operation of a gasified coal system, but it is assumed that losses associated
with less than full output will be-similar to the steam station experience.
Design Considerations - The design characteristics of future steam power
stations have been examined' in some detail.^' The principal factors affect-
ing steam station performance are unit -size, steam conditions, and environ-
mental and fuel factors all of which must be considered in the design.
Unit Size - The overall size of the electric utility system generally fixes
the size of new units or at least limits them to about 10 percent of the
total system size. Therefore, a number of large utilities could be expected
to install units of over 1000 Mw in size. However, a survey '^9.) of about
50 percent of the new plants scheduled for completion in the 1975 to 1978
time period shows no single 'unit having a capacity of over 1000 Mw and shows
an average capacity of ^9^ Mw for conventional steam units. Because of this,
the basic steam plant has been assumed to consist of a pair of 500 Mw units.
Steam Conditions - The general trend in steam conditions has been toward
higher pressure and temperatures to produce power more efficiently. When
combined with reheating and regenerative feedwater heating, significant
efficiency improvements in the basic Rankine cycle have been made possible.
However, material capabilities, additional costs, and diminishing returns on
both initial steam condition increases and number of reheats place a practical
economic limit on the degree to which these changes can be used to improve
performance. A 1000 F peak steam temperature is generally accepted as the
most practical level for the foreseeable future. The increased material costs
and poor operating experience resulting from higher temperatures are primarily
responsible for the limitation. Virtually all of the plants surveyed(^9/ have
steam temperatures at or near 1000 F. Steam pressure, on the other hand, has -
not found a universally accepted level. However, there is a trend downward
toward the 2200-2600 psi level from the 3500 psi supercritical level that was
used in about 50 percent of new plants in the early 1970's. Only 2k percent
of the conventional plants due to start operation through 1978 will use
supercritical pressures.
-51-
-------
The use of a second reheat has also declined in new systems with none
planned for the next four years. This is understandably tied to the decline
in use of supercritical pressure systems as well as the small increase in
performance afforded by the second reheat as shown in Figure lU.
Exhaust Conditions - Performance of steam equipment is highly sensitive to
available cooling water temperatures. In spite of this, over 50 percent of
currently planned stations will rely on cooling towers and about hO percent
will use cooling ponds or lakes. The minimum practical condensing temperature
is about k2 F above the dew point temperature for a wet cooling tower. This
differential results from a 10 F terminal difference in the condenser, a 20 F
drop across the tower and a minimum approach differential of 12 F. These
factors are dis'cussed by Burns and Roe'30) who also identify a wet bulb
temperature of 75 F as the average during the months of June through September
with a peak temperature of 83 F. This results in an average condenser pressure
of 3-5 in- Hg during that period and a peak (exceeded only 1 percent of the
time) of U in. Hg.
The use of cooling towers has undoubtedly had an effect on the design of
new units. The sensitivity of heat rate to changes in condenser pressure can
vary widely with plant design. The basic system selected as a standard shows
a moderate sensitivity, its heat rate increasing by 2.8 percent as condenser
pressure is increased from 2.0 to 3-5 in. Hg. This sensitivity is a function
of the ratio of steam flow to last-stage annular area. The higher the ratio
the lower the sensitivity. However, with a high ratio of flow to annular
area, it is not possible to take advantage of low condenser pressures when
they are available. Thus, it would appear that the use of cooling towers will
cause a trend toward slightly higher flow to area ratios with a lower cost per
unit output and slightly lower efficiency.
Environmental Considerations - Heat rejection and stack gas cleanup for S02
have been the primary areas of concern in preventing the degradation of the
environment. The first of these is highly dependent on the plant location so
that the method of heat rejection will vary widely. Performance character-
istics assume the use of a mechanical draft cooling tower.
Performance of stack gas cleanup systems are much more difficult to
characterize because of the widely varying results of the various test systems.
Therefore, this study did not attempt to select a particular design nor was
the topic reviewed in great detail. The net power losses attributable to the
combination of cooling towers and stack gas cleanup are given ^9 at between
6.h and 6.7 percent of net capability. Another estimate, ^31) places losses
due to stack gas cleanup at between two and seven percent depending on the
process and design philosophy. A value of five percent was chosen to be
-52-
-------
FIG. 14
EFFECT OF REHEATING
1 2
NUMBER OF REHEATS
R04-12-1
-53-
-------
representative and shows good correlation with data for a regenerable scrubbing
system. (32)
Auxiliary Power - While not a major factor in plant efficiency or operating
cost, the power consumed by auxiliary equipment can be between three and six
percent of the net steam cycle output. When comparing different types of
plants having different auxiliary power characteristics it is important that
these be quantified so that a meaningful comparison of plant efficiency and
operating cost can be made. Table 8 presents a typical breakdown of auxiliary
power required for a coal-fired steam plant. To permit uniform application
of these factors to each of the other systems to be considered, the auxil-
iaries have been broken into five categories. The portion shown for coal
transport and pulverization is used only in the basic steam plant. In the
gasified-coal plants a separate estimate of power for coal preparation has
been made. This is included in the bookkeeping for the integrated system
and, in the case of the BCR gasifier , amounts to about 2.27 percent of the
net station output. Boiler fan power has been deducted from both the basic
steam plant and the gasified coal-fired steam plant. All the advanced-cycle
plants run with a back pressure on the gas turbine sufficient to meet the
pressure drop across a waste heat boiler. Condensate pump power has been
deducted from all plants based on the net steam power developed. The power
for cooling water pumps has also been applied to the steam power, but in the
gasified coal systems an adjustment has been made to account for the cooling
water requirements of the cleanup system. The miscellaneous plant losses
include lighting and ventilating as well as copper (electrical resistivity)
losses and apply equally regardless of the primary power cycle.
GASIFIED COAL-FIRED STEAM SYSTEMS
A potential alternative to stack gas cleanup is to clean the fuel prior
to firing. This can be done by first gasifying the fuel and then cleaning it
in one of the many systems described in other sections of the report. The
primary advantage of such a system over stack gas cleanup is the lower quantity
of gas (only about Uo percent of the stack gas quantity) that must be processed.
Also, if the cleanup system is run at pressure, gas volume is further reduced
and the sulfur compoiinds can be removed more readily by a wide variety of
processes. Finally, sulfur leaving the gasifier is mostly in the form of H2S
which, after separation, can be more easily converted to elemental sulfur than
can SOp.
-------
Table • 8
AUXILIARY POWER REQUIREMENTS
Percent of Net Steam Cycle Output
*Coal Pulverizing and Transport
Boiler Fans
Condensate Pumps .22$
Cooling Water Pumps 1.01$
Miscellaneous Plant Los'ses .57$
*Applicable only to direct coal fired boiler. Estimated separately
for gasified coal systems based on gasifier requirements.
-55-
-------
System Configuration
The general system arrangements that would be used in a gasified coal-
fired steam system are shown in Figures 15(a) and 15(b). Witfi a low-pressure
gasifier and cleanup system (Figure 15(a)) the remainder of the plant con-
sists of a gas-fired boiler generating steam for use in a conventional steam
cycle as was shown in Figure 13. The low fuel heating value will require the
burners to handle a volumetric flow of fuel some six times., that of a natural
gas-fired boiler. While conventional boiler design practices are applicable
to a low-Btu system, lower flame luminosity and higher stack gas flows as
compared to coal and natural gas fired boilers will result in a different
split between radiation and convection heat transfer surface required. As
a result, use of low-Btu gas in existing boilers would require some derating
and modification to maintain good control of steam conditions. However, new
designs would display little outward difference from existing boilers, For
pressurized operation of the gasifier and cleanup system it is necessary to
compress the gasifier inlet air as shown in Figure 15(b). After cleanup, the
resultant high-pressure gas can be expanded through a turbine to recover some
or all of the work of compression. If the temperature of the gas entering the
turbine is high enough, power can be extracted from the let-down turbine.
Depending on the ratio of product gas to air flow into the gasifier, a fuel
temperature in the range of 600 to 800 F will be sufficient to sustain the
compressor and a greater temperature will produce net output power from the
let-down turbine.
System Performance - For comparison, the preliminary estimates of the per-
formance of the steam systems are given in Table 9.* Two gasifier/cleanup
combinations were considered; a Bu Mines/low-temperature system and a BCR/
high-temperature system. Steam cycle performance, similar to that in Table 7
was used for all system combinations. Steam boilers firing gasified fuel
have lower efficiency than direct-fired boilers because of the steam injected
into the gasifier during the gasification process. Some of the steam reacts
with coal to produce ^ an& CHr which form H^O upon combustion and results
in higher moisture losses. That steam which does not react with coal and is
not removed from the fuel gas passes through the boiler to the stack and
causes a slight increase in the stack gas sensible heat loss. The boiler
efficiency estimates presented in Table 9 were calculated using standard
boiler practice. ^
Numbers in Table 9 are early estimates and are not entirely consistent
with numbers in Table 7.
-56-
-------
FIG. 15a
GASIFIED COAL-STEAM CYCLE SYSTEMS
COAL
AIR
i
LOSSES
HA9IPIFR
|
I
POWER
_
i
GENERATION
[LOSSES
CLEAN U
P
t '
POWER
STACK GAS
onn CD
STEAM
STEAM & AIR PREHEAT
^
STEAM CYCLE
POWER
^
OUT
LOW PRESSURE GASIFIER
FIG. 15b
STEAM & AIR PREHEAT
»- HEAT FOR STEAM
GENERATION
*•
BOILER
CTPAIV/I PYPI P
POWER
OUT
GENERATOR
POWER
OUT
AIR IN
TURBINE
HIGH PRESSURE GASIFIER WITH LET-DOWN TURBINE
R03-1-1
-57-
-------
TABLE 9
STEAM POWER STATION. EFFICIENCY COMPARISON
I. W. Kentucky No. 9 Coal
Direct Firing
Direct Firing/SGC
BuMines Gasifier/
Selexol Cleanup
BuMines Gasifier/
Cleanup @ 1000 F^ )
II. Illinois No. 6 Coal
Direct Firing
Direct Firing/SGC .
BCR Gasifier/
Selexol Cleanup
BCR Gasifier/Cleanup
@ 1000 F
BCR Gasifier/Cleanup
@ 1800 F^)
BCR Gasifier/Tur"bo-Compressor
Cleanup % 1500 F
Steam, Cycle Boiler^1' Gasifier Station
1*3.5 .
1*3.5
1*3.5
1*3.5
1*3.5
1*3.5
86.9
82.55(3)
8U.8 77.0
85.2 8U. i(5)
88.5
81*.1(3)
35.5
33.8
26.7
29.3
36.2
31*. 1*
1*3.5
U3.5
1*3.5
1*3.5
' 81*.3
85.5
86.8
85.3
88.0
89. 1
87.8
(5)
88.3
20.3
31.3
31.2
31.8
(l) 300 F stack, 59 F ambient, 20% excess air for direct firing, 10$ excess
air for gasified fuel firing.
(2) 6% auxiliary power loss
(3) Includes 5$ loss for SGC system
(U) Gasifier exit temperature
(5) Neglects fuel and steam requirements for cleanup and regeneration
systems
-58-
-------
The principal gasification/cleanup process losses are due to unconverted
carbon from the coal and the need to quench the fuel gas to remove tars and/or
to cool the gas prior to desulfurization. For the purpose of making steam
system performance comparisons, an effective gasifier efficiency was used.
This efficiency is defined as the net energy delivered to the steam boiler
divided by the gross input energy in the coal. The net energy delivered to
the boiler includes fuel gas sensible and chemical energy, plus an adjustment
for heat added directly to (or removed from) the steam cycle. The hot
cleanup system assumed is a fictitious one with no parasitic losses (fuel,
power, or steam) and which removes only HrjS, .COS, and WHg. Direct comparison
of the efficiencies for the two types of gasifiers would be misleading because
of the different types of coals used at the different process operating con-
ditions and assumptions.
The performance estimates presented in Table 9 indicate that the effi-
ciency of a steam station using a Bu Mines gasifier and low temperature
cleanup would be 8.8 points below the efficiency of a direct fired station.
This decrement would be reduced to 6.2 points if a 1000-F cleanup system
could be used. Similarly, the efficiency decrement for a BCR gasifier with
low-temperature cleanup would be 5.9 points and this would be reduced to U.9-
5.0 points for cleanup at 1000-1800 F. The addition of stack gas cleanup
reduces the differential by 1.7-1.8 points.
In the preceding estimates it was assumed that the power required to
drive the booster compressor which feeds the gasifier could be supplied by
an expander turbine with matching output power. An equivalent assumption
would be that the gasifier is operated at low pressure so that the compressor
and turbine powers would be negligible. The validity of this assumption was
tested for the BCR gasifier operating at 500 psi. Fuel gas from the gasifier
was cooled to 1-500 F, desulfurized, and expanded through the turbine to- 20
psi before being burned in the boiler. The 1500 F turbine temperature limit
was imposed because of the difficulty of cooling a fuel expander turbine.
(Cooling air could not be used because combustion would occur in the turbine.)
In this case the expander turbine was capable of delivering net power amount-
ing to approximately 3 percent of the net station- output. The net station
efficiency in this case, was 05-«6 points higher than for the previous case
which did not incorporate a turbo-compressor.
It' appears from this analysis that the benefits to be derived by using
hot cleanup in conjunction with a BCR gasifier and a steam power cycle are
marginal in comparison to low-temperature cleanup, whereas the benefit for a
Bu Mines-type gasifier would be significant. This is because the heat .
recovery in the BCR/low temperature went into steam at essentially cycle
efficiency just as does the sensible heat in BCR/high-temperature. The
-59-
-------
Bu Mines/low- tempera ture system has essentially no heat recovery. In any
event, it would tie extremely difficult to envision such a system with a net
efficiency much over 30 percent. This result is consistent with an estimate
of 27 percent presented by Commonwealth Edison Company.'3*0
ADVANCED COGAS SYSTEMS
Several studies^ ' of advanced combined gas and steam (COGAS) cycles
conducted by the United Technologies Research Center has shown that these
cycles offer the potential of high performance with low capital investment.
They are particularly well adapted to use with gasified fuels where gasifier
a-nd cleanup systems operate at pressure. A simplified schematic of a COGAS
system is shown in Figure 16.
In the above arrangement, the turbine exhaust gas is used directly to
raise steam in an unfired boiler. This arrangement has been shown '3J to
have the potential for highest efficiency and lowest overall cost. For a
given stack temperature, Tgt, the efficiency of a COGAS system is a function
of four primary variables: the gas turbine efficiency, f]gtS the steam cycle
efficiency, 7]s; the turbine exhaust temperature, Tex; and supplemental fuel,
Ws. The relationship is:
n = (3D
'cc
Other terms in Equation 31 are engine fuel flow rate, Wf; and the ambient
temperature, Ta. The performance of a simple waste-heat-recovery system with
no . supplementary firing is then:
n t + [d - ng1
W3 T ~T t
• W T -T ^ ^
f ex a s
w
1 + ^
V
T -T
ex Lst
This equation points out the importance of the basic gas turbine efficiency
and of the cycle temperatures on the ability to utilize the gas turbine
exhaust heat. Exhaust temperature has a two- fold effect on the steam cycle
contribution to combined cycle efficiency. It not only improves the ratio
of available to total heat in the gas stream but higher temperatures allow
the use of higher performance steam cycles. For these reasons, it has been
relatively common practice in currently installed combined-cycle systems to
use supplementary fuel to increase boiler inlet temperature and improve over-
all cycle efficiency.
-60-
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WASTE-HEAT COMBINED GAS AND STEAM TURBINE SYSTEM
ON
AIR
COMPRESSOR
TURBINE
BURNER
FUEL
POWER TURBINE
T-2000 F
P~14 ATM
T~875 F
STEAM
BOILER
T~300 F
TO STACK
T~775 F
P~60 ATM
STEAM
TURBINE
L
PUMP
CONDENSER
ELECTRIC
GENERATOR
70 MW
ELECTRIC
GENERATOR
50 MW
-------
Similar changes in steam cycle contribution to performance could result
from changing either the gas turbine pressure ratio or turbine inlet tempera-
ture. It is the objective of this section to illustrate the interdependency
of the gas turbine and steam cycles and to provide a framework for the
integration of the gasifier and cleanup systems into an advanced power
station.
Gas Turbine Cycle
Considering overall COCAS system performance, the single parameter that
can improve "both gas turbine and steam cycle performance is gas turbine inlet
temperature. In the following paragraphs the various approaches to improving
this important parameter are described.
Turbine Inlet Temperature - The current state of the art in gas turbine design
has seen the aerodynamic efficiencies of the turbine and compressor reach
90 percent or better. While higher aerodynamic efficiencies are desirable,
the single most effective method of increasing the overall performance of the
turbine is to increase the turbine inlet temperature (TIT). The historical
trend in TIT's is depicted in Figure 17. After high-temperature technology
is tested and proven in research and development engines, it is first applied
to military aircraft engines, then to commercial aircraft engines, and finally,
some 6 to 8 years after the E&D engine, to industrial engines. A close look
at Figure 17 will also show that the technology transfer time, i.e., from R&D
engine to industrial engine, is becoming shorter. Thus, the advances in cool-
ing techniques and materials first used on research engines will appear much
sooner in industrial engines than they have in the past. It may be, in fact,
that certain cooling techniques, e.g., water-cooled blades, would bypass the
aircraft application altogether.
A number of blade cooling techniques and improved materials could be
considered for use in future power systems. For this study, five combinations
of materials and cooling techniques were selected as representative of future
technology and their performance was studied in detail for turbine inlet
temperatures of 2200, 2^00, 2600, and 2900 F. The five combinations are:
I. Conventional air cooling
II. Conventional materials but vane cooling air precooled to 300 F
III. Ceramic vanes and conventionally cooled blades
IV. Ceramic vanes and blades
V. Zero bleeds and leakage
-62-
-------
ESTIMATED TURBINE INLET TEMPERATURE PROGRESSION
3600
U)
i
COMMERCIAL
TRANSPORTS
INDUSTRIAL
\PPLICATIONS
1200
1950
1990
-------
Cooling Case I corresponds to using conventional cooling techniques and
conventional materials maintained at a maximum temperature of 1500 F for long
lifetime, baseload use. In a typical gas turbine, a number of bleed streams
of air are extracted from the compressor for cooling purposes in the hot
turbine sections. Most of the cooling air is used in the first pair of vanes
and blades (see Figure 18). This cooling air is at a temperature correspond-
ing to the compressor exit and is a function of cycle pressure ratio. Cooling
air flow for the second pair of vanes and blades is on the order of 1/3 of the
flow to the first pair. The remaining bleed flows for disc, bearing, and
miscellaneous cooling typically amounts to 6 percent. For Cooling Case II the
cooling flow to the vanes is independent of cycle pressure ratio and varies,
for the first vanes, from 7.8 percent at 2200 F to 11.0 percent at 2900 F.
Case III incorporates ceramic vanes (no vane cooling), but the remaining
cooling flows are the same as for Case I. Case IV incorporates ceramic vanes
and blades (no vane or blade cooling), but the required cooling flows to the
discs, bearings, etc. remain at 6 percent. Case V corresponds to the theo-
retical limit with no cooling bleeds of any kind.
For first-generation systems, the use of conventional impingement type
blade cooling at a TIT of 2200 F, is judged to be representative of engines to
be used in the late 1970's. For second-generation systems, the use of ceramic
vanes with conventional blade cooling and a turbine inlet temperature of 2600F
was selected. The effect of these combinations on overall combined-cycle
system performance (distillate oil-fired systems are used for comparison) is
summarized in Figure 19 and is discussed in detail in a later section on COGAS
performance. The effects of temperature and cooling configuration on simple
cycle performance are shown in Figures 20 and 21. In these figures the gas
turbine efficiency is plotted against net power per unit airflow or specific
power. This is a measure of the amount of work which can be done by a given
size machine and is, indirectly, a guide to cost. A machine with a high spe-
cific power will generally have a lower cost in $/kW than a machine of lower
specific power. It is interesting to note in Figure 20 that with conventional
cooling the incremental improvements in efficiency resulting from increases in
turbine inlet temperature are small. The major effect is an increase in spe-
cific power. This same general trend is apparent in Figure 21, however, the
reduced cooling flows associated with ceramic materials result in a more sig-
nificant efficiency improvement. It is only when the gas turbine is used in a
COGAS system that the full benefit in improved cycle efficiency as well as
specific power associated with higher turbine inlet temperatures can be real-
ized.
-------
FIG. 18
EFFECT OF CYCLE PARAMETERS
ON TURBINE COOLING AIRFLOW REQUIREMENTS
CONVENTIONAL AIR COOLING TECHNOLOGY
24
O
QC
O
8
LLJ
DC
Q_
s
O
O
LL
O
5?
I
O
_J
LL
cc
O
O
u
20
TURBINE INLET
TEMPERATURE = 2900F
16
12
FIRST
VANE
COOLING
FIRST
BLADE
COOLING
8 16 24
COMPRESSOR PRESSURE RATIO
32
40
N07-38-2
' -65-
-------
FIG. 19
COGAS STATION PERFORMANCE
DISTILLATE FUEL
56
I
I
I
o
LU
O
LL
LL
LU
2
O
p
co
52
48
44
THEORETICAL LIMIT
CERAMIC VANES AND
BLADES
CERAMIC VANES
PRECOOLED AIR
CONVENTIONAL COOLING
40
2200 2400 2600 2800 3000
TURBINE INLET TEMPERATURE - F
3200
-66-
-------
FIG. 20
GAS TURBINE PERFORMANCE MAP
40
DISTILLATE FUEL OIL
CONVENTIONAL AIR COOLING
35
I
I
O
UJ
y
LL
U.
LU
z
m
cc
CO
COMPRESSOR
PRESSURE RATIO
24
--—V
30
25
2200 240°
TURBINE INLET TEMPERATURE, F
20
80
I
I
I
100 120 140 160
NET POWER PER UNIT AIRFLOW - KW/LB/SEC
180
-67-
-------
PERFORMANCE OF 2600 F GAS TURBINE
DISTILLATE FUEL OIL
co
i
3)
§
I
o
LU
O
u.
"- >
UJ £
LU X
40
35
30
25
20
100
40
32
40
CERAMIC VANES
40
32 CERAMIC VANES
AND BLADES \_-24
16
THEORETICAL
LIMIT
CONVENTIONAL
COOLING
8 8
120
140
160
180
200
220
240
GAS TURBINE POWER PER UNIT AIRFLOW
KW/LB/SEC
-------
Steam System ' • • •'.•'•'..•
Although the efficiency term for the steam-cycle in equation (31) is mod-
ified by several coefficients, it is apparent that it is desirable to achieve
as high an efficiency as practicable. There are, however, constraints on the
steam cycle conditions which are external to the steam power system. These
are gas temperature to the boiler (gas turbine exit temperature) and the stack
temperature. The latter is set by stack gas dew point and is a function of
corrosion and other practical considerations. It is usually not much below
300 F, even when a clean fuel is used. While a lower stack temperature could
be considered, it represents a study that must be done for each specific sys-
tem. For this study a 300 F stack temperature was chosen as a standard and
all systems compared on that basis. It has been found^' that a pinch temper-
ature difference between steam and exhaust gases of 100 F results in the most
acceptable heat exchanger design. Thus, for systems utilizing waste heat re-
covery boilers, the maximum steam temperature is 100 F less than the gas tur-
bine exhaust temperature. At the low-temperature end of the waste-heat-recov-
ery unit, a pinch of 50 F is acceptable in the economizer between exhaust gas
and feedwater. Thus, to maintain a 300 F stack temperature and utilize all of
the available heat in the turbine exhaust, regenerative feedwater heating is
limited to a maximum of 250 F.
High-performance steam cycles such as a 2UOO psi 1000 F/1000 F cycle
would generally require exhaust firing to raise the gas temperature to a level
sufficient (1300-1^00 F) to accommodate both the initial superheat as well as
the reheat section. Without regenerative feedwater heating, such a cycle
would have an efficiency of about 39 percent. As the gas temperature into the
waste-heat boiler is reduced, steam temperature must be decreased and cycle
performance goes down. For example, a 1250 psi 900 F cycle without reheat
would have an efficiency of about 33 percent. However, even at this reduced
steam temperature it is necessary to have a gas turbine exhaust temperature of
1100 to 1200 F to produce a stack temperature of 300 F. Lower exhaust temper-
atures result in higher stack temperatures for a fixed steam cycle. This ef-
fect is shown in Figure 22. When using a single-pressure steam system, the
gas side temperature profile pivots about the evaporator pinch point. Only
with increased gas temperature can the stack temperature be reduced to 300 F.
To relieve this problem, a two-pressure steam cycle is used. As shown in
Figure 22, by adding a second low-pressure cycle it is possible to utilize the
exhaust heat regardless of temperature. This is done at the expense of steam
cycle performance with the efficiency being about 30 percent.
This basic difference in philosophy between operation of a conventional
steam cycle and a COGAS steam cycle may require further clarification. In the
conventional steam station, the use of regenerative feedwater heating results
-69-
-------
TEMPERATURE DISTRIBUTION IN WASTE HEAT RECOVERY BOILER FOR
OIL-FIRED COGAS STATION
MID-1970'S TECHNOLOGY
TURBINE INLET TEMPERATURE = 2200 F
COMPRESSOR PRESSURE RATIO = 16
FIG. 22
1000
800 -
\
SUPERHEATER \
EVAPORATOR
EVAPORATOR
L.P.
EVAPORATOR
H.T. ECONOMIZER
ECONOMIZER
L.T. ECONOMIZER
LIMIT OF HEAT RECOVERY WITH
SINGLE-PRESSURE STEAM SYSTEM -
u- 600 -
UJ
QC
UJ
40 60
% HEAT RECOVERED
100
R09-55-1
-70-
-------
in exhaust gas temperature leaving the economizer section of the boiler in
excess of 500 F and as high as 700 F. The remaining heat (above 300 F) is
removed in an air preheater and, in effect, returned to the boiler. In the
COGAS system it is not possible to use the stack gas in an air preheater
since this would be harmful to gas turbine performance, increasing compressor
power and raising compressor operating temperatures. The only means of re-
ducing stack gas temperature to 300 F is to use the stack gas in an economizer.
As a result, regenerative feedwater heating must be abondoned (at least above
250 F) and the steam cycle efficiency is decreased. However, since the feed-
water heating is now provided by heat that would have been rejected up the
stack, and also the steam that would have been used for that feed heating can
now be expanded in a turbine, the net effect is an improvement in the overall
boiler/steam cycle performance.
The large increase that is possible in steam cycle performance due to the
foregoing is sometimes used to justify the use of supplementary firing. For a
distillate fueled system, the effects of supplementary firing are shown in
Figure 23. It can be seen that at lower turbine inlet temperatures a small
improvement in efficiency is possible, but it is at the expense of added com-
plexity in both the boiler and steam cycle. A similar study was conducted for
a gasified coal system and it was shown that the gains due to supplementary
firing were even less than for the distillate system. This is discussed in a
later section, but it is the basis for considering only waste-heat-recovery-
type COGAS systems.
COGAS Performance
To establish a performance base for the integrated gasified coal/COGAS
stations, a series of calculations were made for a distillate fuel fired sys-
tem. For each of the five combinations of materials and cooling techniques
mentioned earlier, performance was evaluated over a temperature range of 2200
to 2900 F and over pressure ratios of interest. The results are presented in
Figures 2h through 28. In all cases the steam cycle was optimized to give a
stack temperature of 300 F while maintaining a steam temperature of 100 F
less than turbine exhaust to a maximum of 1000 F. The steam cycle condenser
pressure is U in. Hg. abs. With higher gas turbine exhaust temperatures (over
1200 F) a single pressure, non-reheat steam cycle is used since those condi-
tions permit a 300 F stack temperature. Boiler supplementary firing was not
used. In general, the peak efficiency point occurs with a steam temperature
slightly below 1000 F and with a two-pressure steam cycle.
It is interesting to compare simple-cycle and COGAS performance. Refer-
ring to Figures 20 and 2h at peak COGAS efficiency, it can be seen that the
steam cycle adds about 13 points to the corresponding simple-cycle performance.
-71-
-------
FIG. 23
COGAS STATION PERFORMANCE WITH FIRED BOILER
DISTILLATE FUEL OIL
CONVENTIONAL AIR COOLING
50
TURBINE INLET TEMPERATURE, F
2900
45
X
I
i-u
r, 40
UJ
2'
35
30
I
I
20 40 60
FUEL FLOW TO GAS TURBINE - '
80
100
, OF TOTAL
-72-
-------
FIG. 24
COGAS STATK3N PERFORMANCE
WITH CONVENTIONAL AIR COOLING
DISTILLATE FUEL OIL
56
I
X
52
o
z •
cc
o
LL
cc
LLJ
Q.
Z
o
CO
<
C3
O
O
COMPRESSOR PRESSURE RATIO
48
44
2600
40
18
2900
2400
2200
TURBINE INLET TEMPERATURE, F
„• I 1
100
150 200 250
NET POWER PER UNIT AIRFLOW - KW/LB/SEC
300
350
N07-36-1
-73-
-------
COGAS STATION PERFORMANCE
WITH PRECOOLED AIR TO VANES
DISTILLATE FUEL OIL
56
-p-
i
o
*sj
CO
00
_ 52
>
I
I
LU
o
DC
O
LL
DC
00
to
O
o
48
44
40
100
COMPRESSOR PRESSURE RATIO
40
2200
2400
TURBINE INLET TEMPERATURE, F
150 200 250
NET POWER PER UNIT AIRFLOW - KW/LB/SEC
300
350
P
tvj
01
-------
FIG. 26
COGAS STATION PERFORMANCE
WITH CERAMIC VANES AND
CONVENTIONAL COOLED BLADES
DISTILLATE FUEL OIL
56
I
£ 52
s?
I
UJ
o
CC
O
LL
cc
LLJ
Q_
g
<
CO
CO
o
o
o
48
44
40
COMPRESSOR
PRESSURE
RATIO
2200
2400
2600
8
2900
I
I
TURBINE INLET TEMPERATURE, F
I
100
150 200 250 300
NET POWER PER UNIT AIRFLOW - KW/LB/SEC
350
N07-36-2
-75-
-------
COGAS STATION PERFORMANCE
WITH CERAMIC VANES AND BLADES
DISTILLATE FUEL OIL
56
I
I
52
CJ
cc
o
en 48
uj
Q.
<
CO
44
O
O
40
100
COMPRESSOR PRESSURE RATIO
2200
" 2600
2400
TURBINE INLET TEMPERATURE. F
2900
I
I
150 200 250 300
NET POWER PER UNIT- AIRFLOW - KW/LB/SEC
350
o
P
to
-------
COGAS STATION PERFORMANCE
WITH ZERO BLEEDS AND LEAKAGE
DISTILLATE FUEL OIL
56
I
I
as
o
z
<
2
cc
o
LI-
CC
LU
0.
z
o
I-
co
<
(3
O
O
COMPRESSOR PRESSURE RATIO
40
32
52
48
44
120
2200
2400
TURBINE INLET TEMPERATURE, F
40
I
I
100
150
200 250 300
NET POWER PER UNIT AIRFLOW - KW/LB/SEQ
350
400
2
O
O
u
Tl
P
N>
00
-------
At the lower turbine inlet temperatures, simple-cycle efficiency continues to
improve up to a pressure ratio of more than 24:1. In the combined system,
such a high pressure ratio lowers exhaust temperature to a point where steam
cycle performance falls off badly.. Thus, the COGAS performance peaks at a
much lower pressure ratio, reflecting a compromise between gas turbine and
steam system performance. At higher turbine inlet temperatures, this effect
tends to disappear. Peak simple-cycle efficiency then occurs at a pressure
ratio corresponding to an exhaust temperature that permits good steam cycle
performance. Thus, as turbine inlet temperature is increased, it not only
permits a greater gas turbine output, but allows both gas turbine and steam
cycles to perform at peak efficiency.
These same general characteristics apply to systems using a low-Btu gas.
To permit evaluation of the integrated system, it was necessary to select a
gas turbine pressure since it in turn sets gasifier and cleanup system opera-
ting pressure. Based upon the parametric studies, this was chosen to be the
pressure ratio producing maximum specific power without a significant reduc-
tion in efficiency. For the conventionally cooled, 2200 F, first-generation
system, a value of 16:1 was selected. A 2^:1 pressure ratio was chosen for
the second-generation, 2600 F engine with ceramic vanes and conventionally
cooled blades.
Integrated System Performance Evaluation - The combination of gasifier, clean-
up system, gas turbine and steam cycle give rise to a very large number of
possible configurations, especially when the necessary auxiliaries such as
heat exchangers, booster compressors, boilers, etc. are taken into account.
In order to investigate many of these combinations, a versatile computer model
was developed at UTRC under Corporate sponsorship which allows great flexibil-
ity in analyzing the integrated systems. The State-of-the-Art Performance
Program (SOAPP) is basically a bookkeeping system that permits the user to
assemble individual modules into an integrated power system (see Appendix A).
A number of different configurations were used during the course of the study,
the most important of which are shown in Figures 29 through 31.
The configuration shown in Figure 29 was used extensively in preliminary
integration studies. The left column of components comprises the main ele-
ments of a waste-heat-recovery COGAS power system. The gas generator is a
dual spool machine, i.e., separate high-pressure and low-pressure shafts oper-
ating at different speeds with provision for intercooling and for bleeding the
necessary air for the gasifier. The main heat-recovery boiler also has pro-
vision for supplementary firing and the steam cycle can be selected to suit
the exhaust gas temperature. In general, steam raised elsewhere in the system
is used for process steam; however, provision is made to supplement the steam
-78-
-------
FIG. 29
INTEGRATED COGAS/COAL GASIFICATION POWER STATION
AIR
ELEC. GEN. I POWER TURBINE
SAT.
STEAM
STEAM CYCLE MAIN BOILER
B.F.W.
-*-©
STACK
L.P. BOILER
T
COAL
AFTER
COOLER
BOOSTER COMP.
(OPTIONAL)
STACK
L.P. PROCESS
BOILER (OPT.)
COLD SIDE
HEAT EXCH.
GASIFIER
H.P. BOILER
HOT SIDE
FUEL REGEN.
CLEANUP
BOOSTER COMP.
(OPTIONAL)
AIR
COLD SIDE
FUEL REGEN.
LEGEND
— STEAM/WATER
COAL/FUEL GAS
AIR/COMBUSTION PRODUCTS
-79-
-------
INTEGRATED COGAS/COAL GASIFICATION POWER STATION
WITH LOW TEMPERATURE CLEANUP SYSTEMS
FIG. 30
ELECT.
GEN.
STEAM
CYCLE
HIGH COUP.
(CPHI)
GASIFIER
AIR
EXTRACTION
IFSBNI
BURNER ASSY
IBNINI
(PUMP)
(BURN)
COMP. DRIVE
TURBINE
(TBHII
ITBLOI
PWR. TURBINE
(TRAN)
(TBPT)
(DIFF)
DUCT BURNER
IEXLS)
IFIREI
IAJE)
MAIN
STM. BOILER
ISBI
INZEI
I
STACK
L.P. BOILER
AFTER
IAF
:OOLER
FUEL GAS
fc
FUEL GAS
STEAM
FROM N.P.B.
BFW
TO H.P.B.
STEAM TO
MAIN -^
BOILER
F.W. FROM ^
MAIN BOILER
FLOW SPLIT
(PBLD)
(GXITI
COMPRESSOR
(PIPE)
IBOSTI
GASIFIER
(GASFI
H.P. BOILER
IHPB)
M.P. BOILER
IBMP)
^ COAL 1
-• STEAM
(SOURCE
OPTIONAL)
GAS FOR
COAL TRANSPORT
(TXITI
TRANSP. GAS
REGEN
(HFN3I
FUELREGN.
(HFC6I
FUEL REGEN
IHFH5I
TRANSP GAS
REGEN
IHFNI)
L.P. BOILER
ILPBI
AFTERCOOLER
IBFCLI
PURIFIER
IGPURI
IFSPBI
TRANSP. GAS
- — •. REGEN.
INFC4)
TRANSP. GAS
.-._ REGEN
INFC2I
COMPRESSOR
(BST)
-80-
-------
INTEGRATED BCR/CONSOL GASIFICATION POWER STATION
HIGH TEMPERATURE CLEANUP SYSTEMS
FIG. 31
-------
generated -in the main boiler with steam raised in the gasifier-cleanup part of
the system. The center column represents the gasifier and cleanup system.
Air at compressor outlet conditions is cooled in a boiler and aftercooler to a
temperature that is dependent on boost compressor requirements. If necessary,
an air preheater can be supplied with heat from the secondary burner in the
right hand column. The steam consumed by the gasifier is raised in either
the main or H.P. boiler. Gasifier jacket heat is used for feedwater and steam
generation. Heat from the gasifier effluent can be recovered in the H.P.
boiler and in a fuel gas regenerator when used with a low-temperature cleanup
process. The requirements for steam at various pressures in the cleanup system
are used to control steam generation rates and to determine net clean fuel gas
and power output.
In general, any of the modules can be rendered inoperative resulting in a
great degree of flexibility within a single configuration. For example, Fig-
ure 30 shows the configuration used to represent the BCR/low-temperature
cleanup and both the low- and high-temperature cleanup in conjunction with the
BuMines gasifier. In the case of the BuMines systems, many of the modules are
not used because of the tar in the gasifier outlet stream which prevents heat
recovery from that stream. In this configuration, additional boilers, heat
exchangers and a boost compressor for the BCR transport gas stream (right-hand
column) were added. It was more attractive to regenerate the transport gas as
a means of satisfying the gasifier heat balance than either raising steam with
a higher degree of superheat than that from the main stream cycle or increas-
ing the gasifier air/fuel ratio.
The BCR/high-temperature cleanup system is represented by Figure 31. The
differences between this and the low-temperature case are due to the higher
temperature cleanup (no heat recovery) and the need to remove CC>2 from the
transport gas stream for use in the regeneration of some of the high-tempera-
ture absorbers. The system is designed to raise process steam in the L.P.
boiler with steam from both the M.P. and H.P. boiler used to supplement the
main steam cycle.
Integrated System Studies - Prior to selection of the representative configu-
rations for detailed evaluation, a number of general studies were performed to
determine trends that will be generally applicable to all systems. Much of
this work deals with the utilization of available process heat and points up
the value of regeneration in the fuel stream and the advantages that accrue
from using energy at the full cycle efficiency. Wherever possible, it is
desirable to extract process heat only after expanding the working fluid
through the gas turbine. Because of the large amount of process heat used in
the gasification and cleanup systems, these factors are magnified as compared
to a distillate-fired COGAS system.
-82-
-------
The analyses discussed in the following paragraphs were based on prelim-
inary data for gasification and were intended to identify the trends in per-
formance due to variations in operating characteristics. Thus the absolute
values of performance may differ from the final values obtained for the de-
tailed analyses of selected systems.
