EPA-600/2-76-153
June 1976 Environmental Protection Technology Series
f
FUEL GAS ENVIRONMENTAL IMPACT
industrial Environmental Research Laboratory
Office of Research and Development
U.S. Environmental Protection Agency
^search 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 policy of the Agency, nor does mention of trade
names or commercial products constitute endorsement or
recommendation for use.
This document is available to the public through the National Technical Informa-
tion Service, Springfield, Virginia 22161.
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EPA-600/2-76-153
June 1976
FUEL GAS
ENVIRONMENTAL IMPACT
by
F.L. Robsonand W.A. Blecher (UTRC)
and C. B. Colton (Hittman Associates, Inc.)
United Technologies Research Center
; , Silver Lane
!East Hartford, Connecticut 06108
Contract No. 68-02-1099
Program Element No. EHB529
EPA Project Officer: William J. Rhodes
Industrial Environmental Research Laboratory
Office of Energy, Minerals, and Industry
Research Triangle Park, NC 27711
Prepared for
U.S. ENVIRONMENTAL PROTECTION AGENCY
Office of Research and Development
Washington, DC 20460
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ABSTRACT
The program carried out under EPA Contract 68-02-1099 (from
July 1, 1973} to November 1, 197^) evaluated the technical and economic
feasibility of: (l) fixed-bed gasifiers (Bureau of Mines) and two-
stage entrained-flow gasifiers (BCR) in combination with low- and high-
temperature fuel gas cleanup systems, (2) advanced technology combined-
cycle power systems, and (3) integrated gasification systems, cleanup
processes and power systems. This follow-on extended the study to cover
atmospheric pressure, oxygen-blown coal gasifiers (Koppers-Totzek) and
pressurized, air-blown, partial-oxidation residual oil gasifiers (Shell/
Texaco). Cleanup process modifications were made on paper to improve the
efficiency of the integrated systems. Processes and systems considered
were those using technology currently available for power station con-
figurations which the Contractor judges could appear in commercial appli-
cations in the 1975-1978 time frame (first-generation systems) and those
using technology potentially applicable in the 1980-decade time period
(second-generation systems). The objective of this analysis 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 con-
straints, 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
gasification, cleanup and power systems are evaluated as a whole. A
Contractor-owned digital computer program was utilized to define the per-
formance of the system from coal or residual oil into kilowatts out. The
modular approach to analysis by this unique analytical tool permits wide
flexibility in" fuel process configurations and power cycle arrangement.
However, lack of substantial data on gasifier operation limited the
approach to design point calculations.
111
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The analyses indicate that high-temperature cleanup systems have
the potential of improving the efficiency and reducing the capital costs
of integrated gasification systems. However, unacceptable emission
levels for NOX could result with some gasifier types due to the carryover
of fuel-bound nitrogen compounds. No viable method of removing these
compounds at high temperature was identified. Suitable process modifi-
cations of commercially available low-temperature cleanup systems resulted
in increases in overall system efficiencies which approached those of the
high-temperature systems; but at higher costs. These systems would
still allow generation of electrical power at costs competitive with con-
ventional steam stations with stack gas cleanup while having emissions
which are far below current EPA regulations for solid fuels.
This report was submitted in partial fulfillment of Project No.
21ADD 27 10U, Contract No. 68-02-1099, by the Research Center of the United
Technologies Corporation, under the sponsorship of the Environmental
Protection Agency. This report covers work performed between
November 1, 19?U and November 1, 1975- A previous report, EPA-600/2-75-
078, covers work performed between July 1973 and November 197^-
IV
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TABLE OF CONTENTS
Page
CONCLUSIONS 1
RECOMMENDATIONS 3
INTRODUCTION 5
SECTION 1
REVIEW OF GASIFICATION AND CLEANUP SYSTEMS 8
SUMMARY 8
GASIFICATION SYSTEMS 9
Coal Gasification - Koppers-Totzek 10
Process Description 10
Oxygen Plant 12
Operating Characteristics 12
Process Selection - Partial Oxidation Oil Gasification System 15
Process Description 18
CLEANUP SYSTEMS . . . 20
Low-Temperature Desulfurization Processes 22
High-Temperature Desulfurization Processes . 25
Particulate Removal Systems 29
Evaluation Criteria 36
NOX Control Systems 38
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TABLE OF CONTENTS (CONT'D)
Page
SECTION 2
EVALUATION OF INTEGRATED GASIFICATION AND CLEANUP PROCESSES ^5
SUMMARY ^5
SELECTION CRITERIA 1+5
Low-Temperature Desulfurization Systems Selection h6
High-Temperature Desulfurization System Selection ..... 1+6
SYSTEM EVALUATION 1+7
Comparison of K-T/Selexol and K-T/B+W Cleanup Systems 1+8
Koppers-Totzek/Selexol Process Description 51
OIL GASIFICATION PROCESSES 58
Oil Gasification/Selexol Process Description 59
Oil Gasification/CONOCO Process Description 6l
SECTION 3
REFINEMENT OF INTEGRATED SYSTEMS ?6
SUMMARY 76
REVISED UTILITY REQUIREMENTS FOR AMMONIA REMOVAL 77
FUEL GAS RESATURATION-BUMINES/SELEXOL 77
REVISED S02 REMOVAL FROM BUMINES/IRON-OXIDE 77
Process 1 - Glaus Plant 86
Process 2 - Lime Slurry for S02 Removal and Disposal 87
Process 3 - Catalytic Oxidation Process 87
Process k - Resox Process 97
Process Comparison 97
vi
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TABLE OF CONTENTS (CONT'D)
Page
CATALYTIC CONVERSION OF COS TO H2S IN THE BCR/SELEXOL SYSTEM 106
BCR/CONOCO WITH WATER SCRUB 107
GASIFIER MODELING 132
Parametric Study of the BCR Two-Stage Gasifier 132
SECTION h
PERFORMANCE AND COST OF INTEGRATED SYSTEMS lU3
SUMMARY 1U3
PERFORMANCE
Coal Fired Steam Station
K-T Selexol Integrated System Performance
Oil Gasifier/Selexol Cleanup System Performance
Oil Gasifier/CONOCO Cleanup System Performance l6o
BuMines/Selexol Performance 160
BuMines/Iron Oxide System Performance 165
BCR/Selexol/Catalysis Performance 168
BCR/CONOCO/Water Scrub Performance 170
SYSTEM COSTS 177
Cost of Hot Particulate Removal Systems 177
K-T Selexol System Costs 180
Oil Gasifier/Selexol System Costs 180
Oil Gasifier/CONOCO System Costs l8l
BuMines/Selexol Costs l8l
vii
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TABLE OF CONTENTS (CONT'D)
Page
BuMines/Iron Oxide Costs l8l
BCR/Selexol Costs 183
BCR/CONOCO/Wet Scrub Costs 183
Comparison of Three BCR-Based Systems 183
SECTION 5
ANALYSES OF ENVIRONMENTAL INTRUSION. l88
SUMMARY 188
Existing EPA Standards and Their Implication to This New Point Source . . 188
OVERVIEW 191
AIR EMISSIONS
Review of Fuel Processing System
Emissions Associated with the Fuel Processing Systems 196
Emissions Associated with the Power System 199
WATER EFFLUENTS , 20?
Waste Water Sources 20?
Water and Waste Water Treatment 2l6
Process Description 218
Chemical Treatment of Circulating Cooling Water 229
Boiler Feedwater Treatment 230
Cost Estimates for the Water System of the BuMines/Selexol Process . . . 231
Vlll
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TABLE OF CONTENTS (CONT'D)
SOLID RESIDUALS 231)-
Summary of Solids Produced 23^
Identification of Types of Solids Produced by Water and Waste Water
Treatment 236
Disposal Options and Their Implications ... 237
REFERENCES
APPENDICES
A - Equilibrium Model for Coal Gasifiers 2*4-8
B-Bituminous Coal Research, Inc. (BCR) Two-Stage Gasifier Model . . . 25^
C - Effect of Pressure Ratio on System Performance 263
IX
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LIST OF FIGUEES
Page
Figure 1 - The Koppers-Totzek Gasifier 11
Figure 2 - The K-T Gasification Process 16
Figure 3 - Partial Oxidation of Residual Oil 21
Figure k - Regenerative Iron Oxide Desulfurizer 28
Figure 5 - Sulfur Concentration vs. Temperature 30
Figure 6 - Removal of HUS as a Function of Temperature 31
Figure 7 - Size Distribution of Fines from COED Pyrolysis 33
Figure 8 - Equilibrium Percentage of Mo in a 3:1 Hydrogen-Nitrogen
Gas Mixture kO
Figure 9 - Ammonia Decomposition Over Cu-Ni-ALpOo Catalyst (8.1$ Cu,
9.2 % Ni) k2
Figure 10 - Poisoning of Ni-Catalysts Used for Adjustment of Equilibrium
CO + 3H2 ^ 0% +
Figure 11 - K-T Gasifier with Selexol Cleanup 50
Figure 12 - Process Flow Diagram Oil Gasifier/Selexol Cleanup System 62
Figure 13 - Process Flow Diagram Oil Gasifier/COWOCO Cleanup System 68
Figure ih - Revised Process Flow Diagram BuMines/Selexol 78
Figure 15 - Lime-Slurry Scrubbing Process for SOo Removal from BuMines/
Iron-Oxide Regeneration Off-Gas 92
Figure 16 - Catalytic-Oxidation Process for Removal of S02 From BuMines/
Iron Oxide Regeneration OFF-Gas 9&
Figure 17 - Revised Process Flow Diagram BuMines /Iron Oxide System 100
Figure 18 - Process Flow Diagram BCR/Selexol System 108
Figure 19 - Process Flow Diagram for BCR/CONOCO System 118
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LIST OF FIGURES (CONT'D)
Page
Figure 20 - BCR Gasifier Steam/Coal Ratio Versus Product Gas HHV 135
Figure 21 - BCR Gasifier Air/Coal Ratio Versus Product Gas HHV 136
Figure 22 - BCR Gasifier Steam/Coal Ratio Versus Product Gas HHV 138
Figure 23 - BCR Gasifier Air/ Coal versus Ratio Product Gas HHV 139
Figure 2h - BCR Gasifier Steam/Coal Ratio Versus Product Gas HHV
Figure 25 - BCR Gasifier Air/Coal Ratio Versus Product Gas HHV
Figure 26 - Power Generation Cost Summary
Figure 27 - K-T Selexol/Integrated Power System 151
Figure 28 - Oil/Selexol Integrated Power System 155
Figure 29 - Raw Fuel Gas Chemical Heating Value ' 157
Figure 30 - Residual Oil Gasifier/Selexol Cleanup Performance 159
Figure 31 - Oil/CONOCO/Integrated Power System l6l
Figure 32 - Revised BuMines/Selexol System 162
Figure 33 - Effect of Water Vapor in Fuel Gas BuMines-Selexol System 1.6k
Figure 3^ - Effect of Stack Temperature on BuMines/Selexol Performance 166
Figure 35 - Effect of Dry Fuel Gas Temperature BuMines/Selexol System 167
Figure 36 - BuMines/Sintered Oxide System 169
Figure 37 - BCR/Selexol Power System 171
Figure 38 - Revised BCR/COROCO System 172
Figure 39 - BCR/CONOCO With Water Scrub - Effect of Regenerator Effectivenss
Figure UO - BCR/COWOCO With Water Scrub - Effect of Regenerator Inlet
Temperature 175
XI
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LIST OF FIGURES (CONT'D)
Page
Figure Ul - Effect of Fuel Gas Chemical and Sensible Heat on Combustion
Temperature 200
Figure U2 - Nitric Oxide Formation in Gas Turbine Burner 201
Figure h3 - NOX Production From Combustors Burning Low-Btu and Medium
Btu Gas 20k
Figure hk - The Effect of Equivalence Ratio on NOX Emissions 205
Figure 1+5 - Waste Water Treatment for the Bureau of Mines/Selexol Process 225
Figure k-6 - Flow Chart for Coal Enthalpy Conversion 250
Figure 1+7 - Dependence of Product Gas Temperature on Enthalpy Input Under
Equilibrium 253
Figure h& - Simplified Mass Flow Diagram for the Two-Stage BCR Gasifier 255
Figure ^9 - Breakdown of Mass Flows If Yield is Specified 257
Figure 50 - Methane Yield Versus Partial Pressure of Hydrogen 260
Figure 51 - First Generation System Performance 2.6k
Figure 52 - Second Generation System Performance 265
XI1
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LIST OF TABLES
Page
Table 1 - Typical K-T Gasifier Data 13
Table 2 - Utility Requirements for 2000 Ton/Day 02 Plant 1^
Table 3 - Fuel Gas Composition from K-T Gasifier 17
Table h - Low-Temperature Cleanup Processes 23
Table 5 - High-Temperature Cleanup Processes 2k
Table 6 - High-Temperature, High-Pressure Particulate Removal Systems 37
Table 7 - Comparison of K/T Gasifier with High and Low Temperature
Cleanup 52
Table 8 - Material Balance for K-T/Selexol System 53
Table 9 - Summary of Utilities for K-T/Selexol System 56
Table 10 - Equipment List for K-T/Selexol System 57
Table 11 - Properties of Venezuelan Residual Fuel Oil 60
Table 12 - Material Balance for Oil Gasifier/Selexol Process 63
Table 13 - Summary of Oil Gasifier/Selexol Cleanup Systems 65
Table lU - Oil Gasifier/Selexol System 66
Table 15 - Material Balance for Oil Gasifier/COWOCO Cleanup System 69
Table 16 - Summary of Oil Gasifier/CONOCO Cleanup System 72
Table 17 - Oil Gasifier/CONOCO System Equipment List 73
Table 18 - Revised Material Balance for Bureau of Mines/Selexol System 79
Table 19 - Revised Utilities Summary of Bureau of Mines/Selexol System 83
Table 20 - Revised Equipment List for Bureau of Mines/Selexol System 8^4-
.XI11
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LIST OF TABLES (CONT'D)
Table 21 - Material Balance for Lime Slurry Processes 88
Table 22 - Lime Slurry Process-Operating Cost 90
Table 23 - Lime Slurry Process - Capital Cost • • 91
Table 2k - Material Balance for CAT-OX Process 93
Table 25 - Catalytic Oxidation Process - Capital Cost 9^
Table 26 - Catalytic Oxidation Process - Operating Cost 95
Table 27 - Comparison of Alternative Sulfur Recovery Methods for Bureau
of Mines/Iron Oxide Process 98
Table 28 - Power Cost for Alternative Sulfur Recovery Methods Bureau of
Mines/Iron Oxide System 99
Table 29 - Revised Material Balance for Bureau of Mines/Iron Oxide
System 101
Table 30 - Revised Utilities Summary for BuMines/Iron Oxide System IDk
Table 31 - Revised Bureau of Mines/Iron Oxide System Equipment List 105
Table 32 - Raw Gas Composition - BCR/Selexol System 109
Table 33 - Revised Material Balance for BCR/Selexol System Using Catalytic
COS Removal 110
Table 3^- - Revised Summay of BCR Gasification/Selexol System Desulfurization
Utilities Consumption Ilk
Table 35 - Revised BCR/Selexol System Equipment List 115
Table 36 - Revised Materials Balance for BCR/CONOCO System 120
Table 37 - Utilities Summary of Revised BCR/CONOCO System with Water
Scrub 126
Table 38 - Revised BCR/CONOCO With Water Scrub System Equipment List 127
xiv
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LIST OF TABLES (CORT'D)
Table 39 - Summary of BCR Operating Conditions 13*4-
Table ho - Integrated Systems Performance Summary ihh
Table Ul - Power Generation Cost Summary
Table U2 - Coal-Fired Steam Station Capital
Table U3 - Coal-Fired Steam Station Power Generation Cost Summary 150
Table hk - Performance Comparison K-T/Selexol and BuMines/Selexol 153
Table ^5 - Effect of Steam Addition on Fuel Gas Chemical Heating Value 158
Table U6 - BCR/CONOCO-Performance Effects of Water Scrub 173
Table U? - Gasifier and Cleanup System Capital Cost Breakdown 178
Table U8 - Power System Capital Cost Summary 179
Table U9 - BuMines/Selexol Cost Summary 182
Table 50 - BuMines/Iron Oxide Cost Symmary l8U
Table 51 - BCR/Selexol Cost Summary 185
Table 52 - Compairson of BCR-Based Integrated Systems 187
Table 53 - Summary of Residuals from Integrated Systems 189
Table 5^ - Air Emissions from Integrated Systems 197
Table 55 - Combustion Temperatures for Fuel Gas 202
Table 56 - Chemical Characteristics of Process Condensate 210
Table 57 - Trace Element Analysis of Illinois Coal 211
Table 58 - Chemical Characteristics of Process C^ndensates and Potential
Control Systems ' 217
Table 59 - Potential Control Technology for Coal Conversion Waste Water 219
xv
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LIST OF TABLES (COWT'D)
Page
Table 60 - Raw Water Analysis 226
Table 6l - Water Balances for the Bureau of Mines/Selexol Process 227
Table 62 - Capital and Annual Operating Costs of a Water System for
the BuMines/Selexol Process 232
Table 63 - Constituents of Coal Ash 235
Table 6k - Comparison of Results from Koppers-Totzek and UTRC Gasifier
Models 252
xvi
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ACKNOWLEDGEMENT
The support work of Mr. W. R. Davison, Dr. E. B. Smith, and Mr. S. J.
Lehman of United Technologies Research Center is acknowledged with sincere
thanks.
The efforts of the Hittman Associates, Inc. carried out under the
direction of Mr. C. B. Colton with the assistance of Mr. M. Dandavati and
Dr. T. Goldschmid provided an expanded coverage of the environmental intrusions
of the complex integrated plants.
The technical direction and patience offered by Messrs. W. J. Rhodes
and T. K. Janes of the Industrial Environmental Laboratory, EPA, are acknowl-
edged with deep appreciation.
xvi i
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LIST OF CONVERSION FACTORS
Btu x 0.252 = Kcal
ft x 0.30^8 = m
in. x 5.+ = mm
F subtract 32 x 0.555 = C
Ib x 0.1+53 = Kg
Ib x 0.1+53 = Kg
scf (@ 60 F & 30 in. Hg) x 0-0281+ = m3 (@ 15.5 c & 762 mm Hg)
Btu/scf x 8.88 = Kcal/m^ (@ 15.5 C & 762 mm Hg)
lb/10 Btu x 1.798 = Kg/10 Kcal
ton x 1.10U = metric ton
xvi 11
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CONCLUSIONS
1. High-pressure gasification of high-sulfur coal and residual oil followed
by either low-or high-temperature cleanup offers the potential for lower
emission levels than conventional coal-fired systems with flue gas desul-
furization as well as the ability to recover sulfur in elemental form.
2. The addition of gasification and cleanup equipment to a conventional steam
system would increase the unit capital cost per kW by more than 50$.
3. Inefficiencies and utility requirements associated with coal gasification
and cleanup processes can effectively reduce the efficiency of power con-
version by as much as 15 to 25 percent over a comparable clean fuel fired
steam power system.
k. Production of electrical power from coal using a combined-cycle generation
system integrated with a low-Btu gasifier and fuel gas cleanup system can more
than offset the inefficiencies attributed to gasification and cleanup processes,
Overall efficiencies some 10 to 15 percent better than a conventional coal-
fired steam plant with flue gas desulfurization appear realistic while
offering the potential for improved sulfur removal capability.
5. By virtue of the low unit cost of gas turbines and the ability to use a low-
pressure steam system at relatively high condenser pressure for the bottoming
cycle, capital costs of the integrated low-Btu gas/combined-cycle system
that are some 10 percent less than for a conventional system are possible.
6. The combined effect of performance and capital cost advantage for a second-
generation integrated combined-cycle/low-Btu gasification system results
in a cost of coal-derived power as much as 20 percent less than for conven-
tional systems with flue gas desulfurization. These estimates assume 1975
prices and do not include development costs.
T. Currently available gas turbine and gasifier technology do not offer any
economic incentive over conventional coal-fired systems. However, antici-
pated advances in gas turbine technology (2600 F turbine inlet temperature)
and gasifier performance (BCR two-stage slagging unit) offer improved perfor-
mance, better emission control, and lower cost then conventional systems.
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CONCLUSIONS CONT'D
8. Some performance and cost improvement can result from the use of
high-temperature sulfur removal systems. However, these processes
do not remove ammonia if present in the fuel gas and preclude (due
to high temperature) premixing of the air/fuel mixture to reduce
nitrogen oxide production.
9. Where ammonia is present in the raw gas, the use of a low-temperature
water scrub for its removal negates virtually all of the perfor-
mance advantage associated with high-temperature cleanup.
10. Significant reductions in the cost of low-temperature cleanup
processes and utility load (for the system studied) can be
achieved by catalytic conversion of COS to H?S prior to cleanup.
11. Although fuel availability is uncertain, integrated partial
oxidation (e.g., Shell or Texaco process) of residual oil/sulfur
removal/combined-cycle systems offers higher performance and lower
capital cost then their coal-fired counterparts. There is also
' a potential for reduced emissions of S02, NO , and particulates
compared to coal-based units.
12. Integrated systems based upon the partial oxidation of coal at
atmospheric pressure (using Koppers-Totzek as an example) were
not competitive in either cost or performance with those systems
operating at pressure.
13. The water-borne and solid effluents from integrated systems can
meet proposed 1980 standards using best available technology.
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RECOMMENDATIONS
1. The potential economic and performance benefits resulting from the
high-temperature cleanup processes warrant early developmental
efforts to establish their operating characteristics on a scale
large enough to more realistically assess their technical and
environmental viability in commercial applications.
2. In order to fully utilize the potential advantages of the high-
temperature cleanup processes, such as the CONOCO half-calcined
dolomite process, it is recommended that an investigation be under-
taken to identify methods of removing fuel-bound nitrogen compounds
at elevated temperatures.
3. Experimental verification of the catalytic hydrolyzation of COS
under conditions of typical fuel gas streams should be carried out.
k. The performance, cost, and environmental effects of the two generic-
type gasifiers not studied in this work, i.e., pressurized fluid bed
and molten salt gasifiers, should be investigated. This would allow
comparable assessments to be made of the potential of the basic
gasifier types, e.g., fixed bed, entrained flow, fluid bed and
molten salt types.
5. Particle size distribution in the raw fuel gas from the various
gasifiers should be experimentally obtained under a variety of
operating conditions. Only when these, data are available will it
be possible to make realistic estimates of turbine requirements
and particulate removal 'system units.
i
6. Although investigations of integrated power systems carried out
under other sponsorship, e.g., Energy Research and Development
Administration and the"Electric Power Research Institute, have
resulted in conclusions similar to those resulting from studies
carried out under the EPA and its predecessor organizations, con-
sideration should be given to better future coordination of efforts
to ensure truly comparable results.
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7. Future work in gasification should recognize the specific needs of
power generation via low-Btu gas and establish the practicality of
operation under conditions favorable to the performance of generating
plants.
8. The ability to utilize high-temperature fuels without excessive NOX
production is dependent on the ability to premix the air and fuel
or to provide very rapid mixing in the primary combustion zone to
avoid thermal NOX production. Further analytical and test work is
needed to establish the limits of fuel temperature that are consis-
tent with allowable NOX production.
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INTRODUCTION
The prior investigations carried out for the EPA under the first
three phases of this Contract, number 68-02-1099» produced the major
conclusion '!' that integrated power systems consisting of coal gasi-
fiers/sulfur cleanup process/combined-cycle power systems have the
potential for generation of electric power at costs competitive with
or, in some cases, less than those of conventional coal-fired stations
with flue gas desulfurization (FGD). Those integrated systems having
a high-temperature sulfur removal process, i.e., a process projected
to operate at or near gasifier exit temperatures, had higher perfor-
mance and lower costs than the integrated systems using commercially
available low-temperature cleanup systems.
These results are in general agreement with those of other inves-
tigations. The NASA, with the cooperation of ERDA, EPRI, NSF, and
the Office of Management and Budget (OMB), is sponsoring the Energy
Conversion Alternative Study (EGAS), a multiphase effort being carried
out by teams headed by General Electric and Westinghouse. At the end
of the Phase I screening study of many alternative energy conversion
systems, both contractors had identified the integrated low-Btu gas/
combined-cycle power system as having the greatest potential attractive-
ness. '2,3) Similiarily, an EPRI-sponsored study(^) described the
attractive performance of the integrated power systems.
While the integrated low-Btu gas/combined-cycle power system is
undeniably attractive, there are several areas in which the studies of
Phases 1, 2, and 3 indicated a need for expansion or where further empiri-
cal verification or analytical definition is required. These are addressed
in the Phase U work which is reported here. One of the needed areas of
expansion is the inclusion of gasifiers other than the fixed-bed and two-
stage entrained flow types considered in the earlier phases. Since it
appears that development of advanced gasifiers, characterized by the
absence of tar in the off-gas and by ability to operate with a wide variety
of coals, may be the pacing technology, it would be worthwhile to investi-
gate processes which have this capability. One such process is the Koppers-
Totzek (K-T) single-stage atmospheric pressure gasifier using oxygen and
steam rather than air and steam as the suspension media. An investigation of
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this process would put into perspective its applicability to integrated
systems as well as identifying the effluents from an integrated system
having an oxygen-blown gasifier. A second gasifier type, somewhat
similar to the K-T process is the partial oxidation gasification process
widely used in the refinery industry for production of Eg from liquid
feed. This process is sometimes called the Texaco process or the Shell
process, since both of these energy companies have developed variations
' of it. A prior study1^' had identified the partial oxidation process using
coal as being most attractive for use with combined-cycle power systems,
but integration of the gasification cleanup and power system was minimal.
Additional data in the depth necessary for this study exist only for an
oil-based process, thus a system based upon the use of heavy residual
oil was selected for evaluation. This has the added benefit of allowing
the environmental effects of oil-based systems to be placed in context
with the coal-based systems.
A second area requiring more detailed investigations was that of
sulfur cleanup. The processes identified in the earlier phases, i.e., the
low-temperature physical absorption (Selexol), the high-temperature iron
oxide (Bureau of Mines) and the high-temperature half-calcined dolomite
(CONOCO, formerly CONSOL), were representative of their respective genre
but suffered from some shortcomings in their application to the overall
system. These processes were further refined by appropriate modifications
to reduce utility requirements and to better utilize the heat available
from the processes. These modifications achieved as much as 10 percent
increase in system efficiency.
The third, and perhaps most significant area needing further inves-
tigation, was that of system effluents. The prior work had quantified
the major system air effluents, but had not attempted to detail the
water and solid effluents. Thus, a more in depth analysis of this area
was carried out in this Phase h program.
Thus, this study was conducted to broaden the coverage of fuel
processes, to refine the integration of the gasification/cleanup/power
system, and to further identify the potential environmental intrusion
of these systems. Section I of this report contains the description of
the operating characteristics of the atmospheric pressure, oxygen-blown,
entrained-flow coal gasifier (the K-T process) and a single-stage,
pressurized partial oxidation of residual oil gasifier (Shell/Texaco
process). This section also contains a brief review of the sulfur and
particulate cleanup systems, both low- and high-temperature, associated
with the gasification processes. In Section 2, the selection criteria
for cleanup systems are reviewed and their applications to the K-T and
partial oxidation gasifier are analyzed. Detailed flow sheets of the
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integrated gasifier/cleanup system for the additional gasifiers using
both low-and high-temperature cleanup are included. Section 3 contains
descriptions of methods to improve system performance. Here are
described modifications to the cleanup processes for those gasifiers
previously studied as well as a discussion of the effort to develop a
computer model of an entrained-flow gasifier for purposes of evaluating
performance and effluents. In Section U, the performance and cost of
the integrated power systems are presented and comparison to previous
results, where applicable, are given. Lastly, definition of the
effluents from the integrated power systems is given in Section 5 , and
the status of cleanup technology is identified.
Throughout the report, reference is made to the Bureau of Mines
(BuMines) stirred bed gasifier and sintered iron oxide cleanup system.
During the initial phases of this contract, that gasifier was selected
as being representative of fixed (as opposed to air entrained) bed
gasifiers. Since that time, the Morgantown Energy Research Center,
where the gasifier is under development, has become a part of ERDA.
However, for the sake of consistency with the previous Phase Report11 '
the name "BuMines Gasifier" or BuMines Iron Oxide Cleanup System",
have been retained and are used in this report.
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SECTION I
REVIEW OF GASIFICATION AND CLEANUP SYSTEMS
SUMMARY
The review of individual systems presented in this section is intended
to complement the data presented in Ref. 1. Characteristics of the
Koppers-Totzek oxygen-blown, atmoshperic pressure gasifier and a pres-
surized air-blown redisual fuel oil gasifier (Shell or Texaco) are pre-
sented. By virtue of their high operating temperaure, nitrogen compounds
are virtually eliminated fron the resultant product gas. This is
especially helpful if used with a high-temperature cleanup process. While
high-temperature cleanup processes tend to be quite efficient, one of
the associated major problem areas that has been identified is the lack
of a suitable high-temperature, ammonia removal process. As a result,
nitrogen compounds are passed through to the combustor where the major
portion form nitric oxide.
• Cleanup systems are reviewed and the B&W Iron Oxide system is dis-
cussed in some detail. Also, potential solutions for two problem areas
associated with high-temperature cleanup systems are examined. These
are particulate and nitrogen compound removal. In the case of particu-
lates, a system consisting of cyclones and filters is selected for use
with high-temperature sulfur removal systems. While no suitable high-
temperature ammonia removal system has been identified, an iron oxide
sulfur removal system could potentially act as a catalyst for ammonia
decomposition. However, it was found that catalysts presently available
for ammonia decomposition would be poisoned by the sulfur compounds pre-
sent in the gas stream.
Throughout the report, reference is made to gasifier efficiency as
a rough means of comparison. It is important to note that these effi-
ciencies are not used in cycle calculations because both gasifier input
and output are a combination of chemical and sensible (or latent) heat
and it is not possible to account for the differences in temperature
8
-------
level or availability in a single parameter such as efficiency. This is
especially true when the gasifier is mated with a combined-cycle gener-
ating system and sensible heat, unless used to heat fuel to the gas
turbine, can only be utilized at steam cycle efficiency which is about
30 percent as opposed to combined cycle efficiency which is on the order
of 50 percent. The other factor that is omitted from efficiency is
auxiliary power which can be significant when considering systems such
as the K-T where it is necessary to produce oxygen for use in the gasi-
fier.
Nevertheless, the efficiency numbers do provide a means of prelim-
inary comparison and where used, they are as defined here:
1. Cold Gas Efficiency
Y = Chemical heating value of Gas
Chemical heating value of coal
2. Hot Gas Efficiency
Y = Chemical plus sensible (above 80 F) heating value of gas
Chemical heating value of coal
GASIFICATION SYSTEMS
This section contains a review of an atmospheric pressure, oxygen-
blown coal gasifier (Koppers-Totzek) and a pressurized air-blown residual
oil gasifier (Shell/Texaco). Process descriptions are included and
operating characteristics are defined for the coal gasification and oil
gasification systems. The gasifiers "were evaluated with respect to
their output for a fixed input of coal or oil, and their thermal effi-
ciency. The first generation gasifier, typified by the fixed-bed type
(e.g., Bureau of Mines or Lurgi) has off-gases containing condensible
tars, phenols, and other organics. One second-generation gasifier has
been investigated previously. This was the entrained-flow, BCR two-
stage gasifier. While this gasifier operated at a temperature high
enough to crack tars etc., the fuel-bound nitrogen compounds were not
cracked and considerable amounts of these compounds would pass through
the high-temperature sulfur cleanup processes and could result in high
NO emissions from the power system. Thus, it is of great interest to
investigate other second generation gasifiers which have the potential
for low fuel-bound nitrogen production. While both the K-T and partial
oxidation gasifiers are commercially available, they can be considered to
be "second generation" in that they produce no condensible tars.
-------
Coal Gasification System - Koppers-Totzek
Process Selection — The Koppers-Totzek process was selected for review
and for integrationinto a power plant. This process was chosen because:
(1) process operating data and information on effluents are available;
(2) the process is commercially proven; and (3) the process has low
carryover of fuel bound nitrogen. The process is flexible with respect
to feeds, and a prior study'"' has indicated that the gasification por-
tion is relatively clean.
Process Description
A schematic diagram of the Koppers-Totzek (K-T) gasifier is shown
in Fig. 1. The K-T gasifier can be operated on coal of different ranks.
Depending on the rank, the coal is dried to between two percent and eight
percent moisture content and pulverized to about 70 percent through 200
mesh. The coal is conveyed with nitrogen from storage to the gasifier
service bins. Controls regulate the intermittent feeding of coal from
the service bins to the feed bins, which are connected to variable speed
coal screw feeders. The pulverized coal is continuously discharged into
a mixing nozzle where it is entrained in oxygen and low-pressure steam.
Moderate temperature and high burner velocity prevent the reaction of
the coal and the oxygen until entry into the gasification zone.
The oxygen, steam, and coal react in the gasifier at a slight posi-
tive pressure and at 3300F, to produce intermediate-Btu (300 Btu/SCF)
fuel gas. The overall gasification process is endothermic primarily due
to the steam-carbon reaction which requires about 5000 Btu/lb C:
C + HgO - CO (g) + H2 (g) (1)
A portion of the coal feed is burnt with oxygen to provide this heat:
C + 02 - C02 (g) (2)
The carbon and volatile matter of the coal are gasified, and the
coal ash is converted into molten slag. Approximately 50 percent of this
slag drops into a water quench tank and is carried from the tank to the
plant disposal system as a granular solid, while the remainder is en-
trained in the gas exiting the gasifier. Low-pressure steam for the
gasifier reaction is produced in the gasifier jacket from the heat
passing through the refractory lining.
10
-------
THE KOPPERS- TOTZEK GASIFIER
FIG 1
HOT FUEL GAS TO BOILER
WATER-COOLED SHELL
PULVERIZED COAL,
STEAM, AND OXYGEN
SLAG
POOL OF WATER
WATER-COOLED
BURNER
76-02-68-2
-11-
-------
Table 1VI/ shows typical gasification data for eastern and western
, m, ,, „„. . (chemical heating value of gas) „ ,,.
coals. The cold gas efficiency ^(chemicai heating value of coal) of thls
process is from 70 to 77 percent. Additional energy is recovered as
high-pressure steam from the available sensible heat in the gas giving
a hot gas efficiency in excess of 90 percent. Very little heat is
required for gasifier steam, and oxygen is fed at low temperature. The
gasifier performance is greatly affected by the power cycle (discussed
in Section ^) and the need to make oxygen.
Oxygen Plant
The oxygen required for the gasification process is produced in an
on-site oxygen plant. Gaseous oxygen for gasification can be supplied
by an oxygen plant based on the low-pressure (100 psig) air separation
cycle principle. In this process, air is compressed to about 100 psig,
cooled by heat exchange with counter-current flow of cold oxygen and
waste gas streams, and finally distilled in high- and low-pressure dis-
tillation columns. Product oxygen is removed as gas from the low-pressure
column at a typical purity of 99.5 percent. The capacity of the biggest
single-train oxygen plants to date is about 2000 tons/day\ '. The
utility requirements for a typical 2000 tons/day oxygen plant'°' are
given in Table 2.
The requirements in Table 2 are for.an oxygen plant supplying Q^ at
250 psig. The base-load oxygen compressor can be eliminated sine it is
for compressing the product oxygen to 250 psig, but the oxygen for the
Koppers-Totzek gasifier is supplies at essentially atmospheric pressure.
The power requirement, therefore, drops to 27,983 kW, for a 200o tons/day
plant. This is equivalent to approximately 0.17 kWhr/lb of 02.
Oxygen purity has an insignificant impact on the power requirements.
For example, reducing the $2 purity from 99-5 percent to 98 percent,
reduces the power required by about three percent. Further reduction of
purity to 90 percent results in about eight percent power reduction from
the 99-5 percent purity case.
Operating Characteristics
The major operating characteristics associated with the K-T process
are:
1. Versatility — The process is capable of continuous operation for
the gasification of a variety of feedstocks, including all ranks of
solid fuels. Coal size is not a limiting factor and caking coals can be
handled without pretreatment.
12
-------
72.7
5.3
1.1
1.0
9.0
8.9
2.0
69.9
U.9
1.3
l.l
7.1
13-7
2.0
TABLE 1
TYPICAL K-T GASIFIER DATA
TYPE OF FUEL WESTERN EASTERN
COAL COAL
Gasifier Feed
Dry feed to Gasifier
Analysis, Wt$
C
H
N
S
0
Ash
Moisture
100.00 100.00
Higher heating value of dry feed,
Btu/lb 13,135 12,6^40
Oxygen, tons/ton dried feed
@9Q% purity 0.878 0.8^9
Process steam, Ib/ton dried feed 8lU 810
Gasifier Products
Jacket steam, Ib/ton dried feed 600 55^
High-pressure steam, Ib/ton dried
feed ©900F/900 psig 2760 2675
Raw gas analysis, vol%, dry
CO 52.55 52.51
C02 10.00 10.00
H2 36.09 35.96
N2 + argon 1.00 1.15
H2s 0.3*4 0.36
COS 0.02 0.02
100.00 100.00
Dry gas make-SCF/ton dried feed 69,690 ' 66,970
Higher heating value, Btu/SCF, dry 287 286
Heating value of gas/heating value feed, % 76.1 75-8
-13-
-------
TABLE 2
UTILITY REQUIREMENTS FOR 2000 TON/DAY 02 PLANT
Electrical Power Kv^
Main air compressor 27,500
Base load oxygen compressor 8,700
Water wash tower pump 270
Instrument air drier 50
Expander lube oil pump 27
Liquid oxygen circulating pump 18
Main air filter purge blower 18
Lighting, instrumentation, and misc. 100
Total Kw 36,683
Cooling Water
Circulating rate: l6,Uoo gpm at 50 psig, 85F
Steam
30 psig, saturated plant steam is required
intermittently as follows:
Use Steam requirements
Reactivation rich liquid filters 2000 Ib/hr for 8 hr/wk
Reactivation guard adsorber 1000 Ib/hr for 8 hr/wk
Plant derime 2500 Ib/hr for 60 hr/yr
-Ik-
-------
2. Simplicity of construction and ease of operation — The only moving
parts at the gasifier are screw feeders for solids. Control of the
gasifiers is achieved primarily by maintaining COp concentration in the
clean gas at a reasonable constatn and predetermined value. Slag
fluidity may be visually monitored. Gasifiers display good dynamic
response.
3. Moderate capacity — K-T units are designed for coal feed rates up
to 850 tons/day, or for a production of about ^5 x 10^ SCF/day of 300
BTU/SCF gas. This is equivalent to a nominal 80 Mw(e) to 100 MW(e)
per unit.
Due to the high operating temperature (3300-3500F), the K-T process
produces slag. No condensible hydrocarbons, phenols, pyridines, or
other organics are produced. Ammonia and cyanide are produced in amounts
well under one volume percent.
A schematic of the K-T process is shown in Fig. 2. A typical com-
position' '' of the gas including trace impurities at the gasifier outlet
is shown in Table 3.
The composition in Table 3 is for a gasifier coupled with a gas
quench section in which the gas is sprayed with water. The exit tempera-
ture is typically 2200 F but may be varied depending upon the character-
istics of the ash carried in the fuel gas stream. The purpose of the
quench section is to bring the ash to a temperature below its softening
point to avoid sticking to the waste heat boilder surfaces.
Process Selection - Partial Oxidation Oil Gasification System (Shell/Texaco)
The partial-oxidation oil gasification process was also selected
because of the availability of data, because it is a proven commercial
process, and because of its low fuel nitrogen. The process was origi-
nally developed by Hydrocarbon Research, Inc. (HRl) in the 1950's and
subsequently commercialized by the Shell Oil Co. and the Texaco Co. A
large number of installations operating on this principle produce synthe-
sis gas for ammonia manufacture, methanol synthesis, refinery use, etc.
The process operates over a wide range of pressures and hence the syn-
thesis gas could be available at high enough pressure for firing in a
combined cycle.
15
-------
K-T GASIFICATION PROCESS
STEAM DRUM
FROM COAL
PULVERIZER
DESULFERIZED
GAS TO TURBINE
LEAN ABSORBENT
SPRAY
WATERs
FROMs/
CLARIFIER
TO SLAG
QUENCH TANK
I
MAKE-UP TO SCRUBBER SYSTEM
FROM SLAG
QUENCHTANK
CONDENSATE (IF ANY)
Jj
Cl
NJ
-------
TABLE 3
FUEL GAS COMPOSITION FROM K-T GASIFIER
Component
CO
C02
0%
H2
H2S
COS
HCN
M3
Ar
so2
WO
Particulates (gr/SCF)
Volume percent
37-36
7.13
0.08
25.17
0.30
0.23
178 pprav
288 ppmv
0.17
29.19
0.32
22 ppmv
7 ppmv
11-57
-17-
-------
Process Description
In the partial oxidation scheme, residual fuel oil or sour crude is
partially burned in noncatalytic reactors (gasifiers) to provide suffi-
cient heat to maintain a high temperature for the gasification reaction.
The gaseous product is composed primarily of hydrogen and carbon monoxide, .
some carbon dioxide, hydrogen sulfide, small amounts of residual methane
and soot. The soot can amount to one to three percent of the feed car-
bon(10).
The process chemistry can be represented by the following reactions:
CnHra +()0 -» nCO + ()H,
Thermal cracking occurs with partial oxidation of heacy hydrocarbons
and forms free carbon.
CnHra -> nC + (7)^2 (M
Free carbon formation results in reduced gas production. Other
reactions that occur are:
CnHra + (n + 2)0 -> nC02 + (5})H20 (5)
Cnllm + 3 C00 •> I'nCO i (!")!!., ('•)
<- i . i..
CnHra + (51)0- -> nC + (5)H90 (?)
When steam is used in the gasification, the endothermic steam-carbon
reaction occurs:
CnHra + n H20 -> nCO + (| + n)H2
Some relatively slow secondary reactions that occur are:
C + C02 -» 2 CO (9)
C + HQ0 -* CO + H2 (10)
18
-------
Normal residence time is insufficient for the completion of
reactions (9) and (10); therefore, some soot is always present. For
heavy fuel oil, soot may be as high as three percent of the feed carbon.
Under the operating conditions (> 2300F) some shift of carbon monoxide
to carbon dioxide takes place.
CO , + HpO -» COp + Hp
(11)
Typical fuel oil and exit gas composition data^ ' are shown below:
Heavy fuel oil Weight %
c 85.70
H 10.73
c 2.65
N + Ash 0.60
0 0.32
Total 100.00
Gas composition (Dry) Volume percent
H Ml.60
c6 U8.30
co2 ^.60
H2S 0.60
COS 300 ppm
CH^ 0.50
N2 0.60
Ar 0.80
Total 100.00
Pressure, atm. 88.0
When the partial oxidation gasifier is used in conjunction with a low-
temperature cleanup system, the hot gas is cooled and the heat is
recovered in a waste heat boilder. Prior to acid gas removal, the gas
is cleaned and it temperature is further reduced in a water scrubber
where the soot is removed from the gas in a carbon/water slurry. To
recover the carbon, two methods are generally used. One uses an inter-
mediate fluid, naptha, which preferentially wets the carbon and the
naptha-carbon mixture can then be separated from the water stream. The
19
-------
naptha-carbon phase is mixed with fresh oil feed and flashed into the
naptha stripper. The natpha stripper separates naptha for reuse leaving
carbon in the oil feed for recycle to the gasifier. This process was
developed for use with very heavy oils where the carbon could not be
transferred directly to the oils and has the disadvantage of requiring
a considerable amount of steam for naptha stripping. A schematic of
another variant of the process as commercialized by Shell is shown in
Fig. 3- In this version, soot is recovered from the water slurry in a
"pelletizer" and mixed with the feed for recycle to the gasifier. The
small amounts of water introduced into the fuel in this process could be
troublesome where the oil must be heated to high temperatures causing
the water to vaporize. However, operation at pressure offers a potential
solution to the problem and may, in fact, be beneficial in that the
entrained water may help in the atomization of the heavy oil.
CLEANUP SYSTEMS
Fuel gas cleanup systems consist of particulate removal systems,
sulfur removal systems, and systems to remove fuel-bound nitrogen com-
pounds (mostly ammonia). Depending ontheir operating temperature, these
can be divided into two broad categories, viz.:
(1) Low-temperature cleanup systems
(2) High-temperature cleanup systems
Low-temperature systems require cooling of the dirty gas to 250 F or
below, whereas high-temperature systems require little or no cooling
of the dirty gas.
A survey of low- and high-temperature desulfurization processes was
done as part of Phases 1 through 3' ' of this contract study. Several
low-temperature desulfurization processes are commercially available,
and have been widely used for natural gas sweetening, and for treating
synthesis gas in the chemical process industry. Ammonia and methanol
manufacture are examples of such applications. To facilitate a compari-
son between the various cleanup processes, a number of criteria were
developed. These criteria are:
(1) Type of absorbent
(2) Operating temperature
(3) Operating pressure
Efficiency of sulfur removal
20
-------
PARTIAL OXIDATION OF RESIDUAL OIL
WASTE HEAT BOILER
HIGH-PRESSURE STEAM
TO TURBINES
CLEAN GAS
IV)
I—'
I
01
o
10
I
I
M
CARBON SLURRY
SEPARATOR
CARBON-FREE
CIRCULATION
WATER
WASTE
WATER
Zk
AIR FROM GT COMPRESSOR
Zk
RESIDUAL OIL
P
CO
-------
(5) Absorbent characteristics
(a) life of absorbent
(b) regenerability
(c) selectivity toward sulfur compounds over C02
(d) makeup rate
(6) Form of sulfur recovery
(7) Status - commercial, developmental, conceptual
The above criteria were used to compare several desulfurization
processes. Tables h and 5 are a listing of low- and high-temperature
desulfurization processes, respectively. It is seen that the most
effective desulfurization processes are those which have a high effi-
ciency of sulfur removal, high selectivity toward E^S, can handle large
volumes of gases (up to 1000 MMSCFD) containing 0.2 to 1.0 volume per-
cent BUS, are easily regenerable, and have low energy requirements.
Low-Temperature Desulfurization Processes
Several low-temperature processes have been described in the pre-
vious Phase Report^1'. For convenience, a listing of various types are
given in Table h and a brief discussion of the operational characteris-
tics is included.
The processes in Table h are best-suited to operate at 250 F or
below and hence require cooling of the dirty gas. Low-temperature pro-
cesses can be subdivided into four categories, according to the principle
of operation. These are:
(l) Chemical solvent processes
(2) Physical solvent processes
(3) Direct conversion processes
(k) Dry bed processes
Chemical Solvent Processes - These processes employ aqueous solutions of
organic and/or inorganic agents to scrub the "dirty" gas. These agents
are capable of forming "complexes" with I^S, COp, and other acid gas
components present in the raw gas stream. The "complex" is then decom-
posed during regeneration at elevated temperatures, thereby releasing
the acid gases for further processing and recovery. The regenerated
solution is recycled for further absorption. These processes may be
subdivided into those based on amine scrubbing solutions, and those
based on alkali scrubbing solutions. These processes generally exhibit
little or no selective absorption of H2S over COj.
22
-------
Table 4
Low temperature cleanup processes
Basis: 8400 tons/day Illinois No. 6 Coal Fed to OCR Gasifier, or 6700 ppm of Influent H2S
Process
Chemical
solvent type
1. ME A
2. DEA
3. TEA
4. Alkazid
5. Benfield
6. Catacarb
Physical
solvent type
7. Sulfinol
8. Salexol
9. Rectisol
Direct
conversion
10. Stret-
ford
11. Town-
send
Drybed type
12. Iron
sponge
Absorbent
Monoetha-
nolamine
Oiethanol
amine
Trietha-
notamine
Potassium
dimethyl
amino .
acetate
Activated
potassium
carbonate
solution
Activated
potassium
carbonate
solution
Sulfolane
+
Dilsopro-
panoamine
Polyethyl-
ene glycol
ether
Methanol
Na2CO, +
anthraquin
one sul-
fonic acid
Triethylene
glycol
Hydrated
Fe,0,
Type of
Absorbent
Aqueous
solution
Aqueous
solution
Aqueous
solution
Aqueous
solution
Aqueous
solution
Aqueous
solution
Organic
solvent
Organic
solvent
Organic
solvent
Alkaline
solution
Aqueous
solution
Fixed
bed
Temp.
0 F
80 to
120
100 to
130
100to
150
70 to
120
150 to
250
150 to
250
80 to
120
20 to
80
<0
150 to
250
70 to
100
Pressure
Insensitive
to variation
in pressure
Insensitive
to variation
in pressure
Insensitive
to variation
in pressure
Insensitive
to variation
in pressure
1 - 80 atm
Insensitive
to variation
in pressure
generally
> 300 psi
High
pressure
preferred
Efficiency of S Removal
%H, Sin-
fluent
99
99
99
99
99
99
99
99
99
99.9
99.9
99
Effluent
H,S
ppm
-100
-100
-100
-100
HjS.
+ COS
-100
H,S
+ COS
-100
H,S
+ COS
-100
H,S
+ COS
-100
-100
-10
-10
H,S
+ COS
-100
Absorbent
Characteristics
Life
Unlim-
ited
No
degra-
dation
Regenera-
tion
Thermal
Thermal
Thermal
With
steam
With
steam
With
steam
Low
pressure
heating
or with
steam
Selectivity
toward
Forms non-
rcgen. comp.
with COS,
CS,
Absorbs CO, ,
does not
absorb
COS, CS,
H,S
H,S
H,S
is high
H,S- par-
tial also
absorbs
COS, CS,
H,S.and
also absorbs
COS, CS,
and mer-
captans
H,Salso
absorbs
COS
H,S
H,S.
H,S
Hg S and
also towards
COS, CS,
and mer-
captans
Make up
rate
50 to
100%
<5%
<5%
<5%
50 to
100%
Form of
Sulfur
Recovery
AsH.S
gas
AsH,S
gas
AsH,S
gas
As H,S
gas
AsH,S
gas
AsH,S
gas
AsH,S
gas
AsH,S
gas
Elemen-
tal
sulfur
Elemen-
tal
sulfur
Elemen-
tal
sulfur
Status
Commercial
Commercial
Commercial
Commercial
Commercial
Commercial
Commercial
Commercial
Commercial
Commercial
Commercial
23
-------
Table 5
High temperature cleanup processes
Basis: 8400 tons/day Illinois No. 6 Coal Fed to BCR Gasifier, or 6700 ppm of Influent H2S
Process
1. Bureau
of Mines
2. Babcock
and
Wilcox
3. CONOCO
4. Air prod-
ucts
5. Battelle
North-
west
6. IGT -
Me is-
sner
Absorbent
Sintered
pellets of
Fe,O3 (25%)
and fly ash
Fe-03
Half calcined
dolomite
Calcined
dolomite
Molten
carbonates
(15%CaC03)
Molten metal
(proprietary)
Type of
Bed
Fixed
bed
Fixed
bed
Fluidized
bed
Fixed
bed
Solution
Splashing
contact
Temp.
°F
1000 to
1500
800 to
1200
1 500 to
1800
1 600 to
2000
1100 to
1700
900
Pressure
Insensitive
to variation
in pressure
Insensitive
to variation
in pressure
~200psia
H2S removal
is high at
low pressure
Insensitive
to variation
in pressure
Atmospheric
H, S removal
is high
at low
pressure.
5-6 psig
Efficiency of S
Removal
%H2 Sin-
fluent
-95
-99
-95
-95
-98
Effluent
H2S
ppm
~350
~75
-350
-350
~150
Absorbent
Characteristics
Life
>174
cycles
Wt loss
<5%
mini-
mum
5-6
cycles
Regener-
ation
With air
10-13%
with
steam
and CO,
80-90%
with
steam
and CO,
With
steam
and CO,
Elec-
troly-
tic
Selec-
tivity
toward
H,S,
COS
H,S,
COS
H,S,
COS
H,S,
COS,
fly ash
H,S,
COS
Make up
rate
<5%
1%of
circula-
tion
rate
Form of
Sulfur
Recovery
As SO,
gas
As
1 2-1 5%
SO, gas
AsH,S
gas to
Claus
process
AsH,S
gas to
Claus
process
AsH,S
gas to
Claus
process-
Energy
Required
Elec.
kw
96.360
9830
Oth-
er
stu
Status
Pilot
Experi-
mental
Pilot
Aban-
doned
Pilot
Concep'
tual
-------
Physical Solvent Processes - These processes use organic solvents to
remove acid gases by physical absorption, rather than by chemical reac-
tion. The extent of absorption is directly proportional to the partial
pressure of the acid gas comp nents. These processes are best-suited to
high-pressure gas treating where appreciable quantities of acid gases
are present. The solvent is then regenerated by heat and/or pressure
reduction, thereby releasing a concentrated stream of acid gases and a
recyclable solvent. These processes exhibit a selective absorption of
HpS over CC^. In addition to removing E^S and COp, these processes are
all capable of removing COS, CS2, and mercaptans without solvent degrada-
tion.
Direct Conversion Processes - These consist of two types of processes.
(1) Those based on oxidation reduction reactions
(2) Those based on the stoichiometric reaction of KpS with SCU in
the presence of a solvent.
In the first type, HoS is absorbed in an alkaline solution contain-
ing oxidizing agents. The HpS is then oxidized to elemental sulfur by
air feeding to the regenerator and the sulfur product is separated from
the regenerated solution by froth flotation. Partial removal of COS,
CSp, and mercaptans is also possible.
The second group of direct conversion processes are those in which
HpS is absorbed in a solvent and converted to elemental sulfur by the
Glaus type reaction with SOp.
(12)
The solvent are usually aqueous soltuions of organic or inorganic agents.
Dry Bed Processes - These are based on absorption of acid gases by a
fixed bed of solid absorbent. Due to their low absorbent loading, they
are best-suited to removing small quantities of acid gases. These pro-
cesses can be subdivided into the historical iron oxide processes and the
various molecular sieve processes.
High-Temperature Desulfurization Processes
Several high-temperature desulfurization processes are currently
under development. A survey of these processes was made, and these are
listed in Table 5. None of the processes has been commercialized;
-------
however, the Bureau of Mines sintered iron-oxide process and the half-
calcined dolomite process of the Consolidation Coal Company Coal Develop-
ment Center, a part of the Continental Oil Company (CONOCO) are rela-
tively advanced in their development and may result in commercialization
sooner than the others.
The principle underlying high-temperature desulfurization is the
formation of metal sulfides by chemical reaction of the absorbent with
sulfur compounds in the gas at high temperatures. The extent of sulfur
removal depends on the chemical equilibria for the particular system.
All of the high-temperature desulfurization processes with the exception
of the IGT-Meissner Process and the Babcock and Wilcox Process, have
been described in the Phase Report^ '. The Air Products Process is also
discussed here.
IGT-Meissner Process - This process is being developed by the Institute
of Gas Technology in conjunction with its U-Gas Process. This process,
still in the conceptual stage, utilizes a splashing molten metal-gas con-
tact to remove H^S from the gas. The contact results in the formation
of a metal sulfide which is then decomposed electrolytically to release
IL^S and regenerate the molten metal for recycle. The operating tempera-
ture is 900 F and a high sulfur removal efficiency (98 percent) is pro-
jected. The molten metal absorbent is proprietary. The estimated costs
and energy requirements given in Table 5 are preliminary^ '. Further
development is being directed toward establishing mass transfer rates.
Babcock and Wilcox Process - This process is similar to the Bureau of
Mines process in that it utilizes iron oxide to remove HpS from the gas
at high temperatures. The difference lies in the material used by the
two processes. While the Bureau of Mines' process uses a sintered
material made from iron oxide and fly ash, the Babcock and Wilcox process
starts out with carbon steel and generates an iron oxide scale on the
steel surface which is then used as the desulfurization agent. Briefly,
the process chemistry is described by the following reactions:
Fe/FeOv + H,S FeS + H,0
X d. X
(13)
At some point in time all of the available iron oxide scale is converted
to the sulfide scale. At that point, the system is regenerated with air
as follows :
Fe/FeS + Air — - Fe/FeO + XSO^
26
-------
The overall process accomplishes two things :
(1) It concentrates sulfur from the raw gas to 10 to 13 volume
percent SC^ in the regenerant gas.
(2) It provides S02 in the regenerant gas that is either oxidized
to sulfuric acid or reduced to elemental sulfur.
A sulfur removal efficiency greater than 90 percent is projected.
Absorption can be carried out at temperatures as low as 675 F; however,
higher temperatures are desirable for effective regeneration. If regen-
eration is performed below 1000 F, the sulfide is oxidezed to FeSO^ and
not to FeOx. Also, higher temperatures help activate the surface by
developing a thick iron oxide layer over which effective absorption
occurs. Hence, operation is usually at temperatures in excess of 1000 F.
The concentration of HgS in the desulfurized gas increases as the
volume of gas desulfurized on a given iron oxide scale increases. There
fore. the hardware design for desulfurization and regeneration is one
(1) Has a large number of compartments at various stages of regen-
eration to give an average ^S concentration in the fuel gas
relatively independent of the regeneration cycle
(2) Gives a maximum S0 concentration in the regenerant gas
The hardware that has been designed uses a number of compartments for
sulfur removal and the so-called countercurrent principle of air regen-
eration. The desulfurizer uses a modified regenerative type air heater
and is referred to as the "regenerative desulfurizer", a schematic of
which is shown in Fig. ^-. The cylindrical unit is segmented into 16 com-
partments. Each compartment is filled with carbon-steel plates oriented
longitudinally with the gas flow. The vessel itself would be constructed
from high alloy steel.
Desulfurization-step - The sulfided iron surface is converted back to
the oxide in three of the 16 compartments. The regeneration air passes
in and upward in the first compartment to a cross-over, then downward
for a second pass, and upward for a third and final pass. At two revo-
lutions per hour, each of the 16 compartments is regenerated twice per
hour. Air enters the first regeneration compartment where it contacts
a partially regenerated surface accomplished in the second and third
pass downstream. At the end of the first pass, the 02 concentration is
27
-------
FIG. 4
REGENERATIVE IRON OXIDE DESULFURIZER
(B&W)
REGENERANT GAS
AIR
RAW GAS
PRODUCT GAS
28
76-02-4-2
-------
well below 21 percent. During the second pass, the Q^ concentration is
further reduced while S0? increases. Purging the third (most FeS 'fouled)
compartment with a gas containing a minimum concentration of C>2 and a
maximum concentration of SO^ insures a maximum SOo concentration of the
final regenerant gas. The regenerant gas should contain 10 to 13 volume
percent SC^.
The effect of reaction temperature on the sulfur concentration in
the desulfurized gas is shown in Fig. 5- This may also be represented
by the amount of E~S removed as a function of the temperature. This is
referred to as "sulfur pickup" in Fig. 6.
The process concept has been demonstrated on bend scale equipment
and a hardware design has been developed. The process has yet to be
demonstrated on a large scale.
Air Products Process - This process, now abandoned, employed a fixed-bed
of fully calcined dolomite to absorb I^S from the raw gas. The sulfided
dolomite was then regenerated with steam and carbon dioxide before being
recycled to the absorber. Poor regenerability of the sulfided dolomite
led to the abandonment of this process by the Air Products Company.
Particulate Removal Systems
Particulates of varied sizes, shapes, and composition are a major
contributor to air pollution, are a healt hazard, and are the target of
statutory limitations. In addition to their effect on human healt,
particulates adversely affect pollution control efforts by fouling
catalysts used for SCL reduction, sulfur recovery, NHo decomposition,
etc. Particulates in fuel used for firing gas turbines may cause
erosion/corrosion of turbine blades. The need for particulate removal
from gas streams where they are present in significant quantities cannot
be overemphasized.
Only particulates from fixed-bed, fluidized bed, and entrained bed
coal gasification are considered here. The primary differences between
the three gasification methods lie in:
(l) The manner in which the coal feed is supported
(2) The rate of gas flow (superficial velocity)
(3) Temperature
(k) Feed size
29
-------
FIG. 5
SULFUR CONCENTRATION VS TEMPERATURE
0.05
0.04
o
2
O
O
<& 0.03
i
tr
LLJ
0.02
0.01
SPACE VEL = 2000-2500 V/V-HR
END POINT = 0.10% SULFUR
600 700
800
900
1000
1100
1200
TEMPERATURE F
76-02-4-4
30
-------
13.0
LLJ
O
<
LL
cr
CO
LL
O
CM
O
O
O
LU
12.0
11.0
10.0
9.0
co
CM
I
O
CO
O
8.0
7.0
6.0
5.0
4.0
3.0 ~
FIG. 6
REMOVAL OF H2S AS A FUNCTION OF TEMPERATURE
14.0
END POINT: 0.10% SULFUR
SPACE VEL: 2000-2500 v/v HR
2.0
600 700 800 900 1000
TEMPERATURE F
1100
1200
76-02-4-3
31
-------
Gas flow rate is the least in a fixed bed and greatest in an
entrained bed. The maximum temperature that can be used in any par-
ticular bed depends on the caking properties of the coal fed, the sur-
face area of the coal particles, etc. The surface area is much higher
in fluidized and entrained beds than it is in fixed-beds, and therefore,
higher temperatures can be employed. Temperatures can generally be
raised as the fluidizing/entraining velocity is raised, and hence, are
highest in entrained beds and lowest in fixed-beds. Fixed-beds use
coarser feeds, whereas fluidized and entrained beds use finer feeds.
Considering the three types of beds, individually, and using tem-
perature, gas flow rates, and feed size as parameters, the following
qualitative analysis can be made of the particulate quantity and size
distribution from each.
Fixed Beds - With relatively low gas flow rates, the coarser (heavier)
particles tend to settle down and the finer particles are carried away
by the gas. These gasifiers operate at temperatures (up to 1500 F)
that are lower than ash slagging temperatures. Therefore, the ash
does not slag and agglomerate. This leads to a higher ash loading in
the gas than if the ash were to slag and agglomerate. The coarse feed
(~l/5")j on the other hand, tends to reduce particle entrainment to
some extent. The net effect of the three factors, therefore, would be
to yield a fairly high particulate loading of finer particles comprised
of ash and possibly unburned carbon.
Fluidized Beds - Gas flow rates here are higher than for fixed-beds and
there is a greater tendency for the bigger particles to be carried away
by the gas. Particulate loading is also increased by the finer feed
size used in fluidized beds (minus 200 mesh). However, temperature
has an opposing effect in gasification, up to a point. The temperatures
used in some fluidized bed gasifiers are higher (2000 F) than typical
ash softening temperatures, causing the ash to agglomerate. This
effect reduces the amount of ash carried over with the gas. The net
effect therefore, would be to yield a slightly reduced particulate
loading of comparatively larger particles than from a fixed-bed gasifier.
This is illustrated by the particle size distribution curves (Fig. 7)
for fluidized-bed pyrolysis used in the COED Process*' '. It should be
noted that conditions in the four pyrolysis stages to which the curves
correspond, change gradually from fixed-bed type conditions in the first
stage to fluidized-bed type conditions in the fourth, with a correspond-
ing increase in average particle size going from the first stage to the
fourth.
32
-------
SIZE DISTRIBUTION OF FINES FROM COED PYROLYSIS
FIG. 7
(a) FIRST STAGE
(b) SECOND STAGE
100
UJ
N
55 90
Q
uj
HI
>
O
CO
a.
I-
I
80
70
60
50
H
m 40
DC
30
20
Qj 10
5
2 5 10 20
MICRONS DIAMETER
50
HI
N
co
Q
UJ
CO
LLJ
O
CD
O
tr
a
UJ
100
90
80
70
60
50
40
30
20
10
0
1 2 5 10 20
MICRONS DIAMETER
50
(c) THIRD STAGE
(d) FOURTH STAGE
UJ
2 5 10 20
MICRONS DIAMETER
CO
Q
LLJ
CO
UJ
>
O
CO
•2.
LLJ
O
a:
UJ
a.
I
a
100
90
80
70
60
50
40
30
20
10
0
2 5 10 20
MICRONS DIAMETER
50
76-02-4-6
33
-------
A quantitative approximation of particulate loading and size
distribution for fluidized bed gasification was obtained by Westinghouse
researchers using a theoretical model developed by Kunii and Levenspiel^ ''
and by comparison with the corresponding figures for fluidized bed com-
bustion. Projected operating conditions for fluidized-bed combustion
were deduced (by Westinghouse researchers) from experimental data
obtained from the National Coal Board (NCB) of England and from EPA
contractors. The projected conditions for fluidized-bed combustion
and fluidized-bed gasification are given below.
Operating Conditions Fluidized-Bed Fluidized-Bed
Combustion Gasification
Pressure (atm) 10 to 20 10 to 20
Temperature, F 1600 to 1800 1500 to 1?00
Gas flow (Ib gas/
Ib fuel) -12.5 -5.5
Projected dust
loading prior to
gas cleanup
(gr/scf) 10 to 30 10 to 30
Projected particle
size 10 to 2% <10n 10 to 2%
5 to 15$ <5^i 5 to J5%
The projected figures for fluidized-bed gasification were derived by
Westinghouse from the corresponding figures for fluidized-bed combustion
and the theoretical model of Kunii and Levenspiel taking into considera-
tion the following differences:
(1) The gasification process has less than one-half the gas flow;
(2) Fluidizing velocity in the coal gasification system design is
about h ft/sec, which is one-fourth to one-half that for
fluidized-bed combustion;
(3) Ash will essentially be concentrated and removed in the gasi-
fication process as agglomerates, and not carried out of the
system in the fuel gas.
-------
Based on a qualitative assessment of the differences in the
combustion and gasification systems noted above, it is reasonable that
a particulate removal system which is designed to handle the dust loading
from a pressurized fluidized bed combustion process that feeds a com-
bined cycle or let-down turbine will be able to cope with that from the
fluidized bed gasification process.
Entrained Beds - These have high flow rates and operate at high tempera-
tures (to 3000 F) with fine feed sizes (minus 200 mesh). In this flow
regime the entire feed could be carried over as with the BCR two-stage
process, and therefore, the particulate loadings could be very high with
the particle size distribution proportional to the feed size. Carryover
in single-stage systems is not well documented although values for the
K-T gasifier after water spraying have been given previously.
Cleanup Requirements - In all three coal gasification systems discussed
above, the particle size distribution and particulate loading are
affected by several other factors such as the ash content of the coal,
the method of coal preparation, the change in particle density due to
chemical reactions, attrition in the bed, and description of particles
due to their history. The effects of these factors cannot easily be
quantified.
Allowable particulate loadings and size distribution that consider
gas turbine requirements as well as air pollution emission regulations
have to be considered for the design of particulate removal equipment.
As a reference point in considering emission levels, the allowable
particulate loading in low (125 Btu/SCF) Btu gas corresponding to cur-
rent coal standards of 0.1 Ib/million Btu of coal is .11 grains/SCF.
From an operational standpoint, allowable loadings are significantly
less. Based on data presented in Ref. 1, the following tabulation pre-
sents current fuel specifications along with a suggested specification
for low Btu gas.
P&WA1 P&WA2 Westinghouse GE1 Low-BtuGas
Loading 0 . 08 lb/L06 f t3 0.01gr/Ft3 O.OOO^gr/ft? 30ppm 0.0012gr/ft3
Size UOp, mas 2^ to lOji . > 2^
1. Aircraft derivative burning methane
2. Projected for high-temperature turbine
35
-------
Evaluation Criteria
The criteria used for evaluating different particulate removal
systems are given below:
(1) Capability to attain particulate loadings governed by gas
turbine specifications at high gas temperatures (1500 to
l800 F) and pressures (10 to 20 atm).
(2) Demonstration of capability: performance, reliability, and
life in commercial or pilot plant operation.
(3) Compatibility: pressure drop, operating pressure, and tem-
perature with the coal gasification processes.
(k) Capital, operation, and maintenance costs for a commercial
system.
Many low-temperature particulate removal systems are commercially avail-
able and descriptions of these are abundant in the literature. Only
high-temperature systems have been considered here. Several high-tem-
perature and high-pressure particulate removal systems that satisfy
some or all of the above criteria are either available or currently
under development. Table 6 lists the different types of high-temperature
particulate removal systems and where available, gives the operating
conditions and efficiencies attainable with each.
It has been stated previously^ ' that the tolerance of the gas
turbine for particulates is not well documented. The consensus in the
industry that barring significant changes in design philosophy, partic-
ulates of < 2|j, can be accepted. It is assumed that particles of this
size and smaller will remain in the gas stream rather than impinge on
turbine surfaces.
Unfortunately, from an applications standpoint it is the particu-
lates in the l}j, range which cause the visible emissions problem. Thus,
it is possible to meet turbine requirements with respect to particle
size, yet have a power system which appears to be "smokey".
The application of the devices given in Table 6 to the removal of
particulates from the high-temperature, high-pressure fuel gas stream,
while not fully defined, does give hope that at least turbine specifica-
tions and, hopefully, environmental regulations can be met. It would
appear that the use of several stages of cyclones, followed by the
-------
Table 6
High temperature particulate removal systems
Type of removal
system
Mechanical
Collectors
Cyclones
Tornado
Bed Filters
Granular
Panel
Rex
Sonic
Agglomeration
Collection
Systems
Alternating
Velocity
Precipitator
Scrubbers
Fused salts
Filters
Metal and
ceramic
Electrostatic
Precipitators
Manufac-
turer
Bucll
Ducon
Aerodyne
Combustion
Power Co.
Ducon
C.U.N.Y.
Rexnord
Braxton
BatteMe
Solas and
others
Capacity
ACFM
50,000
58,000
30,000
20,000
Collection
efficiency
%
80-90
80-90
93-97
>90
95-99
>99
>99
Minimum
particle
size with
efficiency
> 50%, n
5-10
5-10
0.5
2
<0.5
<0.5
Maximum
operat-
ing
temp. °F
1400
1500
1500
1400
> 900
2000
800
Maximum
operat-
ing pres-
sure atm
2
10
10
1
1
1
Maximum
collection
efficiency
%
90-95
90-95
98
>90
>99
>99
>99
Applicable
dust load-
ing range
grains/SCF
<30
<40
Pressure
drop
in.W.G.
4-40
4-40
30
10-15
4-15
<1
Status
Commercial
Commercial
Commercial
Under
Development
Under
Development
Under
Development
Commercial
Under
Development
Under
Development
Commercial
Commercial
-------
Aerodyne-type tornado and finally a metallic or ceramic filter could
achieve 95 to 99 percent removal of < 2^ particles. There remains a
good deal of testing and development, but this aspect of high-temperature
fuel gas cleanup is in a technological state comparable to the more
developed high-temperature desulfurization processes such as the iron-
oxide process.
In the analyses to follow, it is assumed that a high-temperature
cleanup system consisting of cyclones and filters is used. Since the
energy consumption of these is quite small, even when used in series
(< 1 percent of gasifier exit pressure) and there are little other
utility requirements or systems interfaces, the particulate removal
device is assumed not to affect performance and is represented only in
the cost of equipment.
CONTROL SYSTEMS
Nitrogen oxides, collectively referred to as NO, are an important
group of air pollutants. The term NOX refers primarily to NO (nitric
oxide), although similar quantities of NOo (nitrogen dioxide) and ^0
(nitrous oxide) may also be formed. These oxides are interconvertible,
and the equilibrium between them depends on photochemical reactions,
the presence of oxidizing agents, etc. Gas turbines, like other
combustion engines form NOx in the hot combustion zones of engines.
There are two known mechanisms responsible for NOV formation in combus-
.A
tion engines:
(l) Thermal NOX: NOX formed by the reaction of atmospheric Ng
and 02 in the hot combustion zone within the engine. This
is the dominant mechanism when relatively clean fuels are
burnt in the engine. Removal of NOx from flue gases is an
extremely difficult problem. However, it is possible to con-
trol the thermal NOX formation by several techniques, some
of which are:
(a) Off-stoichiometric combustion by modified combustion
chamber design
(b) Water injection
(c) Exhaust gas recirculation
Each of the above techniques results in a lower peak temperature within
the combustion zone, thereby reducing thermal NOX formation.
-------
(2) NO., from fixed nitrogen in the fuel: This source of NCL is
A A.
important only when nitrogen-bearing fuels such as those
derived from coal and residual fuel oil are burned. Dirty
fuels may contain organic nitrogen compounds which are oxi-
dized during combustion to NOX. Gasified fuels, especially
those from gasifiers operating below 2000 F, contain combus-
tible nitrogen compunds such as ammonia, hydrogen cyanide, and
pyridine. Ammonia is the primary nitrogen compound, while the
others are in smaller concentrations. If retained in the gas,
these compounds are oxidized during combustion to NQx. These
compounds are removed by water scrubbing when a low tempera-
ture cleanup system is used. However, when a high-temperature
cleanup system is used, these nitrogen compounds are carried
through to the turbine. To prevent this carry- through,
ammonia and other nitrogen compounds must be removed from the
gas at elevated temperatures.
A potential method to remove ammonia is to decompose it into stable
elemental nitrogen and hydrogen, at elevated temperatures. The decom-
position of ammonia is governed by the following reaction:
kl
2NH3 ?2 N2 + 3H2 («)
The equilibrium constant kj increases as the temperature increases and
the total pressure decreases. Temperature has a greater effect than
pressure as is verified by the fact that k^ is practically constant over
the pressure range, 1 to 50 atmospheres. Hence the higher the tempera-
ture, the greater is the decomposition of ammonia into nitrogen and
hydrogen. The equilibrium percentage of ammonia in a 3:1 hydrogen-
nitrogen gas mixture is shown as a function of pressure and temperature
in Fig. 8^16).
kn
The equilibrium constants for ammonia formation, k^ = _± are tabu-
k2
lated below^-"-'' for a range of temperatures and pressures.
Pressure, atm
Temp. F
660
750
8Uo
10
0.0266
0.0129
0.00659
30
0.0273
0.0129
0.00676
50
0.0278
0.0130
0.0069
39
-------
EQUILIBRIUM PERCENTAGE OF NH3 IN A 3:1 HYDROGEN-NITROGEN GAS MIXTURE
100
o>
o
to
I
100
200
300
400
500
600
700
800
900
1000
PRESSURE ATMOSPHERES
P
oo
-------
The equilibrium constant for ammonia formation is seen to be very low at
high temperatures and low pressures. Under the conditions at the exit
of the BCR gasifier (1750 F, ^75 psia) the ammonia concentration is
1*260 ppm' ''. If allowd to equilibrate, the ammonia composition is
reduced to ^60 ppm while the temperature, due to other reactions, is
reduced to approximately 1600 F. If the only reaction were ammonia
decomposition, the temperature change would be much less and the resul-
tant equilibrium ammonia concentration would be even lower (on the
order of 200 ppm). This suggests that low equilibrium concentrations,
although thermodynamically favored, are not kinetically feasible. There-
fore, the kinetics must be aided by a catalyst.
A literature search yielded considerable recent work in the develop-
ment of catalysts for ammonia decomposition. General Motors^ 'and
Ford Motor Co.* "' have addressed the ammonia decomposition problem with
the goal of ridding automotive exhausts of poisonous nitric oxide, by
first reducing it to ammonia and further decomposing the ammonia to ele-
mental nitrogen and hydrogen. Research in this direction led to the
development of several catalysts suitable for ammonia decomposition.
Among these are Ni, Pt, W, Mo, Re and Ru. Of these catalysts, a Cu-Ni-
ALpOo catalyst was seen to have the highest activity for ammonia decom-
position. The extent of ammonia decomposition over Cu-Ni-AlpO_ catalyst
as a function of temperature, is shown in Fig. 9-
All the above catalysts proved to have a serious drawback. They
are poisoned by even trace quantities of sulfur compounds present in the
feed gas\20j. The poisoning is due to the formation of a metal-sulfide
which deactivates the catalyst. Generally the metal/metal-sulfide
equilibrium favors the formation of the metal-sulfide at low temperatures
and favors its decomposition at high temperatures. As an example, the
poisoning by sulfur compounds of Ni catalyst used in methanation reac-
tions is shown in Fig. 10. The poisoning effect is seen to diminish
only at high temperatures (~ 2000 F). Effectively this is the tempera-
ture above which, for example, a Cu-Ni-Al20o catalyst could be used to
decompose ammonia. It is impractical, however, to use the catalyst
at this temperature for two reasons:
(1) As such high temperatures, sintering would significantly
reduce the activity of the catalyst.
(2) In all the gasification systems considered in this study where
fuel gas ammonia is significant, fuel gas temperatures approach-
ing 2000 F are unattainable. Gasifiers that operate at tem-
peratures in excess of 2000 F will probably not produce
appreciable amounts of ammonia.
-------
AMMONIA DECOMPOSITION OVER Cu-Ni-AI2O3 CATALYST (8.1% Cu, 9.2% Ni)
1000
800
-p-
ro
Q.
Q.
o
<
600
O
O
400
200
100
200
300
400
500
600
700
800
at
o
I
01
CATALYST TEMP C
P
(O
-------
FIG. 10
POISONING OF Ni - CATALYSTS USED FOR ADJUSTMENT OF
EQUILIBRIUM: CO + 3H2 T*~ CH4 + H2O
100.0
10.0
to
o
o
cc
CO
co
<
tr
o
1.0
0.01
200
400 600 800
TEMPERATURE, C
1000
1200
76-02-4-7
-------
The only commercial catalyst potentially capable of decomposing ammonia
is an iron oxide catalyst composed of five percent FepCL mounted on
high-temperature, fired inert alumina spheres. This catalyst could
simultaneously remove ILS from the gas stream, thus combining the ammonia
and sulfur removal operations in a single step. However, the operating
conditions necessary for this catalyst are not known and must be deter-
mined before the catalyst becomes acceptable(21'. An iron oxide system
thus has the potential for catalyzing the ammonia decomposition while
removing sulfur compounts. Unfortunately, no data concerning ammonia
levels across an iron oxide bed were found to be available. It is hoped
that future testing will include provision for this measurement.
-------
SECTION 2
EVALUATION OF INTEGRATED GASIFICATION AND CLEANUP PROCESSES
SUMMARY
The results of the integration of the Koppers-Totzek and the resi-
dual oil gasifier with selected cleanup systems are presented in this
section. Those systems that were selected for integration are:
Koppers-Totzek/Selexol Cleanup
Oil Gasifier/Selexol Cleanup
Oil Gasifier/CONOCO Cleanup
For each of these, a flow sheet, mass balances, utility summary and
equipment list are presented. For the K-T gasifier, the results of a
preliminary evaluation comparing it in combination with high- and low-
temperature cleanup systems are also given. It was determined that
there was no advantage to be gained from the combination of the K-T
gasifier and high-temperature (B&W iron oxide) cleanup.
Before describing the integration analysis, it will be instructive
to briefly review the selection criteria for the cleanup systems as
previously described in Ref. 1.
SELECTION CRITERIA
The primary factors considered in selecting cleanup systems include
the efficiency of pollutant removal, effect on power system performance,
cost considerations, and estimated time of availability for commercial
application.
-------
Low-Temperature Desulfurization Systems Selection
The following factors developed in Ref . 1 were considered in
selecting low-temperature desulfurization systems most likely to be
applicable to treating fuel gas:
(1) Sulfur removal capabilities, not only with respect to H2S
but also other sulfur compounds such as COS and CS .
(2) Selective absorption of sulfur compounds over carbon dioxide.
The latter need not be removed from low-Btu fuel gas intended
for use in advanced power cycles, and therefore its absorp-
tion is undesirable since it represents an increased operating
load on the cleanup system.
(3) Type of absorbent insofar as the treated fuel gas may contain
entrained or volatilized solvent which could be detrimental
to downstream system components such as turbine blades, etc.
(h) The system's tolerance to other contaminants present in the
raw fuel gas such as ammonia, cyanides, phenols and tars.
(5) Overall energy requirements and operating costs.
An arbitrary ranking technique was used to rank the cleanup systems.
Based on the ranking, it appeared that the Benfield chemical solvent
process and the Selexol and Rectisol physical solvent processes were
fairly comparable, and ranked somewhat higher than the others. Therefore,
these were evaluated for integrated system performance using data
obtained from process licensors.
High-Temperature Desulfurization Systems Selection
In selecting the most applicable high-temperature desulfurization
systems, the following factors were developed in Ref. 1:
(l) Operating temperature
(2) Capability for removing sulfur compounds, COS, CS?, as well as
(3) The form in which the sulfur is regenerated, e.g., HpS, S02,
or elemental sulfur. Elemental sulfur is the preferred form
since it can be stored without significant pollution problems.
-------
Regenerability of the absorbent without substantial loss of
activity.
(5) Overall energy requirements and operating costs.
From a qualitative comparison based on the above factors, the
Bureau of Mines, and the Babcock and Wilcox processes appear well suited
for use with first-generation gasifiers. These processes are suited
for sulfur removal at temperatures below 1500 F, preferably around 1000
F, which is the operating range for first-generation fixed-bed gasifiers.
Off-gas from a high-temperature, second-generation gasifier would require
cooling to the operating temperature of the iron oxide process and this
would represent a lower thermal efficiency than for integrated systems
using the dolomite-based processes such as the CONOCO process. A disad-
vantage of the iron oxide process is the regeneration of sulfur as
sulfur dioxide. In order to convert this to elemental sulfur, part of
the sulfur dioxide must be reduced to hydrogen sulfide, and this step
consumes fuel. The IGT-Meissner process, when developed, should be
applicable to first-generation-type gasifiers, since its operating
temperature is 900 F. The efficiency of sulfur removal is estimated at
98 percent, and it is selective toward both t^S and COS over C0?.
Second-generation gasifiers can employ the CONOCO dolomite process
which has an operating temperature of 1500 F and above. The Battelle
molten salt process also operates at temperatures around 1500 F, but
its sulfur removal capability is questionable, particularly in the high-
pressure range.
SYSTEM EVALUATION
After the selection of standard cleanup systems was made, the
evaluation of integrated gasifiers and cleanup systems was considered.
Of the low-temperature desulfurization systems selected earlier, the
Selexol process was chosen for detailed evaluation of the integrated
system. This selection was somewhat arbitrary since both the Benfield
and Rectisol processes showed comparable sulfur removal to the Selexol
process and preliminary estimates of overall system performance were
essentially the same.
A comparison of integrated high- and low-temperature gas purification
system was of interest in assessing the relative advantage of high-
temperature cleanup systems in conjunction with gasifiers. For each of
the entrained-flow gasifiers, viz coal-based Koppers-Totzek and the oil-
based partial oxidation, two standard cleanup systems were selected for
-------
detailed evaluation of the integrated systems; a low-temperature and a
high-temperature system. The Selexol process was the representative
low-temperature desulfurization process, and the CONOCO half-calcined
dolomite and the Babcock and Wilcox (B&Jtf) iron oxide processes were
selected as the representative high-temperature desulfurization process.
While the B&W and the Bureau of Mines iron oxide processes operate on
similar principles, the B&W process appears to be in a more advanced
engineering state and was selected for consideration. This will allow
identification of significant differences, if any, between the B&W and
Bureau of Mines process.
As part of the detailed evaluation, heat and mass balances, util-
ities requirements, investment cost estimates, and definition of
pollutant streams were developed for the various combinations of gasifi-
cation and cleanup systems selected. The evaluations were based on a
coal feed rate of 8UOO .tons/day and an oil feed rate of 6000 tons/day
which roughly corresponds to a 1000-Mw COGAS power station output.
The two coal gasification/cleanup-system combination considered
were:
(l) Koppers-Totzek/Selexol
(2) Koppers-Totzek/Babcock and Wilcox
A preliminary comparison of these two systems showed that there is
nothing to be gained from the use of the high-temperature B&W cleanup
system with the K-T gasifier and a combined-cycle power system. This is
due to the need to cool the gas prior to compressing it to the required
burner inlet pressure; Therefore, the results of that preliminary
comparison are presented and only the K-T/Selexol system is described in
detail with a complete mass balance.
Comparison of K-T/Selexol and K-T/B&W.Cleanup System
For this comparison the gasifier capacity was taken to be 350 tons
per hour of Illinois No. 6 coal having the following analysis and heating
value:
C H S 0 N Ash H20
wt. % 6lM 5.1 3.8 9.6 1.2 8.7 U.2
HHV = 12,200 Btu/lb
-------
The gasifier inputs were 0.832 Ib 02/lb coal, 0.3^ Ib steam/lb coal,
and the raw gas temperature was taken to be 2730 F. It was assumed that
all the nitrogen in the coal evolved as elemental nitrogen because
at the peak temperature (3300 F) and pressure (1 atm) in the gasifier,
the equilibrium constant for ammonia formation is very small. Furthermore,
the sulfur in the coal reacted to give HpS and COS, of which the latter
constituted about six volume percent (of the total sulfur) in accordance
with the chemical equilibrium for the hydrolysis reaction:
COS + H 0 - CO + H2S (16)
The following assumptions were made for the purpose of determining
the gasifier output:
(1) Approximately 10 volume percent of the dry product gas is COp.
(2) All the convertible carbon in the coal goes to CO and C02-
(3) At the conditions present in the K-T gasifier, i.e., 3300 F,
.1 atm, the CO shift reaction and methanation are not favored,
and are therefore negligible.
(k) Illinois No. 6 coal has a carbon conversion of 97 percent.
(7)
The above assumptions are based on actual observations on K-T gasifiers.
Coal at 160 F, 98 percent 02 at 230 F, and low-pressure steam at
250 F are fed to the entrained-flow K-T gasifier operating a 1 atm. Raw
gas leaves the gasifier at 2730 F. After a water quench, gas containing
0.9 lb/1000 scf of particulates is cooled in a waste heat boiler in which
high-pressure steam is generated for subsequent use in the power cycle,
as shown in Fig. 11 for the K-T/Selexol system. For comparison purposes,
it was assumed that the hot gas was used to regenerate the high-pressure
clean gas out of the Selexol system to give a clean product gas temper-
ature of 1000 F. In the case of the B&W iron oxide system, cool down is
not necessary prior to the desulfurization step but is required prior to
compression of the product gas. In the Selexol system, both cool down
and compression must be done before desulfurization since the solvent is
sensitive to both temperature and partial pressure of the acid gas. In
any event, both systems do require cool down and subsequent regeneration
to achieve a product gas temperature of 1000 F. The heat recovered from
the gas was therefore assumed to be about equal in each case.
-------
KOPPERS-TOTZEK GASIFIER WITH SELEXOL CLEANUP
VJ1
O
WATER
COAL
STEAM
WATER
O
STREAM
^ ) PRESSURE.PSIA
[~] TEMPERATURE.?
SULFUR
-------
The characteristics of the two systems are compared in Table 2.
The overall heating value of the product gas is slightly higher for the
Selexol system. This is due to the higher quantity of fuel required for
sulfur recovery from the B&W iron oxide regenerator off gas which is in
the form of SC>2 at relatively low concentration. Product gas volumetric
flow rate differs significantly due to the loss of water vapor and
removal of some CC^ in the Selexol cleanup system. It is interesting to
note that the cool down and compression step following the B&W cleanup
will not produce a similar reduction in water vapor. Because the iron
oxide calalyzes the water gas shift reaction, much of the water vapor
is consumed by that process and the product gas is quite rich in hydrogen.
While desirable from a combustion standpoint, this means that a higher
volume of gas must be compressed. Another undesirable effect of the
shift reaction is that the lower heating value of the hydrogen is signif-
icantly less than the CO that it replaces. While both effects are
estimated to represent less than 1 percent of the total system output,
they are certainly in the wrong direction. Thus, while there appears
to be little performance difference between the high- and low-temperature
systems, any performance advantage is in favor of the low-temperature
system. Since a preliminary system performance evaluation showed neither
to be competitive with other integrated systems, it was judged that
nothing constructive would result from further consideration of the K-T/
B&W system.
Koppers-Totzek/Selexol-Process Description
A schematic flow sheet for the K-T/Selexol system is shown in Fig.
11. The material balance is given in Table 8, a utilities summary in
Table 9 and- an equipment list in Table 10.
Gasifier performance was based on Koppers-Totzek data for a West
Virginia Pittsburgh seam coal. To make the sulfur removal process com-
parable to the other systems studied, the sulfur content of the coal was
increased to 3-8 percent. This had only a minor effect on the output
gas composition other than to increase the sulfur compounds. Cold gas
efficiency was given as ?6 percent .
Coal input was taken to be 350 tons per hour with the following
composition:
C H S 0 N Ash H20
wt percent 7^.8 5-0 3-8 6.1 1.3 7.0 2.0
HHV = 13,600 Btu/lb
-------
Table 7
COMPARISON OF K-T GASIFIER WITH HIGH AND LOW TEMPERATURE CLEANUP
FEATURE K-T/B&W K-T/Selexol
1. Tons/day of Illinois No. 6
coal fed to gasifier
2. Raw gas temperature at
gasifier exit, °F
3. Raw gas pressure at gasifier
exit, atm
k. Cleanup system inlet
temperature, ?F
5. Efficiency of sulfur removal
6. ^S content of product gas,
ppm
7. Product gas volume, mscfd
8. Product gas temperature, °F
9. Product gas pressure
10. Chemical heating value of
product gas (HHV) - MMBtu/hr
11. Overall cold gas efficiency
QkOO
2730
1000
270
769
1000
~ 1 atm
6186
72$
8Uoo
2730
100
1000
250 psia
6377
-------
STREAM
M.W.
Table 8
MATERIAL BALANCE FOR K-T/SELEXOL SYSTEM (FIG. 11)
1 23
LB/HR MOL/HR LB/HR MOL/HR LB/HR MOL/HR
21+1+652 13579.7
LB/HR MOL/HR
590569 181+55-3
10551 376.6
700000
21+1+652 13579.7 601120 18831.9
VJI
-------
Table 8
v/i
STREAM
TOTAL
Ash
Jarbon
M.W.
MATERIAL BALANCE FOR K-T/SELEXOL SYSTEM (.FIG. 11)
9
LB/HR MOL/HR
H2°
co-
«2
co2
°2
N2
COS
*./
18.016
28.01
2.016
U4.01
32.0
28.016
60.076
3^.082
73179
10288^1
1+6951
205271
1981+7
3160
25960
3962
36731.2
23289.2
1+66U.2
708. U
52.6
761.7
11+011+09 70169.3
10
LB/HR MOL/HR
11
LB/HR MOL/HR
12
LB/HR MOL/HR
61+1+97
3580
6882 382 101+ 5.8
102881+1 36731.2 10251+35 36609.6
1+6951 23289.2 1+6880 23253.8
205271 1+661+.2 17221+2 3913.7
1981+7
3160
25960
708. It.
52.6
761.7
19802
2109
130
706.8
35.1
3-8
3580 1336912 66589.3 1266702 6^528.6
STREAM
co2
°2
N2
COS'
V
TOTAL
Ash
Carbon
M.W.
13
LB/HR MOL/HR
Ik 15
LB/HR MOL/HR LB/HR MOL/HR
18.016
28.01
2.016
l+U.Ol
32.0
28.016
60.076
3U.082
loU
10211+86
1+6699
171577
19726
2103
130
5.8
36U68.6
2316!+. 2
3898.6
7Q1+.1
35.0
3.8
39^9
181
665
76
6
ll+l.
88.6
15.1
2.7
0.1
1^59
3^06
71
33030
^5
1051
25831
8l.O
121.6
35.^
750.5
1.6
17.5
757.9
1261825 6^280.1
1*877
2U8.5 6U893 1765.5
16
LB/HR MOL/HR
23827
78^76
lkk.6
2801.1
102303 35^5.7
-------
fable 8
MATERIAL BALANCE FOR K-T/SELEXOL SYSTEM (FIG. 11)
STREAM
M.W.
17
LB/HR MOL/HE
H20
co2
02
N2
so2
TOTAL
Sulfur
STREAM
H20
co2
°2
No
so2
TOTAL
Sulfur
18.016
44.01
32.0
28.016
64.066
32.066
M.W.
18.016
44.01
32.0
28.016
64.066
32.066
17366
46026
3971
78596
2492
1W1
20
LB/HR MOL/HR
1873070 103967
963.9
1045.8
124.1
2805.4
38.9
4978.1
LI
193
18
LB/HR MOL/HR
23617
21
LB/HR MOL/HR
736.5
19
LB/HR MOL/HR
148097 82203
148097 82203
22
LB/HR MOL/HR
23
LB/HR MOL/HR
5319
295.2
-------
Table 9
SUMMARY OF UTILITIES FOR K-T/SELEXOL SYSTEM
Steam, Lb/Hr
@ 65 psia
@ 1370 psia
Cooling Water, GPM
Power , kW
BFW, Lb/hr §125F
S58I4F
Stm. Cond., Ib/hr
Process Water, Ib/hr
Chemicals $/day
Gasification Cl).
96555
5800(2)
21000
151060, .
99220^ '
530080
f
Heat ..Recovery, and
Acid .Gas Sulfur
Fuel Gas Compression Removal Recovery Total
(191705)
0.902770)
9953^
1955^0
19U0820
95150 .0-
(51120) (.1,953,890)
33250 33,250
12100 3 132,637
3U6.600
521^0 1,992,960
(95150) (95150)
530,080
65 65
(l) Includes Coal Processing
(2) Oxygen Plant Consumption
-------
TABLE 10
EQUIPMENT LIST FOR K-T/SELEXOL SYSTEM (FIG.11)
Section 100 - Gasification
Item Description
S-101 Gasifier and Main Heat Recovery
F-101 Cyclone Separator
V-101 LP Steam Drum
V-102 HP Steam Drum
Section 200 - Heat Recovery & Gas Compression
Item Description
C-201 Fuel Gas Compressor - Interceded
E-201 . Main Fuel Gas Regenerator
E-202 LP Waste Heat Boiler
E-203 LP Waste Heat Economizer
E-20^ Auxiliary Regenerator
E-205 Low Temperature Regenerator
E-206 Aftercooler
P-201 Process Water Pump
P-202 LP Boiler Recirculating Pump
V-201 Condensate Knock-Out Drum
Section 300 - Acid Gas Removal
Item Description
E-301 Lean Solvent Cooler
E-302 Rich/Lean Solvent Exchanger
E-303 Selexol Stripper OVHD-Condenser
E-SQif Selexol Stripper Reboiler
P-301 Selexol Stripper BTMS Pump
P-302 Selexol Stripper Reflux Pump
T-301 Selexol Scrubber
T-302 Selexol Stripper
V-301 Selexol Stripper OVHD Accumulator
57
-------
Oxygen used in the gasifier was .8Mf Ib 02/lb 'coal and steam was .35
Ib/lb coal. Hot gases leave the gasifier at 2630 F and are water
quenched to 2200 F to solidify the ash prior to entry into the waste heat
boiler. Feedwater is supplied to the boiler at saturation temperature.
The feedwater heating is done in the main gas turbine waste heat boiler
as determined in the course of integrated system optimization. Steam
raised in the low pressure boiler (E-202) combined with that raised in
the gasifier jacket is sufficient to supply the requirements of both
gasifier and Selexol system. The gasifier steam requirement shown in
Table 9 represents the difference between gasifier input and that raised
in the gasifier jacket. The hot gas leaves the low-pressure boiler at
300 F and is used to regenerate the clean gas to 250 F for delivery to
the power system. Prior to compression, the gases are further cooled
and scrubbed. The compressed gas is sent to the Selexol system for H2S
removal.
As is shown in Appendix A, the gasifier exit temperature of 2730 F
agrees quite well with the result of an equilibrium calculation. Thus,
the assumption of 6 percent of sulfur compounds'as COS appears to be
justified. This is important in the sizing of the Selexol system as the
solvent has a relatively low capacity for COS. In the preliminary com-
parison, the system was sized for COS removal resulting in removal of a
significant amount of C02 from the fuel gas and increasing the utility
requirements. With the system sized only for I^S removal, the resultant
sulfur in the product gas is approximately 600 ppmv. While this exceeds
our general study guideline of 500 ppm, this value was established for
air blown gasifiers with product gas heating value in the 100-150 Btu/scf
range. For the K-T gas an equivalent guideline would be 1000 ppmv.
Therefore, the sulfur removal system was sized for H^S only resulting in
reduced size and utilities. Because of the higher concentration of I^S
in the Selexol off gas, this also results in an improvement in the sulfur
recovery section. The HoS concentration in the gas to the Glaus plant is
only O.U percent of the total gas produced.
Another result of the relatively, low fuel gas mass flow rate is to
minimize the effect of fuel gas regeneration. Cycle studies showed that
while possible, regeneration to 1000 F was not sufficiently attractive
to warrant the changes that would be required in the basic K-T system.
OIL GASIFICATION PROCESSES
The two oil gasification/cleanup process combinations considered were:
(l) Partial Oxidation/Selexol
(2) Partial Oxidation/CONOCO
-------
For each case study, the gasifier capacity was taken to be 250
tons/hour of Venezuelan residual fuel oil (KFO) having the analysis,
physical properties, and heating value given in Table 11.
The gasifier input was 6 Ib air/lb oil. The temperature of the
exit gas was taken to be 2500. The sulfur in the fuel oil is converted
to HgS and COS. Steam may be used to control the reaction temperature,
but is not essential to the gasification process itself. From the che-
mical equilibrium of the hydrolysis reaction:
COS + H20 - C02 + HgS. (1?)
The COS is approximately five volume percent of the total sulfur.
Oil Gasification /Selexol-Process Description
Residual fuel oil at 250 F and air at 600 F are fed to the oil
gasifier with an air/oil ratio of 6.0. Entrained gasification of the
oil occurs at about 2600 F and kOO psia to give a raw gas containing
soot (up to three percent of the carbon in the feed). The gas is cooled
to 1200 F in a waste heat boiler to generate 1370 psig saturated steam
which is sent to the power system. The gas then passes through cyclones
and a series of heat exchangers in which it is cooled by exchanging its
heat with clean product gas. The gas is then scrubbed with water to
remove traces of soot. It is then desulfurized in the Selexol absorber,
where 96 percent of the acid gases are removed. The desulfurized gas
with a residual sulfur content of about 110 ppmv is reheated in heat
exchangers by the incoming raw gas, and is delivered to battery limits
at 1000 F, and 275 psia. Including the sensible heat used to raise the
high-pressure steam generated in the waste heat boiler, the overall ther-
mal efficiency of the gasifier/cleanup-system becomes 91 percent.
(22)
The carbon slurry from the water wash is fed to a "Pelletizer"
in which the carbon is recovered from the slurry and mixed with the oil
to form a carbon-oil slurry. While this system is normally used only for
oils that can be fired at temperatures below 100 C^ ^', it was selected
over the alternative naphtha extraction process because of its low utility
requirements. Available data on the naphtha based soot recovery process
showed the steam required for naphtha stripping to be equivalent to the
heating value of the recovered carbon. In the peletizer, the carbon is
wet by the fuel oil forming pellets that are then homogenized into the
main fuel stream. Where the oil must be at elevated temperature, the
quantities of water that are introduced by the process can cause
59
-------
TABLE 11. PROPERTIES OF VENEZUELAN RESIDUAL FUEL OIL
Composition, weight % (ash free)
Carbon 86.^3
Hydrogen 10. 78
Sulfur ' 2.59
98.8
Ash Content, weight % 0.20
(1)
Metals , ppm
Vanadium ^25. 3
Nickel . U?.2
Iron 19 . 0
Sodium 8.0
Copper 0.3
Chromium . 0.2
500.0
Viscosity, SSU @ 212 F 250
Viscosity, SSU @ 100 F 3700
Net Heating Value, Btu/gal 1^2,000
Net Heating Value, Btu/lb 17,300
Gross Heating Value, Btu/gal 150,000
Gross Heating Value, Btu/lb 18,300
Flash, F (Pen sky-Mart ens Closed Cup) 175 •
API Gravity, deg . 12
Density, Ib/gal 8.229
Characterization Factor 10-12
Stoichiometric Air/Fuel Ratio 13.8
(1) Metals content based on crude and adjusted to RFO specifications
60
-------
foaming. However, operation at pressure would alleviate these problems
(although it has not been done commercially) and appears to be the most
desirable approach.
The rich Selexol solvent is regenerated with steam to give an off-gas
containing 39 percent I^S. This is converted to elemental sulfur in a
vapor phase Glaus plant. About one percent of the product guel gas is
used to provide fuel for the Glaus plant.
A schematic flow sheet of the oil gasifier/Selexol system is given
in Fig. 12. The material balance is given in Table 12, a utilities sum-
mary in Table 13} and an equipment list in Table 1^.
Oil Gasifier/CONOCO-Process Description
Raw gas from the gasifier is cooled in a waste heat boiler to 1650
F. High-pressure steam is generated for use in the power system. The
gas then goes through a high-temperature particulate removal system,
where most of the soot is removed. After particulate removal, the gas
enters the high-temperature desulfurization system. In the fluidized-
bed desulfurizer, the gas contacts a half-calcined dolomite acceptor at
^50 psia. The acceptor enters with 75 percent of the calcium as CaS,
and leaves with 88 percent of the calcium as CaS. Approach to equili-
brium for the CO shift and sulfur absorption reactions is assumed to be
100 percent. The overall reactions occurring during desulfurization are
slightly endothermic, so that the treated gas exits at 1600 F and con-
tains 60 ppmv total sulfur. This exceptionally low level is directly
attributable to the characteristics of the oil gasifier. By using no
steam and minimizing COo production the gas phase absorption products
(C02 and ^0) exist at very low concentration and thereby favor the
absorption.
Sulfided acceptor plus fresh dolomite, equal to one percent of the
circulating solids, is transported to the fluidized-bed regenerator by
the regeneration gas. Regeneration is carried out at 1300 F with an 85
percent approach to HoS equilibrium and the gas exits the regenerator
with a molar ratio of carbon dioxide to steam equal to 2.0. About Ik
percent of the CaS is converted and the regenerated solids are recycled
by gravity to the desulfurizer. The off-gas from the regenerator contains
6.U percent HgS by volume and after cooling to 380 F, is fed to a liquid-
phase sulfur recovery unit.
61
-------
PROCESS FLOW DIAGRAM OIL GASIFIER/SELEXOL CLEANUP SYSTEM
SULFUR
cr\
ro
2
u
u
STEAM B.r.W.
CLEAN
FUEL GAS
L. P. STEAM
V-101
o
D
STREAM
PRESSURE PSiA
TEMPERATURE.F
STM
P
NI
-------
TABLE 12
MATERIAL BALANCE FOR OIL GASIFIER/SELEXOL SYSTEM
(See Figure 12)
Stream
M.W.
H2°
CO
"2
?
A^
CH^
COS
HpS
18.016
28.01
2.016
44.01
32.0
28.016
39.944
16.042
17.032
60.076
34.082
12345
LB/HR MOL/HR LB/HR MOL/HR LB/HR MOL/HR LB/HR MOL/HR LB/HR MOL/HR
90503
937013
^3059
109556
2263916
40536
46
148
1216
13007
5025.
33452.
21358.
21*89.
80808.
1014.
2.
8.
.20.
381.
5
8
8
3
0
8
9
7
2
6
14261
937013
43059
109556
2263916
40536
46
148
1216
• 13007
791.
33452.
21358.
2489.
80808.
1014.
2.
8.
20.
381.
6
8
8
3
0
8
9
7
2
6
234
933907
U299U
91919
2258938
14-0536
k6
51
811
61
13
333^1.9
21326.3
2088.6
80630.3
101U.8
2.9
3.0
13.5
1.8
3000000
3^99000 1UU560.6 3^22758 1^0328.7 3369^97 138^36.1
Oil
Soot
500000
86ko
8640
-------
TABLE 12
MATERIALS BALANCE FOR OIL GASIFIER/SELEXOL SYSTEM
(See Figure 12)
ON
-p-
Sulfur 32.066
HgO 18.016
Soot 12.01
6
8
LB/HR MOL/HR LB/HR MOL/HR LB/HR MOL/HR
LB/HR MOL/HR
10
LB/HR MOL/HR
HO
CO
Hp
COp
Op
N2
AT
CHj^
NHo
COS
H S
WO
so2
18.016
28.01
2.016
44.01
32.0
28.016
39.944
16.042
17.032
60.076
34.082
30.008
64.066
1457
3106
66
17635
--
4978
—
--
97
403
12944
--
—
80.9
110.9
32.5
400.7
--
177.7
—
--
5.7
6.7
379.8
--
—
234
931744
42894
91708
2253714
40443
46
51
811
61
13.
33264.
21277.
2083.
—
80443 .
1012.
2.
3.
13.
1.
7
8
8
5
9
0
5
8
2162
99
211
5225
92
77.
49.
4.
—
186.
2.
--
--
--
'
2
3
8 '
13616
5 • 44316
3 795
9929
26419
425.5 2269
1581.8 54519
19.9 887
177
1236
551.1
—
—
600.3
70.9
1946. o
22.2
—
—
--
—
5.7
19.3
1^0686
3361707 138116.
7789
320.1
58727
2027.2
95^30
11 12 13
LB/HR MOL/HR LB/HR MOL/HR LB/HR MOL/HR
11775
367.2
500000 27753
86UO 719.
500000 27753
1^027
3215.5
lU 15
LB/HR MOL/HR LB/HR MOL/HR
778.6 7621+2 4231.9
-------
Table 13
SUMMARY OF OIL GASIFIER/SELEXOL CLEANUP SYSTEM
UTILITIES CONSUMPTION
STEM, Ib/hr
@ 65 psia
§ 1370 psia
COOLING WATER, gpm
POWER, kW
BFW, Ib/hr
STEAM COND. , Ib/hr
PROCESS WATER Ib/hr
CHEMICALS, $/day
Oil
Gasification
132000
60600
36600
17565
(192600)
Heat
Recovery
(296000 )
(1710900)*
'301900
Acid Gas
Removal
19UU50
67910
2U710
(19UH50)
Sulfur
Recovery
(30U75)
2.
31100
Total
-0-
1650300
10U510
U2277
333000
(378050)
130
130
*Includes gasifier jacket steam
-------
Table m
OIL GASIFIER/SELEXOL SYSTEM
EQUIPMENT LIST
Section 100. Gasification & Soot Recycle
Item Description
V-101 Oil Gasifier
V-102 Heat Recovery Unit
T-101 Water Scrubber
P-101 Recirculating Pomp
E-101 Regenerator
Section 200. Heat Recovery
Item Description
V-201 LP Steam Drum
E-201 Main Fuel Gas Regenerator
E-202 L.P. Boiler
E-203 L.P. Economizer
E-20^ Aux. Fuel Gas Regenerator
E-205 Air Cooler
E-206 Aux. Fuel Gas Regenerator
P-201 L.P. Boiler Recirculating Pump
Section 300. Acid Gas Removal
Item Description
E-301 Selexol Solvent Cooler
E-302 Rich/Lean Solvent Exchanger
E-303 Selexol Stripper Reboiler
E-30U Selexol Stripper OVHD Condenser
P-301 Selexol Stripper BTMS Pump
T-301 Selexol Absorber
T-302 Selexol Stripper
V-301 Selexol Stripper OVHD Accumulator
Section kOO. Sulfur Recovery
66
-------
Spent dolomite, withdrawn from the regenerator via a lock hopper,
is treated before discharge to the environment. This stream, containing
75 percent of the calcium as CaS, is slurried with water in a hydrocy-
clone. The slurry is then processed in a three-stage counter current
reactor system where CC>2 is used to convert all the calcium to the car-
bonate form, thereby rendering the stream suitable for discharge to a
sludge pond. The H2S generated in the spent dolomite system is com-
pressed and fed to the sulfur recovery along with the regenerator off-gas.
The liquid-phase Glaus reactor operates at 310 F and converts 90
percent of the H2S feed to elemental sulfur. Sulfur is produced by the
reaction of H2S with a solution of I^SOg. One third of the sulfur that
is produced is .burned are subsequently absorbed by contact with water
to replenish the I^SO-^ used in the reactor. The overhead gases from
the reactor are recycled to the dolomite regenerator. Thus, with this
system the only sulfur emission from the processing system is that part
of the S02 not absorbed by contact with the lean solution from the reac-
tor and the apparent conversion efficiency of H2S is in excess of 99 per-
cent.
Makeup C02 is required for acceptor regeneration and treating of
spent dolomite. Because of the low C02 content of the fuel gas, the
use of a slipstream from the gas turbine exhaust was selected as the
source of C02. Approximately 1.^ percent of the exhaust stream is fed
to an amine recovery system. The product C02 is compressed and used for
makeup.
Product fuel gas is delivered to battery limits at 1600 F and 395
psia. Heating value of the gas is 122 Btu/scf (HHV) and cold gas
efficiency is approximately 73 percent. Because virtually all of the
sensible heat is recoverable and because hot gas efficiency does not
account for gasifier air and fuel preheat or other utility requirements,
the hot gas efficiency is slightly in excess of 100 percent.
A schematic flow sheet for the oil gasifier/CONOCO system is given
in Fig. 13- The material balance is given in Table 15, a utilities
summary in Table 16, and an equipment list in Table 17.
67
-------
PROCESS FLOW DIAGRAM OIL GASIFIER/COIMOCO CLEANUP SYSTEM
en
I
o
b
I
CO
u
t.
L.P. STEAM COOLING
WATER
-------
Table 15
MATERIAL BALANCE FOR OIL GASIFIER/CONOCO CLEANUP SYSTEM
(See Fig. 13)
Stream
ON
M.W.
H 0
CO
H2
co2
02
W2
Ar
CHlj.
NH3
COS
H?S
18.016
28.01
2.016
1+1+.01
32.0
28.016
39.9^
16.01+2
17.032
60.076
31+.o82
1
LB/HR MOL/HR
MOL/HR
MOL/HR
69559
226387
1+05U
90503
' 937013
U3059
109556
2173.7
8080.6 2263916
101.5 U0536
1+6
ll+8
1216
13007
5025.5
33^52.8
21358.8
21+89.3
80808.0
1011+.8
2.9
8.7
20.2
381.6
68223
891886
1+6307
198582
2263917
1+0535
1+7
ll+8
90
2l+5
3786.8
3181+1.7
22969.9
1+512.2
80808.0
1011+.8
2.9
8.7
1.5
7.2
3000000 10355.8 3^99000 1UU560.6 3509980 1UU953.7
Oil
5000000
Stream
CaS-MgO
Inerts
M.W.
18U.01
112.
100
LB/HR MOL/HR
5557.1
530
60871
6
MOL/HR
MOL/HR
8
MOL/HR
30
5
35
.2
.3
.5
5557
3652
300605
53600
363^1^
30.2
350.8
2673
536
3590
10^51
256397.
53070
1+13918.
6
6
7^3.
2279.
530.
355^.
9
9
7
5
1039
256U
530
1+133
.0
.1
.1
7
22
5
35
.1+
.8
.3
.5
-------
Table 15 (Cont'd.)
Stream
10
11
MOL/HR
12
LB/HR MOL/HR
H20
C02
H2S
Stream
H20
°2
N2 '
Ar
S
Stream
H20
U2
No
d
Ar
S02
M.W.
18.016
i+i+.oi
31+. 082
M.W.
18.016
32
28.016
39.9^
32.066
M.W.
•18.016
32.0
28.016
39- 9w
61+. 066
LB/HR
7720014-. 6
191729.6
11*89. 1+
965223.6
13
LB/HR
22359.6
22359-6
17
LB/HR
25039
25039
MOL/ hK
2378.5
1+356.6
1+3.7
6787.7
MOL/HR
121+1.1
121+1.1
MOL/HR
1398.5
1389.8
IJ.D/ im
36179.7
176761.8
11+110
227051.5
ll
LB/HR
1 R
LB/HR
636.8
63871+8.6
l±nv li
127J4-9.1
65251+1.9
2008.2
1+016.U
Ull+
6^38.6
+
MOL/HR-
390.7
MOL/HR
19.9
813.8
10 2
199
ioi+2.9
3626!+. 1+
17731+7.1
11+887.0
213611.5
2012.9 .
14-029.7
1+36.8
61+79.^
15
LB/HR MOL/HR
19
LB/HR
5l+.o
636.8
63871+8.6
1+07.1+
153.8
6i+oooo.6
199
MOL/HR -
3.0
19.9
813.8
10.2
2.1+
81+9-3
191+96.9
1773^7.1
11+89.!+
198333.!+
1082.2
1+029.7
^3.7
5155.6
16
LB/HR MOL/HR
7001+.8
22799.^
1+07.1+
Qnoin fe
20
LB/HR
23795.5
23795.5
218.9
813.8
10.2
104-2 . 9
MOL/HR
1320.8
1320.8
-------
Table 15 (Cont'd.)
Stream
Stream
co2
H2S
M.W.
21
LB/HR MOL/HR
22
LB/HR MOL/HR
H20
C02
Op
W2
Ar
NO
S02
18.016
l&.Ol
32.0
28.016
39.9V*
30.088
6U.066
6521.8
21626.5
19116.8
1311+39.9
2352.7
3
6.1+
362.0
U91.U
597.1+
1+691.6
58.9
0.1
0.1
128175.2
1+326.2
19116.8
131^39-9
2352.7
3 .
6A
39^.9
98.3
597. U
U691.6
58.9
0.1
0.1
181067.1 1509.9 28514-20.2 5814-1.2
M.W.
1+1+.01
3l+. 082
H20(g) 18.016
CaC03'MgC03 181+.01
Inerts 100
H20(l) 18.016
25
585.3
777.1
81+.7
MOL/HR
13.3
22.8
26
LB/HR MOL/HR
1+0.8
5557.1
530
7^0.6
13527.7
23
LB/HR MOL/HR
99^.5 55.2
114-382.5 326.8
15377.0 382.0
27
LB/HR MOL/HR
3320321.6 U29.3
3320321.6 lj-29.3
201.8
2917.9
MOL/HR
11.2
66.3
3119.7
77.5.
-------
Table 16
SUMMARY OF OIL GASIFIER/CONOCO CLEANUP SYSTEM
UTILITIES CONSUMPTION
Steam, Ib/hr
@ 65 psia
@ 1370 psia
Cooling Water, gpm
Power, kv
BFW, Ib/hr
Steam Cond., Ib/hr
Process Water, Ib/hr
Chemicals, $/day
Oil Heat
Gasification Recovery
132000
60600 (1099.700*
30600
17565
(192600)
Sulfur
Removal
(181614)
22360
1U5
18527
Spent Dolomite C02
Treating Recovery
30660
600 2780
85 ^570
(30660)
Sulfur
Recovery
(22162)
3300
2055
22605
Sour Water
Stripper Total
7150 129U8U
(10167UO)
37280
20 2kkhO
U1132
(7150) 23C410
800
800
*Includes steam raised in gasifier jacket
-------
Table 17
OIL GASIFIER/CONOCO SYSTEM
EQUIPMENT LIST
Section 100 - Gasifier and Heat Recovery
Section 200 - Desulfurization
Item
Reactors
R-201
R-202
Description
Sulfur Absorber
Acceptor Regenerator
Vessels
V-201
V-20E-
Dolomite Feed Hopper
Spent Dolomite Hopper
Miscellaneous
F-201
F-202
Absorber Cyclone Separator
Regenerator Ccylone Separator
Compressors
C-201
COp Makeup Compressor
73
-------
Table !7 - Continued
OIL GASIFIER/CONOCO SYSTEM
EQUIPMENT LIST
Section 300 - Spent Dolomite Treating
Item Description
Reactors
R-301 Acceptor Converter 1s* Stage
R-302 Acceptor Converter 2nd Stage
R-303 Acceptor Converter 3rd Stage
Vessels
V-301 Quench Water Surge
Pumps
P-301 Quench Water Pomp
P-302 Dolomite Slurry Pump
P-303 Make-up Water Pump
Exchangers
E-301 Quench Water Cooler
E-302 COp Trim Cooler
Compressors
C-301 C02 Blower
C-302 • Acid Gas Compressor
Miscellaneous
F-301 Hydroclone
Section ^00 - CCL Recovery System
-------
Table 17 - Continued
OIL GASIFIER/CONOCO SYSTEM
EQUIPMENT LIST
Item
Reactors
R-501
Section 500- Sulfur Recovery
Description
Liquid Phase Clause Reactor
Towers
T-501
Vessels
V-501
V-502
Puraps
P-501
P-502
P-503
S0p Absorption Column
Sulfur Separator Drum
Sulfur Storage Drum
Sulfur Pump
As id Pump
Acid Circulating Pump
Compressors
C-501
C-502
Recycle C00 Compressor
Air Compressor
Exchangers
E-501
E-502
E-503
E-505
Recycle C0g Reheater
Feed/Bottoms Exchanger
Weak Acid Cooler
SOp Absorber Intercooler
L.P. Boiler
Miscellaneous
F-531
B-501
Electrostatic Precipitator
Sulfur Burner
75
-------
SECTION 3
REFINEMENT OF INTEGRATED SYSTEMS
SUMMARY
Based upon the results previously obtained,*- ' further refinement of
certain integrated systems was judged to be desirable and achievable.
The refinements desired were those leading to higher efficiencies, lower
emissions, and lower power costs. These were achieved by making appro-
priate process modifications whereby low-grade heat was better utilized
and the utility requirements in different process units were reduced.
The process modifications discussed in this section are summarized below:
BuMines/Selexol System - Resaturation of the clean fuel gas
results in a higher turbine mass flow rate and better performance.
BuMines/Iron Oxide System - Based on more recent data for opera-
tion of the iron oxide cleanup system, significant improvement in
the sulfur recovery process can be made resulting in lower utility
requirements and decreased equipment cost. Several alternative
methods of sulfur recovery are compared.
BCR/Selexol System - Catalytic conversion of COS to H2S upstream of
the cleanup system can permit the use of a smaller Selexol unit,
lower solvent flow rate, lower cost and reduced utility consumption.
BCR/CONQCO System - While reducing overall performance, the addi-
tion of a water scrub for ammonia and particulate removal results in
decreased NOV emission.
•A.
In addition to the process refinements, further performance improve-
ments or at least a better understanding of some of the operating and
effluent relationships of entrained-flow gasifiers could result from the
development of a computer model of the gasification process. A discussion
of the model development and the results of parametric studies using the
model are contained in the following section.
-------
REVISED UTILITY REQUIREMENTS FOR AMMONIA REMOVAL'
A re-evaluation of the ammonia scrubbing requirement 'indicated that
for a practical system only 33 percent of the original value resulted in
a more concentrated ammonia solution (2.l4 percent by weight as against
0.85 percent by weight) off the scrubber. This led to a 58 percent reduc-
tion in the steam and power requirements for the BuMines/Selexol sour
water stripper and a 52 percent reduction for the BCR/Selexol stripper.
The overall effect on system performance of those improvements is given
in SECTION k.
FUEL GAS RESATURATION - BUMINES/SELEXOL
It was found that the performance of the integrated system consisting
of the BuMines gasifier and Selexol cleanup improved when the fuel gas was
resaturated with water vapor before being fired in the gas turbine. The
improved turbine performance described in detail in SECTION k, is attrib-
utable to the increased mass flow through the turbine. The amount of
water required to resaturate the fuel gas, and the resaturation tempera-
ture were determined. The process schematic (Fig. lU) downstream of the
water-quench was altered to utilize the heat contained in the gas for
the sour water and Selexol stripping operations. The resaturation
requirements were found to be:
(l) Saturation temperature: 252°F
(2) Water required: 257, 812 Ib/hr
(3) Water circulation rate: 15,850 gpm
(h) Water temperature at the saturator inlet: 252°F
(5) Water temperature at the saturator outlet: 205°F
A revised mass balance is given in Table 18, a revised utilities
summary in Table 19 and a revised equipment list in Table 20.
REVISED S02 REMOVAL FROM BUMBTES/IRON-OXIDE
Previous estimates of the iron oxide performance during regeneration
had shown a low (5 percent) concentration of S02 in the off-gas. More
recent data^2^) gives a value for S02 concentration in the regeneration off-
gas of 12 percent (by volume). This value was used as a basis to evaluate
four alternative methods of handling the S02 in the off-gas. These alter-
natives were evaluated in light of the higher S02 concentration to
determine their effect on the overall plant efficiency and costs. Each
of the alternatives was arbitrarily chosen as being representative of the
three different categories of S02 removal processes:
77
-------
REVISED PROCESS FLOW DIAGRAM BUMINES/SELEXOL
RECYCLE 1AR
oo
-------
Table 18
STREAM
M.W.
02
N2
CO
co2
H2
CHj,
HoS
COS
NHo
H20
TAR
ASH
TOTAL
STREAM
N
CO
C02
Hp
CHj,
H2S
COS
NHo
•s
~J
TAR
TOTAL
32.00
28.02
28.01
1+U.01
2.016
16. Ok
3k. 08
60.08
17.03
18.02
212
M.W.
28.02
28.01
UU.01
2.016
16. Oil
3l+. 08
60.08
17.03
18.02
212
REVISED MATERIAL BALANCE FOR BUREAU OF MINES/SELEXOL SYSTEM
(see Figure lU )
12 3 U
LB/HR MOL/HR . LB/HR MOL/HR LB/HR MOL/HR LB/HR MOL/HR
700,000
LB/HR MOL/HR
1*91,168
1,617,932
15,3^+9
57,7*+2
2,109,100 73,091
LB/HR
1.6l7;931
7^3.637
29L . 691
35:772
55-319
27.09l|
721
13.^37
17c.;20
77.076
MOL/HR
57,7^2
26,5ll9
6,696
17,7^
3,1+80
795
12
789
9,818
364
283,509 15,733
283,509 15,733
llU,132
8
13,619
i+:r 098 123,989
LB/HR MOL/HR LB/HR MOL/HR
1,617,931 57,71+2
71+3,637 26,5^9
2911,691 6,6-96
35,772 17,7^
55,819 3,1+tO
27,09!+ 795
721 12
13,1+37 7£9
77,076 361+ 805,836 l+l+,719
77,076 361+ 3,59l+, 938 158,526
-------
Table 18 - Continued
MATERIAL BALANCE FOR BUREAU OF MINES/SELEXOL SYSTEM
00
o
STREAM
M.W.
N
CO
C02
CH
COS
NHo
HpO
TAR
TOTAL
STREAM
N
CO
co2
?
HgS
COS
NH
H2e
TAR
28.02
28.01
2.016
16. OU
3^.08
60.08
17.03
18.02
212
M.W.
28.02
28.01
1*1*. 01
2.016
16. Ql*
3k. 08
60.08
17-03
18.02
212
LB/HR MOL/HR
801*593 i*U650
801*593
13
LB/HR MOL/HR
TOTAL
167712 9307
167712 9307
10
LB/HR MOL/HR
168073
77076
9327
36U
21*511*9 9691
Ik
LB/HR MOL/HR
11 12
LB/HR MOL/HR LB/HR MOL/HR
168073 9327
168073 9327
15
LB/HR MOL/HR
1617931
7^3637
29^691
35772
55819
2709^
721
13^37
9208
577^2
265^9
6696
177^
3^80
795
12
789
511
1617931
7^3637
291?I)-2
35772
55819
21675
721
1618
9136
577^2
265^9
6629
177^^
3^80
636
12
95
507
2798310 11^318
636520 35323
636520 35323
16
LB/HB MOL/HR
2993 68
51*52 160
122U5 719
1*52356 25103
2778051 113391* 1*730^ 26050
-------
Table 18 - Continued
MATERIAL BALANCE FOR BUREAU OP MINES/SELEXOL SYSTEM
STREAM
M.W.
17
LB/HR MOL/HR
18
LB/HR MOL/HR
19
20
LB/HR MOL/HR LB/HR MOL/HR
CO
H
N2
CO
co
H2S
COS
NH3
TOTAL
STREAM
CO
C02
H2
H2S
COS
28.02
28.01
1+1+.01
2.016
16. ok
3^.08
60.08
17.03
18.02
1+1+
3k
1+26
1+52290
1
1
25
25099
M.W.
28.02
28.01
1+1+.01
2.016
16.01+
3^.08
60.08
17.03
18.02
1+52791+ 25126
21
LB/HR MOL/HR
161)4372.
7^1173
21+1+781+
35717
55322
102
.'481
562
180
2692693
57615
261+61
5562
17717
31+1+9
3
8.
33
10
110858
3559
21+65
U6959
5k
1+97
21573
2kO
1056
1766
78169
127
88
1067
27
31
633
1+
62
98 7190
2137 7190
399
399
22
LB/HR MOL/HR
.21+1+1
571^
1892
27*+
433
1+1+1+
20l+
k3
136
27
1601931
735U59
21+2891
3^553
5^889
102
1+81
562
256010
57171
26257
5519
17581
31+22
3
8
33
11+207
23 21+
LB/HR MOL/HR LB/HR MOL/HR
29^9
5385
11614-9
9533
67
158
68k
529
kk
68
596
617726
35
31+280
TOTAL
2075^
29278UO 12U201
29516
14-38
3^318
-------
00
ro
STREAM
H2°
Sulfur
TOTAL
H20
Table 18 - Continued
MATERIAL BALANCE FOR BUREAU OF MINES/SELEXOL SYSTEM
M.W.
25
LB/HR MOL/HR
26
LB/HR MOL/HR
27 28
LB/HR MOL/HR LB/HR MOL/HR
N2
CO
C02
H2
H2S
COS
NHo
H2°
TOTAL
STREAM
°2
co2
so2
NO
28.02
28.01
2.016
16.Q1+
3^.08
60.08
17-03
18.02
M.W.
32.00
28.02
1+1+.01
6i+. 06
30.01
11529
11529
29
LB/HR
314301+
113061
677
677
MOL/HR
1072
^035
119
921+1+
9363
LB/HR
681+8
129060
67379
2562
1861
7
513
520
30
MOL/HR
211+
U606
1531
1+0
62
29U9 67
3*+ 1 5385 158
170 10
1651+1+2 9181 288 16
16561+6 9192 8622
31
LB/HR MOL/HR
18.02
32.06
22991+
1276
1U7365
5107
LB/HR MOL/HR
255830
1*1197
230701+ 7729
LB/HR 33 MOL/HR
7930000 1+1+0067
21+205
21+205
755
755
-------
CO
Table 19
REVISED UTILITIES SUMMARY OF BUREAU OF MINES/SELEXOL SYSTEM
Coal
Gasification
STEAM, LB/HR
@ 65 psia
@ 1315 psia 119790
COOLING WATER
GPM
POWER, kW 10500
BRW, LB HR 165355
STM COND., LB/HR
CHEMICAL
$/DAY
Gas Sour Water Ammonia
Cooling Stripping Recovery
* 61+20
50710
7330
1286 297 I2.hh
(57130)
ho
Acid Gas
Removal
*
37090
20214-2
(106200)
70
Sulfur
Recovery
(68365)
h
75060
26
Total
0
170500
33573
(163330)
136
* Heat provided by condensation of water vapor in saturated fuel gas stream.
-------
Table 20
REVISED EQUIPMENT LIST FOR BUREAU OF MINES/SELEXOL SYSTEM
SECTION 100
COAL GASIFICATION AND DUST REMOVAL
F-101 GASIFIER OFFGAS CYCLONE
SECTION 200
GAS SCRUBBCTG AND TAR REMOVAL
P-201 Quench Water Recycle Pump
P-202 Quench Water Pump
P-203 Tar Recycle Pump
P-201* Gas Scrubber BTMS Pump
P-205 Resaturator Pump
E-201 Fuel Gas Reheat Exchanger
E-202 Gas Cooler
T-201 Quench Vessel
T-202 Water Scrubber
T-203 Resaturator Vessel
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-3Q1+ Selexol Stripper OVHD Condenser
V-301 Selexol Stripper OVHD Accumulator
T-301 Selexol Absorber
T-302 Selexol Stripper
-------
Table 20 - Continued
EQUIPMENT LIST
SECTION
SOUR WATER STRIPPING
P-U01 SWS Reflux Pump
P-U02 SWS BTMS Pump
E-U01 SWS OVHD Condenser
E-U02 SWS Reboiler
E-H03 l-'eed/JVi'MG Exchanger
T-lj-01 Sour Water Stripper
V-U01 SV/S OVHD Accuiriulator
SECTION 500
AiMMOWIA REMOVAL
SECTION 600
SULFUR RECOVERY
85
-------
(l) Nonregenerable process
(2) Regenerable process producing su.lfuric acid
(3) Regenerable process producing elemental sulfur
The processes chosen were:
(l) Reduction of two-thirds of the S02 to HpS followed by Claus
plant recovery of elemental sulfur (re-evaluated for 12 percent
S02).
(2) Lime slurry process for the removal of SC>2 from the gas fol-
lowed by disposal of the sludge formed (nonregenerable process).
(3) Catalytic oxidation (Cat-Ox) of S02 to SOo followed by absorp-
tion in water to give sulfuric acid (regenerable process
producing sulfuric acid).
(U) Reduction of SOo (Rsoox) to elemental sulfur using coal as the
reducing agent, followed by recovery of the elemental sulfur
(regenerable process producing elemental sulfur).
Evaluation of the alternatives must be done on the basis of their
effect on overall power system performance and cost. To do this, the
energy accounting system presented in the Phase Report'-^ was used. While
it does not account for changes in the steam cycle caused by the tempera-
ture level of the available heat, it does differentiate between the value
of energy when used at combined cycle vs. steam cycle efficiency and
accounts for all utility requirements. The resultant comparison, while
'not accurate on an absolute basis, is adequate for the selection process.
Process 1 - Claus Plant
The higher SC>2 concentration in the off-gas eliminates the need for
an intermediate step to concentrate the S02 in the off-gas. This results
in a significant reduction in the steam and power requirements for the
sulfur recovery section. Although the SC>2 concentration entering the
Claus plant is now lower than the previous value (after the intermediate
step to concentrate SOo) and, therefore, the Claus plant fuel require-
ment is higher, there is an increase in overall plant efficiency due to
the lower utility requirements. The total power plant efficiency gain
is almost four points to 36 percent. The capital costs of the sulfur
recovery plant are also significantly reduced.
86
-------
Process 2 - Lime Slurry Process for SOp Removal and Disposal
In this process the regeneration off-gas is contacted with a recir-
culating slurry containing slaked lime and reaction products in two
venturi absorbers in series. About 90 percent SC>2 removal is achieved
and the calcium sulfite and sulfate solids formed are disposed of.
Material balances for this process are given in Table 21. Capital and
operating costs' •?' are shown in Tables 22 and 23. A schematic is shown
in Fig. 15.
The lime scrubbing is done in two stages (called scrubber and
absorber in Fig. 15) with a combined SC>2 removal efficiency of 90 percent.
A 15 percent (wt) lime solution is used as a make up and the Ca/S02 ratio
is taken to be 1.1 times stoichiometric. Overall plant efficiency is
37 percent.
Process 3 - Catalytic Oxidation Process
The Cat-Ox process utilizes vanadium pentoxide catalyst to oxidize
S02 to SOo in the off-gas followed by the absorption of the SOo to pro-
duce nominal 80 percent sulfuric acid. Efficient conversion of S02 to
80-3 requires a gas temperature of approximately 850°F to 900°F. This is
achievable in the BuMines/iron-oxide system by partially cooling the off-
gas from the regenerator before introducing it into the oxidation unit.
About 90 percent S02 conversion occurs. A high degree of particulate
removal is required prior to the catalytic oxidation to minimize fouling
of the catalyst. It was assumed that high-efficiency electrostatic
precipitators would be introduced upstream of the oxidation unit. The
oxidation is exothermic and the sensible heat in the gas is used to pro-
duce low-pressure steam before it is sent to the absorber. The
recirculating solution is cooled in circulation acid coolers before being
recycled or sent to product storage. Material balances, and capital and
operating costs*1 •?' are given in Tables 2k, 25, and 26. A schematic is
shown in Fig. 16.
The material balance in Table 2k is based upon cooling the regener-
ator off-gas from'1500°F to 890°F prior to passing it through an electro-
static precipitator. It then undergoes catalytic oxidation followed by
an absorbtion process in which SOo is absorbed by sulfuric acid to gain
80 percent
87
-------
Table 21. MATERIAL BALANCE FOR LIME SLURRY PROCESS
(Stream numbers refer to the flow sheet Fig. 15)
Stream 1 - Cooled regenerator off-gas at 307 Fj 20 psia
Component mols/hr mol^ Ibs/hr
N2 5,3H 83.7U 1^8,708
C02 270 U.26 11,880
S09 761 12,00 U£
Total 6,3^2 100.00 209,292
Stream 2 - Flue gas after first stage scrubber
Component mols/hr Ibs/hr
N2 5,311 lU8,708
co2 270 11,880
S02 202 12,928
H20 625 11,250
Total 6,Uo8 18^,766
Stream 3 - Flue gas after second stage absorber
Component mols/hr Ibs/hr
W2 5,311 lU8,708
S02 270 11,880
S02 76 h,Q6k
H20 625 11,250
Total 6,282 176,702
-------
Table 21 (Continued)
Stream No.
6
8
Material
Ibs/hr
Stream No.
Material
Ibs/hr
Stream No.
Material
Ibs/hr
Stream No.
Material
Makeup
water to
absorber
321059
10
Lime
slurry
to
absorber
903^8
15
Lime
slurry
to
scrubber
316219
20
Settled
used
slurry
Pone
water to
absorber
259^3
11
Recycle
slurry
to
absorber
5*4007
16
Recycle
s lurry
to
scrubber
60001
21
Recycle
pond
water
Lime to
screw
conveyor
146787
12
Discharge
slurry
from
absorber
5*038
17
Discharge
s lurry
from
scrubber
60009
22
Pond
water
to
s laker
Vent
from
s laker
269U
13
Absorber
slurry
to
scrubber
14^6900
18
Used
slurry
to
pump
769573
Grit
to
disposal
h6k
Ik
Makeup
water
to
scrubber
321059
19
Used
slurry to
settling
pond
769573
Lime
slurry
to
system
^05357
Ibs/hi
353767
389223
363006
89
-------
Table ?2
LIME SLURRY PROCESS - OPERATING COST
(On-Site Solids Disposal)
Cost $
* Raw Material 3,602,633
* Labor and Supervision . 238,950
* Steam 571,975
* Process Water 37,700
* Electricity 1,297,856
Labor (maint.) 701 ,802
Analysis 59,850
Total Direct 6,510,766
Average Capital Charges
(14.9% of total capital'investment) 2,333,717
Overhead
Plant, 20% of 2,908,133 581,627
Administrative 10% of 238,950 23,895
Total Indirect 2,939,239
Total Annual Operating Cost = $9,450,005
Unit Costs: Steam: $0.60/M Ib.
Process Water: $0.08/M qal .
Electricity: $0.009/Kwh
Lime: $20.50/ton
Labor and Supervision: $8.00/Man-Hr.
90
-------
Ta'ble .'23
LIME SLURRY PROCESS - CAPITAL COST
(On-Site Solids Disposal)
Investment,$
Lime receiving and storage (bins, feeders, conveyors
and elevators) 1 ,234,1 76
Feed preparation (conveyors,• siakers, tanks, agitators,
and pumps) 589,103
Participate - sulfur dioxide scrubbers and inlet ducts
(4 scrubbers including common feed plenum and pumps) 689,300
Sulfur dioxide scrubbers and ducts (4 scrubbers
including mist eliminators, pumps, and exhaust gas
ducts to inlet of fans) 509,688
Stack gas reheat (4 indirect steam reheaters) 41,609
Fans (4 fans including exhaust gas ducts and dampers
between fan and stack gas plenum) 117,170
Calcium solids disposal (on-site disposal facilities
including slurry disposal pumps, pond, liner, and
pond water return pumps) 5,040,968
Utilities (instrument air generation and supply system,
plus distribution systems for obtaining process steam,
water and electricity from the power plant) 88,267
Service facilities (buildings, shops, stores, site
development, roads, railroads, and walkways) 706,103
Construction facilities 1 ,023,698
Subtotal direct investment 10,040,082
Engineering design and supervision 903,607
Construction field expense 1,004,008
Contractor fees 502,004
Contingency 1,004 ,008
Subtotal fixed investment 13,453,709
Allowance for startup and modifications " 1,104,409
Interest during construction (8%/annum rate) 1.104,409
Total capital investment 15,662.527
91
-------
LIME-SLURRY SCRUBBING PROCESS FOR SO2 REMOVAL FROM BuMINES/IRON-OXIDE REGENERATION OFF-GAS
REGENERATION
VO
ro
S
I
CJ
M
I
u
»> TO STACK
AFTER REHEAT
LIME FROM
STORACiI
FEEDER
SCREW
CONVEYOR
SLURRY
PUMP FEEDTANK
-------
Table 2H
MATERIAL BALANCE FOR CAT-OX PROCESS
(see Figure l6)
Stream No.
Component
02
N2
C02
S02
so3
H20
Total
No
Stream
lbs/hr
1
lbs/hr
121760
lK)0792
522552
5
Acid to
No. 1
Circ.
acid
cooler
11680726
2
lbs/hr
•
1U8708
11880
U870U
209292
6
Acid to
No. 2
Circ.
acid
cooler
1168072
3
lbs/hr
110802
5^9500
11880
1^870
5^792
7318UU
7
Acid to
absorber
11600112
If
lbs/hr
110802
5^9500
11880
U870
677052
8
Acid to
product
storage
8o6lU
93
-------
Table 25
CATALYTIC OXIDATION PROCESS - CAPITAL COST
Investment ,$
Converter and absorber startup bypass ducts
and dampers 191 ,610
Electrostatic precipitators and inlet ducts (4 high
temperature electrostatic precipitators including
common feed plenum) 2,467,567
Sulfur dioxide converters and ducts (4 converters
including catalyst sifter, hopper, storage bin,
conveyors, and elevators) 520,107
Heat recovery and ducts (4 steam/air heaters and 4
fluid/air heaters including ducts between economizers
and air heaters, and combustion air ducts and dampers
between powerhouse and air heaters; investment credit
for use of smaller air heaters included) 638,796
Fans (4 ID fans including exhaust gas ducts and
dampers between ID fans and stack gas plenum) 506,426
Sulfuric acid absorbers and coolers (2.absorbers
including mist eliminators, coolers, tanks, pumps,
and ducts arid dampers between air heaters and ID
fans) 1,983,578
Sulfuric acid storage (storage and shipping
facilities for 30 days production of HoSO^) 128,207
Utilities (instrument air generation ana supply system,
and distribution systems for obtaining process steam,
water, and electricity from power plant) 28,964
Service facilities (buildings, shops, stores, site
development, roads, railroads, and walkways) 269,916
Construction facilities 325,646
Subtotal direct investment 7,06~0,817
Engineering design and supervision 776,690
Construct ion field expense 776,690
Contractor fees . 353,040
Contingency 706,082
Subtotal fixed investment 9,673,320
Allowance for startup modifications 967,332
Interest during construction (8%/annum rate) 776,690
Total capital investment excluding catalyst 11,417,342
Catalyst 313.232
Total capital investment 11 ,730 ,574
-------
Table 25
CATALYTIC OXIDATION PROCESS - OPERATING COST
* Raw material (catalyst)
* Labor and Supervision
* Steam
* Heat credit
* Process water
* Electricity
Labor (mai nt.)
Analyses
Total Direct
Cost $
31 ,323
30,766
33,899
(107,540)
42,606
215,004
247,129
14,003
507,190
Average Capital Charges-
(14.9% of total capital investment)
Overhead
Plant, 20% of conversion costs
Administrative
Total Indirect
Total Annual Operating Cost
1 ,747,856
95,173
124,242
1 ,967,271
$2,474,461
Unit Costs: Steam: $0.70/M Ib.
Process Water: S0.07/M gal.
Electricity: $0.01/Kwh
Catalyst: $1.65/liter.
Labor and Supervision: $8.00/Man-Hr.
Heat Credit: $0.60/MM Btu.
95
-------
CATALYTIC-OXIDATION PROCESS FOR REMOVAL OF SO2 FROM BuMINES/IRON OXIDE REGENERATION OFF-GAS
PRECIPITATOR
START-UP BYPASS
I
FROM
REGENERATOR
2
890 F*
2
'
1 ,_
890 F
CON-
VERTER
. BFW
890 F "•=CONOMIz.En|
IBFW
START-UP
BYPASS
SOLIDS
.DISPOSAL
-*- TO STACK
>
470 F
ABSORBER
/
PRODUCT WATER
ACID COOLER
8
NO. 1 CIRC'N
ACID COOLER
I
DLEFtl
F
fL.NO.2 CIRC'N
ACID COOLER
I —
M
ER
PRODUCT
' STORAGE '
TANKS
PUMP
PRODUCT TO
TRUCK/TANK
CAR
TO STEAM
PLANT
CONDENSATE
HEATER
CONDENSATE
TANK
AIR
o
ro
I
u
M
-------
Process U - Resox Process
This process developed by the Foster-Wheeler Corporation'2"' uses
coal as a reducing agent to produce elemental sulfur from SC>2 contained in
a regeneration off-gas. The S02 in the off-gas stream is reduced to
elemental sulfur which is then condensed out of the gas stream. Crushed
coal is the only material and the only catalyst consumed in the process.
At S02 conversions of 65 percent or lower, and at temperatures below
1100°F, only elemental sulfur is obtained as the product. At higher
temperatures and conversions, I^S is the favored product. The Resox unit
is suited to the BuMines/iron-oxide process and can be located downstream
of the iron-oxide absorber/regenerator with partial intermediate gas cooling
Different types of coals can be used as the reducing agent.
Process Comparison
Performance and costs for each method are given in Table 27 and a
breakdown of capital and power production costs are presented in Table 28.
From Table 27 and 28 it is clear that the better performance and lower
power cost are due to the increased SCv, concentration and there is little
difference between the various methods of sulfur recovery or disposal
(approximately ± 2 percent about the mean cost). It is interesting to
note that there is apparently no significant cost of performance penalty
associated with recovery of elemental sulfur. Accordingly, the limestone
and Cat-Ox processes were not considered further.
For the purposes of this study the Glaus plant approach to sulfur
recovery was selected for use in the final performance and cost estimates.
A revised schematic for this process is shown in Fig. 17, a revised
material balance is given in Table 29, a revised utility summary in
Table 30, and a revised equipment list in Table 31. However, there is
little reason for its selection over the Resox process other than the
fact that it is widely used and information is generally more readily
available. It is possible that improvements in the Resox could result
in higher efficiency or lower cost. The first full-scale plant has only
recently been put into operation for the Gulf Power Company in Florida.
Certainly, a lower SOo concentration 'or a decreased H2 content of the
fuel gas would be reason to reassess the situation. In this system,
reduction of the SOg using hydrogen in the fuel gas results in a further
dilution of the sulfur bearing stream. This is due to the presence of
COp and NO in the fuel and results in an equivalent Glaus feed gas con-
centration of seven percent HgS which is at the lower end of the
practical range. However, this can be improved by shifting the CO in the
fuel to almost double the H2 concentration and thereby decrease the
amount of diluent added during the reduction process.
97
-------
Table 27
COMPARISON OF ALTERNATE SULFUR RECOVERY METHODS FOR BUREAU OF MINES/IRON OXIDE PROCESS
00
New Plant Output - Mw
Overall Efficiency
Heat Rate - Btu/kwhr
. 6
Capital Costs - $10
Sulfur Recovery
Total Plant
Plant Cost - $/kw
Power Cost - Mills/kwhr
Phase , .
Report '
751.5
.320
1,668
10A
289 A
385
20.51
12% S02
Feed to
Glaus Plant
81fl.3
.358
9,529
3.0
282.0
335
17.95
Limestone
Slurry SOg
Removal
875.7
.373
9,155
15.7
29l+. 7
337
17.87
Cat-Ox
Processing
to H2 SO),
895.1
.381
8,957
11. U
290 .k
32k
17.25
Res ox to
Elemental
Sulfur
890.3
.366
9,31^
16.1
295.1
331
17.76
-------
Table 28
POWER COST FOR ALTERNATE SULFUR RECOVERY METHODS BUREAU OF MINES/IRON OXIDE SYSTEM
vo
MD
Phase Report^ '
Power System - $/kw
Gasification & Cleanup - $/kw
Total - $/kw
230
155
385
1 PC? Qn
J_£_ JQ OWO
<_
Feed to
Glaus Plant
206
129
335
Limestone Cat-Ox S02 Resox
Slurry Processing to Produce
SOp Removal to H2 S0[^ Elemental Sulfur
198 193
139 131
337 32U
19U
137
331
Owning Costs - Mills/kvrhr
Operation & Maintenance
Power System
G&C
Fuel at 60^/MMBtu
Total - Mills/kwhr
10.66
1.31
2.1U
6.UO
20.51
9.28
1.17
1.78
5.72
17-95
9.33 8.97
1.13 1.10
1.92 1.81
5.^9 5.37
17.87 17.25
9.17
0
l.ll
1.89
5.59
17.76
-------
REVISED PROCESS FLOW DIAGRAM BUMIIME/IRON OXIDE SYSTEM
H
O
O
CLEAN FUEL GAS
STACK GAS
110
-------
STREAM
Table 29
REVISED MATERIAL BALANCE FOR BUREAU OF MINES/IRON OXIDE SYSTEM
°2
N2
CO
C02
H2
H2S
COS
NH3
H20
TAR
ASH
TOTAL
STREAM
N2
CO
CO-
^
H2
Crih
.tipiD
cos
NH3
H20
TAR
02
TOTAL
M.W.
32.00
28.02
28.01
1*1*. 01
2.016
16. OU
31*. 08
60.08
17.03
18.02
212
M.W.
28.02
28.01
1*1*. 01
2.016
16. 01*
3U.08
60.08
17.03
18.02
212
32.00
mol/hr
1*38,2UO
1,UU3,366
mol/hr Ib/hr
13,695
51,512
mol/hr
mol/hr
700,000
Ib/hr
2UU.OU5 13,5^3
1,881,606 65,207 2UU,OU5 13,5^3
llU.132
llU,132
6
mol/hr Ib/hr
mol/hr Ib/hr
mol/hr Ib/hr
mol/hr
11,625
,M*3,366
639,076
253,278
27,000
1*7,960
25,867
721
11,61*9
175,136
73,819
,697,872
51,512
22,816
5,755
13,393
2,990
759
12
68U
9,719
3U8
107,988
1,328,736
U27,096
1*86,882
36,1*55
1*1* ,158
307
10,729
69,936
67,8Uo
2,1*72,139
1*7,1*21
15,2U8
11,063
18,083
2,753
9
630
3,881
320
99>08
ll*8,8ll* 5,311
1*5,181* 1,1*12
193,998 6,723
-------
o
ro
STREAM
02
N2
CO
co2
H2
CHI,
K2S
COS
S02
NH3
NO
H20
Sulfur
Tar
Dust
TOTAL
M.W.
32.00
28.02
28.. 01
UU.01
2.016
16. OU
3U.08
60.08
6U.06
17.03
30.01
18.02
32.06
212
Table 29 - Cont ' d
MATERIAL BALANCE FOR BUREAU OF MINES/IRON OXIDE SYSTEM
(See Figure 17)
10
'hr
,U3,U22
t63,958
128,912
39,602
U7,96o
3M
ll,6U9
75,972
mol /hr
51,51**
l6,56U
12,018
19,6UU
2,990
10
68U
U,2i6
Ib/hr
8,608
^33,525
81,551
U.612
1,621
U5.032
mol/hr Ib/hr
269
15,^72
1,853
72
5^
2,U99
73,819 3^8
2,685,635 107,988 57^,9^9
11 12
mol/hr Ib/hr mol/hr
20,219
20,518
20,518
6UO
6UO 1,99U
-------
o
U)
STREAM
TOTAL
Table 29 - Cont'd
MATERIAL BALANCE FOR BUREAU OF MINES/IRON OXIDE SYSTEM
(See Figure 17)
13
°2
CO
C02
H2
CH^
H2S
COS
so2
NH3
H20
TAR
TOTAL
STREAM
HP0
M.W.
32.00
28.02
28.01
UU.01
2.016
16. OU
3U.08
60.08
6U..06
17.03
18.02
212
M.W.
18.02
Ib/hr mol/hr
51,6l6 1,613
170,025 6,068
221, 6Ul 7,681
17
Ib/hr mol/hr
61*. 033 3,553
Ib/hr
11U.686
36,861
U2,030
3,lU7
3,801
3U
920
6,037
5,936
213,^52
18
Ib/hr
1,280
mol /hr
"U,093
1,316
955
l,56l
237
1
5^
335
28
8,580
mol/hr
71
6U,033
3,553
1,280
71
15
Ib/hr
5,311
11,883 270
U8,750
761
209,^7 6,3^2
16
mol/hr Ib/hr
mol/hr
65,313 3,621*
65,313 3,62U
-------
Table 30
REVISED UTILITIES SUMMARY FOR
BUMINES/IRON OXIDE SYSTEM
Gasifier
Steam, Ib/hr,
65 psia
1315 psia
27,775
Power, Kw 10,500
BFW, 11)/hr 218,^35
Sulfur
Recovery
(131,280)
133,900
Desulfurization
, 030)
8,3^0
65,310
Total
(195,310)
27,775
18,81+5
199,210
Chemicals,
$/Day
26
26
IQh
-------
Table 31
REVISED BUREAU OF MINES/IRON OXIDE SYSTEM
EQUIPMENT LIST
ITEM
DESCRIPTION
Pumps
P-101
P-120
Boiler Recycle Pump
Boiler Feed Water Pump
Reactors
R-101
Drums
D-101
Sulfur Absorber/Regenerator
(2 Req'd)
Steam Drum
Exchangers
E-101
E-102
Waste Heat Boiler
Economizer
Compressors
C-101
Air Compressor
Separators
F-101
F-102
Gasifier Off-Gas Cyclone
Regenerator Off-Gas Cyclone
105
-------
CATALYTIC CONVERSION OF COS TO H2S IN THE BCR/SELEXOL SYSTEM
One of the problems that arise in the integration of the BCR
gasifier with a Selexol cleanup system is the high proportion of sulfur
assumed to be in the form of COS in the raw gas (« 17 percent of S as
COS). When the Selexol system is sized for H^S removal, the relative
probabilities of H2S and COS result in removal of only one third of the
COS. Since this yields a product gas having a sulfur content of some
700 ppm, the design of the Selexol system must be tailored to COS removal
to meet the basic study guidelines. The net result*- ' was seen in the
relatively high steam and power utility requirements of the BCR/Selexol
system and the relatively large quantity of C02 removed from the fuel
gas as a result of an increased solvent circulating rate. Equilibrium
calculations show that the estimated quantities of COS are well in excess
of their equilibrium level and that a suitable catalyst could be used
to reduce much of the COS to H2S according to the reaction:
. . COS + HgS -» C02 + H2S (18)
A suitable commercial catalyst was found^'' to be a CoMo/Al2Oo catalyst.
The optimum conditions for the catalytic reduction are:
Catalyst: Katalco 20-3 CoMo/A^Og spheres
Temperature: . 550 - 750 F
Pressure: Variable
Vessel L/D: si
Gas hourly space velocity: 5jOOO/hr
Greater than 90 percent conversion is expected under these conditions.
The estimated catalyst life and cost are:
Expected life: , > 1 year
Estimated cost: $65/cu ft
For the purposes of this study, other species were assumed to
remain frozen at the gasifier outlet conditions. Testing would be
necessary to verify the validity of that assumption and to be certain
that no undesirable reactions are encouraged.
After catalytic conversion, the COS is reduced to less than two
percent of the total sulfur with HgS now becoming the key component. This
results in a reduction in solvent flow rate and utilities. The Selexol
stripper off-gas now has a higher H2S concentration,lowering the Glaus
plant fuel requirement. The following reductions in utilities were
identified:
106
-------
Selexol plant: Steam: 6k percent
Power: 59 percent
Glaus plant: Fuel: 55 percent
The COS conversion estimates are based on the use of commercial
Katalco 20-3 CoMo/A^Oo catalyst. The recommended temperature for 90
percent or greater conversion is between 550 F and 750 F. Locating the
catalytic converter between the fuel gas regenerator and the low-pressure
boiler as shown in the revised process flow diagram, Fig. 18 provides
a temperature of 6^5 F in the converter. The original and revised com-
positions of the gas stream entering the Selexol absorber are given in
Table 32. A complete revised mass balance is given in Table 33, the
revised utilities in Table 3^-j and a revised equipment list in Table 35.
With the revised Selexol solvent circulation rate, significantly less
C02 is absorbed resulting in a lower product gas heating value (156.8
Btu/SCF HHV as against 159.2 Btu/SCF).
Cost estimates for the cleanup system were revised and showed a
slight savings in capital cost, principally due to the smaller Selexol
system. A comparison of costs with and without the catalytic converter
is shown below.
Capital Cost Capital Cost
Without converter With converter
($ million) ($ million)
Selexol plant 29.72 ' 23.78
Glaus plant 2.97 2.97
Catalytic unit - 0.^5
Subtotal 32.69 27.20
The estimates include interest and escalation during construction.
Annual operating costs associated with the catalytic unit are estimated
to be $UO,000. The cold gas efficiency of the combined gasifier and
cleanup system increased slightly from 76.k to 76.9 percent. However,
this does not reflect the reduced Selexol utilities which further improve
overall system performance as discussed in SECTION k.
BCR/CONOCO WITH WATER SCRUB
A revised configuration of the BCR/CONOCO system (Fig. 19)
incorporating a water scrub for ammonia and particulate removal was
investigated. Gas from the CONOCO absorber is passed through a boiler
107
-------
PROCESS FLOW DIAGRAM BCR/SELEXOL SYSTEM
ra
I /Y\BFW I FROM POWER SYSTEM I
SULFUR
• AMMONIA
P
CD
-------
Table 32
RAW GAS COMPOSITION - .BCR/SELEXOL SYSTEM
Original Raw Gas Composition
Component mols/hr
Np
CO
C02
H2
CHI).
H2S
COS
NHo
H?0
65^98.6
26151.1
11751.0
18319. ^
5188 .U
687.1
1^2.3
573.3
13953.3
U6.0H
18.38
8.26
12.88
3.65
0.^83
0.10
o.Uo
9.81
Total
1U226U.5 100.00
1+830 ppmv
1000 ppmv
Total 'S' = 5830 ppmv
COS - 17.15fo
Ass\med Composition After Catalytic Conversion
Component mols/hr mol%
N2
CO
C02
H2
H2S
COS
NHo
H20
Total
65^98.6
26151.1
11879.07
18319. ^
5188. k
815.17
lU.23
573.3
13825.23
1U226U.5
h6.oh
18.38
8.35
12.88
3.65
0.573
0.01
o.ko
9.72
100.00
5730 ppmv
100 ppmv
Total 'S1 = 5830 ppmv
COS = 1.715%
109
-------
Table 33
STREAM
M.W.
°9
N2
CO
co2
CiL
H2S
COS
NH
H2d
TOTAL
STREAM
N
CO
co2
H2
H^S
COS
H!O
32.00
28.02
28.01
Mi. 01
2.016
16.0*1
3*1.08
60.08
17.03
18.02
M.W.
28.02
28.01
Ml. 01
2.016
16.04
3*i.o8
60.08
17.03
18.02
REVISED MATERIAL EALAMCE FOR BCR/SELEXOL SYSTEM
US.D5G CATALYTIC COS REMOVAL
LB/HR MOL/HR
(see Figure 18)
2 3
LB/HR MOL/HR LB/HR MOL/HR
177260
70672
*il928
3570
7985
l*i
55
3*i
23
6326.2
2523.1
952.7
1770.6
U97.8
O.k
0.9
2.0
1.3
503392 15731.0
1658190 59178.8
.
700000 (coal)
LB/HR MOL/HR
3015*il 12075. 2161582
7*1909.8
LB/HR MOL/HR
t>a
LB/HR MOL/HR
1835271
732*192
517162
36932
83222
23*116
85*19
9763
251*183
65*198.6
26151.0
11751.0
18319.*!
5188. k
687.1
I*i2.3
573-3
13953-3
1835271
732*192
522799
36932
83222
27782
853
9763
2*19130
65*198.6
26151.1
11879.1
18319. k
5188. k
815.2
l*i.2
573-3
13825.2
LB/HR MOL/HR
396900 22025.5
396900 22025.5
7
8
LB/HR MOL/HR LB/HR ' MOL/HR
TOTAL
60900 (slag)
3*1982*15 I*i226*i.5 3U982U5 1U226U.5
2*il*iU6 13398.8
2*ilU*i6 13393.8
1835271
732^92
517591
36932
83222
2500U
853
976
7875
65^98.6
26151.1
11760.3
18319. k
5188. k
733.7
Ik. 2
57.3
U37-0
32*10196 128160.
-------
Table 33 - Continued
H
H
H
STREAM
CO
CCL
COS
TOTAL
VI. W.
28.02
28.01
44.01
2.016
16. OU
34.08
60.08
17.03
18.02
STREAM
M.W.
LB/H? MOL/HR
5259
2795
119-5
82.0
9050 531.^
347757 19298.4
364861 20031.3
13
LB/HR MOL/HR
10
LB/HR MOL/HR
36
17.
262.
3479^8
0.7
0.5
15.4
19309
3U8258 19325.6
11
LB/HR MOL/HR
1739
1739
96.5
96.5
15
LB/HR MOL/HR LB/HR MOL/HR
12
LB/HR MOL/HR
210000 11653.7
210000 11653.7
16
LB/HR MOL/HR
N2
CO
C02
Hg
C%
HQS
COS
NH3
HoO
28.02
28.01
kk.Ol
2.016
16. OU
3U.08
60.08
17.03
18.02
1831233
730075
^33115
36876
82U81
126
571
3^2
23U
6535^.5
2606U.8
98U2.2
18291.9
51^2.2
3-7
9-5
20.1
13-0
20379
8123
U819
Uio
917
-
6
3
2
727.3
290.
109-5
203.5
57.2
-
0.1
0.2
0.1
163359^
651280
386U08
3^896
73579
112
511
305
209
58301.
23251.7
8780.
16317.7
U587. 2
3-3
8.5
17-9
11.6
U038
2Ul?
8UU16
55
?Ul
24878
282
63^
5902
lUU.l
86.3
1918.1
27.5
U6.2
730.
4.7
37.2
327.5
TOTAL
3115093 1247^1.9
3^659 13S8. 277889^ 111278.9
123363
3321.6
-------
Table 33 - Continued
STREAM
H
ro
M.W.
TOTAL
STREAM
02
N2
co2
NO
H2S
S02
HpO
TOTAL
M.W.
32.00
28.02
Mt-.Ol
30.01
3^.08
6U.06
17.03
18.02
17
LB/HR MOL/HR
18 19
LB/HR MOL/HR LB/HR MOL/HR
6080^6 33526.6
21
LB/HR MOL/HR
8512 ^99.8
8512 ^99.8
20
LB/HR MOL/HR
N2
CO
C02
H2
C%
H2S
COS
WH3
HpO
28.02
28.01
UU.01
2.016
16. OU
3^.08
60.08
17.03
18.02
5259
2795
9050
5909^2
119-5
82.0
531. ^
32793-7
5206
2767
8597
7287
118.3
81.2
50U.8
kok.k
53
27
^53
583655
1.2
0.8
26.6
32389.3
85 5-0
7055 391-5
8597
7287
3857
22
S/HR
5206
2767
232
8205
50U.8
hok.k
1108 . 7
MOL/HR
118.3
81.2
12.9
212 . k
^53
583655
58U188
23
LB/HR
U022U
132507
172731
26.6
32389.3
32U17-9
MOL/HR
1257.
U729.
5986.
85
7055
71^0
LB/HR
8029
156923
11576U
1122
2620
29656
31U11U
•5.0
391-5
396.5
2U
MOL/E
250.9
5600. U
2630. U
37. U
U0.9
16U5.7
10205.7
-------
Table 33 - Continued
STREAM
M.W.
Sulfur 32.06
25
LB/HR MOL/HR
2^850 775.1
26
LB/HR MOL/HR
27
LB/HR MOL/HR
Mt.Ol
3^.08
17.03
18.02
2
22
10
191
235707
k850 775.1 235930
0.5
0.3
11.2
13080.3
13092.3
75^920
HP STM
28
LB/HR MOL/HR
U69270 260^1.6
LP STM
(JO
-------
Table 31*
REVISED SUMMARY OF BCR GASIFICATION/SELEXOL
DESULFURIZATION UTILITIES CONSUMPTION
Coal
Heat
Gas
Acid Gas Trans. Gas Sour Water Sulfur Aumonia
STEAM, Ib/hr
@ 65 psia
@ 1370
@ 1250 SPHT
COOLING WATER, gpm
POWER, kw
BFW, Ib/hr
STM. COND., Ib/hr
PROCESS WATER
Ib/hr
CHEMICALS, $/day
Gasification Recovery Scrubbing Removal Compression Stripper Recovery Recovery
(160U79) 10832U
(1009100)
396900
130315 (78160)
10000
21000
783
37832
13765
3190
163690
210000
100
273
(130315)
5
79725
26
Total
-0-
37750 (971350)
396900
5'46o
926
528U2
39911-2
2U3U15
(^2530) (281169)
30
210000
156
-------
Table 35
REVISED BCR/SELEXOL SYSTEM
EQUIPMENT LIST
SECTION 100 - GASIFICATION
ITEM
F-101
DESCRIPTION
Particulate Removal System
SECTION 200 - HEAT RECOVERY
ITEM
Vessels
V-201
V-202
Exchangers
E-201
E-202
E-203
E-20U
E-205
E-206
E-207
Pumps
P-201
P-202
Compressors
C-201
DESCRIPTION
HP Steam Drum
LP Steam Drum
HP Waste-Heat Boiler
HP Economizer
Main Regenerator
LP Waste-Heat Boiler
LP Economizer
Auxiliary Regenerator
Gas Cooler
HP Recirculating Pump
LP Recirculating Pump
Transport Gas Compressor
115
-------
Table 35 - Continued
REVISED BCR/SELEXOL SYSTEM
' EQUIB4ENT LIST
SECTION 300 - GAS SCRUBBING AND SvfS
ITEM
Towers
T-301
T-302
Vessels
V-301
V-302
Exchangers
E-301
E-302
E-303
E-30U
E-305
Pumps
P-301
P-302
P-303
Reactors
R-301
DESCRIPTION
NH_ Scrubber
NH3 Stripper
Condensate Knock-Out Drum
NHo Stripper OVHD Accumulator
NK, Scrubber OVHD Exchanger
WH_ Stripper BTMS Exchanger
NHo Stripper BTMS Cooler
NH_ Stripper Reboiler
NHn Stripper OVHD Condenser
Absorber BTMS Pump
Stripper BTMS Pump
o Stripper Reflux Pump
COS Converter
116
-------
Table 35 - Continued
REVISED BCR/SELEXOL SYSTEM
EQUIPMENT LIST
SECTION UOO - ACID GAS REMOVAL
ITEM DESCRIPTION
Towers
T-U01 Selexol Scrubber
T-U02 Selexol Stripper
Vessels
V-U01 Condensate Knock-Out Drum 2
V-^02 Selexol Flash Drum
V-^03 Selexol Stripper OVHD Accumulator
Exchangers
E-^01 Lean Solvent Cooler
E-U02 Rich/Lean Solvent Exchanger
E-U03 Selexol Stripper Reboiler
E-UoU Selexol Stripper OVHD Cooler
Pumps
P-401 Selexol Stripper Reflux Pump
P-^02 Selexol Stripper BTMS Pump
Compressor
C-^01 Recycle Gas Compressor
SECTION 500 - AMMONIA RECOVERY
SECTION 600 - SULFUR RECOVERY
117
-------
PROCESS FLOW DIAGRAM BCR/CONOCO WITH WATER WASH
FIG. 19
COAL TRANSPORT GAS
118A
-------
FIG. 19
PROCESS FLOW DIAGRAM BCR/CONOCO WITH WATER WASH
WAtTC MAT! ft
118?.
-------
to drop temperature to the desired level at the inlet of a regenerative
heat exchanger. An aftercooler is required to further reduce temperature
to 120 F for the ammonia scrub and particulate removal process. Resatura-
tion of the fuel gas is incorporated into that block. The gas is then
reheated in the cold side of the regenerator prior to being sent to the
burner.
Inherent in the addition of the water scrub is the need for process
steam in both the sour water stripper and ammonex unit. Also, the regen-
erator temperature and effectiveness will have an effect on system per-
formance. The resultant reduction in fuel gas mass flow rate due to
removal of both ammonia and water vapor also affects performance, with
the loss of water vapor having the most significant effect. Because of
the relatively inexpensive equipment involved and the availability of
low-temperature waste heat, it is apparent that resaturation is desirable,
Regenerator temperature also has a large effect on performance and
while materials are available to withstand temperatures in excess of
1600 F, the practical design and operational problems associated with
thermal stress, operational life and cost make their use questionable.
Therefore, alternate solutions to achieve better performance at more
conventional inlet temperatures (1100 F) were considered.
The high-temperature heat available from the fuelgas as it is
cooled from 1700 to 1100 F can be used to improve steam cycle character-
istics. In essence, it can be used to provide almost all the heat used
in vaporizing the steam while the exhaust gas is used for superheating
and feedwater heating. As a result, a 300 F stack temperature can be
achieved at increased feedwater supply temperatures. If regenerative
feedwater heating to 250 F is used, steam cycle efficiency will improve
by about 6 percent thereby increasing the utilization of the heat
available to the steam cycle and increasing output by 6 percent. This
would provide an increase of 0.8 points in overall cycle efficiency
allowing the use of a more conventional regenerator without serious per-
formance degradation.
However, in order to more directly allow comparison to the
BCR/CONOCO system without the scrubber for ammonia(l), it was decided
to keep a consistent steam bottoming cycle for each.
A revised mass balance is given in Table 36, a revised utility
summary in Table 37, and a revised equipment list in Table 38.
119
-------
Table 36
STREAM
M.W.
02
W2
CO
co2
H2
CHj,
H2S
COS
NHo "
H20
TOTAL
STREAM
32.00
28.02
28.01
Ml-. 01
2.016
16. Oil
3*4.08
60.08
17.03
18.02
-M.V/.
Wp
CO
C02
CH^l
H2S
COS
NHo
H20
TOTAL
28.02
23.01
M4.01
2.0L
16. G]4
60.08
17.0:
18.02
REVISED MATERIALS BALANCE FOR BCR/CONOCO SYSTEM
(see Fig. 19)
LB/HR MOL/HR LB/HR 2MOL/HR LB/HR 3MOL/HR LB/HR SflOL/HR
178053
67395
23537
3761
7919
228
5U
1003
637U
503395
635^.5 1658187
2^06.1
53^.8
1865.7
^93.7
6.7
0.9
58.9
353.7
15731.1
59173.7
700000 (coal)
LB/HR PMOL/HR
28832^4 12075.0 2161582 7^90?.8
LB/HR MOL/HR LB/HR
7MOL/HR
396900 22025.5
396900 22025.5
Q
LB/HR MOL/HR
1839079
716557
513583
3737^
81796
23369
8621
10371
258373
6563^.5
25582.2
11669.7
18538.8
5099.5
685.7
1*43.5
609.0
1^338.1
1839079
696085
58U880
38850
81796
23^1
517
' 10371
256288
6562U.5
214851.3
13239.7
19273.9
5C99.5
63.7
8.6
• 609.0
Ik222.k
1661026
628693
528252
35089
73877
2113
1469
9367
231!47i4
59280.0
22^5.3
12003.0
17^05.1
ii605.8
62.0
7.8
550.0
1281i5.U
60900 (sla.g) 3^89123 1^2301.0 3510207 1^305^.5 3170360 129204.
-------
Table 36 - Continued
MATERIAL BALANCE FOR BCR/CONOCO SYSTEM
ro
t-1
.iroij/un
N2
CO
co2
H2
CHlj.
H2S
COS
NHo
HoO
M.W.
28.02
28.01
hh.oi
2.016
16. ok
3*4.08
60.08
17.03
18.02
LB/HR
178053
67392
56628
3761
7919
228
148
1005
2U8114
9MOL/HR
635*4.5
214-06.0
1286.7
1865.7
1493.7
6.7
0.8
59.0
1377.0
LB/HR 10MOL/HR LB/HR
311071
2814-6
l6lU2 895.8 70588
11 .
MOL/HR
7068.2
83.5
3917.2
TOTAL
STREAM
M.W.
CaC03MgC03 18U.01
/ * M-J-
CaS MgO 112.1
IMERTS 100
3398148 13850.1
LB/HR 13MOL/HR
10636 . 57.8
895.8 3814505 11068.9
LB/HR 12MOL/HR
282U143 6^17.7
26985 791.8
5782U 3208.9
367252
LB/HR ll|MOL/HR LB/HR 15MOL/HR LB/HR l6MOL/HR
1010
10.1
10636
914201
57*4997
101750
57.8
670.9
5H2.9
1017.5
199775
14901438
1007140
1U22.8
14361.0
1007.14
199*4
14903
1010
114.2
143.6
10.1
TOTAL
116U6
67.9
6859.1
790953
6791.2
7907
67.9
-------
Table 36 - Continued
MATERIAL BALANCE FOR BCR/CONOCO SYSTEM
H
ro
M
STREAM
C02
H2S
M.W.
H20
18.02
18^.01
IUERTS 100
TOTAL
STREAI-4
°2
H2
S02
H20
Sulfur
TOTAL
M.W.
32.00
28.02
64.06
18.02
32.06
LB/HR MOL/HR
LB/HR MOL/HR LB/HR 1%OL/HR
5580 126.8 1118 25.4
387 21.5 14780 820.2 148 8.2
5967 148.3 14780 820.2
21
22
LB/HR ~:OL/HR LB/HR MOL/HR
13312 4i6.o
43851 1565.0
57162
1931.0
23962
23962
. k
2752
LB/HR
1210
288
101
77.2
23MOL/HR
37.8
1565.0
5'.!
1612.9
LB/HR 20MOL/HR
789.9
10636 57.8
1010 10.1
25880 857.8
pjl
LB/HR MOL/HR
U8U36 2687.9
U8U36 2687.9
-------
Table 36 - Continued
MATERIAL BALANCE FOR BCR/CONOCO SYSTEM
STREAM
oo
M.V,7.
25
LB/HR MOL/HR
TOTAL
8081.3
26
LB/HR MOL/HR
C02
• H2S
H20
TOTAL
STREAM
N2
CO
C02
H2
CHU
H2S
COS
NH2
H20
UU.01
3^.08
18.02
M.V;.
28.02
28.01
hh.oi
2.016
16. oU
3^.05
60.03
17.03
18.02
283561
28U6
22903
309310
LB/HR
122U1U
U633U
5QkO
2586
5UUU
157
36
690
U6U6
6UU3.1
83.5
1271.0
7797.6
29
MOL/HR
^363.8
165^.2
132.7
1282 . 7
339.^
• U.6
0.6
1+0.5
257.8
27511
1910
29^21
LB/HR
55639
21061
17696
1175
2^75
72
18
313
1728
625.1
106.0
731.1
30
MOL/HR
1985.7
751.9
lj-02.1
583.0
15^.3
2.1
0.3
18.U
95.9
2?
LB/HR MOL/HR
28
LB/HR MOL/HR
25^0.2 210000 11653.7
25^0.2 210000 11653.7
31 32
LB/H?. MOL/HR LB/HR MOL/HR
100177 3993.7
US88C 2712.5 193530 10739-7
US8SC 2712.5 193530 10739.7
-------
Table 36 (Cont'd)
MATERIAL BALANCE BCR/CONOCO
•H
ro
TOTAL
M.W.
33
LB/HR MOL/H8
LB/HR MOL/HR
Np
CO
co2
Hp
ci^
H2S .
COS
NH^
H?0
23.02
28.01
Ml. 01
2 . 016
16. OU
3^.08
60.08
17.03
18 . 02
1661025.6
628692.8
522970.8
35088.7
73877.0
1908.5
U68.6
936.6
8238.7
59280.0
22UU5.3
11883.0
17^05.1
U605.8
56.0
7.8
55.0
U57-2
1661025,6
628692.8
522970,8
35088.7
73877.0
1908.5
U68.6
936.6
128U5.U
59280.0
22MI5.3
11883.0
17^05.1
ii605.8
56.0
7.8
55.0
128U5.li
2933207.3 116195.2 293781U.O 128583.k
35
LB/HR MOL/HR
^597561
^597561
36
LB/HR MOL/HR
5281.2 120
20U.5 6.
8663.2 508.7
255136.6 318575.6 17679
225136.6 33272U.5 18313.7
CO
^n
oU2
H:?
COS
H20
TOTAL
M.W.
37
LB/HR MOL/HR
33
LB/HR MOL/HR
28.02
28.01
] I ) I *~\T
4-4 . 01
2.016
16. ok
3>4.08
60.08
17.03 233.3 ^-7
18.02 318575.6 17679
318808.9 17692.7
8663.2
606389.2
615052. U
39
LB/HR MOL/HR
i+U.O
3.U
1.0
0.1
508.7 U23.5 25A
33650.9 60009U.8 33301.6
3^159.6 6o057it.7 33328.1
-------
M.W.
LB/HR MOL/HR
Table 36 (Cont.'d)
MATERIAL BALANCE BCR/CONOCO
kl
LB/HR MOL/fiR
U2
LB/HR MOL/HR
£
vn
N2
CO
co2
a.
CH^
HgS
COS
NHo
H20
TOTAL
28.02
28.01
Ml. 01
2.016
16. Ok
3*1.08
60.08
17.03
18.02
5193.2
201.1
8230.6
629^.^
199193.3
118.
5.9
U83.3 8iU9 U?8,
3'49.3
956.5 8iU9 U78'.
.5
916369 50852.9
5 916369 50852.9
LB/HR MOL/HR
223235 12388.2
223235 12388.2
-------
Table 37
UTILITIES SUMMARY OF REVISED BCR/ CONOCO SYSTEM WITH WATER SCRUB
STEA.M, 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 Spent Dolomite
Gasification Purification Treating
(135530) (11*7819)
396900
10000 1120
21000 277 155
137190 195530
210000
co2
Removal
58200
390
U863
(58200)
Sulfur Ammonia Sour Water Scrubbing
Recovery Recovery Stripping and Resaturation
(U8880) U780 130315
(37750)
6200
3931 926 273 890
^9370 223235
(13680)
Total
11*1*1*15
396900
17710
30255
(71880)
210000
CHEMICALS, $/DAY
1UOO
11*00
-------
Table 38
REVISED BCR/CONOCO WITH WATER SCRUB 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
Miscellaneous
F-101
F-102
DESCRIPTION
Sulfur Absorber
Acceptor Regenerator
Dolomite Feed Hopper
Spent Dolomite Hopper
HP Steam Drum
BFW Circulation Pump
Waste Heat Boiler
Economizer
Gas Cooler
Absorber Cyclone Separator
Regenerator Cyclone Separator
12?
-------
Table • 38 - Continued
REVISED BCR/CONOCO 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 I nta/^e
Acceptor Converter 211 Lituye
Acceptor Converter 3 Stage
Quench Water Surge
Quench Water Pump
Dolomite Slurry Pump
Make-up Water Pump
Quench Water Cooler
Trim Cooler
C02 Blower
Acid Gas Compressor
Hydroclone
128
-------
Table 38 - Continued
REVISED BCR/CONOCO SYSTEM
EQUIPMENT LIST
SECTION 300 - C00 RECOVERY
ITEM
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
DESCRIPTION
COg Absorber
C02 Stripper
Water Separating Drum
Stripper OVHD Accumulator
Stripper BTMS Pump
Stripper Reflux Pump
C02 Blower
Transport Gas Compressor
Stripper OVHD Condenser
Stripper Reboiler
129
-------
Table 38 - Continued
REVISED BCR/CONOCO SYSTEM
EQUIPMENT LIST
SECTION kOQ - SULFUR RECOVERY
ITEM
DESCRIPTION
Reactors
R-U01
Liquid Phase Clause Reactor
Towers
T-401.
Vessels
V-U01
V-402
SOp Absorption Column
Sulfur Separator Drum
Sulfur Storage Drum
Pumps
P-l+01
P-U02
Sulfur Pump
Acid Pump
Acid Circulating Pump
Compressors
c-Uoi
C-U02
Recycle C00 Compressor
Air Compressor
Exchangers
E-U01
E-^02
Recycle COp Reheater
Feed/Bottoms Exchanger
Weak Acid Cooler
S02 Absorber Intercooler
Preheater
Miscellaneous
F-^01
B-^01
Electrostatic Precipitator
Sulfur Burner
130
-------
Table 38 - Continued
REVISED BCR/CONOCO SYSTEM
EQUIPMENT LIST
SECTION 500 - AMMONIA REMOVAL AND HEAT RECOVERY
ITEM
DESCRIPTION
Towers
T-501
T-502
T-503
Vessels
V-501
V-502
V-503
Exchangers
E-501
E-502
E-503
E-505
E-506
E-507
E-508
Pumps
P-501
P-502
P-503
NHo Scrubber
Fuel Gas Saturator
NHo. Stripper
Condensate Kock-Out Drums
NHo Stripper OVHD Accumulator
H.P. Steam Drum
H.P. Boiler
Regenerator
Resaturation Water Heater
Gas Cooler
NHo Stripper BTMS Cooler
NHo Stripper BTMS-Feed Exchanger
NHo Stripper Reboiler
NHo Stripper OVHD Condenser
NH_ Absorber BTMS Pump
Resaturator Circulating Pump
NHo Stripper BTMS Pump
NHo Stripper Reflux Pump
131
-------
GASIFIER MODELING
One of the major limitations of this and prior studies has been
the lack of data on the operational characteristics of coal gasifiers.
While there are numerous gasifier installations, these are for the
most part providing feedstock for chemical plants and are generally
run at constant conditions. Also, much of the development work has
been directed toward substitute natural gas with the resultant emphasis
on methane production in the gasifier. It is quite likely that a gasi-
fier designed for methane production will not produce optimum perfor-
mance in a low-Btu gas fired combined-cycle power generating system.
Also, when used in conjunction with a power system, it will be desir-
able to operate at other than design point conditions, i.e. at differ-
ent air/coal, steam/coal or pressure levels. Thus, it was apparent
that a gasifier model should be developed to indicate the trends in
operation at alternate design conditions.
The initial approach was based on the use of a chemical equilibrium
model developed nearly 20 years ago as a tool to analyze rocket combus-
tion. That approach described in Appendix A, can be used in an equili-
brium calculation considering in excess of 100 species. A simplified
approach (limited to 10 species) was used to model an oil gasifier.
The results of that analysis are presented in SECTION k showing the
effect of gasifier parameters on system performance and supporting the
need for a model capable of predicting gasifier performance at alternate
operating conditions.
Although the equilibrium approach is applicable to several
gasifier types, for example, the oil and Koppers Totzek gasifiers, it
was found to be unsuitable for the upper stage of the BCR two-stage
gasifier. This type gasifier requires a more complicated approach to
account for the nonequilibrium conditions that are encountered in, for
example, the devolatilization stage of the BCR gasifier. The BCR model
is described in Appendix B while results of a series a parametric vari-
ations of the model are given in the following paragraphs.
Parametric Study of the BCR Two-Stage Gasifier
A series of parametric variations of the major operating variables
of the BCR two-stage gasifier were made in order to see if there were
operating regimes away from the reference design point, i.e., the set
of operating conditions which were selected for use in this study, which
might offer advantages to the overall integrated low Btu coal gasifier
132
-------
combined-cycle system. It was found that some flexibility in design
exists. By decreasing steam while increasing air flow, a slightly dif-
ferent operating point results with the product gas having less chemical
but more sensible energy. The limitations on the choices of operating
conditions are due to chemical kinetics limitations on gasifier opera-
tion.
Variations were made on the design point data from the BCR-Selexol
integrated system. The operating data are summarized in Table 39-
Coal is input at ambient temperature (70 F) carried by a transport gas
in a fixed volume/lb coal ratio. The transport gas temperature is fixed
at the reference design point temperature of 370 F. The transport gas
composition is equal to the gasifier product gas composition with the
sulfur compounds (H^S) removed and varies for each case. Steam tempera-
ture is fixed at 1250 F, and all the parametric variations were calcu-
lated for minimum steam input (no steam into Stage 2). The air temper-
ature was fixed at the reference design point of 800 F. The air and
steam flows are calculated by the gasifier model to satisfy energy and
mass flow balances and to meet the specified Stage 1 and 2 temperatures.
The solids output from Stage 2 of the gasifier is assumed to be recycled
to Stage 1 without mass or energy loss by means of a series of cyclone
separators. The Stage 1 char conversion is conservatively assumed at
60 percent. No operating data on Stage 1 exists since this stage has
not been operated. With these assumptions, the Stage 1 and Stage 2
temperatures were varied over broad ranges both at the design point
pressure (500 psia) and at a higher and lower pressure. The intent was
to identify trends for the resultant changes in gasifier operation under
the major variations in operating conditions.
Figures 20 and 21 show the results of variations of Stage 1 and 2
operating temperatures on the gasifier product gas chemical energy
(HHV-Btu/scf) and the resultant requirements in steam and air. Design
point (T]_ = 2800 F, T2 = l800 F) is indicated by an X. It should be
noted that Figs. 20 and 21 should be plotted on a 3-dimensional curve
since Fig. 20 has various air/coal ratios and Fig. 21 has various
steam/coal ratios. The trends, however, are clear enough so that the
curves as shown present the parametric dependences. Figures 20 and 21
show that in general, lowering the Stage 2 temperature (product gas
temperature) T2 increases the product gas chemical energy and lowers
the steam and air requirements. Thus, there is a trade-off between
sensible and chemical energy in the gasifier product gas, with the lower
steam and air requirements generally favoring chemical energy (HHV).
Raising the Stage 1 temperature T^_ for the same product gas temperature
133
-------
Table 39
SUMMARY OF BCR OPERATING CONDITIONS
Coal Temperature 70 F
Steam Temperature 1250 F
Air Temperature 800 F
Transport Gas Temperature 370 F
Air Composition (mole fraction)
Oxygen 21$
Nitrogen 79$
Transport gas moles/lb coal .0173
Withdrawal 0$
Stage 1 char con version fraction (.YC) 60$
Stage 1 Temperature (variable) 2*4-00 - 3*4-00 F
Stage 2 Temperature (variable) 1*400 - 2*4-00 F
Pressure lU.7, 500, 1000 psia
13U
-------
FIG. 20
BCR TWO-STAGE AIR BLOWN GASIFIER. STEAM/COAL VERSUS PRODUCT
GAS HHV FOR DIFFERENT OPERATING CONDITIONS, P=34 atm (500 PSIA)
X
CONSTANT STAGE 1 TEMPERATURE, T-,
CONSTANT STAGE 2 TEMPERATURE, T2 (PRODUCT GAS TEMPERTURE)
BOUDOUARD CONSTANT K=10
DESIGN POINT T«=2800 F, T9=1800 F
I £. ^
2.0
1.0
0.0
\T
\
\
=2600 \
2400
= 1600
K=10
_L
100
150
HHV (BTU/SCF)
200
76-02-112-1
135
-------
FIG. 21
BCR TWO-STAGE AIR BLOWN GASIFIER. AIR/COAL VERSUS GAS
HHV FOR DIFFERENT OPERATING CONDITIONS. P=34 (500 PSIA)
x
CONSTANT STAGE 1 TEMPERATURE, T.,
CONSTANT STAGE 2 TEMPERATURE, T2 (PRODUCT GAS TEMPERATURE)
BOUDOUARD CONSTANT K=10
DESIGN POINT T1=2800 F, T2=1800 F
4.5
O
4.0
3.0
2.0
100
K=10
"-\^T2=1600
150
HHV (BTU/SCF)
200
76-02-112-2
136
-------
T_ raises chemical energy (HHV) and lowers steam and air requirements,
in general. Since all three are favorable, higher Stage 1 temperatures
are favorable. However, because of kinetic limitations, many of the In
and T2 combinations shown in Figs. 20 and 21 cannot be produced in a
gasifier. The limitations are discussed in APFNEDIX B but are not
thoroughly understood yet, although much research is being done. The
computational approach to the second stage reaction uses a set yield of
CHjj. (YCH^) and CO and/or C02 (YCO) as a fraction of the carbon in the
feed. This is combined with the assumption of equilibrium for the
water gas shift reaction. As is discussed in Appendix B, the methane
yield is based on empirical data relating YCHi, to the partial pressure
of hydrogen in. the second stage. The parameter, YCO, is varied to
achieve the closest approach to equilibrium of the Boudouard reaction
( 2 CO :± C02 + C) as suggested by BCR. The Boudouard Constant, K, is
defined as :
CALCULATED
K —
o
[Xc02/p(XCo) ] equilibrium
Values of K greater than 1 indicate that the calculated carbon monoxide
concentration is less than the equilibrium value. Thus, each point
represents the minimum value of K that can be achieved within the con-
straints that have been imposed. While it is not known if there is a
limiting value of K, the reference design point (K = 5) falls to the
right of the area showing K = 10 on the curves.' Points having K higher
than 10 (to the left of K = 10) are relatively far removed from the
design point and may not be kinetically possible. Further research is
needed to determine if this or other parameters may be used to identify
the kinetic limits on gasifier operation.
In order to assess the variations of gasifier operations at
different pressures, data were obtained for 1 atm and 68 atm (1000
psia). It should be noted that use of the model for the 1 atm case is
not realistic (YCHU = .08 = constant was assumed, for example) for
actual gasifier operation, but trends may nevertheless be deduced.
Figures 22 and 23 show the variation of steam and air requirements
against product gas chemical energy (HHV) for various operating tempera-
tures TT_ and T2 at 1 atm. In summary, the steam and air requirements
are somewhat worse (higher) than in the 500 psia case and the Boudouard
137
-------
FIG. 22
BCR TWO-STAGE AIR BLOWN GASIFIER. STEAM/COAL VERSUS PRODUCT
GAS HHV FOR DIFFERENT OPERATING CONDITIONS, P=1 aim (14.7 PSIA)
X
CONSTANT STAGE 1 TEMPERATURE, TI
CONSTANT STAGE 2 TEMPERATURE, T2 (PRODUCT GAS TEMPERATURE)
BOUDOUARD CONSTANT K=10
DESIGN POINT T.,=2800 F, T2=1800 F
O
0.0
100
150
HHV (BTU/SCF)
138
76-02-112-3
-------
FIG. 23
BCR TWO-STAGE AIR BLOWN GASIFIER . AIR/COAL VERSUS PRODUCT
GAS HHV FOR DIFFERENT OPERATING CONDITIONS. P=1atm(14.7 PSIA)
CONSTANT STAGE 2 TEMPERATURE, T1
CONSTANT STAGE 2 TEMPERATURE, T2 (PRODUCT GAS TEMPERATURE)
BOUDOUARD CONSTANT K=10
DESIGN POINT ^=2800, T2=1800F
4.5
4.0
3.0
2.0
= 2400
VA\\
\g\
V2800—-^e^\
T2=2000
T1=2400
T2=1400
100
150
200
HHV (BTU/SCF)
76-02-112-4
139
-------
constant K = 10 suggests a very limited operation region from a kinetic
viewpoint. (The design point at 500 psia, T]_ = 2800 F, T2 = 1800 F has
K = 5, in fact.)
Figures 2i| and 25 show the variation of steam and air requirements
against predict gas chemical energy (HHV) for various operating tempera-
tures, T, and T2, at 68 atm (1000 psia). The general descriptions of
these variations is the same as for the 3^ atm (500 psia) case. Compar-
ing the 1000 psia case with the 500 psia case shows that! the 1000 psia
cases has lower steam and air requirements and a higher product gas
chemical energy (HHV) for the same operating conditions T-, and T2.
Without a detailed analysis, it is difficult to determine the net value
of higher pressure operation.
-------
FIG. 24
BCR TWO -STAGE AIR BLOWN GASIFIER. STEAM/COAL VERSUS PRODUCT
GAS HHV FOR DIFFERENT OPERATING CONDITIONS. P=68 atm (1000 PSIA)
CONSTANT STAGE 1 TEMPERATURE, T.,
CONSTANT STAGE 2 TEMPERATURE, T2 (PRODUCT GAS TEMPERATURE)
DESIGN POINT T1=2800 F. T2=1800 F
BOUDOUARD CONSTANT K=10
2.0
1.0
0.0
\ \ X T =
=2800\ \ \ 1
2400
T2=1600
.$£ T2=isoo
JP
W'
K=10
100
150
200
HHV(BTU/SCF)
76-02-112-5
-------
FIG.25
BCR TWO-STAGE AIR BLOWN GASIFIER . AIR/COAL VERSUS PRODUCT GAS
HHV FOR DIFFERENT OPERATING CONDITIONS. P=68 atm (1000 PSIA)
CONSTANT STAGE 1 TEMPERATURE, ^
CONSTANT STAGE 2 TEMPERATURE, T2 (PRODUCT GAS TEMPERATURE)
BOUDOUARD CONSTANT K=10
DESIGN POINT T1 =2800 F, t T2=1800 F
O
4.0
3.0
2.0
T.,=3400
T^SOOO
100
T2=2400
T2=2200
^=2800
T.,=2600
T1=2400
1800
K=10
1600
150
HHV (BTU/SCF)
200
76-02-112-6
-------
SECTION 1*
PERFORMANCE AND COST OF INTEGRATED SYSTEMS
SUMMARY
The performance and cost of eight integrated power systems are
presented. These include five variations of the BuMines and BCR
systems previously considered as well as two residual oil partial oxida-
tion systems and a Koppers-Totzek gasifier. A summary of the performance
of these systems is given in Table Uo, and the costs of power for these
systems is given in Table Ul and Fig. 26. For comparative purposes,
the cost of power for a conventional coal fired steam station with FGD
have been developed. All the costs are based upon mid-1975 dollars.
The results show that the high-temperature cleanup systems have
discernible performance and cost benefits over those systems using low-
temperature cleanup. However, improvements in the low-temperature
processes previously considered(l) have narrowed the gap in power costs.
These improvements also make these systems a more viable competitor
to conventional steam with FGD.
PERFORMANCE
The gasification and cleanup system combinations presented in
previous sections have been integrated with either a first generation
(16:1 pressure ratio, 2200 F turbine inlet temperature) or second
generation (2U:1 pressure ratio, 2600 F turbine inlet temperature)
combined cycle or COGAS system. In general, the integration consists
of utilizing engine bleed air for the gasifier air supply and an
interchange of heat usually in the form of steam raised in the gas
turbine exhaust heat recovery boiler, the fuel gas streams, or bleed
air stream and used elsewhere in the process. A versatile simulation
system described in Ref. 1 has been used to represent these systems
and to estimate performance.
3*3
-------
Table Uo
INTEGRATED SYSTEMS PERFORMANCE SUMMARY
Gas Turbine
Turbine Inlet Temperature - F
Compressor Pressure Ratio
Exhaust Temperature - F
Output Power - Mw
Steam Cycle
Steam Temperature - F
Steam Pressure - psia
Condenser Pressure in. Hg Abs
Single or Two Pressure System
Net Steam Cycle Output - Mw
Gasifier and Cleanup System
; Fuel Feed Rate - Ib/hr
Air/Fuel Ratio
Steam/Fuel Ratio
Air Temperature - F
Steam Temperature - F
Steam Pressure - psia
Gasifier Exit Temperature - F
Cleanup System Exit Temperature - F
Fuel Gas HHV - Btu/scf
Integrated Station
Gross Power - Mw
Boost Compressor Power - Mv
Gasifier & Cleanup Aux. Power
Plant Auxiliaries - Mw
Net Plant Output - MW
Net Plant Efficiency (HHV Coal)
Net Heat Rate - Btu/kwh
BuMines/Selexol
With Resaturation
2200 2600
16 21*
927. 1127.
610.U 688.9
BuMines/Iron Oxide
827
1250
U.O
2
226.7
1000
1250
U.O
2
2UU.3
700000 700000
3.013
.Uos
800
58U
1250
1OOO
252
11*1.7
837.1
3U.O
33.6.
8.U
761.1
.32U
10531
3.013
.Uos
800
58U
1250
1000
252
11*1.7
933.1
25.8
27.8
9.2
870.3
.370
9213
2200
16
913
619.1
813.
1250
U.O
2 .
259-3
700000
2.688
.3U9
800
58U
1250
950
1070
165.1
878.9
26.9
18.9
8.0
825.1
.351
9721
2600
2U
1106
701.1
1000
1250
U.O
2
279-3
700000
2.688
• 3U9
800
• 58U .
1250
950
1070
165.1
980. U
20.3
18.8
8.9
932. U
.397
8596,
BCR/Selexol
Without Cos
Converter
2600
2k
1107
727.3
1000
1250
U.O
2
293.3
700000
3.088
.567
800
1000
1250
1800
1000
159.3
1020.6
33. U
56.9
11.0
919.3
.36U
9366
With Cos
Converter
2600
2k
1110.
733.9
1000
1250
U.O
2
320.7
700000
• 3-088
.567
800
1000
12-50
1800 '
1000
I56.it
105U.6
33.it
36.8
11.0
973. k
.386
88U5
BCR/Conoco
No Water
Scrub
2600
2k
1115
857.6
1000
1250
lt.0
2
296.7
700000
3.088
.567
800
1000
1250
1800
1700
135.8
115it.3
38.2
27.2
10.1*
1078.5
,1*2?
798U
With Water
Scrub
2600
2k
1115
769.6
1000
1250
lt.0
2
293.0
700000
3.088
.567
800
1000
1250
1800
1000
136.3
1062.6
38.2
29-3
10.0
985.1
.390
871*0
Oil/Selexol Oil/Conoco
2200 2600
16 2U
908 . 1101 .
653.6 737.2
808 1000
1250 1250
U.O U.o
2 2
U12.9 U30.1
500000 500000
6.0 6.0
.567 -o-
800 800
2U25 2U25
1000 1000
127. U5 127. U5
1066.5 1167.3
U9-7 39-0
30.2 30.2
13.2 lU.l
973. U 108U.O
. 363 . itoU
9U01 8U39
2600
2U
1100.
829.2
1000
1250
U.O
2
Uoi.9
500000
6.0
-0-
800
2U25
1600
122.09
1231.1
37.0
2U.U
12.6
1157.1
.U31
7908
K-T/Selexol
2200
16
920
705-1*
820
1250
U.O
2
378.9
700000
.859*
-3U9
220
275
U5
2200
250
299.35
I08lt.3
99-5
132-3
11.2
8U1.3
.301
113*15
* Oxygen-blown system
-------
TABLE 1*1
POWER GEHERATION COST SUMMARY
GASITIEB/CLEAHUP COMBINATION
POWER SYSTEM - PR/TEMP
Capital Costs - $/ky
Power System Cost - $/kw
Gasification & Cleanup
Cost - $/kw
Total Capital Cost - $/Iw
BUMNES/SELEXOL
WITH RESATURATION
l6:l/g200F
2?8
211*
BUMINES/IRONOXIDE
16:1/2200F
267
168
1*35
BCR/SELEXOL
WITH COS
CONVERTER
2l*:l/2600F
228
207
1*35
BCR/COHOCO OIL/SELEXOL
NO WATER WITH WATER
SCHUB SCRUB
gl*;l/2600F 2l*;l/2600F l6:l/2200F
229
170
399
230
201*
1*31+
283
150
1*33
OIL/CONOCO K-T/ Selexol
2l*;l/2600F 16;1/2200 F
227
109
336
323
31*1
661*
Owning & Operating Costs - Mills/Mi
Owning Costs (17* of Capital) 13-61*
Operation and Maintenance
Power System 1.59
Gasification & Cleanup 2.97
Fuel Cost at 60#/MM Btu Coal 6.32
$2.00/MM Btu Oil
Total Generating Cost 2l*.52
12.06
1.52
2.33
5.83
21.71*
12.06
1.30
2.87
5-31
21.51*
11.06
1.31
2.36
!*.79
19-52
12.03
1.31
2.83
5.21*
21.1*1
12.0
1.62
2.08
18.8
3l*. 5
9-31
1.30
1.51
15.82
27.91*
18.1*1
1.81*
1*.73
6.81
31-79
-------
FIG. 26
40
.c
<£
(/)
E
i
O
u
DC
111
z
111
O
QC
30
POWER GENERATION COST SUMMARY
GASIFIED COAL-COG AS SYSTEM
— DIRECT COAL FIRED STEAM SYSTEM
BUMINES/SELEXOL
FIRST GENERATION
COGAS
BUMINES/IRON OXIDE
FIRST GENERATION
COGAS
BCR/SELEXOL
SECOND GENERATION
20
BCR/CONOCOWITH
WATER SCRUB
SECOND GENERATION COGAS
BCR/CONOCO
SECOND GENERATION
COGAS
NOTE : FOR OIL AT $2.00/MMBtu COST OF POWER GENERATION IS:
OIL/SELEXOL-34.5 MILLS/KWH
OIL/CONOCO -27:9 MILLS/KWH
10
0.50 1.00
COAL COST-S/MM Btu
1.50
2.00
76-02-4-
-------
For each system an optimization of engine pressure ratio based on
cost or performance could, depending on the criteria, result in a
different configuration for a particular turbine inlet temperature.
However, such an optimization was- outside the scope of this study and
a representative pressure ratio was selected for each generation of
turbine inlet temperature. This selection is discussed more fully in
Appendix C. In a like manner, various steam systems could be used,
depending on the characteristics of gas turbine exhaust heat and that
available from gasifier and cleanup systems. However, a simple, low
pressure (1250 psia) nonreheat steam cycle was used as a standard in
all cases. The only variation is the addition of a second low pressure
section, if needed, to achieve a 300 F stack temperature. In the
course of the study, estimates were made to evaluate the performance
benefits associated with reheat steam cycles. In general, an improve-
ment of 20 percent or more in steam cycle performance is possible if
there is sufficient high temperature heat available to support a 1000/
1000 F reheat cycle. At the normal ratio of steam cycle to gas
turbine power this translates into an overall system performance
increase of 7-8 percent. Power system costs would naturally increase
to some extent.
For all systems the effect of a lower fuel control supply pressure
has been factored into system performance. A review of the fuel
control requirements has shown that the fuel gas supply pressure can
be reduced to as low as 3 atmospheres above burner pressure. This
has been reflected in the performance of all systems.
Coal-Fired Steam Station
As is usual in making comparisons of advanced power systems, a
reference system must be defined against which the various advanced
systems are compared. The previous study(l) used as this basis a two
unit 1000-MW coal-fired steam station with a lime-limestone FGD process.
It is unfortunate that the cost of FGD processes have yet to be
Cpc
fully defined, even though such excellent estimates are availablev-;
The problem appears to arise from the confusion as to what should be
considered as part of the FGD and what is chargeable to other parts
of the power system; i.e., higher cost high-temperature electrostatic
precipitators, revised air preheaters, etc. For example, in a study
currently being carried out under NASA sponsorship, the TVA and GE have
given preliminary estimates(29) of a wet scrubber system which added
$26l/kW to the estimated $57^/kW for a 3500 psi/1000 F/1000 F power
station. This system, which uses extraction steam to reheat the stack
-------
gas to allow a 250 F stack, as well as an electrostatic precipitator
operating at 750 F, has an estimated efficiency of about 32 percent.
The net effect of this system is to add over 3^ percent to the cost of
electricity.
While the foregoing may be an extreme, it does appear to account
for all the elements required for effective FGD. In the current study,
the wet lime/limestone system was assumed to cost approximately $92/kW
using costing procedures consistent with Ref. 25. This is consistent
with the previous costs(^-/ escalated to the mid-1975 period.
A significant change in the station costs shown in Table ^2 from
those previously estimated arises from an increase in the time of con-
struction. Until recently, a multi-unit coal-fired station could expect
a construction time of approximately four years. Within the last several
years, however, construction schedules have been extended on this type of
station to about five years. This change alone is equivalent to nearly
$U3/kW assuming a 7 percent escalation and 10 percent interest. In
reality, recent escalation rates have been in excess of 10 percent with
some items approaching 15 percent. Thus the estimates given in Table
U2 may be viewed as being low and therefore conservative in nature when
used as a yardstick to identify the potential benefits of future systems.
The performance of this station is estimated to be 35-1 percent
giving rise to a cost of electricity of 23.5 mills/kWhr (Table 1+3) with
coal at $.60/MMBtu. Figure 26 shows the relation between fuel cost and
cost of power for this station.
K-T Selexol Integrated System Performance
For the configuration shown in Fig. 27 the resultant performance
is summarized in Table Uo. Clearly the K-T coal gasifier, when integrated
with a combined cycle power system shows relatively poor performance when
compared to the high-pressure gasification systems. In general, the
differences can be ascribed to the relatively high power needed to com-
press the product gas and the power consumed by the oxygen plant. The
cold gas efficiency of the K-T/Selexol system is low (76 percent) compared
to the BuMines/Selexol system (83 percent). It is only slightly higher
than that of the BCR/Selexol system (75 percent) and since the product
gas is clean and the sensible heat can be used to raise steam and/or
regenerate fuel gas, the system is quite similar to the BCR/Selexol on
a heat in/heat out basis. However, since the power systems used are
different, a direct comparison is not possible.
1U8
-------
Table ^2
COAL-FIRED STEAM STATION CAPITAL COSTS
Two 500-Mw Units
Mid-1975 Dollars
FPC Account No.
310 Land 52,000
311 Structures and Improvements . 23,9^,900
312 Boiler Plant Equipment 105,216,000
311* Turbogenerator Sets Steam 62,060,000
315 Accessory Electrical Equipment 18,756,800
3l6 .Miscellaneous Power Plant Equipment 878,900
353 Station Equipment 2,6l6,900
Subtotal (Excluding Land) 213,^73,500
Other Expenses b,269,500
Direct Construction Cost 217,7^3,000
Indirects
Contingency- 17,^19,000
Engineering and Supervision 32,661,500
Total Station Costs 267,823,900
Escalation (Five Year Construction; Turbogenerator Firm
for Three Years) 50,1+07,600
Investment Subject to Interest 318,231,500
Total Installed Cost Ul5,373,200
Total Installed Cost with FGD 509,328,200
Cost Per Net KW 532
-------
TABLETS
COAL-FIRED STEAM STATION POWER GENERATING COST SUMMARY
Capital Cost - $/kw
Direct Coal Fired Plant
Stack Gas Cleanup 9*4-
Total Capital Cost 532
Owning and Operating Costs - Mills/kvrti
Owning Cost (17% of Capital) lk.7
Operation and Maintenance
Steam System 1.7
Stack Gas Cleanup 1.3
Fuel Cost at 60^/MMBtu 5.8
Total Cost of Power 23.5
150
-------
K-T/SELEXOL/ INTERGRATED POWER SYSTEM
01
I
o
PROCESS COOLING WATE R
P
N3
-si
-------
In order to identify the factors that are responsible for the poor
performance of the integrated system, the K-T/Selexol was compared with
the BuMines/Selexol system which has the same pressure ratio and turbine
inlet temperature. Also, both have about the same clean fuel temperature.
The parameters of interest are shown in Table kh. To facilitate compari-
son, the K-T/Selexol values have been scaled to a fuel input that is
equivalent to the BuMines system. Thus, the comparison is made between
the first and third columns of the table.
It can be seen that there is nearly a 10 percent difference between
the cold gas efficiencies. However, this is more than made up for by
the net heat recovered which is the difference between the high tempera-
ture heat recovered from the process and the sensible heat required to
raise steam and heat air for the gasifier. Because this sensible heat
can be used only at steam cycle efficiency, its value is approximately
60 percent of an equivalent fuel energy. This relationship has been shown
in Table h6 of Ref. 1. Thus, ,if the clean fuel energy for each system
is adjusted by adding 60 percent of the net sensible heat recovered,
the K-T system shows a total fuel energy available to the power system
of 6652 MMBtu/hr as opposed to 6kyk MMBtu/hr or 2.5 percent more than
the BuMines system for an equivalent coal input.
With a slightly higher energy input to the power system, it would
be expected that power system output would be commensurately higher and
the differences in system output would be the result of power requirements
for fuel gas compression and for the oxygen plant. However, the differ-
ence in power requirements is only 30 MW compared to an overall plant
output difference of 55 MW. The remainder of the difference lies in
the energy that is recovered in the expansion of the fuel gas in the gas
turbine. At the fuel delivery temperature, approximately .67 kW/mol of
gas can be obtained by expanding through the engine pressure ratio. The
differential of 70,000 mol/hr can therefore be equated to, more than UO MW
and the combination of low fuel flow rate and higher utility power
consumption can be shown to account for the difference in output between
the two systems.
In other systems, thermal regeneration of the product gas has been
shown to be very desirable. In the case of the K-T gasifier, the low
flow rate of product gas minimizes the benefits of regeneration. A
comparison was made and showed that regeneration to 1000 F could change
efficiency from 0.301 to 0.307. Consider the changes that this would
necessitate in the basic K-T design, it was judged not to be desirable.
152
-------
Table
PERFORMANCE COMPARISON K-T/SELEXOL
AND BUMINES/SELEXOL
BuMines/Selexol K-T/Selexol
K-T Selexol Scaled To
BuMines Coal Energy
Input
Coal Input-MMBtu/hr (HHV)
Clean Fuel Energy - MMBtu/hr (HHV)
High Temperature Heat Recovered - MMBtu
Sensible Heat to Gasifier - MMBtu
Net Heat Recovered
Air or Gas to Compressor - Mol/hr
Fuel to Burner -Mol/hr
G.T. Bleed Air Power - Mv
Boost Compressor Power - Mw
Fuel Gas Compressor Power - Mw
Oxygen Plant Power - Mw
Total Auxiliary Power - Mw
Net Plant Output - Mw
8,015
6,662
328
608
-280
73,091
12^,201
102
3U
136
T61.1
9,5M*
7,283
1,3U2
266
1,076
70,169
6H,28o
99.5
99.2
198.7
8U1.3
8,015
6,116
1,127
233
89 !»
58,927
53,982
83.6
83.3
166.9
706.5
153
-------
Based on the foregoing analysis of this system, it is apparent
that the K-T gasifier is not well suited for integration with a combined
gas and steam turbine type power plant. While the lack of ammonia and
other troublesome constituents in the product gas makes the K-T system
very desirable from an emissions viewpoint, it appears to be better
suited for use with other, low pressure power systems. Consideration
of such systems was outside the scope of this study and was not pursued.
Oil Gasifier/Selexol Cleanup System Performance
A schematic of the integrated system is shown in Fig. 28. It is
quite similar to the BCR/Selexol except that ammonia removal is not
required; however, a means of recovering soot and returning it to the
gasifier must be included. Because of the high temperature out of the
gasifier, a large amount of high temperature heat is available. By
doing all the feedwater heating in the gas turbine exhaust heat
recovery boiler and using the gasifier system heat only for evaporation,
it is possible to achieve a 300 F stack temperature while having a
single pressure steam system. The resultant performance is given in
Table Uo for both first and second generation power systems. As
previously noted, for all systems considered the steam cycle was simply
a nonreheat 1250 psi single pressure cycle or, where necessary to
reduce stack temperature, a second low-pressure section was included.
For the partial oxidation oil gasifier operating with a second-genera-
tion power system it would be possible to use a reheat steam cycle. A
performance improvement of 8 percent or 3 points could be expected with
such a change.
One of the reasons for investigating the gasification of residual
oil is that the process lends itself to modeling. Because of the high
operating temperature, equilibrium calculations give results that
closely agree with published operating data. The effect of varying
air to oil and steam to oil ratios thus can be examined.
The equilibrium model used was developed at UTRC and is described
in Ref. 30 and Appendix A. Typical feed ratios for commercial residual
fuel oil gasifiers would be between 0.0 and 0.2 for steam/oil and 6.0
to 6.5 for air/oil ratios. The minimum air/oil ratio that would provide
1 atom of oxygen per atom of carbon is approximately 5-0. In practice
it is necessary to increase that ratio to increase raction temperature
and obtain a reasonable reactor size. As would be expected, volumetric
heating value drops sharply with increased air/oil ratios. Since gas
turbine and combined-cycle performance are directly related to fuel
heating value, air/oil ratio is of particular interest in the system
analysis. Because the heating value is on a volumetric basis, it will
-------
OIL/SELEXOL/ INTERGRATED POWER SYSTEM
v_n
v_n
O)
o
to
PROCESS COOLING WATER
FUEL GAS
P
N>
00
-------
increase as water vapor, sulfur compounds and carbon dioxide are removed.
The relation between input and output heating value is shown in Fig. 29,
which presents the chemical heating value of the product gas in terms
of Btu per pound of oil consumed. This amount of energy can be used in
both the gas turbine and steam cycle. The remainder of the initial oil
heating value leaves the gasifier as sensible heat and when mated with a
low temperature cleanup system can only be used in the steam cycle.
The change in volumetric heating value of the fuel gas is very small
over the range of operational steam/oil ratios. When viewed in terms
of chemical heating value per pound of oil consumed, the output of the
gasifier is constant over the range of steam/oil ratios considered and
Fig. 29 therefore applies to a gasifier run both with and without steam
addition. The effects of steam addition on composition are shown in Table
^5 to be an increase in hydrogen and carbon dioxide production coupled
with a decrease in CO. Since each additional molecule of hydrogen brings
with it an oxygen atom which will react with one CO molecule, there is a
one for one correspondence in the increase of hydrogen and decrease of
CO. Because the higher heating value of each molecule is almost the
same, the above results are to be expected. Further examination of the
product gas shows that about 25 percent of the input steam shows up as
hydrogen and the remainder leaves as water vapor in the product gas.
Thus, the net effect of steam addition on the product gas is minimal and
the heat needed to raise the steam is mostly lost since the latent heat
cannot practically be recovered from the water vapor in the fuel gas.
The primary function of steam in oil gasification is to control reactor
temperature and its use will depend on gasifier and heat recovery
equipment design. While the presence of steam in the fuel gas does
increase gas turbine output power due to its-mass, the incremental heat
rate is on the order of 15,000 Btu/kWhr making the use of.steam undesir-
able from the power system viewpoint.
The equilibrium gasifier model was used along with a simplified
model of the Selexol cleanup system and Glaus plant sulfur recovery
performance. Using that model, the result of variations in air/oil
ratio are shown in Fig. 30. Quite clearly, the power system benefits
from lowered air/fuel ratios. In practice, as the air/oil ratio is
desreased below 6.0, a number of, factors including imperfect mixing,
residence time and low temperature (slower reaction rate) combine to
cause nonequilibrium conditions at the gasifier outlet. This in turn
will have a significant effect which will tend to reduce overall
efficiency from the predicted values. However, operation in this area
is certainly desirable and the improved performance may be worth pursuing.
156
-------
FIG. 29
RAW FUEL GAS CHEMICAL HEATING VALUE
STU/LB RESIDUAL OIL CONSUMED
WO CREDIT FOR SULFUR COMPOUNDS)
16000
10000
5.0
6.0
AIR/OIL RATIO - LB/LB
7.0
. R02-41-2
157
-------
Table
EFFECT OF STEAM ADDITION ON FUEL GAS CHEMICAL HEATING VALUE
Fuel - Venezuelan Residual Oil
Air/Oil Ratio =6.0
Fuel Gas Characteristics
Mole Fraction Hg
Mole Fraction HpO
Mole Fraction CO
Mole Fraction COg
Kg SCF/lb Oil
CO SCF/lb Oil
Steam/Oil Ratio
0. 0.2 Q.h
.1^69
.0335
.2335
.015^
16.66
26. U9
.151U
.0612
.215
.02H8
17.83
25.32
.15^8
.0858
.1985
.0328
18.91
2h.2k
HHV Btu/SCF
122.9
118. U
llU.2
Gas Produced SCF/lb Oil
113.^3
11T.T8
122.13
Output Gas HHV - Btu/lb Oil
139^0
139^5
139^7
158
-------
FIG. 30
RESIDUAL OIL GASIFIER /SELEXOL CLEANUP
PERFORMANCE EFFECT OF AIR/OIL RATIO
GAS TURBINE - 24:1! PRESSURE RATIO
2600 F TURBINE INLET
TWO PRESSURE STEAM CYCLE
46
44
42
UJ
u
LL
U.
LU
QC
01
40
38
36
I
5.0
6.0
AIR/OIL RATIO
7.0
R02-46-1
159
-------
The above exercise tends to reinforce the view that for use with
combined cycle power generation, gasifier steam input should be minimized
since this will generally result in minimum air use as well. From a power
system standpoint, minimum steam requirements would be those necessary
to produce a hydrogen content sufficient to maintain proper combustion.
Oil Gasifier/CONOCO Cleanup System Performance
This combination of gasifier and cleanup system represents a very
good match in that there is little ammonia in the raw fuel gas and the
low partial pressure of water vapor and carbon dioxide result in a very
favorable equilibrium concentration of H2S in the cleanup system. A
schematic of the integrated system is shown in Fig. 31 and the resultant
performance is presented in Table hO. Only the second generation power
system was considered and the resultant performance is the best of all
systems investigated.
The use of a reheat steam cycle was also investigated for this
system. It was found that a 900/900 F system could be used with the
2i+:l/2600 F gas turbine. The resulting performance estimate showed
the efficiency to be h7 percent.
BuMines/Selexol Performance
There are several methods that offer the potential for improving
the performance of the BuMines/Selexol system. Some of these, such as
a reduction in the steam to coal ratio are beyond the scope of this
study. However, some improvement in performance can be made by a slight
restructuring of the cleanup system as shown previously in Section 3»
Fig. 18. The revised system is shown in Fig. 32 and the performance
summarized in Table Uo.
Several means of improvement were considered; resaturation of the
fuel gas, reduced stack gas temperature and fuel gas regeneration. Each
of these is discussed in the following paragraphs but only resaturation
was incorporated into the system.
The addition of water vapor to the fuel gas stream increases the
mass flow rate of the fuel gas and decreases the amount of excess air
needed to produce the desired turbine inlet temperature. Thermodynamically,
since the specific heat of water vapor is about twice that of air, one
pound of water vapor decreases the necessary air flow by two pounds.
This results in lower compressor power while the power extracted in the
turbine remains essentially constant. Since the temperature required
160
-------
OIL/CONOCO/INTERGRATED POWER SYSTEM
SULFUR
P
CO
-------
REVISED BUMINES/SELEXOL SYSTEM
CLFAN GAS
TO SULFUR
RECOVERY
-------
for saturation of the gas is generally quite low, it is a good means of
utilizing low-temperature heat.
i
Drawbacks to the use of resaturation are the need for large quanti-
ties of makeup water and the reduced fuel gas heating value that could
result in burner design problems at higher turbine inlet temperatures.
Both the availability of water and the burner design represent potential
problems in system application and were not considered further except
to identify the rate of water usage associated with each alternative
system.
The potential performance improvement to be achieved with the
addition of water vapor to the fuel gas was estimated by varying the
cleanup system output composition. The results are shown in Fig. 33-
In determining performance, no penalty was associated with the
addition of the vapor so the trends shown in that curve are the maximum
that can be achieved, i.e., the need to provide heat to achieve the
humidification will result in less performance improvements. In terms
of water used per megawatt of power,'the incremental power produced
requires about 1^,000 Ib/hr for each additional megawatt of electrical
power. This is nearly constant over the range considered (up to .28
mol fraction of water). Since total pressure just downstream of the
cleanup system will be about 270 psia, a 300 F dew point will produce
a water vapor mol fraction of ,2kQ. It is likely that resaturation to
higher mol fractions will produce limited gains since they will require
the use of heat at a temperature in excess of 300 F. At that temperature
each pound of water needs about 1090 Btu to produce vapor from feedwater
at 120 F. The resultant incremental heat rate is therefore about 15,250
Btu/kWhr which is quite undesirable for use with anything other than
waste heat. Possible sources of heat are the gasifier outlet stream and
the main boiler stack gas. Following the quench and removal of tars
from the gasifier outlet stream, the latent heat of water vapor in the
saturated gas stream greatly increases its heat capacity. In order to
use that heat in resaturating the clean gas, a heat exchanger and
saturator connected by a water heat transport loop were added to the
basic system schematic. The flow sheet including these components
and the revised stream compositions have been presented in Section 3-
The resultant system uses the latent heat in the quenched gas
stream to provide the necessary process heat for both the Selexol and
sour water strippers. Using the quenched stream to supply that heat
results in only an 18 F drop in the stream temperature. Maintaining a
50 F approach, the corresponding saturation temperature that can be
achieved in the product gas is 252 F giving a water mol fraction of
163
-------
EFFECT OF WATER VAPOR IN FUEL GAS BUMINES-SELEXOL SYSTEM
800
790
780
I
ID
Q.
770
760
750
3>
o
ui
M
W
I
U
740
NET OUTPUT MW
EFFICIENCY
0.340
0.336
0.332
I
I
0.328 >
o
z
UJ
o
0.324
0.320
0.04
0.08 0.12 0.16
MOL FRACTION H2O IN FUEL GAS
0.20
0.24
0.28
0.316
P
CO
-------
For the basic BuMines/Selexol configuration, the sensitivity of the
system to stack gas temperature was evaluated. This is shown in Fig. 3^
which gives overall efficiency as a function of stack temperature. The
flattening effect at temperatures below 300 F is a characteristic of
the combined-cycle system where there is no air preheater. The important
part of the curve is the slope above 300 F which shows a decrease in
efficiency of about 0.1 percent for each 10 F increase in stack temperature.
At the conditions of interest, the 10 F increment in stack temperature
is equivalent to about Ml. 5 MMBtu/hr. Comparing this to the benefits
due to rehumidication shows that an additional megawatt output requires
about 15.5 MMBtu/hr and produces an increment in efficiency of 0.0^
percent. If the heat must be extracted from the stack gas this increment
is reduced by about 75 percent making rehumidification under such condi-
tions marginal at best. Again, for the basic BuMines/Selexol configura-
tion the effect of fuel gas temperature was evaluated. This is shown
in Fig. 35- It should be noted that this represents only the benefits
from heating the fuel gas and does not account for the source of the
heat. However, if the heat were taken from either the steam cycle or
stack gas it would be returned for use there with only the turbine work
extracted.
In order to derive a significant performance improvement from
regeneration of the fuel gas against the turbine exhaust, it is necessary
to use only a fraction of the exhaust stream for regeneration. The
remainder of the exhaust stream is then used to raise steam at the high-
est possible temperature. If the heat for regeneration were taken from
the full exhaust stream, not only would that amount of heat be removed
from the steam cycle, but the remainder would be available at a lower
temperature. Both approaches were evaluated and the resulting performance
improvement was approximately 1 percent with regeneration against the
full exhaust stream, while against a bleed stream it was 3 percent or
just over one point. However, the heat transport equipment necessary
to isolate the exhaust gas from the fuel stream was judged to be an
excessive price to pay for the performance improvement.
BuMines/Iron Oxide System Performance
The excessive utilities required by the combination of iron oxide
absorbent and Glaus plant for the production of elemental sulfur from
SOp produced during regeneration of the iron oxide are the focal point
in an attempt to improve performance of this system. Previous performance
estimates had shown a significant amount of oxygen in the off gas during
regeneration, and a low (about 5 percent) concentration of S02. More
recent data'^' show the concentration of S00 in the off gas to be
165
-------
FIG. 34
EFFECT OF STACK TEMPERATURE ON BUMINES/SELEXOL PERFORMANCE
0.318
0.316
I
I
T
u 0.314
UJ
y
H
u_
LLJ
0.312
0.310
240
260
280 300
STACK TEMPERATURE
320
340
360
R05-122-1
166
-------
EFFECT OF DRY FUEL GAS TEMPERATURE BUMINES /SELEXOLSYSTEM
800
790
780
I
1-
o.
O
I-
LU
770
760
750
jo
o
Ul
to
to
740
200
NET OUTPUT MW
• EFFICIENCY
0.340
0.336
0.332
I
i
0.328
0.324
0.320
300
400 500 600
FUEL GAS TEMPERATURE-F
700
0.316
800
P
oo
-------
about 12 percent and to rise sharply to that value on initiation of
regeneration and to fall sharply when regeneration is completed.
Based on the air flow to the bed during regeneration, an 862
concentration of 15 percent would be expected in the off gas. The
difference has been attributed to oxidation of carbon carried into the
bed during absorption. System performance was revised using the follow-
ing assumptions.
1. SC>2 concentration is 12 percent
2. No elemental sulfur out of bed
3. Initial composition is FeS-, <-
k. Regeneration produces all Fe^O?
5. All oxygen is consumed by the sorbent or by oxidation of carbon
6. Regeneration is stopped prior to breakthrough of the
reaction front
The higher SC>2 concentration makes possible the use of various
alternatives to the SC^ concentrator and Glaus system previously used.
Those regenerator off-gas process modifications have already been
discussed and the performance results presented in Section 3- With the
choice of the revised Glaus system without the S02 concentrator (Fig. 17
in Section 3) the integrated system using the modified process as shown
in Fig. 36 has an overall efficiency of 35 •! percent, an increase of 10
percent from the base case.(l)
BCR/Selexol/Catalysis Performance
In an effort to improve the performance of the Selexol plant inte-
grated with the BCR gasifier, catalytic conversion of COS to HgS upstream
of the Selexol unit was used (See Fig. 18 Section 3). It was found that
a commercial cobalt molybdenum catalyst can potentially reduce COS in
the fuel gas to H^S at 650 F, with an efficiency of 90 percent or
greater. Once the COS is converted to I^S the Selexol solvent circula-
tion rate and utilities load decreases significantly. The fuel require-
ment in the Glaus plant is also significantly reduced. The reduction in
major utilities are:
Selexol Plant: Steam 6h%
Power 59$
Glaus Plant: Fuel 36$
168
-------
BUMINES/SINTERED IRON OXIDE SYSTEM
P
CO
O)
-------
The reduction in .auxiliary power is approximately 20 MW. The net
effect of the other changes increase both gas turbine output (up by 3 MW)
and steam cycle output (up by 27 MW). The improvement in steam cycle
output results from both an increase in heat available to raise steam
and a slight improvement in steam cycle performance made possible by an
increase in the ratio of high to low pressure steam raised in the two-
pressure boiler. This, along with the reduced auxiliary power, produces
an increase in net plant output of 5^-1 MW and a resultant efficiency
of 38-6 percent. This is an increase of better than 6 percent over the
previous BCR/Selexol system. Figure 37 shows the revised power system.
BCR/CONOCO/Water Scrub Performance
A revised configuration incorporating a water scrub for ammonia and
particulate removal was prepared for the BCR/CONOCO coal gasification
system. (The flow sheet has been presented previously as Fig. 19 in
Section 3.) The overall system configuration is shown in Fig. 38. Gas
from the dolomite absorber is passed through a boiler to drop its
temperature to the level desired at the inlet of a regenerative heat
exchanger. An aftercooler is required to further reduce temperature
to 120 F for the water scrub and particulate removal process. Resatura-
tion of the fuel gas is also incorporated into that block. The gas is
then reheated in the cold side of the regenerator prior to being sent
to the burner.
Inherent in the addition of the water scrub is the need for process
steam in both the sour water stripper and ammonia recovery unit. Also,
the regenerator temperature and effectiveness will have an effect on
system performance. The resultant reduction in fuel gas mass flow rate
due to removal of both ammonia and water vapor also affects performance.
Table h6 presents the performance effects of increased process steam
requirements and reduction in fuel mass flow rate due to absorption of
ammonia and other constituents and condensation of water vapor. Clearly,
the loss of water-vapor has the most significant effect. Because of the
relatively inexpensive equipment involved and the availability of low
temperature waste heat, it was concluded that resaturation would be
desirable. For a system with full resaturation, the effect of regenerator
effectiveness and operating temperature are shown in Figs. 39 and kO,
respectively. To limit the reduction in performance due to regenerator •
effectiveness, a value of 0.8 was selected for the design. This produces
an approach or minimum temperature difference consistent with the values
used in the study of the economics of regeneration.(1) Regenerator
temperature also has a large effect on performance and while materials
are available to withstand temperatures in excess of 1600 F, the practical
170
-------
BCR/SELEXOL POWER SYSTEM
CI EAN HAS TO BURNER
t
» UUHNfcH
lUMOIISIt
^^
-'
r-
V
PHWE:H
IIJHHlNb
ELt
GE
GAS TUHHINt
/ L.P. TURB. \
P
OJ
-------
REVISED BCR/CONOCO SYSTEM
RECOVERY
s+
so
WA
SI HI
INDENSATE m '
PROCESS COOLING
WATER
-2M-3
101*
-------
Table
BCR/CONOCO—PERFORMANCE EFFECTS OF WATER SCRUB
Overall System Efficiency -
Performance Without Water Scrub - ^3.1
Effect of Increased Process Steam for Sour Water Stripping and
Ammonia Recovery (No Mass Removal) - h2.h
Ideal Regeneration (1700 F Inlet with Effectiveness = 1.0) with Full
Resaturation of Fuel Gas (Tsat = 28l F) (Absorbed Gas Removal Only) - U2.3
Ideal Regeneration with Resaturation of Fuel Gas to T . = 250 F - hi.6
Ideal Regeneration with No Resaturation of Fuel Gas - U0.7
NOTE: Base performance is slightly higher than shown in Table 39 due to
accounting process. Numbers are mutually consistent in all other respects.
173
-------
FIG. 39
BCR/CONOCO WITH WATER SCRUB
EFFECTOR REGENERATOR EFFECTIVENESS
(FULL RESATURATION OF FUEL GAS)
(1.700F REGENERATOR INLET TEMP.)
42
u
z
UJ
O
UJ
2 40
O
38
0.7
0-8 o.9
REGENERATOR EFFECTIVENESS
1.0
R08-106-1
-------
FIG. 40
BCR/CONOCOWITH WATER SCRUB
EFFECTOR REGENERATOR INLET TEMPERATURE
(REGENERATOR EFFECTIVENESS = 0.9)
(FULL RESATURATION OF FUEL GAS)
42
5?
I
>-
U
z
UJ
u
oc
LU
o
40
38
800 1100 1400
REGENERATOR HOT SIDE INLET TEMPERATURE - F
1700
R08-106-2
175
-------
design and operational problems associated with thermal stress, operational
life and cost make their use questionable. Therefore, alternate solutions
to achieve better performance at more conventional temperatures were
considered (1200 F hot side inlet with regeneration to 1000 F).
The high temperature heat available from the fuel gas as it is
cooled from 1700 to 1200 F also could be used to improve steam cycle
characteristics. In essence, it could provide almost all the heat used
in vaporizing the steam while the exhaust gas is used for superheating
and feedwater heating. As a result, a 300 F stack temperature could be
achieved at increased feedwater supply temperatures. If regenerative
(steam) feedwater heating to 250 F is used, steam cycle efficiency will
improve by about 6 percent thereby increasing the utilization of the
heat available to the steam cycle and increasing output by 6 percent.
This would produce an increase of 0.8 points in overall cycle efficiency.
Another, more costly alternative would be to use a reheat steam cycle to
achieve even greater performance improvement. However, to allow a
better comparison of the effects of the cleanup system revision and to
be consistent with the other systems, the basic steam cycle operating
parameters were not changed from those of the BCR/CONOCO system without
water scrub.
176
-------
SYSTEM COSTS
The following paragraphs discuss the costs of the various integrated
systems considered to date. All costs are in mid-1975 dollars for a
North Central location. Costs previously presented(^) have been esca-
lated using recognized escalation procedures.(31>32,33) Capital charges
of 17 percent/yr and a 0.7 load factor were assumed.
The costs of equipment are based upon values found in the literature,
upon vendor quotes and upon Contractor-developed costing procedures. Of
particular use in developing the costs for steam stations and for trends
in equipment costs was the work done by United Engineers for the AEC. (.34)
The cost of coal was assumed to be $.60/MMBtu at the mine mouth.
While coal costs have risen dramatically in the past several years, it
would appear that power plants in the North Central Region could obtain
coal at or near this cost.'35) ^he cost of residual oil was assumed to
be $2.00/MMBtu, a cost typical of high-sulfur residual in barge quanti-
ties. (36)
A summary of the gasifier and cleanup system capital costs is given
in Table Vf and a summary of the power system costs is given in Table U8.
In addition, the discussion of the various systems contains comparative
cost summaries.
Cost of Hot Particulate Removal Systems
In the previous study, the cost of particulate removal from the gas
stream was assumed to be relatively low, in the order of <$2/k¥, based
upon the cost of materials for high-temperature cyclones. Because of
the immature state-of-the-art in this area, definitive costs are diffi-
cult. However, consideration of the need for coarse separation followed
by several stages of fine filtration indicate that the cost of hot par-
ticulate removal could be quite high.
As part of a study of fluid-bed combustors carried on under
Corporate sponsorship, a particulate removal system operating at condi-
tions similar to those of interest, i.e., 1650 F, 250 psi was considered.
This system contained cyclone, multiclone and granular filters and had
installed costs in the range of $75 to $100/ACFM (actual cubic foot per
minute) including high-temperature piping. Thus, for the system consid-
ered herein, total system costs were of the order of $20/kW rather than
$2/k¥. As will be seen in the following paragraphs, this has a signifi-
cant effect on system costs.
177
-------
Table
GASIFIER & CLEANUP SYSTEM CAPITAL COST BREAKDOWN
Millions of Mid-1975 Dollars
Gasification
Gas Cooling
Hot Particulate Removal
Desulfurization
Sour Water Stripping
Ammonia Recovery
Sulfur Recovery
Waste Water Treatment
Boost Compressor & Boiler
Feedwater Treatment
Cooling Tower
Condensate Polishing
Other Expenses
Total Captial Cost
(includes Escalation &
Interest)
BuMines/
Selexol
68.98
16.1*2
— _
26.28
6.56
11.1*9
3.28
5.33
11.98
7.85
1.22
.07
3.19
162.65
BuMines/
Iron Oxide
68.98
16.80
23.0
l*.63
k.lk
11.90
6.76
_.._
.21*
2.53
138.98
BCR/
Selexol
91.97
26.28
__«
26.77
6.56
9.86
3.28
6.83
12.21
10.99
2.01
.28
i*.o6
201.10
BCR/
Conoco
91.97
21*. 1
23.00
1.65
___
9.86
5.06
13.02
10.99
.1*8
.03.
3.12
183.28
BCR/
CONOCO
Water
Scrub
91.97
26.28
» » <.
23.00
6.56
9.86
9.86
5.06
13.02
10.99
.1*8
.03
3.9>
201.05
P.O./
Selexol
&M
29.7*1
1+6.77
2.35
—
3.21*
6.55
16.1*0
—
3.68
___
2.86
11*6.06
P.O./
Conoco
32.1*3
3-52
26.00
2l*.8o
1.78
—
10.63
5.1*6
17.98
—
1.15
—
2.1*7
126.23
K-T/
Selexol
207. 12^
(2)
—._.,
23.55
.
3.28
5.33
33.90
6.76
1.07
.OU
5.62
286.67
NOTES:
(1) Includes $100.7 x 10 for oxygen plant.
(2) Included in gasification cost.
-------
POWER SYSTEM CAPITAL COST SUMMARY •
COSTS - $1,000
Gasifier/
Cleanup Combination
Gas Turbine-PR/Temp
FPC Account
3l+l - Structures and Improvements
3l»3 - Prime Movers (-Gas Turbine)
3l+l» - Electric Generators (Gas Turbine)
312 - Boiler Plant Equipment
3ll* - Steam Turbine Generator Units
31*5 &
353 - Accessory Electrical Equipment
3U6 - Miscellaneous Power Plant Equipment
Other Expenses
Direct Construction Costs
Contingency Engineering & Supervision
Total Construction Costs
Interest & Escalation
Total Capital Cost (Power System
Only)
BuMines/ BuMines/ BCR/Selexol BCR/CONOCO Oil/ Oil/ K-T/
Selexol Iron With No Water With Water Selexol CONOCO Selexol
With Resat. Oxide COS Conv. Scrub Scrub
l6:l/2200F l6:l/2200F 2l*:l/2600F 2l*:l/2600F 2l+:l/2600F 16:1/2200F 2l*:l/2600F l6: 1/2200:
8,558
31*, 027
11,018
31,570
21,518
10,267
399
2,31*7
119,701*
27,532
11*7,236
6k, 592
9,078
3l*,l+05
10 . 768
30,200
26,003
11,101
1*26
_2, 1*1*0
121*, 1*21
28,617
153,038
67,138
12,997
30,730
10,138
31*, 637
29,201*
12,900
1*55
2,621
133,682
30,71*7
16k, 1*29
67,7^8
12,61*8
3U,9>*2
10,6UO
37,753
26,26U
Il»,0li5
U6l
2,735
139,^88
32,082
171,570
75,268
11,579
31,353
9.5>*7
33,875
25,958
12,857
U22
2,511
128,103
29,1*61+
157,567
69,121*
10,972
36,1*35
11,798
-0,755
39,082
13,163
511
3.051*
155,770
35,827
191,597
81*, 053
13,609
33,786
10,288
36,1*1*3
35,51*6
15,112
1*96
2,096
Il*8,l86
31*, 083
182,269
79,961
9,>*8l
39,322
10,937
1*3,303
35,81*5
11,371*
1*1*2
_3.0ll*
153,718
35,355
189,073
82,91*6
211,828. 220,176 232,177 21*6,838 226,691 275,650 262,230 272,019
-------
K-T/Selexol System Costs
As discussed in the performance section, the K-T gasifier does not
mate well with a high-pressure combined cycle power system. The increased
proportion of steam power generation, fuel gas compression and oxygen
plant requirements not only affect performance but result in increased
capital costs as shown in the summary Table Uo. Costs for the K-T gasi-
fier and heat recovery system were taken from the literature.''' The
remainder of the plant costs were assembled using contractor-developed
costing procedures and data obtained during the previous contract phase
from sources such as Allied Chemical for their Selexol process.
As can be seen from the tabulated power system costs (Table U8) and
performance summary (Table Uo), the primary reason for the relatively
high cost per unit of installed power is the large proportion of steam
power generation. On the fuel gas production side, Table ^7 shows the
capital cost breakdown. Gasifier and heat recovery costs by themselves
are considerably lower than those of the other high temperature coal
fired gasifier. However, the oxygen plant costs almost double the capi-
tal investment. Once the fuel gas has been compressed, acid gas removal
costs are comparable to the other systems. However, the high power
required by the fuel compressor (approximately three times that of the
bleed air boost compressor in pressurized systems) shows up as a large
cost increase.
The resultant power costs are sufficiently higher than for the other
systems to conclude that this particular combination of gasifier and
power system is not economically practical. A low pressure power system
could undoubtedly show a real improvement but the investigation of such
a system was outside the scope of this study.
Oil Gasifier/Selexol System Costs_
For this system, fuel cost is clearly the dominant problem area.
From Table Ul it can be seen that capital cost per unit power output is
significantly less than comparable coal fired systems.
Costs for the gasification system were developed from data made
available by the TIM Division of United Technologies Corporation(37) and
reports from equipment manufacturers. '3o)
-------
Oil 'Gasifier/COHOCO System Costs
As with the partial oxidation/Selexol system the cost of fuel
represents over 50 percent of the power generating cost. As a result
the desirable performance, capital cost, and emissions characteristics
of this system are of little benefit.
Cost data were assembled in a manner similar to that for the
previous system. There is little doubt that the resultant comparison
is consistent even though we are dealing with a different fuel.
BuMines/Selexol Costs
The major changes in the BuMines/Selexol system costs are the
result of escalation and improved performance. The improved perfor-
mance results from the addition of a fuel gas resaturator which adds
water vapor and heats the clean gas to the saturation temperature,
eliminating the need for a. regenerative heat exchanger. Estimates
costs show little difference when the heat exchanger is replaced
by resaturation equipment. This is due to the reduction in heat
exchanger size when the cold side fluid is changed from gas to water.
The costs for the BuMines/Selexol with and without resaturation are
shown in Table ^9. As can be seen, the increased power out (6 percent +)
and efficiency results in a reduced cost of power compared to the origi-
nal system.
BuMines/Iron Oxide Costs
The revised BuMines/Iron Oxide system has an estimated cost of
$138.98 million (Table ^5) which is an increase of nearly 8 percent over
the base system^' with both escalated to a common mid-1975 dollar basis.
This increase is due entirely to the additional cost of the hot particu-
late cleanup system, which is only partially offset by other system
component cost reductions.
The increase in SOp concentration to 12 percent allows the
elimination of the hot potassium carbonate scrubber which was used pre-
viously to concentrate the stream (remove CCvj) prior to the Glaus system.
While this requires a slight addition to the fuel used by the Glaus
plant, the reductions in overall utilities requirements more than offset
this increase in fuel. The Glaus plant, is also increased in size, but
the revised estimates indicate a reduction in costs by 60 percent in the
overall sulfur recovery step.
181
-------
Table U9
BUMINES/SELEXOL COST SUMMARY
Without With
Capital Costs - $/kW Resaturation Resaturation
Power System 280 278
Gasification System 12^ 119
Cleanup System 101 95
Total Plant Cost 505
Owning and Operating Costs - Mills/kWhr
Owning Costs (17$ of Capital) lU.OO . 13.6U
Operation & Maintenance
Power System 1.60 1.59
Gasifier and Cleanup 3.12 2.97
Fuel Cost at 60^/MMBtu 6.52 6.32
Total Cost of Power 25.2^ 2^.52
182
-------
The net result of the performance improvement and the cost increase
is a small reduction in the cost of power. This is shown in Table 50
where the costs of the revised and base systems are compared. It should
be noted that this table differs from Table 28, which was for compara-
tive purposes and used as a basis the costs presented in the Phase
Report. (•*-' The data contained in Table 50 reflects mid-1975 dollars and
shows the effect of updated particulate removal system costs.
BCR/Selexol-Cost
The major cost reduction in the BCR/Selexol system is due to the
inclusion of a COS to BUS catalytic converter which results in a reduc-
tion of the solvent flow rate of the Selexol system. The cost reductions
have been discussed in Section 3 and will not be repeated here.
The reduced steam requirement in the cleanup can be used to increase
the power out; thus, there is a second benefit. The two systems are
compared in Table 51 a-nd the cost of power for the revised system is
shown as a function of fuel cost in Fig. 26.
BCR/CONOCO/Wet Scrub-Costs
The BCR/CONOCO system with the wet scrubber requires a good deal of
additional components. At the exit of the CONOCO system, a high-pressure
boiler and economizer would be used to drop the gas temperature prior to
entering the fuel gas to fuel gas regenerator. At the exit of the regen-
erator, the fuel gas is further cooled before passing into the wet scrub-
ber for particulate and ammonia removal. The gas is then resaturated
and sent through the cold side of the regenerator.
Those changes result in an estimated $9 million increase in cost
over the high-temperature BCR/CONOCO and introduce a performance penalty.
However, this system has the capability of meeting the most stringent
forseeable emissions standards with the use of a premixed combustor. The
implications of this are discussed below.
Comparison of Three BCR-Based Systems
Three BCR-based integrated power systems have been investigated:
(l) BCR/Selexol/COS Converter - low-temperature
(2) CBR/CONOCO - high-temperature
(3) BCR/CONOCO/Wet Scrub - high and low temperature
183
-------
Table 50
BUMINES/IRON QKIDE COST SUMMARY
With S02 Without S02
Capital Cost - $/kW Concentrator Concentrator
Power System 2?8 26?
Gasification System 117 106
Cleanup System 55* 62
Total Plant Cost h50 1*35
Owning and Operating Costs - Mills/kWhr
Owning Costs (17$ of Capital)
Operation and Maintenance 12.U7 12.06
Power System 1.59 1.52
Gasifier and Cleanup 2.38 2.33
Fuel Cost at 60^/MMBtu 6.UO 5.83
Total Cost of Power 22.Qk 21.7V
*Does not include high-temperature particulate removal
18U
-------
Table 51
BCR/SELEXOL COST SUMMARY
Without COS With COS
Capital Costs - $/kW Converter Converter
Power System 251 228
Gasification System 130 121
Cleanup System 98 85
Total Plant Cost 1+79
Owning and Operating Costs - Mills/kWhr
Owning Costs (17% of Capital) 13.2? 12.03
Operation of Maintenance
Power System 1.^3 1.30
Gasifier and Cleanup 3-l6 2.86
Fuel Cost at 60^/MMBtu 5.69 5.31
Total Cost of Power 23.56 21.5
185
-------
The costs, performance and emissions of these systems are summarized
in Table 52. It is apparent that the gap in power costs between the
BCR/Selexol and the BCR/CONOCO has narrowed from the previous study,
e.g., 2.0 mills/kWhr versus the previous U.5 mills/kWhr. This change is
attributable to several factors: (l) the additional cost for high-
temperature particulate removal, (2) the reduced cost of Selexol/Cata-
lytic system, and (3) the increased BGR/Selexol system efficiency.
Both the BCR/Selexol and the BCR/CONOCO/water scrub have essentially
the same power cost. As fuel costs increase, the BCR/CONOCO high-
temperature system does show a slight advantage.
There are definite differences in emissions. As will be discussed
in the subsequent section, the use of a premix combustor, one in which
the fuel and air are mixed prior to introduction into the combustion,
can reduce thermal NO by about 80 percent. This type of combustor has
JC
yet to be demonstrated at the operating conditions of interest to this
study. Currently, it does not appear to be possible to premix 1600 F
fuel gas with 800 F air, thus only fuel gas at 1000 F or lower can be
considered. Assuming the use of this type combustor, both the BCR/
Selexol and BCR/CONOCO/wet scrub systems would meet all forseeable
emission standards. The BCR/CONOCO has NOX emissions potentially eight
times the present limit. This emission is about equally divided between
thermal and fuel NOX, thus combustor changes beyond the premix concept
must be considered.
It must again be stated that use of a gasifier having higher
operating temperatures will solve the major portion of the fuel-bound
NOX problem, since most nitrogen compounds will be cracked in such a
gasifier. (See Section 5 for emissions from the partial oxidation
gasifier.)
186
-------
Table 52
COMPARISON OF BCR-BASED INTEGRATED SYSTEMS
Cost-$ Millions
Gasification
De s ill f ur i zat i on
Particulate Removal
Heat Recovery
Ammonia Recovery
Other Process Costs
Power System
Total
Efficiency - %
Electricity Cost - Mills/kwhr
Emissions - Ib/MMBtu
so2
NOX (as N02)
Particulates
BCR/Selexol
BCR/CONOCO BCR/CONOCO/Water Wash
91.97
26.77
*
26.28
9.86
1*6.22
222.18
1*23.28
38.6
91.97
23.00
26.63
1*1*. 21
21*6.81*
1*32.65
1*2.7
91.97
23.00
*
26.28
9.86
1*6.85
226.69
1*2H. 61*
39.0
21.5
19.5
21.
0.388 .
0.292*
< 0.01
0.553
5.50
0.553
0.581**
<0.01
*With Premix Combustor
187
-------
SECTION 5
ANALYSES OF ENVIRONMENTAL INTRUSION
SUMMARY
For each of the integrated systems examined in this study, the air,
water and solid residuals have been identified and quantified to the
fullest possible extent. A summary tabulation of all air, and solid
residuals is made in Table 53. Total recycle of water was assumed and,
therefore, none of the water is discharged. All contaminated water is
treated and reused except for water used for ash wetting which evaporates.
There are water losses from evaporation, drift, etc., but these are not
included in the tabulation under the category of water effluents.
The air emissions have been detailed previously*" ', thus emphases
have been placed on defining the water and solid emissions of these
integrated systems.
Existing EPA Standards and Their Implication Relative to This New Point
Source
EPA Standards exist for residuals from conventional coal-fired power
plants. No such standards have yet been promulgated for the integrated
combined-cycle power plants examined in this study. Therefore, a com-
parison between the quantities of residuals determined in this study and
existing EPA Standards for conventional coal-fired power plants is in
order. The EPA New Source Stack Emission Standards for power plants are:
COAL-FIRED OIL-FIRED
S02 1.2 Ib/MM Btu. 0.8 Ib/MM Btu.
NOX (as N02) 0.7 Ib/MM Btu. 0.3 Ib/MM Btu.
Particulates 0.1 Ib/MM Btu. 0.1 Ib/MM Btu.
188
-------
Table 53
SUMMARY OF RESIDUALS FROM INTEGRATED SYSTEMS
Residual
Integrated
System
Air '
so2
Source Ib/MM
Btu
Ib/MM
Btu
i
Particu- |
lates |
Ib/MM Btu |
Ash
Ib/hr
Slag
Ib/hr
Sulfur
Ib/hr
Solid
Fly
Ash
Ib/hr
Misc.
Wastewater
Residuals
Ib/hr
Spent
Dolomite
Ib/hr
Purge
Ib/hr
1) BuMines/Selexol Turbine 0.088 0.218 0.016 ! 114132
Stack 0.320 0.356 !
2) BuMines/Iron- Turbine 5-06
Oxide Stack 0.575
0.032 i 114132
co 3) BCR/Selexol
4) BCR/Conoco
5) K-T/Selexol
6) K-T/B-W
7) Oil/Selexol
8) Oil/Selexol
Turbine 0.080 1.43
Stack 0.306 0.201
Turbine 0.520 2.939
Stack 0.033
Turbine 0.057 3.41
Stack 0.490
Turbine 0.168
Stack
0.468
NA
Turbine 0.108 0.320
Stack 0.142
0.029
0.057
0.0007
Turbine 0.060 0.631 O.OOlU
Stack 0.017
24183 13019 100 to 500
20518 13619 *
60900 24086 - * _
60900 23952 - * 15244
36540 24228 24115 *
36540 23160 23873 *
350 12317 544 *
350 12249 537 * 5952
* Wastewater solid residuals were calculated only for the worst case, the BuMines/Selexol system.
+ Weight considering all oxides as N0p
7185
-------
Sulfur Dioxide Emission - From the emissions listed in Table 53, it is
evident that all the cleanup systems evaluated for first- and second-
generation application will comply with the current EPA standard for SCU
(conventional coal-fired plants). With the exception of the systems
using the half-calcined dolomite (Conoco) process, the sulfur plant stack
gas accounts for 60-90 percent of the total SOo emission. Commercially
available tailgas treating processes may be used to further reduce sulfur
emissions in these cases. Tailgas treatment in the BCR/Conoco and oil
Gasifier /Conoco cases will not significantly reduce the sulfur emission,
since at least 91 percent of the S02 results from fuel gas combustion.
Nitrogen Oxide Emission - The NOX emission from an integrated system
depends on the type of gasifier and the type of cleanup system used.
It is known that a substantial part of the nitrogen contained in the
coal is converted to ammonia during gasification. The low-temperature
Selexol system removes traces of ammonia that get past the water scrub-
bing operation. The high-temperature iron oxide and Conoco processes do
not significantly affect ammonia at all. Therefore, assuming essen-
tially conventional gas turbine combustors, even if only 50 percent of
the ammonia were converted to NOX during combustion, these systems would
not meet the current EPA standard. Thus, high-temperature desulfuriza-
tion systems used in conjunction with gasifiers operating below the
cracking temperature of ammonia are inadequate from a NOX control view-
point, although they yield higher thermal efficiencies for the integrated
systems.
Particulates - The current EPA Standard for particulates is 0.1 Ib/MM
Btu, which is equivalent to approximately 0.1 gr/SCF for a typical low-
Btu gas. This level of particulate loading is easily achieved by all
the low-temperature cleanup systems considered here. Although defini-
tive operating data are lacking for high-temperature particulate removal
systems and for the particulate loading/size distribution in the fuel
gas produced by various gasifiers, it is expected that these systems
will be capable of meeting the existing EPA Power Plant New Source
Standard.
However, the concern for particulate removal is dictated by gas
turbine operating requirements, rather than emission standards. With
the potential for turbine blade erosion in mind, the allowable partic-
ulate loading for the clean fuel gas was established to be 0.01 gr/SCF.
Even this stringent goal can easily be met by low-temperature cleanup
systems with their inherent water scrubbing operations. However, it is
not certain at present whether high-temperature particulate removal
systems will be able to meet this stringent operating requirement. A
number of systems such as high temperature electrostatic precipitators,
190
-------
metallic mesh filters, and granular bed filters claim a high degree
of particulate removal (95 percent) but have yet to be proven on a
commercial scale.
Water - With total water treatment and reuse, liquid effluents can be
completely eliminated, and most impurities that were transferred from
the raw fuel gas to the process water can be removed after treatment,
and disposed of in solid form. Though theoretically possible for this
type power plant facility, this has to be demonstrated and zero liquid
effluents have to be proven.
Solids - Considerable quantities of solids including ash, slag, sulfur,
etc. are generated from the integrated systems. Most of these solids are
stable at ambient conditions and can be utilized or disposed of. Those
that are unstable are first treated to render them stable before disposal.
From an environmental consideration however, it seems that the quantities
of these solid residuals rather than their nature might present the
greater problem.
OVERVIEW
To this point, this report has described the nature of the unit
operations and systems associated with low- and intermediate-Btu
gasification of coal and the subsequent cleanup of the fuel gas pro-
duced by the gasification process. This section presents results
extracted from the material balances and other appropriate data
information concerning the effluents, emissions, and solid wastes
produced by low- and intermediate-Btu gas-fired combined-cycle power
plants. Further, this section will describe possible options assoc-
iated with the ultimate disposition of the residuals produced by these
plants.
From the standpoint of overall environmental considerations, it
appears that advanced low- and intermediate-Btu gas-fired combined-
cycle power plants can potentially produce fewer insults to the
environment than the conventional coal-fired power plant with flue
gas desulfurization (FGD). This advantage stems from two major fea-
tures of the combined-cycle power plant. First, because a major
portion of the electricity being produced by this power plant is pro-
duced by the gas turbine portion of the combined cycle, the amount of
heat rejected to water and hence, the amount of cooling water required
for the combined-cycle power plant is about half that required for the
conventional coal-fired plant using some form of water cooling. Second,
with the integrated combined-cycle power plant, the sulfur that origi-
nates in the coal and ends up in the fuel gas can be removed without
191
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producing any solid waste other than the elemental sulfur itself. With
the conventional coal-fired plant, either the not-easily-disposed-of
scrubber sludge is produced, or a regenerative scrubbing process is used
with a significant reduction in plant efficiency. While the scrubber
sludge can be disposed of by a number of proven processes, it is a
costly undertaking which the power companies would rather avoid. Unfor-
tunately, if a regenerable FGD process is employed, the plant is derated,
requiring more coal to be burned to produce the same amount of power.
This results in greater air emissions, water effluents and solid residuals.
From the previous discussions and analyses it is apparent that there
are numerous low- temperature and several high-temperature fuel gas clean-
up systems that are efficient in removing I^S from the fuel gas, just as
there are scrubbers that are efficient in removing SC>2 from flue gas. The
resulting sulfur oxide emissions from the integrated combined-cycle power
plant, therefore, could be controlled to a level comparable to conven-
tional coal-fired plants with stack gas scrubbing. Since there is less
fuel gas than flue gas volume (there is both less fuel gas mass and its
pressure is higher) for an equivalent size power plant, the fuel gas
cleanup process would probably be more easily controlled and more reliable .
The magnitude of these advantages though, would have to be assessed by
comparing specific systems. Also, the benefits of sulfur and ammonia
recovery units must be weighed against added complexity and additional
emission streams.
In actual practice, the particulate emissions from a conventional
coal- fired plant would be higher than those from the integrated combined-
cycle power plant. This is not through any inherent advantage of the
integrated combined-cycle power plant, but rather because of the limited
particulate loading that can be tolerated by the gas turbine. Stated
another way, because of the potential for erosion damage to turbine
blades, the particulates entering the turbine, and hence the particulates
in the exhaust of the combined-cycle power plant must be controlled to
levels lower than the particulate emission levels of the conventional
coal- fired power plant. On the other hand, it may be necessary to reduce
allowable emissions since the particulates from the combined cycle tend
to be smaller and more harmful.
While it appears ' that the NOX emissions from the low-Btu gas-fired
combined cycle power plant utilizing a low- temperature fuel gas cleanup
system can be controlled to levels comparable to or less than the conven-
tional coal- fired power plant, it is not clear as to whether this would
also be true in the case of high-temperature fuel gas cleanup systems
used in conjunction with gasification processes that operate below the
192
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cracking temperature of ammonia. This is because the ammonia produced
by these gasifiers cannot be removed at elevated temperatures with
existing technology. As has been discussed earlier in this report, it
appears as though this problem could be remedied if a suitable catalyst
for promoting the decomposition of ammonia could be developed. With
higher temperature gasification processes, like the Koppers-Totzek pro-
cess and oil-based partial oxidation process, ammonia formation is
negligible, thereby eliminating the necessity of its removal.
Thermal NOX, i.e., NOX formed from atmospheric nitrogen during
the combustion process may also be increased becuase the use of high-
temperature fuel gas raises the stoichiometric flame temperature.
However, thermal NOX production is subject to reduction by careful
combustor modifications and will not be discussed here other than to
be evaluated as part of the total NOX emissions.
As will be described later in this chapter, all of the water
emissions from this integrated power plant facility can be treated so
that there will be no release of water borne pollutants to a natural
body of water. It will be shown that the integrated power plant could
be operated in full compliance with the effluent controls currently
mandated for 1983 with respect to the Best Available Technology
Economically Achievable.
While at this time it appears as though there are no national
shortages of the materials that might be produced by the combined-
cycle power plant as solid .wastes, there are constructive ways in
which these residuals can be utilized. Further, research is being
conducted that strives to broaden the range of applications for these
residuals. Ash and slag utilization have been the subjects of study
for quite a few years and recent environmental regulations limiting
sulfur emissions have prompted considerable study in the area of con-
structive uses for sulfur. A number of these existing and proposed
applications for ash, slag, and sulfur are identified and discussed in
this section.
193
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AIR EMISSIONS
The air emissions to the surroundings can be divided into two
groups, those from the fuel processing system and those from the power
system. However, it must be remembered that the power system emissions
are very much a function of the fuel processing and therefore, prior to
the discussion of power system emissions, it would be worthwhile to
briefly review the overall fuel processing system to identify potential
pollutent sources.
Review of Fuel Processing System
Gasifiers - As has been noted earlier, there are two generations of
gasifiers. The first generation (e.g., BuMines, Lurgi) has off-gases
condensible tars, phenols, and other organics. The second-generation
gasifiers (e.g., BCR-two-stage, Koppers-Totzek and Texaco/Shell partial
oxidation) have off-gases without condensibles. The principal differ-
erences in the performance of these gasifiers is directly related to the
operating temperatures within the gasifier. The operating temperature
in a fixed-bed gasifier is lower (< 1200 F) than in an entrained-flow or
in a fluidized-bed gasifier (2500 F) and therefore the carbon conversion
is realtively low. The temperature is low enough to allow the formation
and preservation of tar, phenols, and other organics. Organics produced
during gasification do not undergo thermal cracking and therefore emerge
with the product gas. The lower temperature also favors the formation
of ammonia.
In the entrained-flow and fluidized bed gasifiers, not only are the
temperatures higher, but the residence times are shorter. As a result,
the formation of organics and tars is not favored and therefore, these
are not present in the raw gas.
The conditions in the three types of gasifiers viz, fixed-bed,
fluidized-bed and entrained-flow are different and the quantities and
size distributions of the particulates off each gasifier type are cor-
respondingly different. The conditions differ in:
• The manner in which the coal feed is supported
• The rate of gas flow (superficial velocity)
• Temperature
• Feed size
19U
-------
Based on the above differences, a qualitative estimate of
particulate size and quantity can be made. The particulate loading of a
fixed-bed is estimated to be fairly high and comprised of fine particules
of ash and unburned carbon. Particulate loading in the gas from a
fluidized bed is lower and consists of comparatively larger particles
than off a fixed-bed gasifier. The particulate loading off an entrained-
flow gasifier is very high with its size distribution proportional to
the feed size.
Cleanup Processes - This study has addressed the following gasifiers:
• Bureau of Mines (stirred, fixed-bed) gasifier
• BCR (pressurized entrained-flow) gasifier
• Koppers-Totzek (atmospheric entrained-flow) gasifier
• Partial Oxidation (Texaco/Shell entrained-flow) gasifier
Of these, only the BuMines gasifier is a low-temperature gasifier
and has condensible tars, phenols, ammonia and organics besides particu-
lates in the off-gases. The simplest scheme of removal of these conden-
sibles involves a water quench followed by the separation and recycle of
the condensed tar to the gasifier in order to improve thermal efficiency.
The quench also removes phenols, organics and some ammonia from the gas.
Further scrubbing with water removes remaining traces of ammonia and
some hydrogen sulfide from the gas. Thus, a low-temperature cleanup
system is a logical choice for the BuMines gasifier. If however, a
high-temperature cleanup system is used to improve overall thermal effi-
ciency, the product gas will contain tar and ammonia which upon combus-
tion will give S00 and NOV.
C. A
All of the other gasifiers studied belong to the second generation
(high temperature) so that their product gases contain no condensible
tars or organics. Therefore, high-temperature cleanup systems can be
used in conjunction with them and higher thermal efficiencies can be
attained than with low-temperature cleanup systems. Ammonia is present
in smaller quantities, but none the less presents a removal problem at
elevated temperatures. With current state-of-the-art cleanup systems,
ammonia can be removed only at lower temperatures by scrubbing with
water.
From some gasifier types, a significant amount of sulfur is present
in the form of COS which could constitute a removal problem at lower
temperatures. Specifically, the Selexol solvent is not selective to COS
and therefore the addition of a catalytic conversion process changing
195
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the COS to ftjS prior to desulfurization was considered. The COS is
fairly efficiently removed by high-temperature desulfurization processes
such as dolomite absorption.
Although there are currently no commercially available particulate
removal processes which could operate at the temperatures and pressures
required for high-temperature cleanup and meet the projected removal
requirements for particulates < 10 [j,, there are several processes in the
early development stages which show great promise. Therefore, it is not
unrealistic to assume that either high- or low-temperature particulate
removal systems may be used with equal success depending on the operating
temperature of the cleanup system downstream.
Emissions Associated with the Fuel Processing Systems
Glaus Plant Tail Gas - The efficiency of a Glaus plant is generally less
than 95 percent, therefore, some of the EpS or S0? fed to it remains
unconverted and has to be vented to the atmosphere as S02- Unconverted
I^S is incinerated to S02. Much of the fuel for incineration is needed
to raise the temperature of the noncombustiles in the Glaus plant feed
gas. Thus, fuel requirements are largely a function of the cleanup sys-
tem and increase as the H>>S concentration in the feed decreases. For a
typical system (BCR/Selexol with Catalytic Reduction of COS) having a 23
percent concentration of HpS in the Glaus plant feed, fuel requirements
are approximately 1 percent of the total gas produced. Other components
of a Glaus plant tail gas are NO, C02 and H^O. The extent of NOX emis-
sion depends primarily on the quantity of ammonia in the Glaus plant
feed which in turn is a function of the quantity of ammonia removed from
the fuel gas. The quantities of C02 and BpO are a function of the quan-
tity of fuel used in the Glaus plant. This requirement is higher for a
S02 Glaus feed as compared to an HpS feed. The Glaus plant emissions
are shown in Table 5^4- for the different systems.
Commercially available tail gas cleanup process such as the SCOTT
process or the Strethford process could reduce the Glaus plant emissions
but at additional capital and operating costs. Since the overall plant
is below the 1.2 Ib/million Btu limit for coal-fired plants, these pro-
cesses were not included although estimates of the reduced emissions are
included in Table jk.
It should be noted that the process changes in the BuMines/iron
oxide and the BCR/Selexol have resulted in somewhat different emissions
for the Glaus plants associated with those cleanup processes than have
been reported previously.'^-' These changes are noted in Table 5U.
196
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Table $k
AIR EMISSIONS FROM INTEGRATED SYSTEMS
(Ib/MMBtu)
SYSTEM
Pollutent/Source
S02
Fuel Processing
Power System
Total S02
Total With
Claus S?2 F2D
NOX
Fuel Processing
Power System
Thermal
Fuel Bound
Total 110
BuMines/
Selexol
0.320
0.088
0.1+08
0.137
0.356
O.OM+
0.17k
0.57U
BuMines/
Iron Oxide
0.575 (0.535)2
0.33^ —
0.909
0.372
0.310
1.273
3-79
5.17
BCR/
Selexol
0.306 (0.1+87)2
0.080
0.388
0.103
0.201
1.1+23
0.012
1.633
BCR/
Conoco
0.033
0.520
0.553
2.563 (l.^lJ
2.7k (.296)*
5.50
K-T/
Selexol
O.U90
0.057
0.51+7
0.09U
' i3
3.U1
K-T/
BEW
O.U68
0.168
0.636
0.203
NA
-___
Oil /
Selexol
0.108
0.250
0.3073
0.013
0.320
Oil/
Conoco
0.017
0.060.
0.077
0.5923
0.039
0.631
Total Particulates
<0.0l6
<0.032
<0.01
-------
S02 Scrubber Flue Gas - In the BuMines/iron oxide system, the regeneration
of the iron oxide absorbent yields S02 in the off-gas. The concentration
of S02 in the off-gas can be low (5 percent) depending on the amount of
excess air used to regenerate the sulfided bed. If the S02 concentration
is very low, the Glaus plant efficiency drops greatly. To circumvent
this, the S02 must first be concentrated and fed to the Glaus plant.
This represents an additional step in the process. One alternative to
using a Glaus plant and an additional unit for concentrating the flue
gas is to use a flue gas scrubbing system. A flue gas scrubber may also
be used to scrub the Glaus plant tail gas to further reduce S02 emis-
sions. Such systems are generally limited in their efficiency to about
90 percent S02 removal. The remaining S02 along with C02, 1^0, and NOX
are discharged to the atmosphere.
Coal Feed System (Lock Hopper) Releases - The coal feed system to a
pressurized gasifier typically consists of a weigh hopper, a pair of
lock hoppers and a pressurized feed hopper. Pulverized coal from the
storage bin is fed to the weigh hopper which discharges a measured
quantity alternately into two lock hoppers. When one of the lock hop-
pers is filled to capacity it is pressurized with a coal transport gas
(this may be a portion of the product fuel gas). The coal is then con-
veyed to the feed hopper from which it is fed to the gasifier. The
emptied lock hopper is then vented and refilled while the second hopper
is pressurized. This sequence of operations allows a continuous flow of
coal into the gasifier. The feeding operation results in emissions to
the air.from the lock hoppers when they are vented. These emissions con-
sist of coal fines, as well as some small amount of the transport gas
(which may contain pollutants).
Gas Released by Fuel Gas Quench Water Sent to Water Treatment Facilities -
When pressurized fuel gas is quenched, some amount of gas dissolves in
the quench water. The quench water also picks up particulates and sol-
uble organic and inorganic compounds such as phenols, ammonia, etc. that
may be present in the fuel gas. The contaminated quench water is then
sent to water treatment facilities which operate at atmospheric pressure,
so that some of the gases that had dissolved under pressure are now
released. Therefore, this constitutes a potential source of air emis-
sions including HgS, NHo? C02, etc.
Gases Released from/with Gasifier Bottom Ash and Slag - The ash (from •
the BuMines gasifier) is removed from the gasifier through a pressurized
lock hopper system. When the lock hopper is filled to capacity it is
depressurized and the ash is removed for cooling and disposal. The ash
is accompanied into the lock hopper by some raw gas containing pollu-
tants which upon venting the lock hopper, is emitted to the atmosphere.
198
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The second-generation gasifiers produce slag which is collected in a slag
pot and is water quenched. The quenched slag is then removed from the
gasifier via two slag hoppers. Once again, some raw gas escapes with the
slag and is vented to the atmosphere. Some raw gas is also dissolved in
the slag which is evolved once the slag is removed from the gasifier.
Emissions Associated with the Power System
Sulfur Dioxide - As in the integrated systems previously discussed*1 ',
the SC>2 emissions from the K-T and the partial oxidation residual oil-
fired systems now being considered are within the regulation for S02
from coal- and oil-fired steam stations. The emissions from those inte-
grated systems using low-temperature cleanup are lower, especially in
the case of the fixed-bed gasifier (BuMines) since, when using the high-
temperature cleanup, the tars containing significant amounts'of sulfur
are passed through the cleanup to the combustor. Values of SC>2 emissions
are given in Table 5^.
Nitrogen Oxides - The production of NO is through two independent
mechanisms, thermal NOX which is a function of local temperature and
time, and fuel WO which is a function of fuel-bound nitrogen.
X
Since thermal NOX is directly proportional to combustion temperature,
anything that increases combustion temperatures would increase the NOX
production; a decrease in temperature would decrease the NO production.
It has been assumed that the gas turbine combustors used in the integrated
systems are of the conventional type, i.e., a local stoichiometric flame
zone in the front of the can followed by rapid quenching with dilution
air. The thermal N0__ for this type system is thus a function of the
X
stoichiometric flame temperature. The factors affecting this parameter
are l) fuel heating values, 2) fuel sensible heat (fuel temperature), and
3) combustion air temperature. Previous work at UTRC^' 30 > an'3- 39) has
resulted in the preparation of working charts for the determination of
thermal WO as a function of combustion temperature and gas turbine fir-
ing temperature. In Fig. 4l, for example, the effects of both chemical
and sensible heat on combustion temperatures are given. Thermal NOV
X.
emissions as a function of combustion temperature can then be estimated
by using Fig. 4-2. Unfortunately, the values of rate constants used to
develop Fig. 42 are not well defined and it is estimated that the values
shown are probably only within a factor of +2 of actual values. As
examples of the use of these tools, Table 55 has been constructed.
Without regeneration, the fuel gas from the BCR/Selexol system would
have a temperature of about 250 F giving a final combustion temperature
of about 3680 F. This would result in a NO emission of approximately
199
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FIG. 4,1
EFFECT OF FUEL GAS CHEMICAL AND SENSIBLE HEAT ON COMBUSTION TEMPERATURE
REFERENCE FUEL HHV = 120 BTU/SCF
REFERENCE FUEL TEMPERATURE = 80F
STOICHIOMETRIC FUEL - AIR RATIO
INITIAL AIR TEMPERATURE = 825F
4400
GC
I
LU
tr
DC
LU
Q.
g
jo
00
O
o
o
1-
co
Q
4200
4000
3800
3600
3400
INCREASE IN FUEL
TEMPERATURE
INCREASE IN FUEL HHV
100
120 140 160 180
FUEL CHEMICAL PLUS SENSIBLE HEAT - BTU/SCF
200
76-02-91-2
200
-------
FIG. 42
NITRIC OXIDE FORMATION IN GAS TURBINE BURNER
1000
CONSTANT BULK GAS FLOW RATE
FIXED BURNER VOLUME
STOICHIOMETRIC FUEL/AIR RATIO IN RECIRCULATION ZONE
BURNER PRESSURE = 12.5 ATM
100
QL
Q.
CO
CO
01
g
X
o
o
cc
10
1.0
0.1
^. HIGH-BTU +_
FUELS
3400
3600 3800 4000 4200
MAXIMUM COMBUSTION TEMPERATURE
4400
- R
4600
76-02-91-1
201
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TABLE 55
COMBUSTION TEMPERATURES FOR FUEL GAS
System BuMines/Selexol BCR/Selexol
Fuel Gas Heating Value, Btu/SCF lUl.7 156.5
(HHV)
Fuel Gas Temperature, F 252 1000
(Regeneratively heated)
Air Temperature, F 755 903
Combustion Temperature, F 3^50 3580
(Fuel at 80 F, Air at.825 F)
Correction for Fuel Temperature, F + 6k + 360
Correction for Air Temperature, F - 1^1 + k-6
Approximate Combustion Temperature, F 3^-75 3980
Turbine Inlet Temperature, F 2200 2600
NO.. Emission, ppm 13 500
Jt
NOX Emission, Ib/MMBtu (as N02) 0.0k 1.^2
202
-------
0.25 Ib/MMBtu. However, without the regenerative fuel heating, the
overall performance of the system would decrease approximately five
percent. When the fuel-bound nitrogen is added, the total NOX emission
would be (90 percent fuel-bound nitrogen conversion) approximately 0.33
Ib/MMBtu. If the regulation of 0.7 Ib/MMBtu were to be equaled, then
the regeneration would be limited to approximately 750 F.
There is a second approach to reducing the thermal NOX. This
approach involves combustor modification. While the details of these
modifications are beyond the scope of the present study, briefly they
are aimed at burning at off-stoichiometric conditions, i.e., at lower
than stoichiometric flame temperatures. The Turbo Bower and Marine
Systems subsidary of United Technologies Corporation has been carrying
out a series of tests on low- and medium-Btu buel gases produced by an
experimental gasifier at the Texaco Development Company's Montebello,
California research facility. Of particular interst are the results of
the use of premixed (fuel and air mixed prior to the introduction into
the combustion) burners. Figure k3 shows the NO emissions as a func-
tion of source temperature rise for several values of fuel chemical
heating value. In the test series, the fuel gas was delivered at essen-
tially ambient temperature. The theoretical stoichimetric temperatures
are given. Also shown in Fig. h3 are the approximate emissions for
burners having premix conditions. While no significant reduction was
noted for low-Btu gas (the emissions were already low), use with the
medium-Btu gas indicated a large reduction. The reduction as a function
of equivalence ratio (local fuel/air ratio divided by stoichiometric
fuel/air ratio) is shown in Fig. hh.
While it will require experimental verification at the appropriate
operating conditions, it appears that a potential thermal NX) reduction
X
of approximately 80 percent may be realized by premixing. Unfortunately,
premixing cannot be applied to fuel gases much above 1200 F because
of self ignition. However, the lower temperature gases may be regener-
ated to 1000 F and premixed at off-stoichiometric conditions thereby
allowing high performance without undue NOV problems. Because these
X
values need to be experimentally verified, the NO emissions of Table 53
do not reflect any of the improvements felt possible.
Unfortunately, the fuel-bound nitrogen does not appear to be as
ameanable to treatment by combustor modification. As was done in the
previous study^', the assumption of 90 percent conversion of the ammonia
to NO was made. Thus, the values of NO given in Table 5^- indicate
higher than acceptable levels of NOV from those combinations of gasifiers
X
and cleanup systems not having sufficiently high-operating temperature
to decompose ammonia or without an aqueous scrubbing system.
203
-------
NOX PRODUCTION FROM COMBUSTORS BURNING
LOW-BTU AND MEDIUM-BTU GAS
FIG. 43
60
50
40
IMOX (PPM) 30
20
X - OFF STOICHIOMETRIC MEDIUM BTU FUEL
O - OFF STOICHIOMETRIC LOW-BTU FUEL
MEDIUM-BTU DATA BAND
(300 BTU/SCF)
NATURAL GAS DATA BAND
(1000 BTU/SCF)
LOW-BTU DATA BAND (100 BTU/SCF)
I I I
900
1000
1100 1200
BURNER AT
1300
"FLAME
= 4450 R
'FLAME
= 4320 R
TFLAME
= 3650 R
1400
76-02-91-4
-------
FIG. 44
THE EFFECT OF EQUIVALENCE RATIO ON NOX EMISSIONS
Crt
OL
o
to _
D CC
CO O
QC
<
g
X
o
z
90
80
70
60
50
O 40
30
20
0.4
I
I
0.5
0.6 0.7 0.8
EQUIVALENCE RATIO,
0.9
1.0
76-02-91-3
205
-------
Particulates - It was stated in an earlier portion of this report
(Section 2) that the ability of the turbine to operate satisfactorily
for reasonable periods of time was very much a function of particulate
removal. It was also pointed out that because there is very little
interfacing between the particulate removal device (either high- or low-
temperature) and the remainder of the system, little definition of
operating characteristics is necessary. Thus, based upon data reported
previously^) for systems having aqueous scrubbing, particulate carry-
over in the fuel gas meets the stringent turbine requirements given
previously, and in turn would easily meet the 0.1 Ib/MMBtu EPA limit.
Based upon the limited data available on high-temperature and high-
pressure cleanup systems, these systems are capable of removing small
particles to levels approaching the turbine requirements. For example,
small-scale cleanup systems at Argonne National Laboratory^ ' operating
at 8-10 atm and 1500 F - 1700 F shows that grain loadings of < 0.001 gr/
SCF could be attained. This was accomplished with two stages of
cyclones and two stages of final mesh filter. Submicron particulates
were removed. It should be noted that metallic trace elements such as
Pb, Na, Ca, etc. which are harmful to gas turbines tended to agglomerate
on the fine particulates and, thus, were removed from the process stream
with the particulates. While this removal feature has yet to be demon-
strated on a large scale for high-temperature (1600 F) cleanup, it is
hoped that this phenomena will continue to occur.
206
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WATER EFFLUENTS
As a preliminary to any discussion on wastewater treatment, it is
necessary to identify the sources of discharge, to characterize the
water to be treated, and to define the end use of the treated water. To
address the last problem first, it is generally agreed that in the con-
text of minimizing water consumption, all effluent water streams should
be treated and reused within the boundaries of the plant. As a corollary,
utilization/treatment schemes should not be primarily designed to return
water to a river or to the land in any other way.
Coal contains many trace impurities, which if concentrated are
toxic. Our knowledge of the trace elements is imperfect, and regulations
relating to the discharge of many elements do not exist at this time.
Faced with a lack of regulatory guidance and an incomplete picture with
regard to the dangers that might be associated with liquid waste streams,
the disposal of waste material should be, as far as possible, under con-
trolled conditions. Every effort should be made to remove wastes as
solids. This is not an absolute requirement, but an ideal against which
various water treatment schemes can and should be rated. Since all coal
conversion processes are net consumers of water which leaves the plant
as vapor, hydrogen gas or as hydrocarbons, total wastewater reuse is
theoretically possible.
Waste Water Sources
Wastewater effluents produced by a low-Btu coal gasification
combined-cycle electric power plant can result from a number of unit
operations. Some wastes are discharge continuously as long as the
plant is operating. Some wastes are produced intermittently on a fairly
regularly scheduled basis, such as daily or weekly, but are still associ-
ated with the production of electrical energy. Other wastes are also
produced intermittently, but at less frequent intervals and are generally
associated with either the shutdown or startup of coal processing or
electricity generating units. Additional wastes generated are essentially
unrelated to production, but depend on meteorological or other factors.
Wastewaters are produced relatively continuously from the following
sources (where applicable): gasifier raw gas cleanup systems, cooling
water systems, ash handling systems, wet scrubber air pollution control
systems, and boiler blowdown. Intermittently, but on a regular basis,
wastewater is produced primarily by water treatment operations which
utilize a cleaning or regenerative step as part of their cycle such as
ion exchanger regeneration, filter backwashing, and clarifier blowdown.
207
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Wastewater effluents are also produced by the cleaning of major units
of equipment on a scheduled basis either during maintenance shutdown or
during startup of a new unit. The efficiency of coal gasification and
electricity generating plants is largely dependent on the cleanliness of
their heat transfer surfaces. Internal cleaning of this equipment is
usually done by chemical means and requires strong chemicals to remove
deposits from these surfaces. Moreover, the cleaning is not successful
unless the surfaces are cleaned to bare metal which in turn means that
some metal has to be dissolved in the cleaning solution.
Finally, rainfall runoff results in drainage from coal piles in the
storage area, from floor and yard drains, and from construction activity.
Process Condensates - Process condensates is the name given to wastewaters
that have contacted coal or tar. They are produced in the raw gas clean-
up system when the gas is cooled and cleaned to remove impurities and by-
products associated with the gasification of coal or oil. Process con-
densates are generated only from processes which utilize low-temperature
gas cleanup systems. When a high-temperature gas cleanup system is used,
• only sulfur containing compounds and entrained solid impurities are
removed (unless a future high temperature nitrogen compound removal sys-
tem is used); the remainder is discharge to the atmosphere after the
product gas and combustible impurities have been oxidized in the turbines
(unless additional cleanup proves necessary).
Low-temperature gas cleanup systems, on the other hand, are designed
to remove all the materials generated in the gasification that are not
compatible with the product gas. Since such systems operate at or below
the ambient temperature, condensible materials are removed from the gas
stream and are discharged from the process as liquid effluents. The
organic phase which consists primarily of tar and oil is returned to the
process, whereas the aqueous phase is conveyed to the wastewater treat-
ment plant for by-product recovery and water purification.
Low-temperature gas cleanup systems have the greatest potential for
water pollution. The gasifier output may contain all of the products
commonly associated with pyrolysis, carbonization, and coking of coals
in addition to oxygenated products associated with partial combustion.
Hence a broad spectrum of heavier materials present may be classified as
tar, including phenols and cresols, pyridines, anilines, dihydric phenols,
intermediate and high boiling aromatics, saturates, olefins, and thio-
phenes. Another grouping termed light oil and/or naphtha, include B-T-X,
naphthalene, thiophene, and condensible hydrocarbons and carbon disul-
fide. Ammonia, hydrogen cyanide, coal, char, ash fines and trace metals
will also be present.
208
-------
The particular distribution of compounds which will be present in
the raw fuel gas will, of course, depend on the composition of feed coal
and on the particular conditions of the gasification. The composition
of the raw gas will determine the characteristics of the wastewater
effluents. In general, gasification processes are classified into three
categories according to their operating temperatures. These include the
low-operating-temperature fixed-bed gasifiers, the intermediate-operat-
ing-temperature fluidized-bed gasifiers, and the high-operating-tempera-
ture entrained-bed gasifiers. Since the amount and variety of undecom-
posed organic matter that will escape with the raw gas are largely
dependent on the gasifier operating temperature, it is evident that
fixed-bed gasifiers generate the "dirtiest" raw gas. Table 56 shows the
chemical characteristics of wastewater effluents produced by the raw gas
cleanup systems of the Synthane, Lurgi, and Bureau of Mines coal gasifi-
cation processes. The numerical values of the Bureau of Mines/Selexol
effluent were estimated and represent water which has been steam stripped
to recover ammonia. For comparison purposes, Table 56 also shows a
representative chemical analysis of weak ammonia liquor from a coke plant.
Trace elements which are present in 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 of the
process condensate from gasification of an Illinois No. 6 coal via the
Synthane Process^ '(Table 57). These data were presented previously(1)
and are repeated here for convenience. Of particular concern are those
elements identified by the EPA as hazardous to human health: beryllium,
fluorine, arsenic, selenium, cadmium, mercury, and lead. These elements
are all volatile and can be expected to appear in the raw gas and ulti-
mately in the wastewater stream.
It should also be noted that some of the polynuclear hydrocarbons
which may be present in raw gas have exhibited carcinogenic properties
in animal studies. Control of such materials will generally be required
in connection with evaporation from the wastewater treatment system, in
plumes from cooling towers if leakage from the process train occurs, in
the direct handling of separated tar or oil products, and in the flue
gases from coal or tar combustion.
Cooling System Blowdown - In the operation of a closed cooling system,
the bulk of the warm circulating water returning to the cooling system
is cooled by the evaporation of a small fraction of it. The amount of
water lost due to evaporation is a function of the temperature difference
of the water between the inlet and outlet of the cooling system.
209
-------
Table 56. CHEMICAL CHARACTERISTICS OF PROCESS CONDENSATE
COMPONENT
Total ammonia
Total sulfur
phenol
Thiocyanate
cyanide
Fatty acid
chloride
carbonate
COD
BODc:
j
sulfides
Heavy metals
PH
COKE PLANT
1800-4300
0-50
410-2400
100-1500
10-37
1200-2700
2500-10,000
8.3-9.1
SYNTHANE LURGI
7000-10,000 1050
1400
2600-6600 . 500
20-200
0.1-0.6
1750
500 500
17,000
15,000-38,000
8.6-9.2
BuMINES/SELEXOL
200-400
500
100
1-10
500
250
2500
10-100
10-20
9
1. Values are ppm except for pH and do not represent a complete
characterization of the condensate.
210
-------
Table 57. TRACE ELEMENT ANALYSIS OF ILLINOIS COAL
1 ,2
ELEMENT ILLINOIS COAL WASTEMTER
Al 1.29% 1000 ppb
Ca 0.77 4000
Cl 0.14
Fe 1.92 3000
K .0.16 160
Mg 0.05 2000
Na 0.05
Si 2.49
Ti 0.07
As 14.0 ppm 30 ppb
B 102.0
Be 1.6 130
Br 15.4
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 49.4 40
Mo 7.5
Ni 21.1 30
P 71.1 90
Pb 34.8
Sb 1.3
Se 2.1 360
Sn 4.8 20
V 32.7 3
Zn 272.3 60
Zr 72.5
1. Mean value for 101 coals analyzed.^1'
2. Process condensate from gasification of Illinois No. 6 coal.
211
-------
Roughly, evaporation losses amount to approximately one percent of the
circulating water for each 10 F drop assuming a latent heat for water of
1000 Btu/lb. Additional water is lost to the atmosphere as a result of
entrainment of water in the air draft (drift loss). The amounts of
drift losses depend on the cooling system used varying from up to five
percent of the circulating water for spray ponds to approximately 0.1
percent for forced draft cooling towers. Because of the water losses
due to evaporation, the remaining water becomes more concentrated with
dissolved solids. If the concentration level of any of the soluble
salts exceeds its solubility level, the salt will precipitate. Some
of the salts are characterized by reverse solubility, that is, their
solubility decreases with increasing temperature. When cooling water
saturated with such a salt is heated in the process condensers, the salt
will deposit as a scale on the condenser tube walls and hinder heat
transfer across the tubes.
Scale formation is usually controlled by discharging a portion of
the circulating water from the cooling system to prevent a buildup of
high dissolved solids concentration. This bleeding process, which is
referred to as cooling system blowdown, is carried out either continuously
or intermittently. The amount of blowdown is a function of the number
of concentration cycles, that is, the ratio between the content of the
critical component in the circulating water and the makeup water. This
is also known as the number of concentrations. If it is assumed that all
components in the feedwater must leave the system in the blowdown and
enter only through the makeup, then the makeup flow, MI times its criti-
cal component concentration must equal the blowdown, Bj times its criti-
cal component concentration. This results in the following relation
where C, the number of concentrations is equal to the ratio of critical
component concentration in the blowdown to its concentration in the feed:
C - I (17)
The makeup, M, is the sum of the water lost due to evaporation, drift and
blowdown. Blowdown can be calculated as the sum of water
withdrawn for that purpose plus drift losses although this latter quan-
tity is quite variable and in practice a conservative value of zero
drift may be assumed.
A variety of chemical additives may be used to treat water circulat-
ing in the cooling system to control scaling, erosion, and fouling.
These additives will appear in the blowdown along with matter originally
present in the makeup stream. Biological growth in the circulating
water is usually inhibited by chlorinating the water. Cooling waters
212
-------
axe very often acidified with sulfuric acid to increase the solubility
of the dissolved solids, and subsequently, to lower the makeup require-
ments due to blowdown. Pentachlorophosphate is cometimes added to
cooling water to inhibit fungi attack on wooden cooling towers.
There may be particular problems associated with leakage into the
cooling system from the high pressure gas processing train. Such leak-
ages, if they occur, will also be present in the cooling system blowdown.
Steam Cycle Blowdown - A major problem associated with the operation of
boilers or waste heat recovery systems is the formation of scale. The
primary cause of scale formation is the reverse solubility of many of the
scale forming salts. The higher the temperature and pressure of boiler
operation, the more insoluble the scale forming salts become. Calcium
and magnesium salts are the most common ingredients of boiler scales.
Calcium deposition is primarily due to the thermal decomposition of
calcium bicarbonate according to the following equation:
Ca(HC03)2 = CaCO-(S) + C02 + HgO (18)
Deposits of iron oxide, copper oxide and other metallic oxides are
frequently found in boilers operating with very pure feed water. The
source of these deposits is corrosion caused by the action of dissolved
oxygen and carbon dioxide.
Boiler blowdown is the most widely used control method against scale
formation. The amount of blowdown required is a function of the allow-
able concentration of scale forming or other undesirable components in
the boiler and the degree to which the makeup water is cleaned. As for
the cooling towers, the rate of allowable concentrations to makeup con-
centration of the critical component determines the number of concentra-
tion cycles which defines the ratio of makeup to blowdown (Eq. 17).
High pressure boilers have quite stringent limits on contaminants. For
example, the allowable concentration of silica varies from 125 ppm at
pressures under 300 psig down to 0.5 ppm at pressures in excess of
2000 psig. As a result, C, the allowable number of concentrations can
be quite low in a high pressure steam system. At pressures above 600
psi, silica (Si Q^} percent in the boiler will vaporize along with other
contaminants and escape with the steam. To eliminate this condensation
and resultant fouding of the turbine, it is necessary to maintain
extremely low silica concentrations in the boiler which can result in
a high amount of blowdown. Other methods, such as steam washing can
be used to reduce the contaminant vapor content of the steam permitting .
213
-------
higher boiler water concentrations and reducing the required blowdown
quantity or makeup water quality.
Boiler blowdowns contain all of the additives to boiler feedwater
as well as the soluble matter originally present in the boiler feedwater.
Scale formation is usually inhibited by adding chemicals such as phos-
phates which precipitate scale forming salts to form sludge. Chelating
agents which complex with scale forming metal ions, thus increasing their
solubility, are also widely used. Sodium sulfite or hydrazine are often
added to boilder feedwater in order to inhibit corrosion from dissolved
oxygen.
Boiler blowdown is alkaline with a pH of 9-5 to 10 for hydrazine
treated water and a pH of 10 to 11 for phosphate treated water. Hydra-
zine treated boilers produce blowdowns containing up to 2 ppm ammonia
and those treated with phosphate may contain up to 50 mg/<2 phosphate and
up to 100 mg/^ hydroxide alkalinity.
Water Treatment Wastes - Water treatment waste streams are usually des-
cribed by three parameters: pH, suspended solids concentration, and
concentration parameters typical of processes involved or toxic elements
involved in the process.
Clarification wastes consist of clarifier sludge and filter washes.
Clarifier sludge could be either alum or iron sludge from coagulant
chemicals. If the clarifier is a lime softener, the sludge would con-
taind calcium carbonate and magnesium hydroxide. Filter washes would
contain suspended solids either as light carryover floe from the clari-
fier or naturally occurring in unclarified raw water.
Ion exchanger wastes are either acidic or alkaline except for sodium
chloride solutions which are neutral. Usually, such wastes do not con-
tain suspended matter. They may, however, contain calcium sulfate and
calcium carbonate precipitates because of the common ion effect.
Equipment Cleaning Wastes - A variety of cleaning formulations are used
to clean scale and corrosion deposits from boilers and condensers. The
cleaning program is usually dependent on the composition of the surface
adhering materials. Cleaning solutions are usually grouped in three
principal categories according to their composition. The first category
includes the alkaline cleaning mixtures with an oxidizing agent for copper
removal. These solutions contain an oxidizing agent and copper chelating
compound, usually ammonia. The oxidizing compound converts metallic
-------
copper deposits to divalent copper ion which then reacts with ammonia
to form a soluble complex. The wastewater effluents from such cleaning
contain ammonium ion, oxidizing agents, and high levels of dissolved
copper and iron, andhave high alkalinity.
The second category includes acidic cleaning mixtures. These mix-
tures are effective in removing scale due to water hardness. They con-
tain a strong acid and a fluoride salt to remove silica. Waste streams
from such mixtures may contain phosphates, fluorides, BOD, and acidity
as well as large quantities of iron, copper and hardness forming salts.
The last group of formulations include solutions containing alkaline
chelating agents and anticorrosion additives. These cleaning mixtures
may be used alone or after acid cleaning to neutralize residual acidity
and to remove additional amounts of scale forming materials. Their use
generates wastewater containing alkalinity, BOD, phosphate and scale
forming components.
In addition to these three categories, there are a large number of
proprietary formulations which have been developed and are manufactured
by companies specializing in cleaning chemicals. Most of these chemi-
cals are similar to those described earlier and the resulting wastes con-
tain: alkalinity, BOD, phosphate, ammonium compounds, and scale forming
compounds such as iron, copper, and hardness.
Coal Pile Runoff - Coal pile runoff is the water drainage from the coal
storage area which occurs during periods of rain. Such runoffs present
a potential danger of water pollution if allowed to drain into waterways
or to seep into ground aquifiers. The nature of coal pile runoffs
depends on the type of coal used. Generally, there are two groups of
coal pile runoffs. The first has a neutral or slightly alkaline pH and
contains ferrous ions. Such runoffs are obtained from coal containing
large amounts of alkaline materials or small amounts of pyrite. The
second group of runoffs is highly acidic containing large amounts of
dissolved iron and aluminum. These runoffs are produced from pyrite
rich coal. Pyrites, or iron sulfides, are oxidized by atmospheric oxygen
and hydrolyzed to form ferrous sulfate and sulfuric acid according to
the following reaction:
2Fe'S2 + 01 + H20 + 2F2 SOU + 2H2 SO^ (19)
Additional sulfuric acid may be formed if the ferrous ions are
further oxidized to the ferric state. When rain falls on coal piles,
the acid is washed out and eventually winds up in the coal pile drainage.
215
-------
At the low pH produced, other metals, such as aluminum, copper,
manganese, zinc, etc., are also dissolved to further degrade the water.
Floor and Yard Drains - The floor drains, generally, contain dust and
fines, and floor scrubbing detergents. This stream also contains lub-
ricating oil or other oils which are washed away during equipment
cleaning, oil from leakage of pump seals, and oil collected from spillage
around the storage tank area of oil processing gasifiers.
Water and Wastewater Treatment
Treatment Technology - The water treatment scheme adopted in this study
is designed for maximum water reuse and zero water discharge. The pro-
cess was developed for the Bureau of Mines/Selexol system, but is also
applicable, with only minor modifications, to the other processes. The
principal difference between the various gasification processes, as far
as the wastewater treatment is concerned, is the chemical characteris-
tics of the process condensates generated in the raw gas cleanup systems.
The chemical nature of the remaining streams are expected to be identical
for all processes because each of the stream will originate from a unit
which is common to all processes. Moreover, except for process conden-
sates and water used for slag quenching the characterisitcs of the
various wastewater effluents will be similar to those produced from a
conventional coal-fired power plant.
Table 58 shows the'chemical characteristics of process condensates
produced in the various processes and potential control techniques. Of
the eight integrated systems studied, four utilize a high-temperature
cleanup system and therefore, will not generate process condensates.
The remaining four processes will produce process condensates from the
raw gas cleanup system, but of different water qualities. Condensates
produced by the BCR and Koppers-Totzek gasification systems, because of
their high operating temperatures, are expected to be free of organic
matter. Such condensates will contain suspended solids, ammonia, sul-
fides and possibly small quantities of cyanides. In the partial oxida-
tion/Selexol process naphtha can be used to remove soot from the gas water
wash and, therefore, the condensates from this process can contain small
amounts of organic matter. The most contaminated wastewater effluent
will be generated by the Bureau of Mines/Selexol system. Condensates
from the BCR or the Koppers-Totzek gasification systems can be treated
by air stripping at pH 11 to remove ammonia followed by clarification to
remove suspended matter. The offgas from the stripping tower may require
incineration to prevent air pollution. The clarified effluent can be
used as cooling water makeup. Condensates from both the partial
216
-------
TABLE 58. CHEMICAL CHARACTERISTICS OF PROCESS CONDENSATES
AND POTENTIAL CONTROL SYSTEMS
Processes
Pollutants
BCR/Selexsol ammonia, sulfides, cyanides, suspended solids
BCR/CONOCO
BuMines/Iron Oxide
no condensate
no condensate
BuMines/
Selexol
ammonia, sulfides, cyanides, suspended solids,
penoles, tar, oil, dissolved inorganics
K-T/Selexsol ammonia, sulfides, cyanides, suspended solids
no condensate
K-T/Iron Oxide
Oil /Selexsol ammonia, sulfide, cyanides, naphtha
Oil /CONOCO no condensate
Treatment
air stripping at
pH 11, clarification
biological oxidation
air stripping at
pH 11, clarification
biological oxidation
217
-------
oxidation/Selexol and the BuMines/Selexol gasification systems will
require a treatment step for the removal of organic compounds. Con-
densates from the partial oxidation/Selexol process can be treated in an
oxidation pond, whereas those generated by the BuMines/Selexol process
would require a far more extensive treatment. This is discussed in the
subsequent paragraphs. . . . .
It should be emphasized that the treatment of wastewater effluents
can be accomplished by a great number of processes. However, in order •
to select and implement an efficient waste management program it is
necessary to evaluate the control and treatment techniques against
specific factors applicable to each case. Table 59 is a list of control
techniques for potential pollutants from coal gasification plants. The
information in Table 59 is based, in part, on a work plan for environ-
mental study (^3) of coal conversion processes prepared by Hittman Associates
for the Federal Energy Research and Development Agency. The table con-
tains information relevant to the principles of the methods, their
limitations, the concentration range of their applicability, the effi-
ciency of the methods and the extent of their industrial usage. It
should be emphasized, however, that many of the listed methods have been
developed specifically for the purpose of product recovery and as such
are- not applicable to pollutants present in the wastewater in low concen-
trations.
Process Description
Figure U5 shows a simplified flow diagram of a water treatment pro-
cess for the Bureau of Mines/Selexol combined-cycle power plant. The
chemical composition of the raw water(^)is shown in Table 60. The analy-
sis represents the upper limits of the concentration rangeof the constit-
uents in 95 percent of the fresh surface water in the United States.
The water balance for the process is shown in Table 6l.
Raw water is initially pumped to a storage reservior which also
serves as a flow equalizer and a clarifier for removal of naturally
occurring suspended solids. The water withdrawn from the storage
pond is split into two streams, one of which is conveyed to the cooling
system as makeup after the water has been chemically conditioned to con-
trol scaling, corrosion, and fouling. The second stream is demineral-
ized by ion exchangers and deaerated to remove dissolved oxygen and car-
bon dioxide. The demineralized water is conveyed to the various boilers
and waste heat recovery systems for steam generation.
218
-------
Table 59. POTENTIAL CONTROL TECHNOLOGY FOR COAL CONVERSION WASTEWATER
Pollutant
Hexavalent (1}
chromium
(2)
Treatment Method
Reduction to chromium
(III ) with S02,NaHS03
or FeSC>4 at pH below
3 followed by preci-
pitation at pH 8.5-9.5
Adsorption on anion
exchanger
Limi tations
Reduction is not
complete. Rate
depends on pH,
reducing agent
and contact time-
Recovery process
Appl i cab! e
Concentration
Range
100-500 mg/1 .
< 200 mg/1
Level After
Removal
0.05-1 mg/1
Removal to
0.05 mg/1
Industry
Usage
Common
Moderate
(3) Evaporative recovery Recovery process >500 mg/1
MD
Cyanide (1) Oxidation to cyanate
with chlorine at pH
above 10
(2) Oxidation to cyanate
with chlorine at pH
above 10 followed by
acid hydrolysis to
C02 and N2 at pH 2-3
(3) Decomposition to C02
and N2 via cyanate
with chlorine at pH
8-8.5
(4) Electrolytic decom-
position to C02
N2 via cyanate at
200°F
100-1000 mg/1
Increases total
dissolved solid
and treatment
costs
100-1000 mg/1
Toxic cyanogen
chloride may be
liberated, a
large excess of
chlorine is
required
Interference by >1000 mg/1
sulfate
Removal to
0.1 mg/1
Complete
removal
Not
Practiced
Not
Practiced
Moderate
Common
0.1-0.4 mg/1 Common
after 7-18
days
(5) Ozonation
Only partial
decomposition
to C02 and N2
100-1000 mg/1
Complete
removal
Moderate
-------
Table 59.
(Continued)
ro
ro
o
Pollutant
(6)
(7)
(8)
(9)
(10)
Fluoride (1)
(2)
(3)
Treatment Method
Storage of waste
at ambient tempera-
ture
Precipitation as
ferro f erri cyan 1de
with iron salt
Adsorption on
activated carbon
Biological treatment
Oxidation with hydro-
gen peroxide to
cyanate (Kastone
process )
Precipitation with
lime as cal ci urn
fluoride at pH 11
Coagulation by alum
Adsorption on
Limitations
Incompl etc
treatment
Incomplete
treatment
Incomplete
treatment
Propri etary
information
Slow rate of
pred pitation
Applicaole only
to low hardness
water
Presence of
Appl icable
Concentration
Range
100-1000 mg/1
100-1000 mg/1
.
100-1000 mg/1
> 100 mg/1
100-1000 mg/1
720 mg/1
<20 mg/1
<20 mg/1
-Level After
Removal
70-90% re-
duction after
4-8 days
storage
0.5-12.3 mg/1
depending on
the concen-
tration 1n
influent
• '0.6-1.4
70-90%
removal
10-20 mg/1
Removal to
1 mg/1
0.5-1.5 mg/1
Industry
Usage
Practiced
by coking
Industry
Not
practiced
Not
practiced
..
Common
Not
practiced
Water
hydroxylapatite bed
chlori ne i n-
creases cost of
bed regeneration
treatment
-------
Table .39. -(Continued) .
Applicable
Pollutant
(4)
(5)
Treatment Method
Adsorption on
aluminum saturated
cation exchanger
Adsorption on
activated
alumina bed
Limitations
Expensive
4% of bed is
lost in each re-
generation cycle
Concentration
Range
^20 mg/1
<20 mg/1
Level After
Removal
Removal to
1 mg/1
Industry
Usage
Not
practiced
Not prac-
ticed-water
treatment
technology
ro
ro
Iron (11) (1)
(2)
(3)
Tar
Oil
Oxidation to Fe(lll)
by aeration followed
by precipitation as
Fe(OH,3 at pH 7
Oxidation to Fe(lll)
by chlorine followed
by precipitation as
Fe(OH)3 at pH 7
Deep well disposal
and (1) Gravity separation
Does not
emulsion
remove
Concentrated
waste
Primary
treatment
Below 0.5
mg/1
Removal to
0.5 mg/1
60-99% of
floated oil
Common
Moderate
Practiced
by steel
industry
Common
(2) Centrlfugation
(3) Heating
Secondary
treatment
Secondary
treatment
Common
Not
practiced
-------
P 59... (Continued)
ro
ro
ro
Pollutants
pH Control
Phenols
(4)
(5)
(6)
(1)
(1)
(2)
(3)
(4)
(5)
(6)
(7)
Treatment Method
Precoat filtration
Coagulation or de-
musi f icati on with
chemicals, followed
by air flotation or
settl 1 ng
Biological treatment
Neutral ization with
chemi cal s
Benzene-caustic
dephenol Ization
process
Counter-current
extractor
(Chemizon process)
Pulsed column
extractors
Phenosol van
dephenol ization
(Lurgi)
IFAWOL dephenoliza-
tion (Carl still)
Light oil extrac-
tion (Koppers)
Inci neration
Appl icable
Concentration
Limitations Range
Secondary
treatment
Addition of alum- Secondary
forms sludge treatment
which are diffi-
cult to dewater
Secondary
treatment
Cost depend on
buffer capacity
of waste
> 500 mg/1
> 500 mg/1
> 500 mg/1
> 500 mg/1
> 500 mg/1
> 1500 mg/1
> 7000 mg/1
Level After
Removal
5-20 mg/1 •
50-90*
Removal to
15 mg/1
Neutral pH
210-240
mg/1-.
Removal to
100 mg/1
Removal to'.
30 mg/1
4.5-10 mg/1
Removal to
40 mg/1 -
10-30 mg/1
Compl ete
Industry
Usage
Common
Common
Common
Common
Common
Common
Common
Common
Common
Common
Not
practiced
-------
Table 59. (Continued)
Applicable
ro
ro
(JO
Pollutants Treatment Method
(8) Oxidation ditch
(9) Trickling filter
(10) Activated sludge
(11) Oxidation with
ozon
(12) Activated carbon
bed
(13) Oxidation with
Limitations
Expensive when
waste contains
more than
5 mg/1
Concentration
Range
50-500 mg/1
50-500 mg/1
50-500 mg/1
<50 mg/1
<50 mg/1
<50 mg/1
Level After
Removal
99%
98%
99%
Removal to
0.35 mg/1
Removal to
0.005 mg/1
Industry
Usage
Common
Common
Common
Limited
usage
Common
Common
chlorine
Dissolved (1) Concentration and
Solids evaporation
(2) Reverse osmosis
(3) D1st1llat1«n
Efficiency
depends on
membrane
condi tlon'
> 50000 mg/1 Complete
removal
50-95%
60-90%
Not generally
1n use-de-
salination
technology
Not prac-
tlced-de-
salination
technology
Not prac-
ticed-de-
sal1 nation
techno!ogy
-------
Table 59- .(Continued)
ro
ro
-p-
Pollutants
Suspended
Solids
Ammonia
Chloride
Sulflde
(1)
(2)
(3)
(1)
(2)
(3)
(1)
(2)
(1)
Thlocyanate(l)
(2)
Treatment Method
Sedimentation
Chemical coagulation
Filtration
Stripping at pH
of 10-11
Biological
nitrification
Ion exchange
Deep well injection
Evaporation ponds
Biological oxidation
to sulfate
Biological oxidation
Ion exchange
Applicable
Concentration
Limitations Range
Water adsorbs
C02-may lead to
scale formation
Nutrient may be <1250 mg/1
requi red
>60 g/1
Limited by geo-
graphical location
and land avail-
ability
Excess ammonia
lower efficiency
Excess ammonia
lower efficiency
Level After
Removal
90-95%
95-99%
95%
50-90%
Removal to
2 mg/1
80-95%
Ultimate
disposal
Complete
removal
Complete
oxi dation
90%
90%
Industry
Usage
Extensive
Moderate
Not Prac-
ticed-waten
treatment
technology
Extensive
Extensive
Not
practl ced
Moderate
Extensive
Moderate
Not
practiced
-------
WATER AND WASTE WATER TREATMENT FOR THE BUREAU OF MINES/SELEXOL PROCESS
RAW WATER
ro
ro
VJ1
CORROSION INHIBITOR
COOLING
WATER
DEFLOCCULANT
CHLORINE
12
MAKE-UP CIRCULATING 14
PHOSPHATE
HYDRAZINE
AMMONIA
COOLING WATER
DRIFT
t.. 1
EVAPORATION
16
COOLING
SYSTEM
SODA ASH
T17
COOLING SYSTEM
BLOWDOWN
LIME
BOILER FEED WATER
O2+CO2
CATION
EXCHANGER-
1
IDEAERATOR
ANION-^
DEMINERALIZER
BLOWDOWN
CLARIFIER
BLOWN DOWN
EXCHANGER
o
W
SLUDGE FOR
DISPOSAL
FLOOR
AND
RUN-OFF
CLEANING WASTE,
I
L
SLUDGE FOR
DISPOSAL
EFFLUENT TO COAL
PREPARATION WASHER
_J
P
.U
cn
-------
Table 60. RAW WATER ANALYSIS
pH 7.6
Total dissolved solids 400 mg/1
Bicarbonate (HCO^) 180 mg/1
Sulfate (SoJ~) 90 mg/1
Chloride (c 1") 170 mg/1
Nitrate (N0~) . 4.2 mg/1
Calcium (Ca) 52 mg/1
Magnesium (Mg) 14 mg/1
Sodium and Potassium (Na, K) 85 mg/1
Iron (Fe) 0.7 mg/1
Silica (Si02) 8.8 mg/1
Dissolved Oxygen (02) 9.8 mg/1
Ammonia (NH3) 2.5 mg/1
Specific conductivity at 25°C 1.1x10 mho
226
-------
Table 61
WATER BALANCES FOR THE BUREAU OF MINES/SELEXOL PROCESS
Stream No. Description Flow Ib/hr
1. Raw water input 2,635,784
2. Raw water to demineral ization unit 611.,435
3. Acid, regenerant and rinse water 14,604
4. Alkaline regenerant and rinse water 14,604
5. Boiler feed water to process 582,227
6. Demineralizer blowdown 29,208
7. Ash from gasifier 114,132
8. Water for ash wetting 52,704
9. Wet ash for disposal 166,836
10. Process condensate 174,686
11. Boilers blowdown 40,906
12. Cooling water makeup 2,910,280
13. Cold cooling water 108,546,280
14. Warm cooling water 108,546,280
15. Drift losses 108,546
16. Evaporation losses 2,099,165
17. Cooling system blowdown 702,569
18. Softened Water 679,073
19. Clarifier blowdown 23,496
20. Biological treatment unit blowdown 8,734
22?
-------
The Slowdown from the demineralizer consists of the waste
regenerants and rinses from both the cation and anion resins. These
streams are combined and conveyed to a neutralization unit.where the pH
is adjusted to within the range of 6.0 to 9«0> on a batch basis, by the
addition of sulfuric acid or sodium hydroxide as required. The neutral-
ized waste will be used to wet the ash from the gasifier prior to its
disposal.
The blowdowns from the boilers and the waste-heat recovery systems
are high-quality waters and, therefore, can be used as a supplement for
almost every water input to the plant. These waters are combined and
conveyed to the cooling system as part of the cooling water makeup.
The cooling water blowdown is of the same chemical quality as the
water circulating in the condenser cooling system. Limits on the water
quality in that system are governed by the need to remain below concen-
trations at which scale forms in the condenser. The blowdown is lime
softened and recycled back to the cooling system after clarification.
The clarifier blowdown will also be disposed of with the coal ash.
The water condensate from the gas cleanup system is. highly polluted
containing a variety of organic and inorganic compounds. Some of the
pollutants, such as phenols and cyanides, are highly toxic to living
organisms. However, indications are that the concentration of the toxic
compounds is below the tolerance limits of the micro-organism population
used in biological based treatment processes. The water condensing from
the Selexol cleanup system is treated in a two-stage biological treatment
unit. The first stage is an activated sludge system whereby zoogleal
bacteria and other aerobic organisms are mixed with the wastewater and
aerated. The activated sludge is subsequently separated from the treated
waste by sedimentation and the treated effluent is conveyed to a polishing
aeration basin, where residual organic matter is further biodegraded.
The sludge from both stages is collected and a portion is returned to
the aeration basin as required to maintain biological activity.
The rest is sent to the ash disposal area for drying and disposal.
The purified water from the polishing-settling basin is filtered and
sent to the cooling system. It should be indicated that the reuse of
process condensate as cooling water may cause odor problems since the
chlorination of the water effluent may result in the formation of
highly odorous chlorophenolic compounds. The biological treatment unit
can also be used to treat domestic wastewater and any other waste stream
containing biodegradable organic matter. It may be necessary to add
nutrient elements to the wastewater influent if such deficiency occurs
228
-------
in order to maintain the efficiency of the process and to prevent process
upset due to the high load of toxic substances.
The treatment of periodic wastes which are not connected with coal
gasification or electrical power generation can be accomplished in a sin-
gle treatment system. The wastewater effluents from equipment cleaning
operations, coal pile, floor and yard drainage are collected in a storage
pond. Floor and yard runoffs are usually passed through an API oil
separator located ahead of the storage pond to remove nonemulsive oil.
The disposal of the combined waste can be accomplished by evaporation if
the land is inexpensive and the rate of evaporation is higher than the
rate of precipitation. Alternatively, the wastewater effluent is
neutralized with lime to a pH of 6 to 9 an^- the water is clarified to
remove precipitated salts and suspended solids. It may be necessary to
add coagulant in order to remove emulsive oil and colloidal suspensions.
The water effluent can be used in the washing operation of the coal pre-
paration section of a nearby mine. The solid effluent from the clari-
fier is disposed of with other solid wastes.
Chemical Treatment of Circulating Cooling Water
Cooling waters are treated to inhibit scale formation, corrosion,
and fouling. Scale is an adherent layer of foreign material formed on
the water side of the heat exchanger surface. The scale acts as an
insulator reducing the rate of heat transfer, and consequently, the
thermal efficiency of the process. In addition, scale formation
restricts the rates of water flow in the condensers by increasing the
hydraulic friction of the tubing.
Scale is formed as a result of precipitation of inorganic salts
which occur in all natural water. Because of the continuous loss of
cooling water due to evaporation and the addition of makeup to supple-
ment these losses, the concentration of dissolved solids in the cir-
culating cooling water gradually increases. If the solubility of any
combination of cations and anions exceeds their solubility product, the
salt will precipitate. The formation of scale can be controlled by
increasing the solubility of the salts or by precipitating them as
sluge and removing the sludge with the blowdown. The solubility of the
salt can be increased by lowering the active concentrations of the ions.
For example, the concentration of the carbonate ion can be reduced by
lowering the pH of the water. Alternatively, the addition of chelating
agents which form complexes with calcium reduces the concentration of
the free calcium ion. By lowering the concentration of the ions the
cooling system can be operated at a higher concentration cycle and
consequently, with smaller blowdown.
229
-------
Precipitation of scale forming salts as sludge is achieved by adding
dispersants which prevent the agglomeration of solid material. Recent
developments have centered around the use of polyelectrolytes which
adsorb onto the surface of the growing salt crystals and enter the
crystalline structure. This prevents deposition of a uniform adherent
scale causing layer. Instead, irregularly shaped crystals are formed
which are easily sheared or broken off from the surface scale.
Chemical treatment is used most often for corrosion control. Chro-
mates and polyphosphates are used either separately or together to
inhibit corrosion in cooling water recirculating systems. Chromate,
being a strong oxidizing agent, forms a thin passive layer of oxides on
the anodic surface which protects the metal against further oxidation.
When used alone, chromates require concentrations above 700 ppm as
NagCrO. , otherwise corrosion may be even more severe than if no inhibi-
tor had been used. If, on the other hand, chromates are used in combin-
ation with polyphosphate, the concentration of sodium chromate required
is approximately 20 ppm. However, the pH of the cooling water must be
carefully controlled to prevent precipitation of calcium phosphate or
calcium carbonate. Other anodic inhibitors are silicates, ferrocyandies,
and nitrites. Cathodic inhibitors include zinc, nickel, manganese,
and trivalent chromium salts. These inhibitors are also used in combin-
ation with chromates and polyphosphates.
Fouling refers to the deposition of foreign matter on process
equipment surfaces. Fouling can result from deposition of inorganic
matter such as silt and clay or can be caused by algal and bacterial
growth. The latter is by far the most serious source of fouling.
Fouling is most commonly controlled by chlorination, usually in combina-
tion with nonoxidizing biocides such as thiocyanates, copper salts, or
chlorinated phenolic compounds. The deposition of inorganic suspended
solids is controlled by dispersants such as organic polymers which pre-
vent agglomeration and subsequent settling.
Boiler Feedwater Treatment
The treatment of water for the purpose of making boiler feedwater
can be viewed as a two-step process. The first is the external treatment
whereby the raw water is demineralized by ion exchangers, reverse osmosis,
or softening to lower its dissolved solids content. The second step is
the internal treatment involving the addition of various chemicals to
the water to inhibit scale formation and corrosion. Scaling is controlled
by the same methods used for cooling water, that is, either by precipi-
tating the scale forming cations as a sludge, usually as salts of
230
-------
phosphate, and removing the sludge with the blowdown, or by chelating
these ions with complexing agents such as EDTA to increase their
solubility. Corrosion is controlled by chemical deaerators which are
essentially strong reducing agents. These compounds react with dissolved
oxygen to form inert and noncorrosive products. Sodium sulfite is
usually used in low-pressure boilers but not in high-pressure boilers
because it is oxidized to sulfate, an undesirable component in high-pres-
sure boilers. Oxygen corrosion in high-pressure boilers is controlled
with hydrazine which decomposes upon reacting with dissolved oxygen to
water and inert nitrogen.
Cost Estimates For The Water System of the BuMines/Selexol Process
This section discusses cost estimates for the production of process
and cooling water, and for the treatment and reuse of wastewater for the
BuMines/Selexol integrated combined cycle power plant. The cost esti-
mates were determined from published reports dealing with water and
wastewater treatment technologies, and with the capital and operating
costs of such technologies. This is only a preliminary study which
lacks detailed designs of the various processing units involved in the
treatment of water and wastewater, and therefore, the costs should be
considered only as first estimates. The calculated values were deter-
mined from cost estimates of units processing similar flow rates and
chemical compositions of water and wastewater. Very often, however,
such information was not available and the cost estimates were deter-
mined from published data which have been extrapolated to adjust for
differences in both flow rates and chemical compositions. The cost
estimates, so determined, were then revised to adjust for escalation
using the Chemical Engineering Index. The results are listed in Table
62 as the estimated capital costs and annual operating and maintenance
costs for mid-1975.
In calculating the cost estimates, it was assumed that the lime
softening system includes a clarifier, a rapid sand filter, and a slugde
removal system. The demineralizer was assumed to include separate
cation and anion exchangers each consisting of four columns, one of
which is being regenerated at all times. The raw water storage system
includes a pumping station at the raw water source and a storage pond
with a 30-day storage capacity. The latter was assumed to be paved with
reinforced concrete to prevent losses due to water infiltration. In
calculating the cost estimate for the cooling water treatment system,
it was assumed that the principal cost is that associated with the
chlorination of the water effluent from the biological treatment system:
This effluent was assumed to have a 5 ppm phenol content and the chlorine
231
-------
TABLE 62
CAPITAL AND ANNUAL OPERATING COSTS OF A WATER SYSTEM FOR THE BuMINES/SELEXOL PROCESS
Unit
Lime softening system
Demi nera 1i zer
Raw water storage system
Cooling water treatment system
Neutralization system
Activated sludge system
Total
Capital Costs
$ 713,000
$1 ,024,000
$3,312,000
$ 50,000
$ 24,000
$ 500,000
$5,623,000
Annual
Operating Costs
$
$
$
$
$
$
152,000
429,000
5,000
392,000
3,000
46,000
$1 ,027,000
ro
-------
requirement was based on total destruction of the phenol to prevent odor
problems. The neutralization unit was assumed to be completely automated,
neutralizing water effluent of pH k with lime. The neutralized slurry
is conveyed to the ash disposal system for ash wetting. The activated
sludge system consists of an aeration unit, clarifier, and an aerated
polishing pond with a 2h-hour capacity. The latter is also paved with
reinforced concrete to prevent water infiltration.
233
-------
SOLID RESIDUALS
Summary of Solids Produced
Ash - Only first-generation (low-temperature) gasifiers such as the
BuMines gasifier produce ash, because they operate below ash fusion
temperatures ( 1500°F). The ash is a refractory material present in the
coal and has no fuel value. Ranges of typical ash composition from
boilers are given in Table 63.
Sl_ag - Slag is produced by second-generation (high-temperature) gasifiers,
because they operate at slagging temperature ( 2200°F) and cause the
mineral matter to melt. The slag is quenched and removed from the gasi-
fier and may contain some dissolved gases. A typical gasifier slag
composition however, would probably be quite similar to a typical gasi-
fier ash composition.
Spent Limestone (Scrubber Sludge) - If a non-regenerable flue gas scrubber
is used with any of the systems to reduce S02 emissions, a scrubber sludge
is produced. If a lime/limestone scrubber is used, the scrubber sludge
consists of CaSO~ and CaSO^,. This is dewatered and stabilized by conver-
sion to CaSOlj. before disposal.
Spent Dolomite - Spent dolomite is produced in the CONOCO desulfurization
processes. The spent dolomite is removed from the regenerator and con-
sists of CaCOo-MgO and CaS-MgO. This is further treated with H^O and C02
to convert it completely to inert CaCOo-MgC03 which is then disposed of.
Elemental Sulfur - In all the systems utilizing Glaus sulfur recovery ..
with a regenerable scrubbing process, elemental sulfur is produced. The
sulfur is typically 99 percent pure, and, when in solid form, is usually
flaked for bulk shipment to market.
Miscellaneous - Particulates recovered from the fuel gas by particulate
removal devices consist of fly ash, dust and unburned carbon. Processing
steps used to concentrate SOg in off gases produce a solid waste consist-
ing of Na2SOo, which must be oxidized to NapSOij. before disposal. Other
solids produced in smaller quantities, include the spent catalysts which
use alumina supports on which the catalyst is impregnated. Spent iron-
oxide from the Bureau of Mines and B&W iron-oxide processes has to be
regenerated and stabilized before disposal.
23k
-------
Table 63
CONSTITUENTS OF COAL ASH1
Constituent Percent
Si02 30-50
A1203 20-30
Fe203 10-30
Ti02 0.4-1.3
CaO 1.5-4.7
MgO 0.5-1.1
Na20 0.4-1.5
K20 1.0-3.0
S03 0.2-3.2
C and volatiles 0.1-4.0
P • 0.1-0.3
B 0.1-0.6
U and Th 0.0-0.1
Cu trace
Mn trace
Ni . trace
Pb trace
Zn trace
Sr trace
Ba trace
Zr trace
Composition is representative of a fully oxidizing conventional
boiler. Residue from gasification would tend to be in a less
oxidized and possibly sulfided form.
235
-------
Identification of Types of Solids Produced by Water and Wastewater
Treatments
Solid wastes are a major by-product of many of the processing steps
involved in the treatment of water and wastewater. All water supplies
contain varying amounts of naturally occurring suspended solid matter
and dissolved chemical salts. Therefore, raw water must be treated for
removal of the mineral salts and suspended 'matter before being used in
the process. Very often, such treatment produces solid wastes as by-
products. The treatment of wastewater also results in the formation of
solid wastes.
Solid wastes from water and wastewater treatment processes are
usually referred to as sludge. They are formed by the precipitation of
slightly soluble.salts, or due to coagulation and subsequent sedimentation
of suspended matter. They are collected and removed from the water treat-
ment system in clarifiers as slurries with various contents of suspended
solids.
The nature and composition of the sludges depend on the character-
istics of the treated water and the type of treatment used-. Clarifier
sludge from water treatment processes could be either alum or iron salt
sludge, from coagulant chemicals. Alum sludge is a bulky gelatinous sus-
stance composed of aluminum hydroxide, inorganic particles such as clay or
sand, color colloids, micro-organisms including plankton, and other
organic matter removed from the water. The major constituent in sludge
from a lime soda softening clarifier is calcium carbonate. Other consti-
tuents which may be present are .magnesium hydroxide, hydroxides of
aluminum or iron, insoluble matter such as clay, silt or sand, and
organic matter such as algae or other plankton removed from the water. •
The nature and characteristics of the excess sludge from the biological
treatment system will depend to a large extent on the chemical composi-
tion of the waste and the species of micro-organisms that can climatise
themselves to this type of waste. Generally, such waste will contain
dead cells of bacteria and algae, partially decomposed organic matter
and inert soluble and insoluble inorganics. Excess activated sludge is
usually golden brown and flocculent. Uncontrolled disposal of such
sludge may result in the development of septic conditions due to the
decomposition of the residual organic matter by anaerobic organisms.
Sludge is also formed when coal pile runoff and floor and yard drainage
are neutralized. The constituents of such sludge are calcium carbonate
and hydroxides of iron, aluminum, chromium, zinc and manganese. Oil may
be present as well as fines of coal and .dust. Sludge may also be formed
236
-------
as a result of the neutralization of waste regenerants and rinses from
the demineralization system. Since sulfuric acid is used to regenerate
the cation exchanger, calcium sulfate may precipitate due to the common
ion effect.
The lime sludge from the softening unit will be mixed with the ash
to wet it prior to its transport to the disposal site. The sludges from
the activated sludge and demineralizer units would probably be disposed
of with conventional disposal techniques. The activated sludge would" be
dried, usually by vacuum filtration, and then subsequently incinerated.
The sludge from the demineralizer, primarily calcium sulfate, can be de-
watered and disposed of in land fill.
Disposal Options and Their Implications
As with any process or power plant that produces solid waste, there
are several options available for disposing of these residuals, each
depending on a number of site-specific and residual-specific, considera-
tions. These options range from the sale of the residual to the storage
of the residual in an environmentally acceptable manner. The storage may
take the form of on-site and off-site burial or surface storage. The
following deals with the power plant solid residuals, the disposal
options that would likely be implemented and their implications, vis-
a-vis cost and environmental considerations.
Ash and Slag - Activities in the area of ash and slag utilization probably
predate those of any of the other solid residuals one might expect from
low-Btu gasifiers. As a result, the area of by-product utilization is
probably more advanced for ash and slag than that of any of the other
solid residuals. The impetus for this developmental work was the result
of the dominance of coal as an energy source from the beginning of the
Industrial Revolution through the 19^0's. An annotated list of refer-
ences on ash utilization published in the Proceedings of the Second Ash
Utilization Symposium sponsored by the National Coal Association includes
an entry for a U.S. patent covering the production of alumina from coal
ash which was obtained as far back as 1932.
The major use of fly ash is as a concrete additive which serves both
as a mechanical filler supplementing or replacing fine aggregate and as
a pozzolan supplementing or partially replacing cement. There are numer-
ous other, quantitatively less significant, existing uses for fly ash,
bottom ash, and slag. They include use in abrasive cleaning, refrac-
tories, oil well cementing, grouting, snow sanding, mine fire control,
subsidence control, pipe coatings, sand blast grit, etc. Additionally,
237
-------
the Bureau of Mines (now EKDA) has sponsored several studies which
evaluate the use of fly ash in mined-land reclamation as well as ash
utilization from lignite gasification. Most significantly, the cost
per acre for mined-land reclamation can be reduced by a factor of 3 to
5 compared with the conventional methods} 2' It should be noted, how-
ever, that the economics of producing a commercial product from ash or
slag must be examined very carefully in comparison with the storage or
fill alternative, since the necessary equipment and operating staff
required to convert the ash to a commercial product can significantly
increase the cost of ash disposal.
In 1971) 12 percent of the fly ash, 16 percent of the bottom ash,
and 75 percent of the slag produced in the United States were utilized
in the applications previously mentioned. ^°) As can be seen from these
figures, especially those for fly and bottom ash, the major portion of the
ash produced in the United States ultimately ends up in disposal areas. These
disposal options generally consist of ponding and landfilling.
With ponding, ash is transported in the form of a slurry either to
an off-site or on-site ash settling pond, where the ash settles out of
the slurry and the water is removed via weirs or standpipes, thus allow-'
ing continuous operation of the pond unti it is full of ash. The cost
for this method of disposal is in the range of $0.56 to $2.0^ per ton of
ash.( ' This figure includes the operating costs which incorporate
transport and other pond operation expenses. A nationwide survey^ ')
conducted in 1970 of 22 utilities disposing of ash in off-site ponds
showed that it was costing $0.03^- to $1.23 per ton (average of 25 plants
equal to $0.51 per ton) to sluice ash to the disposal area. Trucking
the ash off-site was being done at a higher cost, in the range, of $0.12
to $1.^9 per ton, or at an average of $0.57 per ton for the 10 plants
that utilized truck hauling.
The cost of constructing the pond would depend on a number of site
and design related factors, e.g., the nature of the soil at the pond
location, the size of the pond and the type of liner used. The cost for
a five to ten acre pond with no providions for drainage can range from
$5,000 to $20,000 per acre for a pond with a clay or stabilized poz-
zolan base lining.(^8) On the other end of the scale, the cost of a
drained pond with a plastic liner can range from $25,000 to $30,000
per acre. ^°'
Generally speaking, the viability of this alternative is a func-
tion of the availability of suitable sites at or in close proximity to
the plant. .
238
-------
In the case of the Bureau of Mines Gasifier and Selexol Unit inte-
grated power plant, production (392 x 103 tons/year based on a 0.70 load
factor) of bottom and fly ash would require about 3&0 acres of storage
for ponding the ash associated with 20 years of production, if the ash is
ponded to a depth of 10 feet and compacted to a density of 100
Some plants do not have adequate space for a disposal pond and must
resort to transporting the ash to a land disposal site. In some instances
this site might be within. the plant boundaries, but this is usually
not the case. Care must be taken in the selection of these sites, since
there appears to be some potential for the leaching of contaminants from
the ash causing problems with groundwater. Although there are no veri-
fied instances of groundwater pollution due to leaching of contaminants
in fly ash used in landfills, greenhouse studies have shown that the appli-
cation of fly ash to soils does increase the availability of boron, molyb-
denum, potassium, zinc and phosphorous. (^®> ^9) Also, the constituents of
ash from a gasifier may behave diffferently.
As is the case with ponding, the economics of a landfill can vary
widely depending on the distance to the disposal site, the amount of ash
to be disposed of, the type of transportation used and the landfilling
technique used. The reported range in cost for operating a landfill is
from $0.56 to $2.2U per ton, not including the cost of reclamation. (^°)
Significantly, studies sponsored by the Bureau of Mines (ERDA) have shown a
potential benefit to plant growth through the controlled addition of fly
ash to agricultural soils. (50) Data developed at the Morgantown Energy
Research Center show that approximately 200 tons of fly ash can, on an
average, reclaim an acre of surface mined land. Landfilling spent sur-
faces coal mines might provide a utility, assuming it is conveniently
situated, with a convenient and environmentally acceptable disposal alter-
native. Using the figure developed at the Morgantown Research Center,
it should be noted that the fly ash associated with one year of operation
of a Koppers-Totzek gasifier, of the size addressed in this study, could
have reclaimed approximately three percent of the land disturbed by sur-
face mining for coal in the Central States in 1970. 0^5)
Sulfur - With the advent of environmental regulations limiting the amount
of sulfur that can be discharged to the environment, in particular the
discharge of sulfur oxides resulting from the combustion of fossil fuels
containing sulfur, quantities of sulfur are becoming available that will
far outstrip the demand. As a result, a number of government and private
research groups are currently exploring new applications for elemental
239
-------
sulfur. These applications generally fall into three major groups:
(l) sulfur containing fertilizers, (2) sulfur based construction and
paving materials, and (3) sulfur foams.
Although about 50 percent of the sulfur presently consumed goes into
fertilizer production, the Sulfur Institute foresees an annual added poten-
tial of 2.8 million tons per year of sulfur as a crop nutrient in the
United States and Canada. In addition to its use in fertilizer as a crop
nutrient, sulfur is currently being studied as a coating to urea for its
application as a slow-release fertilizer. Since it is estimated that the
sulfur coated urea can be produced for only about 35 percent more than
the cost of regular urea, it would be much cheaper than other controlled
release products now on the market.(51)
With regard to the utilization of sulfur as a road paving material,
it has been used as a substitute for limestone as the bulk aggregate.
Shell Canada Limited, has been experimenting with the addition of molten
sulfur to hot-mix asphalt paving materials. They claim that this addi-
tion increases the mix workability so that the mix may be placed without
densification. As the mix cools, the sulfur solidifies and imparts a
high degree of mechanical stability to the mix so that high quality mixes
may be produced from poorly graded aggregates and even one-sized sands.
By incorporating sulfur in asphalt mixes, high quality paving materials
can be manufactured using inexpensive, poorly-graded sands. These sand-
asphalt sulfur mixes may be used to construct road bases, surfaces,
curbing, and sidewalks to build platforms over weak subgrades and for
castings of various shapes.
A spray material containing sulfur, talc, fiberglass, and dicyclo-
pentadeiene has been used by the Bureau of Mines to construct block build-
ings as part of a demonstration program. In the demonstration, the blocks
were surface bonded together for structural stability by spraying with
the mixture. This demonstration program was an attempt to show the feasi-
bility of sulfur in coatings as well as in structural materials.
Sulfur foam, as a subsurface insulation, could potentially have an
even higher volume highway application than its use as an aggregate.
Sulfur foam may some day be widely used as roadway or runway subsurface
insulation, either to protect the road subbase from freezing or to pro-
tect a permafrost subbase from thawing. In either case, the foam would
be buried approximately one foot below the surface, deep enough not to
be affected by the daily temperature cycling on a surface. Sulfur foam
might also be used as subbase insulation for homes or cold storage
warehouses.
-------
Miscellaneous Solids - The sodium sulfite resulting from the purge of the
Bureau of Mines Iron Oxide cleanup system is a compound that has already
found some commercial application. Since it s a compound that is easily
oxidized, it can and is being used where a gentle reducing agent is
desired. These applications include its use as a bleach for wool and
silk; as an antichlor after the bleaching of yarns, textiles and papers;
as a preservative for food stuffs; and to prevent raw-sugar solution from
coloring upon evaporation. This material also has wide application in
the preparation of photographic developers, as a preventative of the oxi-
dation of hydroquinone and other agents. To a smaller degree, it has
found acceptance in the field of medicine as an antiseptic and as an
antizymotic for internal use. Recent interest in sodium sulfite has
centered around the discovery that its addition to boiler feedwater will
remove oxygen from the water, and thus help prevent corrosion and scale
formation. In general, the surplus of this material not utilized in the
previously mentioned applications would probably be sold to the sulfate
pulp mills.
The spent dolomite that is produced by the Conoco half-calcinated
dolomite cleanup system might present a disposal problem. Some investi-
gators^ / believe that the spent material will consist of a calcium
carbonate which would not be a particularly troublesome material to dis-
pose of by storage or landfilling. However, the calcium sulfide inner
core component liberates hydrogen sulfide gas very slowly on exposure to
moist air, creating an odor problem and ultimately yielding sulfur
dioxide, sulfite, and sulfate pollutants. Calcium sulfite is, however,
utilized in industry as a depilatory in the tanning industry and in cos-
metics. In a finely divided form, it is employed in luminous paints.
-------
REFERENCES
1. Robson, F. L., A. J. Giramonti, W. A. Blecher and G. Mazzella.
Fuel Gas Environmental Impact. EPA 600/2-75-078, November 1975.
2. Fox, G. R. and J. C. Gorman. A Study of Advanced Energy Conversion
Techniques for Utility Applications Using Coal or Coal-Derived
Fuels - Task 1 Results. ASME Presentation, 1975 Winter Annual
Meeting, December 3, 1975-
3. Anon. Energy Conversion Alternatives Study (EGAS) - Westinghouse
Task I Report Summary. ASME Presentation, 1975 Winter Annual
Meeting, December 3, 1975-
h. Economics of Air vs. ©2 'Pressure Gasification of Coal. Fluor
Engineers and Constructors, Inc. EPRI 239-1? Phase Report, January
1975.
5. Robson, F. L., et al. The Technological and Economic Feasibility
of Advanced Power Cycles and Methods of Producing Non-Polluting
Fuels for Utility Power Systems. NTIS PB-192-398, December 1970.
6. Magee, E. M., et al. Evaluation of Pollution Control in Fossil
Fuel Conversion Processes; Gasification; Section 1: Koppers-
Totzek Process. ESSO Research and Engineering Company, January 197^-,
EPA-650/2-7l|-009a.
7. Farnsworth, J. F., D. M. Mitsak, and J. F. Kamody. Clean Environment
with K-T Process. EPA Meeting, Environmental Aspects of Fuel Conver-
sion Technology, St. Louis, Missouri, May 197^-•
8. Letter Communication January 13, 1975- G. J. Horvat, AIRCO
Industrial Gases, to M. S. Dandavati, Hittman Associates, Inc.
-------
REFERENCES (Cont'd)
9. Proposal for 2000 T/D Oxygen Gas Plant. AIRCO Cryoplants Proposal
No. 73^8.
10. Strelzoff, S.. Partial Oxidation for Syngas and Fuel, Hydrocarbon
Processing. December 197^-•
11. Milner, M. J., and D. M. Jones. 8th World Petroleum Congress,
Moscow, June 1971.
12. Letter Communication November 15, 197^. Dennis Duncan, Institute
of Gas Technology, to M. S. Dandavati, Hittman Associates, Inc.
13. Kertamus, N. J.. Removal of H>>S on Oxidized Iron. Babcock and
Wilcox Research Center.
lU. Char Oil Energy Development, R&D Report No. 73. Office of Coal
Research.
15. Kunii, D. and 0. Levenspiel. Fluidization Engineering. John Wiley
and Sons, Inc. (1969).
16. Mitchell, G. S.. Ammonia, Its Production from Natural Gas. Lion
Oil Company.
17. Maron, S. H. and C. F. Prutton. Principles of Physical Chemistry.
The MacMillian Company, New York (1958).
18. Klimisch, R. L., and K. C. Taylor. Ammonia Intermediacy as a Basis
for Catalyst Selection for Nitric Oxide Reduction, Environmental
Science and Technology, Vol. 7, No. 2, February 1973.
19. Shelef, M. and H. S. Ganhi. Ammonia Formation in Catalytic
Reduction of Nitric Oxide by Molecular Hydrogen. Industrial
Engineering Chemistry, Vol. 11, No. 1, (1972).
20. Pichler, H.. Advances in Catalysis. Vol.. k, Academic Press.
21. Personal communication with Roy Wensink, The Harshaw Chemical Co.
22. Plummer, J. B., C. J. Kuhre, et al. The Generation of Clean
Gaseous Fuels from Petroleum Residues. AIChE, Tulsa, Oklahoma,
March
-------
REFERENCES (Cont'd)
23. Ter Harr, L. W.. Hydrogen and Synthesis Gas Production by Partial
Oxidation. Second Symposium on Large Chemical Plants. Antwerp,
Belgium, November 1973.
2k. Oldaker, E. C., A. M. Poston and W. L. Farrior. Removal of Hydrogen
Sulfide from Hot Low-Btu Gas With Iron Oxide - Fly Ash Sorbents.
Morgantown Energy Research Center Report, MERC/TPR-75/1. February
1975-
25. McGlamery, G. G., et al. Flue Gas Desulfurization Process Cost
Assessment. EPA Contract 68-02-3150, Task 2, May 6, 1975.
26. Steiner, P., H. Juntgen, and K. Knoblauch. Process for Removal of
Sulfur Dioxides from Polluted Gas Streams. Presented at the l6th
National Meeting of the American Chemical Society.
27. Letter communication, June 6, 1975-' John Richmond, Katalco Corp.,
to M. S. Dandavati, Hittman Associates, Inc.
28. Amick, R. S., et al. Flue Gas Desulfurization Process Cost
Assessment. EPA Contract 68-02-3150, Task 2, May 6, 1975-
29. EGAS Review Presentation, ASME Winter Meeting, December 1975.
30. Giramonti, A. J.. Advanced Power Cycles for Connecticut Electric
Utility Stations. UTRC Report L-971091-2, January 1972.
31. Handy-Whitman Index of Public Utility Construction Costs, Bulletin
102. Whitman, Reguardt, and Associates, Baltimore, Maryland,
January 1975.
32. Economic Indicators. Bi-Weekly Issues of Chemical Engineering.
April 1971 through November 1975.
33. Nelson Cost Indexes 'of Refinery Construction. Oil and Gas Journal
Issues. July 1971 through December 1975.
3^. Power Plant Capital Costs, Current Trends and Sensitivity to
Economic Parameters. Atomic Energy Commission, Report WASH-IS^.
October 1975.
35. Steam Coal Prices. Coal Weekly, July 1975 - November 1975.
-------
REFERENCES (Cont'd)
36. Refined- Products Process. Oil and Gas Journal, July 197^-November
1975.
37. Private Communication. Turbo Power and Marine Systems. May 1975.
38. Aron, 60 Mw Power Plant Study Estimates on Shell Gasification,
Adip and Sulfur Recovery. Lummus Nederland N.V. , JN-6667,
August 30, 1971.
39. Giramonti, A. J. . Nitrogen Oxides Emissions from Gas Turbine Power
Systems Burning Low- Btu Fuel Gas. UTRC Report K-170852-2,
August 12, 1971.
ko. Vogel, G. J. , et al. Reduction of Atmospheric Pollution by the
Application of Fluidized-Bed Combustion and Regeneration of Sulfur
Containing Additives. EPA- 650/2- 7*1-10^, September
lj-1. MaGee, E. M. , et al. Potential Pollutants in Fossil Fuels.
EPA-R2-73-2U9, June 1973.
U2. Ruch, R. R. , et al. Distribution of Trace Elements in Coal.
EPA Symposium on Environmental Aspects of Fuel Conversion Technology,
St. Louis, Missouri, May 197^.
k3. Work Plan for Environmental Study of Coal Conversion Processes.
Hittman Associates under ERDA Contract No. E(^9-l8) - 1508.
hk. Manual on Disposal of Refinary Wastes. Volume on Liquid Wastes,
American Petroleum Institute, Washington, D. C., (1969).
^5. Capp, J. P. and D. W. Gillmore. Soil-Making Potential of Powerplant
Fly Ash in Mined- Land Reclamation. Presented at Third International
Ash Utilization Symposium. Pittsburgh, PA, March 1973.
h6. Radian Corporation. Solid Waste Disposal, Final Report.
Environmental Protection Technology Series, E PA-650/2-7^-033,
May 197U.
k"J. Taylor, W. C. . Combustion Engineering's Experience in the Disposal
and Utilization of Sludge from Lime -Lime stone Scrubbing Processes.
Presented at the Flue Gas Desulfurization Symposium, New Orleans,
LA, May 1973.
-------
REFERENCES (Cont'd)
if8. Raymond, S.. Pulvenized Fuel Ash as Embankment Material. Proc.
Inst. Civil Engr., (London), 19, (1961).
1+9. Martens, D. C. and C. 0. Plank. Basic Soil Benefits from Ash
Utilization: Presented at the 3rd International Ash Utilization
Symp., Pittsburgh, PA, March 1973..
50. Martens, D. C.. Availability of Plant Nutrients in Fly Ash. Dept.
of Agronomy, V.P.I., Blacksburg, Virginia.
51. West, J. R.. New Uses of Sulfur. Advances in Chemistry Series,
A.C.S., Washington, B.C., (1975).
52. Brinkley, Stuart R.. Note on the Conditions of Equilibrium for
Systems of Many Constituents. Journal Chemical Physics, Vol. 1^,
No. 9, pp. 563-56U, September 19^6.
53. Brinkley, Stuart R.. Calculation of the Equilibrium Composition of
Systems of Many Constituents. Journal Chemical Physics, Vo. 15,
No. 2, pp. 107-110, February 19*17.
5^. Kandiner, Harold J. and S. R. Brinkley. Calculation of Complex
Equilibrium Relations. Industrial and Engineering Chemistry, pp.
850-855, May 1950.
55. Brinkley, Stuart R.. Computational Methods in Combustion
Calculations in High Speed Aerodynamics and Jet Propulsion. Vol. IV,
Combustion Processes, Ed. B. Lewis, R. N. Pease, and H. S. Taylor,
Princeton University Press, pp. 6^-98, (1956).
56. Brinkley, Stuart R.. Calculation of the Thermodynamic Properties of
Multi Component Systems and Evaluation of Propellant Performance
Parameters. Proc. of the First Conference, Western States Section,
The Combustion Institute, Los Angeles, California, Nov. 2-5, pp.
7^-81, (1959)-
57. Gibbs, J. Willard. Collected Works of J. Willard Gibbs. Vol. I,
Longmans, Green and Co., New York, pp. 63-97, (1928).
58. Keenan, Joseph H., F. Keyes, P. G. Hill, and J. G. Moore. Steam
Tables. John Wiley & Sons, Inc., New York, (1969).
2U6
-------
REFERENCES (Cont'd.)
59. Glen, R. A. and R. J. Grace. An Internally-Fired Process Develop-
ment Unit for Gasification of Coal Under Conditions Simulating Stage
Two of the BCR Two-Stage Super-Pressure Process. 1968 AGA Synthetic
Pipeline Gas Symposium, Pittsburgh, Pa., November 22, 1968.
60. Glenn, R. A.. Status of the BCR Two-Stage Super Pressure Process,
at the Third Synthetic Pipeline Gas Symposium, American Gas Associa-
tion, Chicago, 111., November 17-18, 1970.
6l. Zahradnik, R. L. and R. A. Glenn. Direct Methanation of Coal, Fuel,"
Vol. 50, pp. 77-90, 1971.
62. Zahradnik, R. L. and R. J. Grace. Chemistry and Physics of Entrained
Coal Gasification, Division of Fuel Chemistry Symposium, American C
Chemical Society, Dallas, Texas, April 9-10, 1973- Published by
American Chemical Society as Coal Gasification (ed. Lester G. Massey),
197k.
63. Behrens, H.. Ressbildung und Radikalgleichgewichte in Flammen, Zeit.
Fur Phys. Chemie, Vol. 199, PP. 1-lU, 1952.
6k. Walker, Philip L.. (Ed.) Chemistry and Physics at Carbon, Vol. 1,
pp. 297, 1965.
-------
APPENDIX A
EQUILIBRIUM MODEL FOR COAL GASIFIERS
The equilibrium composition of solids and gases provides a good
estimate for the product gas composition of certain coal gasifiers. If
the gasifier operates at a high temperature (i.e. 3000 F) with the gas-
ifier size such that there is enough residence time for reactions to go
to completion and if stratification may be neglected within the gasifier
then the equilibrium composition will provide a good estimate of the
gasifier product gas composition. Examples of gasifiers for which
equilibrium is a good assumption are the Koppers-Totzek and Kellogg mol-
ten salt gasifiers. It should be emphasized that for gasifiers with
strong stratification and limited residence times, e.g., fluid beds,
and for gasifiers with low temperatures and short residence times, e.g.
BCR's upper stage, the product gas composition is kinetic limited and
not equal to the equilibrium composition. The equilibrium composition
does, however, provide an important guideline for environmental studies.
If the pollutant equilibrium concentration is much lower than the con-
centration in the product gas, then catalytic acceleration of reactions
involving the pollutants toward their equilibrium level can reduce the
pollutant concentration to an acceptable level.
Computer programs to calculate the equilibrium composition of
mixture of gases or gases and a solid have now been in wide use for many
years following the pioneering work of Brinkley (52-56), who developed
very elegant computational procedures for arbitrary mixtures of elements
to solve the basic thermodynamic equations developed by Gibbs (57).
These computer programs were first used in the late 1950's, with
Brinkley as a consultant, and have been updated and improved since then.
Recently the necessary modifications were made to input coal, char,
steam, transport gases and air to the model and to have it compute the
equilibrium composition. Included are solid phase ash and carbon as
well as an arbitrary number of gas compounds and the option to specify
set yields (for kinetic limited products). In Gibbs' model the solids
are assumed to be finely divided and dispersed among the gases--
surface effects are not included. The gases are assumed to follow
the perfect gas law.
-------
The input to the equilibrium program includes the atoms per unit
weight of each element and either the temperature and pressure, the
enthalpy and pressure, the entropy and temperature, the enthalpy and
entropy, the density and pressure, the temperature and density or the
internal energy and density. The unspecified thermodynamic properties
are calculated as well as .the gas mixture molecular weight and mole frac-
tions of the mixture of gases and the solid ash and carbon weight concen-
trations. -The list of compounds to be considered is specified as input
and may be, in principle, arbitrarily long. Of course, the computer
time goes up - quadratically - as the number of compounds.
In simulating coal gasifier operation, the input quantities are
coal, air (or oxygen), steam and possibly transport gases. The enthalpy
of the air and transport gases can be calculated easily using the same
reference states (C (solid), H2 (gas), N2 (gas), 0^ (§as) etc. have
zero enthalpy) and reference temperature (298 K) as for the product gas
using heat of formation and specific heat curves based on JANAF data.
Steam enthalpies can be obtained from the literature^oJ an£ converted
to the reference temperature used in the JANAF tables.
Coal enthalpies involve a slightly more complicated procedure. A
heating value for the coal studied may be determined from a bomb calori-
meter test of the coal. Alternatively a version of Dulong's formula
with approprieate coefficients for the coal can be used. If the coal is
heated prior to its entry into the gasifier, then the specific enthalpy
must be determined by tests on the coal. Specific note should be made
as to whether the coal is superheated so that the energy cost of vaporiz-
ing the moisture in the coal is accounted for external to the gasifier.
Finally, the specific heat and higher heating value of the coal must be
converted to equivalent enthalpies for C (solid), ash (solid), S (solid),
tL, (gas) and Q^ (gas) to be consistent with the other data. This is
done following the scheme in Fig. ^6. The bomb calorimeter test results
in products of f^O (liquid), CC>2 (gas) and 50^ (gas) at 298 K as well as
ash (solid) and excess 02 (gas). After accounting for the latent heat
of the water and converting the higher heating value (HHV) to a lower
heating value (LHV) the heats of formation of C02> ^2> an(^ HgO (aH
gases) may be computed using the JAMAF data heats of formation. Thus
the total enthalpy of the coal (with respect to the reference tempera-
ture and reference states used in the JAMAF data) equals the sensible
heat plus the higher heating value of the coal less the latent heat less
the heat of formation of C02, SC>2, and H^O. It has been noted that the
total enthalpy is not zero because of the specific heat and because of
the energy contained in exceedingly complex coal carbohydrates' chemical
bonds. The question of what enthalpy should be used for char is best
-------
FLOW CHART FOR COAL ENTHALPY CONVERSION
SCHEME FOR CONVERTING COAL SPECIFIC ENTHALPY PLUS
HIGHER HEATING VALUE TO ENTHALPIES OF C (SOLID), ASH (SOLID),
S (SOLID), H2 (GAS) AND 02 (GAS)
FIG. 46
COAL AT TEMPERATURE T
C, S, ASH, H, 0, H20
I
H
sensible
COAL AT 298 K
C, S, ASH, H, 0, H20
HHV
BOMB
CALORIMETER
H20 (LIQUID), C02 (GAS)
S02 (GAS) AT 298 K
latent
H20 (GAS) , C02 (GAS)
S02(GAS) AT 298K
S02
C (SOLID)
AT298 K
S (SOLID)
AT 298 K
H2 (GAS)
AT 298 K
Htotal ~ Sensible + HHV ~ Hlatent ~HFC02 ~ HFS02 ~HFH20
OR
H H
— l_l 1
total sensible
LHV-HFrn -HFqn -HF
CUo oUo
H00
WHERE
HF = HEAT OF FORMATION
76-02-128-1
250
-------
answered by combusting a sample of the char in a bomb calorimeter. If
the carbon in the char were simply fixed carbon, the higher heating value
will simply be the heat of formation of C02« If the higher heating value
exceeds the heat of formation of C02, then energy containing chemical
bonds between carbon atoms were still present in the char. It should be
noted that the variation of the higher heating value of a coal, or char,
determined from bomb calorimeter tests may vary by 100 to 200 Btu/lb
input.
The results of a computer model used by Koppers Co. and the UTRC
model are shown in Table 6k. In this case the coal temperature (l60°F)
and HHV, steam enthalpy (from Ref. 7) and oxidizer temperature (98 per-
cent 02 at 220°F) together with the composition and weight flows yielded
the mixture enthalpy as well as the element weight flows. The equili-
brium program then calculated the equilibrium temperature and composi-
tion of the mixture. As shown in Table 6l the predicted temperature was
85°F below the K-T. model temperature. Therefore, the equilibrium model
was adjusted to the K-T model specified temperature with the resultant
gas composition as shown in Table 6k. The agreement is very good except
for the water-gas reaction components. Therefore, the water-gas shift
equilibrium constant was calculated from both the mole fractions given
by K-T and the equilibrium model and compared with values at that tempera-
ture (2732°F) calculated directly from the JAMAF data for Gibbs free
energies. The water-gas shift equilibrium constant calculated for the
K-T results is 2.62 and from both the equilibrium model and Gibbs free
energies at 2732°F is 3-68. This leads to the conclusion that K-T has
modified the results of an equilibrium calculation to reflect operating
experience. This question has not been pursued since this study did not
focus on the Koppers-Totzek gasifier.
An additional point in using equilibrium gasifiers is the high sen-
sitivity of the operating temperature to the enthalpy input—in particular
to the coal heating value. Varying this parameter for the Koppers-Totzek
gasifier shows the strong dependence, approximately 2.1 F.Btu/lb total
(0.95 F/Btu/lb-coal) of gasifier operating temperature on coal HHV input
(Fig. U7). (The coal amounted to approximately ^5 percent of the weight
input.) The strong dependence is typical of reducing atmospheres, as
opposed to oxidizing atmospheres where a more typical dependence is
0.5 F/Btu/lb. Figure 1+7 illustrates that inaccuracies of 100 Btu/lb in
coal HHV can lead to differences of 95°F in predicted gasifier product
gas temperatures. Since the coal HHV is usually only determined to this
accuracy, the resulting difference in product gas temperature should be
included in system planning.
251
-------
Table 64
COMPARISON OF RESULTS FROM KOPPERS-TOTZEK
AND UTRC GASIFIER MODELS
Koppers-Totzek Model Equilibrium Model __
Enthalpy Specified T Specified
T (F) 2732 2647 2732
Enthalpy (Btu/lb) -1082 -1122 -1082
Product gas
(mole fractions)
co .4969 .5055 .5070
co2 .0631 .0538 .0522
H ' .3179 -3093 .3078
H20 .1056 .1149 .1164
N .0096 .0096 .0096
HgS & COS .0069 .0066 .0064
Others (HS, etc) 0 .0022 ' .0005
Others (H, etc) 0 .0001 .0001
Solids Output
(weight fractions)
Ash .0317 .0317 .0317
C • .0171 .0171 .0170
252
-------
14000 I—
CD
- 13500
O
o
FIG. 47
DEPENDENCE OF PRODUCT GAS TEMPERATURE ON ENTHALPY
INPUT UNDER EQUILIBRIUM
13000
-900
-1000
-1100
I
I-
z
-1200
-1300
2.10
AVERAGE SLOPE
BTU / LB - COAL
0.95
BTU / LB - TOTAL
2000
2500
PRODUCT OF GAS TEMPERATURE
3000
76-02-128-2
253
-------
APPENDIX B
BITUMINOUS COAL RESEARCH, INC. (BCR) TWO-STAGE GASIFIER MODEL
It had been planned to use the equilibrium model described in
Appendix A as a basis for developing a model of the two-stage BCR
gasifier. When the basic equilibrium model was developed and tested on
a single-stage gasifier, contact was made with BCR in order to obtain the
data necessary to proceed with further modeling. After some preliminary
discussions, BCR consented to make available their computer model of the
two-stage gasifier together with advice as to its use contingent upon
the approval of ERDA under whose auspices the model had been developed.
ERDA's subsequent approval is a good example of the benefits of intra-
government cooperation in the energy area. With the availability of the
BCR model, further work on the equilibrium model was suspended and efforts
focused on implementing the BCR model. Following is a description of
this model and of the assumptions necessary to use it.
The model of the two-stage gasifier is based on limited experimental
results of the rapid devolatilization of coal in stage 2 (the upper stage)
in a 100 Ib/hr Process Equipment Development Unit (PEDU). Stage 1 has
not yet been tested. A diagram of the design is given in Fig. ^8. Coal,
carried by a transport gas, and steam are fed into Stage 2 where they
encounter hot gases and char rising from Stage 1. The coal is rapidly
devolativlized, with a high methane yield, at relatively low temperatures
(1700-1900 F) and the resultant gases and char are swept out to a series
of cyclone separators. The clean product gas, possibly with some particu-
late matter (char), indicated as "withdrawal" in Fig. U8 is sent to the
sulfur cleanup system. The precipitated char from the cyclones is then
recycled into Stage 1 of the gasifier where it is burned. An oxidizer
(oxygen or air) and steam are fed in to provide oxygen to combust the
char at high temperatures (2800-3000 F). The steam acts as both a tem-
perature moderator and source of hydrogen since the methane yield in
Stage 2 is strongly a function of the partial pressure of hydrogen^
Ash is slagged out of Stage 1, possibly with the aid of a limestone
catalyst (not shown) and the gases and unburned char are swept up into
Stage 2 completing the cycle. The steam in Stage 1 is used mainly as a
-------
FIG. 48
SIMPLIFIED MASS FLOW DIAGRAM FOR THE TWO-STAGE BCR GASIFIER
COAI
TRANSPORT
GAS
STEAM-
OXIDIZER
STEAM
STAGE 2
1700 -1900F
GAS
CHAR
CYCLONE
PRODUCT
GAS
WITHDRAWAL
SLAG
255
76-02-113-1
-------
temperature modulator, with a secondary effect on the methane yield
via the hydrogen partial pressure.
Experience at BCR with the PEDU has shown that the product gas
produced in Stage 2 is limited by process kinetics, i.e., the gases pro-
duced are very different from the equilibrium mixture at the operating
temperature. In order to estimate the product gas composition BCR speci-
fies the yield of methane and oxidized carbon (CO, COg) from Stage 2
based upon empirical curves derived from both their own experiments with
their PEDU and experiments of other researchers(°1~63). The amount of
carbon which is combusted in Stage 1 is also specified usually at a con-
servative value of 60 or 70 percent. The water-gas shift reaction is
the only reaction assumed to go to equilibrium, in both stages. All sul-
fur is assumed to go to HpS, although Foster- Wheeler has modified this
slightly by introducing a specified"HgS to COS ratio.
These specifications determine algebraically a surprising amount
of the composition of the gas and char streams, as is shown in Fig. ^9.
The gas and solid stream in Fig. hQ are the same as those iri Fig. ^9
with the addition of the gas and solid streams from Stage 1 to Stage 2.
The term "YCHV is the weight .fraction of the C input in the COAL stream
which is converted to CHI), in Stage 2. The term "YCO" is the weight
fraction of the C input in the COAL stream which is oxidized to CO and/or
C02 in Stage 2. The C in the CHAR 1 stream from Stage 1 is assumed not
to react. (This fixed carbon has survived exposure to both Stage 2 and
Stage 1 at least once without reacting.) The term "WDRAWL" is the frac-
tion of CHAR 2 which is not separated out in the series of cyclones,
i.e., it is contained in the product gas. For use in an integrated
power system the 1 percent or less of CHAR 2 would be removed in the
subsequent fuel gas cleanup system. The term "YC" is the fraction of C
in CHAR which is oxidized to CO and/or CX^- in Stage 1, Algebraic
solution of the element mass balances of the solids streams gives a com-
plete solution for all of the solids streams in terms of the COAL stream
input. Similarly, except for the water gas shift,'the gas stream mass
balances can be solved algebraically. It then remains to adjust the
mass flows "OXIDIZER", "STEAM 1" and, if included, "STEAM 2" to achieve
the desired operating temperatures in Stages 1 and 2. This is accom-
plished in an iterative manner by first adjusting the STEAM 1 flow rate
until the desired Stage 1 temperature is achieved and then adjusting the
OXIDIZER flow rate until the Stage 2 temperature is achieved. The energy
balances (enthalpy in equals enthalpy out) in the two stages are the
dependent functions and the flow rates of OXIDIZER and STEAM 1 are the
variables for the iterative solution process.
256
-------
BREAKDOWN OF MASS FLOWS IF YIELD IS SPECIFIED
( MASS FLOWS, LB/HR., YCO, YCH4, YC AND WDRAWL ARE FRACTIONS)
FIG. 49
COAL j.
TRAMQ ^
STEAM 2 ».
G
OXIDIZER »•
STEAM 1 *•
AS 1
STAGE 2
CHAR 1
STAGE I
GAS 2
CHAR 2
CYCLONE
\
CHAR
/
GAS 2
WITHDRAWAL
(CHAR)
SLAG
WHERE CHAR CONSISTS OF
C IN CHAR =C IN COAL X( 1-YCO - YCH4 ) X (1-WDRAWD/D
ASH IN CHAR = ASH IN COAL X ( 1- WDRAWD/D
SLAG CONSISTS OF
ASH IN SLAG = ASH IN COAL X (1- WDRAWL ) X YC/D
CHAR 1 CONSISTS OF
C IN CHAR 1 =C IN COALX(1-YCO-YCH4) X (1-WDRAWL) X (1- YC)/D
ASH IN CHAR 1 = ASH IN COAL X (1- WDRAWLH 1-YC )/D
CHAR 2 CONSISTS OF
C IN CHAR 2 = C IN COAL X (1- YCO-YCH4)/D
ASH IN CHAR 2 = ASH IN COAL/D
WITHDRAWAL CONSISTS OF
C IN WITHDRAWAL = C IN COAL X (1-YCO -YCH4) X WDRAWL/D
ASH IN WITHDRAWAL = ASH IN COAL X WDRAWL/D
WHERE D = 1-(1-WDRAWL)(1-YC)
YC = FRACTION OF C IN CHAR OXIDIZED IN STAGE 1
YCO = FRACTION OF C IN COAL OXIDIZED IN STAGE 2
YCH4 = FRACTION OF C IN COAL CONVERTED TO CH4 IN STAGE 2
. WDRAWL = FRACTION OF CHAR2 CARRIED OVER IN PRODUCT GAS (GAS2)
76-02-113-3
257
-------
Using the foregoing to estimate the energy and mass balances and
operating temperatures, the gasifier operation is reduced to specifying
YC, WDRAWL, YCHU and YCO. As mentioned above, YC, the fraction of the
carbon in the char which is oxidized in Stage 1, is conservatively esti-
mated by BCR at 60 or 70 percent. When operating data for Stage 1
becomes available this will be set accordingly. WDRAWL depends on the
number and efficiency-of cyclone separators. Since excessive carryover
results in both operating inefficiencies and potential downstream damage,
it is preferrable to use high-efficiency cyclones to recycle essentially
all of the char from Stage 2; hence, WDRAWL is 1 percent or less. The
methane yield, YCHU, depends intimately both on the partial pressure of
hydrogen in Stage 2 and on the Stage 2 temperature and has been the sub-
ject of much study at BCR and elsewhere because of the interest in
using coal gasifiers to synthesize conventional pipeline gas. The BCR
method of selecting YCHU is given below. The value of YCO which deter-
mines how much carbon in the coal does not react, i.e., "fixed carbon",
may be expected to be such that YCH4 + YCO = C in the volatiles in the
coal (approximately). In running the model YCH^t- + YCO equals C in the
volatiles within about 30 percent. At the same time, additional carbon
may be formed due to the chemical kinetics toward carbon formation from
the gases. To account for this, consideration must be given to the
Boudouard reaction equilibrium.
It may be noted that YCO provides a conceptually useful and
informative link between the two gasifier operation regimes currently
under consideration — the low temperature, rapid devolatilization
scheme in the BCR gasifier and the high-temperature equilibrium scheme .
in Koppers-Totzek gasifier. If YCO is simplistically increased from 15
percent to 80 percent, the gasifier goes from a favorable region of
operation — high product gas HHV, low steam and oxidizer requirements —
over a "hill" or unfavorable regime — low product gas HHV, high steam
and oxidizer requirements — to another favorable region. The latter
region in which all the C in the coal is gasified (equilibrium) is the
region used for many years by Koppers-Totzek and others. The efficiency
advantage of the BCR gasifier is due to, l) the existence of the first
favorable region of gasifier operating having rapid devolatilization
and low YCO (Stage 2) in which only part (essentially the volatiles) of
the coal is gasified, and 2) recycle of the char (the uncombusted C in
the coal plus ash) to furnish the energy (Stage l) and thereby closing
the gasifier operation loop.
258
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The specification of YCHU is based upon the BCR PEDU experiments
and other laboratory data on the rapid devolatilization of coal at
high pressures'""'^"). The conclusion is that YCHU can be estimated
empirically as follows:
a + b • PH
(20)
where
a = .08
b2
b = b]_ exp[ ] where b]_ and b2 are constants and T is
temperature
Pjjp = partial pressure of hydrogen
There are several methods of determining the values of b^_ and b2 using
both theoretical kinetics arguments and experimental data. Fortunately,
values of YCHU found from these various methods are in relatively close
agreement as can be seen in Figs. 50a and 50b.
Estimates of YCHU for the temperatures and hydrogen partial pressures
of interest reveals quite close agreement between the curves based upon
Refs. 59 and 60. The curve based on Ref. 6l is somewhat different.
Since this curve is based upon experiments at very high hydrogen partial
pressures (up to 500 atm) and the operating conditions for the typical
air blown cases in this study have hydrogen partial pressurs of only
5-10 atm, the curve based on Ref. 6l was not used. In this study the
methane yield YCHU was assumed to lie between the two curves determined
from the data in Refs. 59 and 60.
YCO is specified according to the Boudouard reaction
2 CO ^ C02 + C (solid) (21)
This reaction does not actually occur directly but may occur in two
steps (56). It nevertheless serves as a touchstone to determine whether
the specification of the amount of carbon gasified is correct for the
assumed operating conditions. The basic assumption is that the mole
fraction of CO, XQQ, in the gas being formed in Stage 2 uniformly
increases during the course of the chemical reactions occurring there.
Then, XQQ may be compared to its equilibrium value according to the
Boudouard reaction. If YCO is too high, then XQQ will be above Boudouard
reaction equilibrium and the Boudouard reaction would suggest that solid
carbon is being formed. During the rapid devolatilization of Stage 2,
one would not expect the carbon in the coal to oxidize and then revert
259
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FIG-50
METHANE YIELD VERSUS PARTIAL PRESSURE OF HYDROGEN
(a)
30
20
YCH4
10
T = 1700F
REF
53
USED
10
15
PH2 (atm)
20
25
YCH4
76-02-113-2
260
-------
into solid carbon again, i.e., one would not expect XCQ ^° overshoot
its equilibrium value. Rather, the question is whether XCQ even reaches
its Boudouard reaction equilibrium value. The BCR approach is to choose
YCO such that XQQ is below Boudouard reaction equilibrium (CC>2 amd C tend
to form CO) but that YCO yields X^Q as close to equilibrium as possible.
During the course of this- study hundreds of runs with the BCR model
resulted in only about one-half dozen cases having the Boudouard reaction
constant (see below) below equilibrium (toward solid C), while concurrently
achieving a mass and energy balance in the gasifier under the specific
operating temperatures and pressure. In general, several values of YCO
above equilibrium (toward CO) were possible and the value closest to equi-
librium was chosen in order to determine whether X«Q and ^Op were above
or below Boudouard reaction equilibrium. It is instructive to calculate
an equilibrium constant and determine whether the gasifier would "operate"
at values for XQQ and ^CO^ sucb. that the CO or C02 were above equilibrium.
By classical thermodynamics at equilibrium
(22)
RT
V-»W
where
XCO' XCO = mole
p = pressure in atmospheres
(ipQ, n° , |j,p = Gibbs free energies (chemical potentials) of CO,
2 C02j and solid C
T = temperature
R = gas constant
Thus a simple way to check the Boudouard reaction equilibrium is to
calculate
_ P-x
K.
[XCQ /(p.xc2)]equilibrium
where the denominator is determined from the equation above and the
numerator from the values calculated by the mass and energy balance in
the BCR gasifier computer model. If K ^ 1, then X^Q is below equilibrium
and the reaction is tending away from solid C formation. If K < 1 the
Boudouard reaction tends toward solid C formation. In this study YCO
was chosen to make K as close to 1 as possible, but always greater than
1. As mentioned above, very few conditions were found with K < 1.
261
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In addition to adapting the BCR model to our computer system and
carrying out numerous parametric cases, additional modifications were
made. Enthalpy and Gibbs free energy curves from the UTRC library
were used. Specific heat curves for SiOg were used for ash. Also added
were a check on the Boudouard constant, K defined above, and a check
to see if the specified YCH^ fell between the two curves based on the
empirical data which was supplied by BCR. This addition will allow an
automatic iterative procedure to be used when the gasifier model is
eventually incorporated into the SOAPPW program used at UTRC.
Further work is needed to determine how far from equilibrium the
gasifier can operate, in terms of the Boudouard constant K, defined
above. It was found during parametric analyses that as the operating
conditions were varied from the design point, values of K became several
hundred then several thousand. Clearly, this is inconsistent with the
assumption of water gas shift equilibrium, but it is not clear what
the K cutoff point value is. At design point, K was typically 3-6.
262
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APPENDIX C
During the initial phase of the program, first and second genera-
tion power system characteristics were defined. First generation systems
use conventional cooling techniques with a turbine inlet temperature of
2200°F. Second generation systems use ceramic vanes with conventionally
cooled blades with a turbine inlet temperature of 2600°F. In each case
a turbine pressure ratio was selected to give high specific power (kW
output per unit air flow) while not compromising system performance.
This selection was made initially on the basis of data for distillate
fired systems considered. For the revised systems presented in this
report, the effect of pressure ratio on both first and second generation
systems was again evaluated. The results are presented in Figs. 51 and
52. They show little change from the previous curves and the selection
appears to remain valid.
263
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FIG. 51
FIRST GENERATION SYSTEM PERFORMANCE
TURBINE INLET TEMPERATURE = 2200F
CONVENTIONAL AIR COOLING
45
40
PRESSURE RATIO =24
32
o
z
LLJ
O
35
SIMPLE CYCLE/DISTILLATE
32 24 32
24
30
25
50
COG AS/DISTILLATE FUEL
COGAS/BUMINES/IRON
8 OXIDE
COGAS/BUMINES/SELEXOL
I
I
100 150 200
SPECIFIC POWER - kW/LB/SEC
250
76-02-212-1
-------
SECOND GENERATION SYSTEM PERFORMANCE
TURBINE INLET TEMPERATURE = 2600F
CERAMIC VANES, CONVENTIONAL BLADES
FIG.52
50
45
u
z
UJ
o
£ 40
UJ
35
30
PRESSURE RATIO = 32 24
SIMPLE CYCLE/DISTILLATE
,32
COG AS/DISTILLATE FUEL
16 COGAS/BCR/CONOCO
COGAS/BCR/SELEXOL
100
150 200 250
SPECIFIC POWER - kW/LB/SEC
300
350
76-02-212-2
265
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TECHNICAL REPORT DATA
(Please read Inunctions on the reverse before completing)
1. REPORT NO.
EPA-600/2 -76-153
3. RECIPIENT'S ACCESSION-NO.
4. TITLE AND SUBTITLE
fuel Gas Environmental Impact
5. REPORT DATE
June 1976
6. PERFORMING ORGANIZATION CODE
L Robson and w A Blecher (UTRC) and
C. B. Colton (Hittman Associates, Inc.)
8. PERFORMING ORGANIZATION REPORT NO.
9. PERFORMING ORGANIZATION NAME AND ADDRESS
United Technologies Research Center (UTRC)
Silver Lane
East Hartford, Connecticut 06108
10. PROGRAM ELEMENT NO.
EHB529
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
Final; 11/74-11/75
14. SPONSORING AGENCY CODE
EPA-ORD
is. SUPPLEMENTARY NOTES IERL-RTP Project Officer for this report is W.J.Rhodes, Mail
Drop 61, Ext 2 851.
16. ABSTRACT
As they relate to combined cycle power generation, the report gives
results of an evaluation of the technical and economic considerations of atmospheric-
pressure, oxygen-blown coal gasifiers (Koppers-Totzek) and pressurized, air-blown,
partial-oxidation residual-oil gasifiers (Shell/Texaco). Also presented are refine-
ments of systems reported in an earlier phase report, EPA-600/2-75-078. The
objective of the report is to help define the environmental impact of combinations of:
(1) fossil fuel gasification systems, (2) low- and high-temperature fuel gas cleanup
processes, and (3) advanced cycle power systems.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
Residual Oils
Gasification
Air Pollution
Fuels
Gas Purification
Fossil Fuels
\_oal Gasification
Electric Power Generation
b.IDENTIFIERS/OPEN ENDEDTERMS
Air Pollution Control
Stationary Sources
Fuel Gas
Environmental Impact
Combined Cycle Power
Generation
Emission Control
c. COSATI Field/Group
13B
2 ID
07A,13H
10A
11H
8. DISTRIBUTION STATEMENT
Unlimited
19. SECURITY CLASS (This Report/
Unclassified
21. NO. OF PAGES
281+
20. SECURITY CLASS (This page)
Unclassified
22. PRICt
EPA Form 2220-1 (9-73)
266
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