U.S. Environmental Protection Agency
Office of Research and Deve!opnu;n-
~,v i ronTenTdl Rt'sedfCh
dMilU- ptirk. NorH) Cdrolma 27711
EPA-600/7-77-065
JUtlS 1977
WATER CONSERVATION
AND POLLUTION CONTROL IN
COAL CONVERSION PROCESSES
Interagency
Energy-Environment
Research and Development
Program Report
-------
RESEARCH REPORTING SERIES
Research reports of the Office of Research and Development, U.S.
Environmental Protection Agency, have been grouped into seven series.
These seven broad categories were established to facilitate further
development and application of environmental technology. Elimination
of traditional grouping was consciously plarined to foster technology
transfer and a maximum interface in related fields. The seven series
are:
1. Environmental Health Effects Research
2. Environmental-Protection Technology
3. Ecological Research
4. Environmental Monitoring
5. Socioeconomic Environmental Studies
6. Scientific and"Technical Assessment Reports (STAR)
7. Interagency Energy-Environment Research and Development
This report has been assigned to the-INTERAGENCY ENERGY-ENVIRONMENT
RESEARCH AND DEVELOPMENT series. Reports in this series result from
the effort funded under the ,17-agency Federal Energy/Environment
Research and Development Program. These studies relate to EPA's
mission to protect the public health and welfare from adverse effects
of pollutants associated with energy systems. The goal of the Program
is to assure the rapid-development of domes'tic energy supplies in dn
environmentally—'compatible manner by providing the necessary
environmental data and control technology. Investigations include
analyses of the transport.,of energy^related pollutants and their health
and ecological effects; assessments of, and development of, contrpl
technologies for energy systems; and integrated assessments of a wide
range of energy-related environmental issues*
REVIEW NOTICE
This report has been reviewed by the participating Federal
Agencies, and approved for publication. Approval does riot
signify that the contents necessarily reflect the views and
policies of the Government, nor does mention of trade names
or commercial products constitute endorsement or recommen-
dation for use.
This document is available to the public through the National Technical
Information Service, .Springfield,,Virginia 22161.
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EPA-600/7-77-065
June 1977
WATER CONSERVATION
AND POLLUTION CONTROL IN
COAL CONVERSION PROCESSES
by
D.J. Goldstein and David Yung
Water Purifcation Associates
238 Main Street
Cambridge, Massachusetts 02142
Contract No. 68-03-2207
Program Element No. EHE623
EPA Project Officer: Mark J. Stutsman
Industrial Environmental Research Laboratory
Office of Energy, Minerals, and Industry
Research Triangle Park, N.C. 27711
Prepared for
U.S. ENVIRONMENTAL PROTECTION AGENCY
Office of Research and Development
Washington, D.C. 20460 '
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ACKNOWLEDGMENTS
Stone & Webster Engineering Corporation, under subcontract, assisted in
all phases of this report. Most of the sections represent our joint effort
and we are particularly grateful for this. The Stone & Webster people who
worked on this project are Winthrop D. Comley (Project Manager), John M. Donohue,
Carl H. Jones, William C. Panciocco, Frederick B. Seufert and many others.
We were given the, opportunity to visit most of the operating pilot
plants and many engineering firms actively engaged in coal conversion tech-
nology across the country, and it is our pleasure to acknowledge the kind
reception and time given to us. Among the many places where we were received
we wish to thank Bituminous Coal Research (R. K. Young and R. J. Grace),
C. F. Braun and Co. (Roger Detman, Richard Howell and others), The Bureau of
Mines at Morgantown (Sidney Katell), The Bureau of Mines at Bruceton
(Albert Forney, Sayeed Akhtar, James Gray, Robert Lewis and many others),
Conoco Coal Development Company (Carl Fink), EPA Pacific Northwest Laboratory
(James Chasse), Fluor Engineers and Construction (Max Palmer and
Ralph Beardsley), FMC Corp. Research Center (Haig Terzian; FMC Corp. supplied
water samples for analysis), Hydrocarbon Research, Inc. (F. D. Hoffert),
Institute of Gas Technology (Bernard S. Lee, Louis Anastasia and many others;
IGT supplied basic information on the Hygas process under subcontract and
have collected and shipped many water samples), Koppers Company Inc.
(G. V. McGurl, Reginald Wintrell and Michael Mitsak), Pittsburgh and Midway
Coal Mining Co. (Russel Perrussel and many others; Pittsburgh and Midway
collected and shipped many water samples). In addition, many other people
have been most generous with their time through lengthy telephone calls and
correspondence and have helped us greatly by sending literature often unavail-
able through normal channels.
Many equipment manufacturers did tests and supplied quotations and
cost information without which this report could not have been written.
Among the many firms supplying information it is our pleasure to acknowledge
Dorr-Oliver; Eimco; Ecodyne Graver Div.; FMC Corp. Environmental Equipment
Div.,- B. F. Goodrich; Otto H. York Co.; Permutit Co., Inc; Rohm and Haas;
Union Carbide Corp., Linde Div.; and United States Steel Engineers and
Consultants, Inc. We apologize to those whose help we have neglected to
acknowledge.
Section 18 on Biological Treatment was prepared by Irvine Wei.
-------
CONTENTS
Acknowledgements ii
Figures vii
Tables x
Conversion of American to International (SI) Units xv
1. Introduction
1.1 Purpose of the Study 1
1.2 Organization of the Report 1
1.3 Water Quantities 1
1.4 Water Quality and Treatment 4
2. Conclusions. 5
2.1 Water Consumption 5
2.2 Water Treatment 5
3. Recommendations 8
Part 1 - Water Quantities
4. Coal Conversion Processes 11
4.1 Introduction 11
4.2 Lurgi Process 15
4.3 Bigas Process 17
4.4 C02-Acceptor Process 25
4.5 Agglomerating Burner-gasification Process 29
4.6 Winkler Process 33
4.7 The Stirred-bed Process 35
4.8 Molten Salt Process 39
4.9 The Lurgi Process for Utility Fuel Gas Production 42
4.10 Koppers-Totzek Process 45
4.11 H-Coal Process 50
4.12 Synthoil Process 51
References Section 4 63
5. Hygas Process 69
5.1 Introduction and Summary of Results at Wyoming Site 69
5.2 Material Balance, Wyoming 71
5.3 Heat Balance, Wyoming 78
5.4 Ultimate Disposition of Unrecovered Heat, Wyoming 81
5.5 Water for Flue Gas Desulfurization 82
5.6 Hygas Process at New Mexico and North Dakota 82
References Section 5 87
(continued)
111
-------
CONTENTS (continued)
6. Synthane Process 89
6.1 Introduction and Summary of Results 89
6.2 Material Balance 91
6.3 Heat Balances 91
6.4 Ultimate Disposition of Unrecovered Heat 100
6.5 Water for Flue Gas Desulfurization 100
References Section 6 102
7. Solvent Refined Coal 104
7.1 Introduction and Summary of Results 104
7.2 Design Procedure 105
7.3 Material Balance on Dissolving Section 109
7.4 Effluent Water 112
7.5 Gasification of Carbonaceous Filter Residue 117
7.6 Production of Hydrogen by Steam Reforming of Gas 122
7.7 Total Plant Process Water Streams 122
7.8 Heat Balance on the Dissolving Section 127
7.9 Plant Driving Energy 129
7.10 Thermal Efficiency 129
7.11 Ultimate Disposition of Waste Heat 129
References Section 7 133
8. Other Process Water Needs 135
8.1 Gas Purification 135
8.2 Flue Gas Desulfurization 138
8.3 Water from Drying Coal 146
References Section 8 147
9. Power Generation via Coal Gasification in Combined-cycle
Power Plants 150
9.1 Introduction and Conclusions 150
9.2 Comparison with a Coal Burning Steam-electric Generating
Plant 150
9.3 Description of Combined-cycle Generation 155
9.4 Design Details of Gasifier and Combined-cycle Plants. . . . 160
9.5 Effect of Hot Gas Desulfurization 165
9.6 Cooling in a Power Plant 165
References Section 9 177
10. Cooling 178
10.1 Introduction 178
10.2 The Cost of Water 180
10.3 Cooling Process Streams . 181
10.4 Water Evaporated for Wet Cooling 187
10.5 Cooling in Acid Gas Removal 189
10.6 Characteristics of Steam Turbines 190
10.7 Dry and Wet Cooling Systems for Turbine Condensers 191
10.8 Water Consumption for Turbine Condensers at Specific
Sites 200
10.9 Water Consumption for Interstage Cooling on Gas
Compression 208
References Section 10 223
(continued)
iv
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CONTENTS (continued)
11. Water for Mine Complex and Other Off-site Uses 224
11.1 Introduction and Summary of Results 224
11.2 Road, Mine and Embankment Dust Control 224
11.3 Handling and Crushing Dust Control 227
11.4 Sanitary and Potable Water 228
11.5 Service and Fire Water 229
11.6 Revegetation 229
11.7 Satellite Town 230
References Section 11 , 231
12. Additional On-site Water Streams 232
12.1 Introduction and Summary of Results 232
12.2 Evaporation 232
12.3 Ash Disposal 236
12.4 In-plant Dust Control 237
12.5 Sanitary, Service and Fire Water 238
References Section 12 239
13. Site Studies - 1: Water Consumption 240
Part 2 - Water Treatment
14. Water Analyses 253
14.1 Source Waters 253
14.2 Foul Process Condensate 258
14.3 Clean and Intermediate Process Water 271
14.4 Boiler Feed Water 273
14.5 Cooling Water 273
References Section 14 282
15. Water Treatment Technologies 285
15.1 Introduction 285
15.2 Wet Oxidation 285
15.3 Granular Activated Carbon Adsorption 288
15.4 Freezing 293
15.5 Evaporation 294
15.6 Treatment of Circulating Cooling Water 297
15.7 Reverse Osmosis 305
15.8 Electrodialysis 306
15.9 Ion Exchange 308
References Section 15 311
16. Separation of Ammonia, Carbon Dioxide and Hydrogen Sulfide 314
16.1 Introduction and Results 314
16.2 Separation by Distillation; Calculation of Number of
Theoretical Trays 318
16.3 Vapor-liquid Equilibrium for NH3-C02^H20 324
16.4 Vapor-liquid Equilibrium for NH3-H2S-H20 327
16.5 Deacidification Columns for NH3-C02-H20 330
16.6 Ammonia Concentration 333
16.7 Phosam-W Process 334
16.8 Equipment Size and Capital Cost 337
16.9 Operating Cost and Energy 337
References Section 16 342
(continued)
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CONTENTS (continued)
17. Solvent Extraction of Phenol 343
17.1 Introduction and Summary of Results 343
17.2 Simple Extraction Equation 347
17.3 Base Case Example 350
17.4 Optimization 353
17.5 Selective Adsorption of Phenol 355
References Section 17 360
18. Biological Treatment 361
18.1 Introduction 361
18.2 Equalization 363
18.3 Air Activated Sludge (AAS) System—Scaled Plant 364
18.4 Theoretical Design of Air Activated Sludge Systems 373
18.5 Air Activated Sludge—Nitrification-Denitrification
(AASND) System 381
18.6 High Purity Oxygen Activated Sludge (HPOAS) System 386
18.7 Activated Trickling Filter—High Purity Oxygen
Activated Sludge (ATF-HPOAS) System 396
18.8 Anaerobic Biological Treatment 406
18.9 Additional Considerations and Research Needs 409
References Section 18. 414
19. Site. Studies - 2: Water Treatment Plants 417
19.1 Introduction and Conclusions 417
19.2 Plants in North Dakota 421
19.3 Plants in New Mexico 440
19.4 Plants in Wyoming 447
References Section 19. . . 461
Appendix 1: Analyses of Wastewater Samples. . 462
vi
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FIGURES
Number Page
2-1 Water consumption 6
2-2 Approximate cost and energy requirements of water treatment. . . 7
3-1 Schemes for reuse of foul water 9
4-1 Water streams in a plant to produce pipeline gas from coal ... 12
4-2 Water streams in a plant to produce power gas from coal 14
4-3 Lurgi pressure gasifier 16
4-4 Bigas coal-water slurry preparation system 22
4-5 Design features of a Bigas process reactor 23
4-6 Bigas gasifier 24
4-7 CC>2-Acceptor gasifier block diagram 28
4-8 Burner-gasifier feed and circulation system for Agglomerating
Burner gasification process 32
4-9 Winkler gasifier and heat recovery system 34
4-10 Stirred, pressurized, gas producer 37
4-11 Molten Salt Process; gasifier and combustor in a single vessel . 40
4-12 Flow diagram for Molten Salt gasification section using
two vessels 41
4-13 Flow diagram of an ash removal section 43
4-14 Koppers-Totzek gasifier and heat recovery system 46
4-15 Koppers-Totzek gasification process 48
4-16 H-Coal ebullated-bed reactor 52
4-17 Flow diagram for process water streams in H-Coal process
for production of 50,000 bbl/day of product oils 54
4-18 Block flow diagram for H-Coal process . 55
4-19 Synthoil pilot plant 58
4-20 Flow diagram for process water streams in Synthoil process ... 59
4-21 Flow diagram for hydrogen production in Synthoil process .... 60
5-1 IGT Hygas pilot plant hydrogasification reactor 70
5-2 Flow diagram for Hygas process 75
6-1 Flow diagram for Synthane gasifier 90
6-2 Flow diagram for Synthane process 95
7-1 SRC dissolving section—A 110
7-2 SRC dissolving section—B Ill
7-3 SRC hydrogen production by gasification of filter residue—A . . 118
7-4 SRC hydrogen production by gasification of filter residue—B . . 119
7-5 SRC hydrogen production by reforming 123
9-1 Gasification and combined cycle generation with cold gas
desulfurization 156
(continued)
vii
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FIGURES (continued)
Number Page
9-2 Gasification and combined cycle generation with hot gas
desulfurization 157
9-3 Generating plant details for heat balance 163
9-4 Total annual evaluated costs of wet/dry cooling system as a
percentage of evaporative loss of all evaporative cooling
in New Mexico 175
10-1 Cost of cooling a high pressure methane stream 185
10-2 Cost of cooling a low pressure gas stream 186
10-3 Schematic of wet cooling tower 188
10-4 Turbine heat rates 192
10-5 Turbine condenser cooling loads 193
10-6 Turbine condenser cooling systems 195
10-7 Fan power reduction factor for air coolers 197
10-8 Annual cost of steam turbine condenser cooling at Casper,
Wyoming 207
10-9 Annual cost of steam turbine condenser cooling at
Farmington, New Mexico 210
10-10 Annual cost of steam turbine condenser cooling at Beulah,
North Dakota 211
10-11 The effect of water cost on water consumed for cooling
turbine condensers 212
10-12 Compressor interstage cooling example 213
10-13 Operating cost of air compressor interstage cooling in
New Mexico 221
10-14 Operating cost of air compressor interstage cooling in Wyoming . 222
14-1 Scaling by calcium carbonate 278
15-1 Costs of granular activated carbon adsorption 292
15-2 Cost and energy of evaporation 296
15-3 Clarifier costs 301
15-4 Costs of reverse osmosis 307
15-5 Approximate electrodialysis capital investment as a function
of capacity for various numbers of stages 309
16-1 Phosam-W process for ammonia separation 316
16-2 All-distillation process for ammonia separation 317
16-3 Comparative capital costs for ammonia separation 319
16-4 Operating costs for ammonia separation 320
16-5 Distillation tower nomenclature 321
16-6 Ammonia tower 335
17-1 Phenosolvan process 344
17-2 Idealized solvent extraction 344
17-3 Ideal counter-current liquid-liquid extraction 348
17-4 Extraction equation 351
17-5 Sample optimization 354
17-6 Phenol recovery by adsorption with acetone regeneration 357
17-7 Phenol recovery by adsorption with gasoline regeneration .... 358
18-1 Air activated sludge system (Scheme 1) 370
18-2 Air activated sludge system (Scheme 2) 371
(continued)
viii
-------
FIGURES (continued)
Number
Page
18-3 Typical configuration of an aeration basin for air activated
sludge systems 372
18-4 Capital cost of vacuum filtration and dissolved air flotation. . 375
18-5 Air activated sludge—nitrification-denitrification system . . . 384
18-6 High purity oxygen activated sludge (HPOAS) system 390
18-7 Configuration of step 1 of HPOAS system 393
18-8 Configuration of step 2 of HPOAS system 394
18-9 Activated trickling filter—high purity oxygen activated
sludge system 398
18-10 Schematic anaerobic treatment systems 410
19-1 Water treatment plant block diagram for Hygas process at
North Dakota (Scheme 1) 422
19-2 Water treatment plant block diagram for Hygas process at
North Dakota (Scheme 2) 423
19-3 Boiler feed water treatment schemes in North Dakota 424
19-4 Water treatment plant block diagram for electric generation
at North Dakota (Scheme 1) 433
19-5 Water treatment plant block diagram for electric generation
at North Dakota (Scheme 2) 434
19-6 Water treatment plant block diagram for SRC at North Dakota. . . 438
19-7 Use of untreated brackish makeup to cooling tower of power
plant at New Mexico. 441
19-8 Water treatment plant block diagram for electric generation
at New Mexico 442
19-9 Water treatment plant block diagram for Hygas process at
New Mexico 446
19-10 Water treatment plant block diagram for SRC at New Mexico. . . . 448
19-11 Water treatment plant block diagram for electric generation
at Wyoming 450
19-12 Water treatment plant block diagram for Hygas process at
Wyoming 452
19-13 Water treatment plant block diagram for Synthane process
at Wyoming (Scheme 1, not used) 454
19-14 Water treatment plant block diagram for Synthane process
at Wyoming (Scheme 2, as used) 455
19-15 Preliminary water treatment plant block diagram for Solvent
Refined Coal at Wyoming 458
19-16 Water treatment plant block diagram for SRC at Wyoming 459
ix
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TABLES
Number
Page
1-1 Summary of Conditions for, and Results of, Detailed Studies. . . 3
4-1 Lurgi Process Designs 18
4-2 Water Equivalent Hydrogen Balances for Lurgi Processes 19
4-3 Water Equivalent Hydrogen Balances for Bigas Process 26
4-4 Water Equivalent Hydrogen Balance for the C02-Acceptor Process . 30
4-5 Water Equivalent Hydrogen Balances for the Winkler Process ... 36
4-6 Typical Data for a Stirred-bed Gasifier Fed New Mexico
Subbituminous Coal 38
4-7 Water Equivalent Hydrogen Balance for the Stirred Fixed Bed
Process 39
4-8 Water Equivalent Hydrogen Balance for the Molten Salt Process. . 44
4-9 Water Equivalent Hydrogen Balance for the Koppers-Totzek
Process 49
4-10 Water Equivalent Hydrogen Balance for the H-Coal Process - 1 . . 53
4-11 Water Equivalent Hydrogen Balance for the H-Coal Process - 2 . . 56
4-12 Approximate Coal Analyses for Synthoil Process 61
4-13 Water Equivalent Hydrogen Balance for the Synthoil Process ... 62
5-1 Water Equivalent Hydrogen Balance for Hygas Process Using
Wyoming Subbituminous Coal 72
5-2 Ultimate Disposition of Unrecovered Heat in Hygas Plant Using
Wyoming Coal 73
5-3 Coal and Ash Residue for Hygas Plant at Wyoming Site 74
5-4 Stream Compositions for Hygas Plant Using Wyoming Coal 76
5-5 Hygas Gasifier Heat Balance Using Wyoming Coal 78
5-6 Hygas Production Train Heat Balance Using Wyoming Coal 79
5-7 Hygas Plant Using Wyoming Coal: Driving Energy 80
5-8 Hygas Plant Using Wyoming Coal: Overall Heat Balance 81
5-9 Approximate Analyses of New Mexico Subbituminous Coal and
North Dakota Lignite 83
5-10 Process Coal and Water Streams, and Plant Driving Energy
Requirements at New Mexico and North Dakota 84
5-11 Approximate Thermal Efficiency and Ultimate Disposition of
Unrecovered Heat at New Mexico and North Dakota 86
6-1 Water Equivalent Hydrogen Balance for Synthane Plant 92
6-2 Ultimate Disposition of Unrecovered Heat 93
6-3 Coal for Synthane Example 94
6-4 Stream Compositions for Synthane Plant 96
6-5 Synthane Gasifier Heat Balance 97
(continued)
-------
TABLES (continued)
Number Page
6-6 Char Analysis and Flue Gas Desulfurization Water 97
6-7 Synthane Gas Production Train Heat Balance 98
6-8 Synthane Plant Driving Energy . 99
6-9 Synthane Plant Overall Heat Balance 101
7-1 Total Plant Process Water Streams 106
7-2 Ultimate Disposition of Unrecovered Heat 107
7-3 Analyses of Coal and Solvent Refined Coal 108
7-4 Overall Material Balance for Dissolving Section of the
Plant/ Wyoming Coal 113
7-5 Overall Material Balance for Dissolving Section of the
Plant, New Mexico Coal 114
7-6 Overall Material Balance for Dissolving Section of the
Plant, North Dakota Lignite 115
7-7 Streams in the Production of Hydrogen by Gasification
of Filter Residue ? . . 120
7-8 Approximate Heat Balance for Production of Hydrogen by
Gasification of Filter Residue 121
7-9 Hydrogen Production by Reforming 124
7-10 Heat Balance on Reforming Section 125
7-11 Summary of Reforming Section 126
7-12 Approximate Heat Balance on Dissolver Section 128
7-13 Plant Driving Energy 130
7-14 Plant Fuel Requirements 131
7-15 Plant Conversion Efficiency Calculation 132
8-1 Moles Water/Mole Dry Gas at Saturation 139
8-2 Determination of Flue Gas Volume 141
8-3 Calculated Makeup Water for Flue Gas Saturation 142
8-4 Calculated Makeup Water for Waste Disposal and Total Water. . . . 142
8-5a Weight of Components of Lime Sludge (Dry) and Corresponding
Weights of Sulfur and Water of Hydration 144
8-5b Weight of Components of Limestone Sludge (Dry) and
Corresponding Weights of Sulfur and Water of Hydration 144
8-6 Reported and Estimated Water Requirement for FGD 145
9-1 Coals Used in Analysis of Combined-cycle Plants 151
9-2 Combined-cycle Electrical Plant Water Equivalent Hydrogen
Balances 152
9-3 Combined-cycle Electrical Plant Overall Energy Balances 153
9-4 Comparison of Assumed Energy Balance for a 1,000 Megawatt
Steam-electric Coal Fired Generating Plant, with a
Combined Cycle Generating Plant .... 154
9-5 Comparison of Water for Flue Gas Desulfurization in a
1,000 Megawatt Coal Fired Steam-electric Generating Plant,
with Net Water for a Combined Cycle Plant 154
9-6 Gasifier Mass Balance Using Wyoming Coal 161
9-7 Heat Duties at the Various Points in the Plant 164
9-8 Plant Energy Balances 166
(continued)
XI
-------
TABLES (continued)
Number Page
9-9 Plant Efficiencies 167
9-10 Efficiencies of Various Plant Combinations 168
9-11 Summary of Design Conditions for Optimized Cooling Systems
for Coal Fired Generating Plants 170
9-12 Summary of Annual Evaluated Costs for Optimized Cooling Systems . 173
9-13 Summary of Comparison of Wet and Dry Cooling in Coal Fired
Generating Plants ..... 174
10-1 Representative Formulae and Data Used for Calculating Cost
of Wet and Dry Cooling 182
10-2 Annual Average Water Consumption for Wet Cooling 189
10-3 Water Consumption Rate per Shaft kw for an All Dry System
at Casper, Wyoming 201
10-4 Water Consumption Rate per Shaft kw for an All Wet System
at Casper, Wyoming 203
10-5 Water Consumption Rate per Shaft kw at Casper, Wyoming with
50% Wet Load at Peak Design Condition 205
10-6 Equipment Size and Annual Power Requirement per Shaft kw for
Turbine Condenser Cooling at Casper, Wyoming 206
10-7 Annual Operating Cost for Turbine Condenser Cooling at
Casper, Wyoming 206
10-8 Monthly Average Temperature of Beulah, North Dakota and
Farmington, New Mexico 209
10-9 Basis and Design Conditions for Compressor Interstage Cooling
(All Wet or All Dry) 215
10-10 Summary of Compressor All Wet and All Dry Interstage Cooling
Results in New Mexico 216
10-11 Summary of Compressor All Wet and All Dry Interstage Cooling
Results in Wyoming. 217
10-12 Design Conditions for Compressor Dry-Followed-By-Wet
Interstage Cooling in New Mexico 218
10-13 Design Conditions for Compressor Dry-Followed-By-Wet
Interstage Cooling in Wyoming 218
10-14 Summary of Compressor Dry-Followed-By-Wet Interstage
Cooling Results in New Mexico 219
10-15 Summary of Compressor Dry-Followed-By-Wet Interstage
Cooling Results in Wyoming 220
11-1 Calculation of Water for Mine Complex in Wyoming 225
11-2 Calculation of Water for Mine Complex in North Dakota 225
11-3 Calculation of Water for Mine Complex in New Mexico 226
12-1 Calculation of Additional Water in Wyoming 233
12-2 Calculation of Additional Water in North Dakota • 234
12-3 Calculation of Additional Water in New Mexico 235
13-1 Approximate Water Requirements: --Electric, North Dakota .... 243
13-2 —Hygas-SNG, North Dakota 244
13-3 —Solvent Refined Coal, North Dakota 245
13-4 —Electric, New Mexico 246
(continued)
xii
-------
TABLES (continued)
Number Page
13-5 Approximate Water Requirements: —Hygas-SNG, New Mexico 247
13-6 —Solvent Refined Coal, New Mexico 248
13-7 —Electric, Wyoming 249
13-8 —Hygas-SNG, Wyoming 250
13-9 —Synthane-SNG, Wyoming 251
13-10 —Solvent Refined Coal, Wyoming 252
14-1 Analysis of Water from Lake Sakakawea, North Dakota 254
14-2 Analysis of Brackish Groundwater Near Gallup, New Mexico 254
14-3 Analysis of Sewage at Wyoming Site 255
14-4 Analysis of Water from Yellowstone River Near Hardin, Montana . ., 256
14-5 Exemplary Analyses of Foul Water from SRC and from Gas Plants . ./ 257
14-6 Analysis of Foul Process Condensate, Western Coals 260
14-7 Contaminants in Product Water from Coal Gasification 263
14-8 Contaminants in Product Water from Coal Low-Voltage Mass
Spectrometer Data, ppm by Weight 264
14-9 Analysis of Foul Process Condensate, Central and Eastern Coals. . 265
14-10 Analysis of Foul Process Condensate, Solvent Refined Coal .... 267
14-11 Trace Elements in Condensate from an Illinois No. 6 Coal
Synthane Gasification Test 270
14-12 Analysis of Water from Koppers Coal Gasification,
Kutahya, Turkey 272
14-13 Exemplary Analysis of Condensed Stripping Steam from Acid
Gas Removal in the SRC Plant 272
14-14 Suggested Tolerances in Boiler Water 273
14-15 Control Limits for Cooling Tower Circulating Water Composition. . 275
14-16 Solubility of CaCC-3 at 50°C and 800 ppm TDS 276
14-17 Calcium Phosphate Concentrations at Various pH Values 280
15-1 Analyses - Hygas Wastewater Wet Oxidations 287
15-2 Summary of Costs of Carbon Adsorption 289
15-3 Chemical Costs 297
15-4 Solubility of Magnesium at Different pH 299
16-1 Deacidification Tower for NH3-C02 331
16-2 Costs for Ammonia Separation by All-distillation Process 338
16-3 Size and Cost of Phosam-W Process in a Gas Plant 339
16-4 Size and Cost of Phosam-W Process in SRC Plant 340
16-5 Operating Costs for Ammonia Separation in Gas Plants 341
17-1 Phenol Recovery 346
17-2 Approximate Minimum Annual Cost of Solvent Extraction for an
Ideal Case 355
17-3 Costs of Phenol Recovery by Various Methods 359
18-1 Characteristics of Exemplary Coal Conversion Wastewater 363
18-2 Typical Analysis of Ammoniacal Liquor and Still Waste 365
18-3 Typical Design Criteria for Biological Treatment of
Coke Plant Wastes 367
18-4 Cost Estimation Procedure for Air Activated Sludge Plant 374
18-5 Costs of Air Activated Sludge System (Scheme 1) 376
(continued)
xiii
-------
TABLES (continued)
Number Page
18-6 Costs of Air Activated Sludge System (Scheme 2) 377
18-7 Values of Fundamental Coefficients Used in the Monod Model
Evaluation 379
18-8 Volume of Aeration Basin Based on Monod Model 380
18-9 Concentrations of Nitrogenous Species in Nitrification and
Denitrification Processes. . . , 383
18-10 Costs of Air Activated Sludge—Nitrification-Denitrification
System 385
18-11 Design of the HPOAS System 388
18-12 Cost of High Purity Oxygen Activated Sludge (HPOAS) System ... 395
18-13 Cost of Oxygen . 396
18-14 Preliminary Design Calculation of Activated Trickling Filters. . 402
18-15 Costs of Activated Trickling Filter - High Purity Oxygen
Activated Sludge (ATF-HPOAS) System 405
18-16 Unit Costs of Biological Treatment Using ATF-HPOAS System. . . . 406
19-1 Approximate Cost and Energy Requirements for Total Plant
Water Treatment 418
19-2 Assumed Analysis of Biotreatment Effluent Water 421
19-3 Analysis of Cooling Tower Streams for Hygas in North Dakota,
Scheme 1 425
19-4 Cost of Cooling Water Chemicals for Hygas in North Dakota,
Scheme 1 426
19-5 Analyses of Cooling Tower Streams for Hygas in North Dakota,
Scheme 2 428
19-6 Cost of Cooling Water Treatment for Hygas in North Dakota,
Scheme 2 429
19-7 Boiler Feed Water Compositions in North Dakota 43°
19-8 Approximate Influent to and Effluent from Lime-Soda-Silica
Treatment in SRC Plant at North Dakota 439
19-9 Compositions at Points in the Boiler Feed Water Treatment in
New Mexico 444
19-10 Cooling Water Makeup to the Electric Power Plant in Wyoming. . . 449
xiv
-------
CONVERSION OF
LENGTH
AREA
VOLUME
MASS
WEIGHT RATE OF FLOW
VOL. RATE OF FLOW
ENERGY
POWER
SPECIFIC ENERGY
PRESSURE
WATER FOR ENERGY
HEAT RATE
AMERICAN TO INTERNATIONAL SYSTEM (SI)
To convert from
ft
ft2
acres
ft3
gallons
Ib
tons
103 Ib/hr
tons/day
gallohs/min
gallons /min
10 gallons/day
Btu
kw-hrs
H.P.
kw
106 Btu/hr
Btu/lb
lb/in2
gal/106 Btu
Btu/kw-hr
to
meters
meters
2
meters
meters
meters
kilograms
megagrams
kg/sec
kg/sec
3
meters /sec
3
millimeters /sec
3
meters /sec
kiloj oule
(» Newton x meter)
megajoules
Joules/sec
Joules /sec
kilo joules/sec
kiloj oules/kg
kilopascall _
(«= kilonewton/m )
m /mega joule
Joules/kw-sec
UNITS
multiply by
0.305
0.0929
4047
0,0283
0,00379
0.454
0,907
0.126
0.0105
6.309 x 10~5
6309
0.0438
1.055
3.60
746
1000
293
2.324
6.895
3.592 x 10"6
0.293
TEMPERATURE
HEAT TRANSFER
COEFFICIENT
/•J
Btu/hr.ft .F
Joules/sec.m . K
0.556 ( F + 459.7)
5.674
*Standard for Metric Practice, American Society for Testing and Materials,
E380-76, 1976.
xv
-------
SECTION 1
INTRODUCTION
1.1 PURPOSE OF THE STUDY
The purpose of this study is to develop strategies and recommend
measures to minimize water pollution and consumption by coal conversion
complexes. An important part of environmental assessment of large plants
to convert coal is the determination of the quantity of water consumed. In
this report the determination of water quantity is discussed in great detail
and water treatment is described in enough detail to show that full recycle
of effluent waters is possible.
1.2 ORGANIZATION OF THE REPORT
The report is divided into two parts entitled "Water Quantities" and
"Water Quality."
1.3 WATER QUANTITIES
In the first part of the work, water consumption is studied. Brief
descriptions and process water balances are given for a wide variety of
coal conversion processes. The Lurgi, Bigas, CO -Acceptor and Agglomer-
ating Burner processes convert coal to pipeline gas, also called "substi-
tute natural gas" (SNG) which is more than 90 percent methane and is a
direct substitute for natural gas in all applications.
The Winkler, Stirred-Bed, Molten Salt and Koppers-Totzek processes
convert coal to power gas, also called fuel gas, low-Btu.gas or
medium-Btu gas. This gas has a higher heating value in the range 150 Btu/scf
-------
(for low-Btu gas) to 350 Btu/scf (for medium-Btu gas) compared to more than
920 Btu/scf for pipeline gas. These heating values are too low for the gas
to bear the cost of long-distance transmission in pipelines and are made for
local use. Natural gas burners can use a medium-Btu gas, but modification or
down-rating is needed for the lowest-Btu gases. Low-Btu gas is made using
air as the oxidant. Medium-Btu gas requires oxygen to avoid diluting the gas
with nitrogen and the investment is higher.
H-Coal and Synthoil processes convert coal to a liquid fuel which may be
an ash-free heavy fuel oil sufficiently free of nitrogen and sulfur or, with
process changes, may be refinery feed stock. These are processes in which
coal is hydrogenated directly with hydrogen, and the plant complex also has a
hydrogen production section in which hydrogen is made by reducing water with
coal (gasification of coal with oxygen and steam). Hydrogenation always
converts some of the coal to light hydrocarbon gas and liquid which may be
sold or burnt as plant fuel.
To determine the quantity of water consumed for any conversion process
at any site, one must total the water consumed for the following reasons:
Net process water
Water for cooling
Water for mining and land reclamation
Water lost in evaporation, water lost with solids and other uses.
Four processes are described in more detail so that cooling water
requirements can be determined; these are the Hygas and Synthane processes
for pipeline gas, the Lurgi process to make low-Btu fuel gas which is burnt
on site to generate electricity by the combined cycle method and Solvent
Refined Coal (SRC). SRC is a low ash, low nitrogen, low sulfur fuel which has
a fusion temperature in the range of 350°-450°F.
The choice and economics of evaporative and dry cooling is studied in
detail. At three sites (North Dakota, New Mexico and Wyoming) the water
needs for mining, coal handling, etc., are analyzed; these quantities are
generally dependent on the site, not on the process.
The conditions and results of the detailed studies are summarized
on Table 1-1.
-------
TABLE 1-1. SUMMARY OF COBDITIOHS TOR, ASP RESULTS OF, DETAILED STODIES
Plant Capacity:
Major fuel
Byproduct
Coal analysis as received
C
H
N
S
0
Ash t
Moisture t
10 Btu/hr
Overall conversion
efficiency, %
Set water consumed {see Sec-
tion 13 for details!
10 gals/day
Ar-prcxiaate energy consuaed
for water treatment (see
Section 19 for details)
4 of product energy
Approximate cost for water
treatment (see Section 19
for details)
S/106 Btu product fuel,
or 4/kw-hr power
Kygas
Wyoming
250 scf/day
10.1x10* Btu/hr
0
54.2
4.0
0.8
0.6
14.5
6.0
19.9
1.52
14.2
Hygas
New Mexico
2SO Ecf/day
lO.lxlO9 Btu/hr
0
53. «
3.9
0.9
0.7
11.1
13.5
16.3
1.52
14.1
Hygas
North Dakota
250 scf/day
lO.lxlo' Btu/hr
0
42.5
2.9
0.7
0.6
13.4
5.1
34.8
2.11
14.7
Synthane
Wyoming
250 scf/day
9.8X109 Btu/hr
1.6x10* Btu/hr
54.0
3.5
0.8
0.6
14.1
7.1
19.9
1.92
17.1
SBC
Wyoming
10,000 tons/day
13.3X109 Btu/hr
l.OxlO9 Btu/hr
54.2
4.0
0.8
0.6
14.5
6.0
19.9
1.93
17.9
SRC
New Mexico
10,000 tons/day
13.3X109 Btu/hr
l.SxlO9 Btu/hr
53.5
3.9
0.9
0.7
11.1
13.5
16.3
1.95
18.3
SRC
North Dakota
10,000 tons/day
13.3X109 Btu/hr
0
45.2
3.1
o.a
0.7
14.2
5.4
30.6
2.41
17.8
Electric
Power
Wyoming
1000 Hue
3.4x10* Btu/hr
1.9x10* Btu/hr
52.0
3.8
0.8
0.6
13.9
5.7
23.2
1.31
12.0
Electric
Power
New Mexico
1000 MHe
3.4x10* Btu/hr
1.7x10* Btu/hr
53.6
3.9
0.9
0.8
11.0
13.5
16.3
1.23
12.0
Electric
Power
North Dakota
1000 MHe
3.4x10* Btu/hr
2x10* Btu/hr
42.5
2.9
0.7
0.6
13.7
5.1
34.8
1.66
12.2
•Byproduct fuel used to generate electricity at 10,000 Btu/kw-hr.
-------
The consumptive water requirements of Part 1 have been calculated
on the assumption that no liquid water leaves the plant. Any water
clean enough to be sent to a receiving stream is clean enough for use in
the plant, so the assumption of no liquid water effluent is economically,
as well as environmentally, correct.
1.4 WATER QUALITY AND TREATMENT
In the second part of the report, water treatment for source water,
for recycle water and for reuse water is discussed.
Three sites have been chosen for detailed study: 1) in the vicinity
of Beulah, North Dakota using lignite coal and water from Lake Sakakawea.
This site is regarded as a standard, or basic, example in that water of
good quality is assumed to be cheap and available. 2) The four corners
area of New Mexico using Navajo subbituminous coal and a brackish ground
water. This site is used to study the effect of a brackish, but plentiful,
water supply. 3) In the vicinity of Gillette, Wyoming using subbitum-
inous coal. It is at this site that the water supply for a satellite
town has been integrated into the coal conversion plant. Town sewage is
taken as the main water supply to the plants, and any other water taken
in is assumed to be very expensive because it must be brought from a
long way away.
-------
SECTION 2
CONCLUSIONS
2.1 WATER CONSUMPTION
Total water consumption depends on the product, on the site and,
possibly to a lesser extent, on the process. Details are tabulated in
Section 13 and summarized on Table 1-1 and Figure 2-1.
2.2 WATER TREATMENT
Water treatment costs and energy requirements are given in Figure 2-2
and Tables 1-1 and 19-1. They are much less well defined than water
consumption quantities, and more study is required. In deriving Figure 2-2,
unproven technologies were used; indeed, unproven technologies had to
be assumed because an adequate experimental base is not available for most
water treatment technologies on these waters.
-------
KEY
FLUE GAS DESULFURIZATION
Y/A DUST CONTROL a ASH DISPOSAL
[ | COOLING
NET PROCESS
gol/min
6000
N.D.
N.M. WYO. N.D. N.M. WYO. WYO.
N.O.
N.M. WYO.
POWER
HYGAS
SYNTHANE
— 4000
2000
SRC
Figure 2-1. Water consumption. (Plant sizes given on Table 1-1.)
-------
I -
ENERGY FOR WATER TREATMENT
( % OF PRODUCT ENERGY )
APPROXIMATE COST OF WATER TREATMENT
Btu PRODUCT FUEL OR 4/kw GENERATED)
N.D.| N.M. |WYO.
POWER PLANTS
IWYO.I
SYNTHANE
N.D. | N.M. IWYO.
SOLVENT REFINED
COAL
Figure 2-2. Approximate cost and energy requirements of water treatment.
(Plant sizes given on Table 1-1.)
-------
SECTION 3
RECOMMENDATIONS
In estimating net water consumptions this study provides a method
as well as some answers. We recommend that this method be used when
extending the determinations to other processes at other sites. There is
a need to know regional total water consumptions when planning water uses,
particularly in the Western States, where water is scarce. When totaling
the consumptions from many plants to find the regional consumption,
considerable accuracy is required for each plant if the absolute error in
the total is not to be so large that the total is useless. Attention to
detail is necessary to attain accuracy.
The quality of the plant effluent waters, and the response of these
waters to treatment is not well understood, and properly designed experi-
ments are needed. In designing treatment experiments it is necessary that
real wastewater be used. Analyses available are insufficient to enable
the laboratory to produce a simulated wastewater adequately. Furthermore,
when a wastewater is to pass through several treatments in series so that
the effluent from one treatment is the influent to the next, then all of
the treatments must be simultaneously studied. Experimentation on one
treatment, isolated from others, is unlikely to be good enough because the
composition of the feed water to the isolated treatment is unlikely to
be correct. Figure 3-1 shows four simple schemes for reuse of foul process
condensate. The cost and procedure for reuse, whether in th« cooling
tower or in the boiler, will depend on the treatment and cannot be
investigated without complete treatment. The design of biological treat-
ment depends on the degree, if any, of ammonia separation and solvent
extraction that preceded it; so does the effluent from the biological
treatment depend on preceding treatments. Separation of ammonia'is the
8
-------
SOLVENT
••
EXTRACTION
SEPARATION
OF AMMONIA
1
USED FOR ASH
SALE OF SOLVENT
>-
PHENOLS
SALE OF
AMMONIA
DISPOSAL
AND COOLING TOWER MAKEUP
SEPARATION
i
OF AMMONIA
SALE OF
AMMONIA
BIOLOGICAL OXIDATION
OF PHENOLS AND OTHER
ORGANICS
1
USED FOR ASH
AND COOLING
DISPOSAL
TOWER MAKEUP
w
EXTRACTION
1
SEPARATION
OF AMMONIA ^
SALE OF
PHENOLS
SALE OF
AMMONIA
BIOLOGICAL REMOVAL
OF ORGANICS
\
USED FOR ASH DISPOSAL
AND COOLING TOWER
SOLVENT ^
EXTRACTION
SEPARATION
__^^te
OF AMMONIA
2FFECT
EVAPORATION
FINAL REMOVAL
OF ORGANICS
*
BOILER FEED WATER
MAKEUP
SALE OF
PHENOLS
SALE OF
AMMONIA
Figure 3-1. Schemes for reuse of foul water.
-------
most massive energy-consuming part of the water treatment plant. This energy
consumption will depend to some degree on whether phenol has been extracted
first; some phenol is always stripped in ammonia separation if it is not
extracted first. Solvent extraction is expensive, but sale of extracted
phenols can offset some of the cost. The quality, marketability and value
of extracted phenols is not known.
These are examples to explain why experiments should be done on complete
water treatments, not on individual aspects.
10
-------
PART 1 - WATER QUANTITIES
SECTION 4
COAL CONVERSION PROCESSES
4.1 INTRODUCTION
In this chapter the quantities of water fed to, and evolved from, various
processes to convert coal to other fuels are surveyed where possible. In
doing this the literature has been reviewed; new designs have not been made.
These water requirements, which we have called process requirements, are not
the largest requirements of the plant. However, knowledge of these quantities
is important in the design of a water treatment plant because the quality of
process water is extreme. Most of the water is fed as steam and so must be of
the highest purity. Most of the water evolved is a condensate that has con-
tacted coal, tar, and gas and is particularly dirty.
A diagram of particular plants to make pipeline gas is given for the
Hygas/Oxygen process (Figure 5-2) and for the Synthane process (Figure 6-2).
Both are described in detail and will be used for the site studies. A simpler
and more general diagram showing points of water inflow and outflow is given
in Figure 4-1 and used for less detailed descriptions of the Lurgi, Bigas,
and C0_-Acceptor. The Agglomerating Ash process is also described, but there
is not enough information to define the water streams.
Pipeline gas is more than 90 percent methane (CH.). Although some of the
hydrogen for the methane is available in combined form in the coal, most of
the hydrogen comes from water which is reduced by carbon. A methane output of
250 x 10 standard cubic feet per day contains hydrogen equivalent to
987,000 Ib/hr of water (1,976 gpm). In fact, water in excess of that needed
for chemical reaction is usually added to the gasifiers to control the
11
-------
cool
moisture
steam
water vapor
In off- gat
t
COAL PREPARATION
a PRETREATMENT
OXYGEN PLANT
HYDROGEN PLANT
INDIRECT HEAT
n
i
i
t
.t.
WA'
1
1
1
1
moisture
in coal
steam or steam or
water water
SCRUBBI
GASIFIER *• &
COOLI
tt
t
1
ewn
NG SHIFT
NG *" CONVERTER
**
process
condensate
ALTERNATIVES
wash water
or steam
PURIFICATION
1
condensate
METHANATION
1
product
pipeline
gas
condensate
Figure 4-1. Water streams in a plant to produce pipeline gas from coal.
-------
temperature and moderate the reaction.
The surplus water comes out of the process in several places. Some water
is condensed out of the raw gas leaving the gasifier. This water has been
called process condensate. In some processes the raw gas is scrubbed with
water. This cools the gas and transfers tar, dust, and soluble organic mater-
ials from the gas to the water. Cooling also causes condensation of water
vapor which left the gasifier. The scrubbing water is circulated and only the
surplus water from condensation is removed. The quantity called process con-
densate is the surplus water actually removed from the process; it is not the
quantity of water circulated within the process.
Although many coals contain a high percentage of moisture, this water
is not usually available to enter into chemical reaction. Often the coal must
be dried before it is fed to the gasifier. Even if coal is not predried, it
is usually fed near the top of the reactor where water vaporizes and leaves
before it has time to react. A wet coal will usually increase the amount of
process condensate but not reduce the steam requirement to the gasifier.
Sometimes additional cooling causes water to condense in the gas purifi-
cation liquid. Such water must then be boiled out so that the gas purifica-
tion liquid can be recycled. The quantity of this water is shown in the
tables as leaving the purification stage.
Water is formed during the methanation stage and is condensed and knocked
out of the product gas. This water is shown in the tables as leaving the
methanator.
Figure 4-2 shows the stages in a plant to produce power gas (low-Btu gas)
from coal. The processes described are Winkler, Stirred Fixed Bed, Molten
Carbonate, Lurgi, and Koppers-Totzek. The Lurgi process is the process chosen
for combined-cycle electric power production and is described in more detail
in that section. The Koppers-Totzek process is further described as part of
the design of the Solvent Refined Coal plant as it is the process chosen to
produce hydrogen from the carbonaceous filter residue.
Processes to make clean liquid and solid fuels cannot be shown on a
single block diagram. Brief descriptions of the Synthoil and H-Coal processes
are given.
13
-------
moisture in
prepared coil
steam
& water
1
GASIFIER
^
SCRUB
COOLl
CHAR Rl
\
wate
ash £
steam
1
BING
NG &
EMOVAL
PURIFIC
, fc. c
COMPR
• \
r with conde
k char
:ATION
*
ESSION
r
nsate
product
^- power
gas
Figure 4-2. Water streams in a plant to produce power gas from coal,
-------
4.2 LURGI PROCESS
Brief descriptions of the Lurgi and many other processes will be found
in Reference 1. The Lurgi commercial plant operation experience is given in
References 2-5. At least four full-sized commercial plants have been des-
cribed: the Wesco plant using Navajo coal in the "Four Corners" area of New
Mexico; the El Paso plant in the same region; the Wyoming-Rochelle
plant in Campbell and Converse counties Wyoming, and the Michigan-Wisconsin
plant using lignite in Mercer county, North Dakota.
The Lurgi gasifier shown in Figure 4-3 consists of a number of chambers
stacked vertically. From top to bottom they are: coal bunker, coal lock
chamber, water jacketed gas producer chamber, ash lock chamber and ash quench
chamber. Coal crushed and screened to 1.5 to 0.19 inch flows through the lock
chamber to the pressurized reactor maintained at 350 to 450 psi. Steam and
oxygen are introduced through a revolving grate at the bottom of the reactor.
The moving bed of coal, which is the volume between the inlet and outlet
grates, has several distinct zones. Devolatization occurs at the top and gas-
ification begins lower down where the temperature reaches 1150°F and 1400°F.
The minimum residence time of coal at the desired temperature level of 1400°F
to 1600°F is about one hour. The bottom of the coal bed is the combustion
zone in which about 14 percent of the coal fed to the gasifier is burned with
oxygen to supply heat for the endothermic gasification reaction. Solid ash
is removed through the ash lock chamber at the bottom.
The crude gas leaves the gasifier at temperatures between 700°F and
1100°F, depending upon the type of coal. It contains carbonization products
such as tar, oil, naphtha, phenols, ammonia, etc., and traces of coal and ash
dust. Quenching with circulating water moves the contaminants from the gas
into the water. A portion of the cooled crude gas goes to a shift conversion
reactor where CO and H_0 are catalytically converted to H and CO so that
the H./CO ratio of the mixed gas is adjusted to greater than 3 suitable for
the methanation step. After shift conversion the acid gases are usually
removed by the Rectisol process, although other acid gas removal processes
are feasible. Finally, the purified gas passes through the methanation stage
consisting of fixed bed reactors with pelleted reduced nickel type catalysts.
15
-------
FEED COAL
DRIVE
GRATE
DRIVE
STEAM*
OXYGEN
SCRUBBING
COOLER
•=>
GAS
x^F WATER JACKET
3
Figure 4-3. Lurgi pressure gasifier ,
16
-------
Table 4-1 compares some of the operating conditions for the four commer-
cial designs studied. For these designs hydrogen balances are given in Table
4-2. This table shows the range of variation to be expected between designs
for a chosen process. All the numbers are scaled to a production rate of
250 x 10 SCF per day. Table 4-2 shows clearly that moisture in the coal does
not enter into reaction but is evolved as condensate. Coal moisture is the
major difference among the plants shown on Table 4-2. The major difference
among the plants shown on Table 4-1 is cooling water. The El Paso and Wesco
plants, while both on the same site and using the same coal, differ. For El
Paso, coal is gasified to produce a low sulfur fuel gas to provide driving
energy for the plant. Many of the drive turbines are gas turbines not requir-
ing steam condensers. For Wesco, washed coal is burned to raise steam to
drive the plant. The needs for cooling water are not, therefore, comparable
in the two plants. (The needs for cooling water are discussed more fully in
Section 10.) At both plants, treated dirty condensate supplies all the cool-
ing water makeup and a lot of direct air cooling is used. Cooling water is
evaporated to remove about 26 percent of the heat in the feed coal not recov-
ered as product gas or byproducts. At the Michigan-Wisconsin plant in North
Dakota, cooling water is evaporated to remove about 46 percent of the heat in
the feed coal not recovered as product gas or byproducts. Even more water is
evaporated than is recovered as condensate. This is the designer's choice, it
is not a necessity.
4.3 BIGAS PROCESS
The Bigas process for the production of pipeline gas from coal is under
development by Bituminous Coal Research, Inc., under the sponsorship of the
Office of Coal Research and the American Gas Assocation. More detailed des-
criptions of this process can be found in References 17 to 23. Development
work has proceeded from batch-type experiments to a fully integrated 5 ton/hr
gasification pilot plant at Homer City, Pennsylvania currently under construc-
tion.
Coal will be received from the mine and wet ground until 70 percent of
it will pass through a 200 mesh screen. The coal will be then pumped to
17
-------
TABLE 4-1. LURGI PROCESS DESIGNS
All plants scaled to 250 x 10 scf/day.
GASIFIER INFORMATION
El Paso Wesco Michigan- Wyoming-
Wisconsin Rochelle
10 Ib/hr
,9
10 Btu/lb
.3
Coal feed
Coal feed
Oxygen feed 10" Ib/hr
Oxygen feed Ib/lb (carbon + hyd.)
Steam and bfw, Ib/lb (carbon + hyd.)
1683
14.6
400
0.44
1.7
1822
15.1
473
0.48
1.8
2161
14.7
426
0.46
2.4
1750
17.8
450
0.44
(1)
THERMAL EFFICIENCY, %
HHV (product gas + by-products)
HHV coal feed
67
68
65
COOLING TOWERS
Water evaporated, 10 Ib/hr
1217
880
2849
(1) Not Known
18
-------
TABLE 4-2. WATER EQUIVALENT HYDROGEN BALANCES FOR LURGI PROCESSES
All plants scaled to 250 x 10 scf/day.
Units: 103 Ib/hr.
El Paso Wesco Michigan- Wyoming-
Wisconsin Rochelle
IN
Moisture in as-received coal
Water equivalent of hydrogen in coal
(2)
Steam and boiler feed water
OUT
Dirty condensate from scrubber,
shift reactor and purification
Clean condensate from methanation
Water equivalent of hydrogen
in Naptha and other byproducts
Water equivalent of hydrogen in
product gas
273
:oal 545
1635
2453
<3> 1084
,(3) 274
206
943
2507
226
590
1991
2807
1488
311
40
960
2799
778
529
* 2226
3533
1997
% 290
88
963
3338
411
630
(1)
^ 1870
•>• 290
^ 90
^ 960
(1) Not known
(2) Clean water inflow
(3) Effluent water
(continued)
19
-------
TABLE 4-2. (continued)
Units: gals, water/10 Btu in product gas.
El Paso Wesco
IN
Moisture in coal
Water equivalent of hydrogen in coal
Steam and boiler feed water
3.3
6.6
19.7
29.6
2.7
6.9
23.3
32.9
Michigan- Wyoming-
Wisconsin Rochelle
9.2
6.2
^26.2
41.6
M.9
OUT
Dirty condensate from scrubber,
shift reactor and purification
Clean condensate from methanation
Water equivalent of hydrogen in
naptha and other byproducts
Water equivalent of hydrogen in
product gas
13.1 17.4 23.5
3.3 3.7 %3.4
2.5
0.5
1.0
11.5 11.3 11.3
30.4 32.9 39.2
20
-------
pressurized storage hoppers in a water slurry of up to 50 percent coal. The
water used for slurrying will be vaporized in the pressurized storage and may
be passed on to the shift converter. Figure 4-4 shows one conceptual design.
The Bigas reactor is a two-zone, refractory-lined and water-cooled gasi-
fier operating at approximately 1200 psig (Figures 4-5 and 4-6). The coal and
steam enter the upper section (Zone 2) of the gasifier through four upward-
directed, concentric injection nozzles which produce an entrained flow. The
feed coal is heated very rapidly by the hot synthesis gas and steam rising
from Zone 1 and in a matter of a few milliseconds is devolatilized to produce
methane plus a highly reactive carbon char. Some 50 percent of the product
methane is produced in this stage.
The coal residence time in Zone 2 is 8 to 10 seconds with an upward gas
velocity of about 5 ft/sec. The gases with entrained char leave Zone 2 at
1800°F. In the pilot plant water will be added to the leaving gases which
will vaporize and cool the stream to about 900°F. A waste heat boiler might
be possible. The gases then pass to cyclone separators. The gases leaving
the separators are further quenched and washed before going to the shift con-
verter at about 650°F.
The char is separated and drained back to Zone 1 of the gasifier where it
reacts with steam and oxygen. In Zone 1 the temperature is 2800°F and the
carbon is gasified to synthesis gas and the ash residue forms a molten slag.
The Zone 1 coal residence time is about 2 seconds and the upward gas velocity
is about 6 ft/sec. As a result of the short residence times the capacity of
the Bigas gasifier is very much higher than that of a fixed bed gasifier. The
hot synthesis gas with entrained char passes upwards to Zone 2. The molten
slag is dropped out of Zone 1 and quenched.
Steam is added to the raw gas to adjust the steam-to-dry-gas mole ratio
to unity. The gas then enters an adiabatic shift converter having a fixed-bed
catalyst and the H /CO mole ratio is shifted from about 1:1 to about 3:1, a
ratio suitable for methanation. The hot exit gas is then cooled down through
a series of heat exchangers and process condensate, formed on cooling, is dis-
charged. Process condensate water is low in organic material because of th_
high temperature of the gasification.
The pilot plant will use the Selexol process to remove acid gases.
21
-------
CONDENSATE
NJ
NJ
GAS FROM GASIFIER
TO GASIFIER
213, 300 Mv
QUENCH WATER
Figure 4-4. Bigas coal-water slurry
preparation system
COAL FEED
HOPPER
TO GASIFIER
-------
OUTLET 1700°F
L
£
\
c
[
4
r j S
i
i
—
i i
i
•-
f~L, IT^
V
si
"^
^
COOLING WATER
-yS OUTLET
Jl
— - ZONE II
^ SUPPORT LUGS
-S
REFRACTORY
TWO COAL
^Ss__ INJECTION NOZZLES
K?-
^^~ THREE CHAR
JS^' BURNERS
^3 ^^ ZONE 1
^ COOLING WATER INLET
""""" SLAG TAP BURNER
i AND VIEW PORT
^- SLAG QUENCH ZONE
TWO SLAG
k OUTLET NOZZLES
Figure 4-5. Design features of a Bigas process reactor
23
23
-------
QUENCH WATER
COAL FEED
HOPPERS
2700 F .X
3 Y
^vj Lx'
RAW GAS
TO SCRUBBER
CYCLONE
SEPARATOR
CHAR
HOPPERS
COAL ENTRAINING
GAS
OXYGEN
SLAG QUENCH
WATER
SLAG
Figure 4-6. Bigas gasifier.
24
-------
Bituminous Coal Research is developing a high temperature fluidized bed metha-
21
nator, in which 90 percent or more of the carbon monoxide is converted in a
single pass. The final conversion of carbon monoxide will occur in a standard
fixed-bed cleanup methanator.
In Table 4-3 is presented a water equivalent hydrogen balance of the
relevant process streams for a Bigas commercial gasification plant using Ken-
tucky coal. The quench water may enter the system in several places as shown
in Figure 4-6. Piston feeding of coal used in the Air Products and Chemicals,
18
Inc. design is superseded, the thermodynamical feasibility of using a slurry
24
feed system having been established and incorporated into the pilot plant
19
designs. Table 4-3 presents a similar water equivalent hydrogen balance
using a Montana subbituminous coal-water slurry feed system design made in
1974.
4.4 C02-ACCEPTOR PROCESS
The CO2-Acceptor fluidized-bed process to convert lignite or subbitumi-
nous coal to pipeline gas is being developed by Conoco Coal Development Co.
The basic feature of this process is to provide heat for the endothermic coal
gasification reaction by reacting CO_ and CaO. The lime-bearing material
dolomite is the acceptor of the CO_-Acceptor process. Currently, a demonstra-
tion program sponsored by ERDA is under way in a 40-ton per day pilot plant
located in Rapid City, South Dakota. More information can be found in Refer-
ences 26-30.
Raw lignite is crushed to approximately 1/4 - 1/8 inch in a hot-gas-swept
impact mill and lifted with hot flue gas to the preheater. The lignite, ori-
ginally about 30 percent moisture, enters the preheater with about 5 percent
moisture. The preheater in a large plant would be fluidized with flue gas
for energy conservation. The preheater operates at atmospheric pressure and
400-500°F. Pyrolysis of the coal does not occur. In the pilot plant the
residence time in the preheater is 16 hours.
Preheated lignite is fed to the bottom of the char phase of the fluid bed
gasifier via lock hoppers (Figure 4-7). Steam is introduced and reacts with
the lignite. The gasifier and regenerator operate at 150 psig. Hot dolomite,
25
-------
TABLE 4-3. WATER EQUIVALENT HYDROGEN BALANCES FOR BIGAS PROCESS
Basis: 250 SCF/day
Units: 10 Ib/hr
Kentucky Montana
18 25
Coal Subbituminous
PROCESS INFORMATION
Coal feed to gasifier at 1.3% moisture 946 1089
HYDROGEN BALANCE
IN
Moisture in coal 12 14
Water equivalent of hydrogen in coal 428 446
Steam fed 520 692
Water to gas quench 1256
Water to slurry preparation 1089
Water to char quench 213
2216 2454
OUT
Condensate from shift conversion
and purification 996 1269
Clean condensate from methanation 242 206
Water equivalent of material removed
in purification 26 4
Water equivalent of hydrogen in
product gas 941 937
2205 2416
{continued)
26
-------
TABLE 4-3. (continued)
Units: gallons/10 Btu
IN
Kentucky Montana
Coal Subbituminous
Moisture in coal 0.1 0.2
Water equivalent of hydrogen in coal 5.2 5.4
Steam fed 6.3 8.4
Water to gas quench 15.3
Water to slurry preparation 13.2
Water to char quench 2.6
26.9 29.8
OUT
Condensate from shift conversion
and purification 12.2 15.4
Clean condensate from methanation 3.0 2.5
Water equivalent of material removed
in purification 0.3 0.1
Water equivalent of hydrogen in
product gas 11.5 11.4
27.0 29.4
27
-------
FLUE GAS
PRODUCT
GAS
LIGNITE
--AIR
STEAM
REJECT ACCEPTOR
STEAM
LIFT GAS
Figure 4-7. CO -Acceptor gasifier block diagram.
28
-------
or acceptor from the regenerator, showers through the lignite bed absorbing
CO2 which drives the water-gas shift reaction to completion so that the gas
leaving the gasifier is rich in hydrogen. The absorption of CO- also supplies
heat to maintain the gasifier at 1480-1520°F or hotter.
Acceptor accumulates at the bottom of the gasifier where it flows out.
Char (about 33 percent of the carbon in the feed coal) is allowed to flow from
the top of the gasifier. Separate exit points for acceptor and char allow a
concentration ratio of char to acceptor in the gasifier that is not the same
as the ratio of the circulation rates. In the pilot plant char is withdrawn
at 600 Ib/hr and acceptor is withdrawn at 1100 Ib/hr, but the weight ratio of
char to acceptor in the gasifier is much higher than 6:11. The residence time
of acceptor in the gasifier is less than the residence time of char. The char
residence time is sufficiently long that phenols, tars and oils are not
formed.
In the pilot plant acceptor is lifted to the regenerator by flue gas.
Air may be used if the lift line is made resistant to corrosion. Fresh accep-
tor is added and used acceptor is withdrawn at a rate required to maintain the
activity of the acceptor. Char is lifted to the regenerator with nitrogen in
the pilot plant. Flue gas may also be used. In the regenerator air is added
to burn the char and raise the temperature to 1840-1860°F. This reverses the
acceptor reaction and drives off C0_. Ash is elutriated from the regenerator
A
and recovered in cyclones from which it is released via lock hoppers. The ash
contains less than 10 percent carbon.
The gas leaving the gasifier contains, as H_S, 10-20 percent of the sul-
fur in the feed coal. The balance of the sulfur, 60-80 percent, is converted
to CaS. The ratio of H_ to CO exceeds 3$1 and a shift reactor is not
required.
Table 4-4 presents a water equivalent hydrogen balance scaled to
250 x 10 sCF/day of pipeline gas from the conceptual design for
6 28
260 x 10 SCF/day.
4.5 AGGLOMERATING BURNER-GASIFICATION PROCESS
The Agglomerating Burner-Gasification process as described by
29
-------
TABLE 4-4. WATER EQUIVALENT HYDROGEN BALANCE FOR THE CO--ACCEPTOR PROCESS
Ai
Coal feed: 2,248,565 Ib/hr of North Dakota lignite as received (33.67 wt.%
of moisture). 1,281,215 Ib/hr of preheated, moisture-free lignite fed
to the gasifier (heating value 11,120 Btu/lb).
Product pipeline gas: 250 x 10 SCF/day at heating value of 953 Btu/SCF.
IN
Water equivalent of hydrogen in preheated
lignite to devolatilizer
Steam to gasifier
10 Ib/hr
462
1035
1497
gals/10 Btu
5.6
12.5
18.1
OUT
Water vapor in flue gas from regenerator
Process condensate from cooling and
spraying
Process condensate from purification
Condensate from methanation
Water equiv. of hydrogen from
purification (H S)
Water equiv. of hydrogen in product gas
23
1
918
1468
0.3
155
115
256
1.9
1.4
3.1
11.1
17.8
30
-------
Battelle/Union Carbide31 is a pressurized, two-stage, fluidized-bed system
involving combustion of coal or char in one fluidized bed and the steam gasi-
fication of coal in a separate fluidized bed. The heat for the gasification
reactions is provided by circulation of ash from the burner to the gasifier.
The self-agglomerating method of fluidized-bed combustion, a key feature to
the process, consists of burner operation conditions which cause the ash in
the coal to agglomerate into free-flowing, inert solid pellets.
Initial development of this process was conducted by Battelle in the
early and mid-sixties under sponsorship of union Carbide. Current objectives
include a large-scale development at higher-than-atmospheric pressure. A
25-ton-coal/day Process Development Unit is presently under construction at
West Jefferson, Ohio. It will produce between 800,000 and 1,200,000 SCF/day
of synthesis gas. The gasification section of the PDU will operate under
about 100 psig of pressure.
The burner-gasification section is shown in Figure 4-8. For the Eastern
bituminous coal to be used in the PDU, the burner will operate at 2050°F and
the gasifier at about 1800°F.
Crushed coal (minus 100 mesh) and conveying air will enter the burner by
a pipe which passes through, and is flush with, the surface of the air distri-
butor plate. Additional air will enter from below the distributor plate.
During initial operation the fluidized bed in the burner will be main-
tained below the point where the diameter increases. Hot, agglomerated ash
particles will overflow the burner and be continuously transferred to the gas-
ifier by means of a steam lift.
Coal (minus 8 plus 100 mesh) will enter the gasifier on the same horizon-
tal plane as the feed location for ash agglomerates. Superheated steam will
enter the gasifier below and above the fluidized-bed distributor plate. There
will be a net downward flow of agglomerated ash particles through the gasi-
fier. in passing through, they will transfer a portion of their sensible heat
to support the gasification reactions. Ash agglomerates will be continuously
stripped of carbon by the upward flow of steam and will be then continuously
returned to the burner by means of an air lift for reheating. In passing
through the gasifier, the ash agglomerate temperature drops from 2000°F to
1500-1600°F. Ash agglomerates will be removed from the loop at a rate equal
31
-------
FLUE GAS
BURNER
y
SCRUBBER
FLY ASH
COAL FEED
(INERT GAS )
RAW GAS
AIR
COAL 8 AIR
GASIFIER
ASH
t
'FLY ASH
ASH STEAM
CHAR FOR
RECYCLE
SUPERHEATED
STEAM
AIR
ASH
REMOVAL
Figure 4-8. Burner-gasifier feed and circulation system for Agglomerating
Burner gasification process
32
-------
to that of the ash in the fuel fed to the system.
No material balance information has been published.
4.6 WINKLER PROCESS
The Winkler process to produce low- or medium-Btu power gas has been in
operation since the twenties. The heating value of the product gas depends on
whether air or oxygen is used as the gasifying medium. The main difference
between the oxygen blown Winkler process and the various processes described
to produce high-Btu pipeline gas is the absence of a shift converter and a
methanation stage. Details of the Winkler process can be found in References
32 and 33.
Crushed coal (smaller than 3/8 inch) is fed to the gasifier (Figure 4-9)
through a variable speed screw feeder. Operating experience shows that drying
the coal before gasification is not necessary so long as the surface of the
coal is not wet (approximately 18 percent moisture content or less). In the
fluidized bed of the Winkler gasifier, the coal particles react with oxygen
(or air if low-Btu gas is desired) and steam resulting in a gas rich in hydro-
gen and carbon monoxide. The gasification temperature is in the range of 1450
to 1850°F depending on the reactivity of the coal used. The pressure is about
one to three atmospheres. Unreacted carbon entrained in the gas is gasified
with secondary oxygen and steam in the space above the fluidized bed, so that
the maximum temperature occurs above the fluidized bed. This helps to
decrease the production of tars and oils.
To prevent molten ash from forming deposits blocking the gas exit dust,
the gas is cooled in a radiant boiler immediately above the gasification zone.
About 70 percent of the ash is carried over by gas and 30 percent is removed
from the bottom of the gasifier. After the sensible heat of the raw gas leav-
ing the gasifier is recovered, fly ash and char are removed by cyclones fol-
lowed by a wet scrubber and, finally, an electrostatic precipitator.
Purification (desulfurization) of the raw gas may be achieved by any of
the commercially available acid gas removal systems. As examples, a combina-
)lys
32
tion of COS hydrolysis and Alkazid absorption is recommended, as is the
Stretford system.'
33
-------
L.P. FEED BUNKER
LOCK HOPPERS
H.P. FEED BUNKER
FECO SCREWS
ASH CONVEYOR
GASIFIER
WASTE HEAT RECOVERY TRAIN
ASH BUNKER
ROTARV LOCKS
ASM CONVEYOR
Figure 4-9 Winkler gasifier and heat recovery system
34
34
-------
Table 4-5 presents a process water equivalent hydrogen balance for an
integrated oxygen-blown Winkler process. Also in Table 4-5 is a similar bal-
ance of relevant process streams for an integrated air-blown Winkler process.
The tables are based on the preliminary process designs for production of
10.5 x 10 SCF/day of medium-Btu gas and 23.4 x 10 SGF/day of low-Btu gas
and we expressed the results in gals/10 Btu of product gas. The use of oxy-
gen requires more steam to the gasifier to control the reaction. One third
of the extra steam reacts and two thirds is recovered as dirty condensate.
4.7 THE STIRRED-BED PROCESS
The Stirred-Bed Producer is under development at the Bureau of Mines in
35—38
Morgantown, West Virginia. Figure 4-10 shows the pilot plant producer.
The producer operates at about 205 psi. In a typical run, coal having the
analysis shown on Table 4-6 is fed to the reactor at a rate of 1800 Ib/hr.
Air and superheated steam were introduced below the grid and passed upwards
through the descending coal producing the gas which is withdrawn through an
off-take in the top cover. 0.4 Ib steam is used per pound of coal (and 2.6 Ib
of air). The temperature at the exit is held to about 1000°F by the
steam/coal/air ratio. Gas having the composition of Table 4-6 was produced
at a rate of 77,700 SCF/hr.
In the pilot plant, deep continuous stirring is provided within the pro-
ducer by a water-cooled stirrer equipped for compound motion that is a hori-
zontal rotation and vertical reciprocation. The stirrer prevents caking and
promotes good gas to solid contact. Pressure in the gasifier is maintained
by the resistance to gas flow through the off-take line containing a primary
orifice of diameter 1.067 inches and a secondary orifice of .800 inch dia-
meter. Located between the two orifices there is a cyclone separator to
remove part of the dust entrained in the gas.
For startup, anthracite is used to avoid depositing tar in the cold sys-
tem. Steam flow is started and then air flow increased slowly. Untreated,
bituminous, crushed coal feed is introduced when the gas reaches an off-take
v T*K •
temperature of 1000°F. Approximately four hours is required for the anthra-
cite to burn out and be completely replaced by a bituminous bed.
35
-------
TABLE 4-5. WATER EQUIVALENT HYDROGEN BALANCES FOR THE
WINKLER PROCESS (KENTUCKY COAL)
gals/10 Btu of product gas
oxygen air
blown blown
IN
Moisture in coal 0.7 0.6
Water equiv. of hydrogen in coal 4.7 5.6
Steam to gasifier 12.1 5.3
Process water to gasifier 1.0 0.9
18.5 12.4
OUT
Settler waste water from char removal 7.6 3.7
Moisture in wet char from char removal 0.7 0.6
Water vapor in vent gas from char
removal 0.1 0.1
Water equiv. of hydrogen in product
18.5 12.5
gas 10.1 8.1
36
-------
COUNTER
WEIGHT
COAL
HOPPER
30 FT3
27'-9
46'-0'
COAL
HOPPER
1130 FT3
25'-10'
HIGHEST
6T-0"
POSITION
L
LOWEST
OPERATING
POSITION
1
.
12'
\
1 V
' \
(«
•
-0" M
r
i
'.-,-
i
L
-^
S
4 v.
$
l'-0"
K-AIR-STEAM
3'-6"
ASH HOPPER
Figure 4-10. Stirred, pressurized, gas producer.
37
-------
TABLE 4-6. TYPICAL DATA FOR A STIRRED-BED GASIFIER
FED NEW MEXICO SUBBITUMINOUS COAL38
RUN-PERIOD
A. Coal
FSI
H2° %
Ash
C
H
S
N
O
Btu/lb
B. Bottom Ash,
%, (Dry Basis)
Ash
C
H
S
C . Tar , %
(Dry Basis)
Ash
C
H
S
N
Btu/lb
62-2
2
8.8
24.2
50.8
5.1
1.0
1.6
8.5
8900
94.6
4.1
.6
.2
8.1
77.6
6.6
1.3
1.6
13565
Pressure (psig)
Input, Ib/hr
Coal
Air
Steam
Input Ratios, Ib/lb
Steam: Coal
Air: Coal
Output , Ib/hr
Ash
Cyclone Dust
Gas
Tar
Water
Gas Yield
MSCFH
SCF/LB coal
Gas Analysis (vol. %)
CO
co2
N2
H2
CH4
C,H
2 6
145
1800
4583
763
.4
2.6
330
86
5784
53
889
77.7
43.4
Heating Value
15.3
12.7
59.3
10.7
2.1
0.0
.2
0.0
104 Btu/Cl-
38
-------
TABLE 4-7. WATER EQUIVALENT HYDROGEN BALANCE FOR
THE STIRRED FIXED BED PROCESS
Actual coal feed: 1,490 Ib/hr
IN
gals/10 Btu of product gas
Moisture in coal 1.6
Water equivalent of hydrogen in coal 7.9
Steam to gasifier 11.5
21.0
OUT
Water in product gas 10.9
Water equivalent of hydrogen in gas 10.3
21.2
Table 4-7 presents the water equivalent hydrogen balance of the relevant
process streams for the stirred bed producer gasification process. The num-
bers are based on the results of the Bureau of Mines research studies con-
verted to units of gal/10 Btu of product gas.