Boiler Firing Studies - Using preliminary data for a BCR type gasifier with a
representative low-temperature cleanup system (Selexol), a series of perfor-
mance analyses were conducted. One part of that series deals with the use of
clean product gas to increase steam boiler gas inlet temperature. As dis-
cussed previously the evaluation of fired vs. unfired boilers is relatively
complex and cannot be answered simply. The flow diagram used in these calcu-
lations is basically that shown in Figure 29 with the secondary burner stream
omitted. It was unnecessary to aftercool following the M.P. boiler and the
boost compressor is used to raise the air temperature to satisfy gasifier in-
let air temperature requirements of 800 F. The performance of a first-gener-
ation COGAS system with various steam cycles is presented in Figure 32. With
no boiler firing and a two-pressure steam cycle, peak efficiency occurs at an
exhaust gas temperature to the boiler of about 925 F. This corresponds to a
pressure ratio of 16:1. Steam conditions for the other cycles considered are
noted on the figure. In all cases, supplementary firing was required to bring
the boiler gas temperature to a high enough level to maintain the designated
steam conditions and to fully utilize the heat in the gas stream by maintain-
ing a 300 F stack gas temperature. At boiler gas temperatures lower than this
required level the stack temperature would exceed 300 F at which point overall
efficiency falls off quite rapidly. It is apparent that because of the sup-
plementary fuel that must be burned to utilize the high-performance steam
cycles, the resulting station efficiency is no better than when a relatively
simple, low-cost steam cycle is used. As a result of this analysis, integra-
tion efforts were concentrated on the use of simple steam cycles and unfired
boilers.
Steam Conditions - Using the above system configuration, the effect of varying
steam cycle pressure was investigated. The net efficiency for the station
having an unfired boiler with a two-pressure steam system with the high pres-
sure set at 1250 psi is shown in Figure 32 to be 33-7 percent. Performance
at this condition was selected as a base-line point for comparing other cycle
variations. The cycle parameters corresponding to that point are given in
Table 10. The effect of changing gas turbine compressor pressure ratio and
the throttle pressure levels of the high-pressure portion of the steam cycle
from the reference level is also shown in Figure 33. Station efficiency is
relatively insensitive to compressor pressure ratio about the reference point.
Similarly, efficiency is relatively insensitive to steam pressure as it is in-
creased from 1250 to l800 psi. The l800 psia level represents an upper limit
-83-
-------
FIG. 32
EFFECT OF STEAM CYCLE ON STATION PERFORMANCE
q
i
_i_
-vO
o~-
!
O
111
y
LL
LL
LU
O
H-
H
00
37
35-
33
31
29 -
27 ._
'ion
TURBINE INLET TEMPERATURE - 2200F
COMPRESSOR PRESSURE RATIO = 16. UNLESS NOTED
SINGLE PRESSURE STEAM SYSTEM. UNLESS NOTED
STACK TEMPERATURE = 300 F
STACK TEMPERATURE >300 F
NO BOILER FIRING
VARIABLE COMPRESSOR PRESSURE RATIO
STEAM TEMP. = GAS TEMP - 100
1250PSI H-P STEAM
VARIABLE PSI L-P STEAM
,950F/950F/l80n PSI/40.5%
, 1000F/1000F/1800 PSI/40.9%
IOOOF/1000F/2400 PSI/4I.2%
L
1000 1200 MOO
BOILER GAS TEMPERATURE.F
1600
1800
NCK.t -78-2
-Qh-
-------
Table 10
DESIGN CHARACTERISTICS FOR REFERENCE INTEGRATED COGAS POWER STATION
Gas Turbine-
turbine inlet temperature
compressor pressure ratio
air bleed for gasifier
Steam Cycle (Unfired Boiler)
steam temperature (non-reheat)
superheater pinch temperature difference
high steam pressure
h-p evaporator pinch temperature difference
low steam pressure
1-p evaporator pinch temperature difference
economizer pinch temperature difference
stack temperature
Gasifier and Cleanup Systems
BCR gasifier exit temperature
air supply temperature (no preheat)
steam pressure to gasifier
temperature of feed-water to gasifier h-p boiler
booster compressor pressure ratio
Selexol operating temperature
fuel regenerator effectiveness
regenerator exit temperature (hot side)
fuel supply temperature (regenerator exit, cold side)
fuel sensible heat
fuel higher heating value
Integrated Station
auxiliary consumption
2200 F
16
11.8-5
82? F
100 F
1250 psia
50 F
33^ psia
50 F
50 F
300 F
1800 F
800 F
1250 psia (sat)
250 F
2
95 F
0.9
300 F
677 F
12 Btu/SCF
157 Btu/SCF
-85-
-------
FIG. 33
INTEGRATED COGAS/BCR GASIFICATION STATION PERFORMANCE
TURBINE INLET TEMPERATURE = 2200 F
TWO-PRESSURE STEAM SYSTEM
STACK TEMPERATURE = 300 F
FUEL REGENERATOR EFFECTIVENESS = 0.9
UTILIZE ALL SENSIBLE HEAT ABOVE 300 F
36
8
35
I
34
32
< 30
co
28
HIGH STEAM PRESSURE, PSIA
800
800
1800
I
I
10 20 30
COMPRESSOR PRESSURE RATIO
40
N03-54-1
-86-
-------
and higher pressures cannot be used without reheat or excessive moisture would
form in the low-pressure section of the turbine. Because of the relatively
small influence of steam pressure, the 1250 psi level was selected as a stan-
dard for the integrated systems. Thus, if it appears desirable to raise steam
pressure, it would be possible to re-evaluate any of the selected configura-
tions without changing gas turbine compressor pressure ratio or gasifier and
cleanup system operating pressure levels.
Fuel-Gas Temperature Effects - For the base-line system the effects of high-
temperature cleanup and fuel-gas regeneration were evaluated. The performance
improvements possible by using elevated fuel supply temperatures are shown in
Figure 3^. Two methods were investigated for achieving elevated fuel tempera-
ture: fuel regeneration (using hot, dirty fuel to reheat clean fuel from the
low-temperature system) and hot desulfurization (using a fictitious high-tem-
perature cleanup system which removes only I^S, CC>2, NHg, and particulate
matter). The performance of the reference station with fuel regeneration to
677 F is noted in the lower part of Figure 3^« The performance improvement
depends heavily on the temperature level below which sensible heat in the hot
fuel gas stream is not recovered (i.e., minimum temperature at exit of boiler
or hot side of fuel regenerator). Recovery of this low-temperature heat is
made more difficult by the presence of water vapor, sulfur compounds and ammo-
nia in the dirty gas which results in a mass flow rate that is higher than for
the clean gas on the cold side of the regenerator. As a result, the tempera-
ture drop on the hot side of the regenerator is significantly less than on the
cold side» One of the undesirable aspects of a low-temperature cleanup system
is the need to condense most of the water vapor in the dirty gas stream. This
will result in the need for costly heat exchange equipment to withstand the
weak acids which will be present. However, since no alternative is apparent
it appears desirable to utilize as much of the heat in the dirty gas stream
as is possible to ease the duty of the regenerator.
The foregoing analysis placed no restrictions on regenerator performance
or design characteristics. A highly effective regenerator will result in
excessive size and cost. Also, one that is capable of operating at tempera-
tures up to 1000 F will require more expensive materials. To evaluate the
desirability of regeneration to 1000 F, a cost study was performed for the
BCR-Selexol configuration. Figure 35 shows the effect of regeneration on
overall efficiency and on the heat exchanger area requirements as a function
of the regenerator outlet temperature. Because of the rapidly increasing
area, the high-temperature exchanger shows a capital cost increment due to
both materials and size. Based on the area requirements of Figure 35> a com-
parison of regenerator cost was made between a 750 F outlet temperature and a
1000 F outlet temperature. This is presented in Table 11. Equipment costs
-87-
-------
FIG. 34
PERFORMANCE IMPROVEMENT DUE TO ELEVATED FUEL SUPPLY TEMPERATURE
TURBINE INLET TEMPERATURE =2200 F
COMPRESSOR PRESSURE RATIO = 16
TWO-PRESSURE STEAM SYSTEM
HIGH STEAM PRESSURE = 1 250 PSIA
STACK TEMPERATURE = 300 F
BCR
GASIF.
BLR
I REGEN.
1
COOL -
I I
I REGEN.
36
34
_ 32
_i
<
O
I
I
11 30
u
01
O 36
UTILIZE ALL FUEL GAS SENSIBLE HEAT ABOVE 650 F
<
34
32
30
UTILIZE ALL FUEL GAS SENSIBLE HEAT ABOVE 300 F
~&&*
SELEXOLSYSTEM
TEMPERATURE
GASIFIER EXIT-
TEMPERATURE
500 1000 1500
CLEAN FUEL SUPPLY TEMPERATURE, F
2000
N03-54-2
-------
FIG. 35
EFFECT OF FUEL TEMPERATURE
BCR/SELOXOL SYSTEM
37
O
2
UJ
O
35
700 800 900
FUEL TEMPERATURE, °F
1000
R04-35-1
-89-
-------
Table 11
BCR-SELEXOL FUEL GAS REGENERATOR COSTS
Fuel Supply Temp - F 1000 750
2
Required HX Area - ft 88,500 U8,500
exchanger Cost_$ 2.U3 x 10° 0.85 x
Total D & E Cost-$ ^.38 x 10^ 1.53 x
-90-
-------
were estimated to be $17.50/ft2 for the low-temperature unit and $2?.50/ft2
for the 1000 F unit. These are increased by a 1.8 factor to arrive at total
erected cost. In the high-temperature case, 18-8 stainless steel and chrome
alloy steel are employed for the tube and shell materials. While low alloy
carbon steel can be used for the shell in the low-temperature case, stain-
less steel would still be required for the tube service due to the H2-H2S
environment at the operating temperature. The resultant cost differential is
still quite substantial due to the significantly larger area required for the
high-temperature regeneration. At the net plant output, the incremental cap-
ital cost per point improvement in efficiency is $5.?2/kw. This can be re-
lated to a fuel cost using Figure 36 where yearly fuel savings due to improved
efficiency have been related to an incremental capital investment. The high
temperature regeneration is justified for fuel costs in excess of 600/MMBtu.
Therefore, the BCR/Selexol system includes fuel regeneration to 1000 P.
Gasifier Heat Balance - The initial performance estimates for the BCR-type
gasifier assumed the use of steam at 970 F which is about 150 F higher than
would be supplied to the steam turbine. With the configuration of Figure 29
the alternative means of satisfying the gasifier heat balance were evaluated
and generally found to be undesirable because of the relatively large penal-
ties that resulted from the need to provide this small amount of additional
heat. As an example, a decrease in efficiency of between 0.6 and 0.9 points
results if the air preheater is used to provide the additional heat by raising
inlet air temperature. A summary of the methods considered is presented in
Table 12. It is interesting to note that the most costly (in terms of effi-
ciency lost) means of providing the additional heat was to change the gas-
ifier air-fuel ratio. The most desirable approach uses regeneration of the
transport gas to a temperature of 69^- F. This again serves to illustrate the
value of regeneration in the fuel gas stream. As a result, the configurations
shown in Figures 30 and 31 include provision for regeneration of the transport
gas.
Effect of Pressure Ratio on Power Cost - The effect of gas turbine pressure
ratio on the cost of power produced was evaluated for an integrated system
based upon an entrained-flow residual oil gasifier. The characteristics of
the system are presented in Table 13. The analysis of this system results in
trends similar to coal-based systems. The constraints imposed on the power
system were the same as those presented earlier with a two-pressure steam sys-
tem optimized to produce a 300 F stack temperature. Steam temperature was set
at 100 F below turbine exhaust with a maximum temperature of 1000 F. A low-
temperature cleanup system was assumed. As pressure ratio was varied from 8
to UO with a 2200 F turbine inlet temperature, efficiency varied from about
26.2 percent to 29.2 percent. An estimate of relative power costs is pre-
sented in Figure 37•
-91-
-------
FIG. 36
CAPITAL INVESTMENT EQUIVALENT
OF PLANT EFFICIENCY
17% YEARLY OWNING COST
0.7 LOAD FACTOR
FUEL COST
$1.50/MM BTU
FUEL COST
$ .50/MM BTU
30 35 40
INITIAL PLANT EFFICIENCY -%
-02-
R04-35-2
-------
Table 12
EFFECT OF STEAM EXTRACTION FOR GASIFIER ON STATION PERFORMANCE
Transport Efficiency
Location of
Steam Extraction
1.
2.
3.
It.
5.
6.
7.
8.
9.
Gasifier HPB
Gasifier HPB
Main Boiler
Superheater
Gasifier HPB
Main Boiler
Superheater
Main Boiler
Low-Pressure
Evaporator
Main Boiler
Superheater
Main Boiler
Superheater
Main Boiler
Superheater
Steam Conditions Gasifier Air
T,F p , psia
581* 1250
969 1250
826 1250
58U (Sat) 1250
969 1250
l<67 (Sat) 500
950 900
969 1250
950 900
Temp, F
800
800
859
1015
800
995
800
800
800
Exit Pressure
Temp, F Ratio
926 16.0
926 16.0
926 16.0
926 16.0
1069 9-6
926 16.0
1050 10.3
1069 16.0
1050 16.0
Gas Differential
Temp. F %_
150 Base
150 -0.32
150. -0.59
150 -0.87
150 -0.93
150 -1.10
150 -1.16
150 -I.!t5
150 -1.72
10. Gasifier HPB 581* (Sat) 1250
11. Gasifier HPB
12. Main BLR
969 1370
82l» 1250
800
800
800
926 16.0
92lt 16.0
150
926 16.0 150
69!*
-1.79
-0.55
-0.25
Comments
Gasifier heat "balance
violated
Only steam for gasifier
superheated in HPB, high-
risk-boiler
Fired air preheater
Fired air preheater
Special gas turbine design
Cannot maintain 300 F
stack, fired air preheater
Steam conditions per
original BCR data, special
gas turbine design
Fired boiler
Steam conditions per
original BCR data, fired
boiler
Base.Case 1 with gasifier
air/coal ratio altered to
satisfy heat balance.
Steam for gasifier super-
heated in special HPB
Additional gasifier heat
obtained by preheating
transport gas. -Revised
configuration also allows
higher fuel temperature
to engine.
-93-
-------
Table 13
CHARACTERISTICS OF OIL GASIFICATION SYSTEM
Gasifier;
Boiler
Cleanup
(Selexol)
Gas Analysis
Component
H2
CO
H20
C02
A
COS
Residual oil heating value
Air preheat temperature
Air-to-oil ratio
Steam-to-oil ratio
Exit gas temperature
Pressure drop
Steam pressure (saturated)
Exit gas temperature
Pressure drop
Operating temperature
Fuel gas consumption
Process steam consumption
(300 F saturated)
Pressure drop
Fuel gas heating value (HHV)
Gasifier Exit
(mole fraction)
.1287
.2066
.0009
• 5592
.0682
.0266
.0072
.0024
.0002
18,300 Btu/lb '
955 F
6.39 lt> air/lb oil
0.15 Ib steam/lb oil
2*150 F
50 psia
1250 psia
650 F
40 psia
100 F
.08 Ib steam/lb gas
50 psia
121 Btu/SCF
Cleanup System Exit
(mole fraction)
.1426
.2289
.0010
.6195
.0000
.0000
.0080
.0000
.0000
-------
FIG. 37
EFFECT OF PRESSURE RATIO ON POWER COST
TURBINE INLET TEMPERATURE = 2200 F
TWO-PRESSURE STEAM SYSTEM
HIGH STEAM PRESSURE = 1250 PSIA
STACK TEMPERATURE = 300 F
2.0
1.5
w 1.0
QC
UJ
LL
LL
o
o
oc
UJ
I 0.5
Q.
o •—
0
DIRTY FUEL COST, C/106BTU
10 20 30
COMPRESSOR PRESSURE RATIO
40
-95-
-------
The power cost presented is the estimated busbar cost minus the corre-
sponding estimate for a pressure ratio of 16. In going from low- to high-
pressure ratio, the oil consumption and gasifier output must increase, causing
the gasifier cost to increase by 0.7 exponential cost factor. The effect of
pressure on gasifier cost is neglected. In addition, the gas turbine and
steam system costs also increase resulting in minimum power cost for pressure
ratios near 16.
.Boost Compressor Location - Varying the location of the boost compressor
yielded interesting results. For the oil gasifier system described above, the
results of varying gasifier boost compressor location are presented in Figure
38. The efficiency range for compression after cleanup results from two pos-
sible assumptions. If the same absolute pressure drops are maintained through
the system the net efficiency would be 30.2 percent. If the same fractional
pressure drops are maintained the resultant efficiency would be 30-8 percent.
A number of factors tend to mitigate and actually reverse this effect in a
coal fed system. Because of the lower air to fuel ratio with coal, there is a
greater increase in mass flow rate in going from the air stream to the clean
fuel stream. For the BuMines/low-temperature system, the fuel gas compressor
uses about 50 percent more power than does the air boost compressor. While
heat of compression appears in the product gas stream, locating the compressor
in the low-temperature gas stream actually displaces one of the most desirable
fuel gas regeneration locations resulting in a need to reject an additional
amount of low-temperature heat. Thus, the boost compressor has been located
in the air stream in all of the final coal systems. However, the possibility
of a better location should be considered in oil-fired systems.
Parametric Analysis of Nitrogen Oxide Emissions From Gas Turbine Power Systems
Burning Low-Btu Fuel Gas
Oxides of nitrogen, commonly lumped together as NOX are receiving in-
creasing attention as air pollutants. The various oxides are easily inter-
converted in the atmosphere, their ratio depending on the action of sunlight,
oxygen, and other oxidizing or reducing agents present. A major contributor
of these pollutants are the nitrogen oxides formed in the hot reaction zones
of all air-breathing combustion engines. They are formed primarily as NO
(nitric oxide), although small quantities of N02 (nitrogen dioxide) and I^O
(nitrous oxide) may also be formed.
Two mechanisms are known to contribute to the formation of nitric oxide
in combustion systems. The most important mechanism for gas turbines,'and
other systems which burn relatively clean fuels, is referred to as the ther-
mal or hot air mechanism. In this mechanism, nitrogen and oxygen from the
-96-
-------
INTEGRATED COGAS/OIL GASIFICATION STATION PERFORMANCE
TURB. INLET TEMP. = 2200 F
TWO-PRESSURE STEAM SYSTEM
HIGH STEAM PRESSURE = 1250 PSIA
STACK TEMPERATURE = 300 F
VO
-o
40
I
I
a?
l
o
z
HI
o 30
HI
O
1
20
.BOOSTER COMPRESSOR BEFORE GASIFIER
BOOSTER COMPRESSOR AFTER PURIFICATION
10
20
COMPRESSOR PRESSURE RATIO
30
40
P
oo
00
-------
atmosphere react in the hot combustion zone to form nitric oxide. The second
mechanism is important when relatively dirty fuels such as coal and residual
fuel oil are burned. Most dirty fuels contain small but significant quan-
tities of organic nitrogen compounds. Because nitrogen-carbon and nitrogen-
hydrogen bond energies are so much lower than that for molecular nitrogen,
much of the fuel becomes oxidized during combustion. Experimental studies(35)
of the formation of nitric oxide from fuel nitrogen indicate that the forma-
tion rates are very rapid, occurring on a time scale comparable to that of the
hydrocarbon combustion reactions. This mechanism is strictly fuel dependent
and proceeds at lower temperatures than needed for the thermal mechanism.
Fuel nitrogen could be a problem in systems using gasified fuels. During
gasification of dirty fuels, some fuel nitrogen would carry over into the raw
fuel gas as combustible nitrogen compounds (primarily ammonia, with smaller
concentrations of hydrogen cyanide, pyridine, pyridine bases, and acidic
nitrogenous compounds). If retained in the fuel gas, these compounds could
result in excessive emissions of nitrogen oxides. It appears that low-tem-
perature cleanup systems would adequately remove undesirable nitrogen com-
pounds, but the same cannot be said for high-temperature cleanup systems.
Status of NOX Pollution Modeling - The chemical kinetics of NO formation via
the thermal mechanism are fairly well understood^°~3yj t jn ^he gas turbine,
the local temperature, residence time, and species concentrations which govern
NO production are controlled by engine operating conditions, the combustor
internal flow field, fuel nozzle characteristics, and the air addition sched-
ule to the burner can. Lack of an adequate analytical description of the
combustor flow field and the fuel/air mixing characteristics has prevented
accurate estimation of the temperature-time concentration history which is
essential for reliable estimation of NO formation. At the present time, a
number of engineering and research establishments, including several groups
within United Technologies Corporation, are attempting to develop comprehen-
sive gas turbine combustor models. The results of this modeling'work have
been very encouraging and are leading to a better understanding of NO emis-
sions .
Despite the rather primative state of combustor/pollution models, at
least one model exists which can be used to predict trends in emission char-
acteristics under varying operating conditions for typical types of gas tur-
bine combustors. This model (described in Appendix B) is a semiempirical
simulation developed by United Technologies for use with conventional liquid
fuels and subsequently modified for use with low-Btu gaseous fuels. The model
uses engineering approximations of the flow field in the combustor combined with
physical and chemical combustion kinetic models.
-------
Nitric Oxide Emissions From Gas Turbine Combustors - The combustor model des-
cribed in Appendix B was used to simulate the combustion of low-Btu fuel gas
in a can-type combustor similar to that used in a high-performance, utility
turbine (Pratt & Whitney FT^). A parametric study was made of the effect of
varying operating conditions on thermal NOX emissions for four fuel gas com-
positions produced by four coal gasifier-cleanup system combinations. A com-
puter simulation was also made using test rig conditions burning gasified oil
(see Table lU). These rig tests were run at atmospheric pressure with the
fuel gas at ambient temperatures and are the only known rig tests of gas tur-
bine combustors burning fuel gas from an air-blown gasifier. These rig tests
were made using the Pratt & Whitney FT1*- industrial gas turbine combustor can
and gasified oil produced by Texaco Oil Company's pilot gasifier^40' ^ '. The
tests went very well with no problems in burning the fuel and with low NOX
production — so low that the NOX measurements were only slightly higher than
the error margin in the chemical measurements. The model simulation results
showed good agreement with the rig tests. More, and greatly expanded, rig
tests have been run in 1975> and it is anticipated that a complete correlation
of the model simulation and rig test conditions can be undertaken.
Fuel gas characteristics for the two-stage gasifier being developed by
Bituminous Coal Research, Inc.(°> ?/ coupled with the Consolidation Coal
Company half-calcined dolomite cleanup system are summarized in Table lU.
When the simulated FTk combustor (Figure 39a) is modelled under these condi-
tions, the thermal WOX concentration in the combustor is predicted as shown in
Figure 39b in which half of a cross-section of-the annular combustor is shown.
An FTU gas turbine, which develops approximately 30 Mw of electricity at peak
rating, contains eight of these combustor cans. Each can contains six
swirlers, shown at left in Figure 39&> symmetrically arranged about the annu-
lus and a seventh swirler at the end of the center tube. This extends more
than halfway down the combustor can and is used to inject dilution air into
the main combustion area of the combustor can downstream of the recirculation
zone. The recirculation zone, held by the vortical flow behind the swirlers,
acts as a flame holder and ignition source for the main body of fuel in the
combustor can.
While the NOX concentration in the recirculation zone is high, because
the temperatures there are high and the fuel-air ratio is near stoichiometric,
the recirculation zone is_ not the major contributor to NOX emissions from gas
turbine combustors. As may be seen by examining Figure 39, the NOX concentra-
tions are higher downstream of the recirculation zone. In Figure kO the
cumulative NOX production versus axial distance along the combustor is shown.
It may be seen that virtually all of the thermal NOX emissions are produced
downstream of the recirculation zone. The small contribution of the recircu-
lation zone to NOX emissions is due to the relatively small flow through the
-99-
-------
Table lit
LOW-BTU GAS CHARACTERISTICS AT DESIGN POINTS FOR FOUR GASIFIER-CLEANUP SYSTEMS
Gasifier
Gasified. Coal
BOM
BOM
BCR
BCR
Cleanup
Selexol
Iron Oxide
Selexol
Consol
p
( atm . )
15.8
15.8
23.8
23.8
T •
•'-air
(F)
75U.6
754.6
903.1
903.1
Tfuel
(F)
265
1070
1000
1700
H20
.0001
.0391
.0001
.0994
CO
.2387
.1534
.2123
.173T
H2
.1598
.1819
.1437
.13147
C02
.0502
.1113
.0649
.0929
N2
.5197
.4770
.5318
.4588
Ar CH1+ KH3
COS
H2S
Ash
.0311 .0003 .0001
.0277 .0063 .0001 .0032
.0421 .0001
.0356 .0043 .0001 .0005
i Gasified. OiJ
H
o
o
1 Texaco
Rectisol
1.3
732
95
.2218 .1579 -0001 .6108 .0079 .0015
-------
FIG. 39
NOX IN GAS TURBINE COMBUSTORS
AIR
FUEL
FUEL NOZZLE
SWIRLER
V
CENTER TUBE
OUTER LINER
FT4 COMBUSTOR CONFIGURATION
(PPM BY WEIGHT)
PRATT & WHITNEY FT4 COMBUSTOR
BCR CONSOL PROCESS FUEL GAS
P = 23.8 ATM
TFUEL = 1700°F
NOX CONCENTRATION ISOLINES
N03-80-4
-101-
-------
NOX VS AXIAL DISTANCE
FIG. 40
END OF RECIRCULATION ZONE
I
6.0 10.0
AXIAL DISTANCE (IN.)
14.0
18.0
N03-80-2
-102-
-------
recirculation zone. Since the function of the recirculation zone is to pro-
vide a heat source to stabilize the flame, only enough fuel and air are added
to maintain its temperature. The main fuel combustion takes place downstream.
It may be seen that the NOX production in gas turbine combustor cans is di-
rectly related to combustion in the main burning zone downstream of'the re-
circulation zone and, in particular, to the local temperatures since NO and
NC>2 production rates are extremely dependent on temperatures.
The production of thermal NOX for a series of fuel gas temperatures are
given in Figure Ul for four gasifier-cleanup systems (see Table lU). The com-
bustion air temperature was set to correspond to the appropriate compressor
discharge conditions as determined from the cycle calculations. All runs were
made with- air flow schedules and geometry for the Pratt & Whitney FT^ scheme
2^-2DB combustor. In order to obtain a set of comparable runs, fuel flows
were chosen with the objective of obtaining turbine inlet temperatures of
2200 F for the Bureau of Mines gasifier cases and 2600 F for the BCR gasifier
cases. The units used in Figure kl are Ib NOX/10 Btu where Btu refers to the
higher heating value of the fuel gas. To convert to Ib NOX/10° Btu coal, it
is necessary to multiply by the effective cold gas efficiencies which are 83-2
percent, 78.7 percent, 76.2 percent, and 77-1 percent for the BuMines-Selexol,
BuMines-Iron Oxide, BCR-Selexol, and BCR-Consol processes respectively.
The parametric thermal NOX estimates presented in Figure Ul reveal impor-
tant trends. Increases both in fuel temperature and in cycle pressure cause
dramatic increases in NOX emissions. This comes as no surprise since both
trends had been predicted some time ago'^' ^). it has previously been pre-
dicted that the temperature rise associated with one Btu of fuel gas sensible
heat was twice that associated with one Btu of chemical heating value. The
reason for this is simply that increased chemical heating value is caused by
a higher concentration of combustibles in the fuel and, thus, requires more
combustion air. Since the additional combustion air consumes most of the
incremental heating value while being heated up to the initial combustion
temperature, little heat remains for further temperature rise. Sensible heat,
however, simply adds to the temperature rise. Since the kinetics of NOX pro-
duction is strongly temperature dependent, the trends in Figure hi seem appro-
priate.
Fuel Nitrogen - As has been previously pointed out, fuel nitrogen compounds
would add to the thermal NOX estimates presented in Figure kl. For the four
combinations of gasification/cleanup systems shown in Table lw, the only fuel
nitrogen compound of consequence is NHo with concentrations ranging from 0.03
to 0.65 percent (Vol.). The mechanism whereby NHo is converted to NO is not
well understood at this time, although there appears to be ample
-103-
-------
FIG. 41
NOX PRODUCTION VS. TFUEL AT VARIOUS PRESSURES FOR COAL-GAS FROM FOUR
GASIFIER-CLEANUP SYSTEMS
NO
X
LB
\106Btu
2 -
1000 2000
TFUEL (F)
BUR. MINES-SELEXOL
NOX
LB
106Btu
P=28 ATM
P=16ATM
1000
TFUEL
BUR.MINES-IRON OXIDE
2000
NOX
\106Btu
1000
TFUEL
BCR-SELEXOL
2000
NOX
/_L5_\
\ 106Btu/
P=40 ATM
P=4 ATM
1000
TFUEL(F)
BCR-CONSOL
2000
-------
which suggest that if the concentration of NHo in the fuel gas is low (say
< 0.5 percent, vol.), all of it will be converted to NO ; whereas if the con-
centration is high, only a fraction will be converted to NOX. While it is
difficult to assign a specific fraction of conversion, the mechanism is such
that the addition of NHo, for example, to a low concentration assumed to have
essentially 100 percent conversion to NOX, will not result in a net decrease
in NOX production although less than 100 percent of the total fuel-bound
nitrogen compounds are now converted to NOX.
One of the more plausible attempts to correlate fuel nitrogen conversion
data is presented by Fenimore^^'. A mechanism is postulated whereby all the
fuel nitrogen reacts through an intermediate species I according to the reac-
tion:
I + R - NO + (33)
I + NO -» N2 + (3*0
where R is a nitrogen-free species which causes the eventual oxidation of I to
NO. The second reaction denotes the decomposition of NO, which could become
important when the NO concentration becomes high. The postulate further sug-
gests that R might correspond to the OH radical and that I might correspond to
NH2.
In order to utilize the above model, it would be necessary to combine it
with a thermal and combustor NOV model so that local concentrations of the
.A.
reactive species could be used. No meaningful attempt has yet been made to do
this. Although using average species concentration can be misleading, this
was done using the correlations in reference kk to obtain an order of magni-
tude estimate of the NOX conversion yield for fuel nitrogen. The effective
yield was nearly 100 percent from the two cases with low NHo concentration
(BuMines/Selexol and BCR/Selexol). For the BuMines/hot iron oxide and BCR/
CONSOL cases, the effective yield appeared to be between 50 and 100 percent.
Lacking more definitive information, 90 percent effective yield will be as-
sumed in estimating the NOX produced by nitrogen-based fuel constituents.
Discussion of Results of NOX Modeling - Despite the above uncertainties in
thermal and fuel NOX predictions, some generalized conclusions can be made.
It appears that NOX models predict correct trends, with the values predicted
within about a factor of 2 or 3 of actual production. High-temperature sul-
fur removal systems which do not remove NHo would result in NOX emissions in
excess of the EPA standards for coal-fired powerplants of 0.7 lb/10° Btu.
-105-
-------
Since at the time of this writing no recognized standard exists, the value
for a coal-fired station is used. Whether low-temperature cleanup systems
could meet the NCL. standard would depend on the use of fuel gas regenerators
J\.
to raise the fuel supply temperature.
Fuel gas temperatures in excess of approximately 1000 F, and pressures
over approximately 16 atm, could lead to NOX emissions in excess of the 0.7
lb/10° Btu standard. For temperatures and pressures below these values it
should be possible to design gas turbine combustors which conform to EPA
standards.
It must be noted that the foregoing analysis was based upon a simulated
FT^-type burner can. This combustor was designed for use with liquid fuels
and high-Btu gas. Tests subsequent to the original Montebello tests (pre-
vious rig tests) indicate that various modifications to these combustors
could result in lower NOX formation. Thus, future efforts in modelling NOX
formation could result in estimates more closely resembling actual conditions,
-106-
-------
SECTION 3
COMPARISON AND SELECTION OF CLEANUP SYSTEMS
Section 3 contains comparisons of the various low- and high-temperature
cleanup systems reviewed in Section 1 and describes the criteria and methods
used to select these cleanup systems for integration with the gasifier and
power system.
COMPARISON OF CLEANUP SYSTEMS
The comparison of the low- and high-temperature cleanup processes iden-
tified in Section 1 includes as the primary criteria the efficiency of
pollutant removal, effect on power system performance, cost considerations,
and estimated time of availability (first- or second-generation) for com-
mercial application.
Low-Temperature Systems.
The majority of the approximately Uo low-temperature desulfurization pro-
cesses identified in Section 1 are commercially operative and, therefore, could
be applicable to both first- and second-generation integrated power systems.
In selecting those likely to be most applicable in treating coal derived fuel
gas, the following factors were taken into consideration;
a. Sulfur removal capabilities, not only with respect to HpS, but also
other sulfur compounds such as COS and CSp.
b. Selective absorption of sulfur compounds over carbon dioxide. Since
COg need not be removed from fuel gas intended for use in advanced
power cycles, absorption of CO represents an increased operating load
on the system.
c. Type of absorbent insofar as the treated fuel gas may be contaminated
by entrained or volatilized solvent which could be detrimental to
downstream system components.
-107
-------
d. The system's tolerance to other contaminants present in the raw fuel
gas such as ammonia, cyanides, phenols, and tars.
e. Overall energy requirements and operating cost.
A review of the literature of sour gas stripping was performed and those
processes which appeared to be commercially important were identified. However,
since the scope of this study was limited to published information, the availability
of performance data allowed only ten systems to be selected for further analysis.
Of the chemical solvent processes, the Alkazid, Benfield, and Phosphate
processes were selected on the basis of their capability for partial selective
H S absorption. The Sulfinol, Selexol, Fluor Solvent, Purisol, and Rectisol
systems were included as they are commercial physical solvent processes. The
Stretford and Giammarco Vetrocoke systems represented commercially operating
direct conversion processes. Evaluation of these ten processes was based upon
treatment of UOO MMSCFD of dry feed gas having the following composition:
Ng CO H2 C02 CH H2S COS NH3
Volume % ^5.0 20.0 19.0 11.0 U.O 0.5 0.1 0.5
This composition is typical of that expected from air gasification of coal
after tar, naphtha, and particulate removal. It was assumed that the gas would
be available at 300 F and UOO psig for low-temperature cleanup.
Based solely on information available in the literature, the Fluor Solvent
and Purisol processes were eliminated from further consideration since they
are best suited for cases where the acid gas partial pressure is high, e.g.,
above 75 psia.^ ' The Giammarco Vetrocoke and Stretford processes were
similarly judged unsuitable since their application is economically limited to
sulfur recovery of about 20 tons/day. (^> 13) The remaining processes were
ranked in terms of the following characteristics and corresponding weighting
factors:
Sulfur removal 1.5
Energy Consumption 1.0
Investment cost 0.5
Absorbent type 0.5
Operating temperature 0.5
According to this arbitrary ranking techniaue, the lower the rank the more
attractive the system. While it would have been more desirable to have li-
censor operating and cost data for a given degree of sulfur removal as a means
-108-
-------
of comparison, the ranking procedure employed gives a Qualitative indication
of the relative attractiveness for these cleanup systems.
The degree of sulfur removal, both with respect to H2S and COS, was given
primary importance for pollution considerations. Given secondary considera-
tion were the steam and power requirements since these would directly relate
to the overall efficiency of the integrated power system. Absorbent or sol-
vent type was a factor insofar as it could effect downstream units, e.g.,
carry over of alkali salts in the treated fuel gas is detrimental to high-
temperature gas turbine blade life and, therefore, protective systems for
assuring complete removal would have to be incorporated.
Based on the ranking of the six processes shown in Table 15, it appeared
that the Benfield chemical solvent system and the Selexol and Rectisol physi-
cal solvent systems were comparable. These were chosen for evaluation in
integrated system performance. For this purpose, preliminary evaluation data
were obtained from the process licensors.
High-Temperature Systems
In evaluating the relative merits of the six high-temperature processes
identified in Phase I, the following factors were considered:
a. Operating temperature
b. Capability for removing sulfur compounds, COS as well as
c. The form in which the sulfur is regenerated, H-^S, S0?, or elemental
sulfur. Elemental sulfur is the preferred form since it can be stock-
piled without presenting significant pollution problems.
d. Ability to regenerate the absorbent without substantial loss in
activity.
e. Overall energy requirements and operating costs.
The high temperature processes were compared relative to desulfurizing
the following off-gas from a coal gasification unit operating at 1800 degrees
F and ^50 psia;
W2 H2 CO C02 CH^ H20 H2S ' COS NH3
Volume % ^5-5 13-1 18.1 8.3 3.6 10.k 0.5 0.1 O.U
-109-
-------
Table 15
COMPARISON OF SELECTED LOW-TEMPERATURE CLEANUP PROCESSES
H
H
O
I
PROCESS
Benfield
Selexol
Rectisol
Sulfinol
ABSORBENT
TYPE
Catalyzed Y^CO
Solution
Polyethylene
Glycol Ether
Methanol
Sulfolane +
OPERATING
TEMPERATURE
°F
230
ko
-ko
110
SJMFUR
CONCENTRATION
PPM
100
100
10
100-200
ENERGY
REQUIRED
MMBTU/HR
230
90
53
lUO
INVESTMENT
COST
MM $
3
8
7
3-5
BANK
7
6
6
8
Di-Isopropanolamine
TPP Aqueous Tri-
Potassium Phosphate
Alkazid Potassium Dimethyl
Amino Acetate
100
80
300
300
270
150
3-5
3-5
11
10
-------
The thermodynamic equilibrium absorption of I^S and COS were calculated
using the equilibrium constants developed during Phase I and assuming that
CO shift equilibrium was attained at the absorption conditions. Table 16
summarizes the residual sulfur levels theoretically obtainable.
For operation above 1600 F, which would be compatible with a second-
generation, entrained-flow gasifier, both the Consol and Air Products dolomite-
base processes appear attractive. The latter is capable of reducing the sulfur
level to around 100 ppm compared to UOO ppm for the Consol process but requires
additional fuel for calcination. The Battelle molten-salt process also oper-
ates in this temperature range but its ability to remove sulfur compounds to
acceptable levels is questionable, particularly in the range of 30-^0 atm
total pressure. At low salt loading, below 25 percent, residual H S can be
reduced to 150-1000 ppm depending on the equilibrium constant, which remains
to be defined experimentally. Moreover, there is no available data on COS
removal capability for this process and there is the potentially serious prob-
lem of alkali metal volatility at the high operating temperature.
The Bureau of Mines iron oxide process appears most suitable for sulfur
removal at temperatures below 1500 F, preferably around 1000 F. This is the
operating range for first-generation fixed-bed gasifiers for which this process
was originally developed. Off-gas from a high-temperature, second-generation
gasifier would require cooling to the operating temperature of iron oxide and
would result in a lower thermal efficiency than for integrated systems using
the dolomite-based processes. An inherent disadvantage of the iron oxide
process is that, upon regeneration, the sulfur is released as sulfur dioxide.
This can readily be converted to sulfuric acid for sulfur recovery purposes.
However, conversion to the preferred elemental sulfur form would necessitate
additional processing steps to reduce part of the SOP to H^S.
Since there are no commercial high-temperature desulfurization systems and
hence no available operating data, a preliminary comparison of their energy re-
quirements and operating costs was not possible. The economics of the high-
temperature systems could be significantly influenced by replacement cost for
the acceptor, resulting from loss of activity over repeated regeneration
cycles. Acceptor life for the dolomite system has yet to be satisfactorily
demonstrated. In this respect, the iron-oxide process has an advantage.
Based on the above qualitative comparison of the high-temperature pro-
cesses identified in Phase I, the Bureau of Mines process appears well suited
for first-generation application. Second-generation systems could equally
employ the Consol or Air Products dolomite processes.
-Ill-
-------
Table 16
COMPARISON OF HIGH-TEMPERATURE CLEANUP PROCESSES
PROCESS
ABSORPTION CONDITIONS
RESIDUAL SULFUR, PPM
REGENERATION CONDITIONS
SULFUR
,
I-1
ro
Con sol
Air Products
Bureau of
Mines
Battelle
Northwest
AGENT
Ca C03-Mg 0
Ca 0-Mg 0
Fe203 + Fly
Ash
Molten CaC03
F
1750
1650
1000
1500
1600
H2_S
3^0
100
130
1300
150-1000
COS
50
10
20
200
9
AGENT
Steam + C02
Steam + C02
Air + Fuel
Air
Steam + C02
F
1200
1200
1900
1500
1100
RECOVERY
H2S
HgS
so2
HgS
ICT-Meissner
Proprietary
800
(100)*
Sulfur
•^Estimated from Literature
(27)
-------
Preliminary Performance of Integrated Cleanup Systems
Preliminary performance estimates of the overall, integrated power systems
were used as a basis of comparison of selected high- and low-temperature clean-
up processes. Mass and energy balances together with utility requirements were
developed for the individual desulfurization processes for integrated front end
systems producing clean fuel gas from coal. For the purpose of this comparison,
the entrained-flow BCR gasifier was used under the following operating
conditions:
Coal Type
Illinois No. 6
Feed, Ib/hr 2000
Gasifier Operation
Temperature, F 1800
Pressure, psia 500
Air, Ib/lb, coal 3.1*22
Steam, Ib/lb coal 0.56?
Gasifier Production
Net Gas, SCF/lb coal 7^-53
Slag, Ib/lb coal 0.087
Raw Gas Analysis, Volume %
N2 CO CO H2 CH^ E2S COS NH3 t^O
W.70 16.7^ 8.8U 11.98 3.1^ O.U6 0.10 0.38 10.66
The gasifier heat and material balance for these conditions is given in
Table 17- There is a net heat rejection of 0.38U MMBtu/hr, available for high-
pressure steam generation, resulting from cooling the recycle char stream. Raw
producer gas is then processed through the gas purification section in which
sulfur is recovered in its elemental form.
Integrated Low-Temperature Cleanup Systems - Based on the preliminary evalua-
tion of low-temperature desulfurization processes, the relative performances of
the Benfield, Selexol, and Rectisol processes were analyzed. Preliminary
operating data obtained from process licensors (^o-4o) forme(j the basis for
estimating the cleanup system performance. In each case, the acid gas removal
-113-
-------
Table 1?
MATERIAL BALANCE FOR BCR GASIFIER
(see Figure
STREAM
M.W.
°2
No -•
CO
C02
Hp
CHk
H2S
COS
NHo
H;)b
ASH
32.00
28.02
28.01
kk.Ol
2.016
16. ok
3^.08
60.07
17.03
18.02
1
LB/HR
2
MOL/HR
16.82
6.65
2.6k
1.88
3
MOL/HR
1+9.60
186.60
u
LB/HR
5
LB/HR
•0
6
MOL/HR
203.^5
71-38
37.69
51.08
13-39
1.96
o.Us
1.62
17^.0
TOTAL
F
PS IA
M.W.
2000.0
100
3^.20
300
550
2^.53
236.20
800
550
28.98
113^.0
578
1300
17^.0
2800
1750
U50
2^.96
-------
unit was designed to treat the raw gas to a residual sulfur content of 100 ppm.
The acid gas from the regenerator was designed for a high (> 10 percent) H2S
concentration enabling the recovery of elemental sulfur via a conventional
Glaus plant.
A block flow diagram for the integrated cleanup systems is shown in
Figure h2. Raw producer gas from the BCR gasification system at 1750 F is
first cooled in the heat recovery section to about 300 F. The heat extracted
is available for regenerative heating of the clean fuel gas, boiler feed water
preheating, and steam generation. This section was optimized to give the
highest overall COGAS plant efficiency for each case. The gas is further
cooled below the dew point to about 120 F via direct water scrubbing which also
removes at least 85 percent of the ammonia present In the raw gas. Sour water
from this scrubbing operation is first steam stripped and the stripped gases
are sent to an ammonia recovery section. The cooled, ammonia free, producer
gas is desulfurized in the low-temperature acid gas removal section producing
a clean fuel gas and a regenerated acid gas stream containing 13-22 percent
hydrogen sulfide. Sulfur is recovered from the latter stream together with the
ammonia recovery off-gas stream in a vapor phase Glaus unit, complete with a
tailgas treating section for recovering a minimum of 99 percent sulfur.
Material balances for the Selexol, Benfield, and Rectisol acid gas removal
processes are given in Tables 18-20 respectively. Performance of these systems
in producing a clean, low Btu fuel gas are summarized in Table 21 together with
the utilities requirements.
As expected for a chemical solvent system, the Benfield process requires
2.5 to 3 times more low-pressure steam for solvent regeneration than do the two
physical solvent systems. This is partially offset by the higher power con-
sumption required by the Selexol and Rectisol processes for mechanical refrig-
eration to obtain their subambient operating temperatures.
Using these data, a preliminary analysis of the integrated power 'stations
gave the following relative performance characteristics:
-115-
-------
INTEGRATED LOW-TEMPERATURE CLEANUP SYSTEM
I
H
H
I
COAL
<•>
T
STM
©-
0
' BCR
I GASIF1ER I
HEAT
RECOVERY
COAL TRANSPORT GAS
ACID GAS
REMOVAL
»- AMMONIA
WASTE
-------
Table 18
COMPONENT
H2S
COS
SOo
NO
Sulfur
M.W.
34.08
60.07
6k. 06
17.03
30.01
18.02
32.06
TOTAL, MOL/HR
PSIA
M.W.
MATERIAL BALANCE FOR SELEXOL CLEANUP SYSTEM
(.see Figure 42)
STREAM STREAM STREAM STREAM STREAM
7 8 9 10 11
°2
N2
CO
co2
H2
CH.
32.00
28.02
28.01
Mi. 01
2.016
16. 04
203.45
71.38
37.12
51.08
13.39
1.71
0.43
0.24
10.11
.388.91
200
440
25.58
9.28
9.28
120
0.40
.371.89
100
430
25.47
0.43
0.04
7.50 34.20
120 100
30 430
40.76 25.47
STREAM
12
203.45
71.38
32.22
51.05*
13.35
-
o.o4
4.90
0.03
0.04
1.71
0.39
18.72
6.56
2.96
4.69
1.23
182 . 07
63.89
28.84
45.69
11.95
o.o4
0.36
332.84
STREAM STREAM
13 14
0.57
0.25
1.38
0.31
2.51
100 150
430 30
25.47 24.98
1.37
0.07
1.44
110
-------
Table 18 - Continued
MATERIAL BALANCE FOR SELEXOL CLEANUP SYSTEM
CO
OMPONEI
°2
CO
co2
H2
City
H2S
COS
so2
NH3
NO
HgO
W STREAM
M.W. 15
32.00
28.02
28.01
44.01
2.016
16.04
34.08
60,07
64.06
17.03 o.oi
30.01
18.02 0.24
STREAM
16
0.57
0.25
• STREAM
17
5.47
0.03
0.04
1.96
0.39
0.24
0.43
Sulfur 32.06
TOTAL,
°F
PSIA
M.W.
MOL/HR 0.25
125
0.82
125
28
40.98
8.56
120
28
40.12
STREAM STREAM
18 19
2.66
0.93
0.42
0.67
0.17
0.01
2.34
2.34 4.86
300 100
25.47
STREAM
20
0.68
.15.42
7-43
0.24
3-88
27.65
130
31.03
STREAM STREAM
21 22
3.40
12.79
35.03
35-03 16.19
125 100
28.86
-------
Table 19
MATERIAL BALANCE FOR BENFIELD CLEANUP SYSTEM
(see Figure 1*2)
I
H
M
VO
I
COMPONENT
°2
N2
CO
co2
H2
CEh
H2S
COS
so2
NH3
NO
H20
M.W.
32.00
28.02
28.01
• 1*1*. 01
2.016
i6.dk
3k. 08
60.07
61*. 06
17.03
30.01
18.02
STREAM
. 7
203.1*5
71-38
37-12
51.08
13-39
1.71
0.1*3
0.21*
10.11
STREAM STREAM
. 8 9
203.1*5
71-38
27.29
51.08
13-39
0.02
0.01
0.21*
(12.85) 22.00
STREAM
10
9-83
1.69
0.1*2
0.96
STREAM
11
17.89
6.28
2.1*0
U.i*9
1.18
0.02
1.9l*
STREAM
12
182.89
61*. 17
21*. 53
1*5.92
12.01*
0.02
0.01
0.22
19.78
STREAM STREAM
13 Ik
0.57
0.25
1.38 1.37
0 . 31 0 . 07
Sulfur
32.06
TOTAL, MOL/HR
PSIA
M.W.
388.91
200
kkQ
25.58
(12,85)
. 200
388.86
250
1*30
2i*. 81
12.90
130
30
Ul.30
3^.20
250
1*30
2U. 8l
3^9-58
250
1*30
21*. 81
2.51
150
30
2U.98
1.1*1*
110
-------
Table 19 - Continued
MATERIAL BALANCE FOR BENFIELD CLEANUP SYSTEM
COMPONENT
°2
N2
CO
co
CH,
STREAM
M.W. 15
32.00
28.02
28.01
44.01
2.016
16 . 04
STREAM
16
0.57
STREAM
17
10.40
STREAM STREAM
18 19
2.67
0.9^
0.36
0.67
0.18
STREAM
20
0.58
13.62
12.28
STREAM STREA1
21 22
2.92
10.98
ro
o
H2S
COS
so2
NHo
NO
H 0
34.08
60.07
64.06
17-03
30.01
18.02
0.01
0.24
0.25
1.94
0.42
0.96
0.29
0.01
4.20
22.18
Sulfur
32 . 06
2.34
TOTAL, MOL/HR
PS I A
M.W.
0.25 '
125
0.82
110
28
40.98
13-72
130
28
41.28
2.34
300
5.11
250
24.81
30.69
130
33. lU
22.18
125
13-90
80
28.86
-------
Table 20
ro
Sulfur
MATERIAL BALANCE FOR RECTISOL CLEANUP SYSTEM
(see Figure H2)
:OMPONENT
°2
N2
CO
co2
H2
CEh
H2S
COS
so2
NH
NO
M.W.
32.00
28.02
28.01
Ml. 01
2.016
16. OU
3^.08
60.07
6k. 06
17.03
30.01
18.02
STREAM STREAM
7 8
203.^5
71.38
.37-12
51.08
13-39
1.71
0.^3
0.2k
10.11 10.11
STREAM STREAM
9 10
203. ^5
71.38
29. oil 8.08
51.08
13.39
1.71
O.U3
0.2k
STREAM
11
18.87
6.63
2.70
U.75
1.2k
0.02
STREAM STREAM
12 13
181.87
63.81
25.96 0.57
1+5.65
11-97
0.25
0.21 1.38
0.31
STREAM
1U
1.37
0.07
32.06
TOTAL, MOL/HR
OF
PSIA
M.W.
388.91
200
hkO
25.58
10.11
90
368.58 10.22
3^.21
90
90 120
^30 . 30
25.31 ^3.03 25.31
329.117 2.51
90 150
^30 30
25.31 2^.98
110
-------
Table 20 - Continued
MATERIAL BALANCE FOR RECTISOL CLEANUP SYSTEM
COMPONENT
ro
ro
°2
CO
CO.
V
COS
SOo
NO
Sulfur
TOTAL, MOL/HR
"F
PSIA
M.W.
STREAM
M.W. 15
32 . oo
28 . 02
28 . 01
kk.OI
2 . 016
16. OU
3^.08
60.07
6U. 06
17.03 o.oi
30.01
18.02 0.21*
32 . 06
0.25
125
STREAM STREAM STREAM STREAM
16 17 18 19
2.71
0.95
0.57 8.65 0.39
0.68
0.18
0.25 1.96
2.38
0.82 11. Ok 2.38 U.91
110 120 300 90
28 28 1*30
1*0.98 U2.88 25.31
STREAM
20
0.59
13.78
10.79
0.01
2.63
27.60
130
33.29
STREAM STREAM
21 22
- 3-^0
12.79
35-03
35-03 16.19
125 30
28.86
-------
Table 21
SUMMARY OF LOW- TEMPERATURE HOTEGRATED SYSTEMS
Process
Feed Streams
(1)
BCR Gas
' Flow, mph
- T F
P psia
Product Stream
Sulfur
Flow, Ib/hr
T F
P psia
Transport Gas
Flow, mols/hr
T F •
P psia
Product Gas
Flow, mols/hr
T F
P psia
C0
H2
Selexol
U26.1+5
7^-9
300
Benfield
U26.1+5
1750
1+50
300
H20
H2S
M. W.
HHV Btu/scf
Utilities
Cooling Duty, MMBtu/hr
Steam @ 1300 psia, Ib/hr
@ 65 psia, Ib/hr
Electric Power, kw
Boiler Feed Water, Ib/hr
Steam Condensate, Ib/hr
Feed Gas Cooling, MMBtu/hr'
(l) Based on 2000 Ib/hr Coal Feed to Gasifier
(2) Available for STM Generation and/or Gas Reheat
Rectisol
1+26.U5
1750
t+50
76.2
300
3^.20
300
550
332.81+
100
1+30
5^.70
19.20
8.66
13-73
3-59
0.11
100 ppm
25.^7
1^2.6
3^-20
300
550
3^9 A?
250
U30
52.32
18.36
7.02
13-1^
3-^
5.66
0.06
100 ppm
2lf.8l
136.5
3k . 21
300
550 •
329 A6
. 90
U30
55.20
19-37
7.88
13.86
3-63
0.06
10 ppm
25-31
lUU.o
2.858
106.5
1020 . 0
60.8
219.2
1233.3
' 1+.862
3.287
106 . 5
2501+ 7
25-5
222.8
2718.3
5.156
3.008
106.5
779.6
1+1.8
223.1
99!+. 8
1+.616
-123-
-------
Fuel Gas Gasifier Overall
Clean Fuel Regen. Cold Gas Station
Cleanup HHV Temp. Efficiency Efficiency
System Btu/SCF F % Coal HHV % Coal HHV
Selexol 1U2.6 685 73 A 31.2
Benfield 136.5 685 7^.9 30.5
Rectisol lM*.0 685 73.h 31.k
The above efficiencies do not necessarily represent the optimum power
cycle configuration and fuel gas regenerative heating that can be achieved.
However, in terms of their relative magnitudes for a particular system con-
figuration, it can be concluded that the three low-temperature desulfurization
processes can give comparable performance.
Integrated High-Temperature Cleanup Systems - Evaluation of integrated high-
temperature cleanup systems included four desulfurization processes which
appeared to have potential for the 1980 decade. These were the Consol and Air
Products dolomite systems, the Bureau of Mines iron oxide process, and the
Battelle molten salt process. Schematic flowsheets and corresponding material
balances are shown in Figures ^3-^o an(j Cables ~ ,respectively.
The flow scheme for the half-calcined dolomite process is basically that
which is currently proposed by Consolidation Coal Company. ^0) Both the
absorption and regeneration of I^S is conducted in fluidized beds and recovery
of elemental sulfur is via the liquid phase Wackenroeder reaction. Character-
istic absorption of HgS from the BCR gasifier raw gas is illustrated in Figure
UT. The partial pressure of CC>2 in the product gas is shown as a solid line for
each of two different operating pressures (the variation with temperature is due
to the water gas shift reaction). The limiting or minimum partial pressure
shown by the dashed line is the equilibrium partial pressure of CC^ above CaCOg.
The partial pressure of C02 must be greater than this to prevent decomposition
of the CaCOj to CaO. The intercept (Point A) indicates that at the absorber
operating pressure of ^50 psia, the maximum operating temperature could be 1780 F
and the residual sulfur would be 325 ppm (Point A'). However, the product gas
enters the absorber at only 1750 F and heat balance constraints limit the maxi-
mum temperature to about l650 F since the combined chemical reactions are endo-
thermic to the extent of 112 Btu/mol of feed gas. The residual sulfur content
at this temperature is around 600 ppm (Point B in Figure U7).
In the absorption step, 100 percent approach to equilibrium was assumed
for the H S and COS absorption reactions and for the CO shift reaction. Re-
generation of the dolomite acceptor occurs at 1300 F with an 85 percent approach
to HgS equilibrium and a 2:1 molar ratio of C02 to steam. The COg required for
regeneration is obtained by treating a portion of the clean fuel gas in a
-------
INTEGRATED CONSOL HIGH-TEMPERATURE CLEANUP SYSTEM
ro
P
4^
00
-------
INTEGRATED CALCINED DOLOMITE HIGH-TEMPERATURE CLEANUP SYSTEM
i
to
-------
INTEGRATED BUREAU OF MINES HIGH-TEMPERATURE CLEANUP SYSTEM
ro
-o
i
PRODUCT
GAS
«-• STACK
P
-P.
CJI
-------
INTEGRATED BATTELLE MOLTEN SALT HIGH-TEMPERATURE CLEANUP SYSTEM
TRANSPORT
GAS
ro
00
i
BCR
GAS
CONDENSATE SULFUR
CONDENSATE
P
*>.
CT)
-------
Table 22
ro
vo
I
STREAM
C02
Ho
M.W.
MATERIAL BALANCE FOR CONSOL CLEANUP SYSTEM
(see Figure Us)
2 3 ' U 5 6
MPH
MPH
MPH
MPH
H2S
COS
H20
TOTAL
STREAM
N2 28.02 '
02 32.01
C02 44.01 0.413
H2S 34.08
so2 64.07
H20 18.02 0.027
s 32.07
CaC03 MgC03 184.01
CaC03 MgO l40.4l
CaS MgO 112.46
INERTS 100
TOTAL " 0.440
0.084
o.oi4
0.017
0.029
0.005
0.939
0.003
0.003
0.038
0.009
0.029
0.007
0.007
0.098 o.o45 o.o45 0.051 0.945
7
MPH
0.940
2.883
0.666
8
MPH
28.02
28.01
44.01
2.016
16.04
34.08
60.08
17.03
18.02
M.W.
47.708
16.738
18.838
11.978
3-i4o
0.46o
0.100
0.380
10.658
100.000
9
MPH
47.708
16.009
10.156
12.707
3-l4o
0.055
0.008
0.380
10.334
100.497
10
MPH
3-845
1.290
0.778
1.024
0.253
o.oo4
0.001
0.031
0.793
8.019
11
MPH
2.917
0.979
0.621
0.777
0.192
0.003
0.023
0.632
6.144
12
MPH
2.917
0.979
0.124
0.777
0.192
0.003
0.023
0.143
5.158
13
MPH
43.863
14.719
8.881
11.683
2.887
0.051
0.007
0.349
9-052
91.492
14
MPH
4.606
0.520
'2.303
7.429
15
MPH
4.623
0.055
2.744
7.422
16
MPH
0.443
3.380
0.666
4.489 4.489
-------
Table 23
(JO
o
STREAM
MATERIAL BALANCE FOR CALCINED DOLOMITE
(see Figure UU)
23^5
M.W.
MPH
MPH
MPH
MPH
MPH
6
MPH
MPH
N2
CO
co2
H2
H2S
COS
T\rw
TJ /~\
TOTAL
STREAM
0
2
C02
S02
H20
Kr>
SULFUR
CO
H2S
NO
Wa SO 3
TOTAL
28.02
28.01
1+1+.01
2.016
16. ok
3k. 08
60.08
17.03
18.02
M.W.
32.00
28.02
1+1+.01
6k. 06
18.02
2.016
32.06
28.01
3^.08
30.01
126.05
if 7. 708
16.738
8.838
11.978
3.11+0
0.1+60
0.100
0.380
10.658
100.000
9
MPH
1.385
1.385
If7.708
16.376
9-255
12.339
3.11+0
0.029
0.005
0.380
10.728
99.960
10
MPH
0.976
0.976
0.1+1+6
2.398
3-881+
1.333
0.638
1.005
0.256
0.002
0.031
0.870
8.019
ik
MPH
2.1+70
1.122
0.597
0.019
0.067
^.275
1.355
0.1+65
0.222
0.350
0.089
0.011
0.303
2.795
16
MPH
0.390
1.1+68
1.858
9.161
3.1^5
1.778
2.369
0.603
0.006
0.001
0.073
2.060
19.196
17
MPH
0.1+28
1.610
2.038
9.161
3.1^5
0.356
2.369
0.603
0.006
0.001
0.073
2.023
17-737
18
MPH
0.056
1.9V7
1.111+
0.002
1.558
0.003
i+.68o
1+2.1+69
li+,578
6.973
10 . 981+
2.795
0.027
0.005
0.338
9-518
87.687
19
MPH
0.522
0.522
8
MPH
1.1+22
0.037
20
MPH
0.002
0.002
-------
Table 23 - Continued
STREAM
M.W.
CaC03 MgC03 18*4.. 01
11
MPH
0.0^0
0.619
0.122
12 13 15 21
MPH MPH MPH MPH
0.013
0.011 0.56*4-
0.566
0.002 0.093 0.095
0.002 ' 0.002 0.122 0.122
22
MPH
0.013
0.553
0.093
0.122
o MgO
CaO MgO 112.39
CaS MgO 112. h6
H Na0SOo 126.05
(jo 23
Inerts 100
TOTAL 0.781 0.015 0.015 0.781 0.781 0.781
-------
Table 24
uo
ro
i
STREAM
02
CO
CO^
COS
NH
NO
S
MATERIAL BALANCE FOR BUREAU OF MINES CLEANUP SYSTEM
(see Figure 45)
M.W.
32.00
28.02
28.01
44.01
2.016
16. 04
34.08
60.08
64.06
17.03
18.02
32.06
126.05
1
MPH
47.708
16.738
8.838
11.978
3.i4o
0.460
0.100
0.380
10.658
2
MPH
3.826
0.936
1.123
1.367
0.252
0.001
0.030
0.485
3
MPH
40.835
9-993
11.984
14.585
2.687
0.008
0.326
5-174
4 56789
MPH MPH MPH MPH MPH MPH
2.945 2.121 2.659
11.079 11.079 20.804 3.047
0.746
2.209 0.894
1.089
0.201
0.001
0.550 0.002
0.024
2.296 0.386
0.029
0.507
0.037
TOTAL:
100.000 8.020
85.592 14.024 13-750 27-999 0.037 0.507 6.388
-------
Table 25
STREAM
M.W.
MATERIAL BALANCE FOR BATTELLE MOLTEN SALT CLEANUP SYSTEM
(see Figure 46)
12 3456
MPH MPH MPH MPH MPH" MPH
N2
CO
C02
H2
c%
HgS
COS
NHo
H20
28.02
28.01
44.01
2.016
16. ok
3k. 08
60.08
17.03
18.02
47.708
16.738
8.838
11.978
3.140
o.46o
0.100
0.380
10.658
47.708
16.107
10.097
12.609
3.140
0.032
0.380
10.455
3.8U6
1.298
0.771
1.016
0.253
0.003
0.031
0.801
3.119
1.053
0.660
0.824
0.205
0.002
0.025
0.683
3.119
1.053
0.132
0.824
0.205
0.002
0.025
0.154
43.862
14.809
8.798
11.593
2.887
0.029
0.349
9-125
7 _ 8
MPH 'MPH
0.528 k.579
0.587
0.034 18.314
(JO
I
TOTAL
100.000
100.528
8.019
6.571
5.514
91.452
0.562 23.480
STREAM
N
C02
H2S
so2
H20
Sulfur
CaS
3
Other
M.W.
28.02
44.01
34.08
64.07
18.02
32.07
72.14
100.09
100
9 10
MPH MPH
4.579
0.059
11.871 2.712
11
MPH
i.oo4
•
0.003
0.003
12 . 13 14
MPH MPH MPH
16.096
0.525
0.192
7.008
4o.8oo
15
MPH
0.720
6.480
4o.8oo
16
MPH
3.728
TOTAL
11.871
7.350
1.010
0.525
16.096 48.ooo
48.ooo
3.728
-------
o
cr
D
CO
CO
LU
CC
CL
o.
CM
8
SULFUR ABSORPTION BY HALF CALCINED DOLOMITE
50
40
30
20
10
450 PS IA
300 PSIA
FIG. 47
Q.
0.
OC
D
to
CO
01
cc
600
500
400
300
1600
1650 1700 1750 1800
TEMPERATURE, °F
R1_19_14
-------
hot-potassium carbonate scrubbing unit. Spent dolomite, about 1 percent of
the circulating solids, is reacted with CC>2 in a water slurry to convert the
sulfide to the carbonate form before disposal to the environment In the sulfur
recovery system, 90 percent of the HgS feed is converted to elemental sulfur.
The conceptual flow scheme for the integrated Air Products calcined-
dolomite system is similar to the Consol process with the addition of a
calcination step. Effects of temperature on the characteristic absorption of
I^S and COS were calculated for the BCR raw gas case as shown in Figure h&.
Low residual sulfur levels, on the order of 100 ppm, are obtainable by operating
at temperatures of 1600-1650 F. However, at these conditions a substantial
amount of C02 is also absorbed by the calcined dolomite acceptor. Having an
exothermic heat of absorption of about 73>000 Btu/mol, absorption of C02
required heat removal facilities in the absorption cycle and a corresponding
amount of fuel for subsequent calcination of the carbonated acceptor. The
alternative is to operate at higher temperatures where the partial pressure of
C02 is less than the equilibrium pressure over CaO so that no C02 absorption is
possible. At this condition, above 17.50 F, the lowest residual sulfur content
attainable is 330 ppm.
Two case studies were therfore made for the calcined dolomite process at
absorption temperatures of 1600 F and 1770 F. It was assumed that complete
equilibrium was reached for sulfur and C02 absorption and for the CO shift
reaction at the respective absorption temperature. In the low-temperature
case, high-pressure steam generation is employed for removing the heat of
absorption. Regeneration is accomplished with steam and C02 at 1150 F and,
again, the required C02 is stripped from a portion of the product gas using a
potassium carbonate scrubber. The regenerator operated at 85 percent approach
to HoS equilibrium and a 1:1 steam to C02 molar ratio. Under these regenera-
tion conditions, the off-gas HgS concentration, about 18 percent, is high
enough for use as feed to a vapor phase Clause unit to recover elemental
sulfur. Sulfur recovery of 99 percent is obtainable in the Glaus plant
equipped with a tailgas treating system. Calcination is conducted at 1900 F
using combustion gases produced by burning a portion of the product fuel gas
with air. The following comparison, based on 2000 Ibs/hr coal feed to the
gasifier, illustrates the effect of operating temperature on the calcination
fuel requirements and product fuel gas characteristics:
Case 1 Case 2
Absorption temperature, F 1600 1770
Process fuel gas, mol/hr. 95.8 11.9
Product fuel gas, mol/hr. 2^1.5 373-9
Product temperature F 1550 1630
-135-
-------
FIG. 48
SULFUR ABSORPTION BY CALCINED DOLOMITE
Q
111
ui
u.
co
O
o
o
Q
111
m
oc
CO
CM
O
o
12 -
450 PSIA
400
450 PSIA
CL
Q_
of
D
D
CO
D
Q
CO
LU
DC
300
200
100
1600
1650
1700
1750
1800
TEMPERATURE, °F
-136-
R1-19-13
-------
Case 1 Case 2
HHV, Btu/SCF . 1^3 127
Sulfur, ppm 105 360
High-temperature operation of the calcined dolomite system, which minimizes
the process fuel requirement, would be expected to yield the more efficient
performance for an integrated COGAS plant.
Based in part on process description in the literature } a, conceptual
flow scheme for the integrated Bureau of Mines iron oxide system was developed.
Preliminary evaluation of this system indicated that it was most suitable for
desulfurization at temperatures below 1500 F, especially around 1000 F. The
case study, therefore, was based on sulfur absorption at 1000 F, and the BCR
raw gas, available at 1750 F, was first cooled to operating temperature by
generating high-pressure steam. Equilibrium absorption of H2S and COS from the
raw gas was calculated according the the following chemical reactions;
FeO + H2S - >- FeS + J^O (35) •
FeO + COS - *- FeS + C02 (36)
With CO shift eauilibrium assumed at 1000 F, the sulfur content in the treated
gas is reduced to about 170 ppm. The overall absorption reactions are exo-
thermic by 900 Btu/mol of feed.
The iron oxide system is operated in a cyclic fixed-bed fashion employing
an absorbent made up of 75 percent fly ash and 25 percent Fe^^. Sulfur
capacity was taken at 6.0 lb/100 Ib absorbent and regeneration was assumed to
be 80 percent efficient. Regeneration of the sulfided absorbent is via air
with a maximum temperature of 1500 F and the off -gas was assumed to have an
average SC^ concentration of h percent (vol). Recovery of elemental sulfur
with this system is complicated by having the regenerated sulfur in the form
of sulfur dioxide in low concentrations. The conceptual process, herein con-
sidered, first concentrates the'S02 "via a liquid scrubbing process on the
regenerator off -gas followed by partial reduction to H2S and then using the
Glaus reaction to produce elemental sulfur, i.e.
S02 + 3E£ - «- H2S + 2H20 (37)
2H2S + S02 - »- 3S + 21^0 .(38)
Reducing gas for the first reaction is supplied by a portion of the clean
product fuel gas. Recovery of elemental sulfur by this process sequence is
estimated at 92 percent of the S02 feed.
-137-
-------
Although the preliminary evaluation of the Battelle molten salt system
suggested that it was perhaps the least developed of the high-temperature
processes, an estimate of its performance in an integrated system was. made to
assess its potential. The conceptual flow scheme was totally derived from
/pt>\
bench scale experimental work \^°> . Sulfur absorption takes place in a single
stage venturi scrubber where molten salt concurrently contacts the raw gas.
The molten salt was assumed to contain 15 mol percent calcium salts in the
form of CaCOo and CaS. Because there was no data on COS absorption, it was
assumed as 100 percent and the H2S absorption was estimated according to the
reaction:
Ca C03 + H2S - — CaS + H20 + C02 (39)
where , ,
H?S = — : - -- x fCOpl [H201 x
'
(Ca C03) ' K
At ^50 psia, the following HpS absorption characteristics were calculated
assuming complete CO shift equilibrium and neglecting the non-ideality of the
salt solution:
Temperature F
1600 1650 1700
Equilibrium Constant, K 70 90 115
HoS residual in fuel
gas, ppm
@ 10$ CaCOo conversion ^50 390 270
@ 20$ CaC03 conversion 1020 880 620
@ 30$ CaCOo conversion 1750 1500 1060
For the purpose of this evaluation, H5S absorption was assumed at 1650 F
with 10 percent conversion of calcium carbonate resulting in a residual sulfur
level in the treated gas of 390
Regeneration of the sulfided salt melt is carried out at 1100 F with steam
and carbon dioxide in the molar ration of k:l. Salt regeneration was assumed
to be 75 percent efficient with an off-gas containing 2.5 percent HoS. As
with the dolomite processes, C02 for regeneration is extracted from a portion
of the clean fuel gas via hot potassium carbonate scrubbing. Elemental sulfur
is recovered from the dilute H?S regenerator off-gas in a liquid phase system
similar to that employed by the Consol process;
A summary of the estimated operation of the four high-temperature cleanup
processes is given in Table 26 for the case of treating raw gas from a BCR
-138-
-------
Table 2.6
OJ
MD
SYSTEM
Feed Streams
BCR Gas, mph
Flow
T
P
Product Streams
Sulfur
Flow
T
P
Transport Gas, mph
Flow
T
P
Product Gas
Flow
SUMMARY OF HIGH-TEMPERATURE INTEGRATED SYSTEMS
CONSOL AIR PROD. I AIR PROD. II BUREAU OF MINES
67.15
310
7l*. 82
300
CO
C02
H2
CIfy
H20
H2S+COS
HHV, Btu/scf
M.W.
Utilities
Cooling Duty, MMBtu/hr
Steam @ 1300 psia, Ib/hr
@ 65 psia, Ib/hr
Electric Power, MW
Boiler Feed Water, Ib/hr
Steam Condensate, Ib/hr
Process Water, Ib/hr
Natural Gas, scf/hr
(1) Based on 2000 Ib/hr Coal Feed to Gasifier
71-36
300
69.3
300
BATTELLE
MOLTEN SALT
1*26. 1*5
1750
1*50
1*26.1*5
1750
1*50
1*26.1*5
1750
1*50
1*26.1*5
1750
1*50
1*26.1*5
1750
1*50
71.8
310
31*. 20
300
550
390.17
1610
M5
1*7.91*
16.09
9-71
12.77
3-16
9-89
0.38
630 ppm
125.2
2l*.85
0.376
(61*6.9)
23.8
783.0
(122.0)
Ul. 6
31*. 20
300
550
21*1.1*7
1550
U20
5^-77
12.16
2.52
20.80
3.60
5.70
O.UU
105 ppm
1^3.1
21.96
1.039
(8102.6)
(322.0)
89^1. U
(U29-9)
18U.2
3k. 20
300
550
373-9^
1630
1*20
1*8.1*3
16.63
7.95
12.53
3-19
10.85
0.39
360 ppm
126.5
2l*.53
1.052
(585-1)
U.9
1252 . 9
(3^8.1*)
1^9-3
3^.20
300
550
372.85
1000
1*20
1*7-71
11.68
ll*.00
17..0U
3- 11*
6.01*
0.38
170 ppm
121*. 8
2l*.80
1.890
(11*90.0)
98.1
371*5.9
(2182.6)
3l*. 20
300
550
389.99
1610
1*10
1*7.96
16.19
9.62
12 68
3-16
9-98
0.38
390 ppm
125-3
2l*.81*
1.220
275-1
2l*.9
1113.0
129.6
-------
gasifier having a coal capacity of one ton per hour.
integrated station performances are as shown below;
Preliminary estimates of
Cleanup
System
Consol
Air Products
Case 1
Case 2
Bureau of Mines
Battelle Molten
Salt
Clean Fuel
HHV Temp.
Btu/SCF F
125.2 1610
1U3.1 1550
126.5 1630
12k.8 1000
125.3 1610
Gasifier
Cold Gas
Efficiency
% Coal HHV
76.1
53-8
73-7
72.5
76.1
Overall
Station
Efficiency
% Coal HHV
36.0
29.1
35.5
31.6
The relative performance of the integrated systems indicates that the Consol,
Air Products and molten salt systems are fairly comparable. The low relative
efficiency of the Bureau of Mines integrated system is due to the low absorp-
tion temperature required for efficient desulfurization. Cooling the raw
producer gas from 1750 F to 900 F, even though the heat is recovered by high-
pressure steam generation, represents a loss of efficiency in terms of the
relative power output from steam turbines as compared to higher efficiency gas
turbines. The Bureau of Mines process is therefore more suitable for
first-generation gasifiers which operate at 900-1200 F.
-------
SELECTION OF STANDARD INTEGRATED SYSTEMS
Based on the preliminary comparison of gas purification systems, those
processes which looked most attractive were selected for detailed evaluation
with integrated power systems both for first- and second-generation time
periods. Standard first-generation systems, judged to be commercially opera-
tive by 1980, were based on the fixed-bed gasifier concept typified by the
Bureau of Mines stirred-bed gasifier. Of the low-temperature desulfurization
processes currently available, the Selexol process was chosen for detailed
study of the integrated system. This selection was somewhat arbitrary, since
both the benfield and Rectisol processes showed comparable sulfur removal
performance with Selexol in the preliminary evaluation of integrated low-
temperature systems.
Although it is questionable whether any high-temperature cleanup system
will be commercialized by 1980, a comparison of integrated high- and
low-temperature gas purification systems was of interest in assessing the
potential performance benefit of high-temperature cleanup in conjunction with
first-generation gasifiers. The Bureau of Mines sintered iron oxide process
was selected as the most suitable high-temperature desulfurization' for this
application. This selection was made on the basis of operating temperature
for the first-generation gasifier, about 1000 F, and the fact that the process
has been demonstrated on a pilot, stirred-bed gasifier.
Standard integrated systems for second-generation operation, during the
1980 decade, employ the two-stage, entrained-flow gasifier of the BCR type.
Two standard systems were selected for detailed evaluation; one incorporating
a low-temperature cleanup process and the other based on high-temperature
desulfurization. For the latter, the Consol half-calcined dolomite process
was selected on the basis of the preliminary comparison of integrated high-
temperature cleanup systems. Although the Air Products fully calcined dolomite
process appeared comparable in performance, the Consol process was judged to
be further developed towards commercialization in the 1980s. Subsequent
analysis has indicated that the Consol process offers slightly better overall
performance when integrated into the power station (Section k).
There do no appear to be significant developments in low-temperature
desulfurization technology which would alter the types of processes available
for second-generation application. Therefore, the Selexol process was again
selected as being representative of the low-temperature cleanup technology
which would be applicable for future integrated systems.
As part of the detailed evaluation of first- and second-generation
standard systems, heat and mass balances, utilities requirements, investment
-------
cost estimates, and definition of pollutant streams were developed for the
various combinations of coal gasification and gas cleanup systems selected.
The evaluations were based on a coal feed rate of 8UOO tons/day which roughly
corresponds to a 1000-Mw COGAS power station output. The following gas turbine
operating conditions, judged to be attainable for the respective time periods
were specified, together with the corresponding process air inlet conditions
and fuel gas delivery pressure for the gasifier/purification system plant
section.
Generation First . Second
Gas turbine inlet, F 2200 2600
Compressor pressure ratio 16:1 2^:1
Process air inlet, F 800 800
Process air inlet, psia ^65 600
Fuel Gas Delivery, psia 355 U?0
First-Generation Standard Systems
The two gasifier/cleanup system combinations considered for this time
frame were:
(1) Bureau of Mines/Selexol
(2) Bureau of Mines/Sintered Iron Oxide
For each case study, the gasifier capacity was designed for 350 tons per hour
of Western Kentucky No. 9 coal having the following analysis and heating value:
CHS ON ASH H,0
Wt % . 6k.hl 5.U8 3.90 U.l8 1.58 15.hh U.95
HHV = 11^50 Btu/lb
The gasifier operating conditions were 2.7 lb air/lb coal plus tar feed, 0.35
Ib steam/lb coal plus tar, and the raw gas temperature was set at 1000 F. It
was assumed that all the nitrogen contained in the coal feed was converted to
ammonia, except for that present in unreacted tar. Furthermore, the raw gas
was assumed to contain 0.01 percent by volume of COS in accordance with the
chemical equilibrium for the hydrolysis reaction:
COS + H20 - C02 + H2S
Process flow schemes for the low-temperature and high-temperature purification
-------
cases are given in Figures ^9 and 50, respectively. The corresponding process
equipment lists and material balance summaries are given in Tables 27 through
30.
Bureau of Mines/Selexol System - Process Description - Coal and recycle tar
are fed to the stirred bed gasification unit together with 800 F air and 1315
psia saturated steam. Under reaction conditions, analysis indicates about 98
percent of the coal carbon (including carbon in the recycled tar) is gasified.
Raw producer gas leaves the gasifier at 1000 F and U20 psia and, after dust
removal in the cyclone separator F-101, it is cooled to 338 F by water quench
in T-201 where heavy tar components are condensed. The tar is separated from
the water phase in V-201 and recycled to the gasifier. Quenched gas, saturated
with water, is further cooled to 120 F by heat exchange in the sour water
stripper reboiler E-U02, the product gas reheater E-201, and air fan cooler
E-202. Condensate consisting of water plus light oils is separated in V-202
and sent back to the quench system. The cooled gas is then water scrubbed in
T-202 at 120 F to remove 90 percent of the ammonia content before being
desulfurized in the Selexol unit. Bottoms from the ammonia scrubber are
combined with other condensate streams from the acid gas removal system and
steam stripped in T-U01 where 95 percent of the ammonia and essentially all of
the dissolved acid gases are removed from the water stream before discharge to
the environment. The overhead gas from the stripper is cooled to 200 F and
then processed in the ammonia recovery unit where 99 percent of the ammonia
is recovered in the anhydrous form by extraction with an aromatic hydrocarbon
solvent. Off -gas from this unit contains HoS which is sent to the sulfur
recovery system. Relatively ammonia free raw gas from T-202 is desulfurized
by counter-current contact with Selexol solvent in absorber T-301. The
treated gas, desulfurized to 100 ppm by volume, leaves the top of the absorber
and is reheated to 265 F, via exchange with raw gas in E-201 and delivered to
battery limits at 355 psia. Rich Selexol solvent, containing dissolved HoS,
C02, and COS, is heated against lean solvent in E-302 and then steam regener-
ated in T-302. Acid gas, coming overhead from the regenerator, contains about
30 percent I^s by volume and is sent to the sulfur recovery unit.
The sulfur contained in the acid gas streams from the Selexol regenerator
and ammonia recovery unit is recovered as elemental sulfur via vapor phase
Glaus reaction. The ciaus sulfur recovery unit, shown in Figure 51, is a
split flow, three stage conversion unit in which 95 percent sulfur recovery
is obtained. About 1 percent of the clean product fuel gas is used to supply
necessary fuel for the sulfur plant. Tailgas from the last Glaus conversion
stage is incinerated and discharged to the atmosphere. This results in an
emission of about 0.25 Ib/MMBtu of coal input. A tailgas cleanup system
as the Strethford process, could reduce this at an increase in system cost.
-------
PROCESS FLOW DIAGRAM
BUREAU OF MINES/SELEXOL SYSTEM
RECYCLE TAR
STEAM
»-• STEAM COND.
c-
I
CLEAN FUEL GAS
STACK GAS
P
j^
CO
-------
PROCESS FLOW DIAGRAM
BUREAU OF MINES - SINTERED IRON OXIDE SYSTEM
-p-
ui
i
»- CLEAN FUEL GAS
»- STACK GAS
!
PURGE SULFUR
P
(Jt
o
-------
Table 2?
BUREAU OF MINES/SELEXOL SYSTEM
EQUIPMENT LIST
'SECTION 100
COAL GASIFICATION AND DUST REMOVAL
F-101 GASIFIER OFFGAS CYCLONE '
SECTION 200
GAS SCRUBBING AND TAR REMOVAL
P-201-Quench Water Recycle Pump
P-202 Quench Water Pump
P-203 Tar Recycle Pump
P-20U Gas Scrubber BTMS pump
P-201 Fuel Gas Reheat Exchanger
P-202 Gas Cooler
T-201 Quench Vessel
T-202 Water Scrubber
V-201 Tar/Water Separator
V-202 Gas/Liquid Separator
V-203 Oil/Water Separator
SECTION 300
SELEXOL ACID GAS REMOVAL SYSTEM
P-301 Selexol Stripper BTMS Pump
E-301 Selexol Solvent Cooler
E-302 Rich/Lean Solvent Exchanger
E-303 Selexol Stripper Reboiler
E-30U Selexol Stripper OVHD Condenser
V-301 Selexol Stripper OVHD Accumulator
T-301 Selexol Absorber
T-302 Selexol Stripper
-11+6-
-------
Table 2? - Continued
EQUIPMENT LIST
SECTION
SOUR WATER STRIPPING
P-U01 SWS Reflux Pump
P-402 SWS BTMS Pump
E-U01 SWS OVHD Condenser
E-U02 SWS Reboiler
E-U03 Feed/BTMs Exchanger
T-U01 Sour Water Stripper
V-U01 SWS OVHD Accumulator
SECTION 500
AMMONIA REMOVAL
SECTION 600
SULFUR RECOVERY
-------
Table 2b
-p-
co
STREAM
M.W.
02
Wo
CO
co2
H2
CHj,
H^S
COS
NH3
H20
TAR
ASH
32.00
28.02
28.01
kk.Oi
2.016
16. oU
3^.08
60.08
17.03
18.02
212
MATERIAL MIAIMCE FOR BUREAU OF MIWES/SELEXOL SYSTEM
(see Figure Uj?)
1 2 3 k
LB/HR MOL/HR LB/HR MOL/HR LB/HR MOL/HR LB/HR MOL/HR
U9l,l68 15,3^9
1,617,932 57,7^2
TOTAL
700,000
283,509 15,733
2,109,100 73,091 283,509 15,733
11^,132
STREAM
M.W.
w2
CO
C02
Hg
H2S
COS
3
TAR
TOTAL
28.02
28.01
Ml. 01
2.016
3k. Oo
60.08
17-03
18.02
212
LB/HR MOL/HR
13,619
LB/HR
1,617,931
7^3,637
29^,691
35,772
55,819
27,09^
721
13,^37
176,920
77,076
3,0^3,098
MOL/HR
57,7^2
26,5^9
6,696
17,7^
3,^80
795
12
789
9,818
36k
123,989
7 8
LB/HR MOL/HR LB/HR MOL/HR
77,076
77,076
36k
36k
1,617,931
7^3,637
29^,691
35,772
55,819
27,09^
721
13,^37
813, Mu
3,602,5^3 158,9^8
57,7^2
26,5^9
6,696
17,7^
3,^80
795
12
789
-------
VO
STREAM
M.W.
N
CO
C02
Cft
COS
NH3
TAR
TOTAL
STREAM
N2
CO
co2
0%
H2S
COS
NH
HpO
TAR
TOTAL
28.02
28.01
1+1+.01
2.016
16.01+
3l+. 08
60.08
17.03
18.02
212
M.W.
28.02
28.01
2.016
16.01+
3*+. 08
60.08
17.03
18.02
212
Table 28 - Continued
MATERIAL BALANCE FOR BUREAU OF MINES/SELEXOL SYSTEM
LB/HR MOL/HR
10
LB/HR MOL/HR
801*593 ^650
801+593 UU650
13
LB/HR MOL/HR
167712 9307
167712 9307
11 12
LB/HR MOL/HR LB/HR MOL/HR
168073
77076
21+511+9
11+
LB/HR
1617931
7^3637
291+691
35772
55819
2709!+
721
13^37
9208
9327
361+
9691
MOL/HR
577^2
265U9
6696
177^
31+80
795
12
789
511
168073
168073
15
LB/HR
1617931
7^3637
29171+2
35772
55819
21675
721
1618
9136
9327
9327
MOL/H
577^2
265^9
6629
177^
31+80
636
12
95
507
636520 35323
636520 35323
16
2798310 111+318
291+9 67
159
11819 691+
1370277 7601+2
2778051 11339^ 1390^6U 76962
-------
Table 28 - Continued
MATERIAL BALANCE FOR BUREAU OF MINES/SELEXOL SYSTEM
VJl
o
STREAM
M.W.
25
LB/HR MOL/HR
26 27 28
LB/HR MOL/HR • LB/HR MOL/HR LB/HR MOL/HR
CO
C02
?
COS
NH,,
H20
TOTAL
STREAM
°2
C.
CO
S°2
NO
H20
Sulfur
28.02
28.01
M*.oi
2.016
16. oi*
3^.08
60.08
17.03
18.02
M.W.
32.00
28.02
M*.01
61*. 06
30.01
18.02
32.06
11101* 652
iiioi* 652
29
LB/HR MOL/HR
3l*30l* 1072
113061 U035
119
921*1*
9363
LB/HR
681*8
129060
67335
2562
1861
22991*
-
7
513
520
30
MOL/HR
211*
1*606
1530
1*0
62
1276
M* l 2905
68 2 5351
596 35
1651*1*2 9181 288
166150 9219 85M*
31
LB/HR MOL/HR
21*183 75^
66
157
16
239
TOTAL
1^7365 5107
230660 7728
2^183
75^
-------
I
H
vn
H
STREAM
STREAM
Table 28 - Continued
MATERIAL BALANCE FOR BUREAU OF MINES/SELEXOL SYSTEM
M.W.
17
LB/HR MOL/HR
18
LB/HR MOL/HR
19
LB/HR MOL/HR
20
LB/HR MOL/HR
N2
CO
co2
Hg
CH,
H2S
COS ' •
NHo
^2°
TOTAL
28.02
28.01
1*1*. 01
2.016
16. Ok
34. 08
60.08
17.03
18.02
1611*372
741173
2l*l*78l*
35717
55322
102
l*8l
562
1370205 76038 . 180
1370205 76038 2692693
57615
261*61
5562
17717
3kk9
3
8
33
10
110858
3559
21*65
46959
5!*
497
21573
2kO
1056
1766
78169
127
88
1067
27
31
633
i*
62
98 7190
2137 7190
399
399
M.W.
21
LB/HR MOL/HR
22
LB/HR MOL/HR
23 2^
LB/HR MOL/HR LB/HR MOL/HR
N2
CO
co2
%
CH,
HS
COS
NHo
HpO
28.02
28.01
44.01
2.016
16. Ok
3^.08
60.08
17.03
18.02
1241*1
571^
1892
274
^33
kkk
20k
1*3
136
27
1601931
735^59
21*2891
3^553
51*889
102
1*81
562
108
57171
26257
5519
17581
3
8
33
10
2905
5351
11223
9533
66
157
659
529
1*1*
68
596
1535646
1
. 2
35
85219
TOTAL
2075^
85k
2671938
11000U
29012 lUll 153635^ 85257
-------
Table 29
BUREAU OF MINES/IRON OXIDE SYSTEM
EQUIPMENT LIST
ITEM
Pumps
P-101
P-120
Reactors
• R-101
Drums
D-101
Exchangers
E-101
E-102
E-103
Compressors
C-101
Separators
F-101
F-102
DESCRIPTION
Boiler Recycle Pump
Boiler Feed Water Pump
Sulfur Absorber/Regenerator
Steam Drum
Waste Heat Boiler
Economizer
Gas Cooler
Air Compressor
Gasifier Off-Gas Cyclone
Regenerator Off-Gas Cyclone
-152-
-------
Table 30
STREAM
M.W.
MATERIAL BALANCE FOR BUREAU OF MINES/IRON OXIDE SYSTEM
(see Figure
1 -2
LB/HR MOL/HR LB/HR MOL/HR
—
LB/HR MOL/HR
LB/HR MOL/HR
I
H
U)
C2
N2
CO
CQ2
H2
COS
NH3
H20
TAR
ASH
TOTAL
STREAM
CO
C02
H2
H2S
COS
NH3
TAR
°2
32.00
28.02
28.01
2.016
3^.08
60.08
17.03
18.02
212
M.W.
28.02
28.01
2.016
16.014.
3^.08
60.08
17.03
18.02
212
32.00
114.14.3366
13695
51512
700000
1881606
65207
6
LB/HR MOL/HR LB/HR MOL/HR
214.1+014-5 135^3
2M4O14-5 135^3
LB/HR MOL/HR
1114-132
1114-132
8
LB/HR MOL/HR
TOTAL
11625
11^3366
639076
253278
27000
U7960
25867
721
n6i<-9
175136
73819
2697872
51512
22816
5755
13393
2990
759
12
68^
9719
3U8
107988
13^2382
1*31^66
i4-91900
36830
1*14607
307
10831
70656
68651
2*4.97630
U7908
I5k0k
11177
18269
2781
9
636
3921
32U
100^29
337865 12058
102528 320U
1&0393 15262
-------
Table 30 - Continued
MATERIAL BALANCE FOR BUREAU OF MUTES/IRON OXIDE SYSTEM
vn
-F-
i
STREAM
TOTAL
M.W.
13
LB/KR MOL/HR
Ik
LB/HR MOL/HR
02
N2
CO
co2
H2
CEk
H2S
COS
so2
NHo
H28
TAR
32.00
28.02
28.01
Ml-. 01
2.016
16. 04
34.08
60.08
64.06
17.03
18.02-
212
67776 2118
223291 7969 ioio4o
32492
37012
2772
3352
3k
8l7
5316
5168
3606
1160
84l
1375
209
1
48
295
24
291067 10087 188003
7559
15 16
LB/HR MOL/HR .LB/HR MOL/HR
59936 1873
337865 12058
48750
761
141148
446551 14692 l4ll48
7833
7833
STREAM
M.W.
18.02
126.05
17
LB/HR MOL/HR
TOTAL
138376
138376
7679
7679
18
LB/HR MOL/HR
2775
2775
154
154
19
LB/HR MOL/HR
7185-
7185
57
57
-------
Table 30 - Continued
MATERIAL BALANCE FOR BUREAU OF MINES/IRON OXIDE SYSTEM
i
H1
vn
v_n
i
STREAM
M. W.
LB/HR MOL/HR
10
LB/HR MOL/HR
11
LB/HR MOL/HR
12
LB/HR MOL/HR
°2
W2
CO
C02
H2
CH1+
H2S
COS
S02
NH3
WO
H20
Sulfur
TAR
DUST
32.00
28.02
28.01
M+.Ol
2.016
16.01+
3*+. 08
60.08
61+. 06
17.03
30.01
18.02
32.06
212
11+1+3H22
1163958
528912
39602
1+7960
3^1
1161+9
75972
73819
5151^
1656U
12018
196UU
2990
10
681+
1+216
31+8
77632
662169
H3i91+
1+292
1621
1+3969
21+26
23632
2572
67
^
21+1+0
20518
61+0
199^
TOTAL
2685635 107988
902877 31191
20518
6i+o
-------
PROCESS FLOW DIAGRAM
SULFUR RECOVERY UNIT CLAUS PROCESS
H
VJl
ON
ACID
GAS
FEED
FUEL BFW
C-101
STACK
D
P
U1
-------
About 3 percent of the coal heating value is removed in the gasifier
cooling jacket via generation of 1315 psia saturated steam. This internally
generated steam supplies 50 percent of the steam required for the gasification
reaction.
Bureau of Mines/Iron Oxide System Process Description - In this system, coal
is the only carbonaceous feed to the gasifier. Unreacted tar in the raw
producer gas is not condensed at the cleanup system operating temperature and,
therefore, cannot be recycled. Reaction of the coal with 800 F air and 1315
psia saturated steam results in about 8^ percent gasification of the coal
carbon with the majority of the remaining carbon released as a component of the
tar. Raw producer gas leaves the gasifier at 1000 F and UOO psia and contains
1.8 Ib tar per 1000 SCF and 0.3 Ib dust per 1000 SCF. The gas passes first
through a cyclone separator and then through an undefined high-temperature
particulate removal system to reduce the dust loading to an acceptable level.
Desulfurization takes place over a fixed bed of sintered iron oxide/fly ash
absorbent in R-101 at a space velocity of 1000 reciprocal hours (measured at
60 F and 1.0 atm). A 90 percent approach to shift equilibrium and 100 percent
approach to sulfur removal is assumed giving a residual sulfur content in the
treated gas of 100 ppm. The overall desulfurization reactions are exothermic
so that the fuel gas product is delivered to battery limits at 1070 F and
355 psia. When the sulfur loading reaches 6 lb/100 Ib absorbent, the absorp-
tion cycle is discontinued and the spent absorbent is regenerated with air at
a maximum bed temperature of 1500 F. Regeneration proceeds at a high rate,
6-10 percent S02 in the gas stream, for the first 30 minutes and decreases to
about 1 percent S02 after one hour. During the one hour regeneration period
80 percent of the sulfur is recovered as SC>2 with the off-gas having an average
concentration of k percent by volume S02. The regenerator off-gas is cooled
to 150 F via generation of 65 psia steam in waste heat boiler E-101, boiler
feed water preheater E-102, and air fan cooler E-103.
Recovery of elemental sulfur from the off-gas is complicated by the fact
that sulfur is present as S0p in variable concentration. The S0? is first
concentrated by absorption in an alkali solution from which it is steam .
stripped. About 93 percent of the feed S02 is recovered as a 50 percent by
volume S02 stream. Two-thirds of the recovered S02 is converted'to H S using
a portion of the product fuel gas as the reducing agent. The remaining SOp is
then reacted with the J^S to form elemental sulfur in a typical vapor phase
Glaus unit. Overall recovery of elemental sulfur present in the coal is Qk
percent, and 7 percent of the fuel gas product is used in the sulfur recovery
section. Of the sulfur lost, 1.3 percent is in the powerplant stack, 8.7 per- .
cent in the Glaus plant exhaust and 7.0 percent in the Glaus purge stream. The
65 psia steam generated in E-101 represents 20 percent of the low-pressure
steam requirements for the SOo concentration unit.
-157-
-------
The tar present in the fuel gas product has the following analysis:
C H 0 N S
% wt. 8U.02 7.77 h.66 2.00 1.55
Second-Generation Standard Systems
Gasifier/cleanup system combinations evaluated for application in second-
generation integrated power systems were:
(1) BCR/Selexol
(2) BCR/Consol
The two-stage BCR gasification system operates on a 35Q tons per hour of
Illinois No. 6 coal having the following analysis;
C H S 0 N ASH HpO
Wt. % 67.^ 5-1 3-8 9.6 1.2 8.7 U.2
HHV = 12200 Btu/lb
Operating .conditions for the gasifier were 3-1 lb air/lb coal, 0.57 lb steam/
lb coal with first stage and second stage exit temperatures of 2800 F and 1800
F, respectively. A portion of the clean fuel gas product, 6.6 SCF/lb of coal,
is recycled to the gasifier to transport .the pulverized coal into the second
stage. About 95 percent of the coal nitrogen is converted to ammonia and the
COS content in the raw gas corresponds to the equilibrium for the hydrolysis
reaction to HpS. Heat is recovered from recycle char cooling in the amount of
190 Btu/lb of coal and is. available for steam generation and boiler feed water
heating.
Process flow diagrams for the low- and high-temperature gas purification
schemes are given in Figures 52 and 535 respectively. Corresponding process
equipment lists and material balance summaries are given in Tables 31 through
BCR/Selexol System - Process Description - Coal is gasified using 800 F air
and 1250 psia superheated steam. Under the reaction conditions, essentially
complete carbon conversion is obtained. Molten slag drained from the bottom
stage is water quenched and sent to battery limits for disposal. The slag
quench water is clarified and sent to a sludge pond for additional solids
removal before being reused. Raw producer gas leaving the top stage of the
-158-
-------
PROCESS FLOW DIAGRAM
BCR/SELEXOL SYSTEM
FIG. 52
STACK
COAL
AIR
SLAG
SULFUR
SLUDGE
POND
AMMONIA
WASTE
WATER
R1-19-25
-159-
-------
Fig. 53
PROCESS FLOW DIAGRAM BCR/CONSOL SYSTEM
COAL TRANSPORT OAS
-160-
-------
Table 31
BCR/SELEXOL SYSTEM
EQUIPMENT LIST
SECTION 100 - GASIFICATION
ITEM.
DESCRIPTION
F-101
Particulate Removal System
SECTION 200 - HEAT RECOVERY
ITEM
DESCRIPTION
Vessels
V-201
V-202
HP Steam Drum
LP Steam Drum
Exchangers'
E-201
E-202
E-203
E-20U
E-205
E-206
E-20?
HP Waste-Heat Boiler
HP Economizer
Main Regenerator
LP Waste-Heat Boiler
LP Economizer
Auxiliary Regenerator
Gas Cooler
Pumps
P-201
P-202
Compressors
C-201
HP Recirculating Pump
LP Re circulating Pump.
Transport Gas Compressor
-161-
-------
Table 31 - Continued
BCR/SELEXOL SYSTEM
EQUIPMENT LIST
SECTION 300 - GAS SCRUBBING AND SWS
ITEM
Towers
T-301
T-302
Vessels
V-301
V-302
Exchangers
E-301
E-302
E-303
E-305
Pumps
P-301
P-302
P-303
DESCRIPTION
NHo Scrubber
NHo Stripper
Condensate Knock-Out Drum
NH Stripper OVHD Accumulator
NHo Scrubber OVHD Exchanger
NHo Stripper BTMS Exchanger
NH3 Stripper BTMS Cooler
NHo Stripper Reboiler
NH3 Stripper OVHD Condenser
NHo, Absorber BTMS Pump
NH^ Stripper BTMS Pump
NH Stripper Reflux Pump
-162-
-------
Table 31 - Continued
BCR/SELEXOL SYSTEM
EQUIPMENT LIST
SECTION kOO - ACID GAS REMOVAL
ITEM DESCRIPTION
Towers
T-H01 Selexol Scrubber
T-U02 Selexol Stripper
Vessels
V-U01 Condensate Knock-Out Drum 2
V-402 Selexol Flash Drum
V-403 Selexol Stripper OVHD Accumulator
Exchangers
E-U01 Lean Solvent Cooler
E-402 Rich/Lean Solvent Exchanger
E-U03 Selexol Stripper Reboiler .
E-^OU Selexol Stripper OVHD Cooler
Pumps
P-U01 Selexol Stripper Reflux Pump
P-U02 Selexol Stripper BTMS Pump
Compressor
C-U01 Recycle Gas Compressor
SECTION 500 - AMMONIA RECOVERY
SECTION 600 - SULFUR RECOVERY
-163-
-------
Table 32
STREAM
1
H
ON
°2
N2
CO
co2
H2
H2S
COS
NH3
N20
32.00
28.02
28.01
i&.Ol
2.016
16. (A-
3^.08
60.08
17.03
18.02
TOTAL
STREAM'
N?
CO
co2
H2
H2S
COS
WH3
H20
28.02
28.01
2.016
16. oh
3^.08
60.08
17.03
18.02
MATERIAL BALANCE FOR BCR/SELEXOL SYSTEM
(see Figure 52)
M.W. LB/HR MOL/HR LB/HR 2MOL/HR LB/HR MOL/HR LB/HR MOL/HR
17991^
71809
3^91
3621
8152
3
66
2
23
6^20. 9
2563.7
783.7
1795.9
508.2
0.1
l.l
0.1
1.3
503392 15731.0
1658190 59178.8
700000
298081 12075.0 2161582 7U909.8
396900 22025.5
396900 22025.5
M.W. LB/HR 5MOL/HR LB/HR 6MOL/HR LB/HR 7MOL/HR LB/HR 8MOL/HR
1835271
732^92
517162
36932
83222
23^16
85^9
9763
251^38
65^98.6
26151.1
11751.0
18319. ^
5188. h
687.1
1^2.3
573.3
13953.:
TOTAL
60900
314.982^5
13528.5
13528.5
1835271
732^92
511990
36932
83222
21075
85^9
976
78U6
3238353
65^98.6
26151.1
11633.5
18319. h
5188. U
618. k
lU2.3
57.3
^35. U
12801^. U
-------
Table 32 - Continued
MATERIAL BALANCE FOR BCR/SELEXOL SYSTEM
STREAM
I
H
vji
M.W.
LB/HR ^MOL/HR
LB/HR 10MOL/HR
LB/HR 1:LMOL/HR
LB/HR "^MOL/HR
CO
C02
H2
H2S
COS
NHo
N23
TOTAL
STREAM
N2
CO
C02
H2
CHi,.
H2S
COS
WH3
H20
28.02
28.01
144.01
2.016
I6.0l4-
34.08
60.08
17.03
18.02
M.W.
28.02
28.01
144.01
2.016
16. (A
34.08
60.08
17.03
18.02
5171
2341
8787
1017521
1033820
117.5
68.7
516.0
56466.2
57168. 4
LB/HR 13MOL/HR
1835271
732492
351838
36932
83155
Ul
. 655
22
234
65498.6
26151.1
799^.5
18319.^
5184.2
1.2
10.9
1.3
13.0
1017712
1017712
5614.76.8
561+76.8
LB/HR ^MOL/HR
32265
12876
6183
61+9
114-61
12
4
1151.5
459.7
11+0.5
322.1
91.1
0.2
0.2
1710
1710
9^.9
9^.9
210000
210000
11653.7
11653.7
-, r
LB/HR
1623092
647807
311164
32662
735^2
37
577
20
207
5MOL/HR
57926.2
23127.7
7070.3
16201.4
4581+.9
l.l
9.6
1.2
11.5
LB/HR MOL/HR
160152"
67
21034
7895
950
5902
3639.0
U.2
617.2
131.4
56.0
327.5
TOTAL
12317^.2
53^50 2165.3 2689108 108933.9
-------
STEEAM
Table 32 - Continued
MATERIAL BALANCE FOR BCR/SELEXOL SYSTEM
M.W.
LB/HR 17MOL/HR
18.
LB/HR MOL/HR LB/HR 19MOL/HR
LB/HR 2°MOL/HR
N2
CO
C02
H2
H2S
COS
NHo
H20
TOTAL
STREAM
°2
N2
C02
NO
H2S
S02
NH3
H20
28.02
28.01
if if. 01
2.016
l6.0if
3if.08
60.08
17.03
18.02
M.W.
32.00
28.02
if if. 01
30.01
3if.o8
6if.o6
17.03
18.02
5171 117.5
23ifl 68.7
8787 516.0
1263015 70089.6
I2793lif 70791.8
LB/HR 21MOL/HR
8265 U85.3
5118
2317
83if8
7287
23070
LB/HR
5118
2317
232
116.3 53
68.0 2if
if 90. 2 if39
Uoif.if 1255727
1078.9 I2562if3
22MOL/HR LB/HR
if5011
lif8271
116.3
68.0
. 12.9
1.2
0.7
25.8
69685.2
69712.9
23MOL/HR
Iif06.6
5291.6
83
7055
7138
LB/HR 2
9002
180536
201671
1681
if!96
29237
k.9
391.5
396. if
W/HR
281.3
6ifif3.l
if 582. if
56.0
65.5
1622.5
TOTAL
8265 if85.3
7667. 197.2 193282 6698.2
if 26323 13050.8
-------
&
STREAM
M.W.
Table 32 - Continued
MATERIAL BALANCE FOR BCR/SEIEXOL SYSTEM
25
LB/HR MOL/HR
26
LB/HR MOL/HR
Sulfur
co2
HgS
NH3
H20
TOTAL
32.06 24086 751.3
44.01
34.08
17.03
18.02
24086 751.3
53
24
439
238015
238531
1.2
0.7
25.8
13208.4
13236.1
27
LB/HR MOL/HR
754920
754920 41893.4
28
LB/HR MOL/HR
469270 26*1.6
469270 .26041.6
-------
Table 33 .
BCR/CONSOL SYSTEM
EQUIPMENT LIST
SECTION 100 - DESULFURIZATION
ITEM
Reactors
R-101
R-102
Vessels
V-101
V-102
V-103
Pumps
P-101
Exchangers
E-101
E-102
E-103
Mi s c ellane ous
F-101
F-102
DESCRIPTION
Sulfur Absorber
Acceptor Regenerator
Dolomite Feed Hopper
Spent Dolomite Hopper
HP Steam Drum
BRW Circulation Pump
Waste Heat Boiler
Economizer
Gas Cooler
Absorber Cyclone Separator
Regenerator Cyclone Separator
-168-
-------
Table 33 - Continued
BCR/CONSOL SYSTEM
EQUIPMENT LIST
SECTION 200 - SPENT DOLOMITE TREATING
ITEM
Reactors
R-201
R-202
R-203
Vessels
V-201
Pumps
P-201
P-202
P-203
Exchangers
E-201
E-202
Compressors
C-201
C-202
Miscellaneous
F-201
DESCRIPTION
Acceptor Converter
Acceptor Converter
I st
5nd
Stage
Stage
Acceptor Converter 3rd Stage
Quench Water Surge
Quench Water Pump
Dolomite Slurry Pump
Make-up Water Pump
Quench Water Cooler
C02 Trim Cooler
C02 Blower
Acid Gas Compressor
Hydroclone
-169-
-------
Table 33 - Continued
BCR/CONSOL SYSTEM
EQUIPMEOT LIST
SECTION 300 - C02 RECOVERY
ITEM . DESCRIPTION
Tower
T-301
T-203
Vessels
V-301
V-302
Pumps
P-301
P-302
Compressors
C-301
C-302
Exchangers
E-301
E-302
C02 Absorber
CO Stripper
Water Separating Drum
Stripper OVHD Accumulator
Stripper BTMS Pump
Stripper Reflux Pump
COp Blower
Transport Gas Compressor
Stripper OVHD Condenser
Stripper Reboiler
-170-
-------
Table 33 - Continued
BCR/CONSOL SYSTEM
EQUIPMENT LIST
SECTION ^00 - SULFUR RECOVERY
ITEM
Reactors
R-U01
Towers
T-401
Vessels
V-401
V-U02
Pumps
P-U01
P-l+02
P-U03
Compressors
C-U01
C-U02
Exchangers
E-tol
E-U02
Miscellaneous
F-401
B-U01
DESCRIPTION
Liquid Phase Clause Reactor
o Absorption Column
Sulfur Separator Drum
Sulfur Storage Drum
Sulfur Pump
Acid Pump
Acid Circulating Pump
Recycle CO Compressor
Air Compressor
Recycle C02 Reheater
Feed/Bottoms Exchanger
Weak Acid Cooler
S02 Absorber Intercooler
BFW Preheater
Electrostatic Precipitator
Sulfur Burner
-171-
-------
Table
ro
i
STEEAM
02
W2
CO
co2
H2
CHi,
H2S
COS
NHo
H20
TOTAL
STREAM
32.00
28.02
28.01
Ml. 01
2.016
16. Oil
3^.08
60.08
17.03
18.02
MATERIAL BALANCE FOR BCR/CONSOL SYSTEM
(see Figure 53)
M.W. LB/HR MOL/HR LB/HR 2MOL/HR LB/HR 3MOL/HR LB/HR SdL/HR
15731.1
59178.7
M.W.
N?
CO
C02
Hp
CH^l
H2S
COS
WHo
H20
TOTAL
28.02
28.01
UU.oi
2.016
16. Oil
3*1.08
60.08
17.03
18.02
700000
LB/HR MOL/HR
178053
67395
23537
3761
7919
228
5^
1003
637^
28832^
LB/HE
1839079
716557
513583
3737^
81796
23369
8621
10371
258373
635^.5
2^06.1
53^.8
1865.7
^93.7
6.7
0.9
58.9
353.7
12075.0
6 /
MOL/HR
6563U.5
25582.2
11669.7
18538.8
5099.5
685.7
1^3.5
609.0
1U338.1
503395
1658187
2161582
LB/HR
1839079
696085
58^880
38850
81796
23^1
517
10371
256288
7^909.8
7
MOL/HR
6563^,
2U851,
13289,
19270.9
5099.5
68.7
8.6
609.0
396900 22025.5
396900 22025.5
o
LB/HR MOL/HR
1661026 59280.0
628693 22^5.3
528252 12003.0
35089 17U05.1
73877 1*605.8
2113 62.0
1*69 7.8
9367 550.0
60900
3^89123 1^2301.0 3510207 li)-305^.5 3170360 129201^.
-------
u>
I
Table 34 - Continued
MATERIAL BALANCE FOR BCR/CONSOL SYSTEM
STREAM
M.W.
LB/HR MOL/HR
LB/HR 10MOL/HR
11
LB/HR MOL/HR
N2
CO
C02
H^
CHI,.
H2S
COS
WHo
H20
TOTAL
STREAM
CaCO^MgCOo
CaC03MgO
CaS MgO
INERTS
TOTAL
28.02
28.01
44.01
2.016
16. 04
34.08
60.08
17.03
18.02
M.W.
184.01
l4o.4i
112.46
100
178053
67392
56628
376i
7919
228
48
1005
24814
339848
LB/HR
10636
1010
11646
6354.5
24o6.o
1286.7
1865.7
493.7
6.7
0.8
59.0
1377.0
13850.1
13MOL/HR
57.8
10.1
67.9
l6l42
i6l42
LB/HR
10636
94201
574997
101750
781584
895.8
895.8
ll|MOL/HR
57.8
670.9
5112.9
1017.5
6859.1
311071
2846
. 70588
384505
LB/HR
199775
490438
100740
790953
7068.2
83.5
3917.2
11068.9
15MOL/HR
1422.8
4361.0
1007.4
6791.2
LB/HR ^MOL/HR
282443 6417.7
26985 791.8
57824 3208.9
367252 10418.4
LB/HR l6MOL/HR
i^. 2
4903 43.6
1010 10.1
7907 67.9
-------
Table 3!* - Continued
MATERIAL BALANCE FOR BCR/CONSOL SYSTEM
-p-
STREAM
M.W.
C02 1*1*. 01
H2S 3**.08
18.02
l8i*.01
INERTS 100
TOTAL
STREAM
°
M.W.
32.00
N2 28.02
S02 61*. 06
H20 18.02
Sulfur 32.06
TOTAL
LB/HR 17MOL/HR
5967 l**8.3
13312 1*16.0
1*3851 1565.0
57163 1981.0
18
LB/HR MOL/HR
11*780
820.2
LB/HR 21MOL/HR LB/HR 22MOL/HR
23962 7^7. ^
23962 71*7.1*
LB/HR 19MOL/HR LB/HR 20MOL/HR
5580
387
126.8
21.5
1^780
820.2
1118
11+86
ll*8
25. U
^3.6
8.2
2752
LB/HR
1210
288
101
77.2
3MOL/
HR
37.8
1565.0
u-§
5:6
1612.9
789.9
10636 57.8
1010 10.1
25880 857.8
LB/HR MOL/HR
1+8U36 2687.9
1*81*36 2687.9
-------
Table 34 - Continued
MATERIAL BALANCE FOR BCR/CONSOL SYSTEM
STREAM
VJl
i
M.W.
25
LB/HR MOL/HR
26
LB/HR MOL/HR
C02
H2S
H20
TOTAL
STREAM
N2
CO
C02
H2
CHU
H2S
COS
NH2
H20
44.01
34.08
18.02
M.W.
28.02
28.01
44.01
2.016
16.04
34.08
60.08
17.03
18.02
283561
2846
22903
309310
LB/HR
122414
46334
58^0
2586
5444
157
36
690
4646
6443.1
83.5
1271.0
7797.6
29
MOL/HR •
4363.8
1654.2
132.7
1282.7
339-4
' 4.6
0.6
40.5
257.8
27511
1910
29421
LB/HR
55639
21061
17696
1175
2475
72
18
313
1728
625.1
106.0
731.1
30
MOL/HR
1985.7
751-9
402.1
583.0
154.3
2.1
0.3
18.4
95-9
27
LB/HR MOL/HR
28
LB/HR MOL/HR
45774 2540.2 210000 11653.7
45774 2540.2 . 210000 11653.7
31 32
LB/HR MOL/HR LB/HR MOL/HR
TOTAL
188147 8081.3 100177 3993.7
48880 2712.5 193530 10739.7
48880 2712.5 193530 10739.7
-------
gasifier at 1800 F and 560 psia passes through a series of cyclone separators
to remove entrained char and is then cooled to 120 F in the regenerative
exchangers E-203 and E-206, generating 1370 psia and 65 psia steam in the waste
heat boilers E-201 and E-20U, and preheating boiler feed water in E-202 and
E-205. Cooled gas is scrubbed with water in T-301 to remove 90 percent of the
ammonia and is then sent to the Selexol acid-gas removal system. Condensate
from separator drums V-301 and V-U01 is combined with the bottoms from the
water scrubber and fed to the sour water stripper, T-302, where steam stripping
removes 95 percent of the dissolved ammonia and essentially all of the acid
gases. Part of the stripped water is recycled to the water'scrubber, T-301,
and the remainder is delivered to battery limits as waste water.
The overhead gas from the sour water stripper is fed to an ammonia
recovery unit where 99 percent of the ammonia is recovered in the anhydrous
form via absorption in an aromatic hydrocarbon solvent.
Ammonia-free producer gas is desulfurized in 1-kOI by countercurrent con-
tacting with Selexol solvent at subambient temperature. The treated gas, con-
taining 100 ppm total sulfur, is reheated against raw producer gas and is
delivered to battery limits at 1000 F and 1^75 psia. Rich Selexol solvent,
from the bottom of T-Uoi, is flashed, heated against lean solvent in E-U02, and
finally stripped in the regenerator T-H02. The stripped acid gases from the
regenerator are cooled to 120 F and combined with the off-gas from the ammonia
recovery unit to provide the feed to the sulfur recovery plant. The combined
acid gas stream contains Ik percent H S so that a split flow, three-stage
vapor phase Glaus system (Figure 51, previously shown) is used to recover
95 percent of the sulfur in the elemental form. About 2 percent of the product
fuel gas is required as process fuel for the sulfur recovery unit.
Low-pressure steam generated in E-20h plus that exported from the Glaus
sulfur plant provides 98 percent of the 50 psig steam requirement for the
Selexol regenerator and the sour water stripping system. The high-pressure
steam generated in E-201 is delivered to the steam generator portion of the
combined-cycle power system for superheating.
BCR/Consol System - Process Description - As in the low-temperature cleanup
case, coal is gasified in the two-stage BCR gasifier using 800 F air and 1250
psia superheated steam. Again, essentially complete carbon conversion is
achieved. Heat from the char recycle system is used to generate 1300 psia
saturated steam. After high-temperature particulate removal, the raw producer
gas enters the desulfurization system at 1750 F. In the fluidized-bed
desulfurizer, R-101, the gas contacts half-calcined dolomite acceptor at ^95
psia. The acceptor enters with 75 percent of the calcium as CaS and leaves
with 88 percent" in the sulfided form. Approach to equilibrium for the CO shift
and sulfur absorption reactions is assumed to be 100 percent. The overall
-176-
-------
reactions occuring during desulfurization are slightly endothermic so that
the treated gas exits at 1700 F and contains 5^0 ppm total sulfur.
Sulfided acceptor plus an amount of fresh dolomite equal to slightly more
than 1 percent of the circulating solids, are transported to the fluidized-bed
regenerator, R-102, by the regeneration gas having a molar ratio of carbon
dioxide to steam approximately equal to 2. Regeneration is carried out at
1300 F with an 85 percent approach to H S equilibrium. About lU percent of the
CaS is converted and the regenerated solids are recycled by gravity to the
desulfurizer. The off -gas from the regenerator contains 7$ H S by volume and,
after cooling to 380 F, is fed to a liquid-phase sulfur recovery unit.
Spent dolomite, withdrawn from the regenerator via lock hopper V-102, is
treated before discharge to the environment. This stream, containing 75 per-
cent of the calcium as CaS, is slurried in water in a hydrocyclone F-201 and
then processed in a three-stage countercurrent reactor system where CC>2 is used
to convert all the calcium to the carbonate form thereby rendering the stream
suitable for discharge to a sludge pond. The HgS generated in the spent dolo-
mite system exits from R-201, is compressed, and fed to the sulfur recovery
unit along with the regenerator of f -gas .
The liquid phase Glaus reactor, R-^01, operates at 310 F and converts 90
percent of the HoS feed to elemental sulfur. Feed gas is contacted counter-
currently by a solution of H SO whereby sulfur is formed via the reaction
23
2H S + H SO - 3S + 3H 0 (-U2)
Liquid sulfur is separated from the aqueous phase in V-Uoi and stored in the
sulfur accumulator 7-^02. A third of the sulfur product is burned with air
in B-^01 to produce SO which in turn is absorbed by contact with water in
T-^01. The rich H SO solution from T-U01 is heated against the aqueous phase
from V-^01 and fed to the sulfur converted R-U01. Excess water formed by the
reaction in R-401 is removed from the system and sent to a sour water stripper.
The overhead gas from the sulfur reactor is compressed, combined with
makeup C02 and steam, and recycled to the dolomite regenerator. Makeup COp,
required for spent dolomite treating and acceptor regeneration, is extracted
from a slip stream of product fuel gas . This stream, representing 10 percent
of the product gas, is cooled from 1700 F to 200 F by 1300 psia steam genera-
tion in E-101, boiler feed water heating in E-102, and air cooling in E-103.
Part of the cooled gas is then processed in a hot potassium carbonate absorp-
tion/stripping system where 85 percent of the C02 is removed. The treated gas
from the CO absorber T-301, is combined with the remainder of the product gas
slip stream and recycled to the BCR gasification system as coal transport gas.
-177-
-------
Carbon dioxide taken overhead from the "hot pot" regenerator, T-302, is
compressed and used as makeup
Product fuel gas is delivered to battery limits at 1700 F and ^75 psia.
Low-pressure steam generated in the sulfur burner, B-UOl, supplies 70 percent
of the 50 psig steam required for the CO removal regenerator and the sour
water stripping system. The 1300 psia steam generated in the coal gasification
unit and gas purification system is sent to the steam generator portion of the
combined-cycle power system.
Utilities and Energy Balances
The utilities summaries for the four standard gasification/cleanup system
combinations are given in Tables 35 through 38. For each system, the summary
gives a section-by-section breakdown of the utilities requirement.
The energy balances around the four gasification/cleanup system combina-
tions are shown in Table 39- The basis for these balances were the higher
heating value at 77 F plus the sensible heat referred to 77 F with water in
the liquid state. Electric power input was calculated on the basis of
equivalent coal heating value assuming 35 percent overall efficiency on the
HHV. Generally, the power consumption does not appear as process stream
energy since pratically all of it is used in coal pretreatment and to provide
mechanical refrigeration.
Based on the total energy input to each system, the BCR/Consol system has
the highest .recovery of useful energy; 90-6 percent contained in the product
fuel gas and export steam. The other three systems have useful energy
recoveries in the range of 76-79 percent, as illustrated in Table 39- The
bulk of this differential is directly related to the energy' lost in process
cooling.
Investment Cost
The battery limits investment costs for the integrated gasification and
gas purification systems were estimated using currently available cost data.
Due to the lack of actual cost data on the high-temperature desulfurization
processes, investment costs were developed from conceptual plant designs for
the Consol and Sintered Iron Oxide processes. Other sections were estimated
from cost curves developed from published data.
The 197^ estimated investment costs for the four standard systems are
summarized in Table ho. Considering the variety of sources from which the
costs were developed, an accuracy of not better than ± 25 percent should be
-178-
-------
Table 35
SUMMARY OF BOM GASIFICATION/SELEXOL EESULFURIZATION
UTILITIES CONSUMPTION
Coal
Gasification
STEAM, LB/HR
@ 65 psia
@ 1315 psia 119790
COOLING WATER
GPM
POWER, KW 10500
BFW, LB HR 165355
STM COND., LB/HR
CHEMICALS
$/MY
Gas Sour Water Ammonia
Cooling Stripping Recovery ,
6U20
50710
7330
1286 716 12UU
(57130)
ko
Acid Gas
Removal
106200
37090
13^95
(106200)
70
Sulfur
Recovery
(68365)
u
75060
26
Total
170500
272U5
(163330)
136
-------
Table 36
SUMMARY OF BUREAU OF MINES/SINTERED IRON OXIDE
UTILITIES CONSUMPTION
Steam, Lb/Hr.
65 psia
1315 psia
Cooling Water
GPM
Power , KW
BFW, Lb/Hr.
STM Cond.
Lb/Hr.
Chemicals
$/Day
SULFUR
GASIFLER RECOVERY
646,590
27,775
42
10,500 17,650
218,435
(646,590)
6,186
DESULFURIZATION ' TOTAL
(138,400) • 508,190
27,775
42
8,3^0 36,490
1^1,100 359,535
(646,590)
6,186.
-180-
-------
Table 37
SUMMARY OF BCR GASIFICATION/SELEXOL
DESULFURIZATION UTILITIES CONSUMPTION
STEAM, Ib/hr
@ 65 psia
@ 1370
@ 1250 SPHT
COOLING WATER, gpm
£
H POWER, kw
BFW, Ib/hr
STM. COND., Ib/hr
PROCESS WATER
Ib/hr
Coal Heat Gas Acid Gas Trans. Gas Sour Water
Gasification Recovery Scrubbing Removal Compression Stripper
Sulfur Ammonia
Recovery Recovery
Total
396900
10000
21000
762^70
?1 OOOO
(U69270) 303500
(75^920)
62950
783 33650
U73960
(303500)
271^90 (96780) ^780
37750
5^60
3190 569 5 926
10^730
(271^90) (^2530)
13720
(717170-)
396900
78U10
60123
13^1160
(617520)
210000
CHEMICALS, $/day
100
26
30
156
-------
Table 38
SUMMARY OF BCR GASIFICATION/CONSOL DESULFURIZATION
UTILITIES CONSUMPTION
i
H1
OO
ro
i
STEAM, LB/HR
@ 65 PSIA
@ 1300
@ 1250 SPHT
COOLING WATER,
GPM
POWER, KW
BFW, LB/HR
STM. COND., LB/HR
PROCESS WATER
LB/HR
Coal Gas
Gasification Purification
(135530) (1^7819)
396900
10000
. 21000 277
137190 195530
5/HR
210000
Spent Dolomite C02
Treating Removal
58200
1120 390
155 ^863
(58200)
^
Sulfur
Recovery
(U8880)
6200
3931
^9370
Sour Water
Stripping Total
13680 23000
(2833^9)
396900
17710
29 30255
382090
(13680) (71880)
210000
CHEMICALS, $/DAY
1^00
-------
SYSTEM
Table 39
GASIFICATION/CLEAN-UP SYSTEMS ENERGY BALANCES
EOM/SELEXOL BOM/IRON OXIDE BCR/SELEXOL
ENERGY IN. MMBTU/HR.
CX>
UJ
I
Coal 8015.0
Gasifier Air 378.9
Sulfur Recovery Air 0.1
Boiler Feed Water 36.8
Make-Up Water
Electric Power (Coal HHV) 265.?
Process Steam, 65 Psia 50.3
1300 Psia 193.2
TOTAL 89UO.O
ENERGY OUT, MMBTU/HR.
Product Fuel Gas 6821.2
Export Steam
Sulfur 97.6
Ammonia 98-5
Stack Gas U9.8
BFW Blowdown 1.0
Process Cooling 1203.^
Process Heat Loss 17^-5
Steam Condensate 36.k
Waste Water 7.6
Ash & Dust 226.2
Power Consumption 93-0
Power Generation Loss 172.7
TOTAL . 8981.9
BALANCE % 100.k6
8015.0
338.0
0.5
55-0
355.8
577-0
31.5
9372.8
7265.0
82.8
53.5
l.U
105^.9
218.2
1UU.2
0.1
226.3
231.3
9^02.2
100.31
8610.0
388.3
0.6
173. U
U.8
586.3
15.6
57*+. 5
10353.5
7327.1
812.6
97.2
73.3
76.2
5.1
1117.7
98.1
137.7
32.0
205.2
381.1
10363.3
100.09
BCR/CONSOL
8610.0
388.3
0.1
58.5
5.2
295-0
26.1
57^.5
9957.7
8705.8
321.0
96.7
0.3
1.8
328.8
173.5
16.0
22. k
103.3
191.8
9961.k
100.03
-------
Table 39 (Cont'd)
' GASIFICATION/CLEAN-UP SYSTEM ENERGY BALANCES
SYSTEM
Total Energy In. MMBTU/HR.
Coal In. MMBTU/HR.
°lo Recovery:
Product Fuel Gas
Export Steam
Sulfur
Ammonia
Stack Gas
BFW Slowdown
Process Cooling
Process Heat Loss
Power Generation Loss
Steam Condensate
Waste Water
Ash & Dust
Power Consumption
TOTAL
BCR/SELEXOL
10353.5
8610.0
70.77
7.85
0.9U
0.71
0.7^
0-.05
10.80
0.95
3.68
1.33
0.31
1.98
BCR/CONSOL
9957.7
8610.0
87.1*3
3-22
0.97
0.02
3.30
1.71*
1.93
0.16
0.22
100.10
BOM/SELEXOL
891*0.0
8015.0 .
76.30
BOM/IRON OXIDE
9372.8
8015.0
77.51
O.I
100.^6
100.32
-------
Table Uo
SUMMARY OF INTEGRATED GASIFICATION/CLEANUP SYSTEM
INVESTMENT COSTS
(Millions of Dollars)
BASIS: 197^ Gulf Coast
System:
GASIFIER
CLEANUP
Gasification
Gas Cooling
Desulfurization
Sour Water Stripping
Ammonia Recovery
Sulfur Recovery
TOTAL, MM$
BOM
SELEXOL
>*
10
16
1+
7 •
2
81
BOM BCR
IRON OXIDE SELEXOL
1+2 56
16
lU 20
k
6
7 2
63 . 10U
BCR
CONSOL
56
--
11+
1
—
6
77
185-
-------
expected. Of the total investment costs, the coal gasification system
represents 50-70 percent. The available data on fixed-bed gasification
systems costs is very sparse since no commercial plants have yet been con-
structed in the U.S. Reliable data for second-generation gasifiers is
essentially nonexistent since these systems are only in developmental stages
where only preliminary cost estimates have been made on conceptual pilot plant
designs.
The only significant conclusions that can be drawn from the cost figures
is that, for a given gasification system, the low-temperature cleanup scheme
is more expensive, possibly by 30 percent. The costs for the desulfurization
sections appear to be comparable for both the low- and high-temperature
schemes, but the former scheme must also bear the additional costs associated
with gas cooling, sour water stripping, and ammonia recovery.
Comparison of Standard Integrated Systems
Comparison of first- and second-generation gasification/cleanup system
characteristics is given in Table hi. In terms of the percentage recovery of
input energy in the product fuel gas and export steam, the second-generation
BCR gasification system appears to be about 2.5 percent more efficient than the
first-generation Bureau of Mines stirred-bed gasifier. For the first-genera-
tion gasification system, high-temperature gas cleanup does not offer signifi-
cant advantage over low-temperature desulfurization. The performance of the
sintered iron oxide process could be improved somewhat by a more efficient
sulfur recovery system. The present scheme involves the use of large
quantities of steam to recover and concentrate the SO regenerator off-gas.
In addition, 7 percent of the fuel gas product is used to reduce the S0? to
elemental sulfur.
With second-generation gasification systems, high-temperature cleanup
shows a decided advantage, $0.6 percent energy recovery versus 78.6 percent
for the BCR/Selexol system. At least part of this differential is due to the
more stringent desulfurization duty designed into the Selexol system. If the
low-temperature systems were designed for a comparable 500 ppm total sulfur
in the treated fuel gas, instead of 100 ppm, the energy requirement for
stripping steam and refrigeration would be substantially lower. It should be
noted that the Consol high-temperature process does not have the capability
to reduce the sulfur content to less than 500 ppm under the operating
conditions imposed. In this respect, low-temperature systems are more
versatile.
-186-
-------
Table
COMPARISON OF GASIFIER/CLEANUP SYSTEM PERFORMANCE
GASIFIER SYSTEM
CLEANUP SYSTEM
BUREAU OF MINES
SELEXOL
IRON OXIDE
BCR
SELEXOL CONSOL
Coal:
STPD
MMBTU/HR.
W. Kentucky No. 9
QkQO
8015
Illinois No. 6
8610
Product Fuel Gas
MMSCFD
°F
PSIA
HHV TR
LHR BT
VOL. i
Ammonia STPD
Sulfur LTPD
Slag/Ash STPD
Waste Water, GPM
so,
NO!
Energy Recovery %
Product Fuel
Export Steam
'SCF
'SCF
N2
CO
co2
Hg
CH^
HjS-t^OS
HH-
H23
TAR
)
GPM
Coal
Coal
ay %
Lei Gas
iam
1003.2
265
355
160.2
11*9.1
51-97
23.87
5.02
15.98
3.11
0.01
0.03
0.01
--
100.00
133.2
259.1
1533.0
351
0.^08
0.5^5
76.30
0.00
915-9
1070
355
. 167.9
151*. 2
1*7.70
15. 3^
11.13
18.19
2.77
0.01
0.63
3.91
0.32
100.00
--
219.8
1533.0
--
0.872
l*.52l*
77.51
0.00
993.5
1000
1*70
159-2
1^7.5
53.18
21.23
6.1*9
1^.87
U. 21
0.01
__
0.01
--
100.00
99.2
258.1
730.8
897
0.567
0.306
70.77
7.85
1178.3
1700
^75
137.1
126. h
1*5-88
17.37
9.29
13.^7
3-56
0.05
O.U3
9-95
--
100.00
__
256.7
730.8
129
0.553
2.939
87.1*3
3.22
76.30
77.51
78.62
90.65
-187-
-------
SECTION -
INTEGRATION OF GASIFICATION, CLEANUP AND' POWER SYSTEMS
Section h contains a description of the integration of those gasification
and cleanup processes previously identified as most attractive into both con-
ventional steam systems (2^00 psi/1000 F/1000 F) and into an advanced power
generation system having a corresponding time of maturity. First-generation
systems are based on the use of the fixed-bed gasifier concept typified by the
Bureau of Mines stirred-bed gasifier. Compatible high-and low-temperature
cleanup systems were considered and the Selexol and sintered iron oxide pro-
cesses were selected as the most suitable representatives in their respective
categories. These two systems were combined with a first-generation COGAS
power system having a turbine inlet temperature of 2200 F and using conventional
cooling techniques.
Second-generation systems were focused on the use of a BCR-type entrained-
flow gasifier. Low-temperature cleanup systems were judged to be sufficiently
mature to preclude any significant developments in that technology; therefore,
the Selexol process was again chosen to represent low-temperature cleanup. For
the high-temperature process, a choice was made between the Consol half-
calcined dolomite and the AFC fully-calcined dolomite process. While there is
only a small difference in estimated overall system performance, the fully
calcined system tends to absorb C0? at temperatures below 1750 F causing a
rapid increase in the fuel requirement for subsequent calcination. Therefore,
the Consol process was selected as the representative high-temperature clean-
up system. Both the BCR/Selexol and BCR/Consol systems were combined with a
COGAS power cycle with a 2600 F turbine inlet temperature, ceramic vanes and
conventionally cooled blades.
The salient characteristics of the major components of each system as
they affect the overall system can be summarized as follows:
-188-
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First Generation
Gasifier: Fixed - bed type - Bureau of Mines
Low temperature (1000 F outlet)
Ammonia and tar in product gas
High cold gas efficiency
Low-Temperature Cleanup - Selexol
Includes ammonia removal
Necessitates tar removal by water quench of gas
Effective H_S removal
High-Temperature Cleanup - Sintered iron oxide
Effective HS removal below 1100 F
Regeneration produces SO
Difficult to produce elemental sulfur
No provision for ammonia removal
No water scrub for particulates
Power System - Combined Cycle
2200 F turbine inlet temperature
, Conventional engine cooling
Two-pressure steam bottoming cycle
Power System - Steam
2^00 psi throttle pressure
1000 F throttle temperature with reheat to 1000 F
Second Generation
Gasifier: Entrained Flow - BCR type
High temperature - two stage (1800 F outlet)
Ammonia and COS in product gas
Needs coal transport gas stream
High percentage of sensible heat in gas
-189-
-------
Low-Temperature Cleanup - Selexol
Includes ammonia removal
Requires fuel gas regeneration to recover fuel gas sensible heat
High solvent flow for COS removal
Effective sulfur removal
High-Temperature Cleanup - Consol
Potentially good COS removal
Limited sulfur removal ability
Low utility requirements
No provision for ammonia removal
No water scrub for particulates
Power System - Combined cycle
2600 F turbine inlet temperature
Ceramic stator vanes
Conventional turbine blade cooling
Two-pressure steam bottoming cycle
Power System - Steam
2^00 psi throttle pressure
1000 F throttle temperature with reheat to 1000 F
STEAM SYSTEMS
The performance of four different gasifier/cleanup steam power plant sys-
tems was evaluated for the configuration using a fuel gas let-down turbine as
previously shown in Figure 15(b). The results of these evaluations are sum-
marized in Table k2. Performance of-the BG-R- gasifier-based systems can be
compared directly to the coal-fired power plant characteristics given previous-
ly in Table 7 since they both are based on firing an Illinois Seam #6 coal.
The Bureau of Mines gasifier uses Western Kentucky #9 coal as the feed. A
boiler using that coal would have an efficiency of 86.9 percent as opposed to
the 88.5 percent value used in the reference plant. The systems using the
Bureau of Mines gasifier therefore should be compared to a direct-fired plant
having an efficiency of 3^-5 percent with stack gas cleanup while the BCR-based
systems should be compared to 35-1 percent.
It is apparent from Table U2 that with the exception of the BCR-Consol
system, the performance of these systems is quite unattractive when compared
to the reference coal-fired plant with stack gas cleanup. The reasons for this
-190-
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Table U2
Gasified Coal - Steam Cycle Performance Summary
Gasifier-Cleanup Combination
Fuel System
Coal Feed Rate - Ib/hr
Clean Fuel Gas Temp. - F
Fuel Gas HHV - Btu/scf
Boiler & Steam Cycle
Boiler Efficiency
Net Steam Cycle Heat Rate - Btu/kwh
Net Steam Cycle Output - MW
Let-Down Turbine
Net Output - MW
Auxiliary Power - MW
Plant Performance
Net Output Power - MW
Efficiency (HHV)
BuMines/
Selexol
700,000
265
160.3
.81*8
7836
6M5.1
0
50.0
596.1
.25^
BuMines/
Iron Oxide
700,000
1,070
165.9
.852
7836
620.0
1*3.3
58.0
605.3
.258
BCR/
Selexol
700,000
1,000
159.3
.855
7836
736.9
27.7
87.1
677.5
.269
BCR/
Consol
700,000
1,700
135.8
.867
7836
753.^
166.6
58.3
861.7
.3^2
-------
become quite clear when each of the major parts of the system are examined in
detail considering both thermal and auxiliary power requirements. This was
done in the preparation of the performance estimates and is presented in
Table ^3. In essence, this table summarizes the performance calculations and
an explanation of the table requires a description of the general approach as
well as the procedures used in the calculation.
Performance Calculations
To mate the gasifier/cleanup systems with a steam cycle is outwardly a
simple matter requiring only a source of compressed air for the gasifier and
utilization of the product gas in a steam boiler. These two factors can be
considered separately.
The product gas utilization involves not only raising steam for power
generation, but also for use in the cleanup process and the gasifier. To keep
track of the various uses of this system heat requires an accounting system.
The approach that has been chosen is to relate all process heat requirements
to an equivalent fuel consumption in Btu/hour by dividing by boiler efficiency.
While simple, this approach is also quite'accurate, since by virtue of air
preheaters, a steam system is able to use all the available fuel energy in the
steam power cycle and the heat used to produce low-pressure steam for process
use subtracts directly from that available to the steam cycle.
Boiler efficiency is calculated using an accepted industry method. (33)
These efficiencies are generally two to three points lower than for direct
coal firing due to the increased hydrogen content of the fuel gas caused by
steam injection. Higher fuel gas temperature (sensible heat) increases the
total heat content of the gas while the composition and, therefore, moisture
losses remain constant. This improves boiler efficiency, partially offsetting
the loss due to hydrogen content. The sensible heat content of the fuel is
referenced to 59 ^. It should be noted that even with 10 percent excess air
for gas as opposed to 20 percent for coal the boiler efficiency for the fuel
gas is lower than for coal. In proceeding through the system, the heat re-
quirements of each system component (or the heat exported) are charged against
(or credited to) 'that component as the equivalent fuel energy consumed at the
appropriate boiler efficiency. The net fuel available is then used to cal-
culate heat available to the steam cycle and net steam cycle output, using the
steam cycle performance presented previously for the reference coal fired
steam plant (Net Heat Rate = 7,836).
-192-
-------
Table 1(3
GASIFIED COAL STEAM CYCLE-COMPONENT PERFORMANCE CHARACTERISTICS
ENERGY Bu Mines/Selexol 3u Mines/Iron Oxide BCR/Selexol BCR/Consol
(MM Btu/hr., unless otherwise noted)
Coal Energy Input (HHV) 8,015 8,015
Boiler
Clean Fuel Gas Flow, Ib/hr
Gas Temperature, F
Chemical Energy in Gas (HHV)
Sensible Heat in Gas (Above 59 F)
Total Energy in Gas (HHV)
Boiler Efficiency
Gasifier
Clean Fuel Gas Energy Out (at boiler supply temperature)
Heat Available From Gasifier and Gas Streams
Useful
Rejected
Process Heat Requirements
Added to Fuel Gas Stream
Input to Steam
Gasifier Air Preheat
Net Process Heat Used (Required less useful heat available)
Net Fuel Energy for Process Heat (heat/boiler efficiency)
Net Fuel Energy Out (fuel out less fuel energy for process
Effective Gasifier Efficiency
Cleanup System
Net Process Heat Required by Sulfur Recovery
Net Fuel Energy For Sulfur Process Heat
Fuel Energy Used Directly by Glaus Plant
Total Fuel Energy for Sulfur Removal
Net Process Heat Required for NH^ Removal
Net Fuel Energy for NH3 Process Heat
Total Fuel Energy Required for Cleanup
Net Fuel Energy Out of Cleanup
Steam Cycle
Heat Available From Boiler
Overall Coal to Steam Efficiency
Net Steam Cycle Output, kw
Auxiliaries
Gasifier Auxiliary Pover , kw
Auxiliary Power for Gas Cooling, kw
Auxiliary Power for NH Removal, kw
Sulfur Auxiliaries, kw ,
Subtotal Auxiliary Power Gasifier & Cleanup, kw
Steam Plant Auxiliary Power, kw (3.31? of net steam cycle)
Performance Without Let Down Turbine
Net Plant Output - kw
Overall Plant Efficiency, %
Performance With Let Down Turbine
Net Output From Let Down Turbine - kw
Net Change in Steam Cycle Output - kw
Net Plant Output - kw
Overall Plant Efficiency, %
2,671,938
265
6,670
159
6,829
.81t8
6,882
672
(180)57l(*
190
307
386
211
21(9
heat) 6,633
.828
1(8
57
53
110
382
1(50
560
6,073
5,150
.61(3
• 657,223
(97) 10,500**
(12) 1,286
(20) 2,127
(133) lit. 3"t5
28,258
21, 751)
607,211
25.9
0
-11,158
596,053
25. >t
2,1(97,630
1,070'
6,311
778
7,089
.852
7,623
2lU
7lt
26lt
3UU
1(68
51(9
T.OTH
.883
1(92
577
53>t
1,111
1,111
5,963
5,080
.631*
61(8,290
(97) 10,500
(239) 26,000
36,500
21,1(58
590,332
25.1
1(3,332
-28,308
605,356
25.8
2,689,108
1,000
6.56H
775
7,339
.855
7,1(72
(75)1,973*
(158) 290*
777
555
395
(-21(6)
(-288)
7,760
.901
221t
262
133
395
3>(1
399
T9>*
6,966
5,956
.692
760,082
(278) 2l(,l(i8
(22) 2.U02
(322) 35.090
6l,910
25,159
673,013
26.7
27,736
-23,183
677,566
26.9
3,170,360
1,700
6,636
1,729
8,365
.867
8,365
31(9
65
0
555
395
601
693
7,672
.891
106
122
122
122
7,550
6,5"»6
.760
835,375
(250) 2l(,278
(58) 6.381
30,659
27,651
777,065
30.1
166,583
-81,992.
861,656
3lt.2
•Numbers in parenthesis indicate Latent Heat from Condensation of Gasifier Steam
••Numbers in parentheses indicate Fuel Energy in MM Btu Needed to Produce Equivalent Power
-193-
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Gasifier Performance - From the process data for each system, the total energy
in the clean product gas can be obtained. The energy in the gas used in sul-
fur recovery must be added to this to give the actual clean fuel gas energy
leaving the gasifier. In addition, the gasifier must be credited with heat
extracted from the gasifier and fuel stream that is used elsewhere in the
system and also must be debited for heat that is input to the gasifier and
fuel stream. The quantity of heat rejected to the cooling tower or to air
coolers is indicated. It represents a thermal loss similar to that due to
radiation or slag losses from the gasifier. In all cases, gasifier steam and
air preheat requirements represent a significant fraction of the input coal
energy. The net heat required by the gasifier is then divided by boiler ef-
ficiency to give an effective parasitic fuel loss. This, subtracted from the
actual clean fuel gas energy leaving the gasifier, gives the net fuel energy
out of the gasifier.
In all cases, it can be seen that the sum of actual clean fuel gas energy
leaving the gasifier and rejected from the gasifier and fuel gas stream do not
equal the sum of the coal energy input and net heat input. The difference is
made up of heat losses from the equipment, slag, latent heat of steam in the
product gas, heat rejected from the product gas, and a small loss from the
sulfur which is not credited to the fuel gas.
The relatively poor performance of the Bu Mines/Selexol combination is
due to the need to cool the gas prior to cleanup. This cooling must be done
with a water quench, since the gas contains tars that would foul heat exchange
surfaces. More of the heat could be removed by condensing the water vapor in
the gas stream following the quench and tar removal step; however, that is at
a low temperature and it would be difficult to justify economically.
In the Bu/Mines Selexol system all of the heat extracted from the gas
stream involves condensation of water vapor. Because all of the steam in-
jected into the gasifier is not used in the reaction, about 72 percent of the
steam input to the Bu Mines gasifier (vs. 63 percent for the BCR gasifier) is
present in the gasifier exit stream. The resultant dew point is on the order
of 250 to 280 F. As the fuel gas stream temperature is dropped below this
point, some of the latent heat of that steam is either utilized or rejected.
To indicate the extent to which that happens, the values in Table k3 shown in
parenthesis represent that part of the heat that is removed by condensation of
the gasifier steam. While it may be argued that the effects of cooling the
gasifier effluent for use with a low-temperature cleanup should not be charged
to the gasifier, it is interesting to note that in the BCR/Selexol System the
need for cooling is turned into an advantage by making use of part of the
latent heat of the gasifier steam. Because of this'ability to take advantage
of the latent heat, and also because the process heat requirements for the
-------
BCR/Consol system must be met by burning fuel at boiler efficiency, the net
fuel out of the gasifier is slightly higher for the BCR/Selexol System than
for the BCR/Consol. Note that while 290 MM Btu/hr is rejected to cooling
water or to air from the BCR/Selexol System, it does not represent a major
penalty since a large part of that is latent heat of water vapor that would
be otherwise lost through the stack, were it not condensed here.
It is possible that further use could be found for the low-temperature
heat resulting from the water quench following the Bu Mines gasifier. While
the effective gasifier would be improved, the ability to use other low-tem-
perature heat would certainly be impaired and a significant change in overall
performance would not be expected. As all of these systems show relatively
poor performance potential, such improvements were not investigated. The
fuel energy equivalent of auxiliary power can be obtained by dividing by
boiler and steam cycle efficiency. These numbers are shown in parenthesis in
the auxiliary power tabulation. When charged to the appropriate part of the
system, this gives a measure of overall performance. For the gasifiers, aux-
iliary power requirements reduce the Bu Mines performance by 1.2 points and
the BCR by 0.37 points. Coal preparation is the largest requirement for
auxiliary power.
Cleanup System Performance - Sulfur and ammonia removal have been treated
separately to provide a view of the relative penalty associated with the
removal of each and to permit a better comparison of high- and low-temperature
systems, since the high-temperature systems have no provision for ammonia
removal. In both low-temperature systems, ammonia accounts for a significant
part of the process heat used. When sulfur removal is considered, the Bu
Mines/Selexol System compares quite favorably to the BCR/Consol performance,
requiring less process heat (110 vs. 122 MM Btu/hr), but more power (1^,350
vs. 6,38l kw). The BCR/Selexol System performance is degraded due to the
relatively high concentrations of COS that must be removed. The Selexol
solvent has a lower capacity for COS and the solvent flow rate must be in-
creased to provide high COS removal efficiency. Thus, increases are required
in power for pumping and refrigeration and in heat for regeneration. In
addition, the high solvent flow rates result in a larger amount of CO absorp-
tion which in turn, lowers the concentration of HS in the feed to the Glaus
plant. This causes an increase in complexity and fuel requirements. A pos-
sible alternative would be to design the cleanup system for HpS removal only.
This would result in about 800 ppm of sulfur compounds in the product gas and
would reduce utility requirements to a level comparable to the Bu Mines/Selexol
system.
The high-temperature sintered iron oxide cleanup process shows up poorly
in this comparison primarily because of the energy needed to obtain elemental
-195-
-------
sulfur from the SO that is produced upon regeneration of the iron oxide sor-
bent. The production of SO results in a net export of heat from the sulfur
absorption process and virtually all of the utility requirements for this
cleanup system are associated with the subsequent recovery of elemental sulfur.
The feed gas to the recovery unit is dilute and requires concentration followed
by reduction of a portion of the SO to form the HpS needed in a Glaus reaction
to form elemental sulfur. The system also has a relatively high auxiliary
power requirement about 2/3 of the power required by the sulfur recovery sec-
tion. Another means of disposing of the SO desorbed from the iron oxide would
be necessary to make this system attractive.
The heat available to the steam cycle is the net fuel energy out of the
cleanup process multiplied by the boiler efficiency. The resultant overall
efficiency from input coal energy to boiler output tends to favor the two BCR
systems. However, the high auxiliary power requirements for the BCR/Selexol
combination negate that advantage and the net overall efficiency is about equal
for all except the BCR/Gonsol system which benefits from a high-temperature
cleanup system with very low auxiliary power and process heat requirements.
Steam Cycle Performance - The steam cycle output power neglects the power that
is required to compress the supply air to the gasifier and the power that is
available from the expansipn of the product gas from gasifier pressure to
burner inlet. Therefore, a compressor and turbine must be superimposed on the
system to determine overall performance.
The approach used in estimating performance with a let down turbine was
to approximate turbine power available in an expansion from 500 psia. It was
assumed to be directly proportional to product gas flow rate and absolute fuel
gas.temperature. Turbine output power is used to drive the gasifier air supply
compressor with the excess power available for electric generation. Required
compressor prower was taken to be directly proportional to gasifier inlet air
flow. The net output from the letdown turbine is given in Table ^3 and in-
cludes an allowance for plant auxiliary power. Where turbine power was less
than that needed to drive the compressor, as in the Bu Mines/Selexol system,
it was assumed that the fuel temperature could be increased by regeneration
or by partial combustion to a point where the turbo compressor was self-sus-
taining. For the Bu Mines/Selexol system, that temperature is approximately
?60 F.
The effect of the letdown turbine on steam cycle output was estimated by
assuming that 83 percent of the compressor work was available as thermal
energy and could be used to raise process steam as well as preheat the air to
the gasifier. The remainder is wasted in aftercooling between stages to mini-
mize compressor power and limit the exit temperature to a level compatible
-196-
-------
with common blade materials. All of the turbine power is subtracted from the
heat available to the steam cycle. The net change in plant output was then
calculated based on steam cycle heat rate and correcting for the appropriate
auxiliaries.
The effect of product gas temperature is emphasized by the addition of
the letdown turbine. The 1700 F gas temperature out of the Consol cleanup
system results in an increase of approximately 10 percent in net plant output.
To evaluate the effect of the approximations used in the foregoing analy-
sis and to provide a reference case for a gasified coal - steam cycle system,
the BCR/Consol/steam system shown schematically in Figure 5^- was analyzed
using the SOAPP preprocessing system. (Appendix A). This is basically a
system bookkeeping tool that allows the user to assemble a power system using
a number of standard modules and provides the computation logic necessary to
determine performance of the system. For the BCR/Consol/steam cycle system,
the computer configuration was that shown in Figure 55. The main stream in-
cludes provision for process steam generation iti a low-pressure boiler as well
as a deaerator for boilers external to the main boiler and steam cycle. The
main boiler and steam cycle combines a combustor operating at 10 percent excess
air with a steam boiler and steam cycle having the performance characteristics
previously shown in Figure 13. Stack temperature is- a minimum of 300 F.
The gasification stream consists of three stages of compression with
intercooling and heat recovery where possible. A side stream contains the CO
separation module needed to provide C0? for dolomite regeneration with the
remainder of the gas used to transport coal into the gasifier.
Using this model, the performance characteristics presented in Table hh
were calculated. They differ from the results shown in Table ^3 by about 0.5
points in overall efficiency. This difference is due to the need for cooling
gas in the tubine. At the 1700 F turbine inlet temperature, cpoling would be
required to keep blade temperatures to a level compatible with long life, e.g.,
1500 F. Air cannot be bled from the compressor to be used since there would
be the danger of local combustion at the coolant discharges in the blades.
Therefore, it was necessary to devise a cycle in which a part of the turbine
inlet gas (8.7 percent) is cooled and used instead of compressor bleed air for
engine cooling.
COMBINED-CYCLE SYSTEMS
A summary of the calculated performance characteristics for each of the
four systems is given in Table U5. The performance of the second-generation
gasification and cleanup systems mated with a first-generation power system is
-197-
-------
BCR/CONSOL/STEAM SYSTEM
vo
oo
DO
O
I
KJ
INLET AIR
HP STM.
TRANSPORT GAS
CLEAN GAS TO BURNER
TO CO2 ABSORBER
CYCLONE
PARTICULATE REMOVAL
COM-
PRESSOR
AIR TO GASIFIER
DOLOMITE
ABSORBER
DOMOLITE
REGEN.
FUEL GAS
TO BOILER
BCR
GASIFIER
STEAM TO GASIFIER
CHAR
DOMOLITE
QUENCH MAKE-UPH
REHEAT FLOW
STEAM CYCLE
WASTE
WATER
WATER TO SW
BOILER
NK
x\
"•
^1S
x\
GEIV
A
FEED
COAL &
GAS
WATER-
CaS
•CONVERTER-
. DOLOMITE
SLUDGE
ITRANSPORT
GAS
| COALI
ROCESSING
I SYSTEM I
PROCESS COOLING WATER
GAS
CO2 &
STEAM
STACK
GAS
SULFUR
COOLING WATER
AIR
CO2
SEPARATION
CONDENSATE TO S.W.S.
Tl
P
s
-------
FIG. 55
GASIFIED COAL-STEAM SYSTEM
N11-179-1
-199-
-------
Table
BCR/Consol/Steam Cycle Performance
Fuel System
Coal Feed Rate - Ib/hr 700,000
Gasifier Exit Temp - F 1,800
Cleanup System Exit Temp - F 1,700
Fuel Gas HHV - Btu/scf 135.8
Boiler & Steam Cycle
Fuel - Air Ratio .7195
Air Preheat Temp - F 376
Throttle Press - psia 2,14-00
Throttle & Reheat Temp - F 1,000
Net Heat Rate - Btu/kwh 7,836
Net Stm Cycle Output - Mw 760.5
Let Down Turbine
Net Power Output - Mw 11+3.2
Auxiliary Power
Gasifier & Cleanup - Mw 27.8
Steam Cycle - Mw 25.2
Gas Turbine - Mw .8
Plant Performance
Net Power Output - Mw 81+9.9
Efficiency (HHV) 33-7
-200-
-------
Table 1*5
INTEGRATED SYSTEMS PERFORMANCE SUMMARY
Gas Turbine
Turbine Inlet Temperature - F
Compressor Pressure Ratio
Exhaust Temperature - F
Output Power - MV
Steam Cycle
Steam Temperature - F
Steam Pressure - psia
Condenser Pressure In. Hg. Abs.
iJet Steam Cycle Output - Mw
Net Steam Cycle Efficiency
Gasifier and Cleanup System
Coal Feed 'Rate - Ib/hr.
Air - Coal Ratio
Steam - Coal Ratio
Air Preheat Temperature - F
Steam Temperature - F
Steam Pressure - psia
Gasifier Exit Temperature - F
Cleanup System Exit Temperature - F
Fuel Gas higher heating Value Btu/SCF
Integrated Station
Gross Power - Mw
Boost Compressor Power - Mw
Gasifier.& Cleanup Aux. Power - Mw
Plant Auxiliaries - Mw
Uet Plant Output - Mw
Net Plant Efficiency (KHV-Coal)
First
BOM/
Selexol
2,200
16
916
595. ^
816
1,250
1*.0
223.8
.280
700,000
3,013
.1*05
800
58U
1,250
1,000 '
265
160.3
819.2
1*3.»*
28.2
10.6
737. 0
.311*
Generation
BOM/
Iron Oxide
2,200
16
913
626.2
813
1,250
i*.o
208.1
.292
700,000
2.688
.31*9
800
581*
1,250
1,000
1,070
165.9
831*. 3
36.1
36.5
10.2
751.5
.320
Second
BCR/
Selexol
2,600
21*
1,107
726.6
1,000
1,250
i*.o
293.3
.307
700,000
3.088
.567
800
1,000
1,250
1,800
1,000
159.3
1,019.9
1*0.1
58.7
13.6
907.5
.360
Generation
BCR/
Consol
2,600
21*
1,115
857-6
1,000
1,250
lt.0
296.6
.307
700,000
• 3.088
.567
800
1,000
1,250
1,800
1,700
135-8
1,151*. 2
1*0.2
27.6
11*. 5
1071.9
.1*25
Second
Generation
Gasification
First
Generation
BCR/
Selexol
2,200
16
913
61*2.3
813
1,250
i*.o
273.5
.282
700,000
3.088
.567
800
913
1,250
1,800
1,000
159.3
915.8
1*0.1
58.7
12.5
Sol*. 5
.319
Power System
BCR/
Consol
2,200
16
920
757.6
620
1,250
i*.o
271.1*
.279
700,000
3.088
.567
800
920
1,250
1,800
1,700
135.8
1,029
1*0.2
27.6
13.1
91*8.1
.376
-201-
-------
also given to permit a gross comparison of the gasification and cleanup
systems. Looking at overall plant efficiency, it is clear that there are two
significant factors in performance improvement: turbine inlet temperature and
a high-temperature cleanup system. A closer look will show that the benefits
of the high-temperature system are due to a highly efficient process (lower
requirements for process steam and auxiliary power) as well as the ability
to operate at high-temperature. Unfortunately, the performance gain with the
high-temperature system is at the expense of ammonia and particulate removal
capability.
System Descriptions
The general arrangements for each of the four integrated systems are pre-
sented here. Performance characteristics and detailed utility requirements
are given in the following sections.
Bureau of Mines/Selexol - This system is shown schematically in Figure 56.
The gasifier is supplied with air and steam from the COGAS system. Compressor
bleed air is used to generate process steam prior to entering a boost compres-
sor. Saturated steam is extracted from the high-pressure section of the waste-
heat boiler. Because of the tar content, the gasifier exit stream is quenched
and the tars separated and returned to the gasifier. The gas leaving the water
quench is saturated with water vapor at 338 F and as a result, little regenerar
tion of the fuel gas is possible. However, the latent heat of the water vapor
is quite significant and it is used in the reboiler of the sour water stripper.
After regeneration with the product gas, the remaining heat is rejected to
ambient to lower the temperature prior to a water scrub and Selexol cleanup
tower. Gondensate from the Selexol stripper and the gas cooling process along
with water from the scrubber must be processed in the sour water stripper. In
the ammonia recovery process, H S absorbed in the water is recovered and sent
to the Glaus plant. The requirements for each of these processes for heat,
fuel gas or power have been previously discussed.
After regeneration, the clean fuel gas is fed to the COGAS system which
operates in a conventional manner with the exception that steam is extracted
from the high-pressure section of the boiler for subsequent use in the gasi-
fier and the ammonia recovery unit. Also, steam raised in the gasifier jacket
is used to supplement the waste-heat boiler. Because the fuel gas stream pro-
vides heat directly to the sour water stripper, the quantity of low pressure
steam required is relatively small and the steam raised in the gasifier bleed
air stream is sufficient to satisfy this need.
-202-
-------
BUMINES/ SELEXOL SYSTEM
• CLEAN GAS TO BURNER •
CLEAN GAS TO SULFUR RECOVERY
o
uo
i
STACK
GAS
-------
Bureau of Mines/Sintered Iron Oxide - Because the tar in the product gas does
not affect operation of the cleanup system, the water quench can be eliminated
with this high-temperature cleanup system. The flow schematic, shown in
Figure 57> is quite simple due to the lack of regeneration, and the elimina-
tion of the water scrub and tar recycle. Gasifier performance is enhanced by
not recycling the tar since it permits operation at lower air- and steam-to-
fuel ratios. The net cleanup system reactions are exothermic both during
absorption and desorption adding sensible energy to the fuel and enabling the
recovery of heat in the form of process steam from the regenerator off-gas.
However, the sulfur comes off the iron oxide as SO and in the sulfur recovery
system must be concentrated and a portion reduced fo ILS so that it can be
used to form elemental sulfur in a Glaus unit. This process uses a consider-
able amount of fuel and process steam. As a result, even with the steam
raised in the process, a significant amount of low-pressure steam must be
taken from the waste heat boiler to satisfy these requirements. The lower
gasifier steam requirement means that nearly all the heat needed to raise the
gasifier steam can be taken from the gasifier jacket cooling.
The system suffers from the inability to remove ammonia from the gas and
the uncertainties associated with particulate removal. Also, the tars will
contain some nitrogen and sulfur compounds which will not be touched by the
cleanup system.
BCR/Selexol System - With this system, shown in Figure 58, it is possible to
both regenerate and extract useful heat from the gasifier exit stream. The
high gas turbine exhaust temperature resulting from the 2600 F turbine inlet
temperature allows sufficient superheat in the steam to satisfy the gasifier
heat balance without using the transport gas regenerator that has been in-
cluded in the computer model and used in the parametric analyses. Ammonia
and H^S removal are quite similar to the Bu Mines system, but the Selexol
System is complicated by a high percentage of COS in the gas to be processed.
Because COS is about one-third as soluble as is the ILS, its presence in gas
streams requiring sulfur limitations lower than that, accomplished by a 30 to
k-0 percent COS removal, require that the system be designed specifically for
COS. The utilities for this system reflect this situation. A significant
amount of heat is recovered from the gasifier jacket and the high temperature
gasifier exit stream. This heat is used to raise high pressure steam which
supplements the high pressure section of the waste heat boiler by adding about
50 percent to the steam raised there. After the steam for the gasifier is ex-
tracted, the net increase in steam available for the turbine is about 25 per-
cent. All of the low-pressure process steam requirements are met by steam
raised in the boilers located upstream of the boost compressor and downstream
of the fuel gas regenerator.
-20U-
-------
BUMINES/SINTERED IRON OXIDE SYSTEM
CLEAN GAS TO BUHNER
IV
o
VJ1
I
p
CJ1
-------
BCR/SELEXOL SYSTEM
I
ro
o
CLEAN GAS TO BURNER
INLET
AIR
COMPRESSOR
t
"• BURNER
}
TURBINE
^•s^
-'
<
V
POWER
TURBINE
EL
G
GAS TURBINE
/ L.P. TURB. \
P
01
00
-------
BCR/Consol System - In this system, shown in Figure 59? "the gas steam
arrangement is very simple and the omission of the aqueous cleanup system and
its associated stripping unit significantly reduces process steam requirements.
As a result, it is possible to raise high-pressure steam in boilers located
upstream of the boost compressor and in the transport gas stream prior to the
CO removal unit. The low-pressure process steam needed to supplement that
raised within the cleanup and sulfur recovery system can be supplied by a
boiler located between the high-pressure boiler and boost compressor. The
high-pressure steam is used to supplement the waste heat boiler and balances
the amount of superheated steam extracted for injection into the gasifier. As
in the case of the BCR/Selexol system, steam superheat is sufficient to sat-
isfy the gasifier heat balance so that the transport gas regenerator was not
needed.
Performance
The SOAPP computer routine was used to calculate the performance charac-
teristics of the four selected systems. The configuration shown in Figure 30
was used for all but the BCR/Gonsol system which was represented by the con-
figuration of Figure 31. Using the SOAPP routine with the final system con-
figurations, it was possible to re-evaluate some of the parametric integration
work discussed in an earlier section. It was found that the basic results of
that work were valid. As an example, the effects of pressure ratio and tur-
bine inlet temperature on the BCR/Selexol system are presented in Figure 60
and show the same characteristics as those previously presented, thereby
assuring that the selection of operating pressure was valid. It should be
noted that these estimates were limited to the effects of changes in the COGAS
system and associated compressor power variations because of the fixed input-
output model used for the gasifier and cleanup systems.
The COGAS system performance as a function of the gas turbine pressure
ratio for both first- and second-generation turbine systems is shown in Fig-
ures 6l and 62. To provide consistency, both figures include the performance
of distillate fired systems which can be compared to those using the gas
generated in either a BCR/Selexol or BCR/Consol gasifier-cleanup system. From
these figures it is clear that the key to improved system performance is in
the gasifier/cleanup performance. Therefore, it is necessary to understand
the effect of individual component performance in terms of its net effect on
the overall system. This relationship is quite complex due to the interdepen-
dency of the various parts of the COGAS system. An accounting system must
take into consideration the difference between fuel energy and thermal energy
or steam. Whereas the fuel can be used at combined cycle efficiency (about
50 percent), the thermal energy or steam can be used only at ^team cycle effi-
ciency (about 30 percent). Therefore, to make a proper accounting of each
-207-
-------
BCR/CONOCO SYSTEM
CR / ^7 \
I / TURB. \
p,
c/
COAL
SYSTEM
WASTE 1 1 QUENCH DOLO
-* — 1 r**~ WATER MAK
WATER T
COALS
GAS
T
H2
CaS
CONVERTER
WATER to- to- l
^_
DOLOMITE
SLUDGE
•»- WATER TO SWS
|—»- L.P. STEAM
SULFUR
RECOVERY
CO2&
STEAM
STACK
GAS
* COOLING WATE
C02
SEPARATION
CONDENSATE
TO S.W.S.
PROCESS COOLING
WATER
P
en
to
-------
o
o>
(O
BCR GASIFIER-SELEXOL CLEANUP PERFORMANCE CHARACTERISTICS
CERAMIC VANES-CONVENTIONAL BLADE COOLING
I
ro
o
40
38
o
36
UJ
o 34
O 32
<
to
30
100
PRESSURE RATIO
2600
2400
2200
TURBINE INLET TEMPERATURE, F
I
150 200 250 300
NET POWER PER UNIT AIRFLOW kw/lb/sec
350
-------
i
ro
H
o
z
o
00
PERFORMANCE ESTIMATES FOR SECOND-GENERATION TURBINE SYSTEMS
TURBINE INLET TEMPERATURE = 2600F
CERAMIC VANES, CONVENTIONAL BLADES
45
EFFICIENCY -
%(HHV) 40
35
30
100
COGAS/DISTILLATE
VHH
COGAS/BCR/CONSOL (T?FUEL~85%)
24
6
SIMPLE CYCLE/DISTILLATE
40
PR
28
*24
-20
•16
12
20
1 6
COGAS/BCR/SELEXOL (T7FUEL~74%)
I
12
150
200 250
SPECIFIC POWER - KW/LB/SEC
300
350
o>
-------
FIG. 62
PERFORMANCE ESTIMATES FOR FIRST GENERATION TURBINE SYSTEMS
TURBINE INLET TEMPERATURE'= 2200F
'.ONVENTIONAL COOLING
45
> 40
I
X
>
o
LL
Li-
Li' 35
30
50
20 16
24
12
28
32
COGAS/DISTILLATE
PR = 40
SIMPLE CYCLE/
DISTILLATE
28 24
COGAS/
BCR/SELEXOL
16
I
100 150 200
SPECIFIC POWER - KW/LB/SEC
250
-211-
R03-222-1
-------
part of the system it is necessary to separate gas turbine and steam cycle
performance characteristics. This has been done for each major system com-
ponent and the thermal energy gained or lost in each component has been
charged at the effective value of fuel energy. Where a component consumes
fuel it is charged with the full energy content of that fuel. Process steam
or thermal energy consumption is charged only at a rate corresponding to the
loss in steam cycle generating capacity. The multiplier used to convert a
thermal loss to a fuel energy loss is the ratio of steam to combined cycle
efficiency. After all the parasitic losses are subtracted, the output of the
gasifier and the cleanup systems is an effective fuel energy which is all used
at combined cycle efficiency. In the following paragraphs the accounting
system is first described and then applied to the individual systems.
To give a simple mathematical model of the combined-cycle ststem, it is
necessary to configure the system as shown in Figure 63. The boost compressor
is completely divorced from the gas turbine to keep track of the heat of
compression. Because of the general use of higher heating value when dealing
with coal, the combined- cycle equations must be altered accordingly. The
equations used and their interpretations are given in Appendix E.
For each of the selected systems, the performance was calculated and the
data are presented in Table U6. Gas turbine efficiencies and combined- cycle
efficiencies are both higher than the corresponding distillate efficiency.
This is due to the high fuel flow rate and the fact that gas turbine output
includes the compressor work needed to supply the gasifier air. This will be
later subtracted as an auxiliary. The item of interest is the ratio of steam
to combined cycle efficiency which is a measure of the effective fuel loss per
Btu of process heat required. This is used to relate the value of thermal
energy used in process heating to fuel energy, e.g.
Thermal Energy - = Fuel Energy
This factor has been used to show the effect of the combined cycle energy
utilization on each of the four systems in Table k-7. Starting with the total
energy in the clean fuel gas delivered to the burner, the energy consumed in
the cleanup system is added to give a value for energy in the fuel gas from
the gasifier but at burner delivery temperature. (For high- temperature clean-
up systems this is essentially gasified outlet temperature.) Next the net heat
comsumed (or exported) by the gasifier is calculated. This times the ratio of
STM/T] gives the effective energy removed from (or added to) the fuel stream.
The heat rejected from the gas stream is tabulated but does not enter into the
calculation. It gives an indication of the effect of the requirement to cool
-212-
-------
COMBINED CYCLE REPRESENTATION
STACK
I
I
(V)
r
AIR
GASIFIER
CLEANUP
WASTE
BOILER
STEAM
CYCLE
GEN.
I
AIR-
BURNER
EXHAUST
GEN.
D
O
00
I
o>
CO
-------
Table 46
COMBINED-CYCLE ENERGY UTILIZATION CHARACTERISTICS
Product Gas
Temperature - F
Chemical Energy -
MM Btu/hr (HHV)
Chemical Energy -
MM Btu/hr (LHV)
Sensible Energy (Above 59F)
MM Btu/hr
Total Energy to Burner -
MM Btu/hr (HHV)
Total Energy to Burner -
MM Btu/hr (LHV)
Heating Value Ratio (HVR)
Combined Cycle
Gas Turbine Output - SHP
Gas Turbine Efficiency
(71GT) (HHV)
Gas Turbine Outlet Tem-
perature - F
Gas Temperature Out of
Boiler - F
Boiler Heat Removal Efficiency
(Above 59 F) (T1BLR)
Net Steam Turbine Mech.
Output - SHP
Net Steam Cycle Mech.
Efficiency
Combined Cycle Mech.
Efficiency (HHV)
Effective Fuel Loss Per Btu
Process Heat Required/
Vice/
Bu Mines/
Selexol
265
6,670
6,208
159
6,829
6,367
.932
953,200
.355
916
300
ncy
.733
297,9^0
.286
.476
w -601
Bu Mines/
Iron Oxide
1,070
6,311
5,810
778
7,089
6,588
.929
979,880
.352
913
300
.732
284,920
.298
.478
.623
BCR/
Selexol
1,000
6,564
6,084
775
7,339
6,859
.935
1,167,924
.405
1,107
300
.786
401,480
.313
.535
.585
BCR/
Consol
1,700
6,636
.6,133
1,729
8,365
7,862
.940
1,346,582
.4io
1,115
300
• 790
4o6,o4o
. 3l4
. 54l
.580
-2l4-
-------
Table 1*7
EFFECT OF COMBINED CYCLE OH GASIFIER & CLEANUP PERFORMANCE
Energy
(MM Btu/hr, Unless Otherwise Specified) Bu Mine
Coal Energy Input (HHV) • 8
Fuel To Burner
Clean Fuel Gas Flow, Ib/hr 2,671
Gas Temperature, F
Chemical Energy in Gas (HHV) 6
Sensible Heat in Gas (above 59 F)
Total Energy in Gas to Burner (HHV) 6
Gasifier
Clean Fuel Gas Energy Out (at Burner Supply
Temperature ) 6
Heat Available from Gasifier and Gas Stream
Useful(«)
Rejected (»)
Process Heat Requirements
Added to Fuel Gas Stream
Input to Steam
Gasifier Air Preheat
Net Process Heat Used (Required less Useful
Heat Available)
Effective Loss in Fuel Energy (Net Heat
* nsTM/Hcc)
Effective Energy Out (Clean Gas Energy
Less Loss ) 6
Effective Gasifier Efficiency
Cleanup System
Net Process Heat Required by Sulfur Cleanup
Effective Loss in Fuel Energy
Fuel Consumed in Sulfur Recovery
Total Effective Fuel Energy for Sulfur Cleanup
Net Process Heat Required by NH^ Cleanup
Effective Fuel Energy for NHj Cleanup
Effective Fuel Energy Out of Cleanup 6
Effective Fuel Cleanup Efficiency
Combined Cycle Performance
Heat of Compression-Oasifier Air Supply
Effective Fuel Energy From Compression
s/Selexol
,015
,938
265
,670
159
,829
,882
672
57U (180)
190
307
386
211
127
,755
.81*3
1*8
29
53
82
382
230
,1*1*3
.951*
509
306
Total Effective Fuel Energy to Combined Cycle 6,7!*9
Gross Combined Cycle Shaft Power Output - Mw
Total Compressor Power for Gasifier Air - Mw
Gas Turbine Net Electrical Output - Mw
Steam Turbine Net Electrical Output - Mw
Total Generator Output Power - Mw
Auxiliary Power - kw (* Generator Output)
Gasifier • ' 10
Gas Cooling ' 1
NH3 Cleanup 2
Sulfur Cleanup ll*
Steam Cycle Plant Auxiliaries 7
Gas Turbine Plant Auxiliaries 3
Net Plant Output - Mw
91*1.3
11*9.1
552.0
223.8
775.8
,500 (1.35*)
,286 (0.16*)
,127 (0.27*)
,31*5 (1.85*)
,1*08 (0.95*)
,ll*6 (0.1*1*)
737.0
~.-._ -21'
Bu Mines/
Iron Oxide
8,015
2,1*97,630
1,070
6,311
778
7,089
7,623
21 1*
7U
261*
31*1*
1*68
292
7,331
.915
1*92
307
531*
81*1
—
—
6,1*90
.885
1*1*5 .
277
6,767
91*7.7
130.U
590.1
208.1
798.2
'• 10,500 (1.32*)
25,991 (3.26*)
6,889 (0.86*)
3,363 (0.1*2*)
751.5
BCR/Selexol
8,610
2,689,108
1,000
6,561.
775
7,339
7,1*72
1,973 (75)
290 (158)
777.
555
395
(-21*6)
(-11*1*)
T.616
.885
221.
131
133
26U
31*1
199
7,153
.939
589
31*5
7,1*98
1,175.3
172.6
686.5
293.3
979.8
21,228 (2.17*)
2,1*02 (0.25*)
35,090 (3.58*)
9,709 (0.99*)
3.913 (0.1*0*)
907.5
BCR/Consol
8,610
3,170,360
1,700
6,636
1,729
8,365
8,365
31*9
65
555
395
601
3!*9
8,016
.931
106
61
—
61
—
—
7,955
.992
590
3»*2
8,297
1,315.2
172.8
817.1*
296.6
l.lllt.O
21,228
(1.91*)
6,381
(0.57*)
9,819
(0.88*)
1*,659
(0.1*2*)
1,071.9
-------
the gas stream on overall efficiency. The actual heat lost as compared to a
high temperature system with no gas cooling is the indicated value less the
contribution due to condensation of the gasifier steam (given in parenthesis).
A similar accounting is followed in the cleanup process. Because the gasifier
was charged with the energy required to raise the inlet air from 59 F to gasi-
fier inlet temperature, the heat of compression of the gasifier air must be
credited to the system. This is done in the same manner using the ratio of
steam to combined-cycle efficiency. The result is an effective fuel energy
which, when used at combined-cycle efficiency, will produce the gross plant
output.
While this approach may penalize a system more heavily than is warranted
for the use of low-grade steam, it does provide a guide to determining the
areas requiring attention. After subtracting compressor power and accounting
for generator efficiency, the total electrical output corresponds to total
generator output (computer calculations) within 1 percent. The auxiliary
power requirements are separated to show the amount consumed by each part of
the system.
The effect of this approach to bookkeepting is to reduce the penalty
associated with the use of heat extracted from the steam cycle as opposed to
either chemical or sensible energy from the fuel gas stream. Therefore,
whereas the BCR/Selexol and BCR/Consol Systems have almost the same gasifier
net thermal requirements (see net fuel energy leaving gasifier in Table ^3,
which is a straight thermal energy accounting) the effective fuel energy out-
put at combined-cycle efficiency is significantly higher for the Consol System.
This occurs because there is no need to cool the fuel leaving the gasifier and
the fuel sensible energy is used at combined cycle efficiency. Although the
resultant net heat required by the gasifier is quite high, the performance
penalty associated with the use of thermal energy from the steam cycle is low
in comparison to the sensible heat in the fuel gas.
The advantages of high temperature in the cleanup can be readily evalu-
ated by comparing the BCR/Selexol and BCR/Consol Systems. The benefit of high
temperature shows up in the gasifier performance. With the same coal input
rate, the effective gasifier output is some kOO MM Btu/hr greater in the high-
temperature system than in the low-temperature system. This ^00 MMBtu/hr corre-
sponds to a 7 percent increase in net output power and compares to an overall
18 percent increase in net output power for the high-temperature system. The
remaining difference is due to a more efficient sulfur removal system and the
lack of the process requirements of an ammonia scrub. The difference in sul-
fur removal systems accounts for 7 percent and the ammonia removal for another
h percent of the overall difference in net output. Thus, while the high fuel
temperature is important, it is not the major contributor to the improved
-216-
-------
performance of the BCR/Consol System. In fact, if the low-temperature cleanup
system were as efficient, (i.e., had similar process requirements) performance
of the two could be made to match closely by regeneration and resatura-
tion with water vapor of the fuel gas stream.
While a performance value for sulfur removal has not been set, the
Selexol System offers the advantage of 95 percent + sulfur removal as opposed
to a concentration of 5^-0 ppm in the Gonsol product gas. The reason for the
high auxiliary power and process heat in the Selexol process is to permit
virtually complete removal of COS. If it were designed only for H S removal,
the sulfur level in the product gas would rise to about 700 ppm and the util-
ities would drop to the level of the Bu Mines/Selexol System. This is within
about 1 percent of the Consol System in terms of net output power.
Comparing the two gasifiers with low-temperature cleanup, the advantage
of BCR-type over the Bu Mines-type can be attributed to the difference in sen-
sible heat rejected, amounting to 262 MM Btu/hr, and the latent heat in the
gasifier steam (75 MM Btu/hr) utilized in the BCR system. When combined,
these total U.2 percent of the fuel input, the exact difference in effective
gasifier efficiency. The reasons for the difference in Selexol system perfor-
mance has already been discussed. The major factor in the higher performance
of the BCR/Selexol over the Bu Mines/Selexol System is, of course, the turbine
inlet temperature rather than gasifier or cleanup performance.
Considering the two Bu Mines-type systems, the high-temperature cleanup
has a considerable effect on gasifier performance raising the effective effi-
ciency by some seven (7) points. However,- overall system performance shows
little improvement over the low-temperature systems due to the inefficient
sulfur recovery system. By producing sulfuric acid as opposed to elemental
sulfur as in the other systems, significant improvement might result since the
high utilities are entirely due to the production of SO in the iron oxide
regeneration process.
COST COMPARISON
For each of the four selected systems, an estimate of capital and oper-
ating cost has been made. In addition, the cost of power generation in a con-
ventional plant with stack gas cleanup and in a gasified coal-fired steam
plant has been estimated to provide a basis for comparison. A summary of
total power generation cost for the six systems is presented in Figure 6k.
These are based on mid-197^ costs with escalation through the construction
period. It is quite clear that the most attractive gasification/cleanup sys-
tem for use with a COGAS powerplant is the BCR/Consol combination. This
-217-
-------
FIG. 64
COST SUMMARY
40
GASIFIED COAL-COGAS SYSTEM
GASIFIED COAL-STEAM SYSTEM
BU MINES/SELEXOL
(FIRST GENERATION)
30
BCR/CONSOL/STEAM
to
O
o
cc
111
H
LU
O
CC
LU
I
CL
20
BCR/CONSOL
(SECOND GENERATION)
CONVENTIONAL
COAL WITH
STACK GAS
CLEANUP
BCR/SELEXOL
(SECOND GENERATION)
BU MINES/IRON OXIDE
(FIRST GENERATION)
10
I
0.50
1.00 1.50
COAL COST $/MM BTU
2.00
R10-95-1
-218-
-------
second-generation combination appears capable of generating electric power at
much lower cost than any other alternative, including steam with gasification
and steam with stack gas cleanup. The other gasification/COGAS system combin-
ations are only marginally competitive with the steam'alternatives. Unlike a
nuclear system where fuel cost is relatively insignificant, capital cost ef-
fects will not be negated by changes in coal cost and the relative cost com-
parisons do net change greatly over the coal cost range of interest. Consid-
ering the zero coal cost intercept, which is a rough indication of capital
cost, it is seen that a 30 percent increase in cost results in changing the
BCR/Consol power system from COGAS to steam.
Also, for a given gasifier and power system, the low-temperature cleanup
is more costly than the high. This is mostly due to the lack of ammonia re-
moval and associated equipment in the high-temperature cleanup systems. How-
ever, in the case of the BCR/Consol System that effect is.amplified by the
improved efficiency of the gasifier-cleanup combination which results in a
greater plant output over which to amortize the capital cost. Changing from
first- to second-generation systems also tends to decrease initial cost due to
a higher gas turbine specific output resulting in a lower per kilowatt cost of
• the power system.
The proportion of total cost ascribed to the various major system elements
are given in Tables U8-50. These assume an owning cost of 1? percent of the
capital investment spread over a load factor of 70 percent. On a straight
capital cost basis, the coal-fired plant with stack gas cleanup is virtually
as expensive as any of the gasified coal - COGAS systems. However, the
estimated operating costs are about 2 mils more for the low-temperature systems
as opposed to the conventional steam system. The absolute magnitude of this
difference is subject to a fair degree of uncertainty. Lacking specific data
on operating costs for the gasifier and cleanup system, they were taken to be
8.5 percent of the capital cost. For steam plants this factor runs about 2
percent and for combined-cycle plants about 3-5 percent. The 8.5 percent was
chosen to be somehwat conservative since process plants generally run slightly
lower than this.
A major factor that must be remembered in comparing costs is the differ-
ence in the form of the sulfur by-product, the variation in sulfur removal
efficiency, and NO production which is certainly better in low-temperature
systems where all 'ammonia is scrubbed from the product gas.
All of the gasified coal plants include provision for recovery of elemen-
tal sulfur. No credit has been taken for recovery of this resource, nor has
credit (debit) been given to the disposal of the by-products of stack gas
cleanup.
-219-
-------
Table kQ
INTEGRATED SYSTEM COST SUMMARY
First Generation
Second Generation
Capital Costs - $/kw
Power System Cost
Gasification System Cost
Cleanup System Cost
Total Plant Cost
BuMines
Selexol
232
111
88
BuMines
Iron Oxide
230-
107
BCR
Selexol
208
117
_89
UlU
BCR
Consol
190
99
32U
Owning-Plus-Operating Costs-/
Owning Costs (17$ of Capital)
Operation & Maintenance
Power System
Gasification & Cleanup
Fuel Cost at 60<£/MM Btu
Total Cost of Power
11.
10.66
11.47
8.97
_220_
-------
Table ^9
COST SUMMARY FOR
COAL-FIRED STEM WITH STACK GAS CLEANUP
Capital Cost - $/kw
Coal Fired Plant
Stack Gas Cleanup 8l
Total Capital Cost ^26
Owning-Plus-Operating Costs-mills/kwhr
Owning Cost (Y\% of Capital) 11.8
Operation & Maintenance
Steam System 1.1
Stack Gas Cleanup 1.1
Fuel Cost at 60<£/MMBtu 5.8
Total Cost of Pover 19.8
Net Plant Heat Rate - Btu/kwh 9,721
-221-
-------
Table 50
COST SUMMARY FOR
BCR/CONSOL STEAM-CYCLE POWER PLANT
Capital Cost - $/kw
Steam Plant 252
Let-Down Turbine 28
Gasifier 111
Cleanup UU
Total Capital Cost ^35
Owning-Plus-Operating Costs-miLlg/kwhr
Owning Cost (Y\% of Capital) 12.0
Operating Cost
Steam Plant .8
Gasification & Cleanup 2.5
Fuel Cost at 60<£/MMBtu 6.1
Total Power Cost • 21.k
-222-
-------
In preparing the capital and operating cost estimates for these systems
the degree of detail with which each part of the system was evaluated varies
with the available data. As a result, the confidence level is not the same in
the various parts of one system or between different systems. However, the
various component costs are treated consistently and the resultant comparisons
are believed to be valid.
For each system, direct construction costs for the individual sections
were identified which include all equipment and installation costs. To this
is added an 8 percent contingency factor and 15 percent for engineering and
supervision. Escalation is applied at 7-5 percent assuming a linear rate of
expenditure. The same approach is used for interest during construction which
is taken to be 10 percent per year. Using the "General Cost Model" (50) as
discussed in Appendix C, the costs can be related to specific geographic re-
gions. Unless otherwise noted, all cost summaries are typical of an Illinois,
Ohio, Pennsylvania region.
Gasified Coal - COGAS System
The costing procedures described in detail in Appendicies C and D were
followed in estimating gas turbine manufacturing cost and in determining power-
plant costs. The gasifier and cleanup sections were treated like subsystems
of the powerplant. While it is difficult to assign responsibility for the cost
of the boost compressor and heat exchange equipment located in the gasifier
air supply stream, these costs have been charged to the gasifier in the summary
charts but have been carried as a separate item in the cost tabulations. Gas-
ifier and cleanup system utility requirements are treated as an addition to the
pertinent part of the main power system. The cost of feedwater treatment,
cooling tower and condensate polishing equipment are calculated on the basis
of their marginal or incremental cost which is then assigned to the appropriate
subsystem. Since there is no provision for waste water treatment in the main
power system an allowance of k percent of total gasifier and cleanup system
cost was made for that function. This cost was assigned to the cleanup system.
Because plant size was determined by the net output resulting from a
fixed (700,000 Ib/hr) coal firing rate, equipment size considerations would
generally result in some modification of that firing rate. However, this
modification would be slight since plant size is generally sufficient to ac-
commodate 8 gas turbine engines designed for maximum output power corresponding
to their turbine inlet temperature. Two steam turbines would be used when
steam turbine size is greater than 100 Mw. Figure 65 shows the variation in
gas turbine per kilowatt costs as a function of size and operating conditions.
For each generation the minimum cost engine was used, that being 70 Mw for
-223-
-------
EFFECT OF OUTPUT POWER ON GAS TURBINE PRICE
1.0
I
ro
ro
0.8
LU
O
DC
a.
LU
E 0.6
CD
GC
D
I-
co
O
LU
^ 0.4
FIRST GENERATION
TURBINE INLET TEMPERATURE = 2200 F
COMPRESSOR PRESSURE RATIO = 16
SECOND GENERATION
TURBINE INLET TEMPERATURE = 2600 F
COMPRESSOR PRESSURE RATIO = 24
LU
Ol
-------
first generation systems and 100 Mw for second generation systems. Engine
costs for these two designs were determined using the procedure described in
Appendix D.
The detailed compilation of power system costs are presented in Table 51
using the Bu Mines/Selexol system as an example. Since the powerplant cost
data are for the East Coast (New York, New Jersey) area, the General Cost
Model discussed in Appendix C was used to transpose the resultant cost to the
Central (Illinois, Ohio) area. For the Bu Mines/Selexol System this is shown
in Table 52 which gives the relationship between the two areas and also the
total power section costs escalated through a four year construction period.
Costs for the power system sized to operate with each of the four gasifier and
cleanup combinations are summarized in Table 53-
Gasifier and cleanup system costs are summarized in Table ^k. These were
estimated using currently available cost data. For lack of cost data on the
high-temperature desulfurization processes, investment costs were developed
from conceptual plant designs for the 'Consol and Sintered Iron Oxide processes.
Other sections were estimated from cost curves developed from published data.
Considering the sources from which the costs were developed, an accuracy
of not better than + 25 percent should be expected. Of the total investment
costs, the coal gasification system represents about half. The available data
on fixed-bed gasification systems costs are very sparse since no commerical
plants have yet been constructed in the U.S. Reliable data for second-genera-
tion gasifiers are essentially non-existent since these systems are only in
developmental stages where preliminary estimates have been made on conceptual
pilot plant designs.
The only significant conclusion that can be drawn from the cost figures
is that, for a given gasification system, the low-temperature cleanup scheme
is more expensive, possibly by 30 percent. The costs for the desulfurization
sections appear to be comparable for both the low- and high-temperature
schemes, the former scheme must also bear the additional costs associated
with gas cooling, sour water stripping, and ammonia recovery.
Coal-Fired Steam Station
A cost estimate was prepared for the reference coal-fired plant described
in Section 1. Capital costs are summarized in Table 55- These are based on
previous work (3) and have been escalated to a mid-197^- time period and appro-
priate interest and escalation added for a four year construction period.
Stack gas cleanup system costs were added. For the mid-197^- time frame a
-225-
-------
Table 51
POWER SYSTJM COST DETAILS
BuMines/Selexol/COGAS
Account
3U1-18
3^1-20
3^1-23
3U1-2U
Account 3^3
3^3-01
3^3-02
3^3-03
3^3-01*
3U3-05
31*3-07
31*3-08
31*3-09'
31*3-10
.31*3-12
31*3-13
3l*3-lU
3^3-15
3^3-16
343-17
3U3-18
Account
Site Preparation
Administration Building
Turbogeneration Building
Tank Farm
Condensate Polishing System
Stack
$8U8,250
563,830
3,9^8,000
1,236,100
800,6UO
362,000
Total 3Ul: $7,758,820
Gas Turbine (8 ) Includes Installation, Labor
Starter Motor (8 )
Torque Converter
Lube Oil Purifier & Storage (Pumps, Filters, Etc.)
Lube Oil Fire Protection
Air Compressor Services, Instrumentation
Breeching
Expansion Joints'; Not Applicable in COGAS Plant
Inlet Air Filters
Energy Cooling Tank Pump & Piping
Fuel Oil Heaters & Pumps
Miscellaneous Pumps & Tanks
Control Panels
Computer Controls
Fuel Piping
Fuel Pipe Insulation
Total
Generator For Gas Turbine
$17,6^3,500
85,000
8oU,ooo
224,000
160,000
lUo.OOO
2,lUO,600
0
555,760
11,200
123,200
56,000
560,000
560,000
187,500
$2^,500,760
$10,537,630
-226-
-------
Account 312
312-01
312-02
312-03
312-0^
312-05
312-08
312-09
312-10
312-11
312-12
Account
31^-01
Waste Heat Boiler
Boiler Feed Pump
Boiler Feed Tank Deaerator
Water Treatment (Demineralization)
Condensate Storage Tank
Miscellaneous Pumps
Piping
Insulation for Piping
Controls
Computer )
steam Turbine Only
$22,^32,1*00
339,390
133,070
691,960
29,9*10
70,525
3,080,2^0
26U,i|20
308,025
_
Total 312: $27,3^9,970
Steam Turbine and Generator
(Output per Unit - 105,800 kw)
31^-03 Condenser & Tubes
31^-OU Condensate Vacuum Pump & Motor
31^-05 Condensate Pump & Motor
31^-06 Cooling Tower
31^-08 Circulation Water Valves & Expansion Joints
31U-09 Circulation Water Pumps
31^-10 Make-Up Structure: Screens & Pumps •
Total
Accounts 3^5 & 353
Account
Accessory Electrical Equipment
Miscellaneous Power Plant Equipment
$ 9,833,270
1,212,180
13^,795
27^,780
5,9914,080
l,li+6,280
' _
$18,595,385
,768,750
$359,0^0
Other Expenses
Total Direct Construction Costs
$1,957,
Contingency
Engineering & Supervision (15$)
Total Unescalated Cost
$99,827,765
$7,986,220
$lU, 97^,165
$122,788,150
-227-
-------
Table 52
EFFECT OF LOCATION ON POWER SYSTEM COST
BuMines - Selexol - COGAS
East Coast Central
Materials & Equipment $6l,2UU,800 $6l,2UU,800
Wages, Fringes, Administration & Supervision 26,355,680 22,95^,0^0
Engineering 9,186,720 9,186,720
Tools During Construction 2,878.505 • 2,878,505
Total Direct Construction Costs 99,665,705 96,26^,065
Contingency 9,966,570 9,626,U05
Profit & Insurance 13,155,875 12,706,855
Total Before Escalation $122,788,150 $118,597,325
Escalation 20,595,815 19,892,870
Interest During Construction 33,272,250 32,136,6.50
Total Capital Cost $176,656,215 $170,626,81*5
-228-
-------
Table 53
COGAS POWER SYSTEM COST SUMMARY
Costs - $1,000
BuMines/
FPC Account Number
31+1-
3U3 -
31+1+ -
312 -
311+ _
31+5 &
31+6 -
IX)
vo
Structures and Improvements
Prime Movers (Gas Turbine)
Electric Generators (Gas Turbine)
Boiler Plant Equipment
Steam Turbine Generator Units
353 - Accessory Electrical Equipment
Miscellaneous Power Plant Equipment
Other Expenses
Direct Construction Costs
Contingency, Engineering & Supervision
Total Construction Costs
Interest & Escalation
Selexol
7,U9U
23,665
10,178
26,1+17
17,961
8,1+69
31+7
1,891
96,1+21
22,177
118,597
52,030
BuMines/
Iron Oxide
7,U91+
21+ ,637
10,315
26,201+
18.1U8
8,679
351
1,917
97,7^9
22,1+82
120,231
52,71*6
BCR/
Selexol
11,011
20,725
9,51**
29,1*1*5
23,230
10,350
385
2,093
106,75**
21+ ,553
131,308
57,606
BCR/
Consol
11,1+22
23,556
10,077
32,360
22,822
12,009
1+16
2,253
lll+,9ll+
26,1+30
11+1,31+1+
62 .,009
Total Capital Cost (Power System
Only)
170,627
172,978
188,913
203,353
-------
Table 5U
GASIFIER & CLEANUP SYSTEM CAPITAL COST BREAKDOWN
BuMines/ BuMines/ BCR/ BCR/
Selexol Iron Oxide Selexol Consol
Gasification 62.1*2 62.1*3 83.23 83.23
Gas Cooling ll*.86 23.78
Desulfurization . 23.78 20.81 29-72 20.81
Sour Water Stripping 5-9^ 5-9^ 1.1*9
Ammonia Recovery 10.1* 8.92
Sulfur Recovery 2.97 10.1* 2.97 8.92
Waste Water Treatment 1*.82 3.75 6.18 U.58
Boost Compressor & Boiler 10.8 10.72 11.0 11.73
rv> Feedwater Treatment 6.73 5.8 9.1*2 9.1*2
-------
Table 55
COST SUMMARY - COAL -FIRED STEAM PLANT
1000 Mw Nominal Size
FPC Category Cost - $1,000
311 - Structures & Improvements 21,60*+
312 - Boiler Plant Equipment 91,752
31^ - Turbogenerator Units 53,906
315 - Accessory Electric Equipment 16,032
316 - Miscellaneous Power Plant Equipment 792
353 - Station Equipment 2,329
Other Expenses 3,728
Direct Construction Costs 190,
Indirect Costs U3,733
Total Construction Cost 233,876
Escalation and Interest 96, ^4U
Total Capital Investment 330,320
-231-
-------
total investment cost of $8l/kw was used.'^l) Another reference,(52) gives
capital costs ranging from $58/kw for a limestone system to $90/kw for a
Wellman-Lord system. However, it is generally quite difficult to determine
what corrections must be made to these numbers to put them on an equal basis
with those in this report. Reference 32 places the cost of a magnesia slurry
regenerable system at about $UU/kw for a 1972 time period. When adjusted to
197^ and using an approach to interest and escalation that is consistent with
this study, this increases to about $68.U/kw.
The overall effect of stack gas cleanup on power cost is about U.l6 mills
per fcwh. This is made up of some 2.7 mills capital costs, 1.2 mills operating
costs and 0.3 mills additional fuel costs. Because of the wide variations
that are reported in the literature, the coal-fired unit costs have been
presented as a band in Figure 6k, with regenerable systems represented by the
high side and nonregenerable limestones systems falling on the low side of
the band. As an overall check, the data in Ref. 32 was used to calculate the
cost of a magnesia slurry system under the ground rules selected for this
study. The resultant cost was 3-9 mills per kwh , as compared to U.l6 for
this study. Reference 52 gives overall costs of 3-68 and k.66 mills per. kwh
for nonregenerable and regenerable systems respectively.
Gasified Coal-Fired Steam Plant
The BCR/Consol/Steam System previously discussed was used to evaluate
the cost aspects of this type of system. This combination represents the
highest gasifier/cleanup system performance and lowest capital cost per unit
output. Gasifier and cleanup system costs have been presented earlier in
this section. The low Btu gas-fired steam system costs'^' were escalated as
required. The other major parts of this system are the air compressor, let-
down turbine, and the boilers and intercoolers located in that stream.
Because of the relatively high pressure ratio for the let-down turbine-
compressor combination, it was treated as two low-pressure ratio engines in
preparing the estimate. A breakdown of the total cost is presented in
Table 56.
-232-
-------
Ta"ble 56
BCR/CONSOL/STEAM PLANT COST SUMMARY
Cost $1,000
Gasifier & Cleanup System
(Less Boost Compressor & Boiler) 131,730
Low Btu Steam Plant 2lU,l8T
Compressor-Let Down Turbine
(including Heat Recovery) • 23,795
Total Capital Investment
(Includes Interest & Escalation) 369,712
Plant Output Power - Mw- 850
Unit Cost - $/kw 1*35
-233-
-------
SECTION 5
ENVIRONMENTAL IMPACT AND GOALS
The environmental impact of the integrated power system and the cleanup
system goals required to meet environmental and operational constraints are
identified in Section 5- Also, included in this section are discussions of
effects of low- and high-temperature cleanup processes on the economics and
performance of the selected power systems, the technology needed to attain the
selected goals, and further research and development required to meet these
goals.
IDENTIFICATION OF ENVIRONMENTAL IMPACT
The four integrated power systems defined in Section k were used as a
basis to identify the various pollutants which would be associated with the
gasification system and the power system. A summary of the estimated environ-
mental impact of these four systems is given in Table 57. In the following
paragraphs, a more detailed description is given for each of the pollutants
quantified in Table 57, as well as a qualitative description of other pollu-
tants which would be associated with the various processes.
Air Emissions from Gasifier/Cleanup Systems
It is apparent from Table 57 that the S02 emissions from the sulfur re-
covery plant could be significant (60 percent to 85 percent of the total
integrated sulfur emission). However, commercially available tailgas treating
processes may be used to further reduce sulfur levels in these cases. For
example, the addition of an alkali scrubbing system to treat the sulfur plant
tailgas can achieve >99 percent overall sulfur recovery. Tailgas treatment in
the BCR/Consol case will not significantly reduce the sulfur emission since 9^
percent of the S02 results from combustion of the fuel gas product.
Effluent streams from the battery limits plant producing clean, low-Btu
fuel gas from coal are summarized in Table 58 for the four integrated gasifier-
gas purification combinations selected for standard system analysis.
-------
Table 57
ESTIMATED ENVIRONMENTAL IMPACT OF INTEGRATED POWER SYSTEMS
(All Values in lb/106 Btu Coal Unless Otherwise Noted)
IX)
uo
vn
'
BOM-Selexol
2200 F COGAS
Total
BOM-Iron Oxide
2200 F COGAS
Total
BCR-Selexol
2600 F COGAS
Total
BCR-Consol
2600 F COGAS
Total
Waste
Water
219(3)
p?( 3)
2k6
27(1)
27
520(3)
pi, ( 4 )
«— '
5kk
75(3)
22U)
97
Thermal •
0 005
_ _ (5)
0.005
0.018
(5)
0.018
0.016,
(5)
0.016
0 00?
(5)
0.002
S02
0 -,p0(6)
0 088
O.U08
0.535(6)
0.337
0.872
O.U87(6)
o 080
0.567
0 033
0 SPO
0.553
NO( 1° )
0 3S6
VJ • _L 1 1
0.530
0 31
3
-------
Table 58
GASIFIER/CLEANUP SYSTEM EFFLUENT SUMMARY
Gasifier System
Bureau of Mines
BCR
Cleanup System
Elemental Sulfur,
LTPD
Anhydrous Ammonia,
qrnpT,
o j_ jr u
Ash/Slag,
STPD
Other Solids,
STPD
Waste Water ,
GPM
Stack Gas (Sulfur Recovery
mols/Hour
M.W.
Vol.# N2
02
co2
so2
NO
H00
Selexol
259.1
1 oo o
J-O J • ^
1369.6
•163. U
351
Plant )
7728
29.85
59.60
2.77
19.80
0.52
0.80
16.51
Iron Oxide
219.8
1369.6
21*9.6
. 31 ,191
28.95
75-77
7-78
8.25
0.21
0.17
7.82
Selexol
258.1
99.2
730.8
1*77
13,051
32.67
• U9.37
2.15
35.12
0.50
0. 1*3
12.1*3
Consol
256.7
730.8
139.8
129
1,613
28.18
97.03
' 2.31*
0.28
0.35
-236-
-------
In addition to the S02 emissions from the stack gas, there will be NOX.
The NOX is due to both thermal NOX from the incineration process and also to
the carryover of fuel nitrogen compounds, such as NHg, in the sulfur recovery
plant feed gas. Estimates for the S02 with and without stack treatment and
NO emissions are given in Table 59-
X
The only particulate emission which can be associated with the gasifier/
cleanup systems would be those due to the disposal of ash/slag or spent solid
absorbent. No estimates have been made for any entrainment of these products
during disposal.
Water Emissions From Gasifier/Cleanup Systems
The aqueous waste streams from the low-temperature cleanup systems are
potentially hazardous to the environment. These streams will contain non-
strippable contaminants originally present in the raw producer gas. Data on
the nature and quantity of trace contaminants in coal derived producer gas are
not available. At best, a partial qualitative identification can be made
based on information given in the literature.^ ' ' ' Analysis of waste
water streams given for specific coal gasification processes and coke plants
are given in Table 60.
In addition, trace elements present in raw coal may be volatilized during
the gasification process and subsequently scrubbed out in the water washing
steps. An indication of the elements likely to be found in the water stream
is given by the analysis of Illinois coals ^ ' and the process condensate from
gasification of an Illinois No. 6 coal via the Synthane process ^-'l as pre-
sented in Table 6l. Of particular concern are those elements named as hazard-
ous to human health by the EPA '•?7;. beryllium, fluorine, arsenic, selenium,
cadmium, mercury, and lead. These elements are all volatile and can be ex-
pected to appear in the raw producer gas and ultimately in the waste water
stream. These elements may pass through high-temperature cleanup systems es-
sentially untouched and thus appear in the stack gas much the same as from a
conventional coal-fired steam plant.
The high operating temperature of the BCR gasifier, chosen for second
generation use, would preclude the appearance of phenols, tars, and fatty acids
in the raw gas and consequently in the waste water. Aqueous wastes, from the
first generation system should contain these compounds and traces of heavy
metal contaminants. After steam stripping, the waste water from this type sys-
tem could conceivably contain the constituents given in Table 62.
Design of a water treatment system to render the aqueous wastes acceptable
for discharge to the environment would require detailed analysis for the specif-
ic gasification system involved. Such design is beyond the scope of the current
study.
-237-
-------
Table 59
AIR EMISSIONS FROM GASIFICATION/CLEANUP SYSTEMS
S02 N02
lb/MMBtu Ib/MMBtu
BOM/Selexol
Stack Gas W/0 Cleanup 0.320 .356
Stack Gas/Cleanup 0.032
BOM/Iron'Oxide
Stack Gas W/0 Cleanup 0.535 0.310
Stack Gas/Cleanup 0.038
BCR/Selexol
Stack Gas W/0 Cleanup O.U8T 0.300
Stack Gas/Cleanup .036
BCR/Consol
Stack Gas W/0 Cleanup 0.033
238
-------
Table 60
WASTE-WATER STREAM ANALYSIS
(ppm)
Constituent
Total Ammonia
Total Sulfur
Phenol
Thiocyanate
Cyanide
Fatty Acids
Chlorides
Carbonate
COD
PH
Coke Plant
1,800-4,300
0-50
100-1,500
10-37
1,200-2,700
2,500-10,000
8.3-9-1
Coal Gasification Plants
Synthane
7,000-10,000
1,1*00
2,600-6,600
20-200
0.1-0.6
500
17,000
15,000-38,000
•8.6-9.2
El Paso /Lurgi
1,050
500
1,750
239
-------
Table 6l
TRACE ELEMENT ANALYSIS OF ILLINOIS COAL
Element Illinois Coal* -Waste Water**
Al 1.29 % 1,000 ppb
Ca 0.77 U,000
Cl O.lU
Fe 1.92 3,000
K 0.16 .160
Mg 0.05 2,000
Na 0.05
Si 2.U9
Ti 0.07
As lU.O ppm 30 ppb
B 102.0
Be 1.6 130
Br 15. U
Cd 2.5 6
Co 9.6 2
Cr 13.8 6
Cu 15.2 20
F 60.9
Ga 3.1
Ge 6.6 30
Hg 0.2
Mn U9.H UO
Mo 7.5
Ni 21.1 30
P 71.1 90
Pb 3^.8
Sb 1.3 .
Se 2.1 360
Sn k.Q 20
V 32.7 3
Zn 272.3 60.
Zr 72.5
* Mean value for 101 coals analyzed including some 19 non-Illinois coals.
For Illinois coals only, EPA 650/2-7^-05^, July, I$jk contains a similar
table for 82 different coals.
** Process condensate from gasification of Illinois No. 6 coal.
-------
Table 62
POTENTIAL WASTE-WATER STREAM ANALYSIS FOR BOM GASIFIER
ppm
Ammonia (free & fixed) 200-HOO
Carbonates 250
Sulfides 10-100
Phenols 500
Fatty Acids 1,800
Thiocyanates 100
Cyanides 1-10
Chorides 500
Heavy Metals 10-20
BOD 2,500
pH 9
-------
Solid. Wastes From Gasifier/Cleanup Systems
Both first- and second-generation systems will discharge substantial
quantities of solid material for disposal to the environment, primarily ash .
and slag. With the coals selected, the BOM gasifier would produce about 1^00
TPD ash and the BCR gasifier approximately 700 TPD of slag. First-generation,
stirred-bed gasification systems will also discharge l6o TPD of ash and coal
dust recovered from the raw producer gas via particulate removal systems.
There is the possibility of briquetting a portion of this dust for recycle to
the gasifier, providing the carbon content justifies the additional expense.
In addition, there are chemical solid wastes generated in the high-temperature
cleanup systems; 86 TPD of sodium sulfate from the sulfur recovery section in
the first-generation iron oxide system, and 1^0 TPD of spent dolomite from the
second-generation Consol system. The latter may be recycled if the activity
for sulfur absorption can be restored after cold regeneration.
By-product sulfur from all systems, 200-250 LTPD, should present no sig-
nificant environmental problems. In the elemental form, it can be stockpiled
with little or no environmental hazard and is readily marketable. Similarly,
the ammonia by-product, 100-130 STPD, from low-temperature cleanup systems can
be stored in the liquid form with no associated pollution hazard and sold for
its chemical value. In the present study, no credit has been taken for by-
product sales.
Air Emissions From Power Systems
The cleanup systems used for the four systems were easily able to remove
the HgS in the low-Btu gas to levels which are a fraction of the EPA standard
for coal-fired power stations (1.2 Ib S02/MMBtu in coal), the low-temperature
systems could potentially achieve removal to levels that are somewhat lower
than the natural- gas standard (0.1 Ib" S02/MMBtu). It was assumed that the H2S
remaining in the fuel gas was converted to S02 during the combustion process.
Estimates for the SOo emissions are given in Table 63 and Figure 66 shows S02
emissions as a function of H2S in 150 Btu/ft^ gas.
Also shown in Table 63 are the NOX emissions. As was discussed in Sec-
tion 2, the total amount of NOV emitted is a function of combustion tempera-
jC
ture and time, the "thermal" NOX, and of the nitrogen-bound fuel constituents.
Thus, the low-temperature cleanup systems with their associated water washes
and solvents which also remove NIQ; (the main fuel nitrogen compound produced
during coal gasification) have low NOX emissions in which the thermal NOX could
play a significant role. Those systems using high-temperature sulfur removal,
however, let the fuel nitrogen compounds pass through, giving rise to N0x emis-
sions beyond the acceptable EPA limit for coal-fired stations (0.7 Ib NOx/MMBtu
2U2
-------
Table 63
POWER SYSTEM EMISSIONS
(rb/MMBtu)
so.
Particulates
BOM/Selexol
BOM/Iron Oxide
BCR/Selexol
BCR/Consol
0.088
.337
0.080
0.520
0.171*
3.79
0.012
2.65
<0.01
-------
FIG. 66
SO2 EMISSIONS AS A FUNCTION OF H2S IN FUEL GAS
HHV- 150 Btu/ft3
3
^-*
m
CM
O
to
.0
.X- EPA LIMIT FOR COAL-FIRED STATIONS
5000 10000
H2S IN FUEL GAS-ppm
15000
R05-123-1
-------
coal). In order to achieve environmental acceptability with the higher
performance systems using the high-temperature sulfur removal systems, it
will be necessary to identify a method of also removing fuel nitrogen com-
pounds at high temperatures.
The emission of particulates from these power systems is difficult to
quantify. The very limited information (5°) available on particulates in
gasifier off-gas which has been scrubbed by a sulfur removal system indicates
that the loadings vary from 0.25 milligrams/m3 (1.55 x 10~8 lb/ft3) in a K-T
system, to 1.0 milligram/m3 (6.2 x 10 lb/ft3) in a pressurized Lurgi system,
to 2.5 milligram/m3 (15.5 x 10 lb/ft3) in a Winkler system. Assuming that
these loadings are typical and would pass through the combustion process with-
out alteration, the particulate emissions would range from about 7-5 x 10" ^
Ib/MMBtu based on the low value (1.55 x 10~8 lb/ft3 for BOM/Selexol) to about
10"3 Ib/MMBtu based upon the highest value (15.5 x 10"8 lb/ft3 for BCR/Selexol).
These values are well below the EPA standard of 0.1 Ib/MMBtu for coal-fired
stations. In addition to the particulate carryover in the fuel gas, there are
particulates formed during combustion. Measurements '^°' on stationary tur-
bines firing natural gas show emissions of the order of 0.001 g/ft3 in the
exhaust. This is equivalent to approximately 0.75 x 10 Ib/MMBtu. Assuming
the low-Btu fuel gas burns as cleanly as natural gas, the total emissions
would range from 1.5 x 10"^ Ib/MMBtu to slightly over 1 x 10~3 Ib/MMBtu,
again well below EPA standards.
For those systems using high-temperature cleanup, there will be a higher
loading. The dust carryover for the BOM gasifier was given as 0.016 Ib/lb
coal (Table 2). This value, also assumed as carryover from the BCR gasifier,
is equivalent to about I.l6 Ib/MMBtu. To meet the EPA limit for coal-fired
stations, 92 percent removal must be obtained. High-temperature cyclones have
the potential for removing this mass of particles. It is not apparent that
meeting the EPA particulate limit means that the fuel gas and its combustion
products are clean enough for use in high-temperature turbines. This will be
discussed in the section dealing with cleanup system goals. Discussions on
advanced high-temperature removal systems such as panel bed filters indicate
that this type of system could have the potential of 99 percent-plus removal
efficiencies. This would be equivalent to approximately 0.01 Ib/MMBtu. For
purposes of this study, it was conservatively assumed that removal efficiencies
of 95 percent could be attained for first-generation systems and 98 percent
for second-generation systems giving rise to the values in Table 63.
Water Emissions From Power Systems
Because of the use of mechanical draft cooling towers, which eliminates thermal
pollution, the only significant water effluent is from boiler and cooling tower
-------
blowdown. Boiler blowdown is a function of steam conditions and water purity
and can range from 22 to 27 Ib H20/MMBtu coal input. The type of pollutants
associated with these streams has been discussed adequately elsewhere.(30)
DEFINITION OF CLEANUP SYSTEM GOALS
The goals to be attained by the cleanup systems selected in Section 3 of
the study are based upon two constraints - the EPA promulgated emission stan-
dards for coal-fired powerplants and the operational limitations of the gas
turbine. The latter is usually expressed as some type of specification for
fuels for turbine use. In setting goals to be attained, the limitations
proposed by system operating parameters must also be considered, allowing
flexibility in both the goals and the operating conditions so that the effects
of one on the other can be adjusted to minimize overall environmental impact,
EPA Standards
The EPA standards for coal-fired steam stations of 250 x 10 Btu/hr and
above have been discussed in previous sections. The rationale behind using
coal-fired steam station rather than gas turbine standards is twofold; first,
the overall system is, indeed, coal-fired and the environmental impact is a
function of the entire fuel processing, cleanup, and power systems; secondly,
no firm gas turbine standards have been promulgated. It is instructional,
however, to compare the coal-fired steam station, possible EPA gas turbine
standards (°^) and the gas turbine standards as viewed by the industry- through
the American National Standards Institute (ANSI) Committee for Gas Turbine
Procurement. (fol) These three goals are shown in Table 6k. (NOTE; Both the
EPA and ANSI emission standards are suggested and are subject to revision.)
From Table 6^ it is seen that the gas turbine regulations as proposed by
the EPA are essentially the same as those for steam stations. Those suggested
by ANSI are considerably more lenient (except particulates) and are based upon
the industry's recognition that the ability to burn "dirtier" (higher sulfur
content) fuels must be developed if a viable utility market for gas turbines
is to be maintained. Higher NOV limits are supported because it is imperative
J\.
that a dry combustor (no water injection) must be used since water consumption
could become a very major concern in many areas. The current industry efforts
in low-NOx burner research indicate that the 120 ppmv value is attainable dry.
Gas Turbine Fuel Specifications
Gas turbine manufacturers have established stringent specifications for
fuels to be burned in industrial applications. A summary of pertinent
-------
Table 6h
COMPARISON OF EMISSION STANDARDS FOR GAS TURBINE SYSTEMS
(ib/MMBtu)
Pollutant
SO.
Particulates
EPA
(1)
1.2
0.7
0.1
Agency
EPA(2)
0.3$ S Weight -1.2
55 ppm - 0.7
10/» Opacity
ANSI
(3)
.8$ S Weight -3.2
120 ppm - 1.53
Opacity
(10
(l) For coal-fired steam stations
(2) Suggested standards for gaseous fueled system subject to revisions
(3) Not endorsed by ANSI or the institutes or companies of its members
(k) Suggested 1980 standards would be 80 ppm or 1 Ib/MMBtu
-------
specifications is given in Table 65. The major problem in current utility use
of gas turbines is hot end corrosion. This is a result of two factors, the
presence of alkali metal salts in the fuel and of fuel sulfur. The corrosion
agent is the alkali metal sulfate which attacks the oxidation resistant
coatings. Once these coatings have been penetrated, rapid oxidation of the
base alloy occurs. The result is that, although the turbine could withstand
fuels having sulfur contents well above those allowed by environmental con-
straints, the presence of alkali metals reduces the allowable content to a
much greater degree.
The presence of particulates also contributes to corrosion since any
spalling of the coatings due to impingement could result in subsequent oxida-
tion of the base alloy.
Particulate Loading - One of the major concerns in setting fuel specifications
is the particulate content. As can be seen in Table 65, the particulate
loadings are low, varying from less than 1 ppm to about 30 ppm. The reason
for the low allowable loadings is the potential erosion of the turbine blades.
Although a number of attempts have been made to utilize coal directly in
( > ~>, )
a turbine > ~>, &^ have failed to obtain reasonable machine lifetime
due to excessive erosion. Yet, the actual limitations of particulate content
or size 'distribution have not .been adequately defined. In a previous UTRC
study ' -'' on methods of cleaning emissions from jet engine test cells, it
was shown that the mean diameter of the particles emitted from liquid-fuel
burning engines was O.lp, with over 99 percent-plus of the particles being less
than Ip,. Since erosion is not a problem with these engines, it can be con-
cluded that particles of l|j, or lower in the fuel or combustion products are
not harmful to engines. In fact, it was pointed out in this study that the
use of combustion additions to eliminate visible emissions resulted in agglom-
erated particles of greater than Ijj, (above the size range which results in re-
fraction). While deposition did occur, there was no erosion indicating that
the soft compounds which left the airstream and impinged on the blades did
not damage them.
Tests in the United Kingdom on fluid-bed combustion ^ ' also indicate
that soft particles of quite large diameters (~50|j,) do not appear to cause
erosion. However, tests carried out in Australia ^ '' indicate the hard
particles such as would result from coal combustion above the ash fusion point
cause erosion at sizes above 6y,. Perhaps the most useful guide would be the
work done in the petro- chemical industry on turbine erosion. ^ ' Results of
this work are shown in Figure 67, where engine life is shown as a function of
particle size and loading. The shape of Figure 6? indicates that the relation-
ship between lifetimes, particulate loading and particulate size would be of
the form
-------
Constituent
Sulfur
ro
-p-
VD
Particulates
Metals
Vanadium
Sodium and
Potassium
Calcium
Lead
Copper
P&WA Spec. 527
1.8 Mol % H S
(1)
0.08 lb/106 ft3
(0.00056 gr/ft3)
<0.2 ppm (Weight)
<0.6 ppm (Weight)
0.1 ppm (Weight)
0.1 ppm (Weight)
0.2 ppm (Weight)
Table 65
GAS TURBINE FUEL SPECIFICATIONS
(2)
Suggested
<1 Mol % H2S or Less Than
Amount Required to Form
5 ppm Alkali Metal Sulfates
0.01 gr/ft3
Uc y Maximum
see Sulfur Spec
Westinghouse
2% by Weight
(3)
Limits by Material
(U)
0.5 ppm (Weight)
0.5 ppm (Weight)
10 ppm (Weight)
2 ppm (Weight)
(l) For aircraft-derived turbine using gaseous fuels
(2) For industrial turbines; subject to revision
(3) Liquid fuel specifications
(h) Given in Ref. 70 as h x 10~ gr/ft3 of 2p to 10p particulate in gaseous fuels
General Electric^ '
Less Than Amount
Required to form
3 ppm Alkali Metal
Sulfates
30 ppm (Weight)
(0.01 g/ft3)
See Sulfur Spec.
-------
EFFECT OF PARTICLE SIZE ON ENGINE LIFETIME
FIG. 67
100
10
to
LLJ
_l
o
1.0
10
-1
PARTICULATE LOADINGS, gr/ft3
102
103 2 5 104
LIFETIME,hrs
105
RO2-192-1
250-
-------
Constant ,, .
, x C
where
L = lifetime
p, = particle size
C = loading
Thus, for a given lifetime, a reduction in particle size would be accom-
panied by higher allowable loadings, e.g., capture of particles of 2p, and
above would allow twice the particulate concentration that capture of Ify, and
larger particles would allow (assuming linear particle size distribution).
Suggested Low-Btu Fuel Gas Specifications
The values for loadings in Table 65 are based upon methane fuel. For
low-Btu fuel, both the heating value and the density will affect the allowable
loadings. For fuel gases of 150 Btu/ft^, (about 2500 Btu/lb) the value in
Table 65 would be reduced by approximately a factor of eight if the solids
loading to the turbine were to be kept constant. Using the foregoing approach,
Table 66 has been prepared to serve as a guideline for low-Btu fuel gas.
In addition to the usual fuel property limitations, Table 66 includes a
value for fuel nitrogen compounds expressed in terms of NHo. The limitation
is based upon 90 percent conversion of fuel nitrogen to NOX.
COMPARISON OF CLEANUP SYSTEM CHARACTERISTICS
One of the major considerations at the initiation of this study was the
identification of the point(s) at which one type of cleanup process, either
low- or high-temperature, demonstrated technical and economic benefit over the
other. As the study progressed, it became apparent that when the power system,
the gasification system, and the cleanup process were combined into an inte-
grated power station, much could be done to reduce the differences between the
various processes which would result from a cursory examination. The overall
results of the power station integration have been fully documented in Section
h of this report and the advantages/disadvantages of each system has been
identified. However, the effect of the sulfur removal/recovery system alone
on the performance and economics has not been defined.
To determine the penalties associated with the sulfur removal/recovery
processes, each integrated system was broken into four parts; (1) a
251
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Table 66
SUGGESTED LOW-BTU FUEL GAS CLEANUP SYSTEM GOALS
Property
Sulfur
Particulates
Specification
0.05 Mol % or Less Than Amount
to Form 0.6 ppra Alkali Metal
Sulfates
1* ppm (weight)
(0.0012 gr/ft3)
Resulting Emission
S02/MMBtu
<0.01 Ib
Metals
Vanadium
Sodium and Potassium
Calcium
Lead
Copper
Nitrogen Compounds
<0.03 ppm (veight)
See Sulfur Spec
<0.012
<0.012
<0.0025
500 ppm as HrU
0.3 Ib N02/MMBtu
252
-------
quench/cool down section, (2) a sulfur removal system, (3) a sulfur recovery
section, and (4) an ammonia removal/sour water stripper. The thermodynamic
losses associated with each of these sections was identified, either as a
loss in sensible heat which otherwise could be used to generate power or as a
direct requirement for fuel for combustion. The costs of sulfur removal were
also calculated on the basis of $/lb S removed. The comparison of these clean-
up system characteristics are given in Table 67.
Low-Temperature Systems
The coupling of the Selexol process with the BuMines-Type and BCR-type
gasifier presents an opportunity to define the effect of low-temperature clean-
up on performance since there is nearly twice the sensible heat in the BCR gas
stream to be removed before scrubbing. Yet the net effect as shown in Table
67 is nearly the same, i.e., a 2.51 percent heat loss for the BuMines versus a
2.36 percent loss for the BCR. The mechanisms of loss, however, are different
enough to require discussion.
In the BuMines system, nearly 10.5 percent of the total fuel gas heating
value is sensible heat; for the BCR, slightly over 19 percent. Because of the
tar scrub required by the BuMines gasifier, regeneration of the clean fuel gas
is possible to only 265 F, a process which recovers about 21 percent of the
heat (2.19 percent of the total gas stream energy). The BCR gas stream allows
regeneration to 1000 F, thereby recovering 51-2 percent (9-78 percent of
total). Slightly over 6 percent of the BuMines gas stream sensible heat is
used in the Selexol system versus lU.5 percent in the BCR stream (the COS in
the BCR stream requires a higher solvent flow rate). Process heat furnished
by the fuel gas sensible heat for ammonia removal/sour water stripping is k$
percent for the BuMines, 22 percent for the BCR.
In the sulfur recovery step, additional clean fuel is consumed in the
Glaus process. For the BCR-based process, 1.6k percent of the total fuel
energy is consumed; for the BuMines-based process, 0.71 percent. The higher
requirement of the BCR process is due mainly to the more dilute gases result-
ing from the higher solvent recirculation rates needed to remove the COS.
The auxiliaries associated with these systems use power which can be
expressed as equivalent fuel energy. Once again, due to the higher COS con-
tent, power for solvent recirculation results in a use of nearly 3-8 percent
total energy for the BCR versus the approximately 1.7 percent for the BuMines.
Overall, the energy consumption of sulfur removal/recovery associated
with the BCR-type gasifier is 7-78 percent of the total fuel gas energy while
that for the BuMines-type is h.^6 percent. By allowing a higher S02 emission
253
-------
TABLE 6?
COMPARISON OF CLEANUP SYSTEM CHARACTERISTICS
Low-Temperature Systems
Gasifier Outlet - Temperature - F
Cleanup Inlet Temperature - F
Fuel Temperature to Turbine - F
Fuel Gas Energy Utilization - %
Sensible Heat Available
Used to Regenerate
Used in Sulfur Removal System
Used in Ammonia Removal System
Net Heat Loss
Chemical Fuel Value for Sulfur Recovery
Cleanup System Auxiliary Fuel Equivalent
Total Cleanup System Losses - %
Cost of Sulfur Removed - $/lb
(1) All sensible heat referenced to 59 F
(2) Chemical HHV required to supply process heat
High Temperature Systems
BuMines
Selexol
1000
100
265
10.U7
2 1Q
5 13
2.51
0 71
1.7U
U.96
0.059
BCR
Selexol
1800
100
1000
19.10
Q 7ft
7 • 1
-------
for the BCR-based system, the losses- associated with the removal of COS could
be essentially eliminated. This could result in an over all penalty of less
than k percent with a subsequent overall system efficiency increase of around
two points.
The cost of sulfur removal for the BCR-based system is also higher,
$0.076/lbS versus $0.059 for the BuMines-based systems. This higher cost is
attributed to: (a) higher Selexol system costs due to the COS, and (b) the
more costly regenerative heat exchanger (see gas cooling costs in Table 5*0-
High-Temperature Systems
The BuMines-based high-temperature product gas contains 11.83 percent of
the total fuel gas energy as sensible heat while the BCR-based gas stream is
again 19-1 percent. There is no sensible heat loss due to cleanup and essen-
tially all the sensible heat is used at the efficiency of the combined cycle.
The sulfur removal system requires extraction of heat from the power cycle
(or from an auxiliary fired-boiler) to provide some process heat. For the
BuMines/iron oxide system this amounts to an equivalent of h.3 percent of the
total fuel gas stream energy; the BCR/Consol requires an equivalent of 0.7
percent. (Since the heat required is relatively low grade, it is more effi-
cient to extract it from the power system after some work has been performed
by the combustion/expansion of the fuel gas than by removing sensible heat
from the gas prior to combustion. This process energy required has been pro-
rated back to equivalent fuel gas energy as explained in Section U. )
The transformation of the H2S removed to elemental sulfur requires a very
significant amount of energy in the BuMines/iron Oxide Process; approximately
7-5 percent of the total fuel%gas energy. By converting the fuel gas H^S con-
tent to some form other than elemental sulfur, a very appreciable savings could
be realized. For the purposes of this comparison study, however, only elemen-
tal sulfur has been considered.
The BCR/Consol system has no net requirement for process heat for sulfur
recovery.
The total energy consumption penalty, including auxiliaries, associated
with the two high-temperature cleanup systems are 13-7 percent for the BuMines/
Iron Oxide and 1.1 percent for the BCR/Consol. The cost of sulfur removal for
the two systems are nearly equal at $0.052/lb S for the iron oxide and $0.05/
Ib S for the. Consol system.
255
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Comments on Cleanup System Selection
When comparing the various low- and high- temperature cleanup systems it
is necessary not only to look at the energy penalties associated with each,
but also to consider the net effect on overall cycle performance. For example,
the two first-generation systems based upon. the BuMLnes gasifier and the 2200
F COGAS system have essentially equal overall efficiencies (see Table ^5) yet
the penalties associated with sulfur removal indicate that those for high-
temperature systems are roughly three times as great as for the low-temperature
system. When the systems are integrated, however, the net result is that the
sensible heat in the high-temperature systems offset nearly all of these losses
so that the overall effect is that the integrated system with high-temperature
cleanup is approximately 2. percent more efficient (0.6 efficiency points) than
with the low-temperature cleanup.
The second-generation integrated systems show more dramatically the per-
• formance advantages of the high- temperature cleanup system, due mostly to more
efficient utilization of the very sizable sensible heat content of the BCR
product gas; i.e., approximately 9-5 percent more of the total gas energy is
used at combined-cycle efficiency rather than steam cycle or process effi-
ciency. This, coupled with the much lower process heat/auxiliary requirement,
results in a 17 percent (6 points) performance advantage. It should be noted
that a similar advantage (approximately 6 points) is also available if the
advanced gasifier is coupled with a first-generation COGAS system (see Table
As would be expected, the economics are more favorable for the generally
less complex high- temperature systems. The absence of heat- exchangers and
large-scale pumps and compressors for process work more than offsets the ad-
ditional cost for high-temperature materials.
The foregoing must be tempered, however by the realization that both the
performance and cost estimates for the low- temperature system are based upon
extrapolations of data for actual sour gas treatment while those for the high-
temperature systems are based upon extrapolations from bench- scale or small
pilot facilities. In addition, the emissions of Hox and participates from
systems using the high- temperature processes are potentially unacceptable.
Nonetheless, the results of the analyses performed during this study in-
dicate enough potential advantages for the high-temperature cleanup systems
that further development in this area should receive every consideration.
256
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TECHNOLOGY REQUIRED TO OBTAIN FUEL SPECIFICATION GOALS
The suggested fuel specifications given in Table 66 were arrived at
through consideration of environmental aspects (sulfur and nitrogen components)
and turbine requirements (alkali metals and particulates). The following
paragraphs discuss these suggested fuel specification goals and their implica-
tion on the technology required to achieve them.
Technology for Sulfur Removal
The suggested fuel sulfur specification (0.05 Mol percent or 500 ppm) can
be achieved in nearly all of the currently available low-temperature cleanup
systems. These systems are capable of cleaning to 100 ppm or lower and, in
fact, the technical and economic differences between a 500 ppm and 100 ppm
limit are small.
Several of the high-temperature cleanup systems, such as those based upon
iron oxide can clean to the 100 ppm level. However, the half-calcined dolomite
system must have its operating conditions altered to achieve the 500 ppm goal.
Since the operating temperature is fixed by the heat balance around the desul-
furizer, the temperature differential of about 50 F needed to reduce the con-
centration from 630 to 500 ppm residual sulfur can only be obtained by raising
the gasifer operating temperature. Alternatively, since the residual sulfur
content is inversely proportional to total pressure, (see Figure V/ shown pre-
viously) a 10 percent reduction in operating pressure would result in achieving
the goal of 500 ppm sulfur in the treated gas.
The foregoing does not require technology changes, rather it requires
consideration of altered operating conditions. However, the use of high-tem-
perature fuel gas would require changes in the technology of the gas turbine.
Currently, the control system of the gas turbine is based upon the metering
of fuel into the burner. Machines in operation today are capable of handling
fuel to kOO F in these controls. Research is being carried on to develop a
fuel control system which will work to over 800 F. It appears that consider-
able attention must be paid to developing either a fuel control which operates
to 1750 F or above, or to developing new integrated system controls which
would use the gasifier input as the control element. Because of the thermal
inertia represented by the gasifier, it is not apparent that control from in
front of the gasifier will give satisfactory performance under utility operat-
ing conditions.
Technology For Fuel Nitrogen Removal
As with the sulfur removal, it is not necessary to develop new technology
to remove nitrogen compounds, particularly ammonia, from the effluents coming
257
-------
from the gasifier and low-temperature cleanup process. The low-temperature
scrubbing systems utilize water wash which will remove most of the am-
monia. In addition, the sulfur removal system will absorb ammonia, Selexol,
for example is more selective to ammonia removal than to carbon dioxide. The
process is most selective to HCN, another potential fuel nitrogen compound.
Thus, fuel nitrogen content of low-Btu fuel gas from these scrubbing systems
is quite low and does not appear to contribute significantly to NOX formations.
Unfortunately, the high-temperature processes including the iron oxide
or dolomite-based sulfur removal systems do not remove any of the ammonia or
other nitrogen-based compounds from the fuel gas. As can be seen in Table
57 > the fuel nitrogen does contribute to overall NOX production. Catalytic
decomposition of ammonia to ^ and. Hp has the potential to reduce ammonia
content to perhaps 200 ppm for the BCR/CONSOL. However, a catalysis system
capable of working under the conditions assumed present in the high-tempera-
ture stream, e.g. 1600 F to 1700 F at approximately 500 psi has yet to be
identified. This method does show the only promise and should be 'further
investigated.
'Technology for • Particulate Removal
The low-temperature scrubbing systems are capable of removing particles
to the level of cleanliness beyond those specified'in Table 66. ThuSj little
problem is expected to arise from fuel introduced particles with this process.
Once again, the high-temperature system would be subject to particulate based
problems.
A review of the literature indicated very little work had been done at
the temperatures of interest. Some work has been done on developing electro-
static precipitators for operation up to 1500 F although these systems have
not had significant commerical application. (°9) Of greater promise, is the
development of high-temperature filtration systems through beds of granular
solids such as the panel bed sand filters which claim 99 percent fly ash
removal efficiencies. '' ' ' ' This concept has not yet been demonstrated
on a commercial scale.
A previous study ' •*' had investigated various devices for removing par-
ticles in jet engine exhausts at temperature to 1100 F. Dry cyclones, fabric
and ceramic filters and electro-static precipitators were investigated as were
removal methods which were still in an early stage of development such as
acoustical and magnetic separators. None of these were attractive for the
engine exhaust application, either because of the very high flow rates in-
volved or of the high-temperature. These methods were reviewed and it was
determined that cyclones and ceramic filters could handle both the flow and
258
-------
the temperatures associated with the high- temperature fuel gas stream although
their removal effectiveness was not necessarily adequate.
It was pointed out in a previous study ^ ' that the rotary stream dust
collector is a form of cyclone produced by the Aerodyne Corporation and 'known
as a tornado. This technology is well known in Europe. Aerodyne claims es-
sentially 99 percent collection for particles of 5p, and about 70 percent for
Unfortunately, the size distribution of the effluent particles from a
gasifier are not known. For purposes of this study, it was assumed that the
dust carry over for the BOM gasifier, roughly 10 percent of the coal ash,
would be present in the off gas from both the BOM and BCR gasifiers. It was
further assumed that the dust carryover resembled fly ash and demonstrated
weight fractions shown in Table 68.
The ability of the turbine to withstand particulates as a function of
particle size and composition. Unfortunately, there is little documentation
other than that already cited. ' ' ' ' Measurements of particle sizes in gas
turbine exhausts indicate that most (95 per cent -plus)' are less than Ip, with a
median size of 0.1 (°5). This would indicate that particles above l)j, are
fractured either by the high- turbulence within the turbine, or by collsions
with other particles or turbine parts. Since fly ash is a fluffy friable
material, easily broken, a 2^ upper limit va-s set (this is less than the 6^,
limit suggested in the Australian work ' ' ', but would result in a conserva-
tive estimate of capture effectiveness).
Based upon the foregoing, it appears that a cyclone-based system using an
Aerodyne- type final stage capable of removing 98 percent above 2|j, (equivalent
to < O.OU Ib/MMBtu) would be suitable since it is well below the EPA turbine
limit of Table 6U. However, this value (0.05 gr/ft-') is about a factor of
five higher than the gas turbine fuel specifications for methane (Table 65)
and a factor of ho higher than the suggested specification for low-Btu fuel
(Table 66). Thus, the ability to meet and actually surpass the EPA regulation
does not. in this case, satisfy the governing requirement, i.e., the capability
of the turbine to withstand particulate ingestion. A review of the very recent
literature and discussions with Pratt & Whitney Aircraft revealed only one new
bit of data on particulate loading, the estimate by Westinghouse (72) that,
in the pressurized fluid-bed effluent, a loading of 0.15 gr/ft^ would be suit-
able for the turbine. This is roughly 20 percent higher than the EPA limit.
The turbine for the fluid-bed application operates at relatively low tempera-
ture and these values may not be applicable in the higher temperature cases.
Thus, it is probable that more sophisticated particulate removal devices would
be required.
259
-------
Table 68
SIZE DISTRIBUTION OF FLYASH FROM BOILERS
(cumulative weight % above a given size)
Size. II
595 (30 mesh)
297 (50 mesh)
1^9 (100 mesh)
7^ (200 mesh)
50
25
10
5
2
Illinois No. 6
0.0
0.2
6.6
10.9
19.0
5^.0
82.0
9^.0
98.0
100
Wyoming
0.1
0.3
1.8
11.2
20.2
1*9.0
76.0
89.0
95.0
100
260
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RECOMMENDATIONS
The foregoing study has shown that the potential thermodynamic advantages
of high-temperature fuel gas cleanup are reduced by the operating character-
istics and requirements of the cleanup systems and, further, that the poten-
tial emission of NOx due to fuel nitrogen is unacceptably high compared to the
EPA standards. Nonetheless, the overall .efficiencies of those integrated
power stations using high-temperature cleanup are sufficiently higher (20 per-
cent) that efforts aimed towards early commercialization are well warranted.
Specific recommendations for further effort are given in the following para-
graphs .
Gasification
The BOM gasifier and the BCR gasifier were selected as representative of
two generic types of gasifiers, a low-temperature type having tars in the
effluent, and a high-temperature type having tar-free off gas. Actually, the
tar content has little effect on the cleanup system as it can be rather easily
removed from the off gas prior to the low-temperature sulfur cleanup step and
simply passes through the high-temperature cleanup in the vapor phase. As it
turns out, the ammonia (fuel nitrogen) content appears to be the major problem
area. While easily removed in the low-temperature system, the subsequent am-
monia stripping' and recovery are costly both in money and energy. High-tem-
perature removal of ammonia, while potentially possible, has not yet been
demonstrated in a form similar to that which would be required for use in con-
junction with pressurized gasification.
It would be warranted to examine a gasifier which does not produce ammonia,
i.e., a very high-temperature system which would essentially crack fuel nitro-
gen compounds to nitrogen, hydrogen and carbon. One such gasifier presently
operating on coal is the Koppers-Totzek entrained-flow gasifier.
The gasification of high-sulfur residual oil or of other high-sulfur
petroleum based products such as petroleum coke has been considered as an
alternative fuel supply system to power stations currently burning oil.
While still of interest in other areas of the world, particularly the USSR
and Japan, this process is currently in limbo in this country because of the
high costs of oil. Yet, future energy scenarios such as those, considering
the impact of Alaskan oil on the mainland West Coast could provide significant
supplies of residual for potential use in utility powerplants.
The present study has been limited to gasifier operation at a single
point because of the lack of hard, reliable data on gasifier operation. Since
26l
-------
it will be several years before any large-scale gasifier operation will occur,
other than the fixed-bed type, it would be advantageous to develop the capa-
bility to estimate the relationships of the operating conditions such as
pressure, temperature, air and steam flow rates to the efficiency of the
gasifier and to the makeup of the fuel gas. This could be accomplished for
entrained-flow type gasifiers through the use of a computer model based upon
quasi-equilibrium chemical relationships.
It is recommended that further study be made of an integrated power system
based upon the Koppers-Totzek gasifier.
Based upon the foregoing comments, the following recommendations are set
forth:
It is recommended that a study be performed to determine the technical,
economic feasibility and environmental impact of an integrated residual oil,
low- and high-temperature cleanup and combined-cycle power system.
It is recommended that a computer model of the coal gasification process
be developed.
Low-Temperature Cleanup Systems
The low-temperature cleanup systems displayed the ability to remove I^S
from the fuel gas to levels of 100 ppm or below. Fuel nitrogen compounds were
also removed as were particulates. However, the utility loads (process steam, .
electricity) were such that overall system efficiency was compromised more
than had been suggested by previous, but perhaps, superficial studies. These
inefficiencies are due in part to the process requirements for the aforemen-
tioned ammonia stripping as well as for sulfur absorber solvent stripping.
Several other additional process heat requirements for sulfur reclamation are
present.
It is recommended that the low-temperature scrubbing processes be reviewed
to determine the effect on overall system performance of modifications to the
process.
High-Temperature Cleanup Systems
The type of inefficiencies associated with the low-temperature process
are not present with the high-temperature processes. This may be in part due
to the early state of development of these processes with its attendant lack
of operational data. The major challenges associated with the high-tempera-
ture cleanup are the removal of fuel nitrogen compounds and particulates. The
262
-------
latter has a broader technological base and has several potential solutions
undergoing early phases of testing. Ammonia removal does not have an equiv-
alent base on which to build.
It is recommended that programs aimed at removing particles of 2jj, to 5(j,
from high-temperature gas streams be initiated or strengthened.
It is recommended that methods of removing ammonia from high-temperature
gas streams be identified and their technical and economic feasibility be
established.
Advanced Power Systems
The advanced power systems used in the present study are based upon two
levels of technology. The first level is that which is being applied to
turbomachinery currently being developed for introduction in commercial ser-
vice in the 1976-78 time frames. The basis for the performance of the first-
generation turbomachinery is the use of air cooling in the hot section of the
turbine, a concept well proven in current industrial and aircraft applications.
The technology for the second generation turbomachinery is based upon the
contractor's judgment of the advancements.that will be made in advanced cool-
ing techniques and materials.
These advancements will occur at rate directly proportional .to the R & D
effort expended, i.e., the funding made available to bring the concepts to the
demonstration phase.
It is recommended that a program be initiated to demonstrate high-temper-
ature gas turbines which would be capable of attaining 50 percent efficiency
in the combined-cycle mode using clean fuel.
Environmental Impact
The present study give quantitative estimates for the emissions of SC^
NO and participates. Of these, only the S02 emissions are based upon actual
data, i.e., measurements of H^S and COS in fuel gas streams. The estimates of
NO,, emissions are based upon models using a mixture of theoretical and empiri-
.X
cal data, the latter being extrapolations of a small amount of low-pressure
testing done on the off gas from an entrained-flow gasifier. The estimates of
particulates are based on the assumption of similarity of fly ash emissions
from boiler and gasifiers.
263
-------
There are potentially several methods other than cyclones to remove
particles from high-temperature streams. The use of panel-bed filters has
been mentioned previously. While none of these devices has been operated at
the conditions considered typical for the gasifier off gas, some small scale
tests have been conducted. In particular, a sand filter is being developed by
the Combustion Power Corporation under partial EPA sponsorship. ^'^ The test
data available indicates a potential for removal of < 2(j, particles, but several
problems remain to be solved before a successful pilot plant could be operated,
much less a commerical-sized installation.
During the course of investigating the high-temperature particulate re-
moval devices, it was discovered that the Stone & Webster Co. was doing essen-
tially the same type of investigation under the auspices of the Electric Power
Research institute (EPRI). A discussion was held with Stone & Webster in
order to establish the status of their work and to determine if any potenti-
ally attractive systems had been overlooked. It was ascertained that the two
studies had reached roughly the same-conclusions, i.e., a complete lack of
actual, reliable data on the particulate carryover from gasifiers and also
on the ability of the turbine to withstand particulate ingestion. In addition,
it was learned that metallic filters capable of 99-5 percent-plus removal of
particles down to 0.5|i had been developed by the Brunswick Corporation. These
materials are used as seals in high-temperature (2200 F) gas turbine engines.
This material, called Brunsmet, has been tested in an 8000 cfm capacity pilot
facility and has operated at conditions indicating an advantage over normal
bag house materials by a factor of 25 (25 times more air can be handled per
unit area). Costs for this system are not available but Brunswick Corporation
claims an economic advantage over conventional materials.
Stone and Webster indicated that small-scale tests of a panel bed filter
at the BOM in Morgantown were favorable, although efficiencies were of the
order of 99 percent rather than 99-9 percent. Problems with the test appara-
tus other than the actual filter could account for the discrepancy according
to Stone and Webster.
From the above, it can be concluded that the technology for particulate
removal down to 2^ does exist although it has not yet been demonstrated on a
scale suitable for subsequent commercialization. It is recommended that tar-
get-scale testing be carried out at conditions typical of gasification systems,
e.g., 300-500 psig, 1^00-1700 F and that design studies for mounting and sup-
port systems for holding panel beds or metal mesh systems be initiated.
261*
-------
It is recommended that further development of the computer model of NOX
formation arising from the use of low-Btu fuel gas in compact combustors be
carried out using a broader base of empirical data from various combustor
programs currently being performed on low-Btu gas.
It is recommended that measurements be made of particulate carryover
from existing gasifiers in order that more realistic estimates of particulate
cleanup can be made, and that more complete information on the total environ-
mental impact of the total systems be obtained.
265
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272
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APPENDIX A
PERFORMANCE EVALUATION PROGRAM
System performance calculations and gas turbine design and sizing esti-
mates were made using programs developed by the contractor for use in corpo-
rate sponsored programs. These are described in this appendix.
State-of-the-Art Performance Program (SOAPP)
The SOAPP Program is a revolutionary and sophisticated computational sys-
tem which can be used to analyze almost any complex power system configuration
consisting of a large number of components. SOAPP could also be used to esti-
mate the design, economic, and environmental characteristics of advanced-cycle
power systems if desired. The SOAPP system is based on a completely modu-
larized representation of system components. Modularization permits consider-
able versatility in,selecting power system configurations to be analyzed and
allows continual update of the system as improved revisions of each module
became available. Many revisions of these preprogrammed modules with differing
degrees of complexity are stored in an extensive SOAPP library.
The central and unique feature of SOAPP is a preprocessor or precompiler
which establishes the sequential logic required for system calculations and
performs all the detailed programming work necessary for any configuration.
Figure 68 presents a pictoral representation of how SOAPP operates. The de-
sired configuration is specified using a simple alphanumeric code, along with
specific module and performance map identification codes. The preprocessor
then selects the corresponding module and map routines from the library and
writes the main control program in FORTRAN which will include all the mathe-
matical logic necessary to provide complete mass and energy balance for the
desired configuration. The control program performs all calculations in proper
sequence and accounts for iterative balances of design requirements, in and out
bleed or extraction streams, external schedules and controls, and transfer of
data between modules.
273
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FIG. 68
OPERATION OF SOAPP SYSTEM
CONFIGURATION MODE INPUT
-------
The configuration flexibility described above was the primary goal of
SOAPP. Other advantages of SOAPP involve user convenience features such as
automatic cycling of run data, data input/output flexibility, and automatic
balancing of constraints imposed on parameters at intermediate locations in
the system.
Gas Turbine Flowpath Design
A high-speed digital computer program was developed under Corporate spon-
sorship to facilitate parametric performance and sizing studies of gas turbine
designs incorporating simple cycles, intercooling, regeneration, and multiple
shafts. The program also provides a realistic assessment of turbine cooling
flow penalty effects on gas turbine thermal efficiency and specific output.
Once the primary independent variables (e.g., material technology, turbine
inlet gas temperature, compressor pressure ratio, and airflow rate) have been
specified, the program computes the gas flowpath, pertinent dimensions, number
of compressor and turbine stages, and turbine cooling flow requirements.
Gas Turbine Cooling Flow Calculations
One of the major problems facing gas turbine manufacturers today is proper
design of cooling systems to restrict metal temperatures to levels low enough
to ensure long life and minimize maintenance problems. Several computer pro-
grams have been developed which can be used to estimate the amount of cooling
airflow required in terms of local hot gas temperature, maximum allowable
metal temperature, and the geometry of the cooling flow passages. One of
these programs was modified.by the contractor to provide output which is com-
patible with SOAPP, especially when low-Btu gasified fuels are used in the
engine.
275
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APPENDIX B
GAS TURBINE COMBUSTOR POLLUTION EMISSION MODEL
This Appendix contains a description of a gas turbine combustor model de-
veloped by United Technologies Corporation for use in predicting exhaust pol-
lutant emissions. The original version of this modelH' divided the combus-
tor into three zones in which liquid droplet burning was modeled. This origi-
nal model was subsequently revised(75-80) to incorporate improved simulation
techniques for the principal aerodynamic and combustion processes in the com-
plex combustor flowfield. In this improved version, aerodynamic properties of
the flow field and concentrations of carbon monoxide, total unburned hydrocar-
bons, and nitric oxide were predicted as a function of radial and axial posi-
tion within the combustor for a variety of current combustor configurations.
The analysis directly considers details of the combustor geometry, the fuel
injection system, and engine operating conditions.
The approach taken in the development of the original combustor model was
(l) to formulate mathematical treatments for the principal physical and chemical
mechanisms that influence the combustion process, and (2) to integrate these
mechanisms through a sequence of thermodynamic states obtained from the cou-
pling of these mechanisms with the physical combustor flowfield equations to
provide the gas temperature, flow velocity, and chemical species concentra-
tions as a function of position within the combustor which, in turn, influence
subsequent combustion. The principal elements of the analysis are a com-
bustor internal flowfield model, a physical combustion model, a treatment of
hydrocarbon-air chemical kinetics, and a NOX kinetics model.
The original P&WA liquid fuel combustor model was subsequently modified
by the United Technologies Research Center, at Corporate expense, to permit
the model to be used for low-Btu gasified fuel. The combustor internal flow-
field model and the NOX kinetics model were basically unchanged. The physical
combustion and hydrocarbon-air chemical kinetics models were revised to simu-
late the combustion and mixing characteristics of low-Btu fuel gas. The re-
mainder of this appendix deals with the highlights of the low-Btu combustor
model with emphasis on revisions made to the original P&WA combustor model.
2?6
-------
Internal Flowfield Model
The combustor flowfield model defines the physical system on which the
gas dynamic and combustion rate calculations are based. The experimentally
determined internal flowfield for a conventional swirl-stabilized, can-type
combustor is shown in Figure 69. The flowfield is seen to include a region
of highly turbulent, reversed flow in the front of the combustor, surrounded
by a region of relatively uniform downstream flow. The forward region, in-
cluding the recirculating zone, is designated the primary zone and the down-
stream region, the secondary or dilution zone. The primary zone serves the
purpose of stabilizing the combustion process; combustion is largely completed
within this zone. The mixture of high-temperature combustion products and
reactants leaving the primary zone continues to burn and subsequently is mixed
with dilution air in the secondary zone to provide a suitable temperature pro-
file for entrance to the turbine.
The combustor flowfield model employed in the present analysis is shown
schematically in Figure 70 for the case of a can-type combustor. The two-
dimensional internal flowfield has been approximated by a set of coannular,
one-dimensional reacting stream tubes. The recirculating zone boundary, en-
closing region 1 of this figure, defines the location and size of a zero net
flow, one-dimensional stream tube representing the recirculating flow. Air
entering the front of the combustor is assigned to the main flow stream tubes
on an equal basis. Downstream combustion and dilution jet air is apportioned
to the stream tubes by means of a jet penetration and mixing model. All wall
cooling air is assigned to the outer stream tube that begins at the first
cooling air addition site. The airflow distribution to combustion and cooling
holes is specified as model input. The stream tube boundaries are defined by
inner and outer radii and are computed as dependent variables. The outermost
stream tube is bounded by the location of the chamber wall, which is also pro-
vided as input.
Physical Combustion Model
The combustion simulation used in the original F&WA liquid combustor
model considered the complex oxidation of hydrocarbon fuel as occurring in
three broad stages. The first stage provides for vaporization of the liquid
fuel and breaking down the complex mixture of long chain hydrocarbons into
light, unburned, partially oxidized hydrocarbons. The subsequent sequence of
reactions, comprising the second stage of combustion, includes the principal
exothermic reactions and produces large amounts of HpO and CO. The final
stage of combustion is characterized by the conversion of CO to C02.
277
-------
FIG. 69
SWIRL-STABILIZED CAN FLOW FIELD MODEL
COOLING AIR
TURBULENT
EXCHANGE
N03-80-3
2?8
-------
FIG. 70
PRIMARY-ZONE FLOW PATTERN OBSERVED IN A CAN-TYPE COMBUSTOR
SECTION B-B
SECTION A-A
N03-80-1
279
-------
An essential feature of the liquid combustion model described above is
the rate equations which govern breakdown of the long chain hydrocarbons. Be-
cause of this, the liquid combustion mechanism is not appropriate with fuel
gas consisting of carbon monoxide, hydrogen, and methane. Hydrogen in the
fuel gas will ignite and burn over a wide range of conditions as soon as it
comes in contact with oxygen in heated air. The temperature rise from the
hydrogen combustion would then cause the methane and carbon monoxide to start
burning. The controlling parameter in fuel gas combustion, therefore, appears
to be the time associated with mixing hydrogen in the fuel gas with oxygen in
the air.
Turbulent mixing has been successfully simulated using eddy coefficient
viscosity models for point-to-point diffusion in a large number of fluid dy-
namic models for which the difference equation grid was sufficiently small
compared to a mean turbulent eddy size. These models have been successful
both for energy and mass transfer. In the Pratt & Whitney combustor model the
axial grid is very fine (Ax = .05 inch); however, the cross-axial grid is
coarse. (The stream tube widths (see Figure 67) are often over an inch.)
Therefore, the cross-axial mixing of hydrogen and air cannot be simulated by
an eddy viscosity diffusion model. Also, an ignition delay mechanism, keyed
to a specified temperature, would not work because temperature is averaged
over a stream tube. Thus, a local hot region, e.g., bordering the recircula-
tion zone, cannot be distinguished from the colder flow in a stream tube. The
average temperature across the stream tube would be well below the ignition
temperature even though ignition had occurred in part of the stream tube. An.
alternative model to eddy viscosity and ignition delay is needed to simulate
the hydrogen-oxygen turbulent mixing process in each stream tube.
Equation U6 offers a simple model of the axial hydrogen-oxygen mixing in
a stream tube which can be adapted to the Pratt & Whitney combustor model.
The time rate of increase of burned vapor (fuel gas, including hydrogen, plus
air) is assumed proportional to the product of the molar concentrations of un-
burned fuel vapor and oxygen. This assumption is analogous to the probability
of collision theory customarily used in statistical thermodynamics'"-'-).
The difference in the argument is that the time scale chosen corresponds
to a typical turbulent mixing time rather than a thermodynamic time scale re-
lated to average particle velocities. Thus:
-A_ [WBV> = [WUV]* [Op]*
dt* i:-
* denotes nondimensional quantities - that is, quantities scaled so as to
be approximately unity.
280
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•where:
[WBV]* = [WBV]/[WBV] average
[WUV> = [WUV]/[WUV] average
[ 02] = [ Q2]/[ 02] average
t* = t/T
and:
[WBV] = moles of (burned fuel gas + air)/unit volume
[WUV] = moles of (unburned fuel gas + air)/unit volume
[Op] = moles of oxygen/unit volume
T = mixing time = L/ /u'^
L = mixing length
/u '2 . = r.m.s. turbulent velocity fluctuation
The mixing time formula given above is commonly used in fluid dynamic
analyses ^°^"°^' . Choosing representative values for the dimensional parame-
ters identified above leads to a more useful form of Equation ^5 given below:
d^TOv] = R [wuv] [o 1
dt 2 •
where :
R
= ICr criP -mole" -sec ~-*
It was found that the simple mixing model embodied in Equation U6 yielded
primary zone lengths which were always in the expected 2 to h inch range for
the FTU combustor over a broad range of fuels and operating conditions. This
gives confidence that Equation h6 does simulate hydrogen-oxygen turbulent
mixing adequately for this combustor model.
* denotes nondimensional quantities - that is, quantities scaled so as to
be approximately unity.
281
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NO Kinetics Model
A
Nitric oxide (NO) is generally the most prevalent oxide of nitrogen ob-
served in the exhaust gas from gas turbine engine combustors. Small amounts
of nitrogen dioxide (Wg), however, are also observed and are thought to be
formed by the oxidation of the previously formed nitric oxide. Consequently,
an evaluation of nitric oxide production within the combustor flowfield essen-
tially constitutes a complete assessment of gas turbine engine combustor-
generated oxides of nitrogen.
Several researchers have been successful in describing the mechanisms
involved in the formation and depletion of the oxides of nitrogen. Of these,
however, Lavorie, et al'35) identified specific reaction mechanisms involving
the principal reactions responsible for the formation and depletion of nitric
oxide. Consequently, the mechanisms of Lavorie (see Table 69) were used in
the oxides of nitrogen concentration prediction system.
Equilibrium concentrations of the species 0, O^, H, OH, and Nj, corrected
to the local carbon monoxide level, if in excess of equilibrium, are assumed
in applying the reactions shown in Table 69. The deviation of the local CO
concentration from its corresponding equilibrium values was assumed to be a
measure of the proximity to thermodynamic equilibrium of all other species
involved. Since relaxation times of the species N and ^0 are typically
several orders of magnitude less than that for NO, steady-state concentra-
tions for N and ^0 may also be assumed. The expression for the rate of NO
reaction thus reduces to (following Lavorie):
= 2(3 -a
dt 1 +« i + K
where:
a =
282
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R^ is the one-way rate of the i reaction; for instance:
R! = Kn [N]e [N0]e
and the subscript "e" denotes equilibrium concentration.
Reactions 1 through 3 in Table 69 correspond to the Zeldovich^ ' mecha
nism as modified by addition of the reaction between the species N and OH.
The reaction rates shown have been updated from the originally reported work
using more recent data(37-39).
283
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Table 69
NITRIC OXIDE FORMATION KINETICS RELATIONSHIPS
Reaction Forward Rate Constant
Number Equation cm3/mole-sec
1 N2 + 0 = NO + N R1 = k = 1.35 x 10ll;e-37,500/T
2 N + 02 = NO + 0 Rg = k12 = 6.U x lo9Te~3,125/T
3 N + OH = NO + H R3 = k13 = 7 x 1011
5 0 + N20 = 02 + N2 R5 = k15 = 5 x 10l3e-lU,000/T
6 0 + N20 = NO + NO Rg = k-j_£ = 2.5 x 1013e-13,i*50/T
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APPENDIX C
POWERPLANT COST ANALYSIS
Previous studys (3, *0 demonstrated the need for a capability to provide
cost estimates for future power systems that not only would enable power sta-
tions with different operating characteristics be considered, but also would
allow sensitivity studies to be undertaken for differing technical characteris-
tics of the selected systems. As a result, members of the technical staff at •
UTRC, in cooperation with cost estimators at Burns & Roe, Inc., developed a
set of correlations based on histrocial data which enable the total installed
cost of COGAS power stations to be estimated in-house with a high degree of
confidence. The procedure consists of a set of equations, grouped by FPC ac-
count numbers, which can be used to calculate the installed costs of major
power station equipment.
Minor component equipment items are combined in logical groups, the costs
of which are also estimated from correlation equations. The calculation pro-
cedure requires as input data which are normally calculated as part of a rou-
tine technical analysis for electric power stations. Examples of such data in-
clude: gross and net station output power; steam turbine output power; gas
turbine output power; cooling water flow rates; subcomponent input power re-
quirements; and station efficiency. Because the cost correlations have been
developed in sufficient detail, the need to resort to expensive preliminary
system layout drawings is eliminated.
As part of the overall analysis, the costs of a few major components are
estimated by methods other than the correlations described since these alterna-
tive methods make it possible to determine these specific equipment costs in
even greater detail. Two such components 'are the large, industrial gas' tur-
bines and waste heat boilers, the costs of which are calculated using sophis-
ticated computer program analyses proprietary to UTRC.
Furthermore, the costs of large, "standard" items such as steam turbo-
generators and electric generators are obtained directly from published catalog
285
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data, which are corrected by appropriate price multiplier factors to maintain
consistency with the latest industry experience.
A detailed list of all system components whose prices are capable of being
estimated directly is presented in Table 70 for all applicable FPC power sta-
tion account categories. Not all items are used for each COGAS system con-
sidered since often there are specific characteristics of each station design
which distinguish it from other stations. In the UTRC procedure, allowances
are also made for such items as station start-up, temporary buildings, trans-
portation during construction, special tools required during a particular
power station contract, engineering, contingency, escalation (except where in-
cluded in the cost of selected major system components), and interest during
construction. Examples of these are shown to the right of each category for
a 1000 Mw, nominally-rated COGAS power station design.
The calculation includes allowances for profit, insurance, and fringe
benefits directly in the cost of each component or group of components. This
approach was followed: (l) because it protects the confidentiality of pro-
gram profitability at the A&E firm; and (2) because changes in specific sta-
tion details, once a set of estimates are completed, could be made easily
without the necessity of changing an entire analysis set. This approach has
been successfully demonstrated ^ ' at UTRC. Its only restriction is that it
cannot estimate directly the effect geographic location has on overall power
station costs. However, this problem can be rectified easily by using other
cost analysis techniques^50)m This is done by equating the sum of the costs
of groups of components identified in the Ref. 50 analysis with the sum of
costs of the same group of components determined in the UTRC analysis.
(UTRC estimates can be grouped into the general categories of materials and
equipment, labor, fringe benefits, supervision,construction tooling, engineer-
ing,, contingency, insurance, and profit.) Segregating power station costs in
this manner then allows the geographic location problem to be addressed di-
rectly through use of the location (labor) factors for major U.S. cities.
The following example illustrates the relationship between the UTRC
procedure and that of Ref. 50. On the right side of Table 70 are presented
costs of the noted components for an illustrative COGAS power station. These
costs, expressed in 197^ dollar values, have been estimated from the UTRC cor-
relations described in this appendix. The costs are presented by account num-
ber in Table 70; to the total have been added allowances for other expenses
during construction, contingency, engineering and supervision, escalation, and
interest (at a rate of 10 percent per annum) during a four year construction
period. For this example, total turn-key power station cost for an Eastern
Seaboard (New York/New Jersey) location is estimated to be slightly more than
286
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Table TO
SPECIFIC COMPONENTS WHOSE COSTS ARE CALCULATED
IN POWER STATION ANALYSIS
FPC No. 312: Boiler Plant Equipment
Waste Heat Boiler
Boiler Feed Pumps
Water Demineralizer
Condensate Storage Tank
Process Steam Heat Exchanger
Miscellaneous Pumps
Piping and Pipe Insulation
Computer and Associated Boiler Plant Controls
$27,591,900
339,390
U, 630, 590
29,9^0
In Acct 3^3
70,525
308,025
Total, Account 312: $36,UU8,100
FPC No. 3lU: Turbine-Generator
Steam Turbogenerator
Pedistal
Condenser
Condensate Vacuum Pumps and Motor
Condensate Pump and Motor
Cooling Tower
Water Circulating Pump, Valves, and Expansion Joints
Make-Up Structure: Screen and Pump
Chlorination Equipment
Lube Oil Purifier
Total, Account
FPC No. 3^1: Site and. Peripheral Structures
Site Preparation
Administration Building
Turbogenerator Building
Condensate Polishing System
Tank Form
Stack
$ 9,833,270
In Acct 3^1
1,212,180
13^,795
27^,780
6,618,750
l,lU6,28o
In Acct
$19,217,055
$ 8U8,250
563,830
3,9^8,000
837,065
1,236,100
362,000
Total, Account 3^1: $ 7,795,21+5
287
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Table TO (Cont'd)
FPC No. 3^3: Gas Turbine
Gas Turbine . $17,61*3,500
Starter Motor 85,600
Torque Converter 80l*,000
Lube Oil Purifier and Storage 22^,000
Lube Oil Fire Protection l60,000
Turbine Air Coolers 320,000
Air Compressor Service and Instrumentation ll*0,000
Breeching . 2,lUo,600
Expansion Joints N.A.
Inlet Air Filters 555,760
Turbine Air Cooler Enclosure N.A.
Emergency Cooling Water Tank, Pump and Piping 11,200
Fuel Oil Heaters and Pumps 123,200
Miscellaneous Pumps and Tanks 56,000
Control Panels ' 560,000
Fuel Piping and Pipe Insulation 1,1*36,900
Cooling Air Compressor and Motor 22,782,^*85
Total, Account 3^3: $1*7,603,2i*5
FPC No. 3kh: Electric Generator for Gas Turbine $10,537,630
FPC No. 3^5: Accessory Electric Equipment $ 8,768,750
FPC No. 3^6: Miscellaneous Power Plant Equipment $ 359,QUO
FPC No. 353: Station Equipment In Acct. No. 3^5
288
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$200 million. The alternative procedure'^0) provides the capability of
estimating costs through the UTRC estimates defined as DIRECT CONSTRUCTION
COSTS in Table 71. At the bottom of Table 71 is presented a breakdown by
general cost category according to the general cost model(50) for two geo-
graphic areas: the noted Eastern Seaboard New York/New Jersey location; and
a North Central (Illinois, Ohio, Pennsylvania) location. In the North Cen-
tral total, the primary change in overall cost is due to the lower, mid-west
labor rate based on labor factors(50). it can be seen that for the eastern
location, DIRECT CONSTRUCTION COSTS are identical to the same total detailed
in the upper half of this table. Because of the different approaches repre-
sented by the two models, it is difficult to make direct cost comparisons.
However, both models make reasonable estimates of the desired station costs,
and therefore focus on the end results is recommended rather than on the de-
tail differences which often are the results of judgmental estimates.
The UTRC calculation procedure has been written in such a manner that
corrections by single or grouped component(s) can be made as costs vary over
time and with inflation rate. The Handy-Whitman Index of Public Utility Con-
struction Costs'"3)5 a regularly updated set of construction cost indexes
recognized throughout the utility industry, has been applied where necessary
to all cost equations in the UTRC procedure. As noted, all cost values are
presented in terms of constant 197^ dollars. However, a projected future
escalation rate must be selected and applied. In such cases, it is advised
that all efforts proceed with extreme caution because of the volatility
associated with projections in the unsettled financial environment confronting
utilities and manufacturing organizations in the U.S.
289
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Table 71
ILLUSTRATIVE EXAMPLE OF ELECTRIC POWER STATION COST ESTIMATING
Station Location: New York/New Jersey Area
Detailed Background Data in Table 70
312
311*
31*3
31*1*
31*5
Other Expenses
Contingency
Engineering and Supervision
DIRECT CONSTRUCTION COST
Escalation
Interest During Construction
Total (Turn Key) Cost
$36,1*1*8,100
19,217,055
7,795,21*5
1*7,603,21*5
10,537,630
8,768., 750
359 .0^0
$130,729,065
2,6ll*,58o
10,667,1*90
20,001,550
$161* ,012 ,685
18,1*1*1,775
18,21*5,1*50
$200,699,910
EPA General Cost Model Approach
NY/NJ Labor
Factor: 1.875)
Cost, 197^ $
Materials + Equipment
Labor, Fringe, Supervision, Admin.
Engineering
Equipment + Tools During Construction
Contingency
Profit + Insurance
DIRECT CONSTRUCTION COST
$81,806,950
35,20^,260
12,271,01*0
13,312,720
17,572,785
$l6U,012,685
111., Ohio, Pa (Labor
Factor: 1.633)
Cost, 197*4 $
$81,806,950
30,660,560
12,271,01*0
3, 81)1*, 930
12,858,350
16,973,020
$158,l*lU, 850
290
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APPENDIX D
DESCRIPTION OF GAS TURBINE COST MODEL
The gas turbine engines to be incorporated in advanced power stations
similar to those identified in this study will be based on design philosophies
different from those engines currently in production for present industrial
applications. Thus, it was necessary to revise and enlarge a computer proce-
dure previously developed'3) to estimate manufacturing costs of future gas
turbine engines. Briefly, this manufacturing cost analysis program was based
on information obtained directly from experts in the field of gas trubine
design and from vendors who supply gas turbine manufacturers. A large number
of individuals were consulted in an effort to obtain a broad data base of man-
ufacturing cost information on a component-by-component basis. Correlations
were then developed which related these costs to particular engine character-
istics, e.g., physical dimensions, material selection, and production volume.
After exhaustive examinations of less sophisticated approaches, it was esti-
mated that the approach selected would provide the best correlations. In the
resulting computer analysis the costs of all calculated components were summed,
an allowance was made for the miscellaneous small engine parts, and assembly
and test expenses were then added. Identifiable major component manufacturing
costs were estimated to constitute over 80$ of the total engine cost. The manu-
facturing cost computer program described was based on vendor contacts and
cost correlations represented in 1970 dollar values.
A separate, Corporate-funded study was undertaken to revise the former
cost correlation equations and to update the applicability of the calculation
procedure to the 197^-1975 time period. This approach involved investigating
the changes in costs for each of the major components during the intervening
four year period. The data obtained were then separated into purchased part,
raw material, and labor index values and were applied where appropriate to
the cost correlations. This approach generally made it unnecessary to alter
any of the original cost correlations, and program flexibility was accordingly
preserved. By exercising the resulting proprietary computer program with input
data representative of the engines identified during the study described in
291
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this report, it was then possible to estimate directly new engine costs which
reflect the latest cost information available from a wide range of knowledgeable
sources.
292
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APPENDIX E
COGAS PERFORMANCE EVALUATION
The evaluation of COGAS performance estimates to identify inefficiencies
associated with the various system components requires the development of an
energy accounting system. In a combined gas and steam turbine cycle it is
important that process heat requirements be satisfied by using gas turbine
exhaust heat rather than by burning fuel or extracting heat from hot fuel gases
prior to their use in the gas turbine. In this manner, the power generating
capacity lost by the extraction of the process heat is only charged at steam
cycle efficiency. The burning of fuel or extraction of heat from hot fuel
gas results in a power loss that is charged at the full combined cycle effi-
ciency. Therefore, -a simple but effective approach to energy bookkeeping is
to reflect fuel and heat consumption of individual components as an -effective
loss in fuel energy. Thus consumption of fuel gas or extraction of heat, from
the fuel gas prior to its use in the gas turbine is charged at the full ener-
gy value. However,.the use of heat from the turbine exhaust is charged only
at the ratio of steam cycle to combined cycle efficiency.
To determine the ratio of these efficiencies it is necessary to look at
the system configured in Figure 6l. In that figure, the gasifier air supply
is separated from the main gas turbine flow stream in order to account for
the heat of compression which is input to the gasifier. The equations that
represent the system are:
POWER GENERATION
PGT = Gas Turbine Output = T)GTLHY HVR
PSTM = Steam Cycle Output - TlgTM 1]BLR Q^ (50)
QEXH ' = HVR Qjjjjy - TlGTffiv HVR QKHV (51)
Electrical Output - (PGT - PCOMP + PSTM) ^GEN (52)
293
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MECHANICAL EFFICIENCY
cc
(HHV) = ^GTHHV + %TM %LR HVR "
NOMENCLATURE
T]GT = Mechanical Efficiency of Gas Turbine (Output includes
HVR = Ratio of Total Energy in Fuel LHV/HHV
(^jjy = Total Chemical Plus Sensible Fuel Energy (HHV) (Above 59°F)
TL™, = Net Steam Cycle Mechanical Efficiency
blJVl
T)BLR = Boiler Heat Removal Efficiency (Above 59°F)
= Heat in Gas Turbine Exhaust (Above 59°F)
PpnMp = Compressor Power for Gasifier Air Supply
T) = Combined Cycle Mechanical Efficiency
cc
= Generator Efficiency
These differ from the usual distillate representation in that they
include the power required for gasifier air compression and are based on the
higher heating value of the fuel. This latter fact requires the addition of
a term that is simply the ratio of lower to higher heating value of the fuel
at its supply temperature.
29 k
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TECHNICAL REPORT DATA
(Please read Instructions on the reverse before completing)
1. REPORT NO.
EPA-600/2-75-078
3. RECIPIENT'S ACCESSION NO.
4. TITLE AND SUBTITLE
Fuel Gas Environmental Impact: Phase Report
5. REPORT DATE
November 1975
6. PERFORMING ORGANIZATION CODE
7. AUTHOR(S)
Fred L. Robs on, Albert J. Giramonti,
William A. Blecher, and Gerald Mazzella
8. PERFORMING ORGANIZATION REPORT NO
9. PERFORMING ORQANIZATION NAME AND ADDRESS
United Technologies Research Center
400 Main Street
East Hartford, Connecticut 06108
10. PROGRAM ELEMENT NO.
1AB013; ROAP 21ADD-027/104
11. CONTRACT/GRANT NO.
68-02-1099
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
Phase; 7/1/73-11/1/74
14. SPONSORING AGENCY CODE
is. SUPPLEMENTARY NOTES Author Mazzella represents Foster Wheeler Energy Corp.
16. ABSTRACT
The report gives results of an evaluation of the technical and economic
feasibility of: (1) Lurgi-type fixed-bed gasifiers and BCR-type entrained-flow
gasifiers in combination with low- and high-temperature fuel gas cleanup systems;
(2) advanced technology combined-cycle power systems: and (3) integrated gasifi-
cation systems, cleanup processes, and power systems. Processes and systems
considered were those using technology both currently available for power station
configurations which the contractor judged could appear in commercial applications
in the 1975-78 time frame (first generation systems) and potentially applicable in the
1980-decade time period (second generation systems). The results indicate that
high-temperature cleanup systems have the potential of improving the efficiency and
reducing the capital costs of integrated gasification systems.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
Air Pollution
Fuels
Gas Purification
Fossil Fuels
Coal Gasification
Electric Power Generation
b.IDENTIFIERS/OPEN ENDED TERMS
Air Pollution Control
Stationary Sources
Fuel Gas
Environmental Impact
Emission Control
c. COSATl Field/Group
13B
2 ID
07A, 13H
10A
. DISTRIBUTION STATEMENT
19. SECURITY CLASS (This Report)
Unclassified
21. NO. OF PAGES
Unlimited
20. SECURITY CLASS (This page)
Unclassified
22. PRICE
EPA Form 2220-1 (9-73)
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