4.8 MOLTEN SALT PROCESS
The Kellogg Molten Salt process has been studied in many variations with
time. Coal is gasified by steam alone in a bath of molten sodium carbon-
ate. The heat absorbed by the gasification reaction cools the salt bath. The
molten salt, with carbonaceous char and ash, is circulated from the gasifier
to the combustor where air is introduced. Residual carbon burns and reheats
the salt. A sidestream of molten salt is continuously treated to remove ash.
Figure 4-11 shows a system with the gasifier and combustor in a single, parti-
tioned vessel. Figure 4-12 shows a two-vessel system. Typical temperatures
are shown on the figures. Synthesis gas is produced under a pressure of about
400 psi. Pressures up to 1000 psi have been tested. At high pressures oxy-
gen, and not air, would probably be used whether or not pipeline gas is the
desired product because of the expense of compressing the nitrogen in the air.
39
-------
COAL-
COAL
LOCK
HOPPERS
1
STEAM
V
RAW SYNTHESIS GAS
GASIFIER
405 PSIA
22000F
18300F
405 PSIA
COM-
\ BUSTOR
19000F
FLUE GAS
MELT PURGE
/v
AIR
MAKE-UP
SODIUM CARBONATE
RECYCLE
SODIUM CARBONATE
CARBONATE
LOCK
HOPPERS
Figure 4-11. Molten Salt Process; gasifier and conibustor in a single vessel
42
-------
COAL
STEAM
COAL
LOCK
HOPPER
1750 *F
1200 PSIA
1850'F
¥
GASIFIER
COM8USTOR
I200PSIA
RAW SYNTHESIS GAS
FLUE GAS
CARBONATE
LOCK
CIRCULATING
SAU
\
I
f
ASH
OXYGEN
Figure 4-12. Flow diagram for Molten Salt gasification section using two vessels
42
-------
Sulfur in the coal accumulates in the bath as sodium sulfide until the
following equilibrium is established:
Na S + C02 + H20 £ Na2c°3 + H2S
Hydrogen sulfide begins to leave with the gas. Ash accumulates in the molten
salt and a purge stream is removed to hold the ash at about 8 percent. The
purge stream treatment is shown in Figure 4-13. The stream is quenched to
400°F with a solution saturated with sodium bicarbonate at 100°F in the quench
tower. Solid melt particles in the resulting slurry are ground to facilitate
dissolution of,the salt. This stream is then flashed to essentially atmos-
pheric pressure into a holding tank, where sufficient residence time is pro-
vided to dissolve the sodium carbonate. The slurry leaving this vessel is
filtered to separate the ash and carbon from the solution. This residue is
sent to disposal.
The solution leaving the filter flows to a carbonation tower where the
sodium carbonate is reacted with carbon dioxide from the gas purification sys-
tem. The tower operating temperature is about 100°F, At this temperature the
sodium bicarbonate concentration exceeds its solubility limit and is precipi-
tated from solution. The slurry is then filtered; the bicarbonate leaving the
filter is calcined to decompose the bicarbonate to carbonate and is returned
to the combustor, while the solution is recycled to the quench tower.
Reference 39 is a design of a pipeline gas plant. A hydrogen balance for
this plant is presented in Table 4-8. This balance should be treated as being
only an indication because it is based on somewhat outdated technology. For
this reason it has not been adapted to production of power gas.
4.9 THE LURGI PROCESS FOR UTILITY FUEL GAS PRODUCTION
In this section a very brief mention will be given of recent developments
by the British Gas Corp. and Lurgi in running the gasifier in a slagging
j 44
mode.
Lurgi gasifiers for SNG production are oxygen blown and use about
7 moles steam/mole oxygen with Navajo coal and up to 9 moles steam/mole
oxygen with lignite; 45 to 50 percent of the steam is decomposed. Lurgi
42
-------
*>
U)
MELT FROM GASIFIED (16% ASH) VENT GAS
GRINDER
/"-^N
TOWER
r"
4000F
TANK
WASH WATER
2000F
~\\j~ \
ASH
16 PSIA
1200 PSIA
BICARBONATE
FILTER
1000F
SATURATED SODIUM BICARBONATE SOLUTION\Q
SODIUM CARBONATE
TOGASIFIER -«
U CALCINER U
NaHCOa
ASH 4 CARBON
TO DISPOSAL
FILTER
SODIUM CARBONATE
SOLUTION
VENT GAS
CARBONATION
TOWER
CARBON DIOXIDE
FROM
PURIFICATION
SECTION
30 PSIA
Figure 4-13. Flow diagram of an ash removal section
40,42
-------
TABLE 4-8. WATER EQUIVALENT HYDROGEN BALANCE FOR THE MOLTEN SALT PROCESS
Coal feed: 1,100,000 Ib/hr of run-of-mine bituminous coal (2.0 wt.%
moisture and higher heating value of 13,990 Btu/lb).
Product pipeline gas: 250 x 10 SCF/day with heating value of
914 Btu/SCF.
£T
gals/10 Btu
10 Ib/hr product gas
IN
Moisture in coal 22 0.3
Water equiv. of hydrogen in coal 515 6.5
Steam to gasifier 1,000 12.6
Boiler feed water to shift conversion 147 1.9
Clean water to scrubber 360 4.5
Process water to ash removal 607 7.7
2,651 33.5
OUT
Process condensate from shift conversion 163 2.1
Process condensate from scrubber 582 7.3
Condensate from methanation 313 3.9
Water condensate from ash removal 413 5.2
Water in residue from ash removal 3
Water equiv. of hydrogen in vent gases
from purification and ash removal
(CH , H^ and « S) 13 0.2
42 ^
Water vapor to stack from gasification,
purification and ash removal 291 3.6
Water equiv. of hydrogen in product gas 909 11.5
2,687 33.8
44
-------
gasifiers for utility fuel gas production are air blown and use about
3.77 moles steam/mole oxygen in air (this equals 7.5 moles steam plus
nitrogen/mole oxygen) and 45 percent of the steam is decomposed. In the
slagging mode the oxygen/carbon ratio is slightly increased, but the ratio
of steam/oxygen is reduced to about 1.5 and all of the steam is decomposed.
Only the coal moisture appears in the gas. This much reduces the steam need
and water treatment plant size. Also, when the gas is cooled to remove tars,
etc., very little water condenses out. Condensing water is more than half of
the cooling load in the utility gas plants. With little water to condense
there is little irreversible loss on cooling, and slagging gasifiers have a
higher cold gas efficiency than the usual Lurgi gasifier. The experimental
44
gasifier ran at more than four times the throughput of the normal gasifier.
There is no doubt that the slagging gasifier will prove most useful.
4.10 KOPPERS-TOTZEK PROCESS
The Koppers-Totzek is a commercially established high-temperature,
entrained-flow gasifier capable of partially oxidizing a wide variety of feed
stocks. The gasifiers operate at about 8 psig and oxygen, not air, is
used. 5~51 caking coals can be gasified without pretreatment. The coal
receives primary crushing and drying followed by simultaneous drying and
pulverization using a ball, rod or roller mill. The degree of drying, vary-
ing from 2 to 8 percent, depends on the material to be pulverized. The dry-
ing medium, either hot flue gas or product gas combusted with excess air, is
circulated through the mill. The pulverized coal, of which about 70 percent
will pass through 200 mesh, is conveyed with nitrogen from storage to the
gasifer service bin and feed bins. Variable speed coal screw feeders then
continuously discharge the coal into a mixing nozzle where a mixture of
steam and oxygen entrains the pulverized coal. Moderate temperature and
high burner velocity prevent the oxidation of the coal in the nozzle.
The gasifier (Figure 4-14) is a refractory-lined steel shell equipped
with a steam jacket for producing low pressure process steam. A two-headed
burner has heads 180° apart and can handle about 400 tons of coal per day. A
four-headed gasifier, with burners 90* apart, can handle about 850 tons of
coal per day. The burner heads are opposed so that particles escaping from
45
-------
-
r
TO SLOWDOWN
NITROGEN RETURNS
BOILER FEED WATER->
NITROGEN CONVEYED
LOW PRESSURE STEAM
TO BLOW DOWN TANK
COAL FEED SCREW
CONVEYOR
OXYGEN-^
BOILER FEED WATER
SLAG DISPOSAL CONVEYOR
SLAG REMOVAL CONVEYOR
HIGH PRESSURE
SUPERHEATED STEAM
SPRAY WATER
>-GAS TO WASHER COOLER
COAL FEED
SCREW CONVEYOR
WATER OVERFLOW TO CLARIFIER
I SERVICE BUNKER
FEED BUNKER
HIGH PRESSURE
STEAM DRUM
.LOW PRESSURE
STEAM DRUM
(GASIFIER
WASTE HEAT
BOILER
GASIFIER QUENCH
1 TANK
SERVICE BUNKER
FEED BUNKER
Figure 4-14. Koppers-Totzek gasifier and heat recovery system.
-------
one burner will be burnt in the opposite burner. Gasifiers currently under
design will operate at about 8 psig so that the exiting gases have enough
pressure to pass through venturi scrubbers. Carbon is oxidized in the gasi-
fier producing a high temperature flame zone of about 3300-3500°P. The endo-
thermic carbon-steam reactions reduce the exit temperature to around 2700°F.
The coal is gasified almost completely and instantaneously. Carbon conversion
depends on the reactivity of coal and is about 96-98 percent. At the prevail-
ing high operating temperatures, gaseous and vaporous hydrocarbons emanating
from the coal decompose so rapidly that coagulation of coal particles during
the plastic stage doe's not occur. Thus any coal can be gasified regardless of
the caking property, ash content or ash fusion temperature. Also, only gas-
eous products are produced; no tars, condensable hydrocarbons or phenols are
formed. Approximately 50 percent of the coal ash drops out as slag into a
slag gasifier as fine fly ash.
Gas leaving the gasifier may be direct-water quenched to solidify
entrained slag droplets, if necessary, and then passes through a waste heat
boiler where high pressure steam is produced. The gas exiting the waste heat
boiler at 350-500"F is then scrubbed and cooled to approximately 95°F. The
entrained solids are reduced to 0.002-0.005 grain/SCF. Particle-laden water
from the gas scrubbing and cooling system is piped to a clarifier. The recov-
ered clean water is cooled and recirculated through the gas washing system
(see Figure 4-15).
The gas contains approximately 95 percent of the total sulfur in the
coal, and this is removed by any appropriate process.
Table 4-9 presents a water equivalent hydrogen balance of the relevant
process streams for a plant manufacturing 250 x 10 sCF/day of medium-Btu
power gas (from Reference 47 and personal communication from G.V. McGurd of
Koppers Engineering and construction).
For reasons given in Section 7, the Koppers-Totzek process is used in our
Solvent Refined Coal plant to gasify the filter residue. This is a particu-
larly high ash material, and in order to make material balances it was neces-
sary to extrapolate from actual operating experience. A study of References
43, 46, and 49 suggested the following performance with a high ash coal in
which the H/C ratio (Ib/lb) is more than about 0.77:
47
-------
OC
Figure 4-15 Koppers-Totzek gasification process.
-------
TABLE 4-9. WATER EQUIVALENT HYDROGEN BALANCE FOR THE
KOPPERS-TOTZEK PROCESS
Coal feed: 480,605 Ib/hr of as-received bituminous coal
(16.5 wt.% moisture and 4.77 wt.% hydrogen, higher heating
value of 8830 Btu/lb).
Product power gas: 250 x 10 SCF/day with higher heating value of
303 Btu/SCF (32.6 dry mole % hydrogen and 0.1 dry mole %
methane, moisture is estimated as 0.5 mole %).
gal/10 Btu
product gas
IN
Moisture in coal 3.0
Water equiv. of hydrogen in coal 6.2
Steam to gasifier 2.7
Water to gasifier for spraying hot gas 1.6
Water to purification for makeup 0.. 1
13.6
OUT
Water with wet slag from gasifier 0.2
Water with wet ash from scrubbing and cooling 2.0
Condensate from compression 1.1
Condensate in cooling (net) 0.8
Water vapor in vent gas from coal drier 3.1
Water vapor from purification to Claus 0.0
Water equiv. of hydrogen from purification (H.S) 0.1
Water equiv. of hydrogen in product gas 6.1
Water vapor in product gas 0.2
13.6
49
-------
oxygen feed, Ib/lb (carbon + hydrogen) 1.06
steam feed, Ib/lb (carbon + hydrogen) 0.223
Mole fractions of gases leaving the reactor to be such that
(H2) (002)
= (approx.) 0.5
(CO) (H2)
Note that methane is not produced in this gasifier.
4.11 H-COAL PROCESS
Hydrocarbon Research, Inc. (HRI) has developed a process for coal lique-
faction by catalytic hydrogenation in an ebullated-bed reactor. HRI has con-
ducted bench-scale experiments since 1965, using reactors about 3/4 inch ID
and handling about 25 lb of coal per day. Development work was continued on
Process Development Units (PDU) using reactors about 8.5 inch ID and handling
about 2.5 tons/day of coal feed, representing a scale-up factor of about 200.
HRI is now building a prototype demonstration plant. It will have a hydrogen-
ation reactor of 4.5 ft ID capable of processing 250 to 750 tons/day of coal
depending on whether synthetic crude or low-sulfur fuel oil is the desired
product (725 bbl/day of synthetic crude or 2250 bbl/day of fuel will be pro-
duced) . This represents a scale-up factor of 100 from the PDU. The commer-
cial size reactors would represent a scale-up factor of 10 from the prototype
unit. More description on the Process Development Units and experimental data
can be found in References 52-55.
54
Process designs and economic evaluations have been made for converting
an Illinois coal to gasoline and furnace oil, and to gasoline, LPG ar.d ber.-
zene. Ammonia and sulfur are produced as byproducts in all cases. As an
example we consider the liquefaction of an Illinois coal to gasoline and fur-
nace oil. The capacity of the refinery has been scaled to 50,000 bbl/day of
liquid products.
As-received coal is dried, pulverized and slurried with coal-derived oil
50
-------
for charging to the coal hydrogenation unit shown in Figure 4-16. The reactor
operates at 3000 psi and contains an ebullated bed of cobalt-molybdenum cata-
lyst wherein the coal is catalytically hydrogenated and converted to liquid
and gaseous products. In the ebullated bed the upward passage of the solid,
liquid and gases maintains the catalyst in a fluidized state. Catalyst can
be added and withdrawn continuously to keep a constant level of activity.
Reactor temperature is controlled by the preheat temperature of the hydrogen
feed stream. The hot gas is subjected to partial condensation and further
phase separation, after which the remaining gas is passed through an absorber
for removal of as much hydrocarbon as possible and sent to the hydrogen plant.
The synthetic crude produced can be refined to gasoline and furnace oil. The
detail of the refining process is described in References 54 and 55 and is not
included in this report.
About 90 wt percent of the oxygen in the coal is converted to water which
is sent to a sulfur and ammonia recovery section. Much of the remaining oxy-
gen leaves the hydrogenation step as phenols and other oxygenated aromatics.
Hydrogen consumption is approximately 40,000 8CF per ton of coal charged.
This is for the production of gasoline and fuel oil. For the production of
fuel oil alone the hydrogen consumption can be halved. Table 4-10 presents a
water equivalent hydrogen balance supplied by Hydrocarbon Research, Inc. (com-
munication from F.D. Moffert). This balance matches the simplified flow dia-
gram shown on Figure 4-17.
Hydrocarbon Research, Inc. has also considered the production of a low-
sulfur residual fuel oil from a West Virginia,Pittsburgh seam coal. The flow
diagram is shown in Figure 4-18. In this case the vacuum bottoms slurry con-
taining coal residue is used to produce the hydrogen required for hydrogena-
tion by partial oxidation which is similar to our design of a Solvent Refined
Coal plant. The water equivalent hydrogen balance for this case is shown in
Table 4-11.
4.12 SYNTHOIL PROCESS
This process for the hydrogenation of coal to low-sulfur liquid fuel oil
is undergoing laboratory development at the U.S. Energy Research and
51
-------
CATALYST
INLET
SOLID-LIQUID
LEVEL
CATALYST
LEVEL
RECYCLE
TUBE
SOLID
INLET
LIQUID
INLET
VAPOR
OUTLET
10 V
r ^
•m i.wdui_i>
s •?••*_••
CATALYST
OUTLET
A
CLEAR
LIQUID
LIQUID-SOLID
SETTLED
"CATALYST
LEVEL
-DISTRIBUTOR
PLENUM CHAMBER
GAS
INLET
Figure 4-16. H-Coal ebullated-bed reactor.
52
-------
TABLE 4-10. WATER EQUIVALENT HYDROGEN BALANCE FOR THE H-COAL PROCESS - 1
Coal feed: 1,196,665 Ib/hr dry Illinois Coal (4.8 wt. % hydrogen and
9.2 wt. % oxygen as fed to the hydrogenator).
Product oils: Nominal 50,000 bbl/day of product oils (33,051 bbl/day of
gasoline, 16,667 bbl/day of domestic fuel oil and 1,113 bbl/day of
No. 6 fuel oil).
3 gals/10
10 Ib/hr product oil
IN
Water equivalent of hydrogen in dry coal 517 4.9
Steam to hydrogen plant 408 3.9
925 8.8
OUT
Waste water from sulfur and NH., recovery 133 1.2
Water equivalent of hydrogen in byproduct NH3 21 0.2
Water equivalent of hydrogen in coal residue 79 0.7
Water equivalent of hydrogen in self-generated fuel gas 39 0.4
*
Water equivalent of hydrogen in product oil mix 653 6.2
925 8.S
* Determined by difference equivalent to 12.2 wt. % hydrogen
in the oil mix, HHV approximated as 146,000 Btu/gal.
53
-------
off-gat
itcom
Ul
coal
hydrogen
HYDROGEN
PLANT
hydrocarbon
oases
lulfur
and
amcnonio
GAS
PURIFICATION
H2S
SULFUR
AMMONIA
RECOVERY
watt* wafer
hydrocarbon gam
gaus
COAL
PREPARATION
dry cool in
ilurry feed
aqueous
oeld
solution
HYDROGENATION
REACTOR
hydrogen
PHASE
SEPARATOR
synthetic
crude
REFINERY
gasoline
domestic fuel oil
No. 6 fuel otl
recycle oil
residue
SOLIDS
SEPARATOR
coal
residue
Figure 4-17. Flow diagram for process water streams in H-Coal process for production of 50,000 bbl/day
of product oils.
-------
LOW PRESSURE
VENT GAS •«
TO FURNACE
COAL
PREPARATION
RECOVERED
WATER
Ul
U1
WATER TO
COOLING TOWER
MAKEUP
0
>
o;
TANKAGE
0.3 SULFUR
FUEL OIL PRODUCT
PLANT FUEL TO
POWER GENERATION
PLANT FUEL
02
MANUFACTURE
COAL
HYDROGENATION
SOUR WATER
"2
COMPRESSION
VACUUM BOTTOM SLURRY
AIR
MANUFACTURE
RECOVERED NAPTHA
VENT GAS
NH,
H2S
RECOVERY
O
u
i
"2s
REFINERY GAS
CLEAN UP ft
MANUFACTURE
T
MANUFACTURE
502
POWER GENERATION
COOLING TOWER
WATER TREATING
STEAM RAISING
PLANT INSTRUMENT
AIR
FLAIR SYSTEM
ELECTRICAL SUBSTATION
ANHYDROUS
AMMONIA
HIGH 8TU
GAS PRODUCT
SULFUR
PRODUCT
Figure 4-18. Block flow diagram for H-Coal process.
-------
TABLE 4-11. WATER EQUIVALENT HYDROGEN BALANCE FOR THE H-COAL PROCESS - 2
Coal feed: 1,494,208 Ib/hr as-received Pittsburgh seam coal
(5.7 wt. % hydrogen).
Products: 41,041 bbl/day fuel oil sold (33,595 bbl/day slurry used in
production of hydrogen and in plant fuel requirements), and
128.3 x 10 SCF/day gas with lower heating value of 870 Btu/SCF.
IN
Water equivalent of hydrogen in feed coal
Water consumed in hydrogen manufacture
3 gals/10 3tu
10 Ib/hr products
767
459
1226
5.6
3.4
9.0
OUT
Waste water from S and NH_ recovery 134
Water equivalent of H- in product gases 477
Water equivalent of H2 in refinery gases used as fuel 0.6
Water equivalent of H2 in waste gases to flare
Water equivalent of H» in byproduct NH_
Water equivalent of H in product oils
0.6
37
Water equivalent of ^2 ^-n slurry used to produce power
and used as plant fuel 102
513
1264
1.0
3.5
0.0
0.0
0.3
0.7
3.8
9.3
* 10.4 wt. % H_ in the. oil, HHV approximated as 157,400 Btu/gal.
56
-------
Development Administration, Pittsburgh Energy Research Center. Early work
using a 5 Ib (slurry)/hr reactor has been described in References 56-58. The
process gives a mobile fuel oil (7.6 percent hydrogen or more, molar ratio of
C/H about 1/1) with a very low sulfur content from coals having 4 to 5 percent
sulfur. The desulfurization aspects are particularly discussed in References
58-60.
Figure 4-19 shows the present pilot plant which handles 1/2 ton
61
(slurry)/day. Coal of 200 mesh dried to 0.5 percent moisture content is
slurried to about 35 wt percent in recycle oil, mixed with hydrogen, pre-
heated and passed up one or two reactors, each 14.5 ft long and 1.1 inch ID.
The reactors are packed with pellets of cobalt-molybdenum catalyst. Flow
through the reactors is two-phase and turbulent; the liquid is propelled by
the gas. The reactor operates at about 450°C and 2000 psi or more.
Upon leaving the reactor gas is separated from the liquid and recycled.
A steady purge is removed from the gas stream and makeup hydrogen is contin-
uously added. The liquid is centrifuged from suspended solids and some of it
is recycled to slurry more coal.
The only integrated plant design (including hydrogen production) which
we have seen is that of Reference 63 made for a Wyoming coal and made speci-
fically for cost estimating purposes. For the purpose of estimating water
requirements we have, in a separate study (subcontract No. 1916-3 to univer-
sity of Oklahoma, primary contract EPA 68-01-1410), chosen to make our own,
somewhat simplified design. The block diagram of Synthoil liquefaction pro-
cess is reproduced as Figure 4-20. Based on the experience reported in Refer-
ences 61-63, various rules were made for material balance calculations:
1) five barrels of oil were produced from each ton of carbon in coal;
2) 4700 SCF of hydrogen were needed for one barrel of oil; 3) 6.2 percent
of carbon in coal remained in the char to hydrogen plant. About 83 percent
of oxygen in coal was converted to water and discharged from pha^s i;_._. ...-
4) the hydrogen needed for liquefaction was produced from the gasification or
char and as-received coal followed by two water shift reactions. Gasification
information was extrapolated from Reference 63 and the production train is
shown on Figure 4-21.
57
-------
:
Feed tonk
coal + vehicle
Slurry feed
pump
5lb/hr
Gas
meter H2 compressor
500 scfh
2000-4000 psig
Packed bed
reactor
A 68'x 5/16"ID
Furnace
Presenter
Flore
StOCk
Woter
High pressure
receivers
nore
stack
Low pressure
receivers
Main
product
Gas
sampling
Figure 4-19. Synthoil pilot plant.
-------
COAL
(Ji
t
WATER
VAPOR
COAL
PREPARATION
a DRYING
HYDROGEN
HYDROGEN
PRODUCTION
a
COMPRESSION
I
STEAM
a
WATER
OXYGEN WATER
CONDENSATE
RECYCLE OIL
COAL
SLURRY
PREPARATION
230° F
I
_ HEAT ' _
EXCHANGER
I I
I <
RECYCLE GAS
PURIFICATION
WATER
CONDENSATE
CHAR
800°F
PHASE
SEPARATION
CHAR
DE-OILING
REACTOR
t
GAS
OIL
*- SALES GAS
PLANT FUEL
Figure 4-20. Flow diagram for process water streams in Synthoil process.
-------
COAL
aCHAR
01
o
900 V
DIRTY
CONDENSA7E
STEAM
CW
CLEAN
CONOENSATE
Figure 4-21. Flow diagram for hydrogen production in Synthoil process.
-------
TABLE 4-12. APPROXIMATE COAL ANALYSES FOR SYNTHOIL PROCESS
As-received coal composition (%)
New Mexico
Subbituminous
Wyoming
Subbituminous
North Dakota
Lignite
c
H
0
N
S
Ash
Moisture
HHV (Btu/lb)
47.3
3.4
0.8
9.6
0.9
25.6
12.4
8310
49.4
3.4
0.5
12.8
0.3
5.6
28.0
8449
40.5
2.7
0.8
11.9
0.7
7.4
36.0
6822
Table 4-12 shows the approximate coal analyses at three western sites as
used in the separate study mentioned above. Those compositions differ some-
what from coals to be used later in other conversion processes in this report.
Water equivalent hydrogen balances of a Synthoil plant producing
50,000 bbl/day of oil using these three types of coal are shown on Table 4-13.
61
-------
TABLE 4-13. WATER EQUIVALENT HYDROGEN BALANCE FOR THE SYNTHOIL PROCESS
Basis: 50,000 barrels/day oil (about 13.45 x 10
Units: 103 Ib/hr.
PROCESS INFORMATION
As-received coal fed to liquefaction (if}
As-received coal fed to gasification (3j
HYDROGEN BALANCE
IN
Moisture in coal to liquefaction (2)
Water equiv. of hydrogen in coal to liquefaction
Moisture in coal to gasification ^)
o> Water equiv. of hydrogen in coal to gasification
to
Total steam to hydrogen production v./TjX
Quench water to hydrogen production
OUT
From drying coal to liquefaction (4j
Total dirty process condensate
Clean process condensate /NL/ SL/
Moisture in product oil (7)
Water equiv. of hydrogen in product oil (T)
Water equiv. of hydrogen and moisture in gas
produced (8\
9 Btu/hr) .
New Mexico
1766
334
438
(2) 1098
83
(j) 207
529
581
2936
422
635
139
1143
16
567
2922
Wyoming
1691
318
947
1035
178
194
511
528
3393
935
759
132
1143
12
414
3395
North Dakota
2065
554
1486
1008
399
270
467
670
4300
1473
1102
96
1143
13
464
4291
(continued)
-------
TABLE 4-13. (continued)
Units: gals/10 Btu product oil
New Mexico
to
pi
Wyoming
North Dakota
IN
Moisture in coal to liquefaction {2) 3.9
Water equiv. of hydrogen in coal to liquefaction {2j 9.8
Moisture in coal to gasification {3} 0.7
Water equiv. of hydrogen in coal to gasification \3y 1.8
Total steam to hydrogen production 5.1
8.5
9.2
1.6
1.7
4.6
4.7
30.3
8.3
6.8
1.2
10.2
0.1
3.7
13.2
9.0
3.5
2.4
4.2
6.0
38.3
13.1
9.8
,0.9
10.2
0.1
4.1
26.1
30.3
38.2
-------
(This page left blank intentionally)
62b
-------
REFERENCES SECTION 4
1. Hottel, H. C. and Howard, J. B., New Energy Technology - Some Facts and
Assessments, M.I.T. Press, Cambridge, Mass., 1971.
2. Rudolph, P. F. H., "The LURGI Process, The Route to SNG from Coal,"
4th Synthetic Pipeline Gas Symposium, Chicago, Illinois, Oct. 30 and
31, 1972.
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4. Elgin, D. C., "Trials of American Coals in LURGI Pressure-Gasification
Plant at Westfield, Scotland," Proceedings of Fifth Synthetic Pipeline
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63
-------
11. El Paso Natural Gas Company, "Second Supplement to Application of El
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20. Grace, R. J., Brant, V. L. and Kliewer, V. D., "Design of BIGAS Pilot
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64
-------
24. Air Products and Chemicals, Inc», "Feasibility Study of a Coal Slurry
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Symposium on Clean Fuels from Coal, Institute of Gas Technology,
Chicago, Illinois, Sept. 10-14, 1973.
47. Magee, E. M., Jahnig, C. E. and Shaw, H., "Evaluation of Pollution
Control in Fossil Fuel Conversion Processes. Gasification, Section 1:
KOPPERS-TOTZEK Process," EPA Technology Series, EPA-650/2-74-009, 1974.
48. Wintrel, R., "The K-T Process: Koppers Commercially Proven Coal and
Multi-fuel Gasifier for Synthetic Gas Production in the Chemical and
Fertilizer Industries," presented at Natl. Meet. A.I.Ch.E., Salt Lake
City, Utah, August 1974.
49. Mitsak, D. M. and Karmody, J. F., "Koppers-Totzek: Take a Long Hard
Look," presented at 2nd. Ann. Symp. on Coa. Gasification; Best Prospects
Commercialization, Univ. of Pittsburgh, August 1975.
66
-------
50. Farnsworth, J. F. , Mitsak, D. M, and Kamody, J. P., "Clean Environment
with K-T Process," delivered at EPA Symposium on Environmental Aspects
of Fuel Conversion Technology, St. Louis, MO., May 1974.
51. Kamody, J. F. and Farnsworth, J, F., "Gas from the Koppers-Totzek
Process for Steam and Power Generation," presented at the Industrial
Fuel Converence, Purdue Univ., October 1974.
52. Johnson, C. A., Chervenak, M. C., Johanson, E. S., Statler, H. H.,
Wolk, O. W. and Wolk, R. H., "Present Status of the H-COAL PROCESS,"
IGT Symposium on Clean Fuels from Coal, Institute of Gas Technology,
Chicago, Illinois, Sept. 10-14, 1973.
53. Johnson, C. A., Chervenak, M. C., Johanson, E. S. and Wolk, R. H.,
"Scale-Up Factors in the H-COAL Process," Chem. Eng. Progr. 69,
No. 3, 52-54, 1973.
54. Hydrocarbon Research, Inc., "Project H-COAL Report on Process Develop-
ment," Office of Coal Research, R&D Report No. 26 - Final Report,
1968.
55. American Oil Company, "Evaluation of Project H-COAL," Office of Coal
Research RSD Report No. 32 - Final Report, 1967.
56. Akhtar, S., Friedman, S. and Yavorsky, P. M., "Process for Hydro-
desulfurization of Coal in a Turbulent-Flow Fixed-Bed Reactor,"
AIChE 71st Annual Meeting, Dallas, Texas, February 20-23, 1972.
57. Yavorsky, P. M., "Hydrodesulfurization of Coal into Non-Polluting
Fuel Oil," Pittsburgh Energy Research Center, U.S. Bureau of Mines,
October 1972.
58. Yavorsky, P. M., Akhtar, S. and Friedman, S., "Converting Coal into
Non-Polluting Fuel Oil," Chem. Eng. Progr. 69, No. 3, 51-52, 1973.
59. Akhtar, S., Friedman, S. and Yavorsky, P. M., "Low-Sulfur Liquid
Fuels from Coal," ACS Symposium on Quality of Synthetic Fuels,
Boston, Mass., April 9-14, 1972.
60. Akhtar, S., Sharkey, A. G., Jr., Schultz, J. L. and Yavorsky, P. M.,
"Organic Sulfur Compounds in Coal Hydrogenation Products," ACS 167th
National Meeting, Los Angeles, Calif., March 31 - April 5, 1974.
61. Akhtar, S., Mazzocco, N. J., Weintraub, M. and Yavorsky, P. M.,
"SYNTHOIL Process for Converting Coal to Non-Polluting Fuel Oil,"
4th Synthetic Fuels from Coal Conference, Oklahoma State University,
Stillwater, Oklahoma, May 6-7, 1974.
62. Akhtar, S., Lacey, J. J., Weintraub, M., Rezik, A. A. and Yavorsky,
P. M., "The SYNTHOIL Process - Material Balance and Thermal Efficiency,"
presented at 67th Annual Meeting, AIChE, Washington, D.C., Dec. 1-5, 1974,
67
-------
63. U.S. Dept. of the Interior, "SYNTHOIL Process Liquid Fuel from Coal Plant,
50,000 Barrels Per Stream Day. An Economic Evaluation," Report No.
ERDA 76-35, Bureau of Mines, Morgantown, W. Virginia, 1975; summarized
in: Katell, S. and White, L. G., "Economic Comparison of Synthetic Fuels
Gasification and Liquefaction," presented at ACS Natl. Meeting, Div.
of I&EC, New York, April 1976.
68
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SECTION 5
HYGAS PROCESS
5.1 INTRODUCTION AND SUMMARY OF RESULTS AT WYOMING SITE
The Hygas process has been under development by the Institute of Gas
Technology since 1945. The development of the Hygas process has advanced
to the large pilot plant stage with a facility capable of producing
1.5 x 10 SCF/day of pipeline gas from 75 tons/day of coal. This pilot
plant is now being operated in Chicago, Illinois.
Three different versions of the basic Hygas process, namely, Hygas-
Oxygen, Hygas-Steam/Iron and Hygas-Electrothermal, have been designed.
Basically these three processes share the same coal pretreatment, material
handling, hydrogasification, purification, and methanation systems, but
they differ in the technique used to produce the hydrogen-rich stream for
hydrogasification. In this section the Hygas-Oxygen process is des-
cribed.
The run-of-mine coal is crushed to -8 x 100 mesh in cage mills and dried
to 2 percent moisture content. The coal is then slurried to 50 percent solid
concentration (by weight) with recycle slurry oil from downstream in the pro-
cess. The coal-oil slurry is then pumped to the gasifier operating pressure
of 1200 psig and heated in an external heater to 200*F.
The pilot-plant hydrogasifier is 135 ft high, 5.5 ft ID reactor vessel
with five internally connected reaction stages (Figure 5-1). The slurry is
pumped into the fluidized top section as a spray. Sensible heat in the gas-
eous products efficiently vaporizes the light oil and leaves dry coal to be
fed to the second section.
In the dilute-phase second section, which is the first stage of
69
-------
INLET FOR SLURRY
OF CRUSHED COAL
AND LIGHT OIL
FLUIDIZED BED IN
WHICH SLURRY OIL IS
VAPORIZED BY RISING.
HOT GASES AS
COAL DESCENDS
DRIED COAL FEED
FOR FIRST-STAGE
HYDROGASIFICATION
HIGH VELOCITY GAS
FROM SECOND-STAGE
MIXES WITH DRIED COAL
CHAR FROM FIRST STAGE
FEEDS INTO SECOND -
STAGE FLUIDIZED BED
HYDROGEN - RICH GAS
AND STEAM RISE TO
SECOND-STAGE
RAW GAS OUTLET
TO QUENCH CLEANUP
AND METHANATION STEPS
NITROGEN-PRESSURIZED
OUTER SHELL
SLURRY
DRIER
HOT GAS RISING
INTO DRIER
GAS - SOLIDS
DISENGAGING
SECTION
HYDROGASIFICATION
IN COCURRENT FLOW
OF GAS AND SOLIDS
FIRST-STAGE
HYDROGASIFI-
CATION
132
FEET
HOT GAS RISING
INTO FIRST-STAGE
\SECOND-STAGE
' HYDROGASIFt •
CATION
RISING GASES CONTACT
CHAR FOR FURTHER
HYDROGASIFICATION
HYDROGASIFIED CHAR
FROM SECOND-STAGE
FEEDS INTO STEAM-
OXYGEN GASIFIER
STEAM-OXYGEN'
GASIFIER I
STEAM ~:
OXYGEN—:
ASH
NOTE THIS SIMPLIFIED SK
IS NOT DRAWN 10 SCALE
Figure 5-1. 1GT Hygas pilot plant hydrogasification reactor'
70
-------
hydrogasification, hot gases at about 1700°F rising from the second stage of
hydrogasification react with the coal in concurrent flow. About 20 percent
of the coal is converted to methane in this stage which is operated at about
1200°F. In the second stage of hydrogasification, hydrogen reacts exothermi-
cally with the char to produce methane while steam reacts endothermically
with char to produce CO and more hydrogen. An additional 25 percent of the
coal is converted in this stage. The hot char descends to the final dense
phase fluidized bed where the hydrogen-rich gas is produced in the presence
of steam and oxygen.
The Hygas-Oxygen process has been chosen for a detailed study at the
Wyoming, New Mexico and North Dakota sites. The detailed design is presented
only for the Wyoming site. The gasifier details using Wyoming subbituminous
coal have been supplied by IGT and the rest of the plant has been calculated
so as to minimize water consumption. The water equivalent hydrogen balance
at the Wyoming site, giving all process water streams, is shown on Table 5-1.
Table 5-2 shows the ultimate disposition of unrecovered heat from which the
cooling water requirements will be calculated in Section 10. The analysis of
the Wyoming subbituminous coal used is shown on Table 5-3.
5.2 MATERIAL BALANCE, WYOMING
The gas production train is shown on Figure 5-2. Stream compositions
are shown on Table 5-4 and the ash residue composition is shown on Table 5-3.
The streams into and out of the gasifier were supplied by IGT. This gasifier
information represents IGT's best, conservative estimate of the performance
of the Hygas reactor system using Wyoming coal. All other streams have been
calculated.
The raw off-gas contains materials other than fuel gas species. For
example, the oil made in this type of gasifier is expected to be approximate-
ly 85 percent toluene and 15 percent benzene with a small quantity of phenol.
The oil manufactured in the Hygas reactor is significantly lighter than most
other gasification systems. The oil made about equals the oil lost in puri-
fication or left in the product gas. No oil is sold.
Most of the nitrogen in the coal is converted to ammonia. The remainder
71
-------
TABLE 5-1. WATER EQUIVALENT HYDROGEN BALANCE FOR HYGAS PROCESS USING
WYOMING SUBBITUMINOUS COAL
Coal feed: 1,314,500 Ib/hr as-received subbituminous coal (19.9% moisture,
4.03% hydrogen), dried to 2% moisture as fed to gasifier.
Product gas: 250 x 10 SCF/day product gas (95.44 vol % methane,
1.72 vol % hydrogen) with a HHV of 968.6 Btu/SCF.
IN
Moisture in as-received coal
Water equivalent of hydrogen in as-received coal
Steam to gasifier
10 Ib/hr
261
477
1015
1753
OUT
Moisture from coal drying
Condensate in waste heat recovery phase separator
Water equivalent of hydrogen from acid gas removal
Condensate from methanator
Water equivalent of hydrogen in product gas
240
294
60
201
951
1746
72
-------
TABLE 5-2. ULTIMATE DISPOSITION OF UNRECOVERED HEAT IN HYGAS PLANT
USING WYOMING COAL
Basis: 250 x 1Q6 SCF/day (10.09 x 1Q9 Btu/hr)
10 Btu/hr 10 Btu/hr
Coal drying 0.26
Heat in hot condensate water 0.01
Electricity used + slurry pump dissipation 0.10
Boiler stack loss 0.20
Combustibles lost in purification 0.80
Subtotal - Direct Loss 1-37
*
Air cooling of plant process stream 0,55
*
Wet cooling of plant process stream 0.23
Hot ash from gasifier 0.32
Bottom ash quench from boiler 0.003
Total steam turbine condensers 0.65
Total compressor interstage cooling 0.16
Acid gas removal regenerator condenser 0.80
4.08
* The heat assumed for water treatment and other uses (Table 5-7)
is assumed distributed 50% to dry cooling and 50% to wet cooling.
73
-------
TABLE 5-3. COAL AND ASH RESIDUE FOR HYGAS PLANT AT WYOMING SITE
Wyodak coal, wt. %
Stream No. in Fig. 5-2 (l)
C
H
N
S
O
Ash
Moisture
As -Received
54.2
4.0
0.8
0.6
14.5
6.0
19.9
100
Fed to Gasifier
66.32
4.93
0.93
0.74
17.74
7.34
2.00
100
Ash Residue
17.80
0.19
2.01
0.38
-
79.62
-
100
HHV (calculated) 9264 Btu/lb
2722 Btu/lb
74
-------
OXYGEN STEAM
Ul
MOISTURE
RECYCLE
WASH OIL
COAL
IOO*F PRODUCT GAS
»••
BFW
Ficaire 5-2. Flow diagram for Hygas process.
-------
TABLE 5-4. STREAM COMPOSITIONS FOR HYGAS PLANT USING WYOMING COAL
(SEE FIGURE 6-2)
Basis: 250 x 106 SCF/day (10.09 x 109 Btu/hr)
STREAMS OF GASIFIER 3
10 Ib/hr
sl^ As-received coal 1,315
-,2_/ Vaporized moisture 240
<3\ Dried coal in slurry 1,075
\3_) Slurry oil 1,075
Composition: C,H, : 10 wt. %
b b
C-,HD: 85 wt. %
/ o
C0,C_, and Clri aromatics : 5 wt. %
b y 10
PROCESS
CO
C°2
H2
CH4
C H
2 6
Others
H2°
C H
6 6
C H
7 8
) Oxygen
) Steam
Ash residue
STREAMS (10 moles/hr)
5.34
5.02
6.55
3.57
0.27
0,20
6.71
0.40
2.78
-------
appears as cyanide or free nitrogen gas. However, pilot plant experience
indicates that thiocyanate is found in effluent water streams to the exclu-
sion of cyanide even though thermodynamics dictates against formation of
thiocyanate in reducing atmospheres. Also, the appearance of free nitrogen
is not verified by pilot plant experience since all line purging is done
with nitrogen.
Sulfur in the coal is primarily reduced to H_S. The quantity of COS
reported in the off-gas was based upon thermodynamic expectations within the
reactor; these same considerations indicate that CS_ will not be formed, nor
will mercaptans or thiophenes. With lignite coal containing one percent sul-
fur as feedstock, pilot plant data verify the absence of CS2 and thiophenes;
however, trace quantities of methyl mereaptan have been detected.
As in most Hygas process flow sheets, the gasifier product gas is
quenched with oil to about 400°F to cool the gas and remove particulate mat-
ter. We find that no condensation of water or oil occurs.
A portion of the gas next undergoes shift reaction at an equilibrium
temperature of 750°F to adjust the ratio of hydrogen to CO for the downstream
methanation reaction. The shifted gas is cooled to 100°F to ensure condensa-
tion of the oil. Water also condenses at this point. A circulating water
scrub may be used to ensure that all the ammonia, phenol and other soluble
species are removed from the gas. It has been assumed that these species can
be adequately removed by the quantity of water which condenses. Circulating
water has not been shown on Figure 5-2.
A physical-solvent based system, such as the Selexol process, is used
for acid-gas removal to recover the remainder of the BTX stream, dehydrate
the gas, generate an H_S-rich gas for sulfur recovery, discharge a CO,,-rich
gas with minimum H.S concentration, and provide a treated gas of sufficient
purity that only a nominal sulfur guard is required prior to methanation.
Based on the recommendation of IGT, the following losses are assumed to occur
in gas purification: 0.5% loss of H_ and CO, 1% loss of CH and 25% loss of
C0H . The process is assumed capable of reducing CO_ to 1%. All other acid
ft D 2
gases were completely absorbed. In view of the relative solubilities of
14 •• -
various gases in Selexol solvent, these assumptions are reasonable.
77
-------
TABLE 5-5. HYGAS GASIFIER HEAT BALANCE USING WYOMING COAL
6 9
Basis: 250 x 10 SCF/day (10.09 x 10 Btu/hr)
IN
Coal
Steam (1250 psia, 1000°F)
Sensible heat of coal slurry
Sensible heat of oxygen feed
GOT
Gas
Hot ash
Oil made and sensible heat
109 Btu/hr
5.3 HEAT BALANCE, WYOMING
The gasifier information supplied by IGT is in both material and thermal
balance. Approximate gasifier heat balance is shown on Table 5-5. The heat
balance was then extended to the complete gas production train. The heat
exchangers and waste heat recoveries were individually calculated. The bal-
ance is shown on Table 5-6. The heating value of ash residue shown is the
sum of the higher heating value and sensible heat.
The driving energy required for the plant is shown on Table 5-7. In
making Table 5-7 the following rules were used:
(1) To dry coal the water must be evaporated and coal heated to 230°F.
The excess sensible heat in coal is contributed to the coal slurry heater in
raising the recycle oil to 200°F.
(2) The coal slurry pump operates at 70 percent efficiency and requires
3200 kw.
78
-------
TABLE 5-6. HYGAS PRODUCTION TRAIN HEAT BALANCE USING WYOMING COAL
Basis: 250 x 106 SCF/day (10.09 x 109 Btu/hr)
IN
Coal
Steam
Heat for coal drying
Heater for coal slurry
Sensible heat of oxygen
10 Btu/hr
12.20
1.48
.08
.06
.02
13.84
OUT
Product gas
Steam produced
Combustibles lost in gas purification
Sensible heat of condensate
Dry cooling of process streams
Wet cooling of process streams
Hot ash
10.09
2.12
0.80
0.01
0.41
0.09
0,32
13.84
79
-------
TABLE 5-7. HYGAS PLANT USING WYOMING COAL: DRIVING ENERGY
Basis: 250 x 1Q6 SCF/day (10.09 x 109 Btu/hr)
109 Btu/hr
Coal drying 0.33
Slurry pump 0.04
Recycle oil heater 0.06
Oxygen heater 0.01
Gas purification 0.80
Gasifier steam 1.48
Oxygen production 0.56
Electrical production 0.32
Low temperature heat for water treatment and
other uses (arbitrary) 0.29
Driving energy required 3.89
Less steam raised in process 2.12
1.77
Boiler stack loss 0.20
Net heat required from additional fuel 1.97
80
-------
(3) Gas purification was by the Selexol type process as discussed in
Section 8 (steam requirement is 28,400 Btu/lb mole CO,, absorbed).
(4) The energy for oxygen production is the energy to compress air to
90 psia and oxygen from 15 psia to 1250 psia.
9
The steam raised in the process can all be used, and 1.97 x 10 Btu/hr
of additional fuel must be burned to drive the plant. As shown on Table 5-8,
9
the plant conversion efficiency is about 71.2 percent and 4.08 x 10 Btu/hr
low level heat is unrecovered. To determine the cooling water requirement
the ultimate disposition of the unrecovered heat was then investigated.
5.4 ULTIMATE DISPOSITION OF UNRECOVERED HEAT, WYOMING
The ultimate disposition of unrecovered heat is shown on Table 5-2.
Prom the table, using the discussion of Section 10, the cooling water require-
ment will be calculated. Table 5-2 is taken directly from Tables 5-5, 5-6
and 5-7.
TABLE 5-8. HYGAS PLANT USING WYOMING COAL; OVERALL HEAT BALANCE
6 9
Basis: 250 x 10 SCP/day (10.09 x 10 Btu/hr)
10 Btu/hr
IN
Coal to gasification
Coal to boiler
OUT
Product gas
Unrecovered heat
Plant conversion efficiency
71.2%
81
-------
5.5 WATER FOR FLUE GAS DESULFURIZATION
When as-received Wyodak coal of the composition shown on Table 5-3 is
burnt, it produces about 1.3 Ib SO per 10 Btu. Burning North Dakota lig-
nite of the composition shown on Table 5-9 may result in about 1.7 Ib SO_ per
c *
10 Btu. Burning New Mexico subbituminous coal of the composition shown on
Table 5-9 may result in about 1.5 Ib SO per 10 Btu. We assume the flue
gas will be desulfurized. From formulae (3) and (4) of Section 8.2, 0.49 Ib
water will be consumed per 1 Ib of Wyoming coal; 0.19 Ib water will be con-
sumed per 1 Ib of North Dakota coal; and 0.53 Ib water will be consumed per
1 Ib of New Mexico ,coal. The Wyoming site burns about 212,700 Ib/hr coal and
consumes 104,000 Ib water/hr. At North Dakota, about 356,100 Ib/hr coal are
burnt consuming 67,700 Ib water/hr. At New Mexico about 204,600 Ib/hr coal
are burnt consuming 108,000 Ib water/hr.
It is possible to gasify the coal and to remove H S rather than to
remove SO2 from the flue gas. The production of fuel gas, including neces-
sary water evaporated for cooling, consumes very little water and the water
required for flue gas desulfurization can be saved. However, the loss in
efficiency and added cost are high and we have assumed that fuel gas is not
produced.
5.6 HYGAS PROCESS AT NEW MEXICO AND NORTH DAKOTA
For this study it was not possible to ask the Institute of Gas Technology
to repeat their gasifier balances for these sites. The results have, there-
fore, been approximated. For a first approximation the gasifier balances for
Wyoming will serve for these coals if the carbon plus hydrogen (i.e., the
heating value of the coal) fed to the gasifier is unaltered.
Approximate analyses of New Mexico subbituminous coal and North Dakota
lignite are shown on Table 5-9. Table 5-10 shows the coal and water process
streams at New Mexico and North Dakota plants. The quantities of ash resi-
dues from the gasifier were calculated from the ash content in coal at each
site assuming that the composition is that shown on Table 5-3. Also shown on
Table 5-10 is the estimated plant driving energy requirements at two sites.
82
-------
TABLE 5-9. APPROXIMATE-ANALYSES OP NEW MEXICO SUBBITUMINOUS COAL
AND NORTH
DAKOTA LIGNITE
New Mexico subbituminous coal
Coal composition, wt
C
H
O
N
S
Ash
Moisture
HHV (calculated)
% As-received
53.6
3.9
11.1.
0.9
0.7
13.5
16.3
9382 Btu/lb
Fed to gasifier
62.8
4.6
13.0
1.0
0.8
15.8
2.0
North Dakota lignite
Coal composition, wt
C
H
0
N
S
Ash
Moisture
HHV (calculated)
% As-received
42.5
2.9
13.4
0.7
0.6
5.1
34.8
6965 Btu/lb
Fed to gasifier
63.9
4.4
20.1
1.0
0.9
7.7
2.0
83
-------
TABLE 5-10. PROCESS COAL AND WATER STREAMS, AND PLANT DRIVING ENERGY
REQUIREMENTS AT NEW MEXICO AND NORTH DAKOTA
Basis: 250 x 1Q6 SCF/day (10.09 x 109 Btu/hr)
10 Ib/hr
As-received coal
Vaporized moisture from coal drier
2% moisture coal fed to gasifier
Steam to gasifier
Foul cohdensate
Clean condensate from methanator
Ash residue from gasifier
109 Btu/hr
Coal drying energy
*
Other plant driving energy
Total plant driving energy required
Less steam raised in process
Boiler stack loss
Bottom ash from boiler
Net heat required from additional fuel
New Mexico
1,300
190
1,110
1,015
294
201
220
North Dakota
1,752
586
1,166
1,015
294
201
112
0.28
3.56
3.84
2.12
1.72
0.19
0.005
0.79
3.56
4.35
2.12
2.23
0.25
0.003
1.92
2.48
* Taken from the driving energy required in Table 5-7, excluding
the energy for coal drying.
84
-------
These results are similar to the Wyoming plant (Table 5-7) with the excep-
tions of coal drying energy and boiler stack loss which were calculated for
each site according to the moisture content of the as-received coal. Table
5-11 shows the thermal efficiency estimates based on the performance at the
Wyoming site (see Table 5-8). Coal drying accounts for the major changes in
efficiency. Also shown on Table 5-11 is the ultimate disposition of unrecov-
ered heat at both sites. The results at the Wyoming site (Table 5-2) were
used wherever applicable. Only coal drying, boiler stack loss and bottom ash
from the gasifier and boiler were calculated for each site and the rest were
assumed to be unchanged from site to site. Slight imbalances resulted at
both New Mexico and North Dakota sites due to the simplifying assumptions.
85
-------
TABLE 5-11. APPROXIMATE THERMAL EFFICIENCY AND ULTIMATE DISPOSITION
OF UNRECOVERED HEAT AT NEW MEXICO AND NORTH DAKOTA
Basis: 250 x 10 SCF/day (10.09 x 109 Btu/hr)
9
10 Btu/hr New Mexico
Coal to gasification 12.20
Coal to boiler 1.92
Total energy input 14.12
Less product gas HHV 10.09
Unrecovered heat 4.03
Approximate conversion efficiency 71.5%
Disposition of unrecovered heat
Coal drying 0.21
Heat in hot condensate 0.01
Electricity used plus slurry
pump dissipation 0.10
Boiler stack loss 0.19
Combustibles lost in purification 0.80
Total Direct Loss 1.31
Air cooling of plant process stream 0.55
Wet cooling of plant process stream 0.23
Hot ash from gasifier 0.72
Bottom ash quench from boiler 0.005
Total steam turbine condensers 0.65
Total compressor interstage cooling 0.16
Acid gas removal regenerator condenser 0.80
4.43
Unaccounted energy dissipation (0.40)
4.03
North Dakota
12.20
2.48
14.68
10.09
4.59
68.7%
0.66
0.01
0.10
0.25
0.80
1.82
0.55
0.23
0.36
0.003
0.65
0.16
0.80
4.57
(0.02)
4.59
86
-------
REFERENCES SECTION 5
Tsaros, C. L., Knabel, S. J. and Sheridan, L. A., "Process Design and
Cost Estimate for Production of 265 Million SCF/Day of Pipeline Gas
by the Hydrogasification of Bituminous Coal," OCR R&D Report No. 22 -
Interim Report No. 1, 1965.
Tsaros, C. L., Knabel, S. J. and Sheridan, L. A., "Process Design and
Cost Estimate for the Production of 266 Million SCF/Day of Pipeline
Gas by the Hydrogasification of Bituminous Coal-Hydrogen by the
Steam-Iron Process," OCR R&D Report No. 22 - Interim Report No. 2, 1966.
Knabel, S. J. and Tsaros, C. L., "Process Design and Cost Estimate for
a 258 Billion Btu/Day Pipeline Gas Plant-Hydrogasification Using
Synthesis Gas Generated by Electrothermal Gasification of Spent Char,"
OCR R&D Report No. 22 - Interim Report No. 3, 1967.
Tsaros, C. L., Arora, J. L., Lee, B. S., Pimental, L. S., Olson, O. P.
and Schora, F. C., "Cost Estimate of a 500 Billion Btu/Day Pipeline Gas
Plant via Hydrogasification and Electrothermal Gasification of Lignite,"
OCR R&D Report No. 22 - Interim Report No. 4, 1968.
Schora, F. C. and Matthews, C. W., "Analysis of a HYGAS Coal Gasification
Plant Design," Presented at AIChE 65th Annual Meeting, New York,
November 27-30, 1972.
Schora, F., Jr., Lee, B. S. and Huebler, J., "The HYGAS Process," 12th
World Gas Conference and Exhibition, Nice, France, June 5-9, 1973.
Tarman, P. B., "Status of the STEAM-IRON Process," Proceedings of Fifth
Synthetic Pipeline Gas Symposium, AGA, OCR, and IGU, Chicago, Illinois,
Oct. 29-31, 1973.
Lee, B. S., "Status of HYGAS Process - Operating Results," Proceedings
of Fifth Synthetic Pipeline Gas Symposium, AGA, OCR and IGU, Chicago,
Illinois, Oct. 29-31, 1973.
Lee, B. S. and Lau, F. S., "Results from HYGAS Development," Presented
at AIChE 77th National Meeting, Pittsburgh, Pa., June 2-5, 1974.
Lee, B. S. and Tarman, P. B., "Status of the Hygas Program," Presented
at the Sixth Synthetic Pipeline Gas Symposium, American Gas Assoc.
October 28, 1974.
87
-------
11. Lee, B. S., "Slurry Feeding of Coal to HYGAS Gasifier," Presented at
AIChE 67th Annual Meeting, Washington, D,C.f December l-5f 1974.
12. Blair, W. G., Leppin, D. and Lee, A. L. , "Design and Operation of
Catalytic Methanation in the HYGAS Pilot Plant," in Methanation of
Synthesis Gasf ed. by L. Seglin, Advances in Chemistry Series No. 146,
American Chemical Society, Washington, D.C., 1975.
13. Lee, B. S., "Status of the HYGAS Program," Presented at 7th Synthetic
Pipeline Gas Symposium, Chicago, 111., October 27-29, 1975.
14. Valentine, J. P., "Gas Purification with Selexol Solvent in the New
Clean Ene'rgy Processes," Presented at Div. of Industrial and Engineering
Chemistry, ACS 167th National Meeting, Los Angeles, April 1974.
88
-------
SECTION 6
SYNTHANE PROCESS
6-l INTRODUCTION AND SUMMARY OF RESULTS
The Synthane process for the production of high-Btu pipeline gas is
x • 1—9
being developed by the Bureau of Mines. A laboratory scale gasifier has
been in operation at Bruceton, Pennsylvania. It is a 4-inch diameter tube
inside a 10-inch diameter shell. A pilot plant capable of handling 75 tons
Per day of coal or lignite at 1,000 psig is being started up at Bruceton.'
The pilot plant gasifier is 3 ft (inside diameter) by about 90 ft high.2
Run-of-mine coal is crushed to minus 3/4-inch in hammer mills. The
crushed coal is further reduced in ball mills so that at least 70 percent
will pass through 200 mesh. The pulverized coal is fed to the pretreater
through lock hoppers. Coals which tend to cake when heated in a hydrogen
atmosphere are pretreated in a fluid bed pretreater by reacting at 750-800°P
with about 12.5 percent of the total steam and oxygen fed to gasification.
Coal overflows from the pretreater and enters the gasifier which is a single
stage fluid bed reactor operating between 1400 and 1800°F. Gas leaves
through an internal cyclone which returns solid fines to the fluidized bed
via a dip leg (see Figure 6-1).
About 30 percent of the carbon fed to the gasifier is not gasified. It
^s released as char and tar. Part of the char is burnt to provide power for
running the plant. We find that all of the char is not needed and so some is
Presumed to be sold.
The Synthane process has been chosen for a detailed study at the Wyoming
site only. The gasifier details have been taken from the Bureau of Mines
design and the rest of the plant has been calculated so as to be compatible
89
-------
COAL
LOCK
HOPPER
LOCK
HOPPER
FEED
HOPPER
STEAM S
OXYGEN
GAS TO VENTURI
SCRUBBER
QUENCH WAHR
STEAM TO SHIFT
REACTOR
Figure 6-1. Flow diagram for Synthane gasifier.
90
-------
with the design procedure used for the Hygas process and to minimize water
consumption. The water equivalent hydrogen balance, which gives all process
water streams, is shown on Table 6-1. Table 6-2 shows the ultimate disposi-
tion of unrecovered heat from which the cooling water requirements will be
calculated. The analysis of the coal used is shown on Table 6-3. The
details of the design from which Tables 6-1 and 6-2 were made will now be
given.
6.2 MATERIAL BALANCE
The gas production train is shown on Figure 6-2. stream compositions
are shown on Table 6-4. The streams into and out of the gasifier are taken
from Reference 5. The other streams have been calculated. The shift equili-
brium temperature was taken as 750°P as in the Hygas design.
The Bureau of Mines design shows water added to the circulating scrub-
ber water and recovered in Stream 5. This will dilute Stream 5 which is foul
water, it is not certain that this is necessary and it has not been done in
this design. If it is necessary/ water Streams 8 and 9 should be used for
this purpose.
The ash quench system shown in Figure 6-2 is from Reference 9, not Ref-
erence 5. A small amount of the steam needed for the shift reaction is pro-
duced from dirty water rather than from boiler feed water. The char tempera-
ture is assumed reduced from about 1700°F to about 800°F.
6-3 HEAT BALANCES
The gasifier balance is shown on Table 6-5. Some steam is raised in the
gasifier jacket.5 The energy contained in the char and tar were approximated
as follows: 1) the weight of char shown on Table 6-4 was given by the Bureau
c 9
°f Mines; 2) the composition of char was as scaled from Strakey et al to
fit the ash weight and is shown on Table 6-6; 3) the carbon and hydrogen bal-
ances on the gasifier were then used to find the carbon and hydrogen contents
°f the tar; 4) the heating values of the char and tar were calculated from
the composition. In fact, for our purpose, the distinction between char and
91
-------
TABLE 6-1. WATER EQUIVALENT HYDROGEN BALANCE FOR SYNTHANE PLANT
Coal feed: 1,918 x 10 Ib/hr of Wyodak seam subbituminous coal
(3.5 wt. % hydrogen and 19.9 wt. % moisture) as-received.
Dried to 4.3 wt. % moisture (heating value of 10,640 Btu/lb)
as fed to gasif iers .
Product pipeline gas: 250 x 10 SCF/day with heating value of
940 Btu/SCF (92.3 vol. % methane and 1.8 vol. % hydrogen).
10 3 Ib/hr
Moisture in coal 382
Water equiv. of hydrogen in coal 604
Steam to gasifier and shift converter 1167
2153
OUT
Water lost in drying coal 313
Foul condensate after scrubbing 516
Condensate after shift conversion 138
Condensate after acid gas removal 20
Clean methanation condensate 140
Water equiv. of hydrogen in by-products
and lost gas 87
Water equiv. of hydrogen in product gas 920
2134
92
-------
TABLE 6-2. ULTIMATE DISPOSITION OF UNRECOVERED HEAT
Basis: 250 x 10 SCF/day (9.79 x 10 Btu/hr)
Coal drying
Heat in hot condensate water
Losses around gasifier
Electricity used
Char boiler stack losses
Combustibles lost in purification
Subtotal - Direct Losses
*
Air cooling of plant process streams
*
Wet cooling of plant process streams
Bottom ash quench from char boiler
Total steam turbine condensers
Total compressor interstage cooling
Acid gas removal regenerator condenser
10 Btu/hr 10 Btu/hr
0.42
0.14
0.40
0.11
0.34
0.10
1.51
1.53
0.28
0.02
1.04
C.34
1.01
5.73
* The unaccounted for loss in the gas production train balance (Table 6-7)
and the heat assumed for water treatment and other uses (Table 6-8) is
assumed distributed 50% to dry cooling and 50% to wet cooling, for example,
the heat lost by air cooling of plant process streams is here taken to be
1.33 x 109 (from Table 6-7) + 0.11,x,109/2 (unaccounted loss from Table
6-7) + 0.3 x 109/2 (heat for water treatment and other uses from Table
6-8) .
93
-------
TABLE 6-3. COAL FOR SYNTHANE EXAMPLE
Wyoming coal, wt. %.
As-Received Fed to Gasifier
c
H
N
S
O
Ash
Moisture
54.0
3.5
0.8
0.6
14.1
7.1
19.9
100.0
64.5
4.1
1.0
0.8
16.8
8.5
4.3
100.0
HHV (Btu/lb) 10,640
94
-------
vo
Ul
COAL
330»f
DOWTHERM
STEAM HEAT
EXCHANGE*
GAS
CHAR
HMOPSIA
1
A
X
-0J
WATER
METHANATO*
1
•
7S2°F
_ 71«°F
—VOOVY,*-,
_:
'
GAS-GAS
HEAT
EXCHANGER
1«*
WOfSIA
WAHU
93SPSIA
-*-c.
CW-
tiorsiA
IOO°F
WATER
Figure 6-2. Flow diagram for Synthane process.
-------
TABLE 6-4. STREAM COMPOSITIONS FOR SYNTHANE PLANT (SEE FIGURE 6-2)
GASIFIER FEEDS
(l) Coal
(2) Steam
(_3/> Oxygen
10 3 Ib/hr
1,605
978
482
GASIFIER EFFLUENT
PROCESS
CO
co2
Hn
2
CH4
C-H,.
Char
STREAMS (10 moles/hr)
® © ® © <8> ® <
16.70 7.82
25.89 34.77
16.03 24.91
15.24 15.24
1.12 1.12
410
© © @>
-------
TABLE 6-5. SYNTHANE GASIFIER HEAT BALANCE
109 Btu/hr
IN
Coal 17.08
Steam (satd. at 1000 psia) 1.12
18.20
OUT
Gas 12.14
Low pressure steam raised in jacket 0.61
Char (before quench) 4.25
Tar 0*80
Unaccounted loss 0.40
18.20
TABLE 6-6. CHAR ANALYSIS AND FLUE GAS DESULFURIZATION WATER
wt. %
C 63.6
H 1.0
0 1.4
N 0.4
S 0.3
Ash 33.3
100.0
HHV (calculated) 9800 Btu/lb
97
-------
tar is not important.
A small loss was found around the gasifier. This was assumed to be
directly lost to the atmosphere.
The heat balance was then extended to the complete gas production train.
The heat exchangers and waste heat recovery were individually calculated.
The balance (Table 6-7) shows an additional small loss which, in Table 6-2,
was assigned 50 percent to dry cooling and 50 percent to wet cooling.
The driving energy required for the plant is shown on Table 6-8. In
making Table 6-8 the following rules were used:
(1) Coal must be heated to 220°F and the water evaporated to dry the
coal.
4
(2) The lock hopper compressors require 6,800 kw.
(3) The steam requirement for gas purification by the Benfield process
is 30,000 Btu per mole CO™ absorbed.
(4) The energy for oxygen production is the energy to compress air to
90 psia and oxygen from 15 psia to 1015 psia.
TABLE 6-7. SYNTHANE GAS PRODUCTION TRAIN HEAT BALANCE
109 Btu/hr
IN
Coal
Steam
OUT
Product gas
Char
Tar
Steam produced
Combustibles lost in gas purification
Sensible heat of condensate
Loss around gasifier
Dry cooling of process streams
Wet cooling of process streams
Unaccounted loss
98
-------
TABLE 6-8. SYNTHANE PLANT DRIVING ENERGY
109 Btu/hr
Coal drying 0.42
Lock hopper compressors 0.08
Gas purification 1.01
Process steam 1.43
Oxygen production 1.05
Electrical production (31,000 kw) 0.36
Low temperature heat for water treatment
and other uses 0.30
Driving energy required 4.65
Less steam produced in process (1.61)
Net heat required from fuel 3.04
Char fired boiler
Heat recovered 3.04
Stack loss 0.34
Hot bottom ash 0.02
Char feed to boiler 3.40
99
-------
All of the steam raised in the process can be usefully consumed. As
9
shown on Table 6-8, 3.5 x 10 Btu of char (or tar) must be burnt to drive
the plant. The remainder of the char and tar is available for sale. Thus,
as shown on Table 6-9, the plant conversion efficiency is about 66.4 percent
9
and 5.74 x 10 Btu/hr low level heat is unrecovered. To determine the cool-
ing water requirement the ultimate disposition of the unrecovered heat is
investigated.
6.4 ULTIMATE DISPOSITION OF UNRECOVEPED HEAT
The ultimate Disposition of unrecovered heat is shown on Table 6-2.
From this table, using the discussion of Section 10, the cooling water
requirement can be calculated. Table 6-2 is taken directly from Tables
6-5, 6-7 and 6-8.
6.5 WATER FOR FLUE GAS DESULFURIZATION
When char of the composition shown on Table 6-6 is burnt, an emission of
about 0.6 Ib SO- per 10 Btu will result. For this study it is assumed that
desulfurization will be required. From Formula (3) of Section 8.2, 0.8 Ib
water will be consumed per 1 Ib char. About 336,000 Ib/hr char are burnt
consuming 269,000 Ib water/hr.
100
-------
TABLE 6-9. SYNTHANE PLANT OVERALL HEAT BALANCE
g
10 Btu/hr
IN
Coal 17.08
OUT
Product gas 9.79
Char 4.16
Tar 0.80
(Less char feed to boiler) (3,40)
Unrecovered heat 5.73
17.08
Plant conversion efficiency 66.4%
101
-------
REFERENCES SECTION 6
1. U.S. Department of the Interior, "An Economic Evaluation of Fluidized
Gasification at 40 Atmospheres, Followed by Shift Conversion, Purifica-
tion, and Single-stage Tube Wall Methanation," Report No. 68-8
(Alternate), Bureau of Mines, Morgantown, W. Virginia, 1971.
2. U.S. Department of the Interior, "An Economic Evaluation of SYNTHANE
Gasification of Pittsburgh Seam Coal at 1000 psia Followed by Shift
Conversion, Purification, Single-stage Tube Wall Methanation and
Pollution Control," Report No. 74-31, Bureau of Mines, Morgantown,
W. Virginia, 1974.
3. Forney, A. J., Haynes, W. P., Elliott, J. J., Gasior, S. J., Johnson,
G. E., and Strakey, J. P., Jr., "The SYNTHANE Coal-to-Gas Process,"
IGT Symposium on Clean Fuels from Coal, Institute of Gas Technology,
Chicago, Illinois, Sept. 10-14, 1973.
4. U.S. Department of the Interior, "Synthane Gasification at 1,000 psia,
Followed by Shift Conversion, Purification, Single-stage Tube Wall
Methanation and Pollution Control. 250-Million-SCFD High-Btu Gas Plant.
Pittsburgh Seam Coal. An Economic Analysis," Report 75-2, Bureau of
Mines, Morgantown, W. Virginia, 1974.
5. U.S. Department of the Interior, "SYNTHANE Gasification at 1,000 psia,
Followed by Shift Conversion, Purification, Single-stage Tube Wall
Methanation and Pollution Control. 250-Million-SCFD High-Btu Gas Plant.
Wyodak Seam Coal. An Economic Analysis," Report 75-15, Bureau of Mines,
Morgantown, W. Virginia, 1974.
6. Gasior, S. J., Forney, A. J., Haynes, W. P. and Kenny, R. F., "Fluidized
Bed Gasification of Various Coals," Chemical Engineering Progress, 71,
No. 4, 89-92, 1975.
7. Lewis, R., "Coal Gasification: Some Engineering Problems," Chemical
Engineering Progressf 71, No. 4, 68-69, 1975.
8. Forney, A.J., Haynes, W. P., Gasior, S. J., Johnson, G. E, and Strakey,
J. P., Jr., "Analyses of Tars, Chars, Gases, and Water Found in
Effluents from the SYNTHANE Process," Symposium Proceedings; Environ-
mental Aspects of Fuel Conversion Technology (May 1974, St. Louis,
Missouri), Environmental Protection Agency, EPA-650/2-74-118, 1974.
102
-------
Strakey, J. P., Jr., Forney, A. J. and Haynes, W. P., "Effluent Treat-
ment and Its Cost for the Synthane Coal-to-S.N.G. Process" presented
at American Chemical Society 168th National Meeting, Atlantic City,
N. J., September, 1974; Division of Fuel Chemistry reprint Vol. 19,
No. 5, page 94.
103
-------
SECTION 7
SOLVENT REFINED COAL
7.1 INTRODUCTION AND SUMMARY OF RESULTS
The Solvent Refined Coal process was developed by the Pittsburgh and
Midway Coal Mining Co., and two pilot plants are in operation: at Fort Lewis,
Washington and Wilsonville, Alabama. Raw coal is dried, pulverized and
mixed with a coal-derived solvent boiling in the general range 550-800°F.
The coal-solvent slurry is pumped together with hydrogen to 1700-2500 psi
and heated to about 850°F. The coal structure is broken; a small amount of
the carbon reacts with hydrogen to yield light hydrocarbons and most of the
coal hydrocarbon dissolves. The solution is filtered to remove ash, pyrites
and unreacted char, and the solvent is separated for reuse by distillation
under vacuum. The Solvent Refined Coal so produced has a fusion temperature
in the range 350-450°F, has less than 0.5 percent ash and is low enough in
sulfur to give 0.65 to 1.1 Ib SO2/10 Btu.
The Solvent Refined Coal process (SRC) has been chosen for detailed
study. Since SRC is made by the action of hydrogen on coal, the production
of hydrogen must be part of the plant design. Hydrogen is not made in the
pilot plants, it is bought. We have not found an integrated plant design
suitable for our purposes and so have made our own.
The plants have been sized to yield 10,000 tons SRC/day, that is
3.2 x 10 Btu fuel/day at 16,000 Btu/lb. This is similar in size to a
50,000 bbl/day synthetic oil plant yielding about 3.2 x lo11 Btu fuel/day.
The plant produces more fuel than a standard SN6 plant yielding about
2.4 x 10 Btu/day. (Upon completion of the calculation it was found that
the plant in North Dakota would yield only 9,565 tons SRC/day because some
104
-------
of the product had to be burnt to drive the plant. For ease of understanding
and comparison between sites, the calculation tables were not altered. The
summary tables, however, have been scaled to 10,000 tons/day at all sites.
The basis is shown on every table.) For plants of this size the process
water streams are summarized on Table 7-1. The stream called "foul water
from the dissolving section" is as found in the material balance calculations.
These streams may in fact be bigger, as discussed in Section 7.4. Table 7-2
shows the ultimate disposition of unrecovered heat in the plants. From this
table and the discussion on cooling in Section 10 will be derived the cooling
requirements listed in Section 13.
7.2 DESIGN PROCEDURE
Most of the experimental work has been done on bituminous coals from
Pittsburgh, Kentucky and Illinois. Table 7-3 shows three coal analyses
and three average SRC analyses derived from these coals. Very little work
has been done on solvent refining a Baukol Nooan Mine/ North Dakota Lignite
and an Elkol Mine, Wyoming Subbituminous. These experiments were in a small
laboratory bench reactor; the solvent was not in balance and the analyses of
the SRC are only suggestive of what might be obtained on a large scale. How-
ever, the SRC derived from the Western coals seems very similar to that
derived from Eastern coals. In fact, the SRC does not differ importantly
from coal to coal and, based on the various references, the analysis shown
on Table 7-3 has been assumed for all three of the coals being considered
here.
An alternative process is under study, particularly as "Project Lignite"
g
at the University of North Dakota. In this process carbon monoxide or syn-
thesis gas (CO + H ) is used to dissolve the coal instead of hydrogen. Water
is used (with lignite this may be the coal moisture) and the shift gas reac-
tion, CO + H_0 -»• H, + CO- , occurs in the dissolver probably catalyzed by
222 g
coal mineral. It is this process which was studied by Ralph M. Parsons Co.
and Jahnig who, in addition, designed for one third of the SRC product to
be given additional hydro-treatment to further reduce sulfur. Jahnig's
study is not directly comparable to this design.
105
-------
TABLE 7-1. TOTAL PLANT PROCESS WATER STREAMS
Basis: 10,000 tons SRC/day at all sites.
Wyoming New Mexico North Dakota
106 gal gals per 10 gal gals per 106 gal gals per
per day 106 Btu* per day Ip6 Btu* per day 106 Btu
IN
Steam and boiler
feed water to:
Process
Acid gas removal
1.30
0.76
2.06
4.06
2.37
6.43
1.12
0.72
1.84
3.50
2.25
5.75
1.79
0.93
2.72
5.59
2.90
8.49
OUT
Foul water from
dissolving
section 0.61 1.91 0.42 1.31 0.78 2.43
Condensate from
acid gas
removal 0.76 2.37 0.72 2.25 0.93 2.90
Medium quality
condensate from
gasification
section 0.47 1.47 0.46 1.44 0.51 1.60
Clean condensate
from reformer
section 0.25 0.78 0.14 0.44 0.51 1.59
2.09 6.53 1.74 5.44 2.73 8.52
Net water
produced 0.03 0.10 (0.10) (0.31) 0.01 0.03
* Based on 13,340 Btu/hr as SRC only
106
-------
TABLE 7-2. ULTIMATE DISPOSITION OF UNRECOVERED HEAT
6
Basis: 10,000 tons SRC/day {13,340 x 10 Btu/hr) at all sites.
Units: 106 Btu/hr
Stack losses + coal drying
Losses around filter & dissolver
Sensible heat of SRC
Electricity used + slurry pump heat
dissipated
Other losses
Subtotal - Direct Losses
Wyoming
750
320
120
70
90
New
Mexico
670
360
120
80
80
North
Dakota
1210
360
130
80
110
1350
1310
1890
Air cooling of plant process streams
Wet cooling of plant process streams
+ other allowance*
Ash quench
Total drive turbine condenser losses
Total compressor interstage cooling
losses
Acid gas removal regenerator
condenser
860
210
70
670
150
270
810
180
150
640
140
250
1010
260
70
780
180
320
Error (10) (20) (40)
Total unrecovered heat 3,570 3,460 4,470
* The energy for wastewater treatment and other low-level uses, shown on
Table 7-13, has been distributed 50% to direct losses and 50% to wet
cooling.
107
-------
TABLE 7-3. ANALYSES OF COAL AND SOLVENT REFINED COAL
Analyses in wt.
% for dry materials . )
Pittsburgh6 Illinois6
C
H
N
0
S
Ash
HHV(Btu/lb)
C
H
N
O
S
Ash
HHV (Btu/lb)
Moisture as
received.
Coal SRC Coal
75.1 88.4 70.8
5.1 5.5 5.1
1.3 1.7 1.3
7.6 3.3 ' 8.7
2.6 0.9 3.2
8.2 0.1 10.8
16,000
Western Coals
New
Mexico Wyoming
63.9 67.7
4.7 5.0
1.1 1.0
13.2 18.1
0.9 0.8
16.2 7.4
11,180 11,580
16.3 19.9
SRC
87.1
5.6
1.6
4.6
0.9
0.1
16,000
•North Dakota
Lignite
65.2
4.4
1.1
20.5
1.0
7.8
10,620
30.6
4
Kentucky
Coal SRC
72.9 88.5
4.8 5.1
1.2 1.8
10.3 3.7
3.5 0.8
7.3 0.1
16,000
SRC
(assumed)
87.8
5.3
1.2
5.0
0.5*
0.2
16,000
* Somewhat lower when North Dakota lignite is used because less of the
sulfur in the coal is organic.
108
-------
The dissolving section of the plant, based mostly on the pilot plant
design, is shown in simplified form in Figures 7-1 and 7-2. The plant water
requirements were obtained as follows.
(1) From pilot plant results a set of rules was formulated which gives
the material balance around the dissolving section of the plant.
(2) The carbonaceous filter residue is gasified to produce as much
hydrogen as possible.
(3) The remaining hydrogen is produced by reforming some of the product
gas.
(4) Approximate heat balances have been made around the gasification
section, reforming section and dissolving section.
(5) The energy needed to drive the plant was estimated; mechanical
loads (compressors, electricity generation), acid gas removal, water treat-
went and other loads were estimated. It was assumed that the required energy,
not recovered in waste heat recovery units, would be obtained by burning pro-
duct gas or oil. The quantity of gas or oil sold could then be determined
and the approximate thermal efficiency of the plant stated.
(6) Finally, the points of loss of unrecovered heat have been tabulated
so that the cooling water requirement could be calculated in Section 10.
The calculations showed that North Dakota lignite has such a high mois-
ture and oxygen content there was not enough gas and oil to drive the pro-
Posed plant and some of the SRC must be burned. The addition of coal to the
gasifier to produce the extra driving energy was not studied, although this
should probably be done. In reading the following tables please note that
the basis is shown on every table.
7.3 MATERIAL BALANCE ON DISSOLVING SECTION
i
Product yields for Eastern coals are generally reported as fractions of
the moisture- and ash-free coal. Because of the high oxygen contents of
Western coals, yields are based on carbon. Based on the published experimen-
tal results, we have formulated the following rules for material balances
in the dissolving section of the plant.
(1) 70 percent of the carbon in the coal appears as carbon in the SRC
109
-------
COAL'
SSO'F
HP
FLASH
DRUM
Vx'"*^/
U '!
I* L
*
DECANTER
)
LP
FLASH
DRUM
V
00
[7>GAS
ACtD
GAS
REMOVAL
CW
SOLUTION OF
SRC
WATER
OIL TO CLEAN UP TO FILTER
Figure 7-1. SRC dissolving section—A.
-------
SRC
SOLUTION
FROM
SECTION A
I
FUEL
FILTER RESIDUE
STEAM \£j
CW
c
VENT
cw
CW
1
BFW
j450°F
1
/
rn
A 'TEAM.
V
\
VACUUM
A
V
PREHEATER
VACUUM
TOWER
<=x
L
DIRTY
CONDENSATE
_^ WASH SOLVENT
TO FILTER
t STEAM
/^\ 3900F
IBFW
RECYCLE SOLVENT
TO SLURRY BLEND
TANK
SRC
Figure 7-2. SRC dissolving section—B.
-------
and SRC has the composition shown in Table 7-3.
(2) 14 percent of the carbon in the coal appears as carbon in light
liquid hydrocarbon product of composition CH, ,.
-L. D
(3) 5 percent of the carbon in the coal appears as gaseous hydrocarbon
product of composition CH (about 75 percent CH and the balance higher
3 • / 4
hydrocarbons).
(4) 1 percent of the carbon in the coal appears as CO.,.
(5) 10 percent of the carbon in the coal appears as carbon in undis-
solved residue.
(6) The ratio 0/C in the undissolved residue is the same as in the coal.
The balance of the oxygen appears as water.
(7) All of the sulfate sulfur stays in the mineral residue. 50 percent
of the pyritic sulfur is reduced to H S and the balance appears in the ash.
60-70 percent of the organic sulfur is reduced to H_S and the balance is dis-
tributed between the SRC and undissolved residue.
(8) Nitrogen from the coal appears in the SRC and the undissolved resi-
due, with the balance appearing as ammonia. The ratio N/C is the same in the
coal and in the undissolved residue.
(9) Hydrogen is supplied as required and 10 percent of the feed hydro-
gen does not react.
Application of these rules gives the material balances presented on
Tables 7-4 through 7-6. On these tables stream numbers from Figures 7-1
and 7-2 have been entered. For Stream 2 only the hydrogen content has been
stated. In fact, the hydrogen streams produced by gasification and reforming
contain only about 85 percent hydrogen with CO, CO_ and CH.. These extra
gases are assumed to leave the dissolving section with the gas of Stream 7.
7.4 EFFLUENT WATER
The water stream shown on Tables 7-4 through 7-6 is possibly low for the
following reasons: 1) coal will probably still contain 0.5 percent moisture
when fed to the dissolver and this moisture may be recovered as dirty water;
2) these designs show between 58 percent and 69 percent of the oxygen in the
coal converted to water. Coals having a lower oxygen content may have up to
112
-------
Basis:
Units:
IN
TABLE 7-4. OVERALL MATERIAL BALANCE FOR DISSOLVING
SECTION OF THE PLANT, WYOMING COAL
10,000 tons SRC/day; (13,340 x 106 Btu/hr)
103 lb/hr
1• Coal, dry
2. Hydrogen
Total
H
N
TOTAL 1581 1045 115 15.4 279 12.3
Ash
1543 1045 77.2 15.4 279 12.3 114
38 38 '
114
OUT
% Coal
Total
H
N
Ash
3.
4.
5.
6.
7.
SRC
Filter Residue
Water
H2S
NH3
Light oil
Gaseous hydrocarbon
co2
Unconsumed hydrogen
54
16.5
13
0.5
0.3
11
4.5
2.5
0.3
834
255
204
7
5
166
68
38
4
732 44.2
104 7.7
22.7
0.4
0.8
146 19.5
52.3 16.1
10.5
3.8
10.0 41.7 4.2 1.7
1.5 28 1.5 112.3
181
6.6
3.9
27.9
TOTAL 102.6
1581
1045 115
15.4 279
12.3 114
* 6.4% sulfate, 22.6% pyritic, 71% organic.
113
-------
TABLE 7-5. OVERALL MATERIAL BALANCE FOR DISSOLVING SECTION
1.
2.
Basis: 10,000 tons
Units: 103 lbs/hr
Coal . dry
Hydrogen
OF THE PLANT, NEW MEXICO COAL
SRC/day (13,340 x 106 Btu/hr)
Total C H N 0
1635 1045 76.8 18,0 216
31 31.2
S Ash
*
14.7 265
TOTAL
1666
1045 108
18.0 216 14.7 265
% Coal
Total
C
H
N
TOTAL
102.0
1666 1045 108
18.0 216 14.7
Ash
3.
4.
5.
6.
7.
SRC
Filter Residue
Water
H2S
NH3
Light oil
Gaseous Hydrocarbon
co2
Unconsumed hydrogen
51
24
8
0
0
10
4
2
0
.5
.6
.5
.5
.2
.2
.3
.2
834
401
140
8
8
166
68
38
3
732 44.
104 7.
15.
0.
1.
146 19.
52.3 16.
10.5
3.
2 10.0 41.7 4.2
7 1.8 21.6 2.7
6 125
5 7.8
3 6.2
5
1
27.9
1
1.7
263.3
265
* Assume distributed as Wyoming coal, see footnote to Table 7-4.
114
-------
TABLE 7-6. OVERALL MATERIAL BALANCE FOR DISSOLVING SECTION
OF THE PLANT, NORTH DAKOTA LIGNITE
Basisi 9,565 tons SRC/day* (12,760 Btu/hr)
pnita: 103 Ib/hr
IN
1.
2.
Coal , dry
Hydrogen
TOTAL
OUT
3.
4.
5.
6.
7.
SRC
Filter Residue
Water
H^S
2
NH3
Light Oil
Gaseous hydrocarbon
co2
Unconsumed hydrogen
TOTAL
% Coal
51.
17.
15.
0.
0.
10.
4.
2.
0.
102.
3
1
9
6
4
4
2
4
3
6
Total
1603
52
. 1655
Total
833
274
255
9
7
166
68
38
5
1655
C H
1045 70.
52.
1045 122.
C H
732 44.
104 7.
28.
0.
1.
146 19.
52.3 16.
10.5
5.
1045 122.
N 0 S Ash
*
5 17.6 329 16.0 125
2
7 17.6 329 16.0 125
NO S Ash
2 10.0 41.7 3.5 1.7
7 1.8 32.9 3.9 123.3
3 226.5
5 8.6
2 5.8
5
1
27.9
2
7 17.6 329 16.0 125
* Of the SRC produced in the dissolving section some is burned as plant fuel.
** 3.0% sulfate, 42% pyritic, 55% organic
115
-------
103 Ib/hr
275
215
323
10 gal/day
0.8
0.6
0.9
85 percent of the oxygen converted to water. For these reasons, in sizing
the water treatment plant this particular water stream will be assumed to
have a maximum flow of:
Wyoming
New Mexico
North Dakota
The effluent water will contain about 0.8 percent organic carbon as
1 percent phenol (about 3 percent COD) or about 2.2 x 10 lb carbon/hr. The
material balances are limited to this accuracy.
The rate of production of ammonia is very sensitive to the nitrogen con-
tent of the SRC. The nitrogen content of SRC was estimated to be 1.2 percent.
This is lower than that found with Eastern coals and was chosen because these
Western coals have a lower nitrogen content. Should the choice be wrong and
should the nitrogen content of SRC be 1.6 percent, as with Eastern coals, the
ammonia production might be altered as follows:
10 Ib/hr ammonia
Wyoming
New Mexico
North Dakota
The water treatment scheme must take into account unknown and possibly vary-
ing ammonia concentrations.
The ammonia may be assumed to dissolve in the condensed water and leave
with it. Some H2S will also dissolve. The nitrogen and sulfur contents of
this water stream are measured daily at the Port Lewis pilot plant, and Pitts-
burgh and Midway kindly supplied 22 measurements taken between 10/5/75 and
12/9/75. The measured nitrogen contents averaged 1.3 percent with a standard
deviation of 0.7 percent. The measured ratio (moles NH /moles H S) was 2.0
•3 £»
116
1.2%
N in SRC
4.7
7.5
7.0
1.6% N in SRC
0.7
3.5
3.0
-------
with a standard deviation of 0.17. The nitrogen content of the water might
vary as suggested above. However, the assumption of a molar ratio (NH /H-S)
of 2 in the water will be satisfactory for water treatment design. Most of
the H S will probably dissolve and come out in the water stream/ but some
will always remain in :the gas stream.
7.5 GASIFICATION OF CARBONACEOUS FILTER RESIDUE
The filter residue is a fine powder with a high ash content. The
Koppers-Totzek gasifier has been taken to produce hydrogen from this material
because this gasifier has proven usefulness on related materials. This gasi-
fier is described in Section 4.10 where the rules used to obtain the gas com-
position leaving the gasifier are listed.* A simplified sketch of the equip-
ment is in Figures 7-3 and 7-4, and the stream compositions are shown on
Table 7-7. Streams 10, 11 and 12 follow from the gasifier rules used. The
differences between coals are small and are mostly due to changes in oxygen
content. The ash content has only a small effect on the material balance,
although the ash content does affect the thermal balance discussed below.
For Navajo coal, which gives about 66 percent ash in the residue, slag car-
ries away about 8 percent of the higher heating value of the residue.
3
Stream 14 was set at 15 x 10 moles/hr; the shift reactor exit gases
were calculated as though they were in equilibrium at 750°F (K = 11.8) and
so Streams 15 and 16 were recorded. The acid gas removal section was taken
to remove 90 percent of the CO_ fed to it.
The approximate heat balances are presented in Table 7-8. Minor quanti-
ties, such as the sensible heat of influent and effluent streams (other than
slag), are nearly in balance and have been omitted. The assignment of
*0xygen feed, Ib/lb (carbon + hydrogen) = 1.06
Steam feed, Ib/lb (carbon + hydrogen) = 0.223
Composition of gases leaving the reactor are consistent with a shift equili-
brium approximately
(H0)
-------
00
TO
SHIFT
REACTOR
SECTION ft
OXYGEN
PtANT
FILTER
RESIDUE
§FW
SYNGAS
COMPRESSOR
Figure 7-3. SRC hydrogen production by gasification of filter residue—A.
-------
VO
FROM
SECTION A
STEAM
600°F
STEAM
SHIFT
CONVERSION
REACTOR
BFW
7SO°F
SHIFT
INVERSION
REACTOR
750°F
STEAM
\^Jl8Q°F*\ lnnPE^
I *-^-^ y
BFW
cw
215PSIA
" '' 1 >
T ' I r^Jl^OOPSIA
ACID GAS ^|
REMOVAL I
TO DISSOLVING
SECTION
HYDROGEN
COMPRESSOR
Figure 7-4. SRC hydrogen production by gasification of filter residue—B.
-------
TABLE 7-7. STREAMS IN THE PRODUCTION OF HYDROGEN BY
Stream
No.
10 02
11 HO
12 Raw gas
CO
C°2
H2
H2°
13 H20
14 H20
15 H20
16 Product Gas
CO
co2
H2
16 H0
GASIFICATION OF FILTER
Units Wyoming*
10 3 Ib/hr 118
103 Ib/hr 24.8
10 moles/hr
7.80
0.86
4.18
1.00
103 Ib/hr 18
103 Ib/hr 270
103 Ib/hr 146
10 moles/hr
0.90
0.77
11.08
103 Ib/hr 22.16
RESIDUE
New
Mexico*
118
24.8
7.99
0.68
4.36
0.79
14.2
270
147
0.94
0.77
11.04
22.08
North
Dakota**
118
24.8
7.64
1.03
3.97
1.14
20.5
270
148
0.86
0.78
10.74
21.48
* Basis: 10,000 tons SRC/day (13,340 x 10 Btu/hr)
** Basis: 9,565 tons SRC/day (12,760 x 10 Btu/hr)
120
-------
TABLE 7-8. APPROXIMATE HEAT BALANCE FOR PRODUCTION OF
HYDROGEN BY GASIFICATION OF FILTER RESIDUE
Units: 10 Btu/hr
Stream
__No.
4 Filter residue
14 Steam
TOTAL IN
16 Hydrogen product
Total steam generated
Ash and slag
Dry cooling load
Wet cooling load
TOTAL OUT
Wyoming*
1780
350
2130
1470
400
70
110
30
2080
New
Mexico*
1830
350
2180
1470
390
150
110
30
2150
North
Dakota**
1700
350
2050
1430
420
70
110
30
2050
* Basis 10,000 tons SRC/day (13,340 x 10 Btu/hr)
** Basis 9,565 tons SRC/day (12,760 x 10 Btu/hr)
121
-------
process cooling requirements to air cooling and water cooling will be dis-
cussed in Section 10.
7.6 PRODUCTION OF HYDROGEN BY STEAM REFORMING OF GAS
The hydrogen produced by gasification of filter residue (Stream 16) is
not sufficient to supply the inlet to the dissolving section, Stream 2. The
extra hydrogen is made by reforming some of the gas produced in the dissolv-
ing section of the plant. The stream reforming reaction is mostly
CO
and it is followed by a shift reaction
CO
The reforming reaction absorbs a lot of heat and is carried out at a high
temperature in a direct-fired furnace. Equilibrium is reached in both the
reformer and the shift converter, and the compositions can be calculated if
the temperature is known. The temperature, however, depends on the degree
of reaction, and the heat and material balances are solved simultaneously.
The system shown on Figure 7-5 was solved by computer program with the
results shown on Tables 7-9 and 7-10. (A single gas composition was used
to serve for all three coals, although the gas composition is somewhat
dependent on the coal.) From these tables the important information for
the three coals has been tabulated as Table 7-11.
7.7 TOTAL PLANT PROCESS WATER STREAMS
The total plant process water streams are summarized on Table 7-1. In
making Table 7-1, 6,000 Ib steam/hr has been used for the vacuum ejectors of
12
the vacuum tower. The condensate is added to the foul condensate of Stream
5. Steam is not required for direct addition to the towtr. Th« steam
required for acid gas removal is assumed to be used live and is shown
122
-------
STEAM
ro
19 HYDROGEN
*•
ITOgPSIA
Figure 7-5. SRC hydrogen production by reforming.
-------
TABLE 7-9. HYDROGEN PRODUCTION BY REFORMING
Basis:
1000 moles/hr feed gas
A. Stream Flows
CO
co2
CH4
C2H6
H2
H20
TOTAL
Gas Condensate
Feed Steam Feed Water
(moles/hr) (moles/hr) (moles/hr)
601
107
292
2851 1643
1000 2851 1643
Hydrogen
Product
(moles/hr)
84
46
169
0
2686
2985
The H_ in the product = 5.37 x 10 Ib/hr.
B. Hydrogen Balance
IN
Gas feed
Steam feed
moles/hr
1815
2851
4666
OUT
Hydrogen product
Condensate water
3024
1643
4667
124
-------
TABLE 7-10. HEAT BALANCE ON REFORMING SECTION
Basis: 1,000 moles/hr feed gas
6
10 Btu/hr
IN
Reformer feed gas 337
Fuel burned I18
Steam feed "13
528
OUT
Hydrogen product 405
Steam produced 91
Slack loss 12
Air cooler load 1S
Water cooler load 4
528
125
-------
TABLE 7-11. SUMMARY OF REFORMING SECTION
Hydrogen required
Feed gas
Steam Feed
Condensate recovered
Reformer gas
Fuel
Steam produced
Air cooling load
Water cooler load
Units
10 3 Ib/hr
10 moles/hr
103 Ib/hr
103 Ib/hr
106 Btu/hr
106 Btu/hr
106 Btu/hr
106 Btu/hr
106 Btu/hr
Wyoming*
15.8
2.9
149
86
977
342
255
46
12
New
Mexico*
9.1
1.7
87
50
573
201
150
27
7
North
Dakota**
30.7
5.7
293
169
1921
673
501
91
23
6
Basis 10"~,"OW tons SRC/day (13,340 x 10 Btu/hr)
** Basis 9,656 tons SRC/day (12,760 x 10 Btu/hr)
126
-------
separately. The steam requirement is 26 Ib steam/mole CO . On Table 7-1 the
results for North Dakota have been scaled to 10,000 tons SRC/day.
7.8 HEAT BALANCE ON THE DISSOLVING SECTION
An approximate heat balance on the dissolving section is given on Table
7-12. In the dissolver, chemical reactions tend to cause a rise in tempera-
ture. This rise has been taken to be 50°F as shown on Figure 7-1. This is
among the higher temperature rises recorded in Reference 6. Even with this
rise a loss of energy around the dissolver was found and is recorded on Table
7-12. This loss may not occur; in any event, this energy will not go to eva-
porate cooling water. It is the cooling water requirement which is the objec-
tive of this study.
The additional rules used to calculate Table 7-12 are:
(1) The dissolver preheater fired duty was calculated on the assumption
that 2 Ib of solvent are mixed with each 1 Ib of dry coal to make the slurry.
Hydrogen is added and the whole mass is heated through 170°F. Of the fired
duty, 61 percent is used in the radiant section to heat the slurry, 29 per-
cent is recovered as steam and 10 percent goes up the stack. This may be
optimistic with regard to the steam generated.
(2) The flushed gas, water and oil are passed through an air cooler and
then a water cooler in series. The gas is cooled from 550 to 100°F. The
water and oil are condensed. Most of the heat removed comes from condensing
the water in the flash stream, and most of this removal occurs in the air
cooler. Twenty percent of the entire cooling load has been assigned to wet
cooling.
(3) A loss has been assumed around the filter corresponding to a 100°F
drop in temperature.
(4) The vacuum preheater is designed to heat the vacuum tower feed
through 100°F and to vaporize all the solvent assuming 141 Btu/lb for a
latent heat. As with the dissolver preheater, the radiant duty is 61 per-
cent of the fired duty, 29 percent is recovered as steam and 10 percent
goes up the stack. All of the vaporized solvent must be condensed, and
this can be done by an air cooler.
127
-------
TABLE 7-12. APPROXIMATE HEAT BALANCE ON DISSOLVER SECTION
Units: 10 Btu/hr
Stream No.
1 Coal , dry
2 Hydrogen & carbon monoxide
Fuel to dissolver & vacuum
preheaters
TOTAL IN
3 SRC
4 Filter residue
6 Oil
7 Gas (hydrocarbon + H + CO)
Steam recovered
Stack losses
Dry cooling load
Wet cooling load
Losses around filter
Losses around dissolver
Sensible heat in SRC
TOTAL OUT
Wyoming*
17,870
2,640
1,720
22,230
13,340
1,780
2,860
2,170
800
170
700
70
230
90
120
22,330
fi
New
Mexico*
18,280
2,160
1,810
22,250
13,340
1,830
2,860
2,080
840
180
670
50
250
110
120
22,330
North
Dakota**
17,030
3,740
1,800
22,570
13,340
1,700
2,860
2,460
840
180
770
80
240
100
120
22,690
* Basis 10,000 tons SRC/day (13,340 x 10 Btu/hr)
** Basis 9,565 tons SRC/day (12,760 x 10 Btu/hr)
128
-------
(5) The recycled solvent is cooled from 550 to 390°F in a waste heat
recovery boiler.
(6) The SRC is assumed recovered as a liquid at 550°F.
7.9 PLANT DRIVING ENERGY
The auxiliary energy to drive the plant is shown on Table 7-13 and the
following notes apply.
(1) The energy to dry the coal is the energy needed to vaporize the
moisture and heat the coal to 220°F.
(2) The energy for acid gas removal is discussed in Section 8. It has
been taken to be 28,400 Btu/mole CO .
(3) The electric load is 10,000 kw. The heat rate for the generator
and all turbines in the plant was taken as 11,700 BtuAw-hr, as discussed
in Section 8.
(4) The energy needed to produce oxygen is the energy required to com-
Press the equivalent weight of air to 90 psia.
7.10 THERMAL EFFICIENCY
The thermal efficiency is calculated on Tables 7-14 and 7-15.
7.11 ULTIMATE DISPOSITION OF WASTE HEAT
The ultimate disposition of waste heat is presented on Table 7-2 so that
cooling water requirements can be estimated in Sections 10 and 13. About 40
Percent of the unrecovered heat is lost to the atmosphere without heat trans-
fer. About 24 percent of the unrecovered heat will be lost in dry coolers in
the plant. The justification for the use of dry coolers is given in Section
10 and has to do with the high temperatures of SRC plant process streams.
Whether wet or dry cooling is to be used for the other points of cooling will
be discussed in Section 10 and summed up in Section 13.
129
-------
TABLE 7-13. PLANT DRIVING ENERGY
Units: 10 Btu/hr
Coal drying
Acid gas removal (3 places)
Vacuum tower ejectors
Mechanical loads
Electricity
Oxygen plant
Syn. gas compressor-
Hydrogen compressors
(2 places)
Slurry pump
SUBTOTAL - Mechanical loads
Wyoming*
490
260
10
120
180
260
280
110
New
Mexico*
420
250
10
120
180
260
230
120
North
Dakota**
850
320
10
120
180
260
390
120
950
910
1070
Water treatment plus
allowance for other low
level energy and losses
at 10%T
TOTAL
170
160
1880
1750
220
2470
* Basis 10,000 tons SRC/day (13,340 x 10 Btu/hr)
** Basis 9,565 tons SRC/day (12,760 x 106 Btu/hr)
t On Table 7-2 this energy is distributed 50% to wet cooling and 50%
to losses.
130
-------
TABLE 7-14. PLANT FUEL REQUIREMENTS
.6
Units: io6 Btu/hr
EiHLlL driving energy required
Net steam produced in gasification section
Net steam produced in reforming section
Net steam produced in dissolving section
Total steam produced
Fuel for driving energy
puel for reformer
Fuel for dissolver preheater
Fuel for vacuum preheater
Total Fuel
Wyoming
1880
50
50
800
900
1030*
340
680
1040
3090
New
Mexico
1750
40
30
840
910
880*
200
720
1090
2890
North
Dakota
2470
70
100
840
1010
1520*
670
730
1070
3970
To that part of the fuel used for steam production or superheating
has been added 10% for stack losses. This has not been added to the
fuel used for coal drying.
131
-------
TABLE 7-15. PLANT CONVERSION EFFICIENCY CALCULATION
Units : 10 Btu/hr
Wyoming
Coal feed 17,870
SRC 13,340
Oil produced 2,860
Gas + hydrogen produced 2,170
Gas reformed (980)
Fuel burned in plant (3,090)
Total product fuel 14,300
New
Mexico
18,280
13,340
2,860
2,080
(570)
(2,890)
14,820
North
Dakota
17,030
13,340
2,860
2,460
(1,920)
(3,970)
12,750
Conversion efficiency
80%
81%
75%
132
-------
REFERENCES SECTION 7
1. Hydrocarbon Research, Inc., "Solvent Refining Illinois No. 6 and Pitts-
burgh No. 8 Coals," Electric Power Research Institute, Palo Alto, Calif.,
Report EPRI 389, June 1975.
2. Southern Services, Inc., "Status of Wilsonville Solvent Refined Coal
Pilot Plant," Electric Power Research Institute, Palo Alto, Calif.,
Report EPRI 1234, May 1975.
3. Anderson, R. P. and Wright, C. H., Pittsburg and Midway Coal Mining Co.,
"Development of a Process for Producing an Ashless, Low-Sulfur Fuel from
Coal, Vol. II - Laboratory Studies, Part 3 - Continuous Reactor Experi-
ments Using Petroleum Derived Solvent," May 1975. U.S. E.R.D.A. Res. &
Dev. Report No. 53, Interim Report No. 8 NTIS Catalog FE-496-T1.
4. Schmid, B. K., "The Solvent Refined Coal Process," presented at the
Symp. on Coal Gasification and Liquefaction, Univ. of Pittsburgh,
August 1974.
5. Anderson, R. P., "Evolution of Steady State Process Solvent in the
Pittsburg and Midway Solvent Refined Coal Process," presented at
Symp. on Coal Processing, A.I.Ch. E., Salt Lake City, August 1974.
6. Catalytic Inc., for Southern Services Inc., "SRC Technical Report No. 5,
Analysis of Runs 19 through 40, 20 January to 8 August 1974, Wilsonville,
Alabama," unpublished report.
7. Wright, C. H., et al., "Development of a Process for Producing an Ashless,
Low-Sulfur Fuel from Coal, Vol. II - Laboratory Studies, Part 2 - Con-
tinuous Reactor Studies Using Anthracene Oil Solvent," U.S. E.R.D.A.
Res. & Deve. Report No. 53, Interim Report No. 7, September 1975, NTIS
Catalogue FE-496-T4.
8. University of North Dakota "Project Lignite - Process Development for
Solvent Refined Lignite," U.S. E.R.D.A. Report 106, Interim Report No. 1,
1974. NTIS Catalogue FE-1224-T1.
9. The Ralph M. Parsons Company, "Demonstration Plant, Clean Boiler Fuels
from Coal, Preliminary Design/Capital Cost Estimate," U.S. Dept. of the
, Interior, O.C.R., R&D Report No. 82, Interim Report No. 1, Volume II,
1975.
133
-------
10. Jahnig, C. E., "Evaluation of Pollution Control in Fossil Fuel Conver-
sion Processes: Liquefication: Section 2. SRC Process," Report
EPA—50/2-74-009-f, U.S. Environmental Protection Agency, Research
Triangle Park, N. D., March 1975.
11. The Pittsburg and Midway Coal Mining Company, "Development of a Process
for Producing an Ashless, Low-Sulfur Fuel from Coal, Vol. Ill - Pilot
Plant Development Work, Part 2 - Construction of Pilot Plant," May
1975, U.S. E.R.D.A. Res. & Dev. Report No. 53, Interim Report No. 9,
NTIS Catalogue FE-496-T2.
12. Nelson, W. L., Petroleum Refinery Engineering, 4th ed., pp. 252-262,
McGraw-Hill, 1958.
134
-------
SECTION 8
OTHER PROCESS WATER NEEDS
8«1 GAS PURIFICATION
The removal of hydrogen sulfide and carbon dioxide from gases is an
important consumer of energy in coal conversion plants. In this section
selected processes for gas purification will be briefly discussed with the
object of understanding the energy consumption, the cooling water require-
ment and any influent and effluent water streams required.
Most of the sulfur in the crude gas is present as hydrogen sulfide (H S),
However, small but important amounts of carbonyl sulfide (COS) and carbon
Bisulfide (CS ) as well as traces of other organics including thiophene and
Nercaptans are often encountered. The concentration of sulfur in the crude
gas depends on both the coal and gasification process and may vary from negli-
gible to about 3 percent by volume. Acceptable concentrations in the product
gas are generally about 10 ppm for low and medium BTU gas and less than about
2 ppm for pipeline gas production involving a nickel catalyzed methanation
step.
Carbon dioxide (C0~), formed in the gasifier and in the shift converter,
is present in concentrations from more than 20 percent for oxygen-blown pro-
cesses down to about 2 percent for hydrogen-blown or indirectly heated gasi-
fiers. The latter concentration may be taken as the upper limit for pipeline
Duality gas. Carbon dioxide should not be removed from fuel gas intended,
for example, for combined cycle electric generation, and its partial removal
with H s is an efficiency penalty in the plant.
135
-------
Acid Gas Removal in Pipeline Gas Production
When making pipeline gas the CO and H S must be removed. These acid
gases may be removed by adsorption into an alkaline solution. One solution
widely accepted is hot potassium carbonate. The approximate chemistry of
adsorption is
K CO + CO + HO ** 2KHCO
£•$£•£ J
K CO + H S ^ KHCO + KHS
Adsorption takes place at about 270 to 300°F and at 1000 psi. The solution
saturated with CO- and H S can be regenerated for reuse by reducing the pres-
sure and boiling with little change in temperature. Regeneration reverses
the chemical equations shown above. The hot potassium carbonate process is
well documented. The drawings of the Synthane plant show "Benfield type"
gas purification. The operating conditions are
Regenerator steam 30,000 Btu/lb mole CO .
Regenerator design Closed steam boiler, condensate recovered.
Regenerator condenser Air cooled (see Section 10). This is being
practiced.
Liquid pumping energy 0.26 kw-hr/lb mole CO- after use of recovery
turbine. This is included in the plant elec-
trical load.
Physical adsorbents are also used for acid gas removal. The gases dis-
solve in the cold solvent under pressure and are driven off by release of
pressure and steam stripping. Some examples of physical adsorption processes
are the Rectisol processes used in all Lurgi plants, the Fluor process, '
the Sulfinol processes and the Selexol process. ' The Selexol pro-
cess is widely used in designs and has been used for our design of the Hygas
plant. Information of the sort needed is not published, and it is difficult
to generalize because of the interaction between flow scheme, equipment
136
-------
provided and corresponding utility requirements in any particular process
application. However, the following typical acid gas process design require-
ments were estimated for employment in general evaluation work.
Feed gas temperature
Regenerator steam
Regenerator design
Regenerator condenser
Liquid pumping energy
100°F.
28,400 Btu/lb mole CO .
In the SRC plant live steam is used and the
recovered water is sent to treatment. In Hygas
the load on treatment would be too high. The
condensed water is assumed to be reboiled in a
closed steam boiler. This increases the capi-
tal cost.
Air cooled with trim water cooled heat exchanger
to reduce gas temperature to maintain water bal-
ance (see Section 10).
0.3 kw-hr/lb mole CO- after use of recovery of
£»
turbine. This is included in the plant electri-
cal load.
Hot Gas Desulfurization
Increased attention is being given to the use of solids for high tempera-
19
ture bulk HLS removal and increased overall generating plant efficiencies.
the high temperature systems under investigation are calcined dolomite
Panel bed filters and fly ash/iron oxide absorbents.
In the calcined dolomite system the dolomite may be in the form of a
20
Panel bed filter downstream of the gasifier. These filters could operate
at temperatures of 1,300 to 1,500°F and at pressures of about 300 psi and
are expected to achieve 99 percent sulfur removal. The filter may be
Designed to simultaneously remove fume and dust. The absorbent is regen-
erated by reacting it with steam and C02 to yield HjS at a concentration
137
-------
sufficient for the Claus process:
CaS'MgO + HO + CO ** CaCO -MgO + H S
Theoretical steam requirements are 0.56 Ibs H O/lb S removed. This water is
released in the panel bed filter (reverse of the above equation) and will be
carried up the stack with the combustion products. Design and operation con-
siderations are discussed in Reference 20.
The Bureau of Mines is investigating the capabilities of several solid
^ i *y ^N» o 7
adsorbents in the temperative range 1,000 to 1,500°F. ' They obtained
H^S removal efficiencies of 95 percent with sintered pellets of 75 percent
fly ash, 25 percent Fe_O_. An adsorption capacity of 0.04 to 0.06 Ibs sul-
fur per pound of mixed absorbent was obtained, and no deterioration was
detected over extended absorption-regeneration cycles. Regeneration with
oxygen yields a concentrated stream of SO which could be used for sulfuric
acid production.
8.2 FLUE GAS DESULFURIZATION
In the Hygas plant coal may be burnt for energy to drive the plant. In
the Synthane plant char will be burnt. In both of these plants flue gas
desulfurization is assumed to be needed. A wet desulfurization process has
been selected for the design and the following procedure is used to estimate
the water requirements.
In wet scrubbers, water leaves the system as vapor in the flue gas and
with the waste solids as slurry.
Water in Saturated Flue Gas
The water content of saturated flue gas depends on the temperature and
pressure of saturation and is given on Table 8-1.
From Table 8-1 it is clear that lack of knowledge of the pressure of sat-
uration in the range of interest will alter the water content of saturated
gas by not more than 7 percent. However, if the temperature of saturation
138
-------
TABLE 8-1. MOLES WATER/MOLE DRY GAS AT SATURATION
Total pressure, (psia)
(inch water gauge)
Temperature
120 49
130 54
140 60
150 66
160 71
14 . 696
0
0.130
0.178
0.245
0.339
0.476
15.057
10
0.127
0.173
0.237
0.328
0.460
15.418-
20
0.123
0.168
0.231
0.318
0.444
139
-------
is 10°F higher than was assumedr the water content of the flue gas will be
about 40 percent higher than was calculated.
Most experimental scrubbers equilibrate at an exit gas temperature of
125°F. However , to best match published water consumption the conditions
assumed are 120°F (49°C) and 10 inches of water gauge so that the flue gas
contains 0.13 moles water per mole of dry gas. To find the actual water eva-
porated, the flue gas volume must be calculated.
Assuming negligible carbon monoxide and nitrogen oxides the flue gas vol-
ume is given by the formula derived on Table 8-2. The total moles of flue
gas per unit weight of coal, as fired, are
4.76 (1 + a)-+ + (3.76 + 4.76a) - - (1)
The makeup water requirement for gas saturation, which is the water leaving
in the flue gas minus the water entering in the flue gas, is in Ib/lb coal
as fired
4.76(0.13) (1 + a) -- + -^ (18)
+ (0.13) (3.76 + 4.76a) - - (18) - w - (2)
The excess air a varies from about 0.05 to 0.2 and 0.15 has been used.
If this is introduced into Formula (2) , it becomes
Ib makeup water/lb coal - 12.8 (-^ + ~} + 10.5 ( — - •—) - w - — (3)
Table 8-3 is a numerical examination of Formula (3) . The biggest determinant
is the moisture in the coal, w .
Now let us consider water lost with the waste solids.
Water in Waste Solids
The rate of water lost with waste solids depends on the sulfur quantity
and on the slurry concentration. There are two major solid products:
140
-------
TABLE 8-2. DETERMINATION OF FLUE GAS VOLUME
Basis: Unit weight of coal containing:
Element Wt. Formula Moles
Carbon
Hydrogen
Oxygen
Sulfur
Moisture
w
2
32
H2°
c/12
h/2
x/32
s/32
w/18
The fraction excess air is a.
Flue gas component
Carbon dioxide, CO,
Water,
Sulfur dioxide, SO,
Moles in dry gas
c/12
s/32
Moles O- required
c/12
(h/4)-(x/32)
s/32
Oxygen,
Nitrogen,
. c h_ x s
at!2 + 4 " 32 32'
\ / j.
a)(12 + 4 " 32
s \
32*
Total moles
4.76(1 -f a) ( + ) + (3.76 + 4.76a) ( - -~)
141
-------
TABLE 8-3. CALCULATED MAKEUP WATER FOR FLUE GAS SATURATION (SEE TEXT FOR DETAILS)
N5
1.
2.
3.
4.
5.
6.
7.
1.
2.
3.
4.
5.
6.
7.
Coal
Wet li-gnite
Medium wet lignite
Dried lignite
Subbi tuminous
Bituminous
Bituminous
Bituminous
TABLE 8-4.
Coal
Wet lignite
Medium wet lignite
Dried lignite
Subbituminous
Bituminous
Bituminous
Bituminous
c h
0.45 0.03
0.51 0.034
0.58 0.039
0.54 0.04
0.7 0.05
0.7 0.05
0.7 0.05
CALCULATED MAKEUP WATER FOR
lb water/lb coal
in solids in flue
0.05 0.21
0.06 0.38
0.07 0.55
0.08 0.49
0.35 0.82
0.26 0.81
0.17 0.80
X
0.15
0.17
0.19
0.11
0.07
0.07
0.07
WASTE DISPOSAL
s
.006
.007
.008
.009
.04
.03
.02
AND TOTAL
total in solids
0.26 148
0.44
0.62
0.57
1.17
1.07
0.97
152
152
168
546
412
274
lb H20/
w lb coal
.3 0.21
.2 0.38
.1 0.55
.15 0.49
.05 0.82
.05 0.81
.05 0.80
WATER
gpm/106kw
in flue total
570 720
920 1070
1180 1330
1040 1210
1290 1840
1270 1680
1250 1520
-------
Ib solid/lb sulfur Ib water/lb sulfur
CaS04'2H20 5.38 1.13
CaS03«*5H20 4.03 0.28
Assuming a given solids concentration in the waste slurry, the weight of
slurry water per unit weight of sulfur can be calculated. For example, for
40 wt percent solids concentration
Solids composition Ib slurry water/lb sulfur
CaSO *2H O 8.07
Casey *jH2o 6.05
Vacuum filtration could increase the percent solids to 60 percent and vacuum
filtration plus centrifugation could raise this figure to 70 percent. How-
ever, these mechanical dewatering techniques would only be practiced if trans-
Port of the sludge to distant landfill were envisaged, which is not the case
for the Western sites considered in this study.
Tables 8-5a and 8-5b give the compositions in weight fractions of lime
24
and limestone scrubber sludges. (In Reference 24 the calcium sulfite was
assumed to be in the crystalline form CaSO *2H O. It is now generally agreed
•3 ^
that the form in the sludges is CaSO •'jH O. We have therefore corrected the
O £t
weight fractions in Reference 25.) Also shown in these tables are the cor-
^esponding weight fractions of sulfur and water of hydration, calculated
using the values given above.
From the data of Tables 8-5 the weight of solids and of the water of
hydration per unit weight of sulfur for the lime and limestone processes
have been calculated. These figures are
Process Ib solid/lb sulfur Ib water/lb sulfur
Lime 5.2 0.38
Limestone 6.6 0.38
143
-------
TABLE 8-5a WEIGHT OF COMPONENTS OF LIME SLUDGE
(DRY) AND CORRESPONDING WEIGHTS OF
SULFUR AND WATER OF HYDRATION
Component Ib solid Ib sulfur Ib water
CaCO., ' .058 0 0
CaSO -1/2 HO .690 .171 .048
CaSO -2 HO .126 .023 .026
Ca(OH), .126 0 0
2
Total 1.000 .194 .074
TABLE 8-5b WEIGHT OF COMPONENTS OF LIMESTONE
~~SLUDGE (DRY) AND CORRESPONDING
WEIGHTS OF SULFUR AND WATER OF HYDRATION
Component Ib solid Ib sulfur Ib water
CaCO., .367 0 0
CaSO..-1/2 HO .533 .132 .037
J t*
CaSO "2H O .100 .019 .021
4 2
Ca(OH)2 0 0 °
Total 1.000 .151 .058
144
-------
The unit weights of solids are reasonably close and with little error
the average of the two values is representative of any wet lime, limestone
or liire- limes tone scrubbing process, that is, 5.9 Ib solid/lb sulfur. The
water of hydration can represent only a very small fraction of the total
makeup water (slurry water plus water of hydration) so that this contribu-
tion is neglected and the makeup water equals the slurry water. In this case
Ib makeup water _ /
Ib sulfur ~ * '
1 - m
m
(4)
where m is the weight fraction of solids in the waste (i.e., weight of
solids divided by the weight of water plus solids). Note that a change to
30 percent solids from 40 percent solids makes a 50 percent increase in the
water in the waste since (1 - m)/m changes from 1.5 to 2.3. A change to
50 percent solids makes (1 - m) equal to 1 and decreases the water in the
waste by 33 percent.
TABLE 8-6. REPORTED AND ESTIMATED WATER REQUIREMENT FOR FGD
Quantities in gpm/10 kw generation.
Station % S in Coal
EPA/TVA
Shawnee
22
22
EPA/TCA
Cholla, Ariz.
Public Serv.22
Kaiparowits
Reference 23
Calculated
Mohave, Nevada,
S. cal. Edison28
3.4
3.4
0.5
0.4
0.4
Water
in Stack
900
850
430
816
926
870
Water
in Solids
510
810
195*
58
55
47 to 71
Total
Water
1410
1660
630
870
981
- 930
*Scaled to 40% solids
145
-------
Reported Experience
The limited reported experience is given on Table 8-6. In the case of
Kaiparowits, the composition of the coal was given and a calculated result is
also presented on Table 8-6. To compare Table 8-6 with Table 8-4, the units
Ib water/lb coal have been converted to gpm/10 kw as follows. First the
coal heating value is calculated from the Dulong formula:
HHV(higher heating value) = 14,540c + 62,000(h-x/8) + 4,100s Btu/lb
Then, taking 10,000 BtuAw-hr so that coal feed is 10 /HHV in Ib/hr, the
makeup water for flue gas saturation is
Ib water x 10 Ib coal x 1 gal/min _ Ib water 2x10 .
Ib coal HHV hr 500 Ib/hr ~ Ib coal HHV 9al/min
The result of this calculation is shown on Table 8-4.
Comparison of Table 8-6 with Table 8-4 suggests that the estimates are
high for the flue gas even though a low saturation temperature has been
assumed. Nevertheless the procedure seems to yield useful estimates of
water quantity.
8.3 WATER FROM DRYING COAL
In the Synthane, Hygas and SRC processes the coal must be dried before
use. The question arises: can the water be recovered? Water recovered by
drying coal is quite clean as it has been vaporized and condensed under con-
ditions such that few volatile contaminants accompany it. For an example of
quantity, the water removed when North Dakota lignite is dried for feed to
the SRC plant amounts to 707 X 10 Ib/hr = 2 X 10 gallons/day.
We have investigated recovering water from drying coal using several dif-
ferent techniques and find the cost of the recovered water to lie in the
range $1.30 to $1.50 per thousand gallons. In this report we have not assumed
that water will be recovered, but recovery is certainly a serious possibility
when water is particularly scarce.
146
-------
REFERENCES SECTION 8
1. Colton, C. B., Dandavati, M. S. and May, V. B., "LOW and Intermediate
Btu Fuel Gas Cleanup" presented at EPA Symposium on the Environmental
Aspects of Fuel Conversion Technology II, Hollywood, Florida, December
1975.
2. Robson, F. L., Giramonti, A. J., Blecher, W. A. and Muzzella, G.,
"Fuel Gas Environmental Impact, Phase Report," EPA, Research Triangle
Park, N. C., Report EPA-600/2-75-078, November 1975.
3. Riesenfeld, F. C. and Kohl, A. L., Gas Purification, 2nd Edition,
Gulf Publishing Company, Houston, 1974.
4. Maddox, R. H., Gas and Liquid Sweetening, 2nd Edition, John M. Campbell
Co., Norman, Okla., 1974.
5. Maddox, R. N. and Burns, M. D., "Lease Gas Sweetening," 15 articles,
Oil and Gas Journal, Aug. 14, 21, Sept. 18, Oct. 2, 9, Nov. 13, 1967,
Jan. 8, 22, March 11, May 13, June 3, 17, Aug. 12, Sept. 9, 1968.
6. Ahner, D. J., Shelon, R. C., Garrity, J. J. and Kasper, S., "Economics
of Power Generation from Coal Gasification Combined-Cycle Power Plants,"
page 10, presented at American Power Conference, Chicago, April 1975.
7. Field, J. H., Benson, H. E., Johnson, G. E., Tosh, J. S. and Forney,
A. J., "Pilot Plant Studies of the Hot-Carbonate Process for Removing
Carbon Dioxide and Hydrogen Sulfide," U.S. Bureau of Mines Bulletin
No. 597, 1962.
8. Eickmeyer and Associates, Prairie Village, Kansas, "CATACARB, the
Catalytic Process for CO and H S removal."
9. Benfield Corp., Pittsburgh, Pa., "The BENFIELD Process."
10. Parrish, R. W. and Neilson, H. B. (Benfield Corp.,), "Synthesis Gas
Purification Including Removal of Trace Contaminants by the BENFIELD
Process," presented at the 167th National Meeting, ACS Div. of Ind.
and Eng. Chem., Los Angeles, April 1974.
11 ii
11. Lurgi Gesellschaft fur Warme-und Chemotechnik mbh, "Purification of
Industrial Gases, Waste Gases and Exhaust Air."
147
-------
12. Kohl, A. L. and Buckingham, P. A., "The Fluor solvent CC>2 removal
process," Oil and Gas Journal 58, No. 19, 146, May 9, 1960.
13. Buckingham, P. A., "Fluor solvent process plants: How they are working,"
Hydrocarbon Processing and Petroleum Refiner 43, No. 4, 113-116, 1964.
14. Klein, J. P., "Developments in Sulfinol and ADIP Processes Increase
Use," Oil and Gas International 10, No. 9, 109-112, Sept. 1970.
15. Goar, B. C., "Sulfinol Process has Several Key Advantages," Oil and Gas
Journal, 117-120, June 30, 1969.
16. Fisch, E. J., Hill, E. S. and Van Scoy, R. W., (Shell Development Co.),
"Sulfinol Process for Gas Treating," Paper 72-D-7, pp. D-88 to 91,
Operating Section Proceedings, Am. Gas. Assoc. Distribution Conference,
Atlanta, Georgia, 1972.
17. Fisch, E. J., Hill, E. S. and Van Scoy, R. W. , "Sulfinol Process for
Gas Treating," presented at American Gas Association, Atlanta, May 1972.
18. Valentine, J. P., "Gas Purification with Selexol Solvent in the New Clean
Energy Processes," presented at Div. of Industrial and Engineering
Chemistry, ACS 167th National Meeting, Los Angeles, April 1974.
19. Ashworth, R. A. and Switzer, G. W., "Low BTU gasification: High tempera-
ture - low temperature l^S removal. Comparison effect on overall
thermal efficiency in a combined cycle plant," Office of Coal Research,
U.S. Department of the Interior RSD No. 79 Interim #1, Sept. 1973.
20. Ruth, L. A., Graff, R. A. and Squires, A.M., "Desulfurization of fuels
with calcined dolomite." Presented at the 71st National Meeting AIChE,
Dallas, Texas, Feb. 1972.
21. Abel, W. T., Schultz, F. G. and Langdon, P. F., "Removal of H2S from hot
producer gas by solid absorbents." Bureau of Mines; U.S. Department of
the Interior RI 7947, 1974.
22. Bornstein, L. J. and others, "Reuse of Power Plant Desulfurization
Wastewater," pages 61-63, EPA Report 600/2-76-024, 1976.
23. Bureau of Land Management, "Final Environmental Impact Statement Propose^
Kaiparowits Project," Vol. I, pp. 1-89, 105, FES 76-12, March 3, 1976.
24. Cooper, H. B., "The Ultimate Disposal of Ash and Other Solids from Elect*'
Power Generation," in Water Management by the Electric Power Industry.
(E. F. Gloyna, H. H, Woodson and H. R. Drew, editors), pp. 183-195,
Center for Research in Water Resources, The University of Texas at
1975.
148
-------
25. Oldaker, E. C., Poston, A. M., and Farrior, W. L. , "Removal of
Hydrogen Sulfide from Hot Low Btu Gas with Iron Oxide - Fly Ash
Sorbents," Report No. MERC/TPR-75/1, 1975.
26. Oldaker, E. C., Poston, A. M., and Farrior, W. L., "Hydrogen Sulfide
Removal from Hot Producer Gas with a Solid Fly Ash Iron Oxide
Absorbent," Report No. MERC/TPR-75/2, 1975.
27. Farrier, W. L., Poston, A. M., and Oldaker, E. C., "Regenerable Iron
Oxide Silica Sorbent for the Removal of H2S from Hot Producer Gas,"
Paper presented at Fourth Energy Resources Conference, University of
Kentucky, January 1976.
28. weir, A., et al, "Results of the 170MW Test Modules Program, Mohave
Generating Station, Southern California Edison Company," p. 342,
Proceedings: Symposium on Flue Gas Desulfurization, New Orleans, La.,
March 1976, Report No. EPA-600/2-76-136a, U.S.EPA, Research Triangle
Park, N.C.
149
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SECTION 9
POWER GENERATION VIA COAL GASIFICATION IN
COMBINED-CYCLE POWER PLANTS
9.1 INTRODUCTION AND CONCLUSIONS
In this section the water requirements for a plant to convert coal energy
to electricity are discussed. The water requirements using gasification and
combined cycle generation are shown to be less than the water requirements in
a standard coal fired steam-electric generating plant.
There is not yet a reason to choose combined-cycle technology because
this technology today is less efficient than standard plants. However, gasi-
fication followed by combined-cycle generation is expected to improve in effi-
ciency when designs already started are complete, and this new technology may
be used in the next generation of coal fired electric plants.
For the three coals shown on Table 9-1 a gasifier and combined-cycle gen-
erating plant have been designed using cold gas sulfur removal. The resultant
plant water-equivalent hydrogen balances are shown on Table 9-2 which gives
the process water streams. The plant overall energy balances, from which
cooling water requirements can be determined, are shown on Table 9-3. First
it is interesting to compare Tables 9-2 and 9-3 with the corresponding results
for a direct fired coal burning steam-electric plant.
9.2 COMPARISON WITH A COAL BURNING STEAM-ELECTRIC GENERATING PLANT
Table 9-4 shows a comparison between the usual energy balance for a 1000
megawatt coal fired generating plant and the Wyoming combined-cycle plant.
The condenser cooling load is about 130 percent of the load on wet cooling
150
-------
TABLE 9-1. COALS USED IN ANALYSIS OF COMBINED-CYCLE PLANTS
Basis: 100 Ib dry coal
Wyoming
67.7 67.7
New Mexico
64.0
64.0
North Dakota
Dry As fired Dry As fired Dry As fired
65.2 65.2
H
5.0
8.4
4.7
6.9
4.4 10.4
O
18.1 44.9
13.2
30.4
20.5 67.9
N
1.0
1.0
1.1
1.1
1.1 1.1
0.8
0.8
0.9
0.9
1.0 1.0
Ash 7.4
7.4
16.1
16.1
7.8 7.8
100.0 130.2 100.0 119.4 100.0 153.4
HHV
(Btu/lb) 11,948
Moisture
(%) 23.2
11,660
16.3
11,267
34.8
151
-------
TABLE 9-2. COMBINED-CYCLE ELECTRICAL PLANT
WATER EQUIVALENT HYDROGEN BALANCES
Basis: 10 kilowatts
Wyoming
10 Ib/hr
New Mexico
North Dakota
Coal, (H2 + H20)
753
638
1003
Boiler feed water(1) 957
TOTAL IN
1710
970
1608
943
1946
As NH.
17
26
20
As hydrogen in tar,
oil, naphtha and
phenol
65
60
73
Up the stack
908
897
867
Dirty condensate(2) 720
625
986
TOTAL OUT
1710
1608
1946
Net Consumption
(1) - (2)
237
345
(-43)
152
-------
TABLE 9-3. COMBINED-CYCLE ELECTRIC PLANT
OVERALL ENERGY BALANCES
Basis: 10 kilowatts
Wyoming
10 Btu/hr
New Mexico
North Dakota
Coal
12.00
11.99
12.17
Electricity 3.41
Tar, oil 1.85
Subtotal fuel
5.26
3.41
1.72
5.13
3.41
1.98
5.39
Byproducts (NH3/S) 0.13
Stack, other loss,
dry cooling, etc. 3.05
Subtotal-losses
not requiring
wet cooling
3.18
0.19
3.15
3.34
0.10
2.85
3.01
Process streams t
cooled by water
Ash
Steam turbine
condenser
0.63
0.06
0.61
0.14
0.70
0.07
2.85
2.75
2.97
TOTAL OUT
11.98
11.97
12.14
* Arbitrarily includes waste heat load - thus this is points 9, 21 and 23 of
Figure 9-3.
153
-------
TABLE 9-4. COMPARISON OF ASSUMED ENERGY BALANCE FOR A 1,000 MEGAWATT
STEAM-ELECTRIC COAL FIRED GENERATING PLANT, WITH A COMBINED
CYCLE GENERATING PLANT
Flue gas reheating and other
plant losses
Stack losses, byproducts and
dry cooling load
Turbine condenser and other
wet cooling
Combined Cycle Plant
(Wyoming example
Coal Fired Plant from Table 9-3)
Coal
Electricity
Combustible byproducts
Stack losses
109 Btu/hr %
9.74 100
3.41 35
0
0.97 10
109 Btu/hr
12.00
3.41
1.85
%
100
28
15
0.68
4.68
48
9.74 100
3.18
27
3.54 30
11.98 100
TABLE 9-5. COMPARISON OF WATER FOR FLUE GAS DESULFURI2ATION IN A 1,000
MEGAWATT COAL FIRED STEAM-ELECTRIC GENERATING PLANT, WITH NET
WATER FOR A COMBINED CYCLE PLANT (FROM TABLE 9-2)
10 Ib/hr North
Wyoming New Mexico Dakota
Coal Fired Plant
Coal fired (as received)
Makeup water for flue gas desulfurization
975
458
919
505
1,221
366
Combined Cycle Plant
Net water consumed
237
345
(-43)
154
-------
in a combined-cycle plant. Table 9-5 shows the water required for flue gas
desulfurization calculated by the method of Section 8.2 compared to the net
water consumption in combined-cycle plants. The water requirement is more
than the net water consumed in the gasifier of a combined-cycle plant. The
coal burning plant is more efficient and the comparison is incomplete, but a
combined-cycle plant will use less process water and less cooling water than
a direct-fired plant.
9.3 DESCRIPTION OF COMBINED-CYCLE GENERATION
Figures 9-1 and 9-2 show two simplified versions of coal gasification
and combined-cycle electric generating plants. They differ in the tempera-
ture at which hydrogen sulfide is removed from the low-Btu power gas. The
incentive to build a combined-cycle system is that today it offers an effi-
ciency approaching that of a coal burning steam-electric power plant, with
the probability of higher efficiencies in the near future, while at the same
time employing reliable methods of preventing emissions of sulfur compounds.
In the combined-cycle concept, desulfurized coal gas is burned under
pressure using compressed air to support combustion. The pressurized hot
combustion products are expanded through a gas turbine which drives a power
generator and the compressor supplying air for fuel gas combustion and coal
gasification. A pressure of about 250 psig is required at the turbine inlet,
and a Lurgi pressurized gasifier has been chosen for detailed study. As dis-
cussed below, air blown Lurgi gasifiers are used and the discussions of Lurgi
for SNG and Lurgi slagging are not directly applicable.
An additional decision is whether to use oxygen or air in the gasifier.
One advantage of air is that the nitrogen reduces the need for excess steam
to moderate the temperature at the bottom of the gasifier. Although the pre-
sence of nitrogen increases the volume of gas passing through the purification
system, it is not otherwise detrimental and nitrogen will be added with excess
air in the combustor if it is not already present. Studies suggest that gasi-
fication with air is more efficient for use in this application than gasifica-
tion with oxygen, and air has been chosen for this study.
The remarks about air apply to a dry ash discharge. If the gasifier is
155
-------
U1
COAL
STACK •*-
STEAM CONDENSATE
WASTE HEAT
BOILER
BOILER FEED
WATER
STEAM
COMPRESSORS
CONDENSER
Figure 9-1. Gasification.and combined cycle generation with cold gas desulfurization.
-------
COAL
STACK •*•
AIR
DESULFURIZATION
REGENERATION
BOILER FEED
WATER
STEAM
COMPRESSORS
GH2H3
TURBINE I
1 AIR
STEAM TURBINE
'O
CONDENSER
Figure 9-2. Gasification and combined cycle generation with hot gas desulfurization.
-------
operated in a slagging mode, oxygen will be used. The throughput will
increase three- to fourfold and the reduced use of steam will increase the
efficiency. This will more than compensate for any disadvantages that oxy-
gen may have.
Pressurized hot combustion products from burning purified coal gas are
expanded through a gas turbine. As the gas passes through the turbine, it
cools because some of its contained heat energy is converted to mechanical
work. However, the gas leaving the turbine is at sufficiently high tempera-
ture for raising high pressure superheated steam used to generate electricity
through a steam-driven turbine. Steam is exhausted from the steam turbine at
a low temperature. The temperature of the gas leaving the gas turbine depends
on the gas turbine inlet temperature and on the pressure expansion ratio. The
capacity of the exhaust gas for producing high pressure superheated steam and
the combined-cycle power generating efficiency are a function of gas turbine
design.
At a given gas turbine inlet temperature and pressure, increasing the
gas turbine pressure expansion ratio will tend to reduce the gas turbine exit
temperature and the level of high pressure steam raised. The efficiency of
the steam power cycle is reduced as steam pressure and superheat temperature
fall below about 1000°F superheat. However, the efficiency is not important-
ly increased if steam is produced above about 1000°F. Generally, station
efficiency is maximized by using the maximum practical gas turbine inlet tem-
perature and by increasing the pressure expansion ratio as the inlet tempera-
ture increases so as to produce 1000°F superheated steam.
With increasing gas turbine inlet temperatures, less excess air is
required to establish combustor exit temperature at the desired level and
the air compression requirements are reduced. Station efficiency tends to
be increased even though the efficiency of the gas turbine itself may
decrease as more turbine cooling air is required at higher inlet tempera-
ture levels.
Increasing the gas turbine inlet temperature tends to shift the combined-
cycle generating capacity towards the gas turbine, which does not require a
cooled condenser, and away from the steam cycle which does require a cooled
condenser. This lessens the need for cooling water.
158
-------
2
According to information published by General Electric, gas turbine
design is expected to develop as follows:
Current state-of-the-art gas turbines operate on low-Btu gas
at 1950°F inlet temperature and a pressure ratio of 10:1, producing
a gas turbine exit temperature of about 1000°F. The early 1980's
would be about the earliest this technology can be in operation.
By the mid-1980's a low-Btu gas turbine inlet temperature of
2400°F at a pressure ratio of 16:1 is anticipated. The correspond-
ing steam conditions would be 1500 psig at the 1000°F level.
Ultimately a low-Btu gas turbine inlet temperature of 3000°F
at 16:1 pressure ratio is anticipated. The corresponding steam
system would operate at 2400 psig 1000°F.
Based on Lurgi type gasification of a high sulfur coal, the combined
cycle heat rates expected are of the following order:
Gas; Turbine Waste Heat Overall
Inlet Temperature Steam Heat Rate
and Pressure Ratio °F/psig BTU (HHV)/KWH; (efficiency)
1950°F (Current)
10/1 P-ratio 900°F/1250 psig 11,200 (31%)
2400°F
16/1 P-ratio 1000°F/1500 psig 9750 (36%)
3000°F
16/1 P-ratio 1000°F/2400 psig 9000 (38%)
Higher combined-cycle thermal efficiencies will be obtained by employing
more efficient coal gasification systems which have smaller associated heat
losses. Using a low temperature purification system such as Benfield or
Selexol, the best combined-cycle thermal efficiency exhibited by a moving
bed gasifier will be from the one having the highest cold gas thermal effi-
ciency, since in this case the irreversible sensible heat losses on gas cool-
ing are minimized. A slagging gasifier has a high cold gas efficiency.
Entrained bed gasifiers which do not make tars also tend to produce higher
combined-cycle thermal efficiencies because the losses on cooling of the
gas prior to purification are lower than in the moving bed case. This is
159
-------
because quenching, which is necessary when tars are produced, is not required.
Figure 9-2 shows the layout of a plant in which hydrogen sulfide is
removed from the hot gasifier off-gas at, or near, gasifier exit temperature.
This is not a demonstrated possibility but is being intensively studied at
this time. Hot gas purification eliminates cooling irreversibilities and
increases the combined-cycle thermal efficiency. It also increases the water
consumption but may decrease pollution as no water treatment is required and
no waste residues are produced.
The combined-cycle configuration provides the opportunity of desulfuriz-
ing a small volume of relatively high sulfur content fuel gas before its com-
bustion, instead of scrubbing stack gas which has a low sulfur concentration.
Processes for fuel gas H S scrubbing are more proven and developed than are
processes for stack gas SO scrubbing. The combination of improved thermal
efficiency in clean fuel production and reasonable capital related costs are
expected to make this method of power generation attractive.
9.4 DESIGN DETAILS OF GASIFIER AND COMBINED-CYCLE PLANTS
The following rules and procedure were used to design the plants for the
purpose of determining the water streams. At the start a basis of 1000 Ib of
moisture-free coal was used. The coals are shown on Table 9-1. The gasifier
rules used are:
(1) The ash contains 5 wt percent carbon.
(2) 11.08 percent of the carbon on the coal becomes carbon in CH4-
(3) 1.37 percent of the carbon in the coal becomes carbon in C^^.
(4) 11.3 wt percent of the moisture and ash-free coal is converted to
naphtha, tar and oil having the composition shown on Table 9-6.
(5) 0.74 wt percent of the moisture and ash-free coal is converted to
phenols having the composition shown on Table 9-6.
(6) The ratio of CO/CO in the gas is 0.9 mole/mole. The carbon balance
can now be completed.
(7) The ratio steam feed/oxygen in air is 3.77 mole/mole and 45 percent
of the steam fed is decomposed. Moisture in the coal does not react. The
oxygen balance can now be completed.
160
-------
TABLE 9-6. GASIFIER MASS BALANCE USING WYOMING COAL
Basis: 1,000 Ib dry coal
Material
Coal
Air,02
N2
Steam
TOTAL FEEDS
Gas, CH4
C H
2 6
H2
CO
CO.,
2
H2°
H2S
N2
NH3
Naptha , Tar ,
Oil
Phenols, etc.
Ash, ash
C
Total Total C
(moles) (Ib) (moles)
1302 56.37
14.05
52.86
52.97
56.37
6.25 6.25
0.39 0.77
30.33
19.48 19.48
21.65 21.65
45.89
0.22
52.86
0.63
104.64 7.47
6.85 0.43
3.89 0.32
H0 0_ N0 S Ash
222
(moles) (moles) (moles) (moles) (Ib)
41.67 14,03 0.36 0.25 . 74
14.05
52.86
52.97 26.49
94.64 54.57 53.22 0.25 74
12.49
1.16
30.33
9.74
21.65
45.89 22.95
0.22 0.22
52.86
0.95 0.32
3.37 0.19 0.04 0.03
0.22 0.04
74
TOTAL PRODUCTS 56.37 94.63 54.57 53.22 0.25 74
161
-------
(8) The hydrogen in the gas completes the hydrogen balance.
Table 9-6 gives a typical material balance for a gasifier. To complete
the plant overall hydrogen balance and determine the process water streams,
all that remains is to determine how much of the water in the gasifier pro-
duct gas is condensed in gas clean-up. The gas leaving the sulfur removal
system is taken to be saturated at 270 psia and 260°F and therefore to con-
tain 13.1 vol percent water vapor. The sulfur removal system is taken to
remove 6 percent of the CO as well as the sulfur. These rules enable calcu-
lation of the hydrogen balance shown on Table 9-2 scaled to 1000 megawatts.
The next step is to check the gasifier heat balance. The rules used are:
(9) Gas leaves at 900°F and ash leaves at 500°F.
(10) 23 percent of the steam feed is raised in the jacket.
These rules result in a small loss, which is shown on Table 9-7 as point 4.
Figure 9-3 shows a more detailed schematic than Figures 9-1 and 9-2 and is
presented specifically to show points of heat input and output. These points
are numbered and the results tabulated as Table 9-7. The calculations
required to make Table 9-7 are as follows:
(11) Point 1, the coal input, is the basis.
(12) Points 2, 3 and 4 come from the gasifier material and heat balances
already described.
(13) Point 8, which is the heat needed in the Benfield regenerator, was
supplied by the manufacturer. This system removes sulfur only, not carbon
dioxide, and the rules given in Section 8 do not apply.
(14) Points 5, 6 and 7 are calculated from the temperatures given on
Figure 9-3.
(15) In the Glaus plant 95 percent of the H S is converted to sulfur.
The residual gas is incinerated to convert it to SO2. Calculations show that
about 1 percent of the clean fuel is needed to raise the incinerated gas tem-
perature to 1200°F. This fuel is not available to the gas turbine, but the
hot gases are sent to heat recovery and the energy is mostly recovered.
Points 12, 13 and 14 result from calculations on the Glaus plant.
(16) To determine the energy at points 15, 16 and 17, the gas turbine-
combustion-air compressor performance was simulated by using typical informa-
tion. The simulation involves estimating an adiabatic efficiency for the
162
-------
en
w
GAS
TURBINE
CONDENSATE
TAR, OIL, NAPHTHA
AMMONIA
BFW
CLEAN FUEL GAS
HOT INCINERATOR GAS
COMBUSTOR
GAS
1 TURBINE
/
LOSSES
HEAT RECOVERY
STEAM GENERATION
STEAM
RECOVERED STEAM
STEAM AND BFW TO GASIFIER
WASTE STEAM TO WATER TREATMENT, ETC.
ONDENSATE
COLO AIR
BOILER
FEED
WATER
STEAM TURBINE
. CONDENSER
CW
SULFUR
Figure 9-3. Generating plant details for heat balance.
-------
TABLE 9-7. HEAT DUTIES AT THE VARIOUS POINTS IN THE PLANT SHOWN IN FIG. 9-3
Basis: 1000 Ib dry coal
Units: 106 Btu
Point
(see Fig. 9-3)
1.
2.
3.
4.
5.
6.
7.
8.
9.
10.
11.
12.
13.
14.
15.
16.
17.
18.
19.
20.
21.
22.
23.
Coal , HHV
Steam + bfw to gasifier
Ash
Gasifier loss
Scrubber recovery
Scrubber air cooler
Waste heat recovery
Sulfur removal heat
Sulfur removal cooling water
Clean fuel
Fuel to Glaus incinerator
Claus plant recovery
Sulfur
Claus incinerator gas
Gas turbine generator
Gas turbine losses
Turbine exhaust gas
Stack loss
Steam turbine generator
Steam turbine condenser
Waste steam to water treatment
and other uses
Air compressors
Gasifier air compressor
Wyoming
11.95
1.02
0.06
0.42
1.11
0.12
0.11
0.11
0.33
8.99
0.09
0.02
0.03
0.10
2.26
0.23
5.98
2.10
1.14
2.84
0.22
0.43
New
Mexico
11.66
1.01
0.14
0.54
0.56
0.12
0.29
0.11
0.37
8.94
0.09
0.02
0.03
0.10
2.25
0.23
5.94
2.05
1.07
2.68
0.16
0.43
North
Dakota
11.27
0.94
0.07
0.19
1.12
0.12
0.20
0.12
0.43
8.19
0.08
0.02
0.03
0.09
2.07
0.21
5.44
1.90
1.09
2.75
0.19
0.39
interstage cooling 0.08 0.08 0.08
164
-------
turbine and air compressor, and a cooling air flow. The temperature of the
combustion products is held to 1950°F.
(17) To calculate point 18, the stack temperature is taken to be 300°F.
(18) Plant steam balances were made to complete the calculation. The
steam turbine generator efficiency of 28.4 to 28.6 percent represents a com-
bination of high pressure and low pressure turbines.
Table 9-8 is presented as a check on Table 9-7. Table 9-9 shows the
electricity produced so that the designs can be scaled to 1000 megawatts to
make the summary Tables 9-2 and 9-3. In these plants tar, oil and naphtha
are not burnt but are sold. The actual station efficiencies are 28 to 28.5
percent. The efficiencies of 33.4 to 33.7 percent are presented for illus-
tration only. If the tar, etc., were to be burnt on site, the water require-
ments would alter.
9.5 EFFECT OF HOT GAS DESULFURIZATION
Hot gas desulfurization has been discussed in Section 8.1. The use of
such a process will increase the station efficiency by an amount dependent on
the gasifier and the available gas turbine. This is shown on Table 9-10.
The Bigas process yields a much hotter gas than the Lurgi process and the
advantage of hot gas desulfurization is significant. If hot gas desulfuri-
zation is used, condensate will not be recovered. This will increase the
water requirement; but because tars, etc., will pass straight to the combus-
tor, the cost will be reduced and pollution problems may be lowered.
9.6 COOLING IN A POWER PLANT
To convert "wet cooling load" (in Btu/hr) to water evaporated in the
cooling tower (in Ib water/hr), we must determine the cooling rate in the
tower (in Btu dissipated/lb water evaporated). This is not the same in
electric generating plants as in fuel-to-fuel conversion plants (Hygas,
Synthane and Solvent Refined Coal). This is because the turbine character-
istics for conversion plants are those of industrial turbines, not the larger
drives used in generating plants. In the fuel-to-fuel conversion plants the
165
-------
TABLE 9-8. PLANT ENERGY BALANCES
Units: 10 Btu/lOQO Ib dry coal
Point
(see Fig. 9-3)
10-11 Fuel to gas turbine
15 Gas turbine generator
16 + 22 Gas turbine loss and air
compressor drives
17 Turbine exhaust gas
Subtotal-gas turbine outputs
14 + 17 - 18 Heat from gas given to steam
generation
5+7+12 Waste heat recovered
Subtotal-steam produced
19 Steam turbine generator
20 Steam turbine condenser
2 Gasifier Steam
21 Waste steam used
Subtotal-steam used
1 Coal
15 + 19 Electricity
Fuel value of tar, oil, naptha,
phenol, ammonia, sulfur
Sensible heat of recovered fuel
and condensate
4+6+16 Gasifier losses, dry cooling,
+ 18 gas turbine loss and stack loss
9+20+23 Cooling water
3 Ash
21 Waste steam used
Total energy outputs
Wyoming
8.90
2.26
0.60
5.98
8.90
3.98
1.24
5.22
1.14
2.84
1.02
0.22.
5.22
11.95
3.40
1.97
0.17
2.87
3.25
0.06
0.22
11.94
New
Mexico
8.85
2.25
0.66
5.94
8.85
3.99
0.87
4.86
1.05
2.64
1.01
0.16
4.86
11.66
3.32
1.85
0.12
2.94
3.13
0.14
0.16
11.66
North
Dakota
8.11
2.07
0.60
5.44
8.11
3.63
1.34_
4.97
1.09
2.75
0.94
oj£
4.97
11.27
3.16
1.98
0.22
2.42
3.26
0.07
OO9,
11.30
166
-------
TABLE 9-9. PLANT EFFICIENCIES
Basis: 1000 Ib dry coal
Wyoming
New Mexico
North Dakota
10 Btu kw-hr 10 Btu kw-hr 10 Btu kw-hr
Coal
11.95
11.66
11.27
Electricity
3.40 996
3.32 973
3.16 926
Naphtha, tar and oil,
heat value
1.84 184*
1.67 167'
1.83 183*
Other products, heat value 0.13
0.18
0.15
Unrecovered heat
6.58
6.49
6.13
Station efficiency for
electricity
28.5%
28.5%
28.0%
Station efficiency
including use of tar
to generate electricity
33.7%
33.4%
33.6%
* Assuming 10,000 Btu/kw-hr with the tar burned in a boiler to raise steam.
167
-------
TABLE 9-10. EFFICIENCIES OF VARIOUS PLANT COMBINATIONS
(FROM REFERENCE 4, TABLE 45)
Gasifier
Sulfur Removal Gas Turbine
Station
Efficiency
Bu Mines Stirred Bed
Selexol First generation
31.4%
Iron Oxide
(hot gas)
32%
Bigas
Selexol
31.9%
Consol
(hot gas)
37.6%
Selexol Second generation
36%
Consol
(hot gas)
42.5%
168
-------
cooling towers are bypassed in winter months to prevent too low a temperature
in the condenser. Bypassing is not practiced in generating plants. Further-
more, the decision as to whether to use wet or dry cooling in fan turbine
condensers is economic and is made differently in the fuel-to-fuel plants
from the generating plants, because a somewhat different set of costs is
5—8
taken into account.
In calculating water requirements for the combined cycle power plants,
the following conditions have been used:
Cooling rate in tower Turbine condenser
(Btu/lb water evaporated) cooling method
Wyoming 1,401 wet/dry;
25% of all wet
New Mexico 1,357 all wet
North Dakota 1,484 all wet
These results come from a separate study in which cooling in coal fired
steam-electric power plants was investigated. The turbines in these plants
are not quite the same as the steam turbines in combined-cycle plants (the
steam pressure is different), but the water consumption will be close. (The
computations were made by R.W. Beck and Associates for subcontracted work
done by Water Purification Associates for the University of Oklahoma under
Subcontract No. 1916-3, Prime Contract EPA 68-01-1916.) The optimized con-
ditions for an evaporative cooling system and for two dry cooling systems
(one using the same turbine as in the wet system and the other using a pro-
bable design of a future high back pressure turbine) are shown on Table 9-11.
The annual cooling costs are shown on Table 9-12. On Table 9-13 the results
are summarized. Dry cooling will be used if water supply and treatment costs
are more than the breakeven costs used. If wet cooling is used, the quantity
of water evaporated (averaged over the year) can be calculated from the cool-
ing load and the cooling rate in the tower shown on the table.
In the same study a partial analysis of combined wet/dry cooling in New
9
Mexico was made. The result is shown on Figure 9-4. A very recent report
points out that even if wet condensing is apparently more economical the
added cost of dry condensing can be small. In Braintree, Massachusetts (on
169
-------
TABLE 9-11. SUMMARY OF DESIGN CONDITIONS FOR OPTIMIZED COOLING
SYSTEMS FOR COAL FIRED GENERATING PLANTS (10& KW)
Dry Cooling
High Back Pressure Turbine
Beulah, Gillette, Farmington,
N. Dakota Wyoming New Mexico
Average Annual Dry-Bulb
Temperature, F
Average Annual Wet-Bulb
Temperature,°F
Design Wet-Bulb Temperature, °F
Ambient Temperature for
Determination of Peaking
42
46
51
Capacity Requirement, F
Initial Temperature Difference, °F
Design Cooling Range, °F
Design Approach Temperature, F
Design Terminal Temperature
Difference, F
Design Inlet Temperature, °F
Turbine Exhaust Pressure, in Hga
Total Tower Heat Load, 10 Btu/hr
Condenser Duty, 10 Btu/hr
Circulating Water Flow, 10 gpm
Condenser
Surface Area, 10 sq.ft.
Number of Tubes, 10
Tube length, ft
Velocity through tubes, ft/sec.
82.0
29
14.50
5.0
81.6
2.0
4647.0
4587.0
634.9
622
91
26
6.98
82.0
34
15.0
7.0
79.1
2.0
4647.0
4587.0
613.6
512
81
28
6.94
82.0
31
15.5
6.0
79.6
2.0
4647.0
4587.0
593.9
551
43
49
6.95
(continued)
170
-------
TABLE 9-11 (continued)
Dry Cooling
High Back Pressure Turbine
Beulah,
N. Dakota
Average Annual Dry-Bulb
Temperature, °F
Average Annual Wet-Bulb
Temperature, °F
Design Wet-Bulb Temperature, F
Ambient Temperature for
Determination of Peaking
Capacity Requirement, °F
Initial Temperature Difference, °F
Design Cooling Range, °F
Design Approach Temperature, °F
Design Terminal Temperature
Difference, °F
Design Inlet Temperature, °F
Turbine Exhaust Pressure, in Hga
£
Total Tower Heat Load, 10 Btu/hr
Condenser Duty, 10 Btu/hr
Circulating Water Flow, 10 gpm
Condenser
Surface Area, 10 sq.ft.
Number of Tubes, 10
Tube length, ft
Velocity through tubes, ft/sec.
42
Gillette,
Wyoming
46
Farmington,
New Mexico
51
82.0
68.0
34.0
5.0
81.0
3.5
5211
5144
304.2
451
44
39
6.92
82.0
67.0
33.5
5.0
82.1
3.5
5211
5144
308.7
462
46
38
6.68
82.0
68.0
34.0
5.0
85.0
3.8
5212
5146
304.6
452
45
38
6.73
(continued)
171
-------
TABLE 9-11 (continued)
Evaporative Cooling
Average Annual Dry-Bulb
Temperature, °F
Average Annual Wet-Bulb
Temperature, °F
Design Wet-Bulb Temperature, °F
Ambient Temperature for
Determination of Peaking
Capacity Requirement, °F
Initial Temperature Difference, °F
Design Cooling Range, °F
Design Approach Temperature, °F
Design Terminal Temperature
Difference, °F
Design Inlet Temperature, °F
Turbine Exhaust Pressure, in Hga
Total Tower Heat Load, 10 Btu/hr
Condenser Duty, 10 Btu/hr
Circulating Water Flow, 10 gpm
Condenser
Surface Area, 10 sq.ft.
Number of Tubes, 10
Tube length, ft
Velocity through tubes, ft/sec
Beulah,
N. Dakota
42
36
71
27.43
20.0
7.0
91.0
4.0
4722
4661
344.3
427
49
33
6.94
Gillette,
Wyoming
46
37
61
24.55
28.0
7.0
89.0
3.5
4694
4634
382.3
457
56
31
6.75
Farming ton
New Mexico
51
40
64
26.43
28.0
7.0
92.0
4.0
4722
4661
357.3
438
52
32
6.81
172
-------
TABLE 9-12. SUMMARY OF ANNUAL EVALUATED COSTS FOR OPTIMIZED COOLING SYSTEMS
{106 DOLLARS /YR FOR 1000 MEGAWATTS, 6125
Dry Cooling, Conventional Turbine
Annual capital, & operation & maintenance
cost of cooling system
Annual plant fuel cost
Credit for excess generation
Annual replacement capacity & energy cost
Annual auxiliary power & energy cost
Total annual evaluated cost
Dry Cooling, High Back Pressure Turbine
Annual capital, & operation & maintenance
cost of cooling system
Annual plant fuel cost
Credit for excess generation
Annual replacement capacity & energy cost
Annual auxiliary power & energy cost
Total annual evaluated cost
Evaporative Cooling
Annual capital, & operation & maintenance
cost of cooling system
Annual plant fuel cost
Credit for excess generation
Annual replacement capacity & energy cost
Annual auxiliary power & energy cost •
Total annual evaluated cost
Beulah,
N . Dakota
10.786
27.411
-0.343
0.834
2.847
41.535
6.707
29.361
-0.431
1.263
• 1.554
37.454
2.272
27.407
-0.162
0.406
0.595
30.518
STREAM-HR/YR)
Gillette,
Wyoming
10.069
27.414
-0.325
0.834
2.284
40.285
6.060
29.362
-0.436
0.801
1.593
37.380
2.233
27.414
-0.177
0.057
0.615
30.142
Farmington ,
New Mexico
10.555
27.413
-0.317
0.803
2.750
41.204
6.054
29.360
-0.401
1.143
1.613
37.770
2.062
27.406
-0.070
0.490
0.569
30.457
173
-------
TABLE 9-13. SUMMARY OF COMPARISON OF WET AND DRY COOLING
IN COAL FIRED GENERATING PLANTS (10b KW)
Beulah,
N. Dakota
Gillette,
Wyoming
Farmington,
New Mexico
Water evaporated in wet
cooling (gallons per
stream-hr)
381,780
402,180
417,780
Breakeven cost of water
($/1000 gallons)
Conventional turbine 4.71
High back-pressure
turbine
2.96
4.12
2.94
4.19
2.85
Tower heat load
Btu/stream-hr)
4,722
4,694
4,722
Cooling rate in tower
Btu/lb water evaporated 1,484
1,401
1,357
174
-------
8
126
120
110 —
D
f 100 —
90
>
xx
/X
J4.19/1000 GALLONS OF
WATER EVAPORATED
0.25 0.50 0.75 J.O
WATER EVAPORATED AS A PERCENTAGC OF WATER EVAPORATED FOR ALL EVAPORATIVE COOLING SYSTEM
Figure 9-4. Total annual evaluated costs of wet/dry cooling system as
a percentage of evaporative loss of all evaporative cooling
in New Mexico.
-------
the coast south of Boston) an 85 MW combined-cycle generating plant is
planned (65 MW gas turbine and 20 MW steam turbine). Suitable water costs
$0.7O/thousand gallons, and dry cooling has been chosen.
In tabulating water quantities in Section 13, all wet condenser cooling
was assumed in North Dakota and New Mexico? and in Wyoming, because of the
high cost of water and for illustration, combined wet/dry cooling was
assumed with the water consumption being one quarter of the water required
for all wet cooling.
176
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REFERENCES SECTION 9
1. Hebden, D., "High Pressure Gasification.Under Slagging Conditions,"
presented at 7th Synthetic Pipeline Gas Symposium, Chicago, 1975.
2. Ahner, D. J., Sheldon, R. C. , Garrity, J. J. and Kasper, S., "Economics
of Power Generation from Coal Gasification for Combined-Cycle Power
Plants," presented at American Power Conference, Chicago, April 1975.
3. Kydd, P. H. , "Integrated Gasification Gas Turbine Cycles," Chem. Eng.
Progress Ti (No. 10) 62-68, October 1975.
4. Robson, R. L., et al., "Fuel Gas Environmental Impact: Phase Report,"
EPA 600/2-75-078, November 1975, N.T.I.S. Catalog PB-249-454.
5. Rossie, J. P., Mitchell, R. D. and Young,R. 0., "Economics of the Use of
Surface Condensers with Dry-Type Cooling Systems for Fossil-Fueled and
Nuclear Generating Plants," R. W. Beck and Associates, Denver, Colorado,
U.S. Atomic Energy Commission Report No. TID-26714 (UC-12), December 1973.
6. Rossie, J. P., Cecil, E. A. and Young, R. 0., "Cost Comparison of Dry-
Type and Conventional Cooling Systems for Representative Nuclear
Generating Plants," R. W. Beck and Associates, Denver, Colorado, U.S.
Atomic Energy Commission Report No. TID-26007 (Uc-80), March 1972.
7. Mitchell, R. D., "A Method for Optimizing and Evaluating Indirect Dry-
Type Cooling Systems for Large Steam-Electric Generating Plants,"
R. W. Beck and Associates, Denver, Colorado, U.S. Energy Research and
Development Administration Report No. ERDA-74 (UC-12), June 1975.
8. "Heat Sink Design and Cost Study for Fossil and Nuclear Power Plants,"
United Engineers and Constructors, Inc., U.S. Atomic Energy Commission
Report No. WASH-1360 (UC-13 & 80), December 1974.
9. Henderson, M.D., "Feasibility Study for a Direct, Air-Cooled Condensation
System," EPA-600/2-76-178, U.S.E.P.A., Research Triangle Park, N.D., 1976.
177
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SECTION 10
COOLING
10.1 INTRODUCTION
All fuel conversion plants are less than 100 percent efficient. The
higher heating value of the product fuel plus the heating value of combus-
tible byproducts and wastes is less than the heating value of the coal taken
in. Since the first law of thermodynamics is obeyed, that part of the energy
in the coal which is not recovered as fuel must be lost to the atmosphere in
some way. A major part of each of the plant designs made in this study is
the determination of the ultimate disposition of unrecovered heat. In each
of the design sections, the ultimate disposition of unrecovered heat is tabu-
lated. Some of the unrecovered heat leaves the plant as hot flue gas up a
chimney, water vapor from drying coal, convective and radiant losses from
electric motors and container surfaces and similar ways, all called direct
losses. Direct losses are those for which cooling water cannot be used.
The remaining unrecovered heat leaves through a heat transfer surface.
In this study two types of heat transfer equipment are considered: 1) finned
tube heat exchangers with fans for forced air cooling or dry cooling or air
cooling; and 2) shell and tube heat exchangers using circulating cooling
water, itself cooled in a mechanical draft evaporative cooling tower, called
wet cooling or evaporative cooling. For all the major points of cooling
using heat transfer surfaces, a decision must be made on how much water will
be evaporated to dispose of the unwanted heat. The study is economic, and
there are four factors to be accounted for in choosing a cooling system:
(1) The capital cost difference must be determined. In general, dry
coolers cost more than wet coolers for a given heat removal, and the cost
178
-------
difference increases when the heat is removed at lower temperature levels.
The additional capital cost of dry cooling must also include the greater
installed cost of all plant compressors on which dry cooling is used.
(2) The energy needed to operate the cooling system must be determined;
this is the energy needed to operate fans and to circulate cooling water.
(3) In some circumstances, notably steam turbine condensers, the plant
operation consumes more energy if the temperature of cooling is raised. This
fuel penalty must be determined.
(4) Finally, the cost of water for evaporative cooling must be consi-
dered. This is discussed below.
These various cost factors will be discussed at each of the following
important points of cooling in the plants:
cooling process streams
cooling the acid gas removal regenerator condenser
condensers for steam turbines
interstage cooling in gas compressors.
The designs have usually incorporated air cooling to about 140°F and wet
cooling below this.
Cooling Process Streams
The conclusion of the study on cooling process streams is that air cool-
ing should be used to about 140°F and wet cooling below this. This result
has already been incorporated into the previous design section.
Cooling the Acid Gas Removal Regenerant Condenser
Because in the area of the country under study water is scarce and expen-
sive, dry coolers have been used in all acid gas removal systems. This is
discussed below. This is quite common practice in chemical plants and refin-
eries and is incorporated into many coal conversion plant designs.
179
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Condensers for Steam Turbines
The cooling of steam turbine condensers is a very large part of the
total plant cooling load and is complicated to study. This load is discussed
in several sections below, and great emphasis is placed on combined dry and
wet cooling. The choice of cooling varies from site to site. All wet cool-
ing was selected in North Dakota and wet/dry cooling at the other sites.
Interstage Cooling in Gas Compressors
Cooling between stages of gas compression is also complicated to study.
This load is discussed in Section 10.9 below, and it turns out that wet cool-
ing is preferred at New Mexico and wet/dry cooling at Wyoming. At North
Dakota we have assumed wet cooling as at New Mexico.
10.2 THE COST OF WATER
Water is not free._ In most Western states water must be bought from its
source. A modest cost of water rights is in the range of $10 to §20 per
acre-foot, which is 3C to 6<-' per thousand gallons. The cost of moving water
to the sites of interest in this study is about 1.2
-------
10.3 COOLING PROCESS STREAMS
Below 270 to 300°F, useful heat cannot be extracted from a process stream
and further cooling must dissipate the heat to the atmosphere. A simple cal-
culation shows it is probably best to cool to about 130 to 140°F using an air
cooler and to cool below this temperature using a wet system. The economics
prove to be a trade-off between the cost of water and the cost of money. An
example is now given:
A methane stream at 1,000 psig is to be cooled from 280 to T°F. The
ambient air temperature is 90°F. Cooling water is available at 80°F and may
be heated to 105°F. The heat removed in the cooling tower may be assumed to
be 1,400 Btu/lb of water evaporated (see Section 10.4 for a discussion of this
number which depends on the climate) . The conditions of the process stream
are such that the heat transfer coefficient for dry cooling is
70 Btu/(hr) (ft2) (°F) and for wet cooling is 100 Btu/(hr) (ft2) (°F) . The
data used for this and other calculations on cooling is given on Table 10-1.
First, we find the cost of removing 10 Btu/hr by dry cooling using, as
our example, a cooled temperature T = 130°F. The air temperature rise (see
Table 10-1) is
T \
-= — - - 90 J = 40.25 (where T = 130°F)
The high temperature difference, h^ = 280 - (90+t) = 169.75 (where
T = 130°F) .
The low temperature difference A2 = T - 90 = 40 (when T = 130°F) .
The log mean temperature difference = (A - A )/ln (A /A2) = 83.1 .
The heat exchanger area A = Q/(LMTD) (U) = 106/(83.1) (70) = 172 ft
2 2
The charges for the heat exchangers are $18/ft * 0.17/yr x A ft =
$526/yr = $0.0658/hr for 8,000 hrs/yr.
The cost of running the fan is
(0.0175A x 0.7455, kw) x (2<=Aw-hr) = $0.0449/hr
Total cost $0.11/hr or $0.11/10 Btu.
181
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TABLE 10-1. REPRESENTATIVE FORMULAE AND DATA USED FOR
CALCULATING COST OF WET AND DRY COOLING
1. Heat Transfer Coefficients (U)
U, Btu/(hr)(ft2)(°F)*
Dry Wet
Condensing steam from turbine drives 130 400
Condensing steam in presence of noncondensible
gases in gas purification regenerator 75 no
Cooling methane at 1,000 psig 70 100
Cooling water for off-gas scrubbing 120 170
Air and oxygen compressor interstage coolers: psig
10 10 12
50 20 20
100 30 40
300 40 50
>500 50 70
*Based on bare tube area of finned tubes for dry coolers.
2. Heat Transfer Area
(a) Dry Cooling:
The following empirical formula for air temperature rise was used
At = 0.005
Here, t^ is the inlet air temperature and T^, T2 are the process stream
inlet and outlet temperatures. This formula is suitable for estimating
purposes. In actual designs the rise must be optimized. The other
formulae for calculating the heat transfer area are standard.
(b) Wet Cooling:
The rise in water temperature has usually been assumed to be 25 F. In
actual designs this may be controlled by the limitation set on the maxi-
mum turbine back pressure, here taken to be 5 in. Hg. absolute. In
most cooling systems the water chemistry, particularly scaling tenden-
cies, also controls the permitted temperature rise. Standard heat
transfer formulae were used to calculate areas.
(continued)
182
-------
TABLE 10-1 (continued)
3. Energy for Cooling
Dry cooler fans:
Cooling tower fans:
Cooling water circulation
pumps:
Horsepower = 0.015 x Area (U < 50)
= 0.0175 x Area (50< U < 100)
= 0.02 x Area (U > 100)
Horsepower = 0.012 x gpm circulated
Horsepower = 0.033 x gpm circulated
4. Costs
Heat Exchangers*!
Dry cooling
Wet cooling
Other:
Cooling tower
Electrical energy
Steam
Cost
$18/ft
$5.1/ft:
$5.6/ft"
$6.I/ft"
$8.9/ft"
2**
Pressure
(p,psig)
p < 300
300 < p < 450
450 < p < 600
p > 600
$7/gpm circulated
2<:/kw-hr
$1.80/106 Btu
Annual Charges
for Amortization
Plus Maintenance
17%/yr
20%/yr
15%/yr
*Installed but without piping, valving, instrumentation, engineering or
other costs. These costs are for comparison of wet and dry cooling only,
not for absolute plant costs. The energy and cost figures are typical,
but variations are to be expected for actual design.
**Based on bare tube area of finned tubes.
183
-------
Now we find the cost of removing 10 Btu/hr by wet cooling. When
T = 130°, the log mean temperature difference =
(280 - 105) - (130 - 80) _
280 - 105 "
ln
130 - 80
The heat transfer area = 106/(100) (100) = 100 ft2
The charges for the heat exchangers are
$8.9/ft2 x o.2/yr x A ft2 = $178/yr = $0.0223/hr
The circulation rate is
10 Btu/hr x i lb/25 Btu = 40,000 Ib/hr = 80 gal/min
The cooling tower charges are
$7/gpm x 80 gpm x 0.15/yr = $84/yr = $0.0105/hr
Operating the tower fan plus circulating pump costs
0.045 hp/gpm x 80 gpm x 0.7455 kw/hp x 2C/kw-hr = $0.0537/hr
For free water the total cost is $0.0865/hr = $0.0865/10 Btu.
The rate of evaporation of water is
fi "^
10 Btu/hr x i lb/1,400 Btu = 714 Ib/hr = 0.0857 x 10 gal/hr
This calculation and others are plotted on Figure 10-1. On Figure 10-2,
entitled "Cost of Cooling a Low Pressure Gas Stream," the calculations have
been repeated for the case:
U, dry cooling = 40 Btu/(hr)(ft2)(°F)
U, wet cooling = 50 Btu/(hr)(ft2)(°F)
2
cost of wet area = $5.6/ft
184
-------
0.3
I
i
WET, WATER S2.00/103 GAL
COST
§/!06Btu
0.2
200
SI.OO/I03GAL
80.60/IP3 GAL
180 160 140 120
EXIT TEMPERATURE (°F)
-
i
100
Figure 10-1. Cost of cooling a high pressure methane stream.
185
-------
0.3
0.2
O.I
WET, WATER
COST
$/!06Btu
FREE
I
200
180 160 140 120
EXIT TEMPERATURE (°F)
Figure 10-2. Cost of cooling a low pressure gas stream.
186
-------
It is clear that the choice of wet or dry cooling depends on all the
process conditions as well as the cost of water and should be made for each
situation individually. As an adequate estimate, process streams have been
cooled to 140°F using dry cooling and below this using wet cooling.
10.4 WATER EVAPORATED FOR WET COOLING
Given that a certain cooling load in Btu/hr is to be lost through wet
cooling, it is necessary to know how much water must be evaporated in the
cooling tower to remove this load.
In a wet or evaporative cooling tower as shown in Figure 10-3, the warm
water leaving the heat exchanger is pumped to the top of the tower where it
is distributed through an assembly of nozzles. It then passes through the
heat transfer zone and is collected in the basin. The splash packing, or
fill material, inside the tower retards and breaks up the water into fine
droplets creating both a large droplet surface area and high residence time.
Since heat is removed from the water droplets through evaporation and convec-
tion, the larger the surface area of the droplets, the higher is the heat
transfer efficiency. Air is circulated through the tower, either by natural
convection (natural draft tower) or by a mechanical fan (mechanical draft
tower), to remove the evaporated water and sensible heat from the tower.
Makeup water is required to replace the evaporated water, and the drift and
blowdown are required to eliminate the accumulation of solids in the circu-
lating water.
The evaporation rate of the tower to remove a certain heat load depends
on the design of the tower, the relative flow rates of water and air, the
entering warm water temperature, the temperature and humidity of the enter-
ing air and the annual average evaporation rate. Btu/lb water evaporated
depends on how the tower is operated, that is, on whether the tower is by-
passed and what cold water temperature is maintained. For the purpose of
this study the cooling tower is operated to give the best performance when
coupled with all wet cooling of steam turbine condensers as described in
Sections 10.7 and 10.8. As an example, Table 10-4 gives the heat removal
rate (Btu/hr) and the water consumption rate (Ib/hr) for such an all wet
187
-------
:•.
-
WARM WATER
FROM HEAT EXCHANGER
T
SLOWDOWN
AIR
O
COLD WATER
TO HEAT EXCHANGER
EVAPORATION
OO
V////4 V//M
ISZS3 Willl W///A
P/////A EZS21 V7777\ P77773
DRIFT
FILL MATERIAL
AIR
MAKEUP
WATER
Figure 10-3. Schematic of wet cooling tower.
-------
system in Casper, Wyoming. The monthly averages are:
Month Btu/lb Water Evaporated
1, January 1475
2, February 1472
3, March 1438
4, April 1440
5, May 1345
6, June 1308
7, July 1272
8, August 1300
9, September 1328
10, October 1399
11, November 1487
12, December 1499
The annual average is 1397 Btu/lb water evaporated. This number, and
the other numbers given on Table 10-2, will be used to estimate the water
v -
consumed for wet cooling of process streams in Section 13. (AS explained in
Section 9.6, Table 10-2 is not used for power generation.)
10.5 COOLING IN ACID GAS REMOVAL
Removal of the acid gases, carbon dioxide and hydrogen sulfide, is an
important energy-consuming part of the coal conversion plants and the regen-
erator condenser is a large cooling load. (For power generation, carbon
dioxide is not removed and this discussion does not apply.) In the process
designs two acid gas removal systems have been considered: a Benfield hot
Potassium carbonate type and a Selexol physical solvent type. In each case
TABLE 10-2. ANNUAL AVERAGE WATER CONSUMPTION FOR WET COOLING
Btu/lb Water
Evaporated
Wyoming 1397
North Dakota 1420
New Mexico 1375
189
-------
the feed stream to the condenser is a mixture of water vapor and nonconden-
sible carbon dioxide at about 230°F. For condensing steam in the presence
of the noncondensible gas, the recommended heat transfer coefficients are
75 Btu/(hr)(ft2)(°F) when dry condensing and 110 Btu/(hr)(ft2)(°F) when wet
condensing. Condensation proceeds as the vapor mixture is cooled. For a
Benfield type system, cooling to 140°F will condense enough water to keep
the circulating absorbent in balance.
Calculating by the method used for cooling process streams, the cost for
c
Benfield type condensers is $0.116/10 Btu when air cooling and the same cost
when wet condensing if water costs 46C/thousand gallons. Dry condensing is
assumed for this study.
For a Selexol type process it is probably necessary to cool to 100°F to
keep the water in the system in balance. By air cooling to 130°F and then
wet cooling to 100°F, 90 percent of the total cooling load is taken by the
air cooler and 10 percent by the wet cooler. (The feed stream contains about
1.4 moles H O/mole CO . The stream saturated at 130°F contains about
0.18 moles H O/mole CC>2 and the stream saturated at 100°F contains about
0.07 moles H O/mole CO .)
^ A
10.6 CHARACTERISTICS OF STEAM TURBINES
Since the cooling load on steam turbine condensers is a large part of
the plant cooling load, it is necessary to conduct detailed studies to deter-
mine the relative amounts of wet or dry cooling to be employed. Examples of
various types of cooling are discussed in the following two sections. The
numerical results depend on the characteristics of the turbine.
In a steam turbine drive system the steam rate required by the turbine
to produce a certain shaft power output depends on the inlet steam condition,
the condenser pressure and the turbine efficiency. Usually the higher the
inlet steam pressure and temperature, the higher will be the thermal effi-
ciency of the system. In the present application where the steam is partial-
ly produced by waste heat recovery, the usual steam pressure is in the range
from 715 to 915 psia, and the superheated temperature in the range from
600° to 900QF. Also, in the present application where the steam turbine
drive is used mainly for gas compression purposes, the type of turbine drive
190
-------
used usually has a maximum efficiency of about 80 percent when the condenser
pressure is in the range of 3 to 5 in Hg absolute. The corresponding steam
saturation temperatures for the two condenser pressures are 115 and 134°F,
respectively. Above 134°F or below 115°F, efficiency falls.
The heat rates required when the condenser temperature is in the range
of 115 to 134°F have been calculated for the various inlet steam conditions
mentioned and are plotted in Figure 10-4. The calculations were made using
a turbine efficiency of 80 percent and a bearing efficiency of 95 percent.
The results in Figure 10-4 show that the steam rates for the four inlet steam
conditions are quite close and that they can be represented by a single
straight line going from a steam rate value of 11,700 Btu/kw-hr at the con-
denser temperature of 115°F to a value of 12,200 Btu/kw-hr at the condenser
temperature of 134°F.
The increase in steam rate with condenser temperature indicates that
there is a certain fuel penalty to be considered in evaluating the cost of
various cooling systems.
The condenser cooling loads when the condenser temperature is in the
range of 115 to 134°F have also been calculated for the four inlet steam con-
ditions mentioned and are plotted in Figure 10-5. The results indicate that
the condenser loads for the four inlet steam conditions are also quite close
and that they can be represented by a single straight line, going from a
value of 8,200 Btu/kw-hr at the condenser temperature of 115°F to a value of
8,700 Btu/kw-hr at the condenser temperature of 134°F. This typical line
will be used for condenser load calculations when the economics of condenser
cooling systems are evaluated.
10.7 DRY AND WET COOLING SYSTEMS FOR TURBINE CONDENSERS
For a turbine condenser cooling system direct dry cooling, wet cooling
or a combination of both have been considered. The cost and the water con-
sumption rate of each has been analyzed for the climates of Beulah, North
Dakota, Casper, Wyoming and Farmington, New Mexico. The details of the all
dry system, the all wet system and the wet/dry combination will be discussed.
The results of month-by-month calculations for the three sites and the
191
-------
15
* 10
3
*-
CO
10
UJ
<
a:
STEAM
TEMP
(I) 600
(2) 700
(3) 900
(4) 900
STEAM
PRESSURE
(PSIA)
715
915
715
915
ID
I-
(O
ro
in
110
115 120 130 134
CONDENSER TEMPERATURE (°F)
140
Figure 10-4. Turbine heat rates.
192
-------
15
T
T
2 10
CO
10
o
Q
<
O
tr
u
co
z
LJ
O
Z
o
o
STEAM
TEMP
(I) 600
(2} 700
(3) 900
(4) 900
STEAM
PRESSURE
(PSIA)
715
915
715
915
o>
X
z
10
o»
I
z
in
1
1
no
115 120 130 134
CONDENSER TEMPERATURE (°F)
140
Figure 10-5. Turbine condenser cooling loads.
193
-------
analysis of the results will be presented in the next section.
In the all dry system the condenser is assumed to be cooled directly by
air (see Figure 10-6 assuming that the wet cooling section is not present)
and is assumed to be held within the temperature range of 115 to 134°F. The
condenser is designed to have a maximum temperature of 134°F when the ambient
air is at the summer design temperature. As the ambient temperature drops
during the year, the condenser temperature also drops eventually reaching
115°F. If the ambient temperature drops further the power to the fans, which
are of the variable speed or variable pitch type, is reduced. The reduction
of the fan power reduces the air flow rate through the condenser and, thus,
the heat transfer rate so that the condenser temperature can be maintained
at 115°F.
The heat transfer rate in a condensing heat exchanger is given by the
equation
{TC ~ TAC} " (TC ' TAH)
• ""pf / M
n C AC
^ T _ T
C AH
where Q is the heat transfer rate (Btu/hr) for the dry condenser, U is the
heat transfer coefficient (Btu/(hr)(ft }(°F)), A is the dry condenser sur-
2
face area (ft ), T is the condenser temperature (PF), T and T are the
\* fVmf f\Ll
inlet cold air temperature (°F) and the exit warm air temperature (°F) respec-
tively. The air temperature rise (T - T ) is given by the empirical equa-
X*Ul f\v«
tion (see Table 10-1)
(TC ~ V
Finally, from Section 10.6, the condenser load is a function of condenser
temperature T
/Tc - 115 \
\ 134 - 115 /
QQD = 8,200 + 500 I = 5,174 + 26.3 T (3)
194
-------
10
tn
STEAM
CO CO CO CO
FROM
TURBINE
Tc
CONDENSATE
EVAPORATION
11
DRY
AIR RATE, G
WET COOLING
Figure 10-6. Turbine condenser cooling systems.
-------
For a given design temperature T , which in this study is close to peak
summer temperature (see following tables) , these relationships make it pos-
sible to calculate the condenser area A . To find A , T is taken as 134°F.
U is taken to be 130 Btu/(hr) (ft ) (°F) for all air condenser calculations.
For any other nondesign, ambient air temperature, it is then possible to
calculate the condenser temperature T by trial and error. When T comes
VtX V>r
out to be less than 115°F, it is taken to be 115°F and the fan power is
reduced. The reduction in fan power to maintain the condenser temperature
at 115 °F is a function of the ambient temperature drop below that tempera-
ture when the condenser is at 115°F. The relationship used is a typical
example shown on Figure 10-7. Actual variations in fan power and in the
high and low condensing temperature will depend on the type of equipment
selected.
In the all wet system the condenser is cooled by circulating water
which evaporates in the wet cooling tower (see Figure 10-6 assuming the dry
cooling portion is not present) . The cooling water is assumed to be circu-
lated through the condenser at a constant rate. The wet tower has, however,
multiple cells operating at the same water- to-air flow ratio. The condenser
is assumed to operate in the temperature range of 115 to 134°F. At the sum-
mer design condition the condenser is designed to have a temperature of 134°F.
As the ambient temperature drops during the year, the condenser cools even-
tually to a temperature of 115°F. If the ambient temperature continues to
drop in the colder months of the year, the multiple cells in the wet tower
are shut off one by one. Each cell is shut off by turning the fan off so
the water passes to the basin without any cooling. This allows the return-
ing cooling water to the condenser to be such as to maintain the condenser
temperature at 115°F during these colder months (below 115°F, turbine effi-
ciency falls) . Shutting off of parts of the cooling tower has the benefit
of saving fan energy.
The heat transfer rate in the wet condenser is given by the equation
T - T
TC TWH
196
-------
0 20 40 60 80
AMBIENT TEMPERATURE DROP(°F)
Figure 10-7. Fan power reduction factor for air coolers.
197
-------
where Q is the heat transfer rate (Btu/hr) for the wet condenser, U is the
OW Q
heat transfer coefficient, Btu/(hr)(ft )(°F), AW is the wet condenser surface
o
area (ft ), and T and T are the inlet and exit cooling water temperatures
r¥v* Wii
respectively. The heat transfer coefficient U has been taken as
400 Btu/(hr)(ft2)(°F).
The rate of heat removal by the circulating water is
where C is the heat capacity of water, 1 Btu/(lb)(°F), and L is the cooling
water circulation rate (Ib/hr).
Equations (4) and (5) for a wet condenser are similar to equations (1)
and (2) for a dry condenser. From Section 10.6, the condenser load Q_TTf
vJW
Btu/shaft kw, is known as a function of condenser temperature T . The rela-
tionship is Equation (3). For summer design temperatures we have chosen T
(cold water temperature) = 99°F, TWH (hot water temperature) = 124°F and the
condenser temperature TC = 134°F. This enables the calculation of the water
circulation rate L, and wet condenser area A^. For a chosen summer design
wet bulb temperature, the cooling tower can also be designed. The cooling
tower was designed using Reference 4. (The procedure is described in Refer-
ence 4 and other texts on cooling tower design. Fill type H, 30 ft fill
height, 18 ft air travel were used with the wet bulb temperature suitable
for the site, 99°F cold water temperature and 25°F range; this fixes the
tower characteristic KaY/L and the water-to-air ratio L/G.)
When the ambient wet bulb temperature falls below design, the cold water
temperature, the condenser temperature and the hot water temperature fall.
Given an off-design wet bulb temperature, a trial-and-error calculation is
used to find the water and condenser temperatures for the design condenser
and the design cooling tower. Equations (4) and (5) are used along with the
cooling tower curves from Reference 4.
When the ambient wet bulb is such that the condenser would fall below
115°F, some cooling tower fans are shut off to hold the condenser at 115°F.
The bypass is shown on Figure 10-6. The bypass rate L is given by
198
-------
TWC • TWH - T
-------
eventually the wet tower is completely shut down. All the cooling load is
now in the dry condenser, if the ambient temperature continues to drop fur-
ther, fan power in the dry condenser is reduced to maintain the condenser
temperature at 115°F.
The benefits of the wet/dry combination are that the capital investment
is less than that of an all dry system and the water consumption rate is less
than that of an all wet system. In the following section the operating cost
and the water consumption rate for different combinations of wet and dry sys-
tems will be evaluated for the three sites considered. The wet/dry system
chosen for this study should be considered as illustrative only. Other steam
and water flow arrangements should be studied when evaluating a particular
site. Discussions concerning alternative arrangements may be found in Refer-
ences 6 and 7.
10.8 WATER CONSUMPTION FOR TURBINE CONDENSERS AT SPECIFIC SITES
As an example, consider the climatic conditions of Casper, Wyoming.
Table 10-3 shows the results of month -toy-month calculations for an all dry
condenser. The condenser area is first found for the summer design condi-
tion: ambient air T - 96°F; condenser temperature T = 134°F; load Q =
8,700 Btu/kw-hr. From Equation (2)
~ 96 = 0.005 U(T - T) = (0.005)(130)(30) = 24.7
>- *v—
so
TAH - 12°'7
From Equation (1)
' T — T
ln _C AC
CTV U
f\n
200
-------
TABLE 10-3. WATER CONSUMPTION RATE PER SHAFT KW FOR AN ALL DRY SYSTEM AT CASPER, WYOMING
MONTH
DBT (°F)
WBT (°F)
Q (Btu/hr)
TTir <°F)
we
TWH (°F)
T ( F)
R V '
WET FAN POWER
(10 Kw)
PUMP POWER
(ID"3 Kw)
WATER CONSUMP-
TION (Ib/hr)
QQD (Btu/hr)
DRY FAN POWER
(10 Kw)
Tc
FUEL PENALTY
1
24
20
0
—
—
0
0
0
8200
5.10
115
0
2
26
22
0
—
—
0
0
0
8200
5.10
115
0
3
32
27
0
—
—
0
0
0
8200
5.53
115
0
4
41
34
0
—
—
0
0
0
8200
5.95
115
0
5
54
44
0
—
—
0
0
0
8200
8.08
115
0
6
65
51
0
—
—
0
0
0
8200
14.0
115
0
7
71
55
0
—
—
0
0
0
8200
21.2
115
0
8
70
53
0
—
—
0
0
0
8200
19.6
115
0
9
59
46
0
—
—
0
0
0
8200
10.2
115
0
10
47
38
0
—
—
0
0
0
8200
6.80
115
0
11
32
27
0
—
—
0
0
0
8200
5.53
115
0
12
30
25
0
—
—
u
0
0
8200
5.53
115
0
DESIGN DBT = 96°F
DESIGN WBT = 60°F
WET CONDENSER AREA = 0
DRY CONDENSER AREA = 2.85 ft
WET CIRCULATION RATE = 0
-------
so
A = 2.85 ft2
The installed fan power = 0.02 A hp = 0.0425 kw.
Now consider month 7, which is the hottest month of the year. The
monthly average dry bulb temperature T = 71°F. The calculation is trial-
-A\-»
and-error. For a first try assume T = 115°F so that, from Equation (3) ,
Vj»
Q should be 8,200. From Equations (2) and (1) we calculate Q to be
11,000. In fact, QQD = 8,200 when the ambient air temperature T = 79°F.
Since the monthly average temperature is 8°F below 79° F, the monthly aver-
age fan power is (from Figure 10-7) 0.5 x 0.0425 = 0.0212 kw. The other
monthly calculations involve only a calculation of fan power.
Table 10-4 shows the results for an all wet system. The summer design
conditions are QQW = 8,700, T = 134°F, T = 99°F, and T = 124°F. From
Equation (5) the water circulation rate L = 348 Ib/hr = 0.696 gal/min. From
2
Equation (4) the condenser area AW = 1.09 ft . The cooling tower design is
taken from Reference 4, and the following is assumed for an example: the
fill characteristic curve for cross-flow fill type H, 30 ft high with 18 ft
of air travel, crosses the tower operating curve for 60°F wet bulb tempera-
ture (the design wet bulb), 25°F range and 39°F approach when KaY/L =1.1
and L/G = 2.8 , At this point the two curves are simultaneously satisifed,
and this fixes the cooling tower. Finally, from Table 10-1, the water cir-
culation pump requires 0.033 X gpm,hp = 0.0171 kw and the installed tower
fan power = 0.012 X gpm,hp = 0.0061 kw.
Now consider again month 7, the hottest month of the year; the monthly
average dry bulb temperature T is 71°F and the wet bulb temperature WBT is
£\\-*
55°F. The calculation is trial-and-error. If the cold water temperature to
the condenser is assumed to be 98°F, the warm water leaving the condenser,
based on Equations (4) and (5) in Section 10-7, is 123°F and the condenser
temperature is 133°F with a cooling load of 8,680 Btu/hr. Now with the
ambient wet bulb temperature at 55°F and the cooling tower characteristics
fixed at KaY/L =1.1 and L/G = 2.8, the cooling water leaving the tower
would have a temperature of 98°F, the same as the initially assumed tempera-
ture. No bypass is necessary in this month. Also, given L/G = 2.8 in the
202
-------
TABLE 10-4. WATER CONSUMPTION RATE PER SHAFT KW FOR AN ALL WET SYSTEM AT CASPER. WYOMING
MONTH
DBT (°F)
WBT (°F)
QQM (Btu/hr)
TWC (°F)
TWH (°F)
TR <°F)
WET FAN POWER (10~3Kw)
PUMP POWER (10~3Kw)
WATER CONSUMPTION
to (Ib/hr)
o
U)
QQD
-------
tower and L = 348 lb/(hr)(kw), the air flow rate through the tower can be
calculated as 124 lb/(hr)(kw), and the enthalpy of the air is increased by
a value of 70 Btu/lb. Assuming the air coming out from the tower is com-
pletely saturated, the change in humidity of the air is 0.055 Ib H 0/lb air;
the evaporation rate therefore is 6.84 Ib/hr.
The same calculation procedure can be repeated for the other months,
except for months 1 and 2, and the results are shown in Table 10-4 indicat-
ing a gradual drop in condenser temperature and a gradual reduction in fuel
penalty. For months 1 and 2, the calculation procedure can be repeated
except that bypass is required in order to keep the condenser temperature
at 115°P. For example,'in month 1 the cold and hot water temperatures of
the condenser are 82 and 106°F respectively in order to keep the condenser
at 115°F. With a wet bulb temperature of 20°F, the temperature of the cool-
ing water leaving the tower is 77°F. A bypass of the tower, therefore, is
necessary in order to maintain the temperature of the return cooling water
at 82°F.
Table 10-5 shows the results for a typical wet/dry combination. This is
a "50%" example, because the dry area is half the all-dry area, the wet area
is half the all-wet area, and the water circulation rate is half the all-wet
rate. The system is run with full fan power and no bypass for months 7 and
8. For four more months—5, 6, 9 and 10—the cooling tower is partly
bypassed. For the rest of the year the cooling tower is not used, and
the fan energy to the dry condenser is reduced.
From the month-by-month calculations the annual average figures were
taken for various wet/dry combinations and recorded on Table 10-6 as annual
consumption figures assuming 7,008 hrs/yr. The annual cost is given on Table
10-7. Unit costs are given on Table 10-1. The annual cost from Table 10-7
is plotted in Figure 10-8 as a function of water consumption for different
water costs, from free-of-charge to $2/10 gallons. Also shown in Figure
10-8 is the percentage of wet cooling at peak summer design condition, which
is a measure of the relative wet tower size. Figure 10-8 shows that for a
water cost of less than $0.68/10 gallons, it is more economical to have an
all wet system, whereas for a water cost of more than $0.68/10 gallons and
up to a value of $2/10 gallons it is more economical to have a wet/dry
204
-------
TABLE 10-5 WATER CONSUMPTION RATE PER SHAFT KW AT CASPER. WYOMING WITH 50Z WET LOAD AT PEAK DESIGN CONDITION
MONTH
DBT (°F)
WBT (°F)
QQtf (Btu/hr)
Twc (°F)
T,_. ( F)
WH v
TR (°F)
WET FAN POWER (10~3Kw)
PUMP POWER (10~ Kw)
to
0 WATER CONSUMPTION
01 (Ib/hr)
QQD (Btu/hr)
DRY FAN POWER (10~3Rw)
Tc (°F)
FUEL PENALTY (Btu/hr)
1
24
20
0
—
__
—
0
0
0
8200
5.30
115
0
2
26
22
0
—
*.*.
—
0
0
0
8200
5.94
115
0
3
32
27
0
—
^mt
—
0
0
0
8200
8.48
115
0
4
41
34
0
—
^^
—
0
0
0
8200
17.4
115
0
5
54
44
1217
105
112
92
1.14
8.56
0.88
6983
21.2
115
0
6
65
51
2475
95
109
91
2.36
8.56
1.87
5725
21.2
115
0
7
71
55
2957
91
108
91
3.04
8.56
2.28
5243
21.2
115
0
8
70
53
2957
91
108
91
3.04
8.56
2.28
5243
21.2
115
0
9
59
46
1788
101
111
90
1.45
8.56
1.32
6412
21.2
115
0
10
47
38
418
112
114
90
0.30
8.56
0,31
7782
21.2
115
0
11
32
27
0
—
_
—
0
0
0
8200
8.48
115
0
12
30
25
0
—
«• —
—
0
0
0
8200
7.53
115
0
A
DESIGN DBT - 96°F WET CONDENSER AREA - 0.540 ft_
DESIGN WBT - 60°F DRY CONDENSER AREA - 1.42 ft
WET CIRCULATION RATE - 0.348 CPU
-------
TABLE _10-6. EQUIPMENT SIZE AND ANNUAL POWER REQUIREMENT PER SHAFT KW FOR
TURBINE CONDENSER
WET LOAD FRACTION
2
DRY AREA (ft )
DRY FAN POWER (kw-hr/yr)
WET AREA (ft2)
COOLING TOWER SIZE (GPM)
PUMP POWER (kw-hr/yr)
WET FAN POWER (kw-hr/yr)
FUEL PENALTY (10 Btu/yr)
WATER CONSUMPTION (gal/yr)
0%
2.85
65.5
0
0
0
0
0
0
COOLING AT CASPER, WYOMING
25%
2.13
89.4
0.272
0.174
4.91
0.746
0
70.7
50%
1.42
105
0.540
0,348
30.0
6.60
0
627
75%
0.711
74.3
0.816
0.522
89.7
21.1
0.016
2580
100%
0
0
1.09
0.696
120
41.7
0.048
5100
ANNUAL OPERATING HOURS = 7008
TABLE 10-7. ANNUAL OPERATING
COST
($/( shaft
COOLING AT CASPER,
WET LOAD FRACTION
DRY AREA
DRY FAN POWER
WET AREA
COOLING TOWER
PUMP POWER
WET FAN POWER
FUEL PENALTY
TOTAL COST IF WATER COST FREE
TOTAL COST IF WATER COST IS
$1/10J GAL
TOTAL COST IF WATER COST IS
0%
8.72
1.31
0
0
0
0
0
10.0
10.0
10.0
25%
6.52
1.79
0.28
0.18
0.10
0.01
0
8.88
8.95
10.0
kw)(yr)) FOR
WYOMING
50%
4.35
2.10
0.55
0.37
0.60
0.13
0
8.10
8.73
9.35
TURBINE
75%
2.18
1.49
0.83
0.55
1.79
0.42
0.03
7.29
9.87
12.5
CONDENSER
100%
0
0
1.11
0.73
2.40
0.83
0.09
5.16
10.3
15.4
$2/10J GAL
206
-------
20
"T "T T T"
PLANT OPERATES 7008 HOURS A YEAR
$0.68/I03 GAL
Q
Z
O
O
O
co
LJ
UJ
100 °-
50
h-
liJ
LJ
O
<£
UJ
Q_
01 3 5
WATER CONSUMPTION (I03 GAL/KW-YR)
Figure 10-8.
Annual cost of steam turbine condenser cooling
at Casper, Wyoming.
207
-------
combination. Based on this economic consideration, the water consumption
rate for various water costs can be obtained and are shown in Figure 10-11.
Figure 10-11 shows that if the water cost for Casper, Wyoming is more than
$0.68/10 gallons, a wet/dry cooling combination is justified and the volume
of water consumed is reduced by a factor of 10. The month-by-month calcula-
tions and the annual operating costs and water consumption rates shown in
Tables 10-3 to 10-7 have been repeated for the climates of Beulah, North
Dakota and Farmington, New Mexico. The month-by-month and peak design tem-
peratures for these two locations are listed in Table 10-8. Figures 10-9
and 10-10 show results of the calculation with the annual operating cost as
a function of the water consumption rate for various water costs. The opti-
mal annual water consumption rates for various water costs are determined from
these two figures and are shown in Figure 10-11. Figure 10-11 shows that if
3 3
the water cost is more than $0.78/10 gallons and $0.68/10 gallons for
Farmington and Beulah, respectively, wet/dry cooling combination should be
used for the two locations and the annual water consumption rate can be
reduced by a factor of 10.
Figures 10-8 to 10-10 also show that the all wet system and the all dry
3
system would cost the same if the water cost is $0.90/10 gallons for Casper,
3 3
Wyoming, $1.05/10 gallons for Farmington, New Mexico and $1.10/10 gallons
for Beulah, North Dakota.
10.9 WATER CONSUMPTION FOR INTERSTAGE COOLING ON GAS COMPRESSION
The remaining major point of cooling is interstage cooling on gas com-
pressors. Each compressor requires a separate study. All that has been done
for this project is to study a three-stage air compressor designed to pump
air from the ambient temperature and 15 psia to 95°F and 90 psia, in which
condition the air enters the separation plant to be separated into oxygen
and nitrogen. Air compressors for this service are the largest compressors
in SNG plants and any conclusions reached about interstage cooling can safely
be applied also to the oxygen compressors.
Figure 10-12 shows the assumed design conditions for dry, wet and dry-
wet. Trim wet cooling is required in the dry section because the compressed
208
-------
TABLE 10-8. MONTHLY AVERAGE TEMPERATURE OF BEULAH. NORTH DAKOTA
AND FARMINGTON, NEW MEXICO
Month
1
2
3
4
5
6
7
8
9
10
11
12
DESIGN
PEAK
FARMINGTON, N.M.
DBT(°F) WBT(°F)
26
33
42
49
60
70
76
73
64
51
39
27
98
23
28
33
37
45
51
58
57
49
41
32
24
65
BEULAH, N.D.
DBT(°F) WBT(°F)
8
15
24
43
56
65
72
70
58
46
28
18
102
7
13
21
37
47
57
62
60
50
39
25
17
71
209
-------
20
15
PLANT OPERATES 7008 HOURS A YEAR
100
012345
WATER CONSUMPTION (I03 GAL/KW-YR)
Figure 10-9. Annual cost of steam turbine condenser cooling at
Farmington, New Mexico.
210
-------
"T T ~T "T
PLANT OPERATES 7008 HOURS A YEAR
1
I
1
L
o
o
CO
LJ
o
LJ
100 CL
<
Q
O
t-
LJ
50
UJ
o
ir
UJ
Q.
01234
WATER CONSUMPTION (103 GAL/KW-YR)
Figure 10-10. Annual cost of steam turbine condenser cooling at
Beulah, North Dakota.
211
-------
cc 6
ro
o 4
2
O
i= 3
a.
CO
o 2
o
cr
LJ
H
1 I I T | I \ 1 1 | 1 1 T
PLANT OPERATES 7008 HOURS A YEAR
CASPER, WY.
FARMINGTON.N.M.
BEULAH, N.
50 100
WATER COST (CENTS / IO3 GAL )
150
Figure 10-11. The effect of water cost on water consumed for cooling
turbine condensers.
212
-------
AMBIENT TEMP
15 pjlo
AMBIENT
T
AMBIENT
105aF
| 80'F
AMBIENT
AMBIENT TEMP
15 pilo
3 STAGE AIR COOLED AIR COMPRESSOR
I05*F
80*
80'
3 STAGE WATER COOLED AIR COMPRESSOR
I05*F
AMBIENT TEMP
•" i i
15p.Ia
0
-------
air must be reduced to 95°F before being admitted to the separation plant.
Table 10-9 gives the basis, the design conditions and the size of the resul-
tant compressor at New Mexico and at Wyoming for the wet and dry cases. The
2
capital cost of the installed compressors is $135/installed hp, plus $18/ft
installed dry heat transfer area, plus $5.I/ft installed wet heat transfer
area, plus $7/gpm circulated for the cooling tower.
In operation each stage of the compressor is assumed to be regulated at
the design pressure ratio and thus to consume less than design energy when
the temperature of gas entering the stage is below design temperature. For
the climates of New Mexico and Wyoming, month-by-month calculations were made
on the wet and on the dry systems to find the horsepower drawn by the compres-
sor and the water evaporated in the cooling tower. The results are shown on
Tables 10-10 and 10-11. The annual average compressor horsepower and the
annual total water evaporation rate were found. The operating costs, also
shown on Tables 10-10 and 10-11, were calculated as 18 percent of capital/yr
for compressors, 17 percent of capital/yr for dry area, 20 percent of
capital/yr for wet area and cooling tower (these percentages cover amortiza-
tion and maintenance), and also 1.74$/hp-hr ($2/10 Btu) for compressor drive
steam and 1.49C/hp-hr (2$Aw-hr) for auxiliary energy to run the circulating
water pumps and fans in dry coolers and cooling towers. It is clear that if
the only choice were all wet or all dry, then wet cooling would be required
at all sites. However, a combination of dry followed by wet cooling, as
shown on Figure 10-12, is practical.
The wet/dry combination was also studied for the climate of New Mexico
and Wyoming with the results shown on Tables 10-12 to 10-15 and Figures 10-13
and 10-14. When water costs $1.50 or more in New Mexico per thousand gallons
evaporated, interstage cooling should be designed with dry cooling to 140°F
followed by wet cooling to 95°F. The annual average water consumption will
3 3
be 10.9 * 10 gallons compared to 116 x 10 gallons for all wet cooling. In
Wyoming the change to wet/dry should be made when water costs about
$I/thousand gallons evaporated. In fact, wet cooling has been used in
New Mexico but combined dry/wet in Wyoming.
214
-------
TABLE 10-9. BASIS AND DESIGN CONDITIONS FOR COMPRESSOR
INTERSTAGE COOLING (ALL WET OR ALL DRY)
Basis: 2,000 Ib/hr air compressed from 15 psia and ambient
temperature to 90 psia and 95°F or less.
Design Conditions New Mexico Wyoming
Ambient temperature, F 98 96
Entry to stages 2 & 3 when dry cooling, F 140 140
Entry to stages 2 & 3 when wet cooling, F 95 95
2 dry area 136 131
2 28 28
Design Compressor
Dry cooling
Total hp 82.3 82.2
2
Ft dry area
2
Ft wet area
Gal/min cooling water circulation 1.7 I-7
Wet cooling
Total hp 78.1 78.0
Ft2 wet area 206 206
Gal/min cooling water circulation 16.1 16.1
215
-------
NJ
TABLE 10-10. SUMMARY OF COMPRESSOR ALL WET AND ALL DRY INTERSTAGE COOLING RESULTS IN NEW MEXICO
Month
1
2
3
4
5
6
7
8
9
10
11
12
Ambient
Temp (F)
26
33
42
49
60
70
76
73
64
51
39
27
Cooling Tower Heat
Removal Rate
(&tu/lb water evaporated)
1,719
1,643
1,541
1,459
1,346
1,261
1,231
1,273
1,333
'1,459
1,571
1,711
Compressor
Power
Wet Cooling
74.4
74.8
75.3
75.6
76.2
76.7
77.0
76.8
76.4
75.7
75.1
74.5
Operating
(hp)
Air Cooling
71.7
72.7
74.1
75.1
76.7
78.2
79.1
78.6
77.3
75.4
73.6
71.8
Evaporation Rate (10 gal/month
Wet Cooling
7.34
7.89
8.70
9.44
10.63
11.74
12.27
11.75
10.88
9.50
8.44
7.40
Air Cooling*
0
0
0
0
0.12
0.46
0.67
0.55
0.24
0
0
0
Annual average operating horsepower (hp):
Annual total water evaporation rate (103gal/yr):
75.7
75.4
Annual operating chargest
Capital 'related costs
Compressor operating steam costs
Auxiliary energy costs
Total annual operating costs
Dollars/year
Wet Cooling Dry Cooling
2,130.99
10,538.51
69.26
2,446.25
10,489.16
243.82
12,738.77 13,179.23
For equal charges water cost $3.87/10 gal.
116
2.04
•Zero water evaporation means exit air temperature from the third air cooler is lower than 95 F so that the
trim water cooler is not operated.
-------
TABLE 10-11. SUMMARY OF COMPRESSOR ALL WET AND ALL DRY INTERSTAGE COOLING RESULTS IN WYOMING
to
M
-J
Month
1
2
3
4
5
6
7
&
9
10
11
12
Ambient
Temp (F)
24
26
32
41
54
65
71
70
59
47
32
30
Cooling Tower Heat
Removal Rate
(Btu/lb water evaporated)
1,698
1,691
1,643
1,541
1,431
1,340
1,289
1,295
1,371
1,502
1.643
1,647
Compressor Operating
Power (hp)
Wet Cooling
74.35
74.46
74.75
75.22
75.88
76.44
76.75
76.69
76.13
75.52
74.87
74.75
Air Cooling
71.56
71.85
72.74
74.07
76.02
77.65
78.53
78.38
76.76
74.96
72.74
72.44
Evaporation Rate (10 gal/month
Wet Cooling
7.37
7.46
7.86
8.67
9.80
10.87
11.53
11.43
10.40
9.09
7.86
7.78
Air Cooling
0
0
0
0
0
0.34
0.55
0.51
0.15
0
0
0
Annual average operating horsepower (hp):
Annual total water evaporation rate (10^ gal/yr):
Annual operating charges:
Capital related costs
Compressor operating steam costs
Auxiliary energy costs
Total annual operating costs
75.47
Dollars/year
Wet Cooling Dry Cooling
2,128.09
10,505.00
68.05
2,429.96
10,413.21
235.40
12.701.13 13,078.57
74.81
110.12
1.55
For equal charges water costs $3.48/10 gal.
-------
TABLE 10-12. DESIGN CONDITIONS FOR COMPRESSOR DRY-FOLLOWED-BY-WET
INTERSTAGE COOLING IN NEW MEXICO
Basis: 2,000 Ib/hr air compressed from 15 psia and ambient
temperature to 90 psia and 95°F or less.
Design Conditions
Design temperature of air at exit
of the dry cooler, °F:
140
Design ambient temperature, F 98
Design entry to stages 2 and 3, °F 95
Design dry area, ft 115.5
Design wet area, ft 150.2
Design cooling water circulation, gpm 5.2
Design hp 78.1
TABLE 10-13. DESIGN CONDITIONS FOR COMPRESSOR
INTERSTAGE COOLING IN WYOMING
Basis: 2,000 Ib/hr air compressed from 15
temperature to 90 psia and 95°F or
Design Conditions
Design temperature of air at exit
of the dry cooler, °F:
o
Design ambient temperature , F 96
Design entry to Stages 2 and 3, F 95
2
Design dry area, ft 111.3
2
Design wet area, ft 150.2
Design cooling water circulation, gpm 5.2
Design hp 78.0
160 180
98 98
95 95
76.1 48.7
166.3 179.6
7.6 9.9
78.1 78.1
DRY-FOLLOWED-BY-WET
psia and ambient
less.
160 180
96 96
95 95
73.6 47.0
166.3 179.6
7.6 9.9
78.0 78.0
218
-------
TABLE 10-14. SUMMARY OF COMPRESSOR DRY-FOLLOWED-BY-WET
INTERSTAGE COOLING RESULTS IN NEW MEXICO
Basis: 2,000 Ib/hr air compressed
Air exit temperature of
dry cooler (°F):
140
160
180
Operating
HP
72.4
73.4
74.8
75.9
76.4
76.6
76.9
76.7
76.8
76.1
74.4
72.5
Water
Evap
0
0
0
0.4
1.2
2.2
2.7
2.4
1.5
0.5
0
0
Month
1
2
3
4
5
6
7
8
9
10
11
12
Annual avg operating
horsepower (hp): 75.3
Annual total water
evaporation rate
(103 gal/yr): 10.9
Annual operating
charges ($/yr):
Capital related costs 2,411.44
Compressor operating
steam costs 10,475.45
Auxiliary energy cost 5.86
Total annual operat-
ing cost
12,892.75
Operating
HP
73.9
74.5
75.5
76.2
76.1
76.6
76.9
76.7
76.3
75.7
75.1
74.0
Water
Evap
1.1
1.4
1.9
2.3
3.3
4.2
4.7
4.3
3.6
2.5
1.7
1.1
75.6
32.1
2,310.59
10,525.49
18.56
12,854.63
Operating
HP
74.7
75.4
75.2
75.5
76.1
76.6
76.9
76.8
76.3
75.6
75.1
74.8
Water
Evap
2.6
3.0
3.8
4.3
5.3
6.2
6.7
6.3
5.5
4.4
3.5
2.7
75.7
54.3
2,243.55
10,544.10
32.01
12,819.66
219
-------
TABLE 10-15. SUMMARY OF COMPRESSOR DRY-FOLLOWED-BY-WET
INTERSTAGE COOLING RESULTS IN WYOMING
Air exit temperature of
dry cooler (°F): 140
Month
1
2
3
4
5
6
7
10
11
12
160
180
Operating
HP
72.3
-72.6
73.5
74.8
76.0
76.4
76.7
76.6
76.4
75.7
73.5
73.2
Water
Evap
0
0
0
0
1.0
1.8
2.4
2.3
1.3
0.4
0
0
Operating
HP
73.7
73.9
74.5
75.5
75.8
76.3
76.6
76.6
76.0
76.1
74.5
74.3
Water
Evap
1.1
1.2
1.4
1.9
2.9
3.8
4.3
4.2
3.3
2.2
1.5
1.4
Operating
HP
74.6
74.8
75.4
75.1
75.8
76.3
76.7
76.6
76.0
75.4
75.4
75.2
Water
Evap
2.7
2.8
3.1
3.8
4.8
5.7
6.2
6.2
5.2
4.2
3.1
3.0
Annual avg operating
horsepower (hp): 74.8
Annual total water
evaporation rate
(103 gal/yr): 9.2
Annual operating
charges ($/yr):
Capital related costs 2,396.43
Compressor Operating
steam costs 10,431.61
Auxiliary energy cost 5.12
Total annual operat-
ing cost
12,815.16
75.3
29.3
2,300.56
10,485.81
17.53
12,803.89
75.6
50.6
2,235.93
10,526.62
30.85
12,793.39
220
-------
$3.87/l03GAL
0 20 40 60 80 100 120 140
|03GAL/(TON-AIR)(YR)
Figure 10-13. Operating cost of air compressor interstage cooling
in New Mexico (2000 Ib/hr air).
221
-------
13.2
20
40 60 80 100
K>3GAL/(TON-AIRHYR)
120
140
Figure 10-14.
Operating cost of air compressor interstage cooling
in Wyoming (2000 Ib/hr air).
222
-------
REFERENCES SECTION 10
1. Gold, H., Goldstein, D. J. , and Yung, D., "The Effect of Water Treatment
on the Comparative Costs of Evaporative and Dry Cooled Power Plants,"
U.S. ERDA, Division of Nuclear Research and Application, Report COO-
2580-1, UC-12, July 1976.
2. Skamser, R., "Coal Gasification, Commercial Concepts, Gas Cost Guide-
lines," U.S. ERDA, NTIS catalog FE-1225-1, UC-90C, January 1976.
3. Maddox, R. N. , Gas and Liquid^ Sweetening, p. 57, John M. Campbell and
Co., Norman, Oklahoma, 1974.
4. "Kelly's Handbook of Crossflow Cooling Tower Performance," Neil W. Kelly
and Associates.
5. Mickley, H. S., "Design of Forced Draft Air Conditioning Equipment,"
Chem. Eng. Progress 45, 739-745, December 1949.
6. Smith, E. C., and Gunter, A. Y. , "Cooling Systems Combining Air and
Water as the Coolant," ASME paper 72-HT-29.
7' Larinoff, M. W., and Forster, L. L., "Dry and Wet-Peaking Tower Cooling
Systems for Power Plant Application," ASME paper 75-WA/PWR-2.
223
-------
SECTION 11
WATER FOR MINE COMPLEX AND OTHER OFF-SITE USES
11.1 INTRODUCTION AND SUMMARY OF RESULTS
In this section is presented the general methodology for calculating the
amount of water consumed in the mining of coal and in any subsequent land
reclamation required. The water quantities for this category of usage gen-
erally will not be too strongly affected by the choice of conversion process,
except as it determines the actual quantity of material to be mined. How-
ever, the mine location and whether the mining is surface or underground will
be strong determinants of the quantity of water consumed. At all the sites
studied, surface mining is used.
The categories of water use are shown on Tables 11-1 to 11-3 with the
estimates made in the following paragraphs.
11.2 ROAD, MINE AND EMBANKMENT DUST CONTROL
Fugitive dust generated on haul roads and unpaved areas in the neighbor-
hood of the mine such as the mine benches and overburden placement areas must
be controlled. The length of unpaved haul roads and mine bench areas depends
on the mine productivity, as measured by the amount of coal recoverable per
unit area of stripped land. In the present study the following mine yields
are used: 3
Location 10 Ib/acres
Beulah, North Dakota 50,000
Gillette, Wyoming ' 180,000
4
Navajo, New Mexico 74,000
224
-------
TABLE 11-1. CALCULATION OF WATER FOR MINE COMPLEX IN WYOMING
Units: 103 Ib/hr
H
Coal mined 1
Water for road &
mine dust control
Water for handling &
dust control
Sanitary & potable
water
Service & fire water
*
Sewage returned
Revegetation
^Includes sewage from
ygas Syn thane SRC
,527 1,918 1,916
13 16 16
31 38 39
1.7 1.7 1.7
2.6 2.6 2.6
303 303 303
0 0 0
satellite town.
TABLE 11-2. CALCULATION OF WATER FOR MINE COMPLEX
Units: 103 Ib/hr
H
Coal mined 2
Water for road &
mine dust control
Water for handling &
dust control
Sanitary & potable
water
Service & fire water
Sewage returned
Revegetation
ygas S_RC
,108 2,380
52 59
42 47
2.0 2.0
3.0 3.0
1.5 1.5
0 0
Power
1,570
13
31
1.7
2.6
303
0
IN NORTH DAKOTA
Power
1,990
49
40
2.0
3.0
1.5
0
225
-------
TABLE 11-3. CALCULATION OF WATER FOR MINE COMPLEX IN NEW MEXICO
Units: 10 Ib/hr
Coal mined
Water for road and mine dust control
Water for handling and dust control
Sanitary and potable water
Service and fire water
Sewage returned
Re vegetation
Hygas
1,505
34
30
2.6
3.9
1.8
46
SRC
1,853
44
39
2.6
3.9
1.8
60
Power
1,470
33
29
2.6
3.9
1.8
45
In the assumed mine model the mining of 100 acres per year would require
2 miles of 45 ft wide unpaved haul roads to serve as spurs to conveyor belts
that would feed the coal to the plant. Such a belt line operation is
described in Reference 1. The bench area acreage that would have to be wet
down is approximately equal to four times the daily acreage that is mined.
The sum of the two unpaved areas determines the area where dust control must
2
be practiced. This area is 5_,320 ft /(acre mined/yr).
The simplest means of holding down fugutive dust is to wet down the mine
area and haul roads. It is assumed that the roads and mine area can be kept
in a wetted condition through an annual deposition of water equal to the net
annual evaporation rate. Any rainfall is taken to be an additional safety
factor. The annual pond evaporation rates for the areas examined are:
Location inches/year
Beulah, North Dakota 45
Gillette, Wyoming 54
Navajo, New Mexico 61
The lay-down rate can be calculated from the relation,
lay-down rate - disturbed area * evaporation rate
226
-------
That is, for 10 Ib coal mined,
lay-down rate, Ib = (10 Ib coal) * (acres mined/10 Ib coal)
x (5230 ft wetted/(acre mined/yr)) x (i ft/12 inches)
(wetting rate, inches/yr) x (62.4 Ib water/ft )
This equation gives
Location
Beulah, North Dakota
Gillette, Wyoming
Navajo, New Mexico
Water for road, mine and embankment
dust control (Ib water/lp3 Ib cqa_l_)_
24.5
8.2
22.4
For most of the processes the coal mining rate is equal to the coal uti-
lization rate as given in the various process description sections. However,
because the Lurgi gasifiers cannot accept fines, the coal mining rate for the
power plants is equal to 1.2 times the utilization rate. The fines are
assumed to be sold.
H-3 HANDLING AND CRUSHING DUST CONTROL
The water needs associated with the preparation of the coal are a part
°f the estimate of the water requirements for a mining operation integrated
with a synthetic fuel plant. In all coal preparation plants dust is gener-
ated in the stages of loading and unloading, breaking, conveying, crushing,
general screening and storage. The water required to hold down this dust
will be considered here.
The ways of preventing dust from becoming airborne are through the appli-
cation of water sprays or of nontoxic chemicals and the use of dry or wet dust
collectors with partial or total enclosure. It is assumed that the principal
generating sources will be enclosed and that where feasible, air will be
227
-------
circulated and dry bag dust collection employed. Whenever coal pulverization
is necessary this will be done under conditions of total enclosure with no
fugitive dust or hold-down water requirements. In inactive storage the use
of water for holding down dust can be minimized by the use of nontoxic chemi-
cals.
Despite the design precautions indicated, in large-scale plants with many
transfer points, transfer belts, surge bins, storage silos and active storage
sites it is necessary to employ water sprays to wet down the coal. This is
also generally necessary with breaking and primary crushing operations. An
4
examination of. the Wesco Lurgi plant design and the TOSCO oil shale plant
design indicates that a consumptive use of 1 Ib of water for every 50 Ibs
of coal handled and crushed is a reasonably conservative estimate.
11.4 SANITARY AND POTABLE WATER
A number of water requirements in the mine are a direct function of the
number of people employed in the mine. One such obvious requirement is water
for sanitary needs and potable usage. Of course the number of personnel is
related to the tonnage mined, but the number will also depend on the mine type
and location.
In fact, the quantity of coal used for each plant is sufficiently simi-
lar, and the water requirements dependent on people are sufficiently small
that the requirements may be taken to be independent of the process.
Location Mine Personnel (any plant)
North Dakota 270
Wyoming 230
New Mexico 300
The water used per man-shift and the fraction of that water which is not
recovered as sewage is dependent to some extent on the climate, and we have
used
228
-------
Sanitary and potable water used Consumed
Location (gallons/man-shift) (% of used)
Wyoming, North Dakota 30 25
New Mexico 35 30
The average hourly water rates have been calculated from the following
formula and entered onto Tables 11-1 to 11-3
Water used, Ib/hr = (number of men) * (gallons/man-shift)
x (5 shifts/week) x (i week/168 hrs)
x (8.33 Ib/gallon)
H.5 SERVICE AND FIRE WATER
The service water usage in the mine such as for equipment washing, main-
tenance, pump seals, etc., along with the fire water usage through evapora-
tion loss, is a difficult quantity to estimate. However, an analysis of a
number of mine designs indicates that this usage is essentially nonrecover-
able and can be related to the usage of sanitary and potable water.
The estimated ratio for service to sanitary usage for a proposed
*0 x 10 ton/yr surface mine near Gillette is about 1.6. This same figure
for the proposed Kaiparowits undergound mine is about 1.3, based on esti-
mated sanitary water usage. The two values are sufficiently close that the
average serving* water usage for the mine is 1.5 times the sanitary water
H§age. Moreover, all of the water is taken to be consumed since recovery
^ the mine work areas would prove quite difficult.
11.6 REVEGETATION
As part of any reclamation of mined land in arid and semi-arid regions,
there exists a potential requirement for supplemental irrigation water asso-
ciated with the establishment of soil stabilizing plant cover on mine spoils,
229
-------
It is concluded that coal mined areas with greater than 10 inches of mean
8
annual precipitation can be reclaimed without supplemental irrigation. Where
there is less than 10 inches of annual rainfall, partially reshaped coal mine
spoils can be successfully revegetated with supplemental irrigation of about
10 inches during the first growing season, with no further requirement during
g
subsequent growing seasons. Only at the Navajo, New Mexico site is irriga-
tion for revegetation required. The water requirement can be calculated from
the following formula:
revegetation water, Ib/hr = (Ib coal/hr) x (acres mined/74*10 Ib coal)
x (10 inches water) x (43,560 ft2/acre)
x (1 ft/12 inches) x (62.4 lb/ft3)
Revegetation water in New Mexico
30.6 Ib water/103 Ib coal
11.7 SATELLITE TOWN
At the Wyoming site (for an exemplary study) the process plants take in
the sewage water from a satellite town. The total employment in mine and
plant is about 1,000 people (see also Section 12) and the town will have a
population of about 7,000 people. According to Reference 10 the average
daily per capita water requirement in Wyoming is 175 gallons, of which 50
gallons are consumed and 125 gallons are recovered as sewage. About
0.87 x 10 gallons sewage/day is treated for plant intake water.
230
-------
REFERENCES SECTION 11
1. Wyoming Coal Gas Co. and Rochelle Coal Co., "Applicant's Environmental
Assessment for a Proposed Gasification Project in Campbell and Converse
Counties, Wyoming," Prepared by SERNCO, October, 1974.
2. Geological Survey, "Proposed Plan of Mining and Reclamation - Cordero
Mine, Sun Oil Co., Coal Lease W-8385, Campbell County, Wyoming," Final
Environmental Statement No. 76-22, U.S. Dept. of the Interior,
April 30, 1976.
3. North Dakota Gasification Project for ANG Coal Gasification Co., "Environ-
mental Impact Report in Connection with Joint; Application of Michigan
Wisconsin Pipe Line Co. and ANG Coal Gasification Co. for a Certificate of
Public Convenience and Necessity, Woodward-Clyde Consultants," Federal
Power Commission Docket No. CP75-278, Vol. Ill, March 1975.
4. Batelle Columbus Laboratories, "Detailed Environmental Analysis Concerning
a Proposed Gasification Plant for Transwestern Coal Gasification Co.,
Pacific Coal Gasification Co., Western Gasification and The Expansion of
a Strip Mine Operation Near Burnham, New Mexico Owned and Operated by Utah
International Inc.," Federal Power Commission, Feb.l, 1973.
5. Colony Development Operation, "An Environmental Impact Analysis for a
Shale Oil Complex at Parachute Creek, Colorado, Part 1 - Plant Complex and
Service Corridor," Atlantic Richfield Co., Denver, Colorado, 1974.
6« Atlantic Richfield Co., "Preliminary Environmental Impact Assessment for
the Proposed Black Thunder Coal Mine, Campbell County, Wyoming," and
"Revised Mining and Reclamation Plan for the Proposed Black Thunder Coal
Mine," 1974. Also "Black Thunder Mine, 10 Million Ton Per Year Water
Supply," (Personal Communication, Hugh W. Evans), Denver, Colorado,
March 6, 1975.
7- Bureau of Land Management, "Final Environmental Impact Statement Proposed
Kaiparowits Project," Chapter I, FES 76-12, U.S. Dept. of the Interior,
March 3, 1976.
8* National Academy of Sciences, Rehabilitation Potential of Western Coal
Lands. pp. 32,33, Ballinger Publishing, Cambridge, Mass., 1974.
9- Aldon, F. E., "Techniques for Establishing Native Plants on Coal Mine
Spoils in New Mexico," in Proc. Third Symposium on Surface Mining and
Reclamation, Vol. I, pp. 21-28, National Coal Association, Washington, D.C. ,
1975.
10• Metcalf & Eddy, Inc., Wastewater Engineering, p. 25, McGraw-Hill, New York,
1972.
231
-------
SECTION 12
ADDITIONAL ON-SITE WATER STREAMS
12.1 INTRODUCTION AND SUMMARY OF RESULTS
In this section is presented the method of calculating the remaining
plant water streams not discussed elsewhere. The results are presented on
Tables 12-1 to 12-3.
12.2 EVAPORATION
For the particular source waters chosen for this study a settling basin,
which is subject to evaporative losses, is not required. However, all plants
require a reservoir from which evaporation will occur. Net evaporation rates
(pond evaporation minus precipitation) are
Net evaporation (in/yr)
No control With control
North Dakota 30 23
Wyoming 40 30
New Mexico 53 40
The net evaporation rate with control assumes some form of evaporation con-
trol such as the application of monomolecular films to the water surface.
This control may also be a natural one resulting from the presence of impuri-
ties in the local waters. Experience indicates that a reasonable maximum
value to take for the effective reduction in the evaporation minus precipita-
tion rate is 25 percent.
232
-------
TABLE 12-1. CALCULATION OF ADDITIONAL WATER IN WYOMING
Reservoir volume, 10 gallons
Total bottom ash 10 Ib/hr
109Btu/hr
fly ash 103 Ib/hr
Reservoir Evaporation
Water for ash disposal
Water for in-plant dust contol
Service and fire water
Sanitary and potable water
Sewage returned
Hygas
36.5
102
_Syn .thane
36.5
27
0.053 0.02
10
Water
2.9
86
15
9.6
4.8
9.8
109
Requirements
2.9
55
19
9.6
4.8
9.8
SRC
36.5
114
0.07
0
103 Ib/hr
2.9
104
19
9.6
4.8
9.8
Power
50.
74
0.06
0
4.0
82
13
9.6
4.8
9. .8
233
-------
TABLE 12-2. CALCULATION OF ADDITIONAL WATER IN NORTH DAKOTA
Reservoir volume 10" gallons
3
Total bottom ash 10 Ib/hr
109 Btu/hr
fly ash 103 Ib/hr
Reservoir evaporation
Water for ash disposal
Water for in-plant dust control
Service and fire water
Sanitary and potable water
Sewage returned
Hygas
45
116
0.06
15
Water
3.6
98
21
9.6
4.8
9.8
SRC
20
130
0.07
0
requirements
2.1
107
23
9.6
4.8
9.8
Power
50
84
0.08
0
103 Ib/hr
4.0
119
17
9.6
4.8
9.8
234
-------
TABLE 12-3. CALCULATION OF ADDITIONAL WATER IN NEW MEXICO
Reservoir volume,
Total bottom ash
fly ash
10 gallons
103 Ib/hr
109 Btu/hr
103 Ib/hr
Hygas
45
226
0.125
22
SRC
20
265
0.15
0
Power
60
165
0.34
0
Water requirements 10Ib/hr
Reservoir evaporation
Water for ash disposal
Water for in-plant dust control
Service and fire water
Sanitary and potable water
Sewage returned
4.8
198
15
11.2
5.6
11.2
2.1
229
20
11.2
5.6
11.2
6.3
190
12
11.2
5.6
11.2
235
-------
At North Dakota and New Mexico the reservoir will hold the amount shown
on Tables 12-2 and 12-3, which is about one week's supply. At Wyoming, where
water is particularly scarce, the reservoir for the fuel plants will hold six
weeks' supply of treated sewage from the satellite town, that is,
36.5 x 10 gallons. For the power plant a large reservoir is used. Evapora-
tion control will be practiced at Wyoming and New Mexico, where it is parti-
cularly important, and will be assumed not to be practiced in North Dakota.
The reservoirs are assumed to be 30 feet deep, and the evaporation rates have
been calculated from the following formula and entered onto Tables 12-1 and
12-3.
Reservoir evaporation, Ib/hr
(reservoir capacity, gal)
x (8.33 Ib/gal) x (1/30 ft depth)
x (net evaporation, in/yr)
(1 ft/12 in) x (i yr/8760 hrs)
Reservoir evaporation, Ib/hr = 2.64 x 10 (reservoir capacity, gal)
x (net evaporation, in/yr)
12.3 ASH DISPOSAL
The ash which enters the plants with the coal leaves in two forms called
bottom ash and fly ash or top ash. In the power plants all the ash leaves as
bottom ash from the Lurgi gasifiers. In the SRC plants all the ash is
assumed to be bottom ash from the Koppers-Totzek gasifiers (in this case
some of the ash is recovered from the gas but is treated as bottom ash).
In the Synthane plant all the ash remains in the char fed to a boiler. In
this boiler, which is assumed to be similar to a pulverized coal fired dry-
ash furnace, 20 percent of the ash leaves as bottom ash and 80 percent leaves
236
-------
2
as fly ash. In the Hygas plants most of the ash leaves the bottom of the
gasifier. However, these plants also have coal fired furnaces in which 20
percent of the ash leaves as bottom ash and 80 percent as fly ash.
Water is used in bottom ash disposal to quench the ash and leave it wet
to avoid dusting in transportation to the disposal site. In each detailed
process description is given the enthalpy of bottom ash leaving the plant.
(For the Hygas plant gasifier ash residues sensible heats were obtained by
subtracting the higher heating value from the heating values tabulated.)
This enthalpy is measured above 77°F and it is only necessary to quench the
ash to about 200°F. Nevertheless, the quench water requirement has been
conservatively estimated using the stated enthalpies of bottom ash and taking
1000 Btu/lb water evaporated in quenching. It is also assumed that 0.3 Ib
water remains in each pound of bottom ash to prevent dusting, so the wet ash
leaves as 23 wt percent moisture.
Fly ash is assumed to be recovered from flue gases using dry electrosta-
tic precipitators ahead of the flue gas desulfurization scrub. The removal
process is dry, but when the ash is withdrawn from storage hoppers or silos,
water is sprayed into the screw conveyers to prevent dusting. The water spray
rate is 0.25 Ib/lb dry ash so the wet fly ash leaves as 20 wt percent mois-
ture.
12.4 IN-PLANT DUST CONTROL
Within the boundaries of any of the plants considered in the present
study, water will be needed for dust control at a certain number of points.
These points are similar to those described for the mines, namely, transfer
areas, active storage, surge bins, etc.
Somewhat less water would be required in the plants than in the mines,
since many of the operations tend to be enclosed. On this basis a good
assumption is a consumptive use of one half that applicable to the mine
a*eas, specifically 1 Ib of water for every 100 Ibs of coal handled and
jgansferred. This is a little less water than that deduced from the data
°f Reference 3.
237
-------
12.5 SANITARY, SERVICE AND FIRE WATER
These requirements are calculated as for the mine complex described in
Section 11. The plants each employ about 650 people (References 3 to 7).
Service and fire water is assumed here to be two times the sanitary and pot-
able water for plants, of which 65 percent of the water is assumed to be
returned as sewage.
238
-------
REFERENCES SECTION 12
1- Office of Water Resources Research, "Evaporation Suppression, a
Bibliography," WSRIC 73-216, Water Resources Scientific Information
Center, U.S."Dept. of the Interior, Washington, D.C., 1973.
2. Babcok & Wilcox, Steam - Its Generation and Use, 38th Edition, Revised,
Babcok & Wilcox Co., New York, 1975.
3. Bureau of Land Management, "Final Environmental Impact Statement Proposed
Kaiparowits Project," Chapter I, FES 76-12, U.S. Dept. of the Interior,
March 3, 1976.
4. Batelle Columbus Laboratories, "Detailed Environmental Analysis Concerning
a Proposed Gasification Plant for Transwestern Coal Gasification Co.,
Pacific Coal Gasification Co., Western Gasification and The Expansion
of a Strip Mine Operation Near Burnham, New Mexico Owned and Operated by
Utah International, Inc.," Federal Power Commission, Feb. 1, 1973.
5. North Dakota Gasification Project for ANG Coal Gasification Co.,
"Environmental Impact Report in Connection with Joint Application of
Michigan Wisconsin Pipe Line Co. and ANG Coal Gasification Co. for a
Certificate of Public Convenience and Necessity, Woodward-Clyde Consultants,"
Federal Power Commission Docket No. CP75-278, Vol. Ill, March, 1975.
6- Wyoming Coal Gas Co. and Rochelle Coal Co., "Applicant's Environmental
Assessment for a Proposed Gasification Project in Campbell and Converse
Counties, Wyoming," Prepared by SERNCO, October, 1974.
7- Colony Development Operation, "An Environmental Impact Analysis for a
Shale Oil Complex at Parachute Creek, Colorado, Part 1 - Plant Complex
and Service Corridor," Atlantic Richfield Company, Denver, Colorado, 1974.
239
-------
SECTION 13
SITE STUDIES - 1: WATER CONSUMPTION
For convenience in designing water treatment plants, all the water quan-
tity figures have been assembled on Tables 13-1 to 13-10. Except for cooling
water the quantities have been copied directly from Sections 5 to 9, 11 and
12. Cooling water is discussed below.
The steam turbine condensers for electric power generation will be wet
cooled in North Dakota and New Mexico and wet/dry in Wyoming. The wet cool-
ing loads can be found on Table 9-3 as the sum of the "process streams cooled
by water" and the "steam turbine condenser." The cooling rates on the tower
are those shown on Table 9-13 and in Section 9.6.
In North Dakota, for Hygas and SRC, the water is assumed to be cheap and
available. The wet cooling load is taken from the tables of ultimate disposi-
tion of waste heat and is the sum of the "wet cooling of plant process
streams," "total steam turbine condensers," "total compressor interstage
cooling," plus 10 percent of the "acid gas removal regenerator condenser"
load (as explained in Section 10.5). The cooling rate is taken from Table
10-2.
In New Mexico, for exemplary purposes the water is taken to be available
but salty. Furthermore, because of the salt, it is assumed cooling tower
blowdown cannot be admixed with ash for disposal but must be separately con-
centrated and disposed of in a segregated, lined evaporation pond. This
pushes the cost of cooling water to well over 78<=/thousand gallons evapor-
ated. Therefore, from Section 10, the steam turbine condensers will be
designed to be about 45 percent wet and 55 percent dry at the summer design
condition. The annual average water consumption is 10 percent of the all
wet case.
240
-------
In Wyoming water has been assumed to be very expensive if more than the
quantity of municipal sewage is used. All the tables for Wyoming show this
municipal sewage under the heading "sewage return." Maximum water conserva-
tion will be practiced. The annual average water consumption for turbine con-
densers will be 10 percent of the all wet case. The annual average water con-
sumption for interstage cooling on gas compressors will also be 10 percent of
the all wet case. For the Synthane plant the acid gas removal is a Benfield
type and only dry condensing is used.
To record the net water used a formal loss has been taken of 7 percent
of those streams requiring ammonia separation and 1 percent of those streams
requiring biotreatment. A formal loss has been taken of 5 percent of the
water entering boiler feed water treatment in North Dakota and Wyoming, and
10 percent loss in New Mexico where the water is brackish. Clean condensate
and other waters requiring flashing are assumed to have a 1 percent loss.
These are formal and approximate loss figures, and small differences will
occur from plant to plant. The effect of these losses on water consumption
is less than 5 percent.
Net water consumption is shown on Figure 2-1. Much of the most impor-
tant variation is in cooling water. This requirement is higher in the power
Plants than in the fuel plants simply because the power plants are less effi-
cient. In the power plants the load on wet cooling was reduced in Wyoming
over North Dakota and New Mexico because of the great expense of water, and
the result on total water consumption is to reduce it to about one half. In
the fuel plants the load on wet cooling was successively reduced from North
Dakota to New Mexico to Wyoming resulting in large changes in total water
consumption.
For the fuel plants the process water consumption depends on the process,
because all take in dried coal. All the SRC plants are net producers of pro-
cess water; the product fuel contains no more hydrogen than the coal. The
9as plants all consume water in the process as a source of hydrogen. The
Synthane process uses less water than the Hygas because all the coal enter-
ing the plant passes through the gasifier and contributes its hydrogen. On
the other hand, flue gas desulfurization water, assumed to be required only
in the gas plants, is higher in Synthane than in Hygas because bone dry char
241
-------
is burnt to raise steam. Much of the water consumed in wet flue gas desulfur-
ization is vaporized to saturate the flue gas.
The power plants use Lurgi technology and take in wet coal. The process
water consumption is a reflection of the moisture content of the coal. Where
the coal is wet, as at North Dakota, consumption is low; where the coal is
dry, as at New Mexico, consumption is higher.
The water consumed for dust control and ash disposal is mostly a frac-
tion of the ash content of the coal and is highest in New Mexico for all
plants.
242
-------
TABLE 13-1. APPROXIMATE WATER REQUIREMENTS
Plant Electric Site North Dakota
Plant Size 10 kilowatts
(Note that all water rates are for on-stream times. In calculating annual
average multiply by stream factor.)
103lb/hr 106gal/day
PROCESS
Boiler feed water into process 943 2.72
Foul water out of process 986 2.84
COOLING
Water evaporated for cooling 2470 7.12
FLUE GAS DESULFURIZATION 0
MINE AND OFF-SITE
mine and embankment dust control 49 0.14
Handling and crushing dust control 40 0.12
Sanitary and potable water 2 0.01
Service and fire water 3 0.01
Sewage returned 1.5 0.00
Revegetation 0 0
ADDITIONAL ON-SITE STREAMS^
Reservoir evaporation 4.0 0.01
Ash disposal 119 0.34
°ust control 20 0.06
Sanitary and potable water 4.8 0.01
Service and fire water 9.6 0.03
Sewage returned 9.8 0.03
(See Text) 2840 8.2
243
-------
TABLE 13-2. APPROXIMATE WATER REQUIREMENTS
Plant Hygas-SNG Site North Dakota
Plant Size 250 x 10 scf/day ; 10.09 x 10 Btu/hr
(Note that all water rates are for on-stream times. In calculating annual
average multiply by stream factor.)
3 6 gal/106 Btu
10 Ib/hr 10 gal/day product
PROCESS
Boiler feed water into process 1,015 2.92 12.08
Foul water out of process 294 0.85 3.50
Clean water out of process 201 0.58 2.39
COOLING
Water evaporated for cooling 790 2.27 9.40
FLUE GAS DESULFURIZATION 68 0.20 0.81
MINE AND OFF-SITE
Road, mine and embankment dust control 52 0.15 0.62
Handling and crushing dust control 42 0.12 0.50
Sanitary and potable water 2.0 0.01 0.02
Service and fire water 3.0 0.01 0.04
Sewage returned 1.5 o.OO 0.02
Revegetation 0 0 0
ADDITIONAL ON-SITE STREAMS
Reservoir evaporation 3.6 0.01 0.04
Ash disposal 98 0.28 1.17
Dust control 21 0.06 0.25
Sanitary and potable water 4.8 0.01 0.06
Service and fire water 9.6 0.03 0.11
Sewage returned 9.8 0.03 0.12
NET WATER USED
(See Text) 168o 4.8 20.0
244
-------
TABLE 13-3. APPROXIMATE WATER REQUIREMENTS
Plant Solvent Refined Coal
Site
North Dakota
Plant Size 10,000 tons/day
13.34 x 10 Btu/hr
(Note that all water rates are for on-stream times. In calculating annual
average multiply by stream factor.)
gal/10 Btu
10 Ib/hr 10 gal/day
product
PROCESS
Boiler feed water into process
Foul water out of process (1)
Water from gas purification
Water from gasification train
Clean water out of process
944
323
323
177
177
2.72
0.93
0.93
0.51
0.51
8.49
2.91
2.91
1.59
1.59
COOLING
Water evaporated for cooling
880
2.54
7.91
FLUE GAS DESULFURIZATION
0
MINE AND OFF-SITE
Road, mine and embankment dust control
Handling and crushing dust control
Sanitary and potable water
Service and fire water
Sewage returned
Revegetation
ADDITIONAL ON-SITE STREAMS^
Reservoir evaporation
Ash disposal
Dust control
Sanitary and potable water
Service and fire water
Sewage returned
59
47
2.6
3.9
1.8
0
0.17
0.14
0.01
0.01
0.01
0
0.53
0.42
0.02
0.04
1.62
0
2.1
107
24
4.8
9.6
9.8
0.01
0.31
0.07
0.01
0.03
0.03
0.02
0.96
0.22
0.04
0.09
0.09
WATER USED
(See Text)
Section 7.4, not Table 7-1.
1150
3.3
10.4
245
-------
TABLE 13-4. APPROXIMATE WATER REQUIREMENTS
Plant
Electric
Site
New Mexico
Plant Size 10 kilowatts
(Note that all water rates are for on-stream times. In calculating annual
average multiply by stream factor.)
103lb/hr 106gal/day
PROCESS
Boiler feed water into process
Foul water out of process
970
625
2.79
1.80
COOLING
Water evaporated for cooling
2476
7.13
FLUE GAS DESULFURIZATION
MINE AND OFF-SITE
Road, mine and embankment dust control 33
Handling and crushing dust control 29
Sanitary and potable water 2.6
Service and fire water 3.9
Sewage returned 1.8
Revegetation 45
0.10
0.08
0.01
0.01
0.01
0.13
ADDITIONAL ON-SITE STREAMS
Reservoir evaporation
Ash disposal
Dust control
Sanitary and potable water
Service and fire water
Sewage returned
6.3
190
15
5.6
11.2
11.2
0.02
0.48
0.04
0.02
0.03
0.03
NET WATER USED
(See Text)
3400
10.2
246
-------
TABLE 13-5. APPROXIMATE WATER REQUIREMENTS
Plant Hygas-SNG Site New Mexico^
6 9
Plant Size 250 x 10 scf/day ; 10.09 x 10 Btu/hr
(Note that all water rates are for on-stream times. In calculating annual
average multiply by stream factor.)
3 . gal/10 Btu
10 Ib/hr 10 gal/day product
PROCESS
Boiler feed water into process 1,015 2.92 12.08
Foul water out of process 294 0.85 3.50
Clean water out of process 201 0.58 2.39
COOLING
Water evaporated for cooling 389 1.12 4.63
FLUE GAS DESULFURIZATION 108 0.31 1.28
MINE AND OFF-SITE
Road, mine and embankment dust control 34 0.10 0.40
Handling and crushing dust control 30 0.09 0.36
Sanitary and potable water 2.6 0.01 0.03
Service and fire water 3.9 0.01 0.05
Sewage returned 1.8 0.01 0.02
Revegetation 46 0.13 0.55
ADDITIONAL ON-SITE STREAMS
Reservoir evaporation 4.8 0.01 0.06
Ash disposal 198 0.57 2.36
°ust control 15 0.04 0.18
Sanitary and potable water 5.6 0.02 0.07
Service and fire water 11.2 0.03 0.13
Sewage returned 11.2 0.03 0.13
NgTjWATER USED
(See Text) 1493 4.3 17.8
247
-------
TABLE 13-6. APPROXIMATE WATER REQUIREMENTS
Plant
Solvent Refined Coal
Site
New Mexico
Plant Size 10,000 tons/day
13.34 x 10 Btu/hr
(Note that all water rates are for on-stream times. In calculating annual
average multiply by stream factor.) 6
gal/10 Btu
product
10 Ib/hr 10 gal/day
PROCESS
Boiler feed water into process
Foul water out of process (1)
Water from gas purification
Water from gasification train
Clean water out of process
639
215
250
160
49
1.84
0.62
0.72
0.46
0.14
5.75
1.94
2.25
1.44
0.44
COOLING
Water evaporated for cooling
297
0.86
2.67
FLUE GAS DESULFURIZATION
MINE AND OFF-SITE
Road, mine and embankment dust control 44
Handling and crushing dust control 39
Sanitary and potable water 2.6
Service and fire water 3.9
Sewage returned 1.8
Revegetation 60
0.13
0.11
0.01
0.01
0.01
0.17
0.40
0.35
0.02
0.04
0.02
0.54
ADDITIONAL ON-SITE STREAMS
Reservoir evaporation
Ash disposal
Dust control
Sanitary and potable water
Service and fire water
Sewage returned
2.1
165
12
5.6
11.2
11.2
0.01
0.48
0.03
0.02
0.03
0.03
0.02
1.48
0.11
0.05
0.10
0.10
NET WATER USED
(See Text)
665
1.9
6.0
(1) Section 7.4, not Table 7-1
248
-------
TABLE 13-7. APPROXIMATE WATER REQUIREMENTS
Plant Electric ______^ Site Wyoming
Plant Size 10 kilowatts
(Note that all water rates are for on-stream times. In calculating annual
average multiply by stream factor.)
103 Ib/hr 106 gal/day
PROCESS
Boiler feed water into process 957 2.76
Foul water out of process 720 2.07
COOLING
Water evaporated for cooling 989 2.85
FLUE GAS DESULFURIZATION 0
MINE AND OFF-SITE
Road, mine and embankment dust control 13 0.04
Handling and crushing dust control 31 0.09
Sanitary and potable water 1.7 0.00
Service and fire water 2.6 0.01
Sewage returned* 303 0.88
Re vegetation 0 0
ADDITIONAL ON-SITE STREAMS
Reservoir evaporation 4.0 0.01
Ash disposal 82 0.06
Dust control 16 0.05
Sanitary and potable water 4.8 0.01
Service and fire water 9.6 0.03
Sewage returned 9.8 0.03
NET_WATER USED
(See Text) -1200 3.5
*Include3 sewage from satellite town.
249
-------
TABLE 13-8. APPROXIMATE WATER REQUIREMENTS
Plant Hygas-SNG Site Wyoming
Plant Size 250 x 10 scf/day ; 10.09 x 10 Btu/hr
(Note that all water rates are for on-stream times. In calculating annual
average multiply by stream factor.) ,
3 6 gal/10 Btu
10 Ib/hr 10 gal/day product
PROCESS
Boiler feed water into process 1,015 2.92 12.08
Foul water out of process 294 0.85 3.50
Clean water out of process 201 0.58 2.39
COOLING
Water evaporated for cooling 279 0.81 3.32
FLUE GAS DESULFURIZATION 104 0.30 1.24
MINE AND OFF-SITE
Road, mine and embankment dust control 13 0.04 0.15
Handling and crushing dust control 31 0.09 0.37
Sanitary and potable water 1.7 0.00 0.02
Service and fire water 2.6 0.01 0.03
Sewage returned* 303 0.87 3.60
Revegetation 0 0 0
ADDITIONAL ON-SITE STREAMS
Reservoir evaporation 2.9 0.01 0.03
Ash disposal 86 0.25 1.02
Dust control 15 0.04 0.18
Sanitary and potable water 4.8 0.01 0.06
Service and fire water 9.6 0.03 0.11
Sewage returned 9.8 0.03 0.12
NET WATER USED
(See Text) 851 2.5 10.1
*Includes sewage from satellite town.
250
-------
TABLE 13-9. APPROXIMATE WATER REQUIREMENTS
Plant Synthane-SNG Site Wyoming
Plant Size 250 x 10 scf/day ; 9.79 x 10 Btu/hr as gas
(Note that all water rates are for on-stream times. In calculating annual
average multiply by stream factor.) 6
3 6 gal/10 Btu
10 Ib/hr 10 gal/day product
PROCESS
Boiler feed water into process 1167 3.36 14.31
Foul water out of process 516 1.49 6.33
Medium quality water out of process 158 0.46 1.94
Clean water out of process 140 0.40 1.72
COOLING
Water evaporated for cooling 299 0.86 3.56
FLUE GAS DESULFURIZATION 269 0.77 3.30
MINE AND OFF-SITE
Road, mine and embankment dust control 16 0.05 0.20
Handling and crushing dust control 38 0.11 0.47
Sanitary and potable water 1.7 0.00 0.02
Service and fire water 2.6 0.01 0.03
Sewage returned* 303 0.88 3.72
Revegetation 00 0
ADDITIONAL ON-SITE STREAMS
Reservoir evaporation 2.9 0.01 0.04
Ash.disposal 55 0.16 0.67
Dust control 19 0.05 0.23
Sanitary and potable water 4.8 0.01 0.06
Service and fire water 9-6 0.03 0.11
Sewage returned 9-8 °-03 °*12
NET WATER USED
(See Text) 86° 2p^ 10.5
*Includes sewage from satellite town.
251
-------
TABLE 13-10. APPROXIMATE WATER REQUIREMENTS
Plant
Solvent Refined Coal
Site
Wyoming
Plant Size 10,000 tons/day
13.34 x 10 Btu/hr
(Note that all water rates are for on-stream times. In calculating annual
average multiply by stream factor.)
PROCESS
Boiler feed water into process
Foul water out of process (1)
Water from gas purification
water from gasification train
Clean water out of process
10 Ib/hr 10 gal/day
715
275
264
163
87
2.06
0.79
0.76
0.47
0.25
gal/10 Btu
product
6.43
2.47
2.37
1.47
0.78
COOLING
Water evaporated for cooling
229
0.66
2.06
FLUE GAS DESULFURIZATION 0
MINE AND OFF-SITE
Road, mine and embankment dust control 16
Handling and crushing dust control 39
Sanitary and potable water 1.7
Service and fire water 2.6
Sewage returned (2) 303
Revegetation 0
0.05
0.11
0.00
0.01
0.88
0.14
0.35
0.02
0.02
,72
ADDITIONAL ON-SITE STREAMS
Reservoir evaporation
Ash disposal
Dust control
Sanitary and potable water
Service and fire water
Sewage returned
2.9
104
19
4.8
9.6
9.8
0.01
0.30
0.05
0.01
0.03
0.03
0.03
0.94
0.17
0.04
0.09
0.09
NET WATER USED
(See Text)
100
0.29
0.90
(1) Section 7.4, not Table 7-1.
(2) Includes sewage from satellite town.
252
-------
PART 2 - WATER TREATMENT
SECTION 14
WATER ANALYSES
This section begins the second part of the report, which is devoted to
the quality and treatment of water. In this section are given the analyses
of various key water streams in coal conversion plants supply waters, efflu-
ent waters that need treatment for reuse and waters influent to boilers.
Circulating cooling water is also discussed.
Each subsection deals with a different water stream. The source water
at each of the three sites is given on Tables 14-1 to 14-4 as follows:
Lake water for North Dakota site - Table 14-1
Brackish water for New Mexico site - Table 14-2
Sewage from satellite town for
Wyoming site - Table 14-3
Additional river water for
Wyoming site - Table 14-4
The two analyses assumed for foul process condensate are given in Table
14-5. These are composite pictures made from the few analyses that are
available.
14.1 SOURCE WATERS
The effect of a brackish water and of municipal sewage as an intake has
been investigated as part of this study. A different water is assumed at
each site. In North Dakota the water will be drawn from Lake Sakakawea and
have the analysis shown on Table 14-1. This is good quality water.
253
-------
TABLE 14-1. ANALYSIS OF WATER FROM LAKE SAKAKAWEA, NORTH DAKOTA
mg/1
+2
Ca 49
+2
Mg 19
Na 59
HCO~ 180
S0~ 170
Cl~ 9
Silica 7
Suspended solids 2
Dissolved solids 428
TABLE 14-2. ANALYSIS OF BRACKISH GROUNDWATER NEAR
mg/1
Ca+2 12
Mg+2 13
Na+ 893
C0~2 45
HCO~ 408
S0~2 509
Cl" 770
NO~ 1
F~ 2
SiO,, 5.6
mg/1 (as CaCO-)
123
78
129
148
177
13
GALLUP, NEW MEXICO
mg/1 (as CaCO,)
30
53
1947
75
335
529
1086
1
6
Suspended Solids 10 (nominal)
Dissolved solids 2660
254
-------
TABLE 14-3. ANALYSIS OF SEWAGE AT WYOMING SITE
Raw sewage
BOD5
Suspended solids
Total N
Total P
Treated sewage
BOD5
COD
Suspended solids
Total N
Total P
Ca+2
+2
Mg
Na+
HCO~
S°42
el'
Si02
niccn1v*»r} solids
mg/1 mg/1 (as CaCO )
250
300
50 :'
12
30
170
113
20
9
60 150
25 103
150 327
330 270
110 114
145 204
40
830
255
-------
TABLE 14-4. ANALYSIS OF WATER FROM YELLOWSTONE RIVER NEAR HARDIN, MONTANA
mg/1 mg/1 (as CaCO-j)
Ca+2 46 115
Mg+2 16 66
Na+ 45 98
HC03 160 131
S0~2 147 153
Cl 7 10
Silica 12
Suspended solids 371
Dissolved solids 364
256
-------
TABLE 14-5. EXEMPLARY ANALYSES OF FOUL WATER FROM SRC AND FROM GAS PLANTS
BODr
COD
Phenol as C H OH
6 5
Hygas
and Lurgi
Plants
(mg/1)
13,000-18,000
25,000-30,000
4,000-6,600
Synthane
and SRC
Plants
(mg/1)
about 30,000
30,000-40,000
4,000-6,600
NH3 as N
HC03
Sulfide as S
Ca
Mg
Na
Cl
S°4
"These numbers apply to SRC only
are taken to be the same as in
xx Chloride concentrations
Wyoming
Synthane 600
Hygas 600
SRC 600
Lurgi 600
3,500-4,500
about 13,000
low
about 20
about 15
about 80
XX
low
In the Synthane water
the Hygas water.
Navajo
200
200
200
about 12,600
*
about 4,000
about 14,600
about 20
about 15
about 80
XX
low
these components
North Dakota
1200
1200
1200
257
-------
Furthermore, water will be assumed to be available at about 30C/thousand gal-
lons at the site which is about 25 miles from the Lake.
In New Mexico the source water will be brackish groundwater having the
composition shown on Table 14-2. This water will be pumped from a well and
is assumed to cost 40
-------
COD and Phenol
Some experimental analyses for gas plants using Western coals are given
on Table 14-6. In Appendix 1 our analytical procedures are listed. Clearly,
organic contamination is heavy, particularly phenol. There may be, however,
much oxygen demand other than phenol present. Phenol has a theoretical oxy-
gen demand of 2.38 (Ib/lb) and this is close to the measured COD and BOD.
Thus for the two Synthane analyses shown on Table 14-6, phenol represents
33-41 percent of the COD. For the Lurgi analyses phenol represents 44-90 per-
cent of the COD. The phenol in the coke oven water analysis given in Refer-
ence 2 is 68 percent of the COD. From the analyses of Central and Eastern
coals given in Table 14-9, phenol represents 21-46 percent of the COD. Stu-
dies have been made to discover the nature of the organic contaminants in con-
densate from the Synthane process at the Bureau of Mines. Oxygen-containing
molecules predominate. "Fatty acids" regularly reported in analyses of Lurgi
process water are not apparent in the results given. Fatty acids are deter-
mined on a sample steam distilled from acidified liquor by titration to a
phenolphthalein end point.
Although the organic contamination is assumed the same in water from the
Lurgi and Hygas processes, and the Synthane and SRC processes, this assumption
is only a first approximation. Logic, borne out by conversations with pilot
Plant engineers, says that processes having higher temperatures at the top
and longer residence times should give less organic contamination in the con-
densate water. Of the gasification processes the Lurgi and Synthane pro-
cesses might be expected to give the dirtiest water and the Hygas process
water to be notably cleaner. (Not all of the analyses shown here were avail-
able at the time the treatment plant designs were started. The phenol in
Hygas water is probably less than that shown on Table 14-5, which is a guess.)
The CO -Acceptor process, which has a fluid bed of long residence time, gives
a relatively clean condensate. Water from the high temperature Koppers-
Totzek process (discussed in the next section) is quite clean, and water
from the Bigas process would be expected to be similarly clean.
The rank of coal must also be expected to affect the water. Analysis
and experiment shows this to be true. Some analyses from Central and Eastern
259
-------
TABLE 14-6. ANALYSIS OF FOUL PROCESS CONDENSATE, WESTERN COALS
(mg/1 unless noted)
Process
Coal
Ref. or note
pH (units)
TDS
TDS (after ignition)
Total Carbon
Total Organic Carbon
Inorganic Carbon
Bicarbonate (HCO )
HCO~ (as meq/1)
BOD (5 days)
BOD (15 days)
BOD (20 days)
COD
Phenol as CgH OH
Fatty Acids as Acetic
Total Ammonia as N
Free Ammonia as N
Total Ammonia (meq/1)
Cyanide as CN
Thiocyanate
Total Sulfur as S
Sulfide
Chloride
Synthane
Wyoming
Subbit.
2
8.7
43,000
6,000
9,520
0.23
23
Synthane
North Dakota
Lignite
2
8.7
38,000
6,600
7,200
0.1
22
Hygas
Montana
Lignite
a, f
7,800
5,270
2,630f
13,400°
219C
13,000
14,000
17,500
30,000d
Hygas
Montana
Lignite
a, g
6,414
3,936
2,478b
12,600C
207°
13,590
|
3,618
5.5
ND*
*ND = none detected
2,7856
(continued)
260
-------
(TABLE 14-6 continued)
(mg/1 unless noted)
Process
Coal
Note
pH (units)
TDS
TDS (after ignition)
Total Carbon
Total Organic Carbon
Inorganic Carbon
Bicarbonate (HCO~)
HCO~ (as meq/1)
BOD (5 days)
BOD (15 days)
BOD (20 days)
COD
Phenol as C^H^OH
6 5
Fatty Acids as Acetic
Lurgi (ref. 3)
Rosebud, .Montana, Subbituminous
Coarse Graded Coal
h
9.6
4,030
45
I^IO3
5,642°
93°
9,900
22,700
4*200
1,250
Total Ammonia as N 4,385
Free Ammonia as N i 3,990
Total Ammonia (meq/1) j 258
Cyanide as CN
Thiocyanate
Total Sulfur as S
Sulfide
Chloride
2
6
150
122
45
i
8.3
1,765
35
5,307D
26,978°
442°
13,400
20,800
4,400
1,670
14,540
14,015
855
4
16
265
108
40
Fine Graded Coal
h
9.5
10,540
60
194D
982°
16°
9,100
16,600
6,300
1,390
1,720
1,180
101
3
85
160
10
i
9.8
1,650
25
4,284D
21,794°
357°
5,200
19,600
P 4,800
550
14,380
13,990
840
5
75
535
301
80 30
(continued)
261
-------
(TABLE 14-6 continued)
NOTES
a.
b.
c.
a.
e.
f.
h.
i.
j.
Two samples, analyzed by Water Purification Associates.
Calculated as (Total Carbon)-(Organic Carbon) (see Appendix 1).
Calculated as equivalent to inorganic carbon.
Suspect value, high compared to total organic carbon, see Appendix 1.
Total Kjeldahl nitrogen = 2,800 mg/1.
Absorption spectrographic analysis gave
Ca 17
Mg 12
Na 115
Absorption spectrographic analysis gave
Ca 61
Mg
Na
31
84
Emission spectrograph gave
Calcium
Sodium
Magnesium
Barium, Strontium
Aluminum, Boron
Titanium
Manganese, Iron, Zinc
Silicon, Vanadium
Chromium, Silver, Tin, Copper
result
relative
scale
high
low-medium
low-medium
trace-low
trace
faint trace-trace
faint trace
very faint trace
very very faint
trace
lo"1 - 10
10
10
lo
"1
~2
"2
- 10
- 10
10
- 10
-2
^
-4
Sample from inlet tar separator (labelled t in Ref. 3)
Sample from inlet oil separator (labelled o in Ref. 3)
Given as "carbonate as CO2"
in Ref. 3; converted to C for tabulation.
262
-------
TABLE 14-7. CONTAMINANTS IN PRODUCT WATER FROM COAL GASIFICATION
(HIGH-RESOLUTION MASS SPECTROMETER DATA)(REFERENCE 5)
Mass
(nominal)
79
93
94
107
108
110
117
120
121
122
124
129
132
134
135
136
138
143
144
146
148
150
157
158
186
Precise mass, amu
Meas.
79.004
93.058
94.041
107.071
108.055
110.036
117.057
120.056
121.088
122.036
122.073
124.052
129.055
132.055
134.035
134.070
135:144
136.051
136.087
138.066
143.057
144.056
146.069
148.052
150.013
150.069
157.172
158.073
186.066
Calc.
79.002
93.057
94.042
107.074
108.057
110.037
117.058
120.057
121.089
122.037
122.073
124.052
129.058
132.058
134.037
134.073
135.145
136.052
136.088
138.068
143.058
144.058
146.073
148.052
150.014
150.068
157.170
158.073
186.068
"A x io3
2
1
1
3
2
1
1
1
1
1
0
0
3
3
2
3
1
1 -
1
2
1
2
4
0
1
1
2
0
2
Formula
C5H5N
C6H7N
CgHsO
C7HgN
C7H80
C3Hs02
C8H7N
CgHgO
ca Hi i N
C7Hg02
CgHj^O
C-yHgC^
CgHyN
CgHgO
CQ Hg 02
CgHj^O
CgH^gN
CgHgO
C9H,20
CpI^oO
C10HgN
CloH80
C10H100
C9H802
CgHQOS
CgHjoOa
Cl i HI i N
C11HaoO
C12H10°
Possible compound type
Pyridine.
Methylpyridine.
Phenol.
Cs -pyridine.
Cresol.
Dihydroxybenzene.
Indole.
Acetophenone.
C3 -pyridine.
Benzoic acid,
hydroxybenzaldehyde.
C2 -phenol.
G! -dihydroxybenzene.
Quinoline.
Indenol.
Hydroxybenzofuran.
Indanol.
C4 -pyridine.
Cj -benzoic acid,
methoxybenzaldehyde.
C3 -phenol.
C2 -dihydroxybenzene.
Methylquinoline.
Naphthol.
Cj -indenol.
C.^ -benzofuranol.
Hydroxybenzothiophenol .
C2 -benzoic acid, C3-
hydroxybenzaldehyde.
C2-quinoline.
Cj^ -naphthol.
Biphenol.
* Difference between measured and calculated precise mass.
263
-------
TABLE 14-8. CONTAMINANTS IN PRODUCT WATER FROM COAL LOW-VOLTAGE
MASS SPECTROMETER DATA, ppm BY WEIGHT (REFERENCE 5)
Promoter. ....
Cresols
C2 -phenols. . .
C3 -phenols. . .
Dihydrics. . . .
Benzofuranols
Indanols
Ac e t oph enone s
Hydroxy-
benz aldehyde
Benzole acids
Naphthols. . . .
Indenols
Benzofurans. .
Dtbenzofurans
Biphenols. . . .
Benzothio-
phenols
Pyridlnes. . . .
Quinolines. . .
Illinois No. 6 (hvBb)
None
3,400
2,840
1,090
110
250
"70
} 150
/
^J
I 60
J
160
90
-
_
40
110
-
-
-
None
2,660
2,610
780
100
540
100
100
110
110
90
-
.
20
60
60
-
20
2 pet
limestone
1
2,640
1,760
560
70
140
70
100
100
180
70
10
-
-
100
210
20
40
2
2,890
2,720
800
130
60
120
140
40
290
110
10
-
-
150
240
-
60
2 pet
quick-
lime
2,490
2,470
660
90
170
100
140
60
170
90
30
-
-
130
180
20
40
5 pet
quick-
lime
1,130
3,580
1,170
150
300
80
130
210
270
60
10
.
110
20
250
10
70
Air
instead
of 02
2,770
2,860
860
110
80
110
150
40
210
90
20
-
-
120
580
20
40
Coal fed directly
into Rasifier bed
1
1,300
530
140
20
60
30
40
20
110
30
10
-
-
70
130
40
70
2
1,270
890
270
50
20
40
50
10
140
40
10
-
-
90
'270
30
100
3
1,000
930
330
50
20
50
60
10
180
30
80
10
10
120
320
100
110
Montana (sub)
None
3,160
870
240
30
130
80
140
-
160
70
10
-
-
-
270
20
70
Air instead
of Os,
1
4,480
2,100
440
60
70
100
70
-
180
80
10
-
-
-
160
10
70
2
4,230
2,400
560
80
-
120
80
-
160
70
10
-
-
-
270
20
70
N. Dak.
(lie)
None
2,790
1,730
450
60
70
60
110
40
140
50
10
-
-
10
220
10
30
Air
instead
of Os
3,590
1,450
260
40
170
50
50
-
30
30
-
-
-
-
300
-
^
Wyo.
(subj_
None
4,050
2,090
440
50
530
100
110
60
80
60
-
-
40
20
120
-
20
W. Ky.
(hvBb)
None
2,040
1,910
620
60
280
50
90
50
160
80
-
-
20
70
30
-
40
Pgh.
(hvAb]_
None
1,880
2,000
760
130
130
70
120
80
170
20
110
-
60
20
540
10
40
to
-------
TABLE 14-9. ANALYSIS OF FOUL PROCESS CONDENSATE,
CENTRAL AND EASTERN COALS
(mg/1 unless noted)
Process
Coal
Ref . or note
pH (units)
Total Carbon
Total Organic Carbon
Inorganic Carbon
Bicarbonate (HCOj
HCO (as meq/1)
Carbonate
BOD (5 days)
BOD (20 days)
COD
Phenol as C-H OH
o 5
Total Ammonia as N
Free Ammonia as N
Total Ammonia (meq/1)
Cyanide as CN
Thiocyanate
Sulfur
Chloride
Synthane
Illinois
No. 6
2
8.6
ii,ooob
180
6,000b
15,000
2,600
8,100
6,880
. 476
0.6
152
e
Synthane
Western
Kentucky
2
8.9
19,000
3,700
10,000
588
0.5
200
Synthane
Pittsburg
Seam
2
9.3
19,000
1,700
11,000
647
0.6
188
Hyqas
Illinois
No. 6
a
1225
702
523C
2,658d
46
2,682
3,016
3,000
273
7,200
422
f
260
Notes
a.
b.
c.
d.
e.
f.
Analysis by Water Purification Associates.
Not from same analysis (footnote in Ref. 2)
By difference (see Appendix 1) .
Calculated assuming equivalent to inorganic carbon
S= 400
SO* (as S)
SO* (as S)
S^j (as S)
SO= (as S)
120
467
571
118
265
-------
coals are presented on Table 14-9.
There is no reason to suppose that the analysis of water from oil produc-
ing processes (SRC, Synthoil, H-Coal) should be comparable to the analysis of
water from oxygen blown gasifiers (Lurgi, Synthane). However, the analyses
of SRC water given on Table 14-10 (for Kentucky coal) are quite similar to the
Synthane water given on Table 14-6 (except for animonia and sulfide contents),
and when discussing biological treatment a single example has been used for
SRC and Synthane.
Ammonia
Ammonia is a major contaminant in all the waters and it is derived from
nitrogen in the coal. If, however, one tries to calculate the concentration
of ammonia in the foul condensate on the assumption that all of the nitrogen
in the feed coal ends up as ammonia in condensate, the concentrations found
are 34,000-40,000 mg/1 for.Hygas, 30,000 mg/1 for Synthane and 25,000-57,000
mg/1 for SRC. In no process does even the major portion of the coal nitrogen
end up as ammonia in condensate, but in SRC a much higher fraction goes this
route than in gas processes. In the gas plants the calculations are clearly
dependent not only on the quantity of nitrogen in the coal but on the condi-
tions chosen which determine the quantity of condensate as well. At high
temperatures, as in the Koppers-Totzek process, ammonia is not formed and
free nitrogen is released. The appearance of free nitrogen cannot be veri-
fied in the Hygas pilot plant because all line purging is done with nitrogen.
Sulfur
Sulfur in coal can be expected to be mostly converted to H-S. Although
carbon dioxide, present in large amounts in gas plants, is a stronger acid
than hydrogen sulfide, equilibrium is not reached and some hydrogen sulfide
is found in condensate. In Western coals, where the sulfur content is low,
the sulfide content of the water is low and has been ignored in the design
of the ammonia separation equipment. H S is more volatile than C02 and will
be stripped out and removed with the COj. In the SRC plant C02 is not
266
-------
TABLE 14-10. ANALYSIS OF FOUL PROCESS COMPENSATE,
SOLVENT REFINED COAL
(mg/1 unless noted)
Process
Solvent Refined Coal
Coal
Kentucky
Ref. or note
pH (Units)
Total Carbon
Total Organic Carbon
Inorganic Carbon
BOD (5 days)
BOD (15 days)
BOD (20 days)
COD
Phenol as C^H^OH
6 5
Total Kjeldahl N
Total Ammonia as N
Total Ammonia (meq/1)
Cyanide as CN
Total Sulfur as S
Ca
Mg
Si
8.6
9,000
6,600
2,400b
32,500
34,500
>34,500
43,600
5,000
8,300°
7,900
465
10
10,500°
0.47
0.13
<0.5
8.2
8,160
7,390
770b
25,000-
30,000
12,000
15,000°
14,000
14,000
16,200°
Analysis by Water Purification Associates and Pittsburg and Midway.
By difference, see Appendix 1.
22 analyses for N and S made between 10/5/75 and 12/9/75 were
supplied by Pittsburg and Midway. Four of these analyses had
extreme values and were arbitrarily eliminated. For the remaining
18 analyses the average total nitrogen was 12,600 mg/1 with a
standard deviation of 7,000 mg/1 which is very random. The average
ratio (moles NH3)/(moles H2S) was 2.0 with a standard deviation of
0.17 which is quite reproducible.
267
-------
present in large amounts and H S is found in the water, associated with
ammonia.
Alkalinity
The bicarbonate shown on Table 14-5 for gas plants is not the result of
analysis. It is calculated as 0.75 moles C0_ per mole NH and is about as
£* J
much CO as can be expected judging from vapor pressure data given in Section
15.
Chloride
Chloride is of interest particularly if condensate is to be treated to
boiler feed quality. Chlorine in coal mineral is converted to HC1 by the
8,9
conversion process. The chlorine material balances on the Synthane PDU
gasifier reported in Reference 4 do not balance, but they do show that con-
i
densate is the only important way in which chlorine left the plant. Three
tests were done using Illinois No. 6 coals:
wt % Cl in coal mg/1 Cl in condensate
0.0093 300
0.0220 170
0.0093 190
The concentration in condensate is comparable to the value found for Hygas
using Illinois No. 6 (Table 14-9). The ratio of condensate to coal in these
tests is more than twice the ratio found in the material balance of Section 6.
The chlorine content of coals is not always given. It lies in the range
zero to 0.5 wt percent with zero often being found. We have no suggested
value for Navajo coal. For Wyoming coal 0.04 wt percent (dry basis) has been
reported in one case and for North Dakota lignite 0.2 wt percent (dry bas-
12
is) has been reported. If all the chlorine in the coal ended up as chlor-
ide in condensate, then material balances would give
268
-------
Cl concentration (mg/1)
Synthane: Wyoming 1,200
Hygas: Wyoming 1,400
North Dakota 8,000
SRC: Wyoming 3,000
North Dakota 13,000
Lurgi: Wyoming 600
North Dakota 2,000
These are high concentrations. The calculation depends as much on the
condensate rate (which depends on chosen process conditions and whether dry
or wet coal is used) as it does on the chlorine content of coal. Neither of
these numbers is well known. Wide variations in chloride concentration of
condensate are to be expected from coal to coal and from design to design.
The numbers shown on Table 14-5 are illustrative only. The chloride can pre-
vent distillation of ammonia. For biological treatment 450 mg/1 nitrogen
should be left in the water as nutrient. Chloride equivalent to this much
nitrogen has a concentration of 1,100 mg/1, so fixed ammonia may not be
troublesome.
Table 14-11 shows the concentrations in condensate of a wide variety of
elements when using Illinois No. 6 coal in a Synthane gasifier. Additional
analyses of metals are given in notes f and g of Table 14-6 for the Hygas
Process using Montana lignite, and on Table 14-10 for SRC using Kentucky coal,
The analyses are too few to be sure, but there seems to be a difference bet-
ween Montana lignite on one hand and between Illinois and Kentucky coal on
the other hand. There also seems to be no large concentration in any case.
The metal concentrations on Table 14-5 are approximate.
The analyses presented in Table 14-5 will be used for the design of the
water treatment plants.
269
-------
TABLE 14-11. TRACE ELEMENTS IN CONDENSATE FROM AN
ILLINOIS NO.. 6 COAL SYNTHANE GASIFICATION TEST
From Reference 2
No- 1 No. 2 Average (by weight)
Ppm:
Calcium 4.4 3.6 4
Iron 2.6 2.9 3
Magnesium 1.5 1.8 2
Aluminum 0.8 0.7 0.8
Ppb:
Selenium
Potassium
Barium
Phosphorus
Zinc
Manganese
Germanium
Arsenic
Nickel
Strontium
Tin
Copper
Columbium
Chromium
Vanadium
Cobalt
401
117
109
82
44
36
32
44
23
33
25
16
7
4
4
1
323
204
155
92
83
38
61
28
34
24
26
20
5
8
2
2
360
160
130
90
60
40
40
30
30
30
20
20
6
6
3
2
Run No.* 162 163 164
Ppm
S 5,000 5,000 5,000
Total carbon 16,000 16,000 16,000
B 43 1717 82 82
Cl 300 170 190
F 39 37 32
Na 6.6 6.8 5.4
Si 2,8 4.7 6.6
*From Reference 4. (Analyses by spark source mass spectrometer. Only S & C have
satisfactory material balances around the system,)
270
-------
14.3 CLEAN AND INTERMEDIATE PROCESS WATER
In addition to foul condensate the Hygas, Synthane and SRC plants have
other water effluent streams. A stream that has been labelled clean conden-
sate is obtained either from the water of reaction in raethanation, or surplus
water from the reverse of this reaction, namely, reforming in the SRC plant.
This water is very clean as it comes from very clean conditions. A sample of
methanation water is crystal clear and has no odor. Clean condensate will be
saturated with CO , CO, CH and H , but as it is always recovered hot and
under pressure these gases will be assumed to be removed by flashing. After
flashing, clear condensate is assumed to have less than 200 mg/1 dissolved
solids and to be suitable as boiler feed. The Hygas process has no other
stream.
Condensate is recovered from two other points in the Synthane plant:
after shift conversion and after acid gas removal. These streams will be
mixed to give medium quality condensate. Since most of the condensable and
water soluble contaminants in the gasifier off-gas appear in the foul con-
densate, and as no analyses of these downstream condensates are available,
medium quality condensate is taken to have one tenth of the concentrations
shown for foul condensate on Table 14-5.
In the Solvent Refined Coal plant condensate is obtained in the produc-
tion of hydrogen by gasifying residue in Koppers-Totzek gasifiers. Because
of the high temperature, condensate from the Koppers-Totzek process is quite
clean. Table 14-12 shows a particular analysis of condensate water after it
has been used to quench slag. Table 14-12 will be used for the SRC plants
in this study.
Also in the SRC plant live steam stripping is used in Selexol type acid
9as removal sections. There is a small acid gas removal plant for gas made
in the dissolving section, but over 90 percent of the acid gas removal occurs
in the reforming section where the gas is very clean, and in the Koppers-
Totzek section where the gas is quite clean. Condensed stripping steam will
contain carbon dioxide and H-S which must be flashed, but there is no reason
ft
to expect much organic matter, ammonia or salts. We have chosen to use the
approximate analysis of Table 14-13.
271
-------
TABLE 14-12. ANALYSIS OF WATER FROM KOPPERS COAL
GASIFICATION, KUTAHYA, TURKEY^
pH 8.9
Ca+2 159
Mg+2 68
Na 18
NH* 122
Cl" 46
S0~ 109
H2S Not detected
COD 63
Silica 43
TABLE 14-13. EXEMPLARY ANALYSIS OF CONDENSED STRIPPING
STEAM FROM ACID GAS REMOVAL IN THE SRC PLANT
mg/1
Ca 20
Mg+2 15
Na 18
Cl" 40
S0= 10
CO= 50
COD 30
Silica 2
272
-------
TABLE 14-14. SUGGESTED TOLERANCES IN BOILER WATER
Total
Pressure Dissolved Solids
(psig) (rag/1)
Slowdown Concentrations
300-450 3000
1000-1500 1000
Make-up Concentrations for 5%
300-450 150
1000-1500 50
(mg/1 as CaCO_) .
600
200
Blowdown
30
10
^uuao. iidj-uiicaa
(mg/1 as CaCO.)
60
0
0.3
0
Silica
mg/1
40 to 50
2 to 5
2 to 2.5
0.1 to 0.25
14.4 BOILER FEED WATER
Table 14-14 is a partial listing of acceptable concentrations in boiler
feed water modified from Reference 13. We will use Table 14-14. In the
Hygas and Synthane plants boiler feed water must be prepared to 1000 psig
quality. In the SRC plant 400 psig quality will suffice, as it will for
most of the makeup boiler feed water in the power plants. In the power
plants the makeup goes to the low pressure Lurgi gasifiers. The small
amount of water needed as makeup to high pressure steam used for the elec-
tricity generation will require an extra deionization step.
14.5 COOLING WATER
A wet cooling tower is an evaporator, and salts dissolved in the makeup
water will be concentrated, often to the point of precipitation. The preci-
pitate, which usually consists of carbonates, sulfates and phosphates of cal-
cium and magnesium together with silica, tends to adhere to heat transfer
surfaces forming a hard scale and lowering the heat transfer coefficient.
This must be prevented.
Not only may the makeup water contain silt but the circulating water, in
its passage through the tower, scrubs dust out of the air. Circulating water
273
-------
thus contains an ever-increasing amount of suspended matter which will settle
out in stagnant spots in the pipes and heat exchangers. This must also be
prevented.
Circulating cooling water is warm and well oxygenated. It is seldom
sterile when fed to the system and, in any case, receives a steady supply of
air-borne growth. Untreated cooling systems will be subject to fungal rot of
the wooden parts of the tower, bacterial corrosion of iron and bacterial pro-
duction of sulfide, and large growths of algae in the sunlit portions of the
14
tower. Biocidal chemicals must be added to control growth.
Finally, the well oxygenated circulating water can be very corrosive to
heat transfer surfaces.
Treatment of cooling water is intended to prevent the problems of scal-
ing, fouling, microbial growth and corrosion. In applying various treatments'
it is necessary to know the upper permissible limits of concentration of var-
ious dissolved substances so that the treatment procedures can be sized cor-
rectly .
Grits and Glover have published a detailed set of recommendations
which has been copied onto Table 14-15 together with the limits used in this
study.
(a) Suspended solids. We have used a medium figure from Grits and
Glover's recommendations.
(b) Calcium carbonate and bicarbonate ion. Because of the equilibrium
H + C0~ ** HCO~
the concentration of CO in solution depends on the concentration of H , that
is, on the pH. The solubility of CaC03 depends, therefore, on pH as well as
on temperature and the presence of other dissolved ions which affect the ionic
strength of the solution. A procedure for calculating equilibrium solubility
16 17 ^
has been given by Langelier and discussed by Larson and Buswell. Sisson
19
and Sussman have published nomograms based on Langelier's equation. Table
14-16 presents some solubilities calculated by Langelier's procedure at 122°F
(50°C) and 800 ppm total dissolved solids.
It must be emphasized that Table 14-16 gives equilibrium solubilities.
274
-------
TABLE 14-15. CONTROL LIMITS FOR COOLING TOWER
CIRCULATING WATER COMPOSITION
Conventional at
low pH
Suggested at
high pH with
high concentration
and dispersants*
Used in
this study
6.5 to 7.5
7.5 to 8.5
Suspended Solids (mg/1) 200 - 400
300 - 400
300
Ca x 0>3 (as CaC03)
1,200
6,000
**
6,000
Carbonates (mg/1)
Bicarbonates (mg/1)
50 - 150
300 - 400
300
Silica (mg/1)
150
150 - 200
150
Mg x Si02 (mg/1)
35,000
60,000
60,000
Ca x SO (as CaCOj
1.5 x 10 to
2.5 x 106
2.5 x 10° to
8 x 106
2.5 x 10
PO
20
Chlorides
3,000
20
* From Ref. 15
** More data needed to confirm (footnote from Ref. 15).
275
-------
TABLE 14-16. SOLUBILITY OF CaCO AT 50°C AND
*3
800 ppm TDS CALCULATED FROM REF. 16
**
Solubility with
* equal concentrations
Solubility Product of Ca and Alkalinity
6 830,000 911
6.5 262,500 512
7 83,000 288
8 8,300 91
9 830 29
10 83 9.!
* (Concentration Ca as ppm CaCO3l x (Alkalinity as ppm CaCO.,) .
** This is the square root of the solubility product.
276
-------
20
In fact, as is pointed out by Feitler, CaCO- will not form scale when it is
present at many times its equilibrium solubility. Calcium carbonate easily
supersaturates, and when precipitation does occur it is usually as a multi-
tude of very fine crystals which do not grow and which remain in suspension.
Furthermore, the lower the pH the lower the tendency to form scale. Figure
14-1 has been recalculated from Feitler. The critical pH curve for scale
formation was experimentally determined by Feitler for Los Angeles city water
and, as he points out, is not totally representative of other waters.
The upper limit of the Grits and Glover recommendation for the calcium
and carbonate solubility product has been used.
When air is passed through water containing bicarbonate, some carbon
dioxide is evolved converting the bicarbonate to carbonate. For a bicarbon-
ate limit a medium value from Crits and Glover has been used.
(c) Silica. The recommendation of Crits and Glover for the silica con-
21
centration (which is substantiated by Griffin from experience in the power
industry) has been used. Silica is removed from solution by coprecipitation
with magnesia. The Crits and Glover recommendation for the solubility product
Mg x SiO_ has been used.
(d) Calcium sulfate. A detailed review of the solubility of calcium
22
sulfate has been given by Glater. Two of the most complete studies are
23
those of Marshall and Slusher who studied the solubility in seawater con-
24
centrates at temperatures to 200°C, and Denman who studied the solubility
in cooling water. These authors give detailed procedures for calculating
solubility. The solubility of calcium sulfate depends not only on tempera-
ture but on the concentration of other ions present. In pure water, calcium
21
sulfate has a solubility of about 2100 mg/1 as CaCO-, giving a solubility
, 2 24
product of 2.4 x 10 as (mg/1 CaCO3) . According to Denman the solubility
of calcium sulfate in 5000 mg/1 NaCl is about 2800 mg/1 giving a solubility
product of 4.4 x io6 as (mg/1 CaCO3)2. A conservative solubility product of
2.5 x io6 as (mg/1 CaCO3) has been chosen.
The Solubility of Calcium Phosphate
Ortho-phosphate scaling can be troublesome if phosphate is present.
277
-------
107
O
O
D
O
x.
O
E
8
o
u
10s
\04
V N. FEITLER'S20
v x "CRITICAL pH"
\ "USUAL PRACTICE"
(EQUILIBRIUM
CURVE DISPLACED
BY ONE pH UNIT)
EQUILIBRIUM
SOLUBILITY
TABLE 4.2
CHOSEN LIMIT
i
I
6.0 6.5
I
7.0 7.5 8.0
SOLUTION PH
8.5
9.0
Figure 14-1. Scaling by calcium carbonate
(taken from Reference 20).
278
-------
Phosphate will be present if sewage plant effluent is used as makeup. Phos-
phorus and nitrogen must be present as nutrients in biotreatment. Phosphate
is added and ammonia is controlled by not stripping all of the ammonia out of
process condensate. The feed waters to the biotreatment plants contain about
450 mg/1 ammonia and 90 mg/1 P (about 270 mg/1 POJ. When biotreatment is
running smoothly (which cannot be guaranteed) it should be possible to reduce
the feed concentrations of N and P to 5 percent without lessening the carbon
conversion. In this study that biotreated foul condensate has been assumed to
contain about 20 mg/1 ammonia and 15 mg/1 PO . The degree of nutrient makeup
in the biological treatment seems not to have received much study, but it is
important if biologically treated foul condensate is to be used as cooling
tower makeup.
The salts, Ca (POJ-, calcium ortho-phosphate and Ca5(OH)(PO ) , hydroxy-
apatite, are the least soluble form. The salts CaHPO4 amd CaH PO are much
more soluble. Precipitation of calcium phosphate depends therefore on pH.
14
Details have been presented by McCoy and Table 14-17 is taken from this
reference. McCoy points out that precipitation can be slow and is complicated
by complexes of calcium with other phosphates. In fact, much higher concen-
26
trations than suggested by Table 14-17 are used. Kluesner and others found
that a water containing 70 mg/1 Ca as CaCO3 and 0.5 mg/1 ortho-phosphate as
P (1.5 mg/1 as PO.) can be concentrated 15 times. The pH of the sewage was
27
7.5. Harpel reports many successful uses of wastewater in cooling towers.
The organic matter in the waste, particularly in treated sewage, stabilizes
calcium phosphate. Concentrations of 30-50 mg/1 ortho-phosphate have been
successfully used. Phosphates, if stabilized, are usual corrosion inhibi-
tors.28 The limit given on Table 14-15 is arbitrary but seems to be quite
safe according to References 21, 27 and 28.
Phosphate is a nutrient and its presence will encourage yeast and fungus
growth. When phosphate is present biocide dosage will have to be
increased.27'29
Chloride Limitations
At some point the concentration of chloride in the circulating water
279
-------
TABLE 14-17. CALCIUM PHOSPHATE CONCENTRATIONS
AT VARIOUS pH VALUES (TAKEN FROM REFERENCE 14)
Maximum Ca (ppm as CaCO_) at equilibrium
ppm 0-P04 pH=5.0 pH=6.0 pH=6.5 pH=7.0 pH = 7.5
10 1.59 x 104 700 174 49 12
25 8.97 x 104 379 94 26 6.5
50 5.66 x 103 240 59 16 4.0
75 4.30 x 103 182 45 13 3.2
100 3.72 x 103 151 37 11 2.7
280
-------
limits further concentration. There are two limits to chloride concentra-
tion. The first is the corrosive effect on steel even with inhibitors pre-
sent. It has been suggested that a chloride concentration of 3000 mg/1 should
not be exceeded when using stainless steel. The second limitation on chlor-
ide is quite indefinite. Although very small, some drift will leave the tower
and settle on adjacent foliage. If the spray is salty and rainfall to wash
the leaves is limited, then plants can be damaged. For this study 3000 mg/1
chloride has been arbitrarily taken as the upper permissible limit regardless
of the material of the heat exhanger. Reference 25 indicates that for this
concentration the maximum salt deposition rate is approximately 25 Ibs/acre-yr
at about l^j miles from the tower and drops to 10 Ibs/acre-yr at 2 miles and
to 1 Ib/acre-yr at 5 miles. This should be compared to the natural salt dep-
osition rate of approximately 100 Ibs/acre-yr at 5 miles from the ocean and
15 Ibs/acre-yr at 25 miles inland.
Ammonia
All experience with treated sewage tower makeup includes experience with
27 29 30 30
ammonia. ' ' The problems have been summarized by Fleishman and are:
corrosion of copper, inactivation of zinc corrosion inhibitors, chlorine con-
sumption and nutrient for biological growth. The concentration of NH3 in the
circulating water does not increase over its makeup water value. Air strip-
Ping and biological oxidation maintain the NH concentration slightly higher
than in the makeup water. In this study copper is not used. Also, when
ammonia is present quite high values of organic carbon are expected at the same
time, so the use of oxidizing biocides is not planned. Table 14-15 shows a
somewhat high allowable concentration of ammonia, and if this concentration
is present increased biocide usage will be required.
281
-------
REFERENCES SECTION 14
1. "1975 Advanced Wastewater Treatment Seminar Manual" published by Clean
Water Consultants, El Dorado Hills, Calif, 95630.
2. Forney, A. J., et al. "Analyses of Tars, Chars, Gases and Water Found
in Effluents from the Synthane Process," Bureau of Mines (Pittsburgh)
Technical Progress Report 76, January 1974; also in Symposium Proceedings^
Environmental Aspects of Fuel Conversion Technology, St. Louis, 1974,
EPA-650/2-74-118.
3. Woodall-Duckham Ltd., "Trials of American Coals in a Lurgi Gasifier at
Westfield, Scotland," U.S. E.R.D.A. Res. and Dev. Report No. 105, Final
Report, November 1974 (NTIS Catalog FE-105).
4. Forney, A. J., et al. "Trace Element and Major Component Balances around
the Synthane PDU Gasifier," U.S. E.R.D.A., Pittsburgh Energy Research
Center, Technical Progress Report 75/1, August 1975; also in Symposium
Proceedings; Environmental Aspects of Fuel Conversion and Technology II.
Hollywood, Florida 1975. EPA-600/2-76-149, 1976. U.S. E.P.A. Research
Triangle Park, N.C.
5. Schmidt, C. E., Sharkey, A. G. and Friedel, R. A., "Mass Spectrometric
Analysis of Product Water from Coal Gasification," Bureau of Mines
(Pittsburgh) Technical Progress Report 86, December 1974.
6. Farnsworth, J. F. , Mitsak, D. M. and Kamody, J. F., "Clean Environment
with K-T Process," in Symposium Proceedings; Environmental Aspects of
Fuel Conversion Technology, St. Louis, Missouri, May 1974, EPA-650/2-74-H '
U.S. E.P.A. Research Triangle, Park, N.C.
7. Gloyna, E. F. and Ford, D. L., "Petro Chemical Effluents Treatment
Practices," February 1970, NTIS Catalog PB-205-824 (p. VI-7).
8. Schora, F. C. and Fleming, D. K., "Effluent Considerations in Coal Gasi-
fication," in Symposium Proceedings; Environmental Aspects of Fuel
Conversion Technology,11, Hollywood, Florida 1975, EPA-600/2-76-149, 1976,
U.S. E.P.A., Research Triangle Park, N.C.
9. King, J. G., Meries, M. B. and Crossley, H. E., "Formulae for the Calcu-
lation of Coal Analyses to a Basis of Coal Substance Free from Mineral
Matter," Trans. Soc. Chemical Industry (55) 277-81, 1936.
282
-------
10. Abernathy, R. F., and Gibson, F. H., "Rare Elements in Coal,"
Bureau of Mines Information Circular 8163, 1963.
11. SERNCO, Applicants' Environmental Assessment for a Proposed Gasifi-
cation Project in Campbell and Converse Counties, Wyoming, Prepared
for Wyoming Coal Gas Co. and Rochelle Coal Co., p. E-34, October 1974.
12. Michigan Wisconsin Pipe Line Co. and ANG Coal Gasification Co.,
"Application for Certificates of Public Convenience and Necessity,"
Exhibit Z-6, p. 4, Federal Power Commission Docket CP75-278, 1974.
13. Simon, D. E., "Feedwater Quality in Modern Industrial Boilers—A
Concensus of Proper Current Operating Practice," pp. 65-69, in
Proceedings 36th International Water Conference, Engineers' Society
of Western Pennsylvania, Pittsburgh, November 1975.
14. McCoy, J. W., The Chemical Treatment of Cooling Water, Chemical Pub-
lishing Co., 1974.
15. Crits, G. J., and Glover, G., "Cooling Slowdown in Cooling Towers,"
Water and Wastes Engineering, 45-52, April 1975.
16. Langelier, W. F., "The Analytical Control of Anti-corrosion Water
Treatment," J. Am. Water Works Assoc. 28, 1500-1521, 1936.
17. Larson, T. E., and Buswell, A. M., "Calcium Carbonate Saturation
Index and Alkalinity Interpretations," (with published discussion)
J. Am. Water Works Assoc. 34, 1667-1684, 1942.
18. Sisson, W., "Langelier Index Predicts Water's Carbonate Coating
Tendency," Power Eng. 44(2), 44, 1973.
19. Sussman, S., "Fundamentals of Cooling Tower Water Technology," Paper
140A, Cooling Tower Inst. Meet., February 1975.
20. Feitler, H., "The Effect of Scaling Indexes on Cooling Water Treatment
Practice," Proceedings of 35th Ann. Meet. Intl. Water Conf., Pitts-
burgh, Pennsylvania, 1974.
21. Griffin, R. W., "Water Management Trends in Refinery Cooling Systems,"
ASME publication 74-Pet-15, September 1974.
22. Glater, J., "Evaluation of Calcium Sulfate Scaling Thresholds,"
Cooling Towers, AIChE, 138-145, 1973.
23. Marshall, W. L., and Slusher, R., "Solubility to 200°C of Calcium
Sulfate and its Hydrates in Sea Water and Saline Water Concentrates,"
J. of Chem. and Eng. Data 13_ (No. 1) 83-93, January 1968.
24. Denman, W. L., "Maximum Re-Use of Cooling Water Based on Gypsum
Content & Solubility," Ind. Eng. Chem. 53 (10) 817-822, October 1961.
283
-------
25. Roffman, A., et al, "The State of the Art of Saltwater Cooling
Towers for Steam Electric Generating Plants," Westinghouse Electric
Corporation, Pittsburgh, Penn., U.S. Atomic Energy Commission Report
WASH 1244 (February 1973).
26. Kluesner, J., Heist, J., and VanNote, R. H., "A Demonstration of
Wastewater Treatment for Reuse in Cooling Towers at Fifteen Cycles
of Concentration," presented at AIChE Water Reuse Conference,
Chicago, 1975 (Bechtel Inc.).
27. Harpel, W. L., "Wastewater Reuse as Cooling Tower Make-Up," 34th
International Water Conference, Pittsburgh, 1973.
28. Gray, H. J., McGuigan, C. V., and Rowland, H. W., "Sewage Plant
Effluent as Cooling Tower Make-Up—A Continuing Case History,"
International Water Conference, 34th Meeting, Pittsburgh, 1973.
29. Ladd, K., "City Wastewater Re-Used for Power Plant Cooling and
Boiler Make-Up," p. 165 in Water Management by the Electric Power
Industry, Water Resources Symposium Vol. 8, Center for Research in
Water Resources, Route 4, Box 189, Austin, Texas 78757.
30. Fleischman, M. , "Reuse of Wastewater Effluent as Cooling Tower Make-Up
Water," p. 501, Second National Conference on Complete Water Reuse,
1975, AIChE.
284
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SECTION 15
WATER TREATMENT TECHNOLOGIES
15.1 INTRODUCTION
In this section are given brief descriptions of water treatment technolo-
gies not described elsewhere. The subsections are the names of the technolo-
gies. Not all possible technologies are included in this section. If a par-
ticular technology is used in the overall plant design, then it is discussed.
A few technologies are discussed even though they were not used in the plant
designs. Wet oxidation was investigated quite thoroughly; it worked well but
is too expensive compared to biological treatment at the organic loading of a
typical fuel process condensate. The use of oxygen instead of air was not
investigated, and this study is not the last word. Freezing is a treatment
technology potentially extremely useful for obtaining water purified of inor-
ganic and organic contamination simultaneously. Freezing is not developed
and has never been tried on waters of the type of process condensate. Freez-
ing is described here because of its future potential.
In this section a technology is discussed to arrive at its utility and
cost. Design information, to which the study does not contribute, will be
found in the references given.
!5.2 WET OXIDATION
This is a procedure for the destruction of organic matter dissolved or
suspended in water by oxidizing with air at high temperatures. The tempera-
tures used are above the normal boiling point of water (212°F) and the reac-
tion is carried out under pressure to prevent boiling. The pressure is
285
-------
usually 600 psig or above. Description and details will be found in Refer-
ence 1 and other references given therein.
The degree of oxidation achieved depends on the temperature and the
material oxidized. A sample of water from the Hygas pilot plant was tested
by Zimpro, Inc. in Rothschild, Wisconsin with the results shown on Table
15-1. It is apparent that satisfactory results were obtained. At 280°C
with catalyst the COD was reduced by 93 percent and all the phenol removed.
Wet oxidation is a chemical oxidation by 0 and the material destroyed
^
does not have to be biodegradable. Wet oxidation will therefore be particularly
useful when the waste is toxic, when the COD/BOD ratio is high and when the /
concentration is high. The cost of wet oxidation is not greatly dependent on
the concentration of the feed water and the process works well on concentrated
waters. This is quite different from biological oxidation. Costs estimated
4 5
for various throughputs in the range 1.5 x 10 to 6 x 10 gal/day suggest that
the capital cost can be estimated by
cost (?) = 650 (gallons/day)0'7
If 17 percent/yr of capital is charged for amortization and maintenance, this
formula gives
Throughput Amortization Charges
(10 gallons/day) ($/thousand gallons)
0.5 7.20
1.5 5.20
3.0 4.20
The energy requirement for air compression is about 0.3 kw-hr/lb COD
removed which is not high compared to biotreatment. However, the capital
cost is high. The capital cost does not depend on the COD concentration and
wet oxidation is particularly useful for very high COD (about 5 percent or
more). Wet oxidation has not been used in this study.
286
-------
TABLE 15-1. ANALYSES - HYGAS WASTEWATER WET OXIDATIONS
Sample
Temp . , C
Time, Kin.
COD, g/1
% COD Reduction
Total Solids, g/1
Total Ash, g/1
PH
NH3 as N, g/1
TKN, g/1
Total S, g/1
Total Halides as Cl, g/1
co2
Phenol, mg/1
Cyanide, mg/1
Thiocynate, mg/1
BOD5, mg/1
Catalyst
Feed
13.7
13.7
1.75
0.36
8.2
3.25
3.25
0.17
0.1
10.4
740
0
0
wn- —
694-56-1
240
60
4.9
64.2
1.36
0.34
8.2
2.84
3.04
0.13
0.1
5.8
<1,0
0
0
2,350
No
694-55-1
280
60
3. 3
75.9
1.16
0.37
8.1
2.88
3.01
0.17
0.1
6.2
<1.0
0
0
190
No
694-58-1
280
60
1.0
92. 7
2.46
.38
8.0
3.07
3.29
0.43
0.1
7.2
<1.0
0
0
Yes
287
-------
15.3 GRANULAR ACTIVATED CARBON ADSORPTION
Adsorption of the organic matter remaining in biologically treated con-
densate is one possible procedure for reducing the COD if treated condensate
is to be fed to a boiler. The procedure is expensive and will probably not
be used if the reuse of condensate is for anything other than boiler feed.
It is not known to what level soluble COD can be reduced by carbon adsorp-
tion, but it may not be good enough for boiler feed and some oxidative treat-
ment may have to follow, the carbon.
Adsorption with granular activated carbon is a well-known procedure which
will not be described here. Selected information, together with current
quoted costs of carbon, have been used to estimate the cost of treatment as
follows. Because of the very high COD of the feed water, a longer contact
time than usual has been assumed in municipal plants as has the use of
"pulsed" contactors which approach counter-current flow. The costs esti-
mated more nearly approach those common to sugar refineries than those com-
mon to municipal treatments. Thermal regeneration is used and no contribu-
tion is planned from biological regeneration, so long on-stream times for
the carbon are not required.
First, the secondary clarifiers in the biological treatment have been
costed without addition of flocculating agent and without filtration of the
overflow. This is satisfactory if the treated water is used as makeup to the
cooling water because side stream clarification of circulating cooling water
will be practiced. If biologically treated water is to go to carbon adsorp-
tion, then synthetic polymeric flocculants will be added to the clarifier
(at 5 mg/1) and the overflow will pass through some form of automatic back-
wash sand filter.
If Q is the feed flow in gallons/minute, the capital cost of the filter
is $175Q. This will be amortized at 15%/yr with an additional 2%/yr for
maintenance. The filter charge is
(0.17 x 175Q $/yr) x (1 yr/4.32 x 10 mins)
x (10 /Q, mins, thousand gallons) = $0.07/thousand gallons
288
-------
(This and all the other costs are summarized on Table 15-2.) Because all the
charges are proportional to flow, they are expressed in $/thousand gallons.
The charge for the flocculating agent is
($1.60/lb flocculant) x (5 Ib flocculant/106 Ib water)
x (8330 Ib water/thousand gallons)
$0.07/thousand gallons
The carbon contactors are assumed to have an open vessel residence time
of 120 minutes with an extra 120 minutes installed. Makeup carbon costs
3 3 ••
$0.4/lb or, at 35 Ib/ft , $14/ft . When considering installing contactors,
piping, valving, etc., $10 has been added to the cost of one cubic foot of
carbon. This will allow for vessels nearly twice the volume of the carbon
3
that they contain. Thus installed carbon costs $24/ft or $0.686/lb. In
TABLE 15-2. SUMMARY OF COSTS OF CARBON ADSORPTION
Q = flow (gallons/min)
x = COD removal (mg/1)
k = capital cost of regenerating
furnace ($/lb carbon/day)
Flocculation
Prefiltration
Contactors and Carbon
Regeneration Furnace
Fuel for Regeneration
Pumping energy
Replacement Carbon
Capital ($)
Expenses
<$/thousand gallons)
175Q
995Q
O.OSQxk
0.07
0.07
0.44
1.5 x
I0~5xk
1.7 x lo"4x
0.003
4.6 x io"4x
_
1170Q + 0.03Qxk 0.58 + 6.3 x 10 x
+ 1.5 x 10~5xk
289
-------
addition to the carbon in the contactors, an additional 120 minutes of carbon
at makeup carbon price is in the regenerator and piping. The capital cost of
contactors and carbon is
(Q/7.48, ft3/min) x (240 mins) x (24$/ft3) x (Q/7.48, ft3/min)
x (120 mins) x (14$/ft3) = $995Q
The operating charges are taken at 15%/yr amortization plus 4%/yr maintenance
which is $0.44/thousand gallons.
The cost of field-erected regenerating furnaces has been given in Ref-
erence 4 and can conveniently be expressed as
Capacity (Ib carbon/day)
less than 7,000
7,000 to 70,000
more than 70,000
Installed cost [$/(lb carbon/day)]
54
8930/(capacity)
14.3
0.577
Regeneration begins at 0.4 Ib COD/lb carbon. This is lower than might
be obtained if biological regeneration occurred but higher than recently
reported refinery experience. This somewhat high removal is justified
because of the high COD in the feed water and because a counter-current sys-
tem is used. The pounds of carbon regenerated per day are (where x is the
COD in mg/1)
(8.33Q, Ib water/min) x (x/106, Ib COD/lb water)
x (1/0.4, Ib carbon/lb COD) x (1440, min/day) = 0.03 Qx Ib carbon/day
If k is written for the capital cost in S/(lb carbon/day) then the capiji
tal cost of the regenerating furnace is $0.03 Qxk.
The charges for regeneration are taken as 15%/yr for amortization plus
another 7% for maintenance and are therefore 1.5 x 10 xk $/thousand gallons-
The assumed opened vessel capacity of 240 minutes means that 1120Q Ib
290
-------
carbon are installed. The regeneration rate is O.OSQx Ib/day, so the number
4
of days that the carbon is on line between regenerations is 3.7 x 10 /x. If
x = 10,000 mg/1, this means 3.7 days between regenerations which seems ade-
quate .
The energy for regeneration is about 4,500 Btu/lb carbon, divided bet-
ween fuel and steam, and is taken to cost $1.80/10 Btu, or
-4
1.69 x 10 x $/thousand gallons.
The pumping energy is enough to lift the flow through about 15 psi.
2
(With a linear velocity of 5 gpm/ft , a 120 minute contact time means 80 ft
total for the carbon. The pressure drop is assumed to be 2 inches water per
foot of carbon and 7 psi is added for other pressure drops.) At 2C/kw-hr
this costs $0.003/thousand gallons.
Replacement carbon is 5.5 percent of the regenerated carbon and costs
-4
4.58 x 10 x $/thousand gallons.
In keeping with the other cost estimates, labor and laboratory charges
are ignored. Some costs are shown on Figure 15-1. The effect of throughput
on $/thousand gallons is not significant. Above about 3,000 mg/1 COD removed,
the costs become very high. Carbon adsorption is only suitable for low con-
centrations. The assumed carbon loading gives an energy requirement for
regeneration of about 11,000 Btu/lb COD removed. This energy might make
about 1.1 kw-hr of electricity and is about twice the energy needed to remove
1 Ib BOD in an activated sludge plant. The energy is 93.7x(Btu/thousand gal-
lons) and is
Energy for carbon regeneration
COD removed (mg/1) (10 Btu/thousand gallons)
3,000 0.3
10,000 0.9
For high removal rates Figure 15-1 tends to a cost of ll$/lb COD
removed. This can be compared with the cost of biotreatment of 2.1<=/lb BOD
removed. There is very little published experience on the use of carbon for
coke plant wastes. VanStone presents an adsorption isotherm and a cost
estimate made in 1972. Costs tend to 8* to 9$/lb phenol removed, which is
291
-------
10
z
o
o 6
O
o
r T
3XI06 gal/day
O.SxlO6 gal/day
i
j L
0 2,000 4,000 6,000 8,000 10,000
COD REMOVED (mg/€)
Figure 15-1. Costs of granular activated carbon adsorption.
292
-------
about 4«/lb COD removed. This is low. VanStone uses lower unit costs than
here, but the most important difference is the use of a carbon loading of
about 0.3 Ib soluble organic carbon/lb carbon, or about 1 Ib COD/lb carbon.
Other Adsorption Possibilities
Wet oxidation, described in Section 15.1, has been investigated for the
regeneration of powdered activated carbon. It has proved satisfactory. If
this procedure is used to replace granular activated carbon, the cost of con-
tractors is much reduced. The furnace is not required and the fuel is
reduced. The replacement carbon is not much altered and the cost of the wet
oxidation reactor must be added. This replacement does not seem worthwhile
unless the COD removed is more than about 8,000 mg/1.
The possibility of selective adsorption of phenol on a synthetic poly-
meric adsorbent is discussed in Section 17. The general use of synthetic
polymeric adsorbents is not included in this study.
Synthane char is being studied as an absorbent at the Bureau of Mines.
15.4 FREEZING
When dirty water is frozen the ice crystals formed tend to exclude con-
taminants and to be very pure water. A practical consideration is that it is
much easier if the fraction of water recovered as ice is low enough to leave
all of the original contaminants in solution in the unrecovered water. Con-
taminants should not precipitate. With foul process condensate water, about
90 percent of the water might be recoverable. An advantage of freezing is
that all sorts of contaminant molecules are excluded, organic as well as inor-
ganic, ionized and unionized. The energy consumption is expected to be less
than 70 kw-hr/thousand gallons in direct contact freezing; this is equivalent
to 0.7 x 10 Btu/thousand gallons assuming 10,000 Btu/kw-hr to generate elec-
tricity.
However, no large-scale freezing unit is available nor have tests ever
been made on these waters. Freezing is not used in this study, but laboratory
investigatory research to determine applicability is strongly recommended.
293
-------
15.5 EVAPORATION
Evaporation is a procedure for removing water from nonvolatile dissolved
materials. Evaporation is well known and widely practiced. Evaporation of
g
these waste waters has been studied by Bechtel Corporation and seems to work.
8
Skrylov and Stenzel distilled wastewater after phenol extracting and after
steam stripping to remove ammonia and carbon dioxide, and after addition of
alkali to hold down volatile acids. They report that for the lab tests "the
caustic required was substantial." The wastewater altered with age and the
best distillate was obtained on fresh samples. That is, with age the organic
contaminants break up into more volatile fractions. With a 95 percent recov-
ery of water the distillate contained only 4 percent of the total organic
carbon in the feed. (The feed contained 330 mg/1 TOG and the distillate con-
tained 14 mg/1 TOC. This is approximately equivalent to reducing COD from
1100 to 47.) The remaining TOC stayed in the remaining 5 percent of the
water which was not distilled over. If the first 5 percent of the water dis-
tilled was rejected so that a 5-95 percent cut was taken, the organic carbon
in the recovered distillate was halved.
The TOC in the feed is lower than is assumed in some of the waters in
this study and represents what is left after solvent extraction. The TOC in
the distillate is still too high for boiler feed. It was found that it could
be further reduced by oxidation with ozone. Ozonation was slow and 150 min-
utes were required to reduce TOC from 14 to 3. Skrylov and Stenzel concluded
that a good arrangement might be to extract and then strip all of the waste-
water of ammonia. Some of the water would then be evaporated and ozonated
for boiler feed water, and the rest of the water would probably require bio-
logical oxidation to make it fit for makeup to the cooling tower.
In the range of interest of 0.5 x 10 to 3 x 10 gallons /day, costs are
proportional to throughput and independent of concentration and so can be
conveniently measured in $/thousand gallons. Capital cost is about
$2/ (gallons/day) and, taking amortization and maintenance at 17%/yr £ or
300 days/yr the charge is about $1.13/thousand gallons. The fuel
depend on the plant design. Considering as an example a multistage flash
evaporator using waste steam recovered from the coal conversion, and if the
294
-------
6
price of waste steam is Sl.80/10 Btu, then Reference 1 suggests that the
optimum designs are:
Throughput No. of Btu/lb of Fuel cost
(10 gal/day) flash stages water removed ($/thousand gals)
0.5 22 136 2.00
1.0 25 120 1.80
3.0 37 81 1.20
The remaining charges, particularly including the cost of alkali, are not
known. They are taken as $1.00/thousand gallons, bringing the total cost to
between $4.10 to $3.40/thousand gallons depending on throughput.
Throughput Total cost Energy
(10 gal/day) ($/thousand gals) (10 Btu/thousand gals) (10 Btu/day)
0.5
1.0
3.0 .
4.10
3.90
3.40
1.1
1.0
0.7
0.6
1.0
2.0
These numbers are shown on Figure 15-2.
Distillation is energy intensive but, for the foul waters that are being
treated here, so are many other treatments. One of the ammonia separation
possibilities uses about 37 Ib steam per 100 Ib feed water in the deacidifi-
cation and ammonia towers. This is 370 Btu/lb water treated or
3.1 x 10 Btu/thousand gallons. The energy to supply oxygen in an air-
activated sludge biological treatment plant is about 54 kw-hr/thousand gal-
lons depending on the BOD. If it takes 10,000 Btu/hr to generate 1 kw of
electricity, this means 0.5 * 10 Btu/thousand gallons, but this energy must
be supplied at a high level, not as waste heat.
If it works, distillation should be considered when condensate is to be
returned to the boilers. Prerequisites for distillation are ammonia removal
and solvent extraction. Biological treatment is probably not only not
required but may be harmful in that large, nonvolatile organic molecules tend
295
-------
rn
z
m
70
o
-------
to be broken by biological treatment to smaller, volatile molecules.
15.6 TREATMENT OF CIRCULATING COOLING WATER
In this section is a brief description of the usual treatment required
for circulating cooling water with the references in which details are given.
Treatment is intended to prevent scaling, fouling, microbial growth and cor-
rosion and will be described under these headings.
Prevention of Scaling
The usual way to prevent scaling when high cycles of concentration are
used is to remove calcium, magnesium, bicarbonate, phosphate and silica.
Since calcium and magnesium are referred to as "hardness," their removal is
called "softening." Bicarbonate is sometimes referred to as "alkalinity,"
although this term encompasses other alkalis. Scaling can be prevented even
when precipitation occurs by the addition of chemicals which inhibit crystal
growth and hold the microcrystals in suspension. These chemicals are dis-
cussed under the heading "Prevention of Fouling."
The simplest procedure is to add sulf uric acid (H SO ) . This acid is
much stronger than carbonic acid and carbon dioxide is released from the
water according to the equation
H SO + 2HC03" •*• 2H20 + S04= + 2C02 i
One pound of HCO ~ requires 0.8 Ib H2SO4 for treatment. The cost of sulf uric
acid and other chemical costs are given in Table 15-3.
TABLE 15-3. CHEMICAL COSTS
Delivered costs used in this study
Sulf uric acid 3.8
Lime 2 . 7
Soda ash ' 4.0
Dolomite 2.0
Alum 5.3
297
-------
A common softening procedure is precipitation with lime (Ca(OH) ) and
soda ash (Na CO ). The quantities of lime and soda ash required are esti-
9 12
mated from the following equations. '
Calcium and dissolved carbon dioxide and bicarbonate are first precipi-
tated together:
Ca(OH) + CO -> CaCO 4- + HO
£ £, J £.
)2 + Ca(OH)2 + 2CaCO34- + 2H O
The solubility of calcium carbonate is dependent on pH. Lime is usually
12
added to raise the pH to about 9.4 and calcium is reduced to about 12 mg/1
(as Ca) . The concentration of calcium and magnesium left after lime-soda
treatment is less the higher the temperature, and sidestream treatment should
be done preferably on the hot water.
If removal of calcium in excess of bicarbonate is required, soda ash
must be added:
Ca + Na2C03 •*• 2Na+ + CaC034-
If bicarbonate is present in excess of calcium, the magnesium will pre-
cipitate :
Mg(HC03)2 + 2Ca(OH)2 ->•
More magnesium can be precipitated by adding more lime:
Mg++ + Ca(OH)2 -»• Mg(OH) * + Ca++
Since the solubility product (Mg++) x (OH~)2 tends to be constant, the resi-
dual magnesium in solution can be lowered by raising the concentration of
hydroxyl ions, that is, by raising the pH. This is shown on Table 15-4. In
practice a pH of 10.5 or higher is used and magnesium is reduced to about
2.5 mg/1 (as Mg) . This procedure adds calcium to the water, unless soda ash
298
-------
TABLE 15-4. SOLUBILITY OF MAGNESIUM AT DIFFERENT pH (TAKEN FROM REF. 13)
pH * 9 pH - 10 pH = 11 pH - 12
Mg++ (mg/1) 216 2.16 0.0216 0.00216
is used.
Silica coprecipitates with magnesium hydroxide at a rate of about 1 gm
^^
SiO_ per 7 gm Mg . (This figure was supplied by A.B. Mindler, Permutit
Research and Development Center, Princeton, New Jersey, and is also given by
Kluesner et at. If morn silica removal is required then magnesium will
have to be added. The ability to remove silica with magnesium at a cooling
water temperature of 80°P (25°C) or above is illustrated in Reference 11,
page 84.
When the pH is raised to about 11, phosphate removal is quite complete
and the equation is
5Ca (OH) + 3HP0" •»• Ca (OH) (P0> + 60H~ + 3HO
indicating a consumption of 5 moles of lime for 3 moj.es of P. An alternative
view is that the precipitate is calcium orthophosphate Ca3(PO4)2 and not
hydroxyapatite Cac(OH) (PO.) , which indicates a requirement of 4.5 moles of
15
lime for 3 moles of P. In fact, these differences in lime requirement
would not be readily distinguishable.
Calculation of lime requirements must include the lime needed to raise
the pH.
After treatment the water is neutralized with sulfuric acid. Carbon
dioxide is sometimes used, but this decreases the net removal of carbon
dioxide and in sidestream treatment increases the size of the sidestream.
Lime-soda softening involves a small investment in chemical feeders and
mixers and a large investment in clarifiers to settle out the precipitates.
Cost-estimating curves have been kindly supplied by the following companies:
Door-Oliver, Inc.; Envirotechj F.M.C. Corporation; Graver Water Division of
299
-------
Ecodyne Corp.; and Permutit Co., Inc. Installed clarifier costs have been
i
estimated using the cheaper of steel or concrete tanks. The concrete cost
was taken to be $175/cubic yard and includes excavation, backfill, concrete,
concrete forms, rebar and finish. Steel tank costs were in part supplied by
Bethlehem Steel. The cost curve used in this study is given on Figure 15-3.
To aid sedimentation in the clarifiers, synthetic flocculants were
assumed to be added at a rate of 1 mg/1. Flocculant costs were supplied by
Dow Chemical, Rohm and Haas, and the Tretolite Division of Petrolite Corpora-
tion. A cost of 1.3$/thousand gallons of sidestream has been used.
Prevention of Fouling
All cooling towers scrub dust out of the air. In this study we have
used a standard rate of input of dust to the water using a dust concentration
in the air of 10 Ib/ft (Reference 17) and the tower conditions given in
Section 10. The calculation is
1400 Btu 1 Ib water circulated/hr
Ib water evaporated 25 Btu/hr
X 1 Ib air passed x 380 ft3 10~6 Ib dust
2.8 Ib water circulated 29 Ib air _ 3
ft air
-4
= 2.6 x 10 Ib dust transferred/lb water evaporated
17
About 1/5 of the dust transferred is suspended in the circulating water;
the rest settles in the basin or is not trapped. The suspended dust input
rate is 0.052 Ib/thousand pounds of water evaporated. Because of the dust
scrubbed from the air, the concentration of 300 mg/1 suspended solids in
the circulating water dictates a maximum of 6.8 cycles of concentration even
with crystal clear makeup water, unless sidestream clarification is used.
In many cases some form of clarification has proved necessary. When
clarifying the effluent from biotreatment, a filter has been assumed. Effl°*
ent from biotreatment has about 100 mg/1 suspended solids. The installed
costs of automatic backflushed, gravity sand filters have been supplied
300
-------
V)
O
O
0.4 gpm/ft used for dust only
2
0.6 gpm/ft used for lime plus dust
2
1.25 gpm/ft used for lime only
OJ
o
<
600
500
400
300
200
100
23456
FLOW (I0sgpm)
8
10
Figure 15-3. Clarifier costs,
-------
by Graver Water Division of Ecodyne, Environmental Elements Corporation Divi-
sion of Koppers Company, and Permutit Co., Inc. The supplied costs from these
sources, which included tanks, mechanisms, installation and sand, were all
2 2
very close to each other. We have used $150/ft at 2 gpm/ft for a cost of
$75/gpm. Excess capacity is not used because the filters are all multiple
units, and automatic continuous backwash is included in the cost.
In clarifying a cooling tower sidestream where the suspended solids are
300 mg/1, a filter has been used on low flow streams, but the flow rate is
2
reduced to 1 gpm/ft increasing the cost to $150/gpm. Above about 2,000 gpm,
clarifiers (shown on Figure 15-3) are cheaper and are used. The overflow
from a clarifier is assumed reduced to 100 mg/1 S.S.
Synthetic polymeric dispersant chemicals (also called antifoulants) are
added to the circulating water. Information on dosage and cost of dispersants
has been supplied by Drew Chemical Corp. and Hercules, Inc. A cost of
3£/l,000 gals sidestream + blowdown has been used. The cost basis is the
sidestream because the flocculants added to the clarifiers will destroy
the dispersants.
Control of Microbial Growth
Microbial growth must be prevented by the addition of microbicides to
the circulating water. The constant presence of air, nutrients and addition-
al growth introduced from the air renders makeup and sidestream treatments
ineffective. Unrestricted growth of microorganisms can lead to fouling of
heat transfer surfaces and plugging of pipes; it can also lead to corrosion
of wood in the cooling towers and corrosion of metal surfaces, particularly
by bacteria that metabolize sulfur. Descriptions of the microorganisms that
22
are found in cooling water are given in References 18 to 21. Sussman has
recommended prompt disinfection if the total plate count is higher than about
6
10 /ml and states that most systems can operate adequately with plate counts
less than 500,000/ml. A very large number of biocidal chemicals are avail-
able. They are usually used in combinations to prevent acclimatization of
the microorganisms and often used in the form of a "shock treatment" in
which a high dose is added for a short while and then the concentration is
302
-------
allowed to die down. This produces high effectiveness at the lowest cost.
Chlorine is probably the most common microbicide used today. It has
disadvantages. It is most active at acid pH, which is not planned because
of corrosion. Shock treatment will overcome this problem to some extent. It
reacts with ammonia when sewage water is used as makeup, which lowers its
activity. It also reacts with oxidizable organic molecules which will be
present. Chlorine is toxic. Less toxic nonoxidizing chemicals have been
chosen because organic molecules are usually present and because blowdown is
usually dumped with ash.
A very large number of nonoxidizing microbicides have been used ' '
and only a few of the most popular will be described here. Quaternary ammo-
nium salts are cationic surface active agents with a wide range of microbici-
dal activity. They are the least toxic to animal life. Several quaternary
ammonium compounds are approved for use in dairies. The most commonly used
in dairies are alkyldimethylbenzyl ammonium chloride, alkylpyridinum bromide
23 24
and cetylpyridinum chloride. Schultz has described two quaternary ammo-
nium salts and a diamine which have been found to be nonpersistent and to
behave well in cooling towers. He has tested a trialkylbenzyl ammonium chlo-
ride (this group includes the first of the dairy sanitizers mentioned above),
a tetra-alkyl ammonium chloride and an aliphatic diamine. Because quater-
nary ammonium compounds are surface active, they are removed from effective-
ness in solution by oily or turbid waters. Quaternary ammonium compounds
when used with organotin compounds have a synergistic effect and a wide range
of microbicidal activity.
Methylenbisthiocynate has broad range microbicidal activity at low con-
centrations. It is not affected by dirt, hardness and oil. It loses all
activity after one hour at pH 8, which makes blowdown disposal convenient
but makes it useless for some alkaline waters.
Costs of biocides having EPA, USDA and similar approvals as nondanger-
°us materials have been supplied by Buckman Laboratories, Drew Chemical Corp.,
Gamlen Chemical, Hercules, Inc. and Lonza, Inc. Costs lie in the broad range
°f 15C to $1.50 per thousand gallons of blowdown. The major mechanism of
loss of biocide is in blowdown. An average cost of 50C/1,000 gallons of
blowdown has been used in this study for water without a lot of organic
303
-------
carbon, phosphate or ammonia. When these nutrients are present the cost of
biocide has been increased (see Section 19 for the individual cases).
Corrosion Control
18 25 26
Corrosion in cooling systems is discussed in many references.
In this study carbon steel is assumed to be used in the heat exchangers so
corrosion control will be necessary. However, in most of the plants the
cooling tower blowdown is assumed to be used for dust control and ash dis-
posal, so chromium, or indeed any heavy metal, should be available for cor-
rosion control. Experience on controlling corrosion without chromium is
27
varied. In 1975 Brunn made the following statement:
"Our goal is to find a replacement for chromate/zinc which can be
purchased from a supplier. Five years have been spent on this pro-
gram, and it is continuing. None of the materials tested during
the first two years even approached the desired qualities. During
the past two years there have been some promising candidates which
are applicable to some company systems, but none which are univer-
sally acceptable. In general, these fall into the category of
organic phosphate plus some inorganic salt such as zinc or inorgan-
ic phosphates. In most of these materials corrosion inhibition is
found to be relatively good, but there is a significant increase
in deposition on heat transfer surfaces. This is caused by more
demanding microbiological control, precipitation reactions between
treatment and circulating water, and fouling which results from
laydown of the additional corrosion products because of the slight
loss in corrosion protection. Our initial movement to nonchromate
inhibitors is slow, but in the past two years we have made signi-
ficant moves toward nonchromate inhibitors."
It must be understood that the procedures tentatively adopted for this
study are not fully proven. This is the more so because zinc, as well as
chromium, is avoided because of the presence of ammonia.
In using nonmetallic corrosion inhibitors, certain practices seem to be
necessary and are followed in this study. The circulating cooling water is
best kept just alkaline and a lot of effort is made to control scaling and
fouling. Furthermore, throughout the plants the temperature on the opposite
side of the heat transfer surface from the cooling water is below about 300°*'
304
-------
28
Petry has reported on the importance of temperature. Finally, it will be
necessary to passivate the new system at start-up when it is still clean.
Many tests have been reported including those by von Koeppen and
others, Hwa, and Gesick. Harpel and Donahue31 found adequate corro-
sion resistance using orthophosphate in the range 1-5 mg/1, total phosphate
4-10 mg/1 and phosphonates. The conditions necessary for success were
(i) prior passivation with chromate, (ii) PH 7 to 9, and (iii) calcium more
than 50 ppm as CaC03 (Langelier index 1 to 2.5). The phosphonates prevented
scale formation by calcium carbonate and calcium phosphate. Furthermore,
the phosphate concentration was not so high as to increase the growth of
algae.
Based on cost information kindly supplied by Drew Chemical Corp., Gam-
len Chemical and Hercules, Inc., the cost of nonchromium corrosion inhibition
chemicals lies in the range of 35
-------
related to both molecular weight and molecular size, as determined by steric
geometry. Size is the more important; branched chain molecules are better
34
rejected than straight chain molecules. Lindemann has found that reverse
osmosis of sewage plant effluent renders the water suitable as feed to the
boiler makeup demineralizers which says that carbon removal is adequate.
Reverse osmosis membranes will also partly reject silica which is impor-
tant in the treatment of both cooling water and boiler feed water.
A cost estimate is given on Figure 15-4. This shows cost as a function
of throughput and independent of concentration. It is meant to apply when
concentrating to about 2 percent total dissolved solids and the throughput is
a measure of the recovered water, not the feed water. A similar cost curve
(not exactly the same conditions) has been published for 80 percent recovery
from brackish water (2,000-5,000 mg/1 TDS feed, 1-2.5 percent TDS maximum
32
concentration) and this also is given on Figure 15-4 along with the cost
estimates from Reference 32 for sea water desalting in a two-stage plant. In
this study the cost curve from Reference 1 has been used unless the concen-
tration is pushed above 2 percent. For higher concentrations higher costs
have been used which are given for each individual case. The energy require-
ment, which is mostly to pump the feed to about 800 psi, is about
7.2 kw-hr/thousand gallons.
15.8 ELECTRODIALYSIS
Electrodialysis is a membrane separation process in which dissolved
ionic impurities are removed from water through membranes under the driving
force of a dc electric field. Only dissolved, charged particles are
removed and concentrated. Unlike reverse osmosis, electrodialysis has no
effect on uncharged, dissolved molecules and little effect on suspended mat-
ter. In this study electrodialysis has been used for desalting brackish
source water for the plants sited in New Mexico.
Because the energy required depends on the quantity of salt removed,
while the capital depends on the throughput, the cost of electrodialysis is
dependent on both the flow rate and the change in concentration. The cost
therefore has not been presented in graphical form but has been calculated
306
-------
10
O
O
O
REF. 32, SEAWATER
1
REF. I, BRACKISH WATER
REF. 32, BRACKISH WATER
1
4 6 8 10
THROUGHPUT (tO6GAL/DAY)
12
14
(From Reference 1, pages 90-94, two-year membrane life with membranes
costing 15% of the total capital, 8,000 hr/yr.)
Figure 15-4. Costs of reverse osmosis.
-------
for each individual use. The total operating cost is the sum of three fig-
ures: (1) the capital cost read from Figure 15-5 and charged at 17%/yr for
amortization, maintenance and other capital-related items; (2) membrane
replacement, prefiltration and chemical additives at $0.20/thousand gallons
of throughput; and (3) electricity at 2C/kw-hr for
0.4 kw-hr/(thousand gallons)(100 mg/1 removed).
15.9 ION EXCHANGE
Ion exchange is used in this study to remove ionic species and silica
from boiler feed water by exchanging these species with hydrogen, sodium and
hydroxyl ions on solid resins. The resins are then regenerated by H SO , NaCl
and NaOH enabling their continued use and reuse over long periods of time.
Design of an ionic exchange system is straightforward and is explained else-
where. ' ' The boiler feed water treatment areas on the plant designs of
Section 19 have been designed and costed individually, there being no simple
rules such as were used with other water treatment technologies. A very
brief description, limited to resins found useful for this study, follows.
A "weak-acid" resin is used to exchange positive ions in the water with
H . The resin removes only that part of the total cations equivalent in
amount to the bicarbonate alkalinity. This resin is particularly good for
removal of Ca and Mg ions and less satisfactory for removing Na ions. When
H+ has been substituted for Ca and Mg , carbonic acid is formed which
decomposes to free carbon dioxide that is then released in a degasifier. A
great advantage of weak acid resins is the ease of regeneration with sulfuric
acid; 110 percent of stoichiometry is all that is required.
+
A "strong-acid" resin is used to replace all of the cations with H .
Regeneration of strong acid resins with H SO is very inefficient, requiring
^ 44
about 200 percent of stoichiometry with counter-current regeneration. But
in all the designs the regenerant acid is poured first through the strong-
acid resin and then through the weak-acid resin, thus ensuring a high use of
the acid.
A "softening" resin is used to replace Ca and Mg with Na . It is
regenerated with NaCl and requires about 200 percent of stoichiometry to
308
-------
]
2
3
One Stage, approximately 50% demineralization
Two Stages, approximately 75% deruinerallzation
Three Stages, approximately 87.5% demineralization
Four Stages, approximately 93.8% demineralization
1
•
3
1.2
1.0
0-8
0.6
0.4
0.2
0.0
0.5 1
ext
Capacity (10 gal/day)
6 8 10 20 40 60 80 100
Figure 15-5.
Approximate electrodialysis capital investment as a
function of capacity for various numbers of stages.
(Each stage removes approximately 50% of salts in
its feed water.)
309
-------
regenerate in a counter-current manner.
A "weak-base" resin is used to replace SO. with OH . It cannot remove
4
weakly dissociated carbonic acid from the alkalinity or silicic acid from the
silica content in the water. It is easily regenerated with NaOH requiring
about 110 percent of stoichiometry. This resin follows a strong acid resin
to produce a demineralized water. A "strong-base" resin is used to replace
both the weakly dissociated and strongly dissociated acids. In these designs
a strong base resin has been used to replace SiO,_ with OH . It is regenera-
ted with NaOH. The regeneration is inefficient, but the caustic soda is
passed from the "strong-base" to the "weak-base" resin to obtain a high use
of the chemical.
A "mixed-bed" is a mixture of strong acid and base resins designed to
polish demineralized water without large swings in pH. The treatment is
necessary for high pressure boilers.
Regenerating chemicals are taken to cost
NaOH 6
NaCl 2
Capital costs for the ion exchange systems discussed in Section 19 have been
supplied to us by Permutit Co., Inc., Paramus, New Jersey. These systems are
skid-mounted and fully instrumented.
310
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REFERENCES SECTION 15
1. Water Purification Associates, "Innovative Technologies for Water
Pollution Abatement," Report No. NCWQ 75/13, National Commission on
Water Quality, Washington, D.C., December 1975. NTIS Catalog No.
PB-247-390.
2. Swindler-Dressier Company, "Process Design Manual for Carbon
Adsorption," EPA Technology Transfer Manual, October 1971.
3. Hutchins, R. A., "Economic Factors in Granular Carbon Thermal
Regeneration," Chem. Eng. Progress 69_ (No. 11) 48-55, November 1973.
4. Hutchins, R. A., "Cost of Thermal Regeneration," presented at 78th
National Meeting, AIChE, Salt Lake City, Utah, distributed by ICI
United States Inc., Wilmington, Delaware.
5. Loop, G. C., "Refinery Effluent Water Treatment Plant Using Acti-
vated Carbon," EPA-660/2-75—020, June 1975, U.S. E.P.A., Corvallis,
Oregon. NTIS Catalog No. PB-244-389.
6. Nandi, S. P., and Walker, P. L., "Adsorption Characteristics of
Coals and Chars," Office of Coal Research R&D Report 61, Interim 1.
7. Gitchell, W. B., Meidl, J. A., and Burant, W., Jr., "Carbon Regener-
ation by Wet Air Oxidation," Chem. Eng. Progress 7j^ (No. 5) 90-91,
May 1975. This is a capsule report; the full text is available
from Zimpro, Inc., Rothschild, Wisconsin.
8. Skrylov, V., and Stenzel, R. A., "Reuse of Waste Waters—Possibili-
ties and Problems," presented at the Workshop on Industrial Process
Design for Pollution Control, AIChE, New Orleans, October 1974.
9. Strauss, D., "Water Treatment," Power S.2 to S.24, June 1973.
10. Applebaum, B., Demineralization by Ion Exchange, pp. 23-67, Academic
Press 1968.
11. Betz Handbook of Industrial Water Conditioning, Betz Laboratories,
In., Trevose, Pennsylvania 19047.
12. Holden, W. S., Water Treatment and Examination, Williams and Wilkins
Co., Baltimore, Maryland 1970.
311
-------
13. Matson, J. V., and Perry, M. I., "Lime Softening of Cooling Tower
Slowdown," presented at 79th National Meeting AIChE, Houston, Texas,
March 17, 1975.
14. Kleusner, J., Heist, J., and Van Note, R. H., "A Demonstration
of Wastewater Treatment for Reuse in Cooling Towers at Fifteen
Cycles of Concentration," presented at AIChE Water Reuse Conference,
May 1975.
15. Brown and Caldwell, "Lime Use in Wastewater Treatment," page 37,
U.S. E.P.A. 600/2-75-038, October 1975; NITS Catalog No. PB-248-181.
16. Webb, L. C., "Side Stream Treatment of Cooling Tower Systems, A Step
Toward Environmental Improvements," presented at 37th Annual Meeting
American Power Conference, 1975.
17. Crits, G. J., and Glover, G., "Cooling Slowdown in Cooling Towers,"
Water and Wastes Engineering 45-52, April 1975.
18. McCoy, J. W., The Chemical Treatment of Cooling Water, Chemical Pub-
lishing Co. 1974.
19. Smith, A. L., and Muia, R. A., "Identify and Control Microbiological
Organisms in Cooling Water Systems," Power, July 1973.
20. Yost, W. H., "Microbiological Control in Recirculating Water Systems
Avoids Fouling," Oil and Gas Journal, 107-109, April 16, 1973.
21. Shair, S., "Industrial Microbicides for Open Recirculating Cooling
Water Systems," presented at International Water Conference, Engi-
neers' Society of Western Pennsylvania, Pittsburgh, Penn., Oct. 1970.
22. Sussman, S., "Fundamentals of Cooling Tower Water Technology," Paper
140A, Cooling Tower Institute Meeting, February 1975.
23. Clegg, L. F. L., "Disinfection in the Dairy Industry," in
Disinfection, Berarde, M. A., editor, Marcel Dekker, New York 1970.
24. Schultz, R. A. , "Evolution of Non-Polluting Microbicides," Paper
No. 131A, Cooling Tower Institute Meeting, January 1974.
25. Koeppen, A. von, Emerle, G. A., Nishio, K., and Metz, B. A.,
"Applying Pollution Control Technology to Cooling Water Treatment,
Wright Chemical Corp., Chicago, Illinois.
26. Koeppen, A. von, Pasowicz, A. F., and Metz, B. A., "New, Nitrogen-
Containing Organic Non-Chromate Corrosion Inhibitors," Ind. Water
Engr. 9 (3) 25-29 (1972).
27. Brunn, A. F., "Environmental Considerations in Cooling Tower Treat-
ment," Paper No. TP-136A, Cooling Tower Institute Meeting, Feb. 1975.
312
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28. Petrey, E. Q., "New Advances in Organic Cooling Water Programs,"
Paper TP-100A, Cooling Tower Institute, 1972.
29. Hwa, C. M., "New, Non-Chromate Synthetic-Organic Corrosion
Inhibitor for Cooling Water Systems," Paper TP-58A, Cooling Tower
Institute Meeting, June 1968.
30. Gesick, J. A., "A Comparative Study of Non-Chromate Cooling Water
Corrosion Inhibitors," presented at 35th Annual Meeting Interna-
tional Water Conference, Pittsburgh, Penn., October 1974.
31. Harpel, W. L., and Donahue, J. M., "Effective Phosphate/Phosphonate
Treatments Replace Chromate-Based Programs," Paper TP-117A, Cooling
Tower Institute Annual Meeting, January 1973.
32. Reed, S. A., "The Impact of Increased Fuel Costs and Inflation on
the Cost of Desalting Sea Water and Brackish Waters," Oak Ridge
National Laboratory Report ORNL-TM-5070 (Revised) January 1976.
See also, Reed, S. A., and Savage, W. F., same title, in National
Water Supply Improvement Association Journal 3^ (No. 2) 11-15, July
1976.
33. Duvel, W. A., and Helfgott, T., "Removal of Wastewater Organics by
Reverse Osmosis," J.W.P.C.F. 47 (No. 1) 57-65, January 1975.
34. Lindemann, K., "Design of Advanced Desalting Processes (Reverse
Osmosis)," in Water Management by the Electric Power Industry, ed.
Gloyna, Woodson & Drew, Resources Symposium No. 8, Center for
Research in Water Resources, University of Texas at Austin, 1975.
35. Strauss, S. D., "Water Treatment," Power S.2 to S.24, June 1974.
36. VanStone, G. R., "Treatment of Coke Plant Waste Effluent," presented
at the Winter Meeting of the Eastern States Blast Furnace and Coke
Oven Association, February 11, 1972, and printed in Ir_qn_and Steel
Engineer, distributed by Calgon Corp., Pittsburgh, Penn.
313
-------
SECTION 16
SEPARATION OF AMMONIA,
CARBON DIOXIDE AND HYDROGEN SULFIDE
16.1 INTRODUCTION AND RESULTS
Process condensate is laden with ammonia. In the gas plants the ammo-
nia is mostly neutralized by carbon dioxide. In the solvent refined coal
plant the ammonia is partly neutralized by hydrogen sulfide. Depending on
the coal, hydrogen chloride may have been present in the gas phase and this
will pass into the condensate and neutralize ammonia. Any organic acids in
the vapor will also mostly pass into the condensate and neutralize ammonia.
Phenol is itself a weak acid. Ammonia must be removed from condensate.
Ammonia is volatile, but simply venting it to the atmosphere is not accep-
table. Fifty percent of the population will respond to the odor of 21.4 ppm
by volume NH in air. The state of New Mexico permits 25 ppm by volume in
an effluent gas stream. This strength vapor is blown from a solution con-
taining about 13 mg/1 ammonia. Of course the vapor if dispersed, but it is
quite clear that simply blowing ammonia into the atmosphere is not permis-
sible from strong solutions. Ammonia must be collected, and since ammonia
is a salable commodity, it should be collected in a salable form. Also, H S
£
cannot be vented to the atmosphere but must be collected and sent to the sul-
fur recovery plant. Ammonia interferes with the Claus sulfur recovery pro-
cess.
The problem is not the removal of ammonia and gaseous acids from water/
which is the usual problem of stripping or "sour water stripping." For these
wastewaters ammonia and gaseous acids must be separated into three individual
streams.
314
-------
The presence of phenol and volatile organic acids may interfere with
ammonia separation because the procedures considered in this section all rely
on relative volatility to effect part of the separation. In the following
discussion the possible problems caused by phenol and volatile organic acids
are not considered. If phenol is extracted, this is done before the separa-
tion of ammonia. Extraction will remove most of the volatile organic carbon
from the water. That part of the ammonia neutralized by hydrochloric acid
will remain in the water unless alkali is used. If the water goes on to bio-
treatment, 450 mg/1 N is required for nutrient. This represents the quantity
of ammonia fixed by 1140 mg/1 Cl, so fixed ammonia may not be a problem unless
chloride exceeds this value.
It is not difficult to separate volatile gases (NH and volatile acid
gases) by stripping the water in a tray column with steam (live or from a
reboiler) at a rate usually in the range 8-12 Ib steam/100 Ib water. Strip-
2-4
ping is much described in the literature. If this is done then ammonia
must be separated from the CO + H S. One way to do this, not considered in
£ t*
this study, is to absorb the ammonia in sulfuric acid and produce crystalline
ammonium sulfate for sale. Another procedure is to absorb the ammonia in
phosphoric acid or acid ammonium phosphate. It is then possible to strip out
the ammonia with steam and to recycle the acid ammonium phosphate. This is
the Phosam-W process of United States Steel Corporation Engineers and Consul-
tants described in Section 16.7. For this study, this is the preferred pro-
cess. The Phosam-W process is shown in Figure 16-1.
It is possible to separate wastewater into three streams by an all-
distillation procedure as pictured in Figure 16-2. By refluxing clean water
to the "deacidification" column, C02 and H2S are removed overhead while ammo-
nia remains in solution. In the second ("ammonia") column, ammonia is
stripped out of water and concentrated. A procedure similar to the all-
distillation procedure is shown in the drawings of most Lurgi plant water
treatment schemes as part of the Phenosolvan processes. The procedure inves-
tigated in this study is not the Phenosolvan process.
In order to compare the all-distillation process with the Phosam-W pro-
cess, a study has been made of the all-distillation procedure. Section 16.2
describes the basic equations that were used for tray-to-tray calculations
315
-------
U)
/FRACTION-
( ATOR FEED
V TANK
Figure 16-1. Phosam-W process for ammonia separation.
-------
10
DEACIDIFIC
ATION
TOWER
Figure 16-2. All-distillation process for ammonia separation.
-------
down the tower to determine the necessary number of theoretical trays which
control the height of the tower. The vapor liquid equilibrium data for a
mixture of NH., CO- and H-0 is presented in Section 16.3 and applies to gas-
producing plants. Similar data for a mixture of NH , H S and HO, presented
«3 ft f,
in Section 16.4, are held to apply to an SRC plant. Calculated results
applicable to gas plants are presented in Section 16.5. For the Phosam-W
process the design was provided by United States Steel Engineers and Consul-
tants. We have estimated the capital cost of both systems.
On Figure 16-3 are shown the capital costs for the all-distillation pro-
cess to separate water, NH and CO , together with two costs for the Phosam-W
process. The Phosam-W process is cheaper than the all-distillation process,
as assumed in this study, and uses less energy. Furthermore, the all-
distillation process may not work for mixtures of water, NH, and H_S without
the presence of CO.. For ammonia separation in all plants, we have therefore
assumed the use of the Phosam-W process with capital and operating costs as
shown in Figure 16-4. We have limited points on the capital cost versus
throughput curves for the Phosam-W process and have drawn the curves through
these points proportional to the curve for the all-distillation process.
16.2 SEPARATION BY DISTILLATION; CALCULATION OF NUMBER OF THEORETICAL TRAYS
The distillation equations for the deacidification column in the all-
distillation procedure are briefly given here in the form in which they were
used.
The nomenclature is shown on Figure 16-5. All flows are in moles/hr.
The mixture is assumed to be HO, NH and CO . The washwater (reflux) and
feed are saturated liquids. The usual simplifying assumption of constant
molal overflow is made. The calculation is made in the manner given in stan-
dard texts on distillation. The desired condition at the top of the tower—
D, DN, DC—is stated. The reflux or wash water—0—is stated. The feed—F,
FN, FC—is given. The tower bottoms—W, WN, WC—are calculated from the
overall tower material balance. The reboil rate S is stated. The number of
theoretical trays is then calculated. If this number was greater than about
20, calculation was stopped and a new choice of O and S was made.
318
-------
12
10
8
O
~ 6
OT
O
U
T
T
PHOSAM-W TOTAL COST
A SNG PLANT
V OIL FROM COAL PLANT
TOTAL,ALL
DISTILLATION
PROCESS
\
AMMONIA TOWER
(ALL DISTILLATION
PROCESS)
\
DEACIDIFICATION TOWER
{ALL DISTILLATION PROCESS)
I I I i \ I I I L
234
THROUGHPUT (I06GAL/DAY)
Figure 16-3. Comparative capital costs for ammonia separation.
319
-------
r—!>| , ,r
SRC
— COST
2 « — ^AS
o
Q
z
3
-------
WASH
H2O, O MOLES/HR.
LIQUID
H2O, O MOLES/HR.
NH3, ON MOLES/HR
CO2, OC MOLES/HR.
H.O
NH
FEED
F MOLES/HR.
i3, FN MOLES/HR.
L, FC MOLES/HR.
\
"•
I k
r~^
i
i°"
i V
n
Vn + l
k,
•
— j^. i
C
TRAY )
TRAYn
TRAYn + 1
VAPOR
H2O, V MOLES/HR.
NH3, VN MOLES/HR
CO2 , VC MOLES/HR
TOP
H2O, D MOLES/HR.
NH3, DN MOLES/HR.
CO_, DC MOLES/HR.
BOTTOMS
•^ V —~" -~~—~~_
I. \ — - - -
STEAM, S MOLES/HR.
CO2, WC MOLES/HR.
Figure 16-5. Distillation tower nomenclature.
321
-------
The calculation starts at the top of the tower. A material balance
around the top of the tower, below tray n and above tray n+1, when n is above
the feed, gives
V = D = S (la)
n
VN , = ON + DN (lb)
n+1 n
VC ._ = OC + DC (Ic)
n+1 n
Since only pure water is refluxed
ON = OC =0
o o
and
VN = DN
= DC
The molar vapor rates leaving the top tray are now known. The molar
vapor rates are next converted to partial pressures. From the partial pres-
sures the liquid concentrations can be found from the vapor-liquid equilib-
rium relations given in Section 16.3. From the liquid concentrations the
liquid molar rates ON. and OC, can be calculated. Knowing ON and OC , the
vapor rates rising from the second tray VN and VC_ can be found from Equa-
tions 1. Calculation then proceeds from tray to tray down the tower.
The total pressure P_ on the top tray is taken to be 20 psia and the
pressure is assumed to increase by 0.25 psia per tray going down the column-
The total pressure on tray n is
PT = 20 + 0.25(n-l) (2)
Within the narrow range of interest the vapor pressure of water, P .-./is
H2°
322
-------
given by
T°K = 352 + 1.5PH Q (psia)
for 18 < P < 22
H2°
T°K - 0.555(t°F + 459.7) (4)
But,
PH O = PTV/(V+VN+VC) (5)
and so, from the vapor composition the temperature on any tray can be deter-
mined. Also,
PNH = pTtVN)/(V+VN+VC) (6)
and
PCO = VVC>/ (7)
Now, knowing the temperature and the partial pressures of NH and CO ,
the total concentration of carbon dioxide, C moles/1, and the total concen-
tration of ammonia, A moles/1, can be calculated as described in Section 16.3.
Finally, the liquid molar rates are related to the liquid compositions
by the equations
ON = (0.018A )0 (8)
n n
OC = (0.018C )O (9)
Below the feed tray Equations (1) do not apply and Equations (10)
should be used as follows.
323
-------
V = D = S (XOa)
n
VN , = ON - FN + DN = ON - WN (lOb)
n+1 n n
VC , = OC - FC + DC = OC - WC (10c)
n+1 n n
Feed is added to that tray which improves the separation of NH3 and CO_.
When the calculation approaches the feed tray, double calculations must be
made, one using equations (1) and one using equations (10). For each tray
the ratio OC /ON (moles CO_/moles NH_ in liquid) is calculated, and when
n n 2 3
this ratio is lower using equations below than it is using equations above
the feed, the feed should have been put in on the tray one above. Below the
feed a double set of calculations is not necessary. This method of locating
the feed tray seemed to give satisfactory results. It was not checked by
calculating upwards from the bottom. This procedure throws the cumulative
error into the bottom composition.
16.3 VAPOR-LIQUID EQUILIBRIUM FOR NH -CO -H 0
The results of Van Krevelen, Hoftijzer and Huntjens are used to find
the concentrations of total ammonia and total carbon dioxide (A moles NH /I
and C moles CO.,/1) given the partial pressures P and P (psia) .
4* JNtt — v*fL/ —
When carbon dioxide and ammonia are simultaneously dissolved in water
+ — — =
the solution contains six species: NH , NH , NH COO , HCO , CO and C02
"Total concentration" of ammonia means:
A = (NH3) + (NH4+) + (NH2COO~)
"Total concentration " of carbon dioxide means:
(C02) + (HC03 ) + (C03 ) + (NH2COO
(approx.) (HC03) + (C03) + (NH2COO)
324
-------
because (CO ) is very small.
<£
The equations which give the concentrations of the individual species
are:
1) The modified Henry's law equation for NH ,
P (psia, - 0.0193^ 10°-0« \
(16)
3) Because of the equilibrium
2 (NH3)(HC03")
4) Because of the equilibrium
325
NH + HC03 ^ NH2COO +
-------
NH + HCO ~ - NH* + CO/
33 43
(NH +)(CO ) , v
K = —— — = exp I—^ - 18.161 (18)
(NH3)(HC03~) V '
5) To maintain ionic balance
'(NH. ) = (HCO ~) + 2(CO ~) + (NH.COO~) (19)
4 332
The procedure is as follows:
1) Calculate (NH ) from Equation (13). Since the maximum value of
(NH,) used in this work is about 0.58, the maximum value of 10 3 is
1.03. A very good approximation for (NH ) is
H P
0 NH3 -x
= I 0.0193
Since most of carbon dioxide is in the form of bicarbonate, K.^ can be
326
-------
approximated as
7540 12.4(HCO )
K, = exp ill®. - 28.7 + 5_
1 1 + 7.9(HC03
First calculate K with (HCO ) = 0. Then calculate (HCO ) from (21). Then
JL «3 J
recalculate K and recalculate (HCO, ) and so on until two successive approxi-
mations for (HCO "} agree within 0.2 percent.
3) From (17) calculate (NH COO~).
* +
4) Equations (18) and (19) can be combined to eliminate (NH. ) and
=
give a quadratic equation for (CO ).
- 2 « /
2(C0_ ) + (CO. ) (HCO, ) + (NH COO )
O J 1 J £•
- K3(NH3) (HC03") - 0 (22)
Solve (20) for (C03~).
5) Solve (18) or (19) for (NH ).
6) Solve (12) for C.
7) Solve (16) for i^.
8) Return to (21) and solve for (HCO3~).
9) Repeat steps 3 to 8 for successively better approximations of
(HC03~).
10) When (HCO ~) is know within 0.2 percent, determine A and C from
*3
(11) and (12).
16.4 VAPOR-LIQUID EQUILIBRIUM FOR NH3-H2S-H2O
4
Equilibrium data in graphical form have been published by Beychok.
Similar graphs were prepared for the mixture NH -CO -H O. Examination of
•5 f* £
the two sets of graphs shows that CO is sufficiently volatile compared to
JL
NH for the deacidification column to work, that is, for CO to go overhead
while holding NH, in the bottoms. However, H2S seems not to be sufficiently
volatile compared to NH, to be separated in this way and it is doubtful if
327
-------
the all-distillation procedure, as shown in Figure 16-2, will work for the
mixture NH -H S-H O. Equilibrium calculations were not made for the four
component mixture NH -CO -H S-H 0, but if CO is present in large excess over
H S, then H2S would be expected to go overhead with CO.. This is because C0_
is a stronger acid than H S. The use of deacidification towers in Lurgi pro-
cess plants substantiates this supposition.
In fact the Phosam-W process and not the all-distillation process has
been used. For costing purposes it was necessary to be sure that a simple
stripping column would work with the same steam rate for NH -H S-H 0 as for
NH -CO -H 0. It was found that two extra theoretical trays were used when
H S was the acid gas. The analytical procedure used for the vapor-liquid
equilibrium calculations for NH -H S-H O is given below, derived from Refer-
ences 4 and 5.
The only species present are HS , NH , NH and very small concentra-
tions of H_S, so the calculation is very much simpler than for ammonia and
carbon dioxide. The relationships are:
P (pBia) = 0.0193 TI-^T- (23)
£. A
(psia) - 0.0193 ££ 100.025(ft-S) (24)
NH_ H
3 o
where
A = total ammonia in solution, gm-mole/1
S = total hydrogen sulfide in solution, gm-mole/1
As before
HQ = exp- 16.54) (14)
K is given by the equations
328
-------
I = 10a-0.089S (25)
K4
where
0.627
(t'C) (26,
5.95
(Note that Equation (8) of Reference 4 seems to be misprinted in that the
number 0.089 is missing. The above equations do give the vapor pressures
given on the graphs in Reference 4.)
To calculate the liquid composition from the vapor pressures, the pro-
cedure is:
1) The first approximation to A-S is
H0PNH, -(0.025H PMU )/0.0193
fA_qj = i in ° NH3
IA b'l 0.0193 iu J
The second approximation is
H P
IA_SI - _f 3 -0.025 (A-S)
(A S)2 0.0193 10 *
and so on.
2) With A-S established, calculate S from
For the first approximation use
, = 10~a
4
For the second approximation use
and so on.
329
-------
16.5 DEACIDIFICATION COLUMNS FOR NH^CO^H O
The basis for all calculations is F = 100 moles/hr water in feed. The
following feed solutions were studied.
Feed concentrations (mg/1) Feed rates (moles/hr)
NH3
8,000
4,000
4,000
co2
15,500
7,800
3,900
NH (FN)
0.848
0.424
0.424
C02 (FC)
0.636
0.318
0.159
FC/FN
0.75
0.75
0.375
The deacidification tower is a two-variable problem, since the refiuxed
wash water W and the steam rate S can both be varied. If the steam rate was
not high enough it was not possible to strip out CO below the feed, and if
the washwater rate was not high enough it was not possible to prevent strip-
ping ammonia out the top. There is a narrow range of wash and steam rates
within which the tower will function satisfactorily; steam rates should be in
the range 8 to 10 moles/100 moles feed and wash rates in the range 120 to
150 moles/100 moles feed. This high wash rate means that a lot of steam is
needed in the ammonia tower reboiler.
Some results are given on Table 16-1. Steam rates of 6, not shown on
Table 16-1, do not work. With a steam rate of 8, water rates below 100 do
not work.
The feed, in most cases, should be about one half a theoretical stage
above the bottom. The system is quite sensitive. If the bottom takeoff is
too low, the ammonia concentration is too low; if the bottom takeoff is not
low enough, the carbon dioxide concentration is too high. With a feed in
which the molar ratio (CO0)/(NH ) is 0.75, it is not possible to get less
£ *J
than 10 to 15 percent of the feed carbon dioxide out the bottom. With feed
in which the molar ratio (CO )/(NH3> is 0.375, we cannot get less than 20 to
25 percent of the feed carbon dioxide out the bottom. There was no problem
getting only 1 percent of the feed ammonia out the top, and altering this to
2, 4 or 6 percent made little difference, it may be possible to inject some
330
-------
TABLE 16-1. DEACIDIFICATION TOWER FOR NH3-CC>2
Basis 100 moles water in feed.
Bottom Tray
FN FC DN DC S W
(moles/hr) (moles/hr) (* FN) (% FC) (moles/hr) (moles/hr)
0.424 0.318 1 99 10 150
140
120
8 150
140
120
100
95 8 150
140
130
120
90 8 150
140
130
120
2 99 8 150
140
120
4 99 8 150
140
120
6 99 8 150
140
120
0.424 0.318 2 95 8 150
140
120
4 95 8 150
140
120
5 95 8 150
140
130
120
6 95 8 150
140
120
Feed
Tray
12
15
*
7
8
11
»
7
8
9
11
7
8
9
11
6
7
10
5
6
8
5
5
7
6
7
10
5
6
8
5
5
6
7
5
5
7
No.
12
13
15
16
*
8
9
11
12
*
8
9
9
10
11
12
7
8
8
9
9
10
11
12
7
7
8
10
11
5
6
7
8
8
9
8
9
5
6
7
8
6
7
7
8
10
11
5
6
7
8
9
6
5
6
6
7
7
8
7
8
5
0
7
8
WN
(* FN)
134
65
115
30
102
101
120
44
91
90
123
52
115
34
132
77
133
75
112
40
109
22
75
139
93
130
67
125
69
106
24
124
58
lib
56
128
76
127
68
126
67
135
83
126
58
122
63
98
120
51
107
111
37
119
52
110
33
127
72
125
70
124
62
we
(% FC)
30
11
27
6
19
18
30
9
16
15
31
10
28
7
34
13
32
12
29
7
26
5
15
36
17
32
13
34
14
19
6
31
11.5
13
8
34
15
32
13
33
13
34
15
30
10
32
12
17
30
10
18
30
a
30
10
30
7
16
a
32
13
31
11
•Not workable
(continued)
331
-------
TABLt 16-1 (continued)
Bottom Tray
FN FC DN DC S W
(moles/hr) (moles/hr) (% FN) (% FC) (moles/hr) (moles/hr)
0.848 0.636 1 95 10 150
120
99 8 150
140
130
120
95 8 150
140
130
120
90 8 150
0.424 0.159 2 99 8 150
140
120
4 99 8 150
140
120
6 99 8 150
140
120
2 95 8 150
140
120
4 95 8 150
140
120
6 95 8 150
140
120
Feed
Tray
10
*
7
7
a
10
7
7
8
10
7
7
9
14
6
7
11
5
6
9
7
9
14
6
7
11
5
6
9
No.
10
11
*
7
8
8
9
10
11
11
7
8
8
9
10
11
8
9
7
B
11
15
7
8
12
5
6
7
10
7
8
11
15
7
8
12
5
6
7
9
10
WN
(% FN)
112
26
124
59
121
50
132
72
94
119
17
116
39
126
58
147
24
127
67
97
80
91
96
80
122
66
89
73
124
61
76
71
85
88
73
120
61
84
128
68
we
U FC)
31
7
40
15
38
13
40
16
12
37
12
36
9
38.
13
27
3
50
24
26
23
15
30
24
49
24
29
23
48
21
20
20
27
27
22
47
22
27
45
21
*Not workable
332
-------
wash into the lower part of the column as a control.
The tower height increases as the wash rate is decreased from 150 to 120
and as the fraction of ammonia out of the top is decreased. The extreme
range of variation is from 7 to 12 theoretical stages. Doubling the feed
concentration has little effect.
For cost estimating, S = 8 Ib steam/100 Ib feed, a wash (reflux) rate of
130 Ib water/100 Ib feed; 12 theoretical trays (24 actual trays ) above the
feed and 1 theoretical tray (2 actual trays ) below the feed were assumed.
16.6 AMMONIA CONCENTRATION
Aqueous ammonia leaves the bottom of the deacidification tower at less
than half its strength in the feed water and must be concentrated. The ammo-
nia tower calculations involve only two components and were made using the
methods given in standard chemical engineering texts. Vapor-liquid equili-
brium data were taken from Reference 6 at 65 psia total pressure. For most
of the tower the concentration is so low that Henry's law applies, and it
was used in the form
PMH (psia} m 0.0193 A/H
Nn. O
where, as before,
At 65 psia and 300°F (422°K) this becomes
x = 0.074y
where
x = mole fraction NH3 in liquid
y = mole fraction NH3 in vapor
333
-------
The top concentration was set at 30 wt percent ammonia and 99 percent of the
feed ammonia was sent out the top. The tower height was not sensitive to feed
concentration in the range 4,000 to 2,000 mg/1. The results are shown on Fig-
ure 16-6. For cost estimating, 13 Ib steam/lb feed and 11 theoretical stages
(22 actual stages ) with 2 above the feed and 9 below the feed were used. It
must be remembered that for every 100 Ib water treated, about 222 Ib are fed
to the ammonia tower, so the steam rate is 29 Ib steam/100 li> dirty water
treated, which is high.
16.7 PHOSAM-W PROCESS
In this section the Phosam-W process as shown in Figure 16-1 is briefly
described.
The process was developed to remove the free ammonia from waste water
streams containing both ammonia and acid gases such as H S, CO- and HCN. Tra-
^ ^
ditiohally this was accomplished by stripping from the vapor with sulfuric
acid to produce ammonium sulfate crystals. By substituting phosphoric acid
for sulfuric acid, a cyclic process is possible for producing anhydrous ammo-
nia. The basic concept behind the process is that phosphoric acid has three
hydrogen atoms that can be replaced by ammonium ions in water solution.
Therefore there should exist a range of mole ratios, ammonia to phosphoric
acid, over which the ammonia is bound tightly enough for good absorption but
still loosely enough so that it can be stripped back out.
As can be seen in the flow diagram, the process consists of three main
pieces of equipment: the superstill, which is a stripper and absorber, a
Phosam stripper and a fractionator. The liquid feed stream containing acid
gases and ammonia is divided into two streams which are preheated and then
recombined, at the bubble point, to enter the top of the stripping section of
the superstill. Both the ammonia and acid gases are stripped from the water
at an operating pressure of 4 psig.
The stripped water leaves the bottom of the superstill containing less
than 200 ppm free NH3 and is used for heating part of the feed stream in the
feed preheater. Boilup for the superstill stripper is provided primarily by
indirect steam at 60 psig. Direct flashing of the fractionator bottoms in
334
-------
16
14
OT
UJ
o
o
o
LJ
X
I-
8
1
I
10 12 14 16
STEAM (moles/100moles feed)
Figure 16-6. Ammonia tower.
335
-------
the superstill stripper provides the remaining reboiler duty.
The vapor containing the stripped ammonia and acid gases passes upward
into the absorber section of the superstill. In the bottom stage of the
absorber, it is contacted with an intense spray of ammonium phosphate solu-
tion which has been cooled in an external air-cooled exchanger. The vapors
then pass into a tray section where they are further scrubbed by the lean
ammonium phosphate solution in counter-current contact. The vapors then
leave the top of the absorber at a pressure of 3 psig.
The net flow of rich solution from the absorber is purged of acid gases
by vapors generated in a solution exchanger in a contactor compartment at the
bottom of the absorber. The rich solution is then pumped from the absorber
to the stripper, after passing through the upper section of the stripper con-
denser where it is heated by exchange with the overhead vapors from the strip-
per.
The stripper is operated at elevated pressure. This is because the
stripping equilibrium is markedly improved at high pressures. A high operat-
ing pressure in the stripper reduces the steam requirements. The function of
the stripper is to remove NH from the rich solution to regenerate the lean
solution for recycle to the tray section of the absorber. Boilup in the
stripper is provided by steam at 600 psig in the stripper reboiler. The hot,
lean solution leaving the stripper bottom is cooled and enters the top of the
superstill absorber section.
The aqua ammonia vapor, 10 to 20 wt percent ammonia, from the top of the
stripper passes through a two-stage condenser. The top stage cools and con-
denses part of the vapor, thus recovering much of its heat content. In the
bottom section the remaining vapor is totally condensed by heat exchange with
cooling water. The condensate flows to the fractionator feed tank where it
is pumped into the ammonia fractionator column. In the fractionator the aqua
ammonia is distilled, at elevated pressure, into 99.99 percent pure ammonia
leaving the top and less than 0.05 percent ammonia leaving the bottom.
The overhead pure ammonia vapor is condensed in the fractionator con-
denser by heat exchange with cooling water. Reflux is returned to the top
of the column.
By the addition of a very small flow of sodium hydroxide solution at the
336
-------
proper location, traces of acid gases and organic compounds are prevented
from accumulating in the fractionator and eventually contaminating the pro-
duct.
The slightly alkaline, pressurized hot water from the bottom of the
tower is flashed directly into the bottom of the superstill providing a por-
tion of the steam requirement.
16.8 EQUIPMENT SIZE AND CAPITAL COST
For the equipment shown on Figure 16-2, the size and cost for the all-
distillation procedure have been estimated. The column diameters were esti-
7 8
mated using the Glitsch manual. The costs were taken from Guthrie, using
the multipliers given by him for erection and for stainless steel (or clad-
ding) and multiplying by 1.68 to update the costs to early 1976. The results
are given on Table 16-2 and plotted on Figure 16-3.
Equipment sizes for the Phosam-W process were provided by United States
Steel Engineers and Consultants, and the capital cost was estimated similarly
to the all-distillation process with the results shown on Table 16-3 and
Figure 16-3.
For the SRC plant where the separation is between NH , H S and water,
the size and cost for the Phosam-W plant is shown on Table 16-4 and Figure
16-3. The equipment size provided by United States Steel Engineers and Con-
sultants, as reproduced in Table 16-4, was for a coal liquefaction process,
not SRC but similar.
16.9 OPERATING COST AND ENERGY
The all-distillation procedure uses about 2.9 x 10 Btu/thousand gallons.
This high energy need is caused by the large recycle to the deacidification
tower necessitating large amounts of steam in the ammonia tower reboiler. The
Phosam-W process uses about 1.7 x 10 Btu/thousand gallons and is cheaper than
the all-distillation process. Phosam-W is the process of choice. Operating
costs are presented on Table 16-5 and Figure 16-4 for gas plants. For SRC
plants the capital is a little higher, and this is reflected in Figure 16-4.
337
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TABLE 16-2. COSTS FOR AMMONIA SEPARATION BY ALL-DISTILLATION PROCESS
(Material S.S.
Throughput, 10 gal/day
Deacidification column:
Number of columns
Top: ht, ft
dia, ft
Bottom: ht, ft
dia, ft
2
Reboiler, ft
£i
Cost, 10 $
Ammonia Tower:
Number of columns
Top: ht, ft
dia, ft
Bottom: ht, ft
dia, ft
2
Reboiler, ft
2
Air cooler, ft
Cost, 106$
Total Equipment Cost* 10 $
clad, 24" tray spacing)
0
1
55
6
25
7
1,555
0
1
15
6
40
8
5,354
7,902
1
1
.5 1.5
1
55
.0 10.0
25
.5 12.0
4,667 9
.78 1.4
1
15
.5 11.0
40
.5 14.0
16,069 32
23,705 47
.1 2.4
.9 3.8
3.0 4.5
1 2
55
13.8
25
16.6
,222
2.5 3.8
2 2
15
11.0
40
14.0
,141
,410
4.9 7.3
7.4 11.1
*Royalty and engineering not included
338
-------
TABLE 16-3. SIZE AND COST OF PHOSAM-W PROCESS IN A GAS PLANT
(Material S.S. clad)
Throughput, 10 gal/day 4.5
Equipment
Superstill
Top: ht, ft 62
dia, ft 11.5
Bottom: ht, ft 69
dia, ft 15.5
Exchangers
E-l, ft2 7,000
E-2, ft2 120
E-6, ft2 7,330
E-7, ft2 8,270
E-8, ft2 29,690
Cost, 106$ 5.1
Phosam Stripper
ht, ft 67
dia, ft 7.0
Exchangers
E-3: top, ft2 5,940
2
bottom, ft 1,360
E-9, ft2 3,140
Cost, 106$ 1.1
Fractionator
ht, ft 64.00
dia, ft 3.25
Exchangers
E-5, ft2 2,810
E-10, ft2 930
Cost, 106$ 0.69
Approx. royalty 10 $ 1.1
Total Equipment & Royalty Cost, 10 $ 8.0
339
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TABLE 16-4. SIZE AND COST OF PHOSAM-W PROCESS IN SRC PLANT
(Material S.S. clad)
Throughput, 10 gal/day 3.6
Equipment
Superstill
Top: ht, ft 84
dia, ft 14.5
Bottom: ht, ft 61
dia, ft 9.5
Exchangers
E-l, ft2 12,000
E-2, ft2 540
E-6, ft2 6,725
E-7, ft2 11,300
E-8, ft2 5,180
Cost, 10 $ 4.2
Phosam Stripper
ht, ft 70
dia, ft 9.5
Exchangers
E-3: top, ft 8,600
2
bottom, ft 1,760
E-9, ft2 5,440
Cost, 106§ 1.7
Fractionator
ht, ft 65
dia, ft 5.25
Exchangers
E-5, ft2 7,100
E-10, ft2 2,020
Cost, 106$ 1.3
Approx. royalty, 10 $ 1.0
Total Equipment & Royalty Cost, 10 $ 8.2
340
-------
TABLE 16-5. OPERATING COSTS FOR AMMONIA SEPARATION IN GAS PLANTS
Throughput, 10 gals/day
Capital changes including
maintenance at 17%/yr
(330 days/yr):
Energy at $1.80/10 Btu
Caustic soda
Phosphoric acid
0.5 1.5 3.0 4.5
Costs ($/thousand gallons)
1.41 0.95
3.06 3.06
0.04 0.04
0.01 0.01
0.91 0.92
3.06 3.06
0.04 0.04
0.01 0.01
4.52 4.06 4.02
4.03
*Sale of ammonia at $140/ton with 95% recover yields
$3.30/103 gallons of water treated when feed water
contains 6,000 mg/1 ammonia, or generally
$0.55/(103 gallons) (1,000 mg ammonia/1).
341
-------
REFERENCES SECTION 16
1. Arthur D. Little, Inc., "Research on Chemical Odors," prepared for
the Manufacturing Chemists Association, October 1968.
2. Melin, G. A. Niedzwieki, J. L., and Goldstein, A. M., "Optimum Design
of Sour Water Strippers," Chem. Eng. Progress 71 (No. 6), 78-82, June
1976.
3. 1972 Sour Water Stripping Survey Evaluation, American Petroleum Insti-
tute Publication 927, Washington, D.C., 1973.
4. Beychok, M. R., Aqueous Wastes from Petroleum and Petrochemical Plants^
pp. 158-198, John Wiley and Sons, 1967.
5. Van Krevelen, D. W., Hoftijzer, P. F., and Huntjens, F. J., "Composi-
tion and Vapor Pressures of Aqueous Solutions of Ammonia, Carbon
Dioxide and Hydrogen Sulphone," Recueil Trav. Chim. Pays-Bas 68 191-220,
1949.
6. Perry, R. H., Chilton, C. H., and Kirkpatrick, S. D., eds., "Chemical
Engineers Handbook," pp. 3-65, 3-66, 4th edition, McGraw-Hill, 1963.
7. Glitsch,Inc., "Ballast Tray Design Manual," Bulletin No. 4900-Third
Edition, P.O. Box 6227, Dallas, TX 75222, 1974.
8. Guthrie, K. M., "Capital Cost Estimating," Chemical Engineering
pp. 114-142, March 24, 1969.
9. Skamser, R., "Coal Gasification, Commerical Concepts, Gas Cost Guide-
lines," U.S.E.R.D.A. publication FE-1235-1, UC-90c, January 1976.
342
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SECTION 17
SOLVENT EXTRACTION OF PHENOL
17.1 INTRODUCTION AND SUMMARY OF RESULTS
Phenol can be extracted from dirty process condensate and a crude, mixed
product recovered for sale. Alternatively phenol can be destroyed by biolo-
gical treatment. In this section is given a simplified procedure for deter-
mining the cost of solvent extraction so as to determine when, and whether,
it should be used.
Solvent extraction is standard for all Lurgi process plants where the
proprietary Phenosolvan process is used. The Phenosolvan process has been
described: "The liquid is mixed with an organic solvent (isopropyl ether)
in an extractor in order to dissolve the phenol. The phenol solvent mixture
is collected and fed to solvent distillation columns where crude phenol is
recovered as the bottom product and the solvent as the overhead product. The
solvent is then recycled to extractors after removing some of the contained
water. The raffinate is stripped with fuel gas to remove traces of solvent
which are picked up in the extraction step. The fuel gas is scrubbed with
crude phenol product to recover the solvent. Finally, the phenol solvent
mixture is distilled in the solvent recovery stripper to produce the crude
phenol product, and the solvent is recycled to the extraction step. The
solvent free raffinate is heated and steam stripped to remove carbon dioxide,
hydrogen sulfide, and ammonia."
This scheme is illustrated on Figure 17-1, which is also partly taken
from Reference 2. For the discussion in this section we have not designed
the full distillation section but have idealized the situation to that of
Figure 17-2. This means that the cost estimates are low.
343
-------
DIRTY
WATER
RICH
SOLVENT
DIRTY .
WATER
SOLVENT
REMOVAL
LEAN
SOLVENT
GAS +
SOLVENT
«—^2r*—*
PHENOL
RECOVERY
PHENOL + SOL VENT
TREATED
WATER
-FUEL GAS
FUEL GAS
PHENOL + SOLVENT
PHENOL
Figure 17-1. Phenosolvan process,
RICH
SOLVENT
.TREATED
WATER
PHENOL
Figure 17-2. Idealized solvent extraction,
344
-------
To determine cost, an optimization procedure is necessary. The more sol-
vent that is circulated, the smaller the extraction equipment becomes and the
larger the still becomes. There is an optimum solvent circulation rate. In
Sections 17.2 to 17.4 some minimum costs have been found and these are given
on Table 17-1 for the case of benzene as the solvent.
If 95 percent of the phenol is extracted, the minimum possible cost is
6 6
2.9 x 10 $/yr to treat 10 Ib water/hr. This is about $3/thousand gallons
and is not cheap. Furthermore, extraction is not necessary so will only be
done if it pays for itself. Two things offset the cost of extraction:
reduced cost of downstream biological treatment and income from the phenol
sold.
Refined phenol sells for about 26C/lb so the crude mixtxlre extracted from
waste water is probaly not worth more than 6£/lb. Values as low as 1.7* to
2.2C/lb have been suggested. ' As a fuel phenol has a higher heating value
of 11,380 Btu/lb and might be worth 2.3«/lb at $2/10 Btu.
We have very little information on the nature of the non-phenolic bio-
chemical oxygen demand and cannot estimate with accuracy the reduced cost of
biological treatment resulting when phenol is extracted. Referring briefly
to Section 18, the cost of biotreatment is very much dependent on the BOD
removal rate. If phenol is extracted and the BOD is reduced to about half,
the cost of biotreatment may be reduced by about one third; that is, about
$I/thousand gallons treated may be saved on biotreatment. In Lurgi process
there is some possibility that the water, after extraction and ammonia sepa-
ration, can be used as cooling water makeup without any biotreatment at all.
In this study this has not been assumed possible.
The lowest cost shown on Table 17-1 is $2.9 x 10 /yr to recover
45.6 x io6 Ib phenol/yr with 95 percent recovery from a feed of 1000 Ib/hr
(2.88 x io6 gal/day) having 6000 mg/1 phenol. If 0.9 x 10 $/yr are saved
by reduced need for downstream biological treatment the extraction will break
even even if phenol can be sold for 4.4$/lb. However, the cost estimates may
be low.
Some conclusions can be drawn. Extracting more than 95 percent of phenol
in the feed water is unlikely to be worthwhile. A feed concentration of more
than 6000 ppm phenol is probably required before extraction pays. A better
345
-------
TABLE 17-1. PHENOL RECOVERY
Basis: 10 Ib water/hr; benzene solvent (K = 2.3)
(The breakeven value of phenol is the minimum cost divided by
the quantity of phenol recovered.)
Feed phenol % Phenol
(ppm) Recovered
Phenol Recovered
(106 Ib/yr)
Min. cost
(106 $/yr)
Breakeven value
of phenol
C/lb
6000
6000
6000
2000
2000
2000
600
600
600
99.8
99
95
99.8
99
95
99.8
99
95
47
47
45
16
15
15
4
4
4
.9
.5
.6
.0
.8
.2
.8
.8
.6
4
3
2
4
3
2
4
3
2
.18
.62
.90
.18
.62
.90
.18
.62
.90
8.
7.
6.
26
23
19
87
75
63
7
6
4
346
-------
solvent than benzene would increase the usefulness of extraction. Solvent
extraction requires quite a lot of energy. For example, using
0.7 Ib solvent/lb water where the solvent has a latent heat of 170 Btu/lb,
then the energy required is about 1 * 10 Btu/thousand gallons of water
treated.
Solvent extraction removes most of the volatile organic contaminants in
foul condensate. Extraction is a prerequisite to purifying the water by dis-
tillation. Extraction will assist the separation of ammonia.
In the treatment plants designed for this study, solvent extraction has
not been considered except where distillation has been considered. This is
most probably correct for Hygas, probably correct for Synthane, marginal for
the Lurgi process used in the power plants, and should be studied further and
possibly reconsidered for the Solvent Refined Coal plant.
Section 17.5 describes an adsorption process for phenol, as yet untested,
which has a strong potential for use in the future. At the present time the
cost seems to be comparable to solvent extraction, but the adsorption process
is new and still undergoing improvements. In any situation where solvent
extraction is considered, adsorption systems should also be investigated.
17.2 SIMPLE EXTRACTION EQUATION
Consider a simplified, ideal, counter-current extraction using the nomen-
clature of Figure 17-3. Complete immiscibility of solvent and water is
assumed, and at the exit of each theoretical stage it is assumed that
y_
— = k (a constant) (1)
n
Material balances for the solute (phenol) give
Stage 1
x = £k _
2 w 1 1 o
347
-------
OJ
CD
w = water feed rate, 10 Ib/hr
s = solvent feed rate, 10 Ib/hr
y - phenol concentration in solvent, ppm
x = phenol concentration in water, ppm
Figure 17-3. Ideal counter-current liquid-liquid extraction.
-------
Stage 2
X W
sk
W
/skxl
skx.
- Xl + SYo
/ skx \ skx
„ sk - 1 _ + - i +
w \ w 'o/ w 1
-*
= /SIC)' + Sk /Sk+ \
\w/l wl 1 ^oyw /
Now taking y , the concentration of phenol in the feed solvent, to be negli-
gible, one obtains
Stage n
W
so
w x,
w
and, by subtraction,
n+1
w
349
-------
or
log I I — 11 — -
= n + 1 (3)
log (-
Equation (3) relates the number of stages n to the desired extraction
ratio x /x , the extractive ability of the solvent k and the flow rate
ratio of solvent to,water s/w. Figure 17-4 is a graphical representation of
Equation (3). For a chosen situation in which x /x , w and k are fixed, the
number of stages n can be decreased by increasing the solvent rate s. With
fewer stages the extraction column can be smaller and costs less. However,
all of the solvent has to.be evaporated to separate it from the phenol, and
increasing the solvent rate increases the boiler and energy cost. To find
the cost of solvent extraction the design must be optimized. To do the opti-
mization one first finds the cost for a particular example, called the base
case, and then scales from this case to do the optimization. A related tech-
4
nique is described by Scheibel.
17.3 BASE CASE EXAMPLE
The base case example has been arbitrarily chosen as:
(1) Benzene to be the solvent because information is published; k = 2.3
and the latent heat of vaporization = 170 Btu/lb.
(2) x /x - 100 (99 percent extracted).
6 6
(3) w = 1 (i.e., 10 Ib/hr) and s = 1.3 (i.e., 1.3 x 10 Ib/hr), so
sk/w = 2.99.
The number of stages is then n = 3.8.
Otto H. York Co., Parsippany, N.J. has kindly supplied a basis for esti-
mating an erected cost of Scheibel extraction columns. The system will use
10 columns, each 60-70 ft high and 10 ft in diameter in type 304 SS. Each
column has a 30 hp agitator, total (for 8,000 hrs/yr) 1.79 x 10 kw-hrs/yr.
The water flow is about 20 ft/hr and the solvent flow is about 30 ft/hr. The
350
-------
10
o
<
I-
(O
o
UJ
(C
ui 6
i
H
U.
O
UI
CO
5
rs
z
2 -
A*
xl
20
2345
"w~
Figure 17-4. Extraction equation.
351
-------
columns cost $3.75 x 10 . A charge of 17%/yr of capital has been taken for
amortization, maintenance and other capital related items. The annual charge
for the extraction columns is $0.64 x 10 /yr. The agitator energy at
2
-------
17.4 OPTIMIZATION
The extractor cost will depend, but not linearly, on the height of the
columns which in turn depends on the number of theoretical stages n and on
some combination of the number of columns and their diameter which in turn
depends on the throughput measured as (s+w) . In the narrow range of interest,
the cost for case "i" can be related to the cost for the base case "b" by the
equation
°'5
B
cost
cost
. + w. \' n.
- i -i.
b + wb; "b
For optimization calculations take w. = w = 1 (i.e., 10 lb/hx of water)
so Equation (4) becomes
6 /s + 1 \°'5 n
cost of extractors = 0.68 x 10 I 2 3 I 3~a
- 0.116 x lo6 (s + l)°*5n $/yr (5)
The distillation costs are proportional to the solvent rate s and are
cost of distillation = 4.71 x io6 -^
J- • j
= 3.62 x io6 s $/yr (6)
Equations (3) , (5) and (6) are used for optimization.
A sample optimization is shown on Figure 17-5. Figure 17-5 includes the
base case, which is clearly a long way from the optimum. The minimum is
clear and readily determinable. Additional results are given on Table 17-2
as a function of the percent of the feed phenol extracted. The calculations
were also made for better solvents than benzene (for which k - 2.3). When
calculating costs for these imaginary solvents the latent heat of vaporiza-
tion was kept constant at 170 Btu/lb. In using Table 17-2 it must be
353
-------
O
O
X
O
-------
TABLE 17-2. APPROXIMATE MINIMUM ANNUAL COST OF
SOLVENT EXTRACTION FOR AN IDEAL CASE
Basis
% Feed
phenol
extracted
95
99
99.8
95
99
95
99
: 106 Ib feed
Extraction
ratio, x /x
t J-
20
100
500
20
100
20
100
water/hr
Solvent
k
2.3
2.3
2.3
4
4
6
6
(= w)
Solvent
rate, s
(106 Ib/hr)
0.5
0.65
0.67
0.35
0.38
0.25
0.33
sk
w
1.15
1.5
1.58
1.38
1.5
1.5
2.0
No. Of
stage
(n)
8.9
8.7
11.2
5.7
8.7
5.9
5.7
Cost
106 $/yr
2.90
3.62
4.18
2.02
2.54
1.67
1.95
remembered that the costs apply to an ideal case, that is, to a solvent that
is completely immiscible with water and that can be distilled from phenol
without reflux. The costs are low.
17.5 SELECTIVE ADSORPTION OF PHENOL
Rohm and Haas Company manufactures synthetic polymeric adsorbents which
can accept high loadings of phenol and may be sufficiently selective for
phenol that recovered phenol can be sold (the selectivity requires testing).
Once phenol is loaded onto the solid, it can be removed by a solvent which
need not be immiscible with water. The added choice of solvent gives the
possibility of very high concentrations of phenol in the solvent going to
distillation with drastic reduction in the energy required for distillation
which is the largest cost.
Without extensive tests a detailed cost estimate is not possible. Rohm
and Haas has kindly supplied a "ball park" estimate for complete "grass roctr
type of installation to treat 10 Ib/hr water having 6000 ppm phenol with 90
percent phenol removal. The systems use either acetone regeneration as
355
-------
pictured in Figure 17-6 or gasoline regeneration as shown in Figure 17-7. To
complete the scheme of Figure 17-7, a very simple still used for benzene
liquid-liquid extraction has been added. The gasoline rate is
fi o
0.0895 * 10 Ib/hr. There are 14 adsorption columns each with 950 ft of
resin. Six columns are in use at any moment and six are being regenerated.
Two columns are spare. The costs are presented on Table 17-3. The costs
have been adjusted by using the same unit costs as in the rest of the report.
Labor cost has been omitted as in the rest of the report. Table 17-3 also
presents the costs for the optimized liquid-liquid extraction system for 99
percent recovery with benzene solvent (k = 2.3) (details on Table 17-2;
benzene rate 0.65 x 10 Ib/hr).
The cost apparently increases as the completeness of the design and
sophistication of the estimate increases. For the moment it must be assumed
that the systems are comparable. However, the adsorbent systems are very
new; the first commercial installation in the United States was started up
in 1975, and improvements are to be expected. In any situation where sol-
vent extraction is considered, adsorption systems should also be investigated.
356
-------
PHENOLIC
WASTE
POLYMERIC
ADSORBER
#1
(Loading)
TREATED
WASTE
SOLVENT
MAKE UP
WATER
V V
RECOVERED
SOLVENT
POLYMERIC
ADSORBER
#2
(Regenerating)
2
DISTILLATION
COLUMN #1
SOLVENT/WATER
f~lX Y WATER TO
I
-------
GASOLINE
Figure 17-7. Phenol recovery by adsorption with gasoline regeneration.
358
-------
TABLE 17-3 COSTS OF PHENOL RECOVERY BY VARIOUS METHODS
Basis: 10 Ib water/hr with 6000 ppm phenol, 99% removal
Note: See text - different systems are not directly comparable
U)
I/I
vD
Adsorption with
acetone regeneration
Adsorption with
gasoline regeneration
Capital Cost
Adsorption and distillation systems.
For adsorption resin, engineering,
site preparation, development costs
and license fee included
Operating Cost
Amortization and other capital
related expenses at 17%/yr
Steam at $1.80/10 Btu
Electricity at 2
-------
REFERENCES SECTION 17
1. Shaw, H. and Magee, E. M., "Evaluation of Pollution Control in Fossil
Fuel Conversion Processes, Gasification; Section 1 - Lurgi Process,"
EPA Report 650/2-74-009-c, EPA, Research Triangle Park, N.C., 1974.
2. El Paso Natural Gas Company, "Second Supplement to Application for a
Certificate of Public Convenience and Necessity," Docket CP 73-131,
October 1973.
3. Tennessee Valley Authority, "Evaluation of Fixed-Bed, Low-Btu Coal
Gasification Systems for Retrofitting Power Plants," Electric Power
Research Institute, February 1975; NTIS catalogue PB-241-672.
4. Scheibel, E. G., "How to Design Optimum Extraction Columns for Minimum
Over-All Cost," distributed by Fluid Separation Design, Inc.,
Fairfield, N.J.
5. Crook, E. H. and Stevens, B. W., "Removal and Recovery of Phenols from
Industrial Waste Effluents with Amberlite XAD Polymeric Adsorbents,"
presented at New York Water Pollution Control Association Meeting,
January 1975.
6. Skamser, R., "Coal Gasification; Commercial Concepts; Gas Cost Guide-
lines," U.S. ERDA, NTIS Catalog FE-1235-1, UC-90C, January 1976.
360
-------
SECTION 18
BIOLOGICAL TREATMENT
18.1 INTRODUCTION
In biological treatment of wastewater, dissolved and colloidal organic
matter is subjected to action by bacteria and other microorganisms which
remove organic matter from water by converting it, partly to carbon dioxide
and partly to a settleable organic sludge consisting mostly of dead and live
microorganisms. Organic matter is commonly measured as biochemical oxygen
demand (BOD), chemical oxygen demand (COD) and total organic carbon (TOG).
The efficiency of biological treatment is measured by the decrease in these
parameters. Biological treatment can also be used to convert ammonia to
nitrate (called nitrification) and to convert nitrate to nitrogen (called
denitrification).
In this section various biological treatment systems are examined in
terms of their capabilities for treating foul condensate from coal conversion
processes. Also discussed will be field experience, design procedures and
their limitations, and research needed if existing information is inadequate.
Since there is no existing full-scale facility in the United States, most of
the evaluation has to be based on the experience of treating wastewaters with
similar characteristics and on laboratory studies of coal conversion waste-
waters. Under certain circumstances educated guesses were necessary; the
assumptions made are clearly stated and can be tested and modified as more
information becomes available.
A long history exists in the treatment of industrial wastes containing
phenolic compounds, and these wastes are amenable to aerobic biological
1 o 07 TO
treatment. Acceptable treatment may be achieved using trickling
361
-------
filters or activated sludge systems. Most of these studies are related to
2
coke plant wastes, while a limited number of laboratory studies on coal con-
version wastes are available. However, it has also been noticed that appro-
priate safeguards must be provided against process upsets, and inorganic
macro and micro nutrients may have to be added to ensure the success of aero-
3 4
bic biochemical oxidation. ' In this case these include not only the common
nutrient, phosphorous, but also various inorganic trace nutrients such as
metals which are required by activated sludge but are not available in the
4
wastewater.
In the area of anaerobic biological treatment, little is known regarding
its capability for treating phenolic wastes. A similar situation exists with
the biological removal of nitrogenous oxygen demand and nutrients, which
require a combination of nitrification and denitrification following an accep-
table removal of carbonaceous oxygen demand.
In order to identify the optimal biological treatment system for coal
conversion wastewaters, the following systems have been studied and a preli-
minary design provided for costing purposes where possible:
Air Activated Sludge (AAS) System
High Purity Oxygen Activated Sludge (HPOAS) System
Activated Trickling Filter—High Purity Oxygen Activated Sludge
(ATF-HPOAS) System
Air Activated Sludge—Nitrification-Denitrification (AASND) System
Anaerobic Treatment System
Each system has been studied for the exemplary wastewater shown on Table 18-1.
Comparisons of the results in Section 18.3 through 18.8 show that the system
consisting of an activated trickling filter followed by a high purity oxygen
activated sludge process has the lowest cost, has good stability and control-
lability and reduces BOD to the lowest level. Furthermore, the design of the
high purity oxygen activated sludge system was more conservative than the
design of the air activated sludge system, which emphasizes the correctness
of our choice. The activated trickling filter followed by a high purity oxy-
gen activated sludge process is used for all cost estimating on the complete
treatment plants discussed in Section 19. As shown on Table 18-16, cost
estimates were made for this system using several BOD loadings and
362
-------
TABLE 18-1. CHARACTERISTICS OF EXEMPLARY COAL CONVERSION WASTEWATER
Flow 3 x io6 gallons/day = 1.04 x io6 Ib/hr
Phenol 6000 - 6600 mg/1 as C.H.OH
D 5
BOD ~ 18,000 mg/1
COD 26,000 - 30,000 mg/1
NH_ ~ 3,500 mg/1 as N
pH 5.5* to 7.0
Ca + Mg < 100 mg/1
Na ~ 100 mg/1
Cl ~ 500 mg/1
*Approximate pH for 6,300 mg/1 pure C H OH solution.
throughputs. For the Hygas plants, which have the lowest BOD loading and low
throughput, the cost has been estimated as $0.025/lb BOD removed. For all the
other plants the cost has been estimated as $0.021/lb BOD removed. Each sys-
tem requires (0.18 kw-hr + 1900 Btu as steam)/lb BOD removed.
18.2 EQUALIZATION
A feed of nearly constant characteristics can eliminate the upsets due
to shock loads or abrupt changes in composition and is essential for treating
high-strength wastewaters by biological means. Recent batch studies ' on
coke and petroleum refining wastewater treatment by biological processes have
reaffirmed this requirement.
In many respects the exemplary coal conversion wastewater described in
the last subsection is very similar to the coke plant wastewaters, character-
ized by high concentrations of phenol, ammonia, BOD and COD. As a first
approximation the field experience on the treatment of coke plant wastewaters
has been used for coal conversion wastewaters. For coke plant wastewaters?
it has been recommended7 that a storage capacity equal to five or more days
363
-------
of feed liquor be provided for equalization. In an existing activated sludge
plant treating coke plant wastewater, the storage capacity of the equaliza-
Q
tion basin is on the order of 7.4 days. A period of 7.4 days of storage has
been used to calculate costs of the storage basins, which were assumed to be
approximately square, 10 feet deep with concrete walls and base one foot
thick. Cost is then estimated by taking the concrete at $100 per cubic yard.
18.3 AIR ACTIVATED SLUDGE (AAS) SYSTEM—SCALED PLANT
The air activated sludge (AAS) system is probably the most common treat-
ment system used for wastewaters with constituents similar to coal conversion
wastewaters, e.g., coke plant wastes. An extensive literature review on the
9
biological oxidation of coke plant wastes was reported by Barker and Thompson
in 1973. Among the treatment systems discussed, AAS is the predominant
2
treatment system of success. Laboratory studies abroad have also indicated
that AAS systems can satisfactorily treat the coal conversion wastes with the
following characteristics:
total ammonia ~1,500 ppm
total phenols ~300 ppm
thiocyanate ~150 ppm
chloride ~2,500 ppm
Field Experience
Since much experience on the treatment of coke plant wastes will be uti-
lized in the following discussion, it is appropriate to summarize on Table
9
18-2 the typical compositions of coke plant wastes, compositions after ammo-
nia distillation, and after both ammonia distillation and phenol extraction.
In comparison with the exemplary coal conversion wastewater listed on Table
18-1 and the analytical data in Section 14-2, the coke plant waste (excess
ammoniacal liquor) has a comparable ammonia content, larger concentrations
of cyanide, thiocyanate and chloride, and less of phenol.
One of the most important considerations regarding biological treatment
of coke plant wastes is to determine if the waste contains any inhibitory
364
-------
TABLE 18-2. TYPICAL ANALYSIS OF AMMONIACAL LIQUOR AND STILL WASTE
Excess Ammoniacal Undephenolized Dephenolized
Liquor Still Waste Still Waste
Concentration (ppm)
Ammonia
Phenol
Cyanide
Thiocyanide
Sulfide
Chloride
3800
1500
20
600
2
7000
Concentration (ppm) Concentration (ppm)
155
1320
0
0
0
4350
110
158
0
0
0
5400
constituents which may render the biological treatment system totally or par-
tially unfunctional. If these constituents exist it is essential to deter-
mine their threshold concentrations and thus the dilution required for the
influent to the biological treatment system. Field experience with coke
plant wastes at municipal sewage plants has shown that phenols may be suc-
cessfully reduced to less than 1 mg/1 if they are diluted by municipal sewage
at a ratio of between 6 and 19 to 1, i.e., coke plant wastes constituting
about 15 percent to 0.5 percent of the influent to biological treatment sys-
g
tern. Dilution may also be achieved by internal recirculation of treated
water rather than the use of an external dilutant.
The constituents of inhibitory potential to the AAS system include
phenol, ammonia, chloride and refractory organics in the coal conversion
wastewaters. The threshold concentration of phenol has been reported at the
level of 500 mg/1; however, as phenol is readily utilized in the aeration
tank of activated sludge systems, its concentration in a completely mixed
aeration tank is not likely to exceed this level. The inhibitory effect of
ammonia on the AAS system has been well recognized in several full-scale
designs where ammonia concentration in the aeration unit is maintained at
1,2009 to 2,000 mg/1 or less. Although nitrogen is an essential nutrient
for the growth of activated sludge, its consumption in the AAS system is far
less than that of phenol and thus may dictate the dilution requirement. As
365
-------
for choride, no detrimental effect to the AAS system is expected if the chlo-
9
ride concentration does not exceed 2,000 mg/1.
The refractory organics, as commonly measured by the difference between
COD and BOD, may also affect the performance of the AAS system and the dilu-
tion necessary for efficient treatment. Unfortunately this appears to be a
most important, but poorly understood, parameter. Any definite information
in this regard can only be obtained by pilot testing with the specific waste-
water to be treated.
Design Criteria
Because of the inhibition of biological processes by ammonia and the
need to remove ammonia, two alternative schemes have been considered:
(1) Ammonia recovery (probably by distillation) followed by AAS system.
(2) Dilution of ammonia followed by AAS system and further removal of
ammonia in subsequent processes.
Ammonia distillation is discussed in Section 16, and the three-sludge system
for ammonia removal is detailed in Section 18.5.
The design criteria for biological treatment of coke plant wastes are
potentially useful for coal conversion wastewater treatment. Some typical
design criteria are shown on Table 18-3. Based on the information of Table
18-3, the ability of the AAS system to remove phenols to a level of less than
1 mg/1 appears to be consistent, while key design parameters like F/M ratio
vary over a wide range. The most complete data are available under column
(1), representing the design at Bethlehem Coke Plant, Bethlehem, Pennsylvania
where a full-scale biological treatment facility has been in operation since
1962.
Scale-up From Coke Plant Wastewater Treatment
One approach to designing a treatment facility without pilot tests and
adequate data base is to scale from an existing facility on the basis of key
design parameters. The Bethlehem Coke Plant has successfully removed phenols
and BOD for years, and rather complete design and operation data are
366
-------
TABLE 18-3. TYPICAL DESIGN CRITERIA FOR BIOLOGICAL
TREATMENT OF COKE PLANT WASTES
(1) (2) (3) (4)
Influent
Dilution Yes * Yes *
NH Concentration 2,000 * <1,200 *
in aeration unit
mg/1
PH 6-87-8 * *
Weight: ratio 70 * * *
of phynol to P •
Temperature of Q 80 - 95 >70 * *
aerated water, F
F/M, Ibs phenol 0.7 0.2 to 0.25 * 0.16 to 0.3
per Ib MLSS per day
MLSS, mg/1 3,300-4,700 2,500-3,500 * 2,500-3,000
Aeration time, hrs 56a 24 37 114
Influent phenol ^1,400 250-475 260-400 3,000
concentration, mg/1
Effluent phenol <0.1 0.1-0.3 0.8-3.6 0.1
concentration, mg/1
BOD removal, % 85-95 * * *
*
No data available
a. Based on undiluted wastewater flow
(1) Data from Reference 8
(2) Data from Reference 9, for Lone Star Steel Co., Texas
(3) Data from Reference 9, for Dominion Foundries & Steel of Hamilton, Ontario
(4) Data from Reference 3, for Alan Wood Steel Co., Pennsylvania
367
-------
g
available in the literature. As an introduction to activated sludge systems,
such a scaled design has been made and costed. Theoretical design procedures
are discussed in the next section.
As pointed out, two treatment schemes are possible for ammonia removal.
In Scheme 1, ammonia is removed by distillation leaving just sufficient ammo-
nia nitrogen in the influent to the biological system to meet the nutrient
requirement. Scheme 2 will provide adequate dilution so that the ammonia
concentration in the biological reactor will not exceed the inhibitory level
of 2,000 mg/1, and certain further treatment of ammonia (e.g., nitrification
and denitrification) will be necessary following the removal of carbonaceous
BOD. The major difference between the two schemes lies in the hydraulic load
to the clarifiers. Scheme 2 requires larger clarifiers because of the added
dilution flow.
The scaled design is based on the assumption that the biodegradability
of coal conversion wastewaters is identical with that of the coke wastewater.
This assumption may well be open to question. No data on COD of the coke
Q
wastewater is given by Kostenbader and Flecksteiner. However, an analysis
of an average coke plant) waste indicated that the theoretical oxygen demand
due to phenols, which are readily biodegradable, constitute about 68 percent
of the measured COD while for coal conversion wastewater phenol averaged
about 40 percent of the COD. Although the question of biodegradability
can only be answered fully by pilot testing, the above comparison indicates
certain differences in chemical composition between coke plant and coal con-
version wastewaters. It is essentially unknown at this point whether and how
this will affect the design of biological treatment. Should the assumption
of biodegradability become invalid to any extent, there would be correspond-
ing limitation on the usefulness of the preliminary design.
The following rules were used to produce the scaled design.
Nutrients such as N and P are essential for biological treatment. The
required weight ratio is assumed to be invariant and is phenol:N:P = 70:5:1.
Assuming an average phenol concentration of 6,300 mg/1 (6,000-6,600 mg/1),
the concentration of N and P required will be 450 mg/1 and 90 mg/1 respec-
tively. Excess N is available in the wastewater, and for Scheme 1 the ammo-
nia nitrogen concentration will be reduced to 450 mg/1 by distillations prior
368
-------
to biological treatment. Phosphorus will have to be supplied by the addi-
tion of phosphoric acid or equivalent.
The phenol loading rate to the aeration tank is taken to be constant and
has been determined to be 12 Ib phenol/(day)(10 ft ) or less, and
8
0.7 11) phenol/(day)(Ib MLSS) or less. For the actual scale-up the following
2 3
are used: 11.4 Ib phenol/(day)(10 ft ) and 0.45 Ib phenol/(day)(Ib MLSS).
The aerator power requirement is taken to be proportional to the BOD or
phenol removed. At Bethlehem Coke Plant the power requirement is based on
18.2 Ib phenol removed/(day)(hp), or 43.3 Ib BOD/(day)(hp)—calculated on the
basis of 2.38 Ib BOD per Ib phenol—which compares closely with typical
12
values in the literature of 45-50 Ib BOD removed/(day)(hp).
Sludge generation at Bethlehem Coke Plant is reported to be
0.2 Ib sludge produced/lb phenol treated at a phenol loading rate of
1,300 Ib/day and a waste sludge rate of 3,300 gal/day. The solids concentra-
tion in the clarifier underflow can thus be calculated to be about 9,500 ppm.
When these numbers are used to scale, the required plant has a return sludge
flow of about 1.5 x 10 gal/day and a waste sludge flow of 0.42 x 10 gal/day.
The mean cell residence time in such an AAS system will be about 10-11 days.
The quantity of sludge to be disposed of will be about 16.5 tons dry solids
per day.
The surface area of clarifiers is determined on the basis of a hydraulic
loading of about 685 gpd per square foot.
For the treatment of sludge, the following design criteria are used:
20 Ibs dry solids per square foot per day for the dissolved air flotation
(DAF) thickener, and 120 Ibs dry solids per square foot per day for vacuum
28
filters. These values are assumed, not scaled, because Bethlehem Coke
Plant discharges its sludge to a sewage plant and provides no sludge treat-
ment .
The preliminary design of the AAS system for Schemes 1 and 2 is shown on
Figures 18-1 and 18-2. Figure 18-3 shows the general configuration of aera-
tion basins, used for costing purpose. The major difference in the design
between Schemes 1 and 2 lies in the size of clarifiers because of different
hydraulic loads.
The power requirement for aerators is estimated to be 9100 hp and is a
369
-------
9,000hp
AIR
NUTRIENTS
COOLED EFFLUENT
FROM AMMONIA
STILL
SLUDGE
DISPOSAL
82.6
@ 20% solid
3xlO*gali./day
450mg/l NH3-N
EQUALIZATION
22.2xl06gali.
*
1
AERATION
k *^ 10.8x10 gols.
^/ TIARIFIC
I 5,100
|.5xlO%U. /day >
:AT
f,2
r
EFFLUENT
FROM BIOLOGICAL
TREATMENT
RETURN SLUDGE
0.42x10 gali./day
Figure 18-1. Air activated sludge system (Scheme 1)
-------
U>
DILUTION
WATER
1.7x10 go'*-
3x10 gall./day
3,600 ppmN
EQUALIZATION
22.2xl06 ga'
4.3xlO°gols./day
2,000 ppm N
I FOR
_' emergency
RETURN SLUDGE
0.42xl06go's-
0.94 solids
SLUDGE
DISPOSAL
Figure 18-2. Air activated sludge system (Scheme 2).
-------
53'-
©
+
•^—
AERATOR
©
227'
CONCRETE
THICKNESS-I ft.
f 10 parallel basins
||' required
Figure 18-3.
Typical configuration of an aeration basin for air activated
sludge systems.
372
-------
highly significant outlay of capital and operating costs. Since this power
requirement is directly related to the BOD removed and has essentially noth-
ing to do with the ammonia concentration, no difference in power requirement
is expected between Schemes 1 and 2.
According to the Bethlehem Coke Plant experience, about 0.1 mg/1 phenol
or less is found in the effluent when the influent contains about
Q
1,400 mg/1 phenol. Now that the influent contains about 6,600 mg/1 phenol,
phenol in the effluent may be about 0.5 mg/1. The reported BOD removal at
Bethlehem Coke is about 85-95 percent, so 10 percent of the influent BOD of
about 18,000 mg/1 may remain in the effluent, i.e., 1,800 mg/1. No COD data
is available at Bethlehem Coke Plant, but the COD removed will not be less
than the BOD removed. So the COD may be expected to be reduced from about
28,000 mg/1 to 11,800 mg/1. In summary, the effluent characteristics of the
AAS system may be as follows:
Phenol ~0.5 mg/1
BOD ~1,800 mg/1
COD ~11,800 mg/1
Cost of Scaled Plant
The costs of the AAS systems, calculated according to Table 18-4, are
shown on Tables 18-5 and 18-6. The difference in costs between Schemes 1 and
2 appears to be negligible, and the total cost of treatment is about
$3.2/1,000 gallons.
18.4 THEORETICAL DESIGN OF AIR ACTIVATED SLUDGE SYSTEMS
The performance of an activated sludge process depends on biochemical
transformation and subsequent solid-liquid separation. In principle, a
rational design would have to be based upon the mechanism of these processes.
The mechanism in the activated sludge process is generally represented by
kinetic models which invariably involve the application of certain fundamen-
tal coefficients. Two auch models for biochemical transformation, the Monod
model and the first order approximation, have recently been utilized for the
373
-------
TABLE 18-4. COST ESTIMATION PROCEDURE FOR AIR ACTIVATED SLUDGE PLANT
Capital Costs
Aeration Basin: See Figure 18-3
Cost based on volume of concrete at $240/yd
Aerators $600/installed hp
Clarifiers Figure 15-3
Cooling Towers $10/gpm circulated
Dissolved Air Flotation Thickeners Figure 18-4
Rotary Vacuum Filters Figure 18-4
Operating Costs
Amortization and other capital related items 15% of capital/yr
Maintenance: Concrete 1% of capital/yr
Machinery 2% of capital/yr
Electricity 2CAw-hr, 8,000 hrs/yr
Phosphorus 42$/lb P
design of activated sludge process. The first order approximation has been
found valid if the influent BOD does not exceed 500 mg/1 and the effluent
14
soluble BOD is less than about 90 mg/1. ' It is therefore considered not
applicable in the treatment of these exemplary wastewaters which have an
influent BOD far in excess of 500 mg/1. In the following paragraphs the
Monod model will be used to estimate the size of aeration basins and the
range of size variation by assuming reasonable values for fundamental coeff-
cients.
Monod Model Calculations
In a completely mixed activated sludge system, the mean cell residence
time 0 , hrs, may be related to the specific growth rate of the micro-
c -1 -1
organism y, hrs , and the microorganism decay coefficient b, hrs , by
f - u.-b (i)
374
-------
I.W
.9
.6
.7
.6
.6
.4
.3
2.5
~ •«
•
2 1.5
1-"
o J0
0 .09
-J .08
£ .07
rf .06
« .05
.04
.03
.025
.oz
.010
Al
«^
.^r.
>
^X^
x^
TP IND
,-X
^*
>
>X
^
EX 250
—
"2_
^^-<
^_
^
•^
/
\
x
'A
-^
X
:i
^
X
JUM FILTRATION
^
^^
,*
^
^S'
s
X
X
ix'
X
x*
X
X
X
X
^ DISSOLVED AIR
FLO
FAT
ON
^
V
1
100
1000
AREA (sq.ft.)
10000
Figure 18-4. Capital cost of vacuum filtration and dissolved air flotation
(from Reference 12) .
375
-------
TABLE 18-5. COSTS OF AIR ACTIVATED SLUDGE SYSTEM (SCHEME 1}
Capital Costs 10 $
Equalization 1.19
Aeration Basin 2.38
Aerator 5.40
Clarification 0.23
DAF Thickening 0.54
Vacuum Filtration 0.36
TOTAL 10,10
Operating Costs 10 $/yr
Amortization & other
capital-related items
@ 15% of capital/yr 1.52
Maintenance
of Concrete work 0.04
of Machinery 0.12
Electricity @ 7500 kw 1.20
Chemicals:
Phosphorous 0.32
TOTAL 3.20
TOTAL ANNUAL COST =3.2 x 106 $/yr
=3.2 $/1,000 gallons
376
-------
TABLE 18-6. COSTS OF AIR ACTIVATED SLUDGE SYSTEM (SCHEME 2)
Capital Costs 10 $
Equalization 1.19
Aeration Basin 2.38
Aerator 5.40
Clarification 0.29
DAF Thickening 0.54
Vacuum Filtration 0.36
TOTAL 10.16
Operating Costs
Amortization & other capital-related
items @ 15% of capital/yr 1.52
Maintenance
Concrete work 0.04
Machinery 0.12^
Electricity @ 7500 kw 1.20
Chemicals
Phosphorous 0.32
TOTAL OPERATING COSTS 3.20
TOTAL ANNUAL COST = 3.2 x 10 $/yr
=3.2 $/l,000 gallons.
ju
Including requirements for aeration and pumping of dilution water
377
-------
The Monod model then relates \i to the soluble substrate concentration S
which may be expressed in terms of BOD or equivalent
v s
y = -J*L- (2)
H K +S v '
s
where
y = maximum specific growth rate of the microorganisms (hrs )
K = saturation constant, i.e., the substrate concentration when
S
y is one-half of y
m
According to equations (1) and (2) , 0 may be determined by assuming
c
proper values for S, K , y and b. Furthermore, the volume of aeration
s m
basins can be determined from 0 and the following parameters:
o
S = influent soluble substrate concentration (mg/1)
Y = true growth yield (mg MLSS)/(mg BOD removed)
X = microorganism concentration in aeration basin (mg/1)
Q = influent flow rate to aeration basin (gal/hr)
using the equation
Y (S -S)
xe ' u/ej + b »>
where 0 is the hydraulic residence time and equals
(volume of aeration basin)/Q.
However, in the determination of 0 some practical constraints have to
be observed. Experience shows that 0 should be kept in the range of 3 to 14
days to achieve an activated sludge that flocculates well.
Assumptions
In our evaluation using the Monod model, attention has been primarily
378
-------
focused upon the BOD due to phenol, the most significant organic contaminant
in the influent to the activated sludge process. Based on a phenol concen-
tration of 6,600 mg/1, the S in terms of BOD is assumed to be 2.38 x 6,600 =
15,700 mg BOD/1. The microorganism concentration in aeration basins, X, is ...
assumed to be equal to the MLSS (mixed liquor suspended solids) because the
influent concentration of suspended solids is considered insignificant and X
is assumed to be equal to 4,000 mg/1. The influent flow rate is 3 x 10 gpd,
or 2,083 gpm.
The other parameters necessary for our evaluation are S, b, u , K and
ill 5
Y . The value of S depends on the BOD removal percentage and the following
9
values have been used:
case
% phenol removed
S, mg/1 BOD
1.2
2.4
D
99.0
160
For the fundamental coefficients, b, y , K and Y , four sets of values
m s g
have been selected and are shown on Table 18-7.
TABLE 18-7. VALUES OF FUNDAMENTAL COEFFICIENTS
USED IN THE
Coefficients
Set 1*
Set 2
Set 3
Set 4
b
(hr)'1
0.0025
0.0025
0.0025
0.0104**
ym
(hr)"1
0.55
1.65
0.55
1.65
MONOD MODEL EVALUATION
K
s
mg/1 BOD
120
120
120
120
Y
g
mg MLSS produced
per mg BOD removed
0.5
0.5
0.084***
0.084
13
13
*Typical values for domestic sewage.
**Typical value for refinery wastewater.
***Typical value for coke plant wastewater.
8
379
-------
The values of the coefficients in Set 1 are those typical of domestic
sewage and are used as the basis for modifications in the other sets. The
oxidation of phenol appears to be more efficient than the BOD removal from
g
domestic sewage. At Bethlehem Coke Plant more than 99 percent of phenol was
removed at a phenol loading of 0.7 Ibs phenol per day per Ib MLSS, or an equi-
valent BOD loading of about 1.7 Ibs BOD per day per Ib MLSS, which is much
higher than those used for domestic sewage, namely, 0.2 to 0.6 Ibs BOD per
day per Ib MLVSS. For both applications 0 is in the range 5 to 15 days,
C
and the MLSS in the range 3,000 to 6,000 mg/1. This indicates that in com-
parison with domestic sewage the coefficients for phenol-bearing wastes may
include a larger y or smaller Y , or both. This reasoning is reflected in
our modification of the coefficients in Sets 2 through 4. In Set 4 we also
tried a value of b which is typical of refinery wastewater.
Using the values of the fundamental coefficients and S, the volume of
aeration basins was determined in accordance with the Monod model and the
results are summarized on Table 18-8.
TABLE 18-8. VOLUME OF AERATION BASIN BASED ON MONOD MODEL
Set of +6
Coefficients Volume of Aeration Basin, 10 gals
1
2
3
4
ft
Case: A B C_
45 21 *
15 * *
8 4 *
27 * *
D
*
*
*
*
.1.
'MLSS = 4,000 mg/1, S = 15,700 mg/1 BOD.
tt °
Based on concentration of S.
*Unsatisfactory as 0 < 3 days.
C
380
-------
When compared to the volume calculated by scaling the AAS system, namely,
10.85 million gallons, the values on Table 18-8 show a variation from about
one half to four times. It is interesting to note that in most cases the cri-
terion of a 0 larger than three days cannot be met and, presumably, under
c
those conditions poor flocculation of microorganisms would result and the
effluent quality would tend to deteriorate as the suspended solids concen-
tration increases. Alternatively one can say that in these cases the size
of the aeration basin is controlled by 0 and not by the removal rate of BOD.
C
In conclusion it does not seem feasible to size the aeration basin on
the basis of inadequately known fundamental coefficients. The use of limited
information on coke plant wastes as the only basis for comparison is also not
completely satisfactory. In future pilot studies it is highly recommended
that these fundamental coefficients be evaluated so that a rational design
based on their values and the mathematical models presently available can be
made possible.
18.5 AIR ACTIVATED SLUDGE—NITRIPICATION-DENITRIFICATION (AASND) SYSTEM
Scheme 2 of the AAS system requires further treatment to remove ammonia.
Ammonia stripping is ruled out because of the air pollution and odor problem.
One possible alternative is to remove ammonia biologically through nitrifica-
tion and denitrification. An extensive description of these processes has
been given by Barber and Thompson.
Field Experience
Most of the field experience available to date comes from municipal
wastewater treatment. An extensive collection of design criteria and pro-
cedures is available in the literature. The use of an AASND system for
removing carbon and nitrogen compounds from coke plant wastes is limited to
pilot plant scale.9 The information obtained from these two sources has been
used as the primary basis of the preliminary design and subsequent cost esti-
mate.
A continuous-flow pilot plant study to evaluate the feasibility of
381
-------
9
available in this pilot study, and it might at least partly account for the
treating coke plant wastes by the AASND system was reported in 1973. The
results of the study indicate that the AASND system can be used to remove
organic carbon compounds and ammonia from diluted coke plant wastes. The
overall treatment efficiencies obtained include removal of greater than 99.9
percent phenols, 80 percent COD and 90 percent ammonia. The level of ammonia
removal depends strongly on the efficiency of ammonia oxidation to form
nitrate in the nitrification stage as any subsequent conversion of nitrate to
nitrogen is limited by the amount of nitrate available. However, the nitri-
fying organism has been found sensitive to many constituents in the coke
plant wastes. Dilution and efficient operation of the carbonaceous unit are
necessary to prevent inhibition and loss of nitrification efficiency. For
instance, the ammonia removal percentage was found to decrease from 60 per-
cent to 45 percent as the influent ammonia nitrogen concentration increased
9
from 400 to 1,200 mg/1, and ammonia-free water may be necessary for dilu-
tion. The BOD concentration in the influent to the nitrification unit, which
depends on how well the carbonaceous unit operates, can significantly affect
the nitrification efficiency. Unfortunately, this information on BOD is not
9
available in this pilot study, and it m:
fluctuation in nitrification efficiency.
Design
The design criteria developed from municipal experience and summarized
in an EPA technology transfer publication has been used for the preliminary
design. Additional information gained from the pilot study on coke plant
9
wastes has also been utilized. This includes the limiting ammonia nitrogen
concentration in the influent to the nitrification unit based on a certain
level of treatment efficiency. Concentrations of the various nitrogenous
species in the transformation process are summarized on Table 18-9. The
carbon removal unit has already been designed in Section 18-3. In nitrifica-
tion and denitrification, the reaction is the same regardless of the waste
medium, whether it is municipal sewage, coke plant wastes or coal conversion
wastes. However, the waste constituents besides ammonia may differ in their
effects on nitrification and denitrification. In this preliminary design
382
-------
TABLE 18-9. CONCENTRATIONS OF NITROGENOUS SPECIES IN NITRIFICATION
AND DENITRIFICATION PROCESSES
mg/1
N°3~-N Total N
Influent to Nitrification* 400 — 400
Influent to Denitrification 160 240** 400
Effluent from Denitrification 160 24*** 184
*Dilution is required to reduce NH -N concentration from 2,000
to 400 mg/1.
**Assume 60% of NH -N is oxidized in the nitrification unit.
***Assume 90% of NO ~-N is reduced in the denitrification unit.
these differences are assumed negligible, and the experience with municipal
sewage and coke plant wastes is assumed transferable to coal conversion
wastes.
The preliminary design of the AASND system is shown in Figure 18-5. For
the nitrification process, 15.8 x 10 gallons/day of ammonia-free dilution
water is necessary to reduce the influent ammonia nitrogen concentration to
400 mg/1; 86.9 tons/day of CaO and 39.3 tons/day of Na CO are needed for pH
control and provision of inorganic carbon respectively; and 4,035 hp are
necessary to pump air into the water for the nitrification reaction. In the
denitrification process, in addition to the reactors and clarifiers, 201 hp
are needed for mixing, 40 hp for nitrogen gas release and 70.4 tons/day metha-
nol for the provision of organic carbon. The treatment of sludge is essen-
tially the same as that for the AAS system.
Costs
The costs of the AASND system calculated according to Table 18-4
shown on Table 18-10. According to this cost estimate, the total cos,:
383
-------
U)
00
»oo
OtNITOinCATION
«MOV*L
Al«
OIIUTION
WATCH
..woW** (?>UX"%f)
" - N
- N
Figure 18
-5. Air activated sludge-nitrification-denitrification system.
-------
TABLE 18-10. COSTS OF AIR ACTIVATED SLUDGE—
NITRIFICATION-DENITRIFICATION SYSTEM
Capital Costs 10 $
Equalization 1.19
Carbonaceous BOD removal
Aeration Basins 2.38
Aerators 5.40
Clarification 0.29 8.07
Nitrification
Aeration Basins 2.44
Aerators 2.43
Clarification 1-04 5.91
Denitrification
Reactors °-46
Mixers & Aerators 0.14
Clarification 0-51 1.11
DAF Thickening 0.60
Vacuum Filtration 0.68
TOTAL 17.56
Operating Costs —1Q $/Yr
Amortization & other capital related items @ 15% of capital/yr 2.63
Maintenance
Concrete Work 0.06
Machinery °-17
Electricity
Carbonacious BOD @ 7,500 kw
Nitrification @ 3,400 kw
Denitrification @ 200 kw
11,100 kw 1.78
Chemicals
CaO 89.9 tons/day @ $54/ton 1.62
Na CO 39.3 tons/day @ 80/ton 1.05
CH OH 70.4 tons/day @ 127/ton • 2.98
; p 1.1 tons/day @ 840/ton 0.32 5.97
TOTAL 10.61
TOTAL ANNUAL COST 10.61 x 10 $/yr = 10.61 $/l,000 gallons
385
-------
about $10.29/1,000 gallons, with the cost of chemicals constituting more than
50 percent of the total cost. Ammonia separation, with recovery of ammonia
for sale as described in Section 16, is clearly the preferred procedure.
18.6 HIGH PURITY OXYGEN ACTIVATED SLUDGE (HPOAS) SYSTEM
The use of high purity oxygen, rather than air, as the source of oxygen
for biological waste treatment has gained increasing acceptance in the water
pollution control field. Because oxygen is required in many coal conversion
plants, it can be made available for water treatment at the cheapest possible
price. Approximately 3,000 tons/day of oxygen will be needed in a standard
size SNG plant, and the amount of oxygen required for the HPOAS system may be
about 10 percent of that required for coal conversion.
The use of an HPOAS system may have the following advantages over an AAS
system:
(1) Smaller oxygenation basins, and thus lower space requirements
(2) Less susceptible to upsets due to shock loads, and thus more stable
(3) Less power required for mixing because of smaller oxygenation
basins
(4) Less energy and cost for oxygen transfer, particularly when a large
quantity of oxygen is produced on site
(5) Less water loss due to evaporation as the oxygenation basins are
closed reactors.
To evaluate the HPOAS system, suppliers of these systems have been con-
sulted, particularly Linde Division of Union Carbide Corporation. Their
experience in the use of the HPOAS system for the treatment of coke plant
wastes, and in the pilot studies of HPOAS for treating coal gasification
wastes, has been used to establish the criteria for a preliminary design.
Design
The HPOAS system design consists of multitrains in parallel, with each
train consisting of multistages to obtain a quasi-plug flow condition. High
purity oxygen is fed to the space above the liquor level in each stage of the
386
-------
oxygenation basin, and oxygen transfer is accomplished by use of surface aera-
tors or equivalent. The dissolved oxygen concentration in the mixed liquor
will be maintained at about 5 mg/1, rather than 2 to 3 mg/1 as commonly used
in the AAS system.
Because of the high strength of coal conversion wastes, a high level of
treatment will be required. In order to reduce BOD to an acceptable level,
taken to be less than 60 mg/1, two steps of HPOAS treatment are used with
each step achieving about 95 percent removal of BOD.
Two key parameters for the design of activated sludge system are mean
F/M (food to microorganism) ratio and MLVSS (mixed liquor volatile suspended
solids). The F/M ratios for Step 1 and Step 2 differ because of the differ-
ence in BOD loading; F/M is 0.8 in Step 1 and 0.3 in Step 2. The MLVSS will
be substantially larger than that for the AAS system because of improved
settling velocities of the oxygen sludge, and the MLVSS in this case is
assumed to be 7,300 mg/1 in Step 1 and 4,500 mg/1 in Step 2. The clarifiers
2
are designed on the basis of an overflow rate of 400 gals/(day)(ft ) in Step
2
1 and 300 gals/(day)(ft ) in Step 2. The depth of the clarifiers is assumed
to be 15 feet. The design is summarized on Table 18-11.
The oxygen requirement, pounds of oxygen required per pound of BOD
17
removed, is a function of F/M and COD/BOD ratios. The effect of COD/BOD
ratio may be particularly significant in this case as the fate of COD in the
biological treatment of coal conversion wastes is unknown at present. The
oxygen requirement is assumed to be 1.03 Ib/lb BOD removed in Step 1 and
1.21 Ib/lb BOD removed in Step 2. Whenever COD needs to be evaluated in the
biological treatment, the removal of COD is assumed to be equal to that of
BOD; this assumption is conservative and should lead to a design on the safe
side.
The average oxygen utilization in the oxygenation basin depends on the
purity of the oxygen in the gaseous mixture which essentially consists of
feed oxygen and the carbon dioxide produced as a result of the biochemical
oxidation. Therefore the average oxygen utilization percentage will increas
as the feed BOD concentration decreases and is assumed to be 75 percent in
Step 1 and 80 percent in Step 2. Based on the oxygen requirement and ava.-^:
oxygen utilization efficiency, the amount of oxygen to be transferred car. b;
387
-------
TABLE 18-11. DESIGN OF THE HPOAS SYSTEMa
Design Basis
Flow, 106 gal/day. 3
BOD5, Ibs/day 450,000
BOD5, mg/1 18,000
COD, rag/1 28,000
COD/BOD5 1.56
Wastewater Temperature, °F 80°F
pH Adjusted as required
Nutrients Phosphorous to be added
System Design Step 1 Step 2
Flow, QUO6 gal/day) 3 3
Retention Time, hrs (based on feed flow) 74 16
MLSS, mg/1 7,800 5,100
MLVSS, mg/1 7,300 4,500
Sludge Recycle Rate, %Q 35 35
Mean Biomass Loading, Ibs BOD5/(lb MLVSS)(day) 0.8 0.3
Volumetric Organic Loading, Ibs BOD5/(10 ft )(day) 364 84
Average D.O. level, mg/1 5.0 5.0
Oxygen Supplied, tons/day 278.9 16.1
Average Oxygen Utilization Efficiency, % 79 80
Secondary Clarifier Overflow Rate gal/(day)(ft^) 400 300
Recycle Suspended Solids Concentration, wt % 2.0 2.0
Effluent Soluble BOD5 , mg/1 900 45
Utility Requirements Oxygenation Motors
• t. „ Step 1 Step 2 Total
Operating Energy: c— *•—
Brake HP 2,667 161 2,828
KWC 2,162 133 2,295
Connected Load:
Nameplate HP 3,150 180 3,330
Maximum Single Connected Load = 100 hp, across-the-line, starting
a
Preliminary information supplied by Union Carbide on the basis of
assumptions provided by WPA.
Used as basis for determining oxygen requirement.
Q
Includes motor losses.
388
-------
calculated.
The energy requirement is estimated as follows. The surface aerators
consume 1 hp-hr for 7.8 Ib oxygen supplied, or 191 kw-hr/ton oxygen supplied.
Air separation consumes about 330 kw-hrs/ton oxygen (Reference 29). Of this
330 kw-hrs, about 248 kw-hrs is required to drive the air compressors in the
air separation plant. In coal conversion plants it is assumed that the air
compressors are driven by steam turbines at 11,700 Btu/kw-hr, not by electri-
city. This leads to an energy requirement of 2.9 x 10 Btu as steam, plus
273 kw-hr per ton of oxygen. For quick estimating, as is necessary when
doing the complete water treatment plant block diagrams in Section 19, a
relationship is needed between the energy requirement and the BOD removed,
not between the oxygen supplied and the BOD removed. An average figure of
1.31 Ib oxygen produced/lb BOD removed and an energy requirement of 1,900 Btu
as steam plus 0.018 kw-hr per Ib BOD removed have been used.
A major design consideration is the control of water temperature in the
oxygenation basin. It has been recommended that the water temperature in
the aerobic biological treatment of coke plant wastes be 95-100°F throughout
the year. Attention must be paid to the temperature rise as a result of the
biochemical oxidation reactions: the oxidation of about 165 mg BOD/1 as
phenol can theoretically cause 1°F temperature rise at these concentrations.
Considering the various heat losses in oxygenation basins, it is assumed that
the removal of 200 mg/1 BOD will cause an increase in water temperature of
1°F. Since the removal of BOD in Step 1 is 95 percent of 18,000 mg/1, this
will result in a temperature rise of about 85°F. To maintain the proper tem-
perature in the oxygenation basin, it will be necessary to recycle
12 x io6 gallons/day of the mixed liquor at a temperature of about 97°F and
to reduce its temperature to 80°F in a cooling tower, as shown in Figure
18-6. This arrangement will also strip the CO2 produced from the mixed
liquor. The temperature of the 3 x 10 gallons/day feed is assumed main-
tained at 80°F from the equalization basin.
The preliminary design of the HPOAS system includes sizing the oxygena-
tion basins and clarifiers, and determination of oxygen and electric power
requirements. The procedures are outlined in the following paragraphs.
389
-------
NUTRIENTS
CCOUO EfflUtNT
HOM AMMONIA '
Ul
ID
o
SLUDGE
DISPOSAL '
Figure 18-6. High purity oxygen activated sludge (HPOAS) system.
-------
Oxygenation Basins
(1) Determine BOD (mg/1) to be removed on the basis of removal effi-
ciency expected.
(2) Determine F/M and MLVSS (mg/1) from field experience or pilot
studies.
(3) Calculate hydraulic retention time (RT ) by using the following
equation:
BOD (mg/1) x 24 (hrs/day)
(hrs)
Q " ' MLVSS (mg/1) x F/M [Ib BOD/(lb MLVSS)(day)]
(4) Calculate the total volume of the oxygenation basins (V) which is
the product of RT and Q (hydraulic feed rate).
(5) Calculate the total surface area (A) by assuming an economic water
depth in the oxygenation basin, usually 15 feet, A = V/15.
(6) Determine the total number of stages to provide the area, each
stage being a square in cross section.
(7) Design the overall configuration such that there are at least two
parallel trains with at least three stages of the oxygenation basin in each
train.
Recirculation of Mixed Liquor in Step 1
(1) Calculate the BOD removed in mg/1 in Step 1.
(2) Calculate the temperature rise of mixed liquor on the basis of 1°F
temperature rise per 200 mg/1 BOD removed.
(3) Calculate the recycle flow required to hold the overall temperature
rise in the mixed liquor passing through the oxygenation basin to about 15°F,
from 80°F to 95°F. Assume that the cooling tower will lower the temperature
of the recycled mixed liquor from 95°F to 80°F.
Clarifiers
(1) Assume the overflow rates for Steps 1 and 2 to be 400 and
300 gpd/sq ft respectively.
(2) Calculate the total surface area by dividing the hydraulic flow
rate to the clarifiers with the overflow rate.
(3) Determine the diameter of clarifiers by considering the number of
391
-------
clarifiers to be at least two and the diameter of each to be in the range of
30 to 100 feet.
(4) Assume the depth of clarifiers to be 15 feet for costing purposes.
Oxygen and Electric Power Requirement
(1) Calculate BOD removed in each step in pounds of BOD per day.
(2) Determine the oxygen requirement in terms of pounds of oxygen
required per pound of BOD removed.
(3) Determine average oxygen utilization efficiency for each step
according to the feed BOD concentration.
(4) Calculate the amount of oxygen to be transferred in pounds of oxy-
gen per day.
(5) The energy required for oxygen transfer is about 191 kw-hr per ton
of oxygen transferred, and that for oxygen generation is about 330 kw-hr per
ton of oxygen generated (partly as drive steam and partly as electricity).
The flow diagram of the HPOAS system is shown in Figure 18-6. The
design basis, system design and utility requirements are detailed on Table
18-11, which is preliminary information supplied by Union Carbide on the
basis of assumptions provided by us. Step 1 consists of seven trains, each
with five stages of concrete oxygenation reactors, while Step 2 consists of
two trains each with three stages. Their detailed configurations are shown
in Figures 18-7 and 18-8. It should be pointed out that 12 x io6 gal/day of
the mixed liquor from Step 1 will be recycled through a cooling tower to
reduce the water temperature from 95°to 80°F and to strip off CO . The bro-
ken line in Figure 18-6 shows the recycling of the clarified water through
the cooling tower for more flexible operation.
Costs
The costs of the HPOAS system including sludge treatment are summarized
on Table 18-12 showing the unit cost on the order of $3.61 per thousand gal-
lons of wastewater treated. It is interesting to note that the total elec-
tric power requirement including that for oxygen generation and transfer,
about 6,526 kw or 8,752 hp, is slightly less than that for the AAS system
392
-------
OXYGEN
INFLUENT
OJ
i£>
U)
r
4 x ]
5 x ]
and v
ickne
' i
— r
- I*
6
D
D
D
D
D
i
i
f ii
o
D
n
n
n
a
r 1
1
r I'
"A
a
a
a
n
a
a
a
a
a
a
-s —
a
a
n
a
n
r i f «
6
a
a
a
n
a
LO6 gal
LO6 gal
rolumes include no allowance
ssses and weirs.
r \
~i
a
a
D
a
a
48.5'
\
--~N
>RIFICATION
^^
48.
^
j1
1
1
48.5'
2.5'
I7.51
_L
EUVATION VJEW
EFFLUENT
(TO STEP 2)
Figure 18-7. Configuration of step 1 of HPOAS System.
-------
INFLUENT (FROM STEP I)
PIPELINE
OXYGEN
Stage volume 0.33 x 10 gal
System volume 2.0 x 106 gal
NOTE: Dimensions and volumes
include no allowance for
wall thicknesses and
weirs.
*f\
1 17.5'
K. ..H
ELEVATION VIEW
EFFLUENT
Figure 18-8. Configuration of step 2 of HPOAS System.
394
-------
TABLE 18-12. COST OF HIGH PURITY OXYGEN ACTIVATED SLUDGE (HPOAS) SYSTEM
Capital Costs 106$
Equalization 1.19
Step 1 HPOAS:
Oxygenation Basins 1,96
Clarification 0.25
Cooling Tower 0.08
Pumps for Recirculation 0.13
Step 2 HPOAS:
Oxygenation Basins 0,45
Clarification 0.29
Pumps for Recirculation 0.02
Oxygenation Equipment & Related Instrumentation
for Steps 1 and 2* 3.50
Installation of Oxygenation Equipment & Related
Instrumentation* 0.32
DAF Thickening 0.54
Vacuum Filtration 0.36
TOTAL 9.09
Operating Costs 106$/yr
Amortization & other capital-related items @ 15% of
capital/yr 1.36
Maintenance:
Concrete work 0.05
Machinery 0.08
Electricity @ 2,470 Kw** 0.40
Chemicals:
Phosphorous 0.32
Oxygen, 295 tons/day @ $14.32/ton (see 1.40
Table 18-13)
TOTAL 3.61
TOTAL OPERATING COSTS 3.61 $/1000 gal
^Quotation from Union Carbide
**Excluding electricity required for oxygen generation
395
-------
TABLEi 18-13. COST OF OXYGEN (MODIFIED FROM REFERENCE 29)
Capital of $13,000/(ton/day) amortized
at 15%/yr for 329 days/yr 5.94
Power, 330 kw-hr/ton at 2
-------
than two decades been found highly economical and desirable to achieve bio-
22
oxidation and water cooling in a cooling tower structure. Functionally,
the cooling tower in this case is analogous to the trickling filter in terms
of organic removal, although differences exist in their physical configura-
tion and design. We propose merging a trickling filter with a cooling tower
as an integral unit prior to the HPOAS. Since the removal of BOD may be more
significant than the removal of heat in such an application, the new unit
will be designated as an activated trickling filter (ATF). An activated
trickling filter as used here is a trickling filter of plastic medium loaded
continuously with the mixed liquor from the HPOAS units, as shown in Figure
18-9. The ATF is expected to achieve the following objectives:
(1) Reduce BOD by about 30 percent as a pretreatment to the HPOAS sys-
tem.
(2) Reduce the temperature of the recycled mixed liquor from the HPOAS
system from about 95°to 80°F.
(3) Strip off the excessive carbon dioxide from the recycled mixed
liquor.
Qualitatively, the use of an ATF-HPOAS system may be expected to have the
following advantages over the use of an HPOAS system alone:
(1) Less energy required. The energy required to pump water and drive
the air fans in the ATF may be lower than that to transfer the large quanti-
ties of air or to generate and transfer adequate oxygen for the activated
sludge process.
(2) Lower capital and operating costs.
(3) Fewer system upsets and higher treatment reliability. This is due
to the fact that fixed biological growth is less susceptible to loss of the
biota activity through shock loadings of either hydraulic feed, BOD concentra-
tion or toxicants. Recycling of the mixed liquor may also contribute to the
treatment reliability.
The use of an ATF in combination with an HPOAS system in the manner shown
on Figure 18-9 results in an organic loading of about
8,000 Ib BOD/(103ft3 of medium)(day) compared to current practice of having
high organic loadings in the range 1,000-1,400 lbBOD/(10 ft ) (day). This
occurs because the BOD concentration in the feed water is high and, also,
397
-------
NUTRIENTS
COOLED EFFLUENT
FROM AMMONIA
STILL
3xl068al$./do£"
EQUALIZATION
22.2xi06gal«./
-~*
f
F
S)
ID
8'
'
•^
k
2
V
STEP I HPOAS
i
k 9xl06gali./doy
K
^~^ r /-—^
/ STEP 1 \ STEP 2 HPOAS / ST^P 2
-*J CLAHIFIfATlON \ «. 5TfcF 2,HPOAS j-J n ABlcirATin
OJ
«5
CD
7.38x10
SLUDGE
DISPOSAL
Figure 18-9. Activated trickling filter—high purity oxygen activated sludge system.
-------
because the recirculation rate is determined by the cooling requirement of
Step 1 of the HPOAS system and is not adjusted to control the BOD loading of
the trickling filter. Also, there are contaminants in the coal conversion
wastewater other than phenol which may inhibit biochemical oxidation in the
ATF to some extent. For these reasons, as will be seen in the description
of the design procedure given below, the usual trickling filter design equa-
tion has been modified by assuming that the reduction in BOD obtained is only
30 percent instead of the 80 percent found by use of the standard design
equation. Furthermore, forced ventilation is used to avoid oxygen transfer
limitation.
The other assumptions that have been made in the design are (1) that
first order kinetics prevail in both steps of the HPOAS with 95 percent
removal in each step and (2) that without recycle the water temperature
would rise 1°F per 200 mg BOD/1 removed during passage through Step 1 of
the HPOAS system.
According to B.F. Goodrich General Products who manufactures plastic
medium for trickling filters, no difficulty is anticipated in running the
mixed liquor (MLSS ~ 6,800 mg/1) thrdugh the filter medium as long as the
MLSS does not exceed 10,000 mg/1 and the diameter of solid particles is less
than 0.5 inches. The suspended solids in the effluent from ATF will be effec-
tively removed in the HPOAS process, and the suspended solids content in the
effluent from the biological treatment system will be in the range 50 to
500 mg/l.2° In Section 19 a value of 100 mg SS/1 is assumed, and this is
filtered or not, as demanded by the downstream use or treatment.
Design
The BOD removal relationship for trickling filters may be stated as
1
where
L = BOD of the effluent, mg/1
e
399
-------
L = BOD of the influent, mg/1
cl
K = treatability factor
D = fill depth, ft
(T—20)
0 = temperature factor, = 1.035 , where T = water temperature in °C
2
J = water flux through the filter, gallons/(min)(ft )
n = medium factor
The treatability factor K is a measure of the susceptibility of the waste-
water to biological treatment. The medium factor n is a function of the
character of the medium, e.g., its specific surface area and geometry, etc.,
which can affect the contact time between water and biological mass on the
surface of the medium.
If the trickling filter effluent is recycled, then the influent BOD L
a
may be expressed as
L + NL
where
L ~ BOD in the raw wastewater before mixing with recycled flow, mg/1
o
N = recycle ratio = recycle flow/raw water flow
Also, the water flux through the filter J can be expressed in terms of the
raw wastewater flow as
j = (1 + N)j (6)
where
j = raw water flow (flow influent to the system) expressed in
gallons/(min)(ft2 of trickling filter)
Combining Equations (4) through (6) gives
L
o
-^ - (1 + N) exp
L
e
KD0
IjU+N)]
n
- N (7)
400
-------
For the first calculation, the following numerical values have been used:
K = 0.045 for coal conversion wastes
D = 18 ft for ease of ventilation
T = 30°C = 86°F
0 = 1.03510 = 1.41
n - 0.5
The recycle ratio N will be set so as to ensure the correct temperature
control in Step 1 of the HPOAS so Equation (7) relates the raw water flux
2
gallons/(rain)(ft ) to the fraction of BOD remaining. Since the desired treat-
ment rate in gallons/min is given, Equation (7) relates the desired fractional
removal of BOD to the cross-sectional area of the trickling filter.
It should be pointed out that since it is the mixed liquor from Step 1
of the HPOAS being recycled, rather than the filter effluent as used in the
derivation of Equation (7), using L as the effluent BOD from the ATF will
provide a conservative design.
Equation (7) has restrictions. To avoid excessive shear on the biomass,
the total flux J should not exceed about 4 gallons/(ft }(min) (information
from B.F. Goodrich General Products and Reference 22). The flux cannot be
very low or distribution of the water over the medium cannot be accomplished.
Also, it must be remembered that Equation (7) has been used for a BOD loading
of up to 1,000-1,400 Ib BOD/(10 ft filter medium)(day) in present practice.
Higher BOD loadings will tend to reduce the BOD removal efficiency as oxygen
transfer becomes limiting.
To control the undesirable effects of organic overloading, forced draft
ventilation will be provided. In the preliminary design modular units of ATF
designed for ease of counterflow ventilation, each 20 feet in diameter and
18 feet in height, have been used. The capital cost of each module, includ-
ing containment structure, filter medium, rotary distribution, fans and
installation, is estimated at $50,000.
The design is given on Table 18-14 and is detailed below for the HPOAS
and the ATF systems. The results are shown on Figure 18-9.
401
-------
TABLE 18-14. PRELIMINARY DESIGN CALCULATION OF ACTIVATED TRICKLING FILTERS
% BOD Removed
L /L
e o
L /L
o e
T(L /L ) + 3 I
ir ° e
111 L 4 J
j, gpm/sq ft
Total hydraulic loading, J = 4j
Cross-sectional area required = A, ft
Medium volume = ISA ft
Organic loading, Ibs BOD in raw-waste
per (1,000 ft3) (day)
No. of modular ATP units used
90 85 80 75
0.1 0.15 0.2 0.25
10 6.67 5 4
1.18 0.88 0.69 0.56
0.23 0.42 0.68 1.04
0.94 1.68 2.72 4.16
3,063
55,134
8,168
10
70
0.3
3.33
0.46
1.54
6.16
HPOAS System Design
(1) Assume that 30 percent of the BOD will be removed by the ATF and
95 percent of the remaining BOD is removed in Step 1 of the HPOAS;
11,970 mg/1 are removed in Step 1.
(2) Design the HPOAS as described in Section 18.6. In particular, find
the recycle rate for temperature control. There is 1°F rise for 200 mg BOD/1
removed, so the temperature rise in Step 1 is 60°F without recycle. Let R
be the recycle rate (10 gallons/day) when 3 x 10 gallons/day are processed.
The heat evolved which would be enough to heat 3 million gallons through 60°F
must only heat (R+3) million gallons through 15°F, so R = 9 * 10 gallons/day
and the recycle ratio N = 3.
ATF System Design
(1) Assume various percentages of BOD removed and calculate the water
flux from Equation (7). The calculations are shown on Table 18-14.
(2) Choose the column in which the total hydraulic loading is less
402
-------
2
than, and close to, 4 gpm/ft . This is the column with 80 percent BOD
removal.
(3) For the required feed rate (3 x 10 gal/day) find the total cross-
2 2
sectional area. At 0.68 gpm/ft this is 3,063 ft .
(4) Find the medium volume noting that a depth of 18 ft is used. The
volume is 53,134 ft .
(5) Find the organic loading based, in this case, on 18,000 mg/1 influ-
/- o -3
ent BOD = 0.45 x 10 Ib BOD/day. The loading is 8,169 lb BOD/(10 ft )(day).
This loading is six to eight times current practice, and it is assumed that
the trickling filter designed here, with forced ventilation included, will
produce only a 30 percent reduction in BOD.
(6) Find the number of modular units, each having 314 ft cross-
sectional area and costing $50,000.
In this example it is found that ten modules of ATF will reduce the BOD
by 30 percent. This adjustment to 30 percent is an educated guess, and it is
considered a conservative estimate in view of the flexibility provided by the
ventilation and the recycling of the mixed liquor. The verification of above
assumptions can be achieved by pilot studies.
Adjusting the Activated Trickling Filter Design to Different Throughputs
and BOD Loadings
The ATF-HPOAS system turns out to be the system of choice among all bio-
treatment systems in this study. A method of scaling the example to many
site-specific plants is therefore needed. This was done as follows:
(1) Equation (7) was used to adjust K, the water treatability factor,
to obtain a 30 percent reduction in BOD in the exemplary plant. If in Equa-
tion (7) one puts
0 = 1.41°C
n = 0.5
N = 3
D « 18 ft
2
j = 0.68 (from Table 18-14) gal/(min)(ft of trickling filter)
L /L - 0.7
e o
403
-------
then
K = 6.6 x io"3 (ft"1)(°c"1)
(2) For a requirement of 30 percent reduction in BOD, Equation (7) now
becomes
, |1.43 + N1 0.1675
(3) For each design, N is set by the cooling requirement in Step 1 of
the HPOAS system. Equation (8) gives the feed water flux, j.
2 2
(4) If the total water flux J = (l+N)j exceeds 4 gpm/ft , use 4 gpm/ft .
(5) The total area of the ATF, the number of modules and the total cost
are then calculated as for the example above.
Costs
The costs of the complete system are shown on Table 18-15. Step 1 of
the HPOAS consists of five trains,'each having five stages of oxygenation
reactors. Step 2 consists of one train having four stages in series. The
unit cost is about $3.10 per thousand gallons of wastewater treated, 16 per-
cent less than that of the HPOAS system alone (which was about $3.6 per thou-
sand gallons). The total power requirement including that for oxygen genera-
tion, oxygen transfer and pumping is about 4,770 kw and is significantly less
than that for HPOAS alone, which is about 6,526 kw. It may be concluded with
confidence that the use of ATF prior to HPOAS may significantly improve the
cost effectiveness of the biological treatment. It should be pointed out
that the 30 percent BOD removal by ATF is considered a conservative estimate,
as indicated on Table 18-14, and no optimization has been attempted.
The activated trickling filter-high purity oxygen system is the chosen
system wherever biotreatment is used in the complete water treatment plant
designs of Section 19. For ease in rapid cost estimating, the design and
costing procedures detailed above were used to cost the four cases shown on
404
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TABLE 18-15. COSTS OF ACTIVATED TRICKLING FILTER - HIGH PURITY
OXYGEN ACTIVATED SLUDGE (ATF-HPOAS) SYSTEM
106$
Capital Costs
Equal!zation
ATF
Containment Structure
Filter Medium
Rotary Distributor
Ventilation Fans
Installation
Step 1 HPOAS
Oxygenation Basin
Clarification
Pumps for Recirculation
Step 2 HPOAS
Oxygenation Basin
Clarification
Pumps for Recirculation
Oxygenation Equipment and Related Instrumentation
for Steps 1 and 2*
Installation of Oxygenation Equipment and
Related Instrumentation*
DAF Thickening
Vacuum Filtration
TOTAL
Operating Costs
Amortization and Other Capital-Related Items
@ 15% of capital/yr
Maintenance
Concrete Work
Machinery
Electricity @ 1,930 kw**
Chemicals
Phosphorous
Oxygen, 218 tons/day @ $14.23/ton
TOTAL
TOTAL OPERATING COST
BOD Removed
1.19
0.50
1.71
0.18
0.13
0.47
0.20
0.02
3.10
0.28
0.54
0.36
8.68
106$/yr
1.30
0.05
0.08
0.31
3.10
3.I/thousand gallons
224,7 tons/day
*Quotation from Union Carbide.
**Excluding energy for oxygen generation.
405
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TABLE 18-16. UNIT COSTS OF BIOLOGICAL TREATMENT
USING ATF-HPOAS SYSTEM
Coal. Water Flowrate
Conversion .,
Process 10 Ibs/hr 10 gals/yr
Hygas 273 263
SRC 256 246
Lurgi-
Power 581 556
Lurgi-
Power 1,040 999
Cost of
Biological BOD „ . . _
Treatment _ _. „ Removal "n^ Cost
6 Influent $/lb BOD
10 $/yr BOD mg/1 10 Ibs/yr Removed
0.7 13,000 28.4 0.025
1.27 30,000 61.5 0.021
1.68 18,000 83.3 0.020
3.1 18,000 149.7 0.021
Table 18-16. Except in the Hygas plant, which has a low throughput and the
lowest loading, the cost seems to be directly proportional to the BOD removal
rate. In Section 19 the costs used are $0.025/lb BOD removed for the Hygas
plants and $0.021/lb BOD removed in the other plants.
18.8 ANAEROBIC BIOLOGICAL TREATMENT
In the anaerobic biological treatment processes, organic constituents of
the wastewater are converted to methane and carbon dioxide. General explana-
tions are given in Reference 23. The significant advantages of anaerobic
over aerobic biological treatment are:
(1) Less sludge growth will be produced in the anaerobic process, about
0.05 to 0.2 pounds suspended solids per pound BOD. while the corresponding
L 17
figure for the aerobic process may be as high as 0.4. Variations in the
above figure primarily depend on the nature of organic constituents and the
solids retention time (SRT).
(2) The problem of biological sludge disposal and the requirement of
nutrient, phosphorus in this case, will be significantly reduced.
(3) Elimination of aeration results in significant savings in energy
and in the capital and operating costs.
406
-------
(4) The methane gas produced can serve as a source of fuel.
On the other hand, the major disadvantages of anaerobic processes lie in
their relatively high susceptibility to upsets due to toxicants or shock
loads and the lack of field experience with the coal conversion wastewater.
The dry solid residue generated through biological treatment will be on
the order of 10-20 tons/day, while the total ash generated from the coal con-
version facilities may be one hundred times as much, or more. Therefore, the
problem of biological sludge disposal is relatively insignificant in coal
conversion facilities.
As to the methane generation from anaerobic treatment processes, it is
also quantitatively insignificant in comparison with the total gas production
from the coal conversion facilities, about 250 million scf/day. Assuming
that 90 percent of 30,000 mg/1 COD may be converted to methane at 5.61 cubic
C.
foot methane per pound of COD, a 10 gal/day wastewater plant will be able to
produce up to 1.3 million scf/day of methane. While this is small compared
to plant production, it is an important part of the plant driving energy and
can conveniently be used on site (for example, for drying coal, etc.).
The major advantages of utilizing anaerobic biological treatment for
coal conversion wastewater will lie in the cost savings due to the elimina-
tion of aeration, consisting of both capital and energy expenditures, and, to
a lesser extent, the reduction in nutrient requirements.
As to the disadvantages of anaerobic treatment for coal conversion
wastewater, upsets of methane-forming bacteria by toxicants existing in the
wastewater or shock loads are probably the most critical considerations.
Methane-forming bacteria usually exhibit very slow growth and are relatively
sensitive to constituents in the wastewater. Once the anaerobic process suf-
fers an upset, it usually takes weeks or months to have methane-forming bac-
teria re-established.
The first constituent of concern is phenol if anaerobic processes are
used for the treatment of coal conversion wastewater. There have been con-
tradictory reports on the biodegradability of phenol in anaerobic treatment
processes. Some researchers have found phenol to be nonbiodegradable (cor-
respondence with Prof. P.L. McCarty and Reference 24), while others indicated
that phenol may be used as the sole source of organic carbon and, based on
407
-------
gas production data, at least 75 percent of phenol may be used (correspon-
dence with Prof. J.S. Jeris).
At concentrations exceeding certain levels, phenol will inhibit methane
production and thus render the anaerobic process unfunctional. The inhibi-
tory concentration of phenol seems to vary depending on the types of testing
conducted. In laboratory studies, 195 mg/1 and 300 mg/1 of phenol fed to an
24
anaerobic system have been found to be nontoxic. A more recent study
showed that for an active methane-producing, unacclimated, domestic digester
sludge the 50 percent inhibitory concentration of phenol (based on gas pro-
duction) was found to be larger than 1,000 mg/1 under substrate-limiting con-
ditions, and it ranged from 300 to 1,000 mg/1 (averaging about 400 mg/1)
under nonsubstrate-limiting conditions. Through acclimation the inhibitory
concentration of phenol will increase. Under nonsubstrate-limiting condi-
tions with a phenol concentration of about 500 mg/1, the methane-forming acti-
vity of the acclimated biomass was found to be about 95 percent of that for
the noninhibited.
::n a series of packed-bed anaerobic reactors the effect of four mixed
inhibitors, including phenol, formaldehyde, acrylonitrile and ethyl aerylate,
24
was also studied. The mixed inhibitors were found to act synergistically
since they were more detrimental to methane formation as a mixture than would
have been predicted based on individual chemical inhibition tests.
24
Acclimation was found to be unsuccessful in a digester-type reactor
unless complete mixing was provided. No such problem was observed in the
packed-bed type reactor (upflow anaerobic filters). It was also concluded
that acclimation in anaerobic systems was at best a slow process and consi-
derable difficulty would likely be experienced in the start-up of a full-
scale anaerobic process treating an inhibitory waste.
Other phenolic compounds may be more toxic than phenol to biological
treatment processes. For instance, pure cresols have been found to be rela-
tively nontoxic at a concentration of 250 mg/1, while the corresponding con-
centration for pure phenol has been found to be above 500 mg/1. Further-
more, the large ratio of COD to BOD in the coal conversion wastewater indi-
cates that although phenol constitutes most of the BOD, substantial quanti-
ties of other oxygen-demanding substances exist, most of which are probably
408
-------
organic in nature. The effects of these biologically refractory substances
on the anaerobic process are unknown at the present time and deserve future
study. However, in view of the sensitive nature of anaerobic reactions it
is likely that adverse effects on anaerobic reactions may be exerted by these
refractory substances.
Based upon the information presently available, if anaerobic treatment
is found feasible for coal conversion wastewater the most likely system will
probably be in the form of upflow anaerobic filters since it may be more
stable than the anaerobic activated sludge or conventional digester sys-
''3-26
terns. ' These different anaerobic systems are schematically shown in
Figure 18-10. However, the feasibility and performance of these systems
should be fully evaluated by pilot testing before any attempt at preliminary
design can be justified. Therefore, for this report no further discussion
will be made on anaerobic treatment.
18.9 ADDITIONAL CONSIDERATIONS AND RESEARCH NEEDS
It will be apparent from the discussion on biological treatment that all
possible reactor configurations have not been considered and that for those
reactors which were considered, some empirical and subjective design proce-
dures had to be used because of a lack of basic information. In this sub-
section these facts are summed up lest they be forgotten.
The components of a biological treatment system are:
(1) the wastewater
(2) the biological agents
(3) the treatment reactors.
The discussion can conveniently be arranged in this way.
Wastewater
The main consideration here is the biodegradability of wastewater con-
stituents. The use of gross parameters, like BOD, COD and TOC as used above,
is not adequate to answer the following key questions which should be studied
in future research:
409
-------
MIXING
INFLUENT
EFFLUENT
CONVENTIONAL PROCESS
INFLUENT
MIXING
MIXED h^-=H EFFLUENT
LIQUOR
RETURN
WASTE ORGANISMS
ANAEROBIC ACTIVATED SLUDGE PROCESS
.C
INFLUENT
EFFLUENT
UL .«. <_t_t
ROCK OR EQUIVALENT
ANAEROBIC FILTER PROCESS
Figure 18-10. Schematic anaerobic treatment systems
(modified from Reference 23).
410
-------
(1) What constituents are toxic to biochemical degradation, and what
are their threshold concentrations?
(2) What is the expected COD and TOG removal by biological treatment?
(3) What are the constituents of the remaining COD and TOC, and what is
their fate after biological treatment?
Biological Agents
Most biological agents utilized in wastewater treatment are mixed cul-
tures. With respect to the treatment of coal conversion wastes, the follow-
ing questions need to be answered:
(1) What is the feasibility of treating coal conversion wastes by
anaerobic processes which can result in a very high energy saving?
(2) Are there any specific biological strains that could work more
satisfactorily than mixed cultures?
(3) What are the characteristics and environmental requirements of
these strains?
(4) What are the values of the fundamental coefficients in the reaction
between biological agents and substrates (wastewater), for instance, the
coefficients in the Monod model?
Treatment Reactors
In general, for either aerobic or anaerobic treatment processes reactors
may be classified according to the mode of biological growth and the medium
involved:
(1) suspended growth
(2) fixed-bed growth
(3) fluidized-bed growth
For aerobic treatment, according to the discussion in Section 18.7, the
use of a fixed-bed (ATF) appears to be more energy-effective and cost-effec-
tive than the use of suspended growth (HPOAS), while the suspended growth
tends to provide a better effluent quality than the fixed-bed growth does.
A definitive conclusion on such a comparison can only be obtained through
411
-------
pilot studies using actual wastewaters.
The use of a fluidized-bed reactor is not developed enough to have been
costed in this study, but it is discussed below.
Tapered Fluidized-bed Bioreactor
Scott et ai at oak Ridge National Laboratory have experimented with
the use of a tapered fluidized-bed bioreactor for the removal of phenol. The
significant characteristics and findings of this study include:
(1) The experimental bench-scale system utilized fluidized anthracite
coal in the particle size range of 0.15 to 0.18 mm.
(2) A mutant strain of Pseudgmonas bacteria, under the trade name
Phenobac, was found superior to mixed cultures from activated sludge system
and was used for the experiment.
(3) Synthetic feed solutions made from deionized water and reagent grade
of phenol and other additives were used.
(4) The phenol concentration in the feed stream ranges about 20-60 ppm,
while the effluent phenol concentration is mostly less than 1 ppm and, in
certain instances, less than 25 ppb.
(5) Under the experimental conditions when oxygen transport is not
limiting, a multistage tapered fluidized-bed system appears to be at least
ten times more efficient than the conventional stirred-tank reactors in terms
of the rate of phenol conversion. Thus a significant decrease in reactor
volume could be expected with the fluidized-bed system (the difference in
cost has not been estimated).
(6) If higher phenol concentration exists in the influent, oxygen trans-
fer will become limiting.
(7) It was observed that the Pseudomonas used may be stored at 4°C
almost indefinitely for later use, and after Pseudomonas is reintroduced into
the bioreactor the system will be ready for normal operation in a very short
time.
(8) The use of fluidized-bed bioreactors alleviates operational prob-
lems associated with biomass build-up and allows easy removal or addition of
the active materials.
412
-------
(9) The tapered reactor tends to stabilize the fluidized bed while
allowing a much wider range of operating conditions, namely, flow variations.
According to the information outlined above, tapered fluidized-bed bio-
reactors appear to be a promising system to provide polishing treatment to
ensure an effluent phenol concentration below 1 mg/1 or less. However, in
order to improve biological treatment of coal conversion wastes, further
research on tapered fluidized-bed bioreactors is needed in the following
areas:
(1) Pilot testing of tapered fluidized-bed bioreactors to obtain design
information, using actual coal conversion wastewater rather than synthesized
feed solutions.
(2) Aerobic biological treatment of high-strength phenolic wastes may
well be limited by oxygen transfer rather than being limited biokinetically.
Therefore, the oxygen limitation should be fully characterized and means
determined to circumvent it.
(3) The fact that a mutant strain of Pseudomonas may be stored and used
more effectively than mixed cultures from activated sludges deserves further
characterization, probably involving the determination of various fundamental
coefficients in the biochemical transformation reaction.
413
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REFERENCES SECTION 18
1. Alexander, H. C., Blanchard, F. A., and Takahashi, I. T., "Biodegrada-
tion of 1 C-Phenol by Activated Sludge," presented at ACS 168th
National Meeting, Atlantic City, New Jersey, September 1974; Div. of
Fuel Chemistry Preprints 1.9_(5) , 104-112.
2. Cooke, R., and Graham, P. W., "The Biological Purification of the
Effluent from a Lurgi Plant Gasifying Bituminous Coals," Int. J. Air
Wat. Poll, Vol. 9, pp. 97-112, 1965.
3. Jablin, Richard, and Chanko, Gary P., "A New Process for Total Treat-
ment of Coke Plant Waste Liquor," AIChE Symposium Series, Vol. 70,
No. 136, Water-1973, pp. 713-722.
4. Adams, C. E., "Treatment of a High Strength Phenolic and Ammonia Waste-
stream by a Single and Multi-stage Activated Sludge Process," Proceed-
ings of the 29th Annual Industrial Waste Conference, Purdue University,
West Lafayette, Ind., May 1974, pp. 617-630.
5. Adams, Jr., C. E., Stein, R. M., and Eckenfelder, Jr., W. W., "Treatment
of Two Coke Plant Wastewaters to Meet EPA Effluent Criteria," presented
at 27th Purdue Industrial Waste Conference, May 1974.
6. Short, T. E., DePrater, B. L., and Myers, L. H., "Petroleum Refining
Phenolic Wastewaters," presented at ACS 168th National Meeting, Atlantic
City, New Jersey, September 1974; Div. of Fuel Chemistry Preprints,
19_(5), 113-148.
7. Gurnham, C. F., Industrial Wastewate^Contrpl, Academic Press, 1965,
pp. 228-233.
8. Kostenbader, P. D., and Flecksteiner, J. W., "Biological Oxidation of
Coke Plant Weak Ammonia Liquor," Journal Water Pollution Control
Federation, jll(2) , 199-207, February 1969.
9. Barker, J. E., and Thompson, R. J., "Biological Removal of Carbon and
Nitrogen Compounds from Coke Plant Wastes," EPA-R2-73-167, April 1973.
10. Sawyer, C. N., and McCarty, P. L., Chemistry for Sanitary Engineers,
McGraw-Hill Book Co., 1967.
414
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11. "Analyses of Tars, Chars, Gases and Water Found in Effluents from the
Synthane Process," Bureau of Mines Technical Progress Report, p. 3,
January 1974.
12. Process Design Techniques for Industrial Waste Treatment, AWARE, Inc.,
Enviro Press, Nashville, Tenn., 1974.
13. Mynhier, M. D., and Grady, C. P. Leslie, Jr., "Design Graphs for
Activated Sludge Process," Journal of the Environmental Engineering
Division, ASCE, Vol. 101, No. EE5, Proc. Paper 11659, October 1975,
pp. 829-846.
14. Water Quality Engineering for Practicing Engineers, by W. Wesley
Eckenfelder, Jr., Barnes & Noble, Inc., 1970.
15. Wastewater Engineering by Metcalf & Eddy, Inc., McGraw-Hill, 1972.
16. Nitrification and Denitrification Facilities Wastewater Treatment.
EPA Technology Transfer, August 1973.
17. Oxygen Activated Sludge Wastewater Treatment Systems, EPA Technology
Transfer, August 1973.
18. Bryan, Edward H., "Two-stage Biological Treatment Industrial Experi-
ence," Proceeding of Eleventh Southern Municipal & Industrial Waste
Conference, North Carolina State University, 1962.
19. Smith, R. M., "Some Systems for the Biological Oxidation of Phenol-
Bearing Waste Waters," Biotechnology and Bioengineering, Vol. 5, pp.
275-286, 1963.
20. Fleischman, M., "Reuse of Wastewater Effluent as Cooling Tower Makeup
Water," Proceedings of the Second National Conference on Complete Water
Reuse, AIChE and EPA, 1975, pp. 501-514.
21. Eckenfelder, W. W., and Ford, D. L., Water Pollution Control, Pemberton
Press, 1970, Chap. 13.
22. Mohler, E. F., Jr., and Clere, L. T., "Bio-oxidation Process Saves
H-?0." Hydrocarbon Processing, October 1973.
23. McCarty, P. L., "Anaerobic Treatment of Soluble Wastes," Advances in
Water Quality Improvement, E. F. Gloyna and W. W. Eckenfelder, Jr.,
Eds., pp. 336-352, University of Texas Press, Austin, Tex., 1968.
24. "Identification and Control of Petrochemical Pollutants Inhibitory to
Anaerobic Processes," Environmental Protection Technology Series,
EPA-R2-73-194, April 1973.
25. Jennett, J. C., and Dennis, N. D., "Anaerobic Filter Treatment of
Pharmaceutical Waste," J. Water Poll. Control Fed^, 4J7, 104, 1975.
415
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26. Mueller, J. A., and Mancini, J. L., "Anaerobic Filter Kinetics and
Application." Paper presented at 30th Annual Ind. Waste Conf., Purdue
University, West Lafayette, Ind., May 1975.
27. Cousins, W. G., and Mindler, A. B., "Tertiary Treatment of Weak Ammonia
Liquor," J. Water Pollution Control Federation, 44_(4) , pp. 607-618,
April 1972.
28. "Cost Curves for Basin Plans," Division of Planning and Research, State
Water Resources Control Board, State of California, January 1973.
29. Hugill, J. T., "Cost Factors in Oxygen Production," presented at
Symposium on Efficient Use of Fuels in Metallurgical Industries, Insti-
tute of Gas Technology, Chicago, 111., December 1974.
30. Badger, E. H. M., and Jackman, M. I., "Loadings and Efficiencies in
the Biological Oxidation of Spent Gas Liquor," Coke and Gas, Part 1,
pp. 316-323, August 1958, Part 2, pp. 426-434, October 1960.
31. Scott, C. D., Hancher, C. W., Holladay, D. W., and Dinsmore, G. B.,
"A Tapered Fluidized-bed Bioreactor for Treatment of Aqueous Effluents
from Coal Conversion Processes," presented at Symposium on Environ-
mental Aspects of Fuel Conversion Technology II, Hollywood, Fla.,
December 15, 1975, Environmental Protection Agency, Research Triangle
Park, N.C., EPA-600/2-76-149.
416
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SECTION 19
SITE STUDIES - 2: WATER TREATMENT PLANTS
19.1 INTRODUCTION AND CONCLUSIONS
The study is divided into three sections by location, which fixes the
assumed water supply and quality. The first part of the study is concerned
with plants located in North Dakota. The water supply is adequate, cheap and
clean. Because of this, reasonably generous amounts of water are evaporated
for cooling; in all cases the cooling tower makeup rate is several times the
plant effluent condensate rate. The foul condensate only has to be treated
to a suitable quality for the cooling towers or for disposal as dust control
and ash sluice water. Discussion centers on the cooling tower, the largest
water consumer. The assumption of a high chloride concentration in process
condensate leads to the special requirement of desalting this water for the
power plant.
In New Mexico the source water is assumed to be brackish. Therefore
cooling water consumption is reduced from the consumption in North Dakota.
Most of the discussion concerns where and how to apply desalting technologies.
In Wyoming, source water is assumed to be very scarce allowing the study
of an exemplary plant in which foul condensate, or sewage from the satellite
town, should be treated for feed to a boiler.
Various schemes are considered. For the schemes selected the approxi-
mate costs (expressed as $/10 Btu of product fuel or C/kw-hr) are given on
Table 19-1 along with the heat and electricity requirements. A word is
needed about accuracy. The costs in this report are minimum: they do not
include engineering fees or other than the bare minimum investment cost;
land cost is omitted as is all site preparation and most foundation work;
417
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TABLE 19-1. APPROXIMATE COST AND ENERGY REQUIRE-
MENTS FOR TOTAL
Approximate Cost*
$/hr $/1()6 Btu product fuel
or $/kw-hr power
Hygas
North Dakota 287 0.028
New Mexico 404 0.040
Wyoming 291 0.029
Syn thane
Wyoming 664 0.066
SRC
North Dakota 216 0.016
New Mexico 185 0.014
Wyoming 302 0.022
Electric Power
North Dakota 835 0.084
New Mexico 774 0.077
Wyoming 566 0.057
PLANT WATER TREATMENT
Approximate Energy**
10 6 Btu % Product
hr Energy
67 640 0.74
67 2290 0.93
67 665 0.74
136 2890 1.7
84 1670 0.75
56 1820 0.58
124 — 0.92
230 3750 2.7
150 5650 2.1
170 2170 1.9
*Approximate operating plus amortized capital costs with credit
taken for sale of ammonia.
**Total energy is sum of steam and electricity converted at
10,000 BtuAw-hr in power plants and 11,700 Btu/kw-hr in
gas and SRC plants.
418
-------
there is no allowance for a control building or laboratory. In operating
costs the labor, laboratory and overheads have been omitted. However, the
costs are quite good enough to distinguish between alternatives, which is
the major purpose. A maximum cost for operating the overall water treatment
plant will not exceed three times the numbers of Table 19-1 and will probably
cluster around twice the numbers of Table 19-1.
As a process was moved from site to site, very different qualities and
quantities of intake water were assumed. In spite of this the maximum cost
variation from site to site is 1.7 times for SRC and 1.4 times for Hygas and
power plants. The cost is unlikely to exceed 5 percent of the sale price of
the product fuel for any of the plants.
In the Hygas process, boiler feed water treatment is 31 to 32 percent of
the total water treatment cost in North Dakota and Wyoming. However, in New
Mexico, with a brackish water intake, the desalting costs including boiler
feed are 49 percent of the total. The brackish intake water makes the cost
in New Mexico higher than the cost of the other two sites.
For SRC, Power and Synthane, the cost of boiler feed water treatment is
in the range 11 to 23 percent of the total water treatment cost at North
Dakota and Wyoming; at New Mexico the cost of desalting plus boiler feed
treatment is 41 to 46 percent of the total. However, for these processes
other tilings affect the cost even more and the added cost of desalting is
swamped by other factors (discussed in later sections).
The cost of water treatment for the Synthane plant at Wyoming is higher
than for the Hygas plant at the same site. When compared to Hygas, the Synthane
process is assumed to put out about twice the foul condensate and to have
about twice the concentration of contaminants. As a result, 60 percent of the
difference is in the biotreatment and the rest is in other treatments applied
to foul condensate. A fourfold increase in contaminant load may not occur,
but if it does it will increase water treatment costs at least twofold.
Water treatment cost for power plants which employ Lurgi technology is
higher than for the fuel production plants. Note that the units used in
Table 19-1, namely S/106 Btu and «/kw-hr, are comparable: about 10,000 Btu
are used to make 1 kw-hr. However, power plants have only half the effi-
ciency of fuel plants, so water costs based on product energy will be doubled
by this fact alone. Furthermore, the higher fuel condensate effluent rar•
419
-------
caused by an intake of wet coal to the Lurgi gasifiers will tend to raise
costs over those in a gas producing plant using dry coal. Costs in North
Dakota and New Mexico hardly differ and are higher than in Wyoming, because
desalting is needed at both of the first two sites. In North Dakota chloride
is introduced in the coal and in New Mexico salt is introduced in the source
water.
In considering the energy requirements it should be noted that the
direct steam used, tabulated as 10 Btu/hr in Table 19-1, is never less than
2.5 times the steam needed to generate the electricity also tabulated. Since
desalting technologies, reverse osmosis and electrodialysis use electricity
and not direct steam, it is apparent that salt removal plays only a small
part in energy consumption. In fact, the energy required is very much con-
trolled by the energy for ammonia separation which is directly proportional
to the rate of production of foul condensate. Whereas sale of ammonia can
offset cost, there is no offsetting factor to obscure the importance of this
treatment in energy consumption. In SRC plants where foul condensate has a
particularly high contaminant concentration, biotreatment, solvent extraction
and evaporation add importantly to the energy load.
In designing the fuel plants we had to decide how much electricity to
generate on site. The two big users of electricity are the circulation pumps
for the acid gas adsorption liquid and the water treatment plants. In all
the plants, adequate electricity seems to have been produced. In the Hygas
plants 27,350 kw were generated and 8,440 kw were used for the acid gas remo-
val area. In the Synthane plant 31,000 kw were generated and 9,000 kw were
used for acid gas removal. In the SRC plants 10,250 kw were generated and
2,300 to 3,400 kw were used for acid gas removal. It would certainly be an
improvement to vary the electricity generated with the electricity needed for
water treatment, but this would not alter the plant efficiencies very much.
In the gas plants about 300 * 10 Btu/hr were allowed for wastewater
treatment and other unspecified uses. This seems adequate. In the SRC
6 6
plants 160 * 10 to 220 x io Btu/hr were allowed, which also seems adequate.
In the power plants 1.4 to 1.8 percent of the coal feed is allowed for
wastewater treatment. This is about 4.2 to 5.4 percent of the electricity
generated and is adequate.
420
-------
19.2 PLANTS IN NORTH DAKOTA
The Hygas plant uses dry coal and has a low rate of foul condensate
effluent. Furthermore, this effluent is not expected to be high in phenol,
so extraction of phenol is not likely to pay for itself. The water manage-
ment schemes shown on Figures 19-1 to 19-3 have been studied. The first dif-
ference is in details around the cooling tower. In fact, as will first be
shown, Scheme 2 is probably preferred. In Scheme 1 the wastage of biocides
and anticorrosion chemicals in cooling tower blowdown is too expensive.
Consider, first, the cooling tower treatment for Scheme 1. An analysis
of biotreatment effluent water is required and the analysis is shown on Table
19-2. The total makeup water = 1071 x 10 lb/hr. The desired concentration
ratio between makeup and blowdown (desired cycles of concentration) =
1071/281 = 3.8. The analyses of various waters are shown on Table 19-3. The
source water must be acidified to remove bicarbonate (see Section 15.6 for
discussion). The ammonia is assumed not to concentrate because it is vola-
tile.
TABLE 19-2. ASSUMED ANALYSIS OF BIOTREATMENT EFFLUENT WATER
mg/1 mg/1
Suspended solids 100 P04 15
Ca++ * NH3 20
Na+ * Cl~ in North Dakota 1200
Mg++ 15 in New Mexico 200
HCC
SO." 0
HCO " 40 in Wyoming 600
In North Dakota, and probably in Wyoming, much of the chloride enters
biotreatment in the form of NH4C1. As ammonia is consumed in the biotreat-
ment, lime or soda ash will be added to neutralize the chloride. The con-
centration of Ca++ and Na+ will depend on the choice of treatment. A minimum
concentration of 20 mg/1 Ca++ has been assumed.
421
-------
N)
I6/O *"•" BOILER ^^^^
^
^ rccu PRO-
TREATMENT J k KKUC
i CLEAN
CONDENSATE ^
POLISHING "^
19 ^ POTABLE
TREATMENT
1
AS
:ESS
19
19
f
it*
CLARIFIED
FOUL
WATER
9
/SERVICE a\
790 f SANITARY I
EVAPORATION V USES /
A
>
' 79° /COOLING'X^ 281
>( TOWER )^ BIOTRtAH
^^ 1
Y
11
f
AENT
273
CHEMICALS SLUDGE
/FLUE GAS DESULFURIZATIONM
1 >•{ ASH DISPOSAL
V OUST CONTROL t
)
^
f
AMMONIA
SEPARATION
V AA
AMMONIA
UNITS-tO4 lb./V.
Figure 19-1. Water treatment plant block diagram for Hygas process at
North Dakota (Scheme 1).
-------
M
78 869 BOILER FEED
^ 86y ^ TUPATMFMT 8I<
~ SCHEME 2 OF
FIGURE 19-3
SOLUBLE 1 CC
WASTE | PC
T
19 POTABLE
7REATMI
79
EVAPORATION
__». LIME 79° / COOLING V-J*
" SOFTENING ^^ TOWER f*
s*k
l.ntMlV.Au J 1
SLUDGE
-------
^
L
HOT LIME
SOFTENING
\
SOFTENING
IX
{
STONG-ACID
IX
1
WEAK-BASE
IX
• i
MIXED-BED
IX
•^
SCHEME 1
ro
\
1
WEAK -ACID
IX
\
DEC A!
\
No
SOFTENING
^
f
STRONG-ACID
IX
\
f
WEAK-BASE
IX
}
'
STRONG-BASE
IX
\
MIXEC
1
SCHEME 2
Figure 19-3. Boiler feed water treatment schemes in North Dakota,
-------
TABLE 19-3. ANALYSIS OF COOLING TOWER STREAMS FOR HYGAS
IN NORTH DAKOTA, SCHEME 1, FIGURE 19-
Flow
(103 Ib/hr)
SS
Ca
Mg
HC03
S0tt
Cl
sio2
PO,
NH3
Source
Water
790
mg/1
2
49
19
180
170
9
7
0
0
Biotreatment
Effluent
281
mg/1
100
650
15
40
0
1200
0
15
20
Source Water
After
Acidification
790
mg/1
2
49
19
0
312
9
7
0
0
Mixed
Makeup
1071
mg/1
26
207
18
10
230
321
5
4
5
1
Circulating
Water at
3.8 Cycles
n»g/i
245*
790
68
38
874
1220
19
15
5
*See text.
425
-------
The following calculation gives the suspended solids in the circulating
water. Let x be the concentration in mg/1. Then the material balance is
-—• x 281 x io3 Ib/hr blowdown = -~ x 1071 x io3 Ib/hr in makeup
io6 io6
+ 0.052 Ib dust from air x ?9Q x 1Q3 ^ evaporated/hr
10 Ib water evaporated
so x = 245. The circulating water is below all the limits given on Table
14-15. The treatment cost is chemicals only and is given on Table 19-4.
Now consider the cooling tower treatment for Scheme 2, Figure 19-2.
First we find what limits the concentration of acidified source water (analy-
sis given on Table 19-3) using the limits on Table 14-15. The limits are:
silica (mg/1): 150/7 - 21.4 times
Mg x si02 (mg/1): /(60,000/19 x 7) = 21.2 times
Ca x S04 (as CaCC>3) : V[2.5 x 106/(49 x 2.5) (312 x 1.04)] - 7.9 times
TABLE 19-4. COST OF COOLING WATER CHEMICALS FOR HYGAS
IN NORTH DAKOTA, SCHEME 1, FIGURE 19-1
$/hr
"
Biocide: Assume $1/10 gallons of blowdown because of
high concentration of phosphate and carbon:
($1/103 gallons) x (i gallon/8.33 Ib) x (28ixi03lb/hr) 33.7
Anticorrosion chemicals: at $0.5/10 gallons blowdown 16.9
Suspending agent: Assumed not required
Total: 54.9 $/hr
426
-------
The concentration is limited by calcium sulfate. To restrict blowdown
to conserve chemicals, a cold lime softening will be used on the source water.
As described in Section 15.6, lime is added to raise the pH to 9.5 to 10.0.
This reduces Ca to about 12 mg/1, HCO. to a low level and Mg to about 4 mg/1.
Mg in excess of that equivalent to HCO is not removed. Some silica is
removed, but need not be estimated because the water is now limited to 21
times concentration by Ca x so.. This is the system shown on Figure 19-2.
The water analyses are given on Table 19-5.
Note that for Scheme 2 it was necessary to neutralize excess chloride in
biotreatment with soda ash to limit the concentration of calcium (see note at
the bottom of Table 19-2). For Scheme 1, lime could be used. Lime costs
about $l/lb equivalent and soda ash costs about $2.12/lb equivalent, so lime
is preferred where permitted. The total cost for neutralizing the biotreat-
ment depends on the analysis which is not known but will not exceed $9.50/hr
for lime or $20.10/hr for soda ash, neither of which is large. The differ-
ence in cost is much less than the difference between Schemes 1 and 2.
The side stream rate is somewhat more than that needed to remove dust
taken in from the atmosphere - 0.052 x 790 = 41 Ib/hr. If the side stream
filter reduces the concentration from 300 mg/1 to 100 mg/1, the side stream
rate is
(41 Ib dust/hr) x (1Q6/200 Ib water/lb dust) = 205,000 Ib/hr
» 410 gal/min
The costs are shown on Table 19-6. Scheme 2 is apparently cheaper than
Scheme 1. Although approximate costs are used it does seem that using cool-
ing tower blowdown for ash disposal, etc., may not be correct because this
procedure wastes chemicals which are expensive because they must be nontoxic.
For boiler feed water treatment the two schemes shown on Figure 19-3 are
considered. In Scheme 1, Ca++, Hg++, HCO " and SiO are removed in a hot
• •* .J..J. *
lime treatment. This is followed by Na -Ca exchange on a softening resin.
The silica and hardness are now low enough, but sodium has been increased.
Demineralization is carried out in a strong acid exchange followed by a weak
427
-------
TABLE 19-5. ANALYSES OF COOLING TOWER STREAMS FOR HYGAS
Flow
(103 Ib/hr)
SS
Ca
Mg
HC03
SOit
Cl
Si02
POi»
NH3
IN
Source Water
After Lime
Softening
790
mg/l
12
4
~ 0
170
9
7
0
0
NORTH DAKOTA,
Biotreatment
Effluent
40
mg/l
20
15
40
0
1200
0
15
20
SCHEME 2, FIGURE 19-2
Mixed
Makeup
830
mg/l
12
4.5
2
162
66
7
1
1
Circulating Water
at 21 Cycles
—
mg/l
300*
260
95
42
3400
1386
147
21
1
*Limited by side stream clarification.
428
-------
TABLE 19-6. COST OF COOLING WATER TREATMENT FOR HYGAS
IN NORTH DAKOTA, SCHEME 2, FIGURE 19-2
$/hr
3 —'
Lime clarifier for 790 x 10 lb/hr: 1,580 gpm
Install two clarifiers at 65% capacity
costing $520,000; at 17%/yr for 8000 hrs/yr 11.l
Side stream filter for 500 gpm at $150/gpm 1.6
Chemicals:
Lime 2.5
Biocide at $0.50/10 gallons of blowdown 2.4
Anticorrosion chemicals 2.4
Flocculation chemicals 1.6
Dispersant chemicals 0.7 9.6
Total: 22.3
base exchange followed by a mixed bed. In Scheme 2 the water is first
H"+ "H"
treated with a weak acid ion exchange resin to remove Ca , Mg and some
Na , and to replace these ions with H . The hydrogen ions release CC>2 from
HCO , which is removed in the degasifier. Additional removal of Ca++ is
required on a softening resin. Demineralization requires a strong acid resin
followed by a weak base resin. A strong base resin is used for removal of
SiO and a mixed bed system is used for final polishing.
Scheme 2 makes more efficient use of regeneration chemicals and is the
cheaper. A return condensate polishing system will be required, but this is
inexpensive. Compositions of Scheme 2 are shown on Table 19-7.
The treatment shown as Scheme 2 of Figure 19-3 costs about $3.2 x 10
(quotation from Permutit) or $64/hr including resin replacement. Regenera-
tion requires 314 Ib H SO /hr, 148 Ib NaOH/hr and 94 Ib NaCl/hr and costs
$26/hr.
The total system cost and energy requirements as entered on Table 19-1
may now be determined for Scheme 2.
Ammonia separation from Figure 16-4 for 0.847 x 10 gallon/day costs
429
-------
TABLE 19-7. BOILER FEED WATER COMPOSITIONS IN NORTH DAKOTA
Ca + Mg
Na
Cl
HC03
S°4
SiOn
2
After
Weak Acid IX
and Degasifier
rag/1
20
63
9
0
170
7
After
Softening
mg/1
0
91
9
0
170
7
After Strong-Acid
Weak-Base
Demineralization
mg/1
0
4
0
0
2
7
After
Strong-Base IX
Silica Removal
mg/1
0
4
0
0
0
< 1
S4.35/thousand gallons = $154/hr. About $78/hr is recovered from sale of
ammonia. At 1.7 x 10 Btu/thousand gallons the energy required is
60 x io6 Btu/hr.
Biotreatment costs $0.025/lb BOD removed and requires (1900 Btu +
0.18 kw-hr)/lb BOD removed. 13,000 mg BOD/1 are removed from
283 x 10 Ib water/hr, so the cost is $89/hr, the steam required is
7 x 10 Btu/hr and the electrical consumption is 640 kw.
Potable water treatment is probably simple chlorination at an insignifi-
cant cost.
Total cooling water treatment, including extra for soda ash in biotreat-
ment, costs $32/hr and the energy requirement is small.
Total boiler feed water treatment cost is $90/hr and the energy require-
ment is small.
These numbers are summed up below. _
$/hr 10 Btu/hr kw
Boiler feed treatment 90
Cooling water treatment 32
Ammonia separation 76 60
Biotreatment* 89 7 640
*In all plants the cost of biotreatxnent includes sludge dewatering.
430
-------
Electric Generation
Because the Lurgi process accepts wet coal, the foul condensate rate in
this plant is nearly four times the rate in the Hygas plant. We have, how-
ever, continued to assume a chloride concentration of 1200 mg/1 in the foul
water, the same concentration as is assumed at the Hygas and SRC plants.
Because these plants consume comparable amounts of coal and coal is the
source of chloride, it should be the mass flow of chloride in the foul con-
densate and not the concentration which is constant from plant to plant. In
taking a constant concentration it can be said that the coal composition has
been changed. The chloride passes through the ammonia separation and bio-
treatment untouched and there are important consequences which make the
choice of 1200 mg/1 chloride an interesting example to study. If the blow-
down from the cooling tower is disposed of with ash and for dust control,
then 228 x 10 ib water/hr are required. Since the evaporation rate is
2470 x 10 Ib/hr, 11.8 cycles of concentration are necessary. This results
in an excess chloride concentration in the circulating cooling water. Other
ways to use the biotreated effluent for solids disposal and cooling were also
considered and it was found the removal of chloride is necessary no matter
what scheme is envisaged.
Reverse osmosis has been chosen as the desalting procedure because there
is a good probability of simultaneously removing some organic molecules. Pre-
filtration, with addition of flocculating agents, will be necessary. Reported
experience on biotreated municipal waste suggests that membrane fouling can
be coped with. The costs assume a two-year membrane life. Precipitation of
calcium phosphates can be prevented by addition of a sequestering agent; the
costs include a charge for chemicals. As shown on Table 19-2, in the water
leaving the process the chloride is mostly in the form of NH Cl. In bio-
treatment ammonia is converted and lime will have been added to prevent the
solution from going acid. Therefore, the most important material in the bio-
effluent is CaCl,,. The total dissolved solids is about 2000 mg/1 going into
reverese osmosis. Calcium rejection is better than 99 percent by most mem-
branes and electroneutrality encourages the rejection of chloride associated
with calcium. An overall rejection of 97 percent of the TDS is quite likely.
431
-------
Residual ammonia is an exception; ammonia passes most membranes.
Taking a 95 percent recovery of water and a 97 percent point rejection,
the exit concentration in the concentrate stream is
0 97
(feed concentration) x 20 = 36,600 mg/1 TDS
By material balance the treated water should contain about 180 mg/1 TDS. A
probable composition of the treated water is
mg/1
s.s.
Ca
Cl
Si
NH,
J
Others
0
50
86
< 1
20
< 1
This is very good quality water.
If effluent from reverse osmosis is mixed with acidified lake water
(acidified to remove the high bicarbonate level) and the blend is fed to the
cooling tower, the desired concentration cannot be achived because of the
limit on calcium sulfate. Lime softening of lake water is required, as it
was in Scheme 2 for Hygas. Because side stream clarification will be
required for dust removal, the same clarifier can be used for softening. All
this leads to the first practical scheme shown on Figure 19-4.
The boiler feed water treatment, not detailed on Figure 19-4, is the
same as for Hygas (i.e., Scheme 2 of Figure 19-3) except that mixed bed
demineralization is not required. Effluent from the reverse osmosis may
be better than lake water as a source for boiler feed, so Figure 19-5 is
an obvious alternative. The boiler feed treatment in Figure 19-5 will have
to include some scheme for removal of residual dissolved carbon, probably
carbon adsorption.
The choice between Scheme 1 and Scheme 2 will be made by knowing which
of the treated waters behaves better in the boiler, and without experimental
432
-------
2837
BOILER FEED \
TREATMENT \
SCHEME 2 OF /
FIGURE 19-3 J
CONCENTRATED
SOLUTION
DGE
1825
Jl
u
LIME SOFTENING
AND CLARIFICATION
1
, * i -
1
ASH DISPOSAL
DUST CONTROL
UNITS «l03lb./hr.
Figure 19-4. Water treatment plant block diagram for
electric generation at North Dakota (Scheme 1).
433
-------
2837
• >
70
2698
BOILER mD\
TREATMENT \
943
T. SCHEME 2 OF /
\FIGURE 19-3 /
873.
V9. !
I '
I SOLUBLE
{ WASTE
>OTABLE
VATER
TREATMENT
— i rxm-ti* i ' • •
\ . J CIARIFIEP
FOUL
WATER
19 /SERVICE a\
— — H SANITARY I
N"0—^/
CIO
REVERSE ylv
OSMOSIS
AMMOHd* AMMONIA
AMMONIA-* SEPARATION
II
FILTER " B(OIKL>
917
^TMENT
CONCENTRATED
SOLUI'°N
(COOLING\
TOWER J
LIME SOFTENING
AND CLARIFICATION
228
VASH DISPOSAL ^N
^DUST CONTROL^
SLUDGE
UNITS-!03lb.Ar.
Figure 19-5. Water treatment plant block diagram for
electric generation at North Dakota (Scheme 2).
434
-------
results on reverse osmosis this is not known. Any difference in cost between
the schemes cannot be found at our level of accuracy.
The total system cost and energy requirement, as entered on Table 19-1,
has been determined for Scheme 1 as follows.
Ammonia separation from Figure 16-^4 at 2.8 x 10 gallons/day costs
$4/10 gallons = $473/hr. About $260/hr are recovered from sale of ammonia.
At 1.7 x 10 Btu/thousand gallons the energy required is 198 x 10 Btu/hr.
Biotreatment costs $0.021/lb BOD removed. For 917 x IQ lb/hr with
18,000 mg/1 BOD removed, the cost equals $346/hr. Per pound of BOD removed
the energy requirements are 1,900 Btu plus 0.18 kw-hr, so the energy required
is 31.4 x 106 Btu/hr plus 2,970 kw.
Potable water treatment is probably simple chlorination at an insignifi-
cant cost.
Filtration, with a capital cost of $75/gpm, a flow of 1.8 x 10 gal/min,
a charge of 17%/yr for amortization and maintenance plus flocculating chemi-
cals at 3$/thousand gallons costs $6/hr. The energy requirements are small.
Reverse osmosis for 2.6 * 106 gallons/day costs $132/h*. The energy
required is 7.2 kw-hr/thousand gallons, or 780 kw.
The cooling tower side stream clarifier must handle 1,210 x 10 lb/hr to
maintain bicarbonate at an acceptable level. Lime is added only to precipi-
tate bicarbonate. The cost of the clarifier plus lime is $15/hr. The energy
requirement is small.
Cooling tower chemical charges are biocide at $0.50/thousand gallons of
blowdown, corrosion-resistant chemicals at $0.50/thousand gallons of blowdown
(this involves the assumption that side stream treatment with lime does not
remove the anticorrosion chemicals; should this not be true then two separate
clarifiers should be used: one in the makeup stream using lime and one in
the sidestream for dust removal without lime addition), and dispersant chemi-
cals at $0.03/thousand gallons of side stream. The total cost for cooling
water chemicals is $31/hr.
The boiler feed treatment costs about $3.2 x io6 or $92/hr for capital,
replacement resin and regenerating chemicals. The energy requirement is
small. The numbers are summed up below.
435
-------
$/hr 106 Btu/hr kw
Boiler feed treatment 94
Cooling water treatment 46
Filtration and reverse osmosis 138 — 780
Ammonia separation 213 198 —
Biotreatment 346 31 2970
Alternative for Lurgi Technology
The scheme chosen in this study is not the scheme usually presented in
the literature. It is usually stated that for Lurgi plants the foul water,
after solvent extraction and ammonia distillation, can be fed to the cooling
towers without further treatment (see, for example, Reference 1). From Table
17-1, assuming a 95 percent recovery of phenol with phenol at 6,000 mg/1 in
the feed, the cost of solvent extraction is $363/hr and the energy required
is 110 x 10 Btu/hr. About 5,700 Ib phenol/hr are recovered, and if this can
be sold at 2.2C/lb, $125/hr is recovered.
Because of the high chloride, content it must be assumed that solvent
extraction and ammonia separation, without biotreatment, make the water fit
for feed to reverse osmosis. If this is true then biotreatment could be eli-
minated saving $330/hr and 66 x 10 Btu/hr (including the steam needed to
generate the electricity required for biotreatment).
The alternative route apparently costs 72 percent of the route of Scheme
1 and requires 167 percent of the energy. This difference is probably not
due to the low accuracy of cost estimating, and it is worthwhile considering
solvent extraction and dropping biotreatment. The technological feasibility
of this change requires experimental verification.
Residual Disposal
The residue from reverse osmosis contains 3.6 percent soluble salt
(mostly CaCl?). The disposal of this residue has not been studied and no
added cost has been assumed. In New Mexico, where the source water is
436
-------
brackish and the major positive ion is Na , segregation and separate disposal
of soluble waste has been assumed required for all plants and lined evapora-
tion ponds have been costed.
Solvent Refined Coal
The SRC process is a net producer of water, but it takes in water of a
quality suitable for low pressure boilers and puts out four streams:
323 x 10 lb/hr foul condensate, 320 x 10 Ib/hr water (after flashing) from
gas purification having the approximate composition of Table 14-13,
177 x 10 lb/hr condensate from the Koppers-Totzek gasifiers having the
approximate composition of Table 14-12, and 175 x 10 lb/hr clean water
(after flashing) from the reformers which can be returned to the boilers
with other condensate. In addition the plant takes in raw water of the
composition given on Table 14-1. Various ways have been considered to
combine these streams for use in the cooling tower and in the boiler. The
most satisfactory and flexible is to blend the two intermediate quality efflu-
ent streams with raw water and treat the blend to approaching boiler feed
quality using the same quality for cooling tower makeup. This is shown in
Figure 19-6. This procedure results in a low cooling tower blowdown with
savings in chemicals and a large and economical lime clarifier.
The approximate blended water composition is shown on Table 19-8 with
the approximate composition of the water leaving the lime-soda treatment.
Lime, soda ash and additional magnesium as dolomite are needed for the chemi-
cal treatment. The approximate rates are lime 290 lb/hr, soda ash 220 lb/hr
and dolomite 550 lb/hr. Chemical costs are given on Table 15-3 and result
in a total cost of $28/hr. The clarifiers will cost about $560,000 or
$12/hr.
The cooling tower will require a large side stream filter to hold sus-
pended solids to 300 mg/1 because all the makeup (from all sources) is over-
flow from clarifiers and not clear. A side stream of about 1,000 gal/min is
probably required, so the filters cost about $5/hr including flocculating
chemicals. The cooling tower is limited to 20 cycles of concentration by
chloride (because of the high chloride content of the foul condensate) or
437
-------
Co
CD
320
GAS PURIFICATION
CONDENSATE
1141
WEAK-
BASE
IX
944
177
GASIFICATION
CONDENSATE
SRC PROCESS
323
CONDENSATE
POLISHING
175
CLEAN
WATER
. SERVICE a>
*H SANITARY
USES j
80S
880
EVAPORATION
^/ COOLING \<
^^ TOWER j
116
't
44
UNITS-I03 Ib./
CLARIFIED
FOUL
WATER
12
309
0
193
237
ASH DISPOSAL a
DUST CONTROL
D
1
300
SLUDGE
Figure 19-6. Water treatment plant block diagram for SRC at North Dakota.
-------
TABLE 19-8. APPROXIMATE INFLUENT TO AND EFFLUENT FROM
LIME-SODA-SILICA
Ca
Mg
HCO_
3
SO.
4
Cl
SiO0
2
P°4
NH.,
3
COD
TREATMENT IN
Influent
(mg/1)
55
26
125
130
19
~ 10
6
13
7
SRC PLANT AT NORTH DAKOTA
Effluent
(mg/1)
12
2.5
0
130
19
~ 1
0
13
7
25 cycles of concentration by calcium sulfate. The difference is unimportant.
Biocide, anticorrosion chemicals and suspending agents cost about $10/hr.
The boiler treatment costs about $1.8 x 10 or, including resin replace-
ment and chemicals, about $50/hr.
The cost of ammonia separation, taken from Figure 16-4, for
0.93 x 10 gallons/day is $4/6/thousand gallons = $178/hr. There being about
12,000 mg/1 in the foul condensate, about $256/hr is recovered from the sale
of ammonia. The energy required is 66 x 10 Btu/hr.
Biotreatment costs $0.021/lb BOD removed and about 30,000 mg/1 BOD are
removed. The cost is $189/hr. The energy requirement is 18 x 10 Btu/hr as
steam plus 1670 kw. These numbers are summed up below.
Boiler feed treatment
Lime soda softener and chemicals
Cooling water filter and chemicals
Ammonia separation
Biotreatment
$/hr 10 Btu/hr kw
50
40
15
[+78] 66
189 18 1670
439
-------
19.3 PLANTS IN NEW MEXICO
Power Plant
The overriding problem in all the exemplary plants sited in New Mexico
is that the source water is assumed to be brackish and to have the composi-
tion shown on Table 14-2. Desalting procedures are required for all water
uses. For the boiler feed water desalting is needed to save excessive use
and cost of ion exchange demineralizers. Desalting is needed to make the
water fit for revegetation and drinking use. In the cooling towers the deci-
sion has been made that the salt in the source water may not be disposed of
with the ash but must be segregated and separately disposed of in a lined
evaporation pond. This means either that the makeup to the cooling tower
must be desalted or that the blowdown must be segregated.
If raw source water is used as makeup to the cooling tower mixed with
effluent from biotreatment, it must be acidified to remove bicarbonate. The
next limit is excessive chloride in the circulating water which occurs at
4 cycles of concentration. A side stream must be taken from the tower and
treated, first to remove suspended solids which interfere with the desalting
procedure and then to concentrate and reject the dissolved solids. The reject
will be saturated in calcium sulfate which is as concentrated as the reject
can be without precipitation occurring. Precipitation is not permissible in
either electrodialysis or reverse osmosis desalting procedures. This possible
scheme is shown in Figure 19-7. A side stream electrodialysis system as shown
is a device essentially to concentrate salt for removal, not a device to pro-
duce desalted water. The concentrations in and out are not the same and the
cost figures given in Section 15 do not necessarily apply. Electrodialysis
for salt concentration is available from Dow Chemical, who made a cost esti-
mate for this report. The overall plant cost for the systems shown on Figure
19-7, with side stream desalting, and Figure 19-8, with makeup desalting, are
close to the same. We are investigating side stream desalting further but
for this study, because we are unsure of the basis of the cost estimate, the
scheme of Figure 19-8 is the one used.
For simple removal of ionic species at the throughput of this plant,
440
-------
UNTREATED BRACKISH
SOURCE WATER
2454
346
dM
EVAPORATION
2476
TOWER
| CHEMICALS [
D
j
2952
ELECTRODIALYSIS
324
BLOWDOWN TO
EVAPORATION POND
BIOTREATMENT
242
ASH DISPOSAL
a
DUST CONTROL
UNITS =IOJ Ib.Ar.
Figure 19-7. Use of untreated brackish makeup to cooling tower of
power plant at New Mexico.
441
-------
BRACKISH
SOURCE
WATER
3X17
1077
WASTE TO
EVAPORATION
POND
2340
STRONG-ACID
IX
1
WEAK-BASE
IX
1
STRONG-BASE
IX
1
r
ELECTRODIALYSIS
J
WASTE TO
EVAPORATION POND
I"7
CONCENTRATE TO
EVAPORATION POND
POTABLE WATER
TREATMENT
45
(REVEGETATION
EVAPORATION
2476
2155
( COOLING V
TOWER Jj
421
— . J (
II TPD
1950
CHEMICALS
100
SERVICE a
SANITARY USE
13
BIOTREATMENT k
SLUDGE
167
267
ASH DISPOSAL \
a
DUST CONTROL
J
970
PROCESS
)
581
625
AMMONIA
SEPARATION
AMMONIA
UNITS = 10 Ib./V-
Figure 19-8. Water treatment plant block diagram for
electric generation at New Mexico.
442
-------
electrodialysis uses more energy than reverse osmosis, but the total operat-
ing cost is lower. Electrodialysis does not remove any non-ionic species,
while reverse osmosis does partly remove such molecules. In particular,
reverse osmosis partly removes silica. As it turns out, silica is not a
limitation in the operation of the cooling tower and electrodialysis has
been selected for desalting.
A three-stage electrodialysis system will reduce the total dissolved
ionic content to about l/8th (see Section 15.8 for costs). Because the elec-
trodialysis plant will separately handle the boiler feed water and because so
much of the capital cost of electrodialysis is in rectifying equipment, econ-
omy of scale is realized and the capital cost is based on a throughput of
3,400 x 10 Ib/hr = 10 gallons/day. From Figure 15-5 the capital cost is
SO.32/(gallons/day). For that part of the flow used for cooling water makeup
and revegetation, = 2,200 x io3 Ib/hr = 6.3xio gallons/day, the cost and
energy requirement of electrodialysis can now be estimated. Capital charges
at 17%/yr for 8,000 hrs/yr of 0.32 x 6.3 x 10 dollars investment = $42.8/hr.
Membranes, filters and chemicals at $0.2/thousand gallons cost $52.5/hr.
Electricity at 0.4 kw-hr/(thousand gallons) (100 mg/1 removed) and $0.02Awhr
costs $48.9/hr. Acid costs $2.9/hr. The total cost is $147/hr and the total
energy is 2560 kw.
The cooling tower requires a side stream filter to remove suspended
solids from the air and from the effluent of the biotreatment. The filter
will handle about 1,950 x 10 Ib/hr of side stream flow and costs about $12/hr
including flocculating chemicals. Biocide and anticorrosion chemicals, con-
sidering the large amount of effluent from biotreatment used as cooling tower
makeup, cost about $18/hr.
The boiler feed water is first treated with a weak acid ion exchange
resin to remove Ca++, Mg+ and some Na and replace these ions with H . The
hydrogen ion combines with HC03" and C03~ to yield CC>2 which is released in a
degasifier. H SO is used to regenerate the resin at 110 percent of stoichio-
metric to the HCO ~ + CO ~. Removal of the alkalinity is necessary to pre-
vent precipitation in the electrodialysis equipment. However, sulfate must
also be removed downstream so weak acid exchange is used in preference to
direct acid addition.
443
-------
The water then passes through a three-stage electrodialysis unit to
reduce the total dissolved solids to about l/8th through a strong acid ion
exchange and a weak base ion exchange to further remove Na and SO , and
through a strong base ion exchange to remove silica. The composition of the
boiler feed at various points in the treatment is shown on Table 19-9.
The cost of operating the electrodialysis unit for
1340 x 10 Ib/hr = 3.8 x 10 gallons/day is $89/hr and the electric usage
is 1200 kw.
The ion exchange system cost is 3.4 x 10 dollars (quotation from Permu-
tit) or $69/hr, including resin replacement. Regeneration chemical require-
ments are: H SO 660 Ib/hr and NaOH, 200 Ib/hr for a cost of $43/hr. The
total ion exchange system cost is $112/hr.
A return condensate polishing system will be required; the cost is small.
The foul condensate treatment was costed using the procedures described
in Sections 16 and 18 and used in Section 19.2.
The electrodialysis concentrate and the ion exchange regeneration waste
amount to about 275 x 10 Ib/hr and are sent to a lined evaporation pond. At
this site the net annual average evaporation rate after rainfall is
2
53 inch/yr = 33 gallons/(ft )(yr). For sizing the pond an allowance of
TABLE 19-9. COMPOSITIONS AT POINTS IN THE BOILER
Ca
Mg+4
Na+
co3=
HCO.
S°4~
Cl
SiO,
FEED WATER TREATMENT IN NEW MEXICO
Entering
Electrodialysis
(mg/1)
0
0
743
0
0
509
770
6
Leaving
Electrodialysis
(mg/1)
0
0
93
0
0
63
93
6
After Weak
Base Resin
(mg/D
0
0
3
0
0
1
1
6
After Strong
Base Resin
(mg/1)
0
0
3
0
0
1
1
< 1
444
-------
20 percent was made and 26 gallons/(ft2)(yr) = 0.027 lb/(ft2)(hr) was used.
The pond is 10.2 x 10 ft (234 acres) which, at 20C/ft2, costs $2.04 x io6.
This is amortized at 15%/yr or $38/hr.
The total plant costs and energy requirements are:
$/hr IO6 Btu/hr Jew
Ion Exchange 112
Electrodialysis 236 — 3760
Cooling tower filter and chemicals 30
Ammonia Separation 139* 128
Biotreatment 219 21 1890
Evaporation Pond 38
*After recovering $165/hr from the sale of ammonia.
Hygas
The water treatment block diagram is quite similar to that for the power
plant and is shown on Figure 19-9. The cycles of concentration in the cool-
ing tower are low enough that no side stream treatment is required. In this
plantf where the biotreated effluent is used only for disposal with solids,
as high a level of treatment as high purity oxygen may not be necessary and
less expensive alternatives should be considered. The boiler makeup must
meet high pressure specifications, and the block marked "demineralizer" is
like the blocks shown on Figure 19-8 but contains a mixed bed ion exchange
as well.
The cost and energy requirements are:
445
-------
BRACKISH
SOURCE
WATER
1490
WASTE TO T
EVAPORATION
POND
69
491
-»»| DEM1NERALIZER
.__
->•
->•
ELECTRODIALYSIS
WASTE TO
EVAPORATION
POND
^CONCENTRATE TO
EVAPORATION POND
POTABLE WATER
TREATMENT
EVAPORATION
389
CLARIFIED
FOUL
WATER
4
(
COOLING
TOWER
CHEMICALS
SLOWDOWN
FLUE GAS DESULFURIZATION
ASH DISPOSAL a DUST CONTROL
\
J
294
AMMONIA
UNITS -I03lb./hr.
Figure 19-9. Water treatment plant block diagram for Hygas process at New Mexico.
-------
$/hr 106 Btu/hr jw
Ion Exchange 94 — —
Electrodialysis 103 — 1650
Ammonia separation 76* 60
Biotreatment 89 7 640
Cooling water chemicals 19 — —
Evaporation Pond 23 —
*After recovering $78/hr from the sale of ammonia.
SRC
This water treatment block diagram is also similar to that for the power
plant and is shown on Figure 19-10. The cycles of concentration in the cool-
ing tower are about 7, low enough that only a dust filter is required in the
side stream to remove the suspended solids. The cost and energy requirements
are:
$/hr 10 Btu/hr kw
Ion Exchange 40
Electrodialysis 45 — 740
Ammonia separation (49 INCOME)* 44
Biotreatment 126 12 1080
Cooling water filter and chemicals 11
Evaporation Pond 9
*After recovering $170/hr from the sale of ammonia.
19.4 PLANTS IN WYOMING
Electric Generation
At all plants in Wyoming, sewage from a satellite town is assumed to be
447
-------
09
BRACKISH
SOURCE
WATER
663
3BO
GAS PURIFICATION CONDENSATE
GASIFICATION
CONDENSATE
ION EXCHANGE
WASTE TO1
EVAPORATION
POND
248
DEMtNERAUZER
ELECTRODIALYSIS
WASTE TO
EVAPORATION
POND
CONCENTRATE TO
EVAPORATION POND
S3
POTABLE WATER
TREATMENT
l
23 ^/SERVICE ANDA
^\ SANITARY USE )
60 > -T- '
EVAPORATION
»EVEGETAT.ON
13
186
AL5
>(co
v_T^
*^2»7
OLINGl
3WER J
A I M»
FILTE
^n
« -*l
BLOWDOWN ^_
^
VII | «- 200
BIOIktAIMtNl < SEPARATION
j, {
SLUDGE AMMONIA
r
CkSH
?US
DISPOSAL
AND
UNITS -103 lb.A".
Figure 19-10. Water treatment plant block diagram for SRC at New Mexico,
-------
taken into the plant and treated in a small package-type treatment facility,
the effluent from which has the composition of Table 14-3. Additional water,
having the composition shown on Table 14-4 which is the best quality of any
of the sites, is available but is very expensive. The power plant at this
site uses partial dry cooling for the turbine condensers, but the cooling
water requirement is still so large that fresh water is needed. The boiler
is fed with treated fresh water; the treatment is no different from that
described for plants in North Dakota, shown as Scheme 2 of Figure 19-3.
The scheme for this plant is shown in Figure 19-11. The blended cooling
water has the approximate composition of Table 19-10. Because the evapora-
tion rate is lower than at other power plants, the cooling tower is required
to work only at about eight cycles of concentration. At eight cycles chlor-
ide will reach 3,200 mg/1, which is just tolerable. The other limitations are
phosphate and bicarbonate, which must be removed. Also, the makeup has high
suspended solids. A combined lime treatment clarifier is required. This
treatment can be placed in a side stream, in which case a side stream rate of
about 750 x 10 Ib/hr suffices to hold phosphate below 24 mg/1, bicarbonate
below 188 rag/1 and suspended solids below 300 mg/1. Alternatively, the two
biotreatment effluent waters can be lime treated and clarified before going
to the tower makeup. This involves a flow rate of about 930 x 10 Ib/hr and
keeps phosphate and bicarbonate very low. Dust drawn in from the air may
raise the suspended solids concentration in the circulating water undesirably
TABLE 19-10. COOLING WATER MAKEUP TO THE ELECTRIC
POWER
Ca
Mg
HCO,
3
SO.
4
Cl
Si°2
P°4
PLANT IN
mg/1
34
18
133
50
396
12
17
WYOMING
mg/1
SS 100
COD > 200
. NH, 20
3
449
-------
TREATED
IIB4
1007 BOILER FEED 957 . v 720
W TPFATMFKIT >fBBJ--rit\ fc_
SCHEME 2 OF V Kt-"-L*i J ^
FIGURE 19-3 •—* CLARIFIED
r— FOUL WATER
I SOLUBLE
^ WASTE
19 POTABLE l7 /SERVICE AND\
TREATMENT \^ USES J
158 "
SEWAGE 299 W ' k
%^ FVAPOBATIOKJ
989
1 r ^
jf ' -v A7^ A.ln 1 '
c
'
IffWICALS w,/COOLING^<< niOTnCATMEMT 1 AMMONIA
HEMICALI ^\^ TOWER ^^ BIOTRtATMENT J SEPARATION
4 750 T ^
LIME ^
CLARIFIER ^*
1 1 r
^ 142
ASH DISPOSAL^
( AND )
\DUST CONTROL J
ONIA
UNITS-I03 Ib.
Figiore 19-11. Water treatment plant block diagram for
electric generation at Wyoming.
450
-------
high. Side stream, rather than makeup, treatment seems preferable, but the
cost is independent of the location.
The cooling water makeup has a high nitrogen, phosphorus and carbon con-
tent. This may be the most nutrient water used at any site. The cost of bio-
cides will be high; $1.50/thousand gallons of blowdown has been assumed.
Other treatments are as previously used at other sites. The cost and
energy requirements are:
$/hr 10 Btu/hr kw
Ion exchange for boiler feed 98 —
Cooling tower clarifier
and chemicals 55 —- —
Ammonia separation 160* 147
Biotreatment 253 23 2170
*After recovering $188/hr from sale of ammonia.
Hygas
In this plant the sewage water plus the treated process effluent
total close to the whole cooling water makeup requirement. Fresh water is
used for boiler feed. In operating the cooling system the same choice is
available as for the Hygas plant in North Dakota, namely to operate at low
cycles of concentration, to blowdown onto the ash and to pay heavily for bio-
cide and other chemicals, or to operate at high cycles of concentration, to
treat the makeup and to save wasting chemicals. The latter scheme has been
used as in North Dakota. As found in Reference 2, treated sewage can be con-
centrated by 15 cycles without problem and this is what will be done. The
scheme is shown on Figure 19-12.
The cost and energy requirements are:
451
-------
Ui
to
834
840
TREATED
BOILER FEED
TREATMENT
SCHEME 2
OF FIG. 19-3
45
SOLUBLE
WASTE
294
CLARIFIED
FOUL
WATER
k 19 ^_
POTABLE
WATER
TREATMENT
19
\
f
EVAPORATION
279
229
/FLUE GAS DESULFURIZATION
{ ASH DISPOSAL
V DUST CONTROL
52
UNITS-I03 lb./hr.
AMMONIA
Figure 19-12. Water treatment plant block diagram for
Hygas process at Wyoming.
-------
S/hr 106 Btu/hr kw
Ion exchange for boiler feed 94
Cooling water treatment and chemicals 31
Ammonia separation 77* 60
Biotreatment 89 7
*After recovering $77/hr from the sale of ammonia.
Synthane
The Synthane process takes in more boiler feed water and puts out more
dirty water than the Hygas process. Also, 158 x 10 lb/hr of intermediate
quality water is obtained from the process. The composition of this water
is not known, but it is assumed to have one tenth the concentration of the
foul condensate. If this intermediate quality water were to be treated with
the sewage, it would have to be stripped of ammonia. It is best, therefore,
to treat this water with foul condensate where the ammonia can be consumed
and it will not be necessary to pass it through ammonia separation. Since
foul condensate is stripped of ammonia, this can be done to a slightly lower
concentration with little increase in cost. This saves the expense of strip-
ping the intermediate quality water.
The effluent from the biotreatment of foul condensate is assumed to have
600 mg/1 chloride and, even when diluted with treated sewage, this is too much
chloride for cooling water makeup. The scheme shown in Figure 19-13, which is
similar to the Hygas scheme, will not work. Since desalting is required and
since some effluent from one of the biotreatments will go to boiler feed, we
will use the desalted water for boiler feed; that is, the scheme of Figure
19-14 will be used.
The cooling water and lime softening on Figure 19-14 is the same as was
used in the Hygas plant at Wyoming shown on Figure 19-12. The reverse osmo-
sis is a smaller version of the equipment used on the power plant in North
Dakota. However, in the power plant in North Dakota we did not feed the
effluent from reverse osmosis to a boiler, and here we will. The approximate
453
-------
857
TREATED
SEWAGE
297
01
LIME
SOFTENING
2«6
1083
BOILER FEED
TREATMENT
SCHEME 2
OF FIG. 19-3
SOLUBLE
WASTE
19
CLARIFIED
FOUL WATER
245
565
w320
EVAPORATION
299
COOLING
TOWER
J
FILTER
376
397
642
CLUE GAS DESULFURIZATION
ASH DISPOSAL
OUST CONTROL v
INTERMEDIATE
QUALITY WATER
SERVICE AND
SANITARY uss
A
y
11
158
BIOTREATMENT
480
SLUDGE
AMMONIA
UNITS-
Figure 19-13. Water treatment plant block diagram for Synthane process
at Wyoming (Scheme 1, not used).
-------
BOILER FEED
TREATMENT
SCHEME 2
OF FIG. 19-3
SYNTHANE
PROCESS
CLARIFIED
FOUL WATER
CLEAN
WATER
(40
CARBON
ADSORPTION
SOLUBLE
WASTE
REVERSE
OSMOSIS
INTERMEDIATE
QUALITY WATER
SERVICE AND
SANITARY USE
TREATED
SEWAGE
EVAPORATION
COOLING
TOWER
BIOTREATMENT
AMMONIA
N\
, ASH DISPOSAL J
V DUST CONTROL J
/FLUE GAS DESULFURIZATION
I ACH RK.POS4J
UNITS -105 |b./hr.
Figure 19-14. Water treatment plant block diagram for Synthane process
at Wyoming (Scheme 2, as used).
-------
composition is given in Section 19.2. However, COD or organic matter is not
shown in Section 19.2 and is not known. It has been arbitrarily assumed that
1,000 mg/1 is to be removed by carbon adsorption at a cost of $1.5/thousand
gallons, or $42/hr (see Figure 15-1), and an energy requirement of about
11,000 Btu/lb COD removed, or 2.6 x 1Q6 Btu/hr.
The costs and energy requirements can now be tabulated:
$/hr 10 Btu/hr kw
Ion exchange for boiler feed 118
Cooling water treatment
and chemicals
Ammonia separation
Biotreatment
Reverse osmosis
Carbon adsorption
31
115*
312
46
42
—
105
28
—
3
—
—
2680
210
__
*After recovering $136/hr from sale of ammonia.
Solvent Refined Coal
Being given the sewage from the satellite town, this plant has a very
small net consumption of additional water. However, there are four differ-
ent water streams required and five different water streams available.
Before making economic studies some reasonable schemes must be devised. The
four required water streams, taking approximate account of losses in treat-
ment, are:
Feed to ion exchange for boiler after
returning clean process water
Drinking water
Evaporated in cooling tower
For ash disposal and dust control
662 X 10 Ib/hr
3
19 X 10 Ib/hr
3
229 x 10 Ib/hr
3
178 x 10 Ib/hr
Total: 1088 x 10 Ib/hr
456
-------
The five available water streams, again taking approximate account of losses
in treatment, are:
Fresh water 98 x iQ lb/hr
Water from gas purification 261 x 10 lb/hr
Water from gasifiers 163 x 10 lb/hr
Treated sewage from satellite town 3
plus plant complex 310 x 10 lb/hr
Partially-treated foul condensate 256 x 10 lb/hr
Total: 1088 x 1Q3 lb/hr
The next step is to record which available water streams are good for
which requirements without treatment:
Fresh water and water from gas purification are good for all uses.
Water from gasifiers and treated sewage are good only for ash dis-
posal and dust control.
Foul condensate is good for nothing.
Minimum water treatment will be required if the preliminary and incomplete
scheme of Figure 19-15 is used. Broadly speaking, either sewage or foul con-
densate has too high a chloride content (600 mg/1) to go to a cooling tower
without desalting, whereas sewage can be treated and used as cooling tower
makeup without removal of chloride. For boiler feed desalting is required
anyway, so foul condensate will be used. This leads to the scheme of Figure
19-16, which is based in part on work reported in Reference 3 and described
in Section 15.5.
The cost and energy requirements can now be summed up and then briefly
derived.
457
-------
FRESH WATER 98
WATER FROM GAS -,,
PURIFICATION 261'
FOUL CONDENSATE 256
SEWAGE 310
WATER FROM 163 •
GASIFIERS
79
332
15
I7B
/ "N
>\ DRINKING WATER )
SUPPLY TO BOILER
TREATMENT
^
/
COOLING TOWER
MAKE-UP ASSUMING
15 CYCLES CONCENTRATION
BIOWDOWN
)
Dl SPOSAL AND
DUST CONTROL
J
Figure 19-15. Preliminary water treatment plant block diagram
for Solvent Refined Coal at Wyoming.
458
-------
261
01
FROM GAS PURIFICATION
F-RBH —....—.
WATER 98 79 WEAK STRONG WEAK MIXED 629 715 /"TIT
—*\ * ACID * DEGASIFIER r ACID 1 BA5E — i — » 6ED • r. J SRC
"^ WATER FROM GASIFIERS
ix 1 1 ix ix ' ix " \PROCEsy 163
o, CLEAN
00 WATER
CONDENSATE „
(9 nr*, ^_...._ _ POI I1HIKJR
HU'AW ./ -citvicea N POLLHIMC
CLARIFIED FOUL
CONDENSATE
FROM TREATMENT ^SANITARY USESJ CARBON SOLVfNT PHENOLS
TOWN /"sIwAGTi SEWAGE .1 ADSORPT.ON EXTRACTION
1 PLANT i"" '
t , 1 T r,,rix 310 '
C - IKEAItU AMMONIA A«UON,,A
SEPARATION *
M •
— ^^ — ™--to m^TII 1 ATI AM ^ . .,
LIME SOFTENING
CLARIFIED FVAPnRAT,ON
229 |
244 r/COOLING\
^J^*^TOWER y
j CHEMICALS -*" .A
FILTER *
,,15
178 /ASH DISPOSAL\
^DUST CONTROL^
Figure 19-16. Water treatment plant block diagram for SRC at Wyoming.
-------
$/hr 10 Btu/hr kw
Ion exchange 48
Cooling water treatment and chemicals 28
Solvent extraction 147 33
Ammonia separation (63 INCOME) 56
Distillation 132 35
Carbon adsorption 10 0.3 —
The ion exchange system for demineralizing boiler feed is similar to
others used. The mixed bed, which is assumed to handle the full flow includ-
ing distilled water/ has been added as a protection because the quality of
the distilled water is not well defined. It may not be necessary for the low
pressure boiler of this plant. In any event the chemical requirements are
very small for this bed. The chemical requirements are measured by a flow of
only 340 x 10 Ib/hr and are about $10/hr. The capital cost is about
$2.0 x io6, or $38/hr.
The cooling tower treatment is just a little smaller than in the Hygas
and Synthane plants.
For solvent extraction, approximate costs have been found in Section 17
for 10 Ib/hr flow. If one scales down using an arbitrary 0.7 power, a cost
of $182/hr is found. If the feed concentration is 6,000 ppm phenol, about
1,600 Ib phenol/hr are recovered. Assuming a sale price of 2.2$/lb, this is
$35/hr recovered. The energy is taken as 10 Btu/thousand gallons.
Ammonia separation is as used before. Since the feed has 12,000 mg/1
ammonia, $218/hr are recovered from the sale of ammonia.
Distillation requirements are shown on Figure 15-2. For
0.8 * 10 gallons/day these are $132/hr and 3.5 x 10 Btu/hr.
Carbon adsorption costs, from Figure 15-1, are. based on an assumed COD
removal of 50 mg/1 as discussed in Section 15.5
460
-------
REFERENCES SECTION 19
1. Milios, P., "Water Reuse at a Coal Gasification Plant," Chemical
Engineering Progress 71 (6), 99-104, June 1975.
2. Kleusner, J., Heist, J. and Van Note, R. H., "A Demonstration of Waste-
water Treatment for Reuse in Cooling Towers at Fifteen Cycles of Con-
centration," presented at AIChE Water Reuse Conference, May 1975.
3. Skrylov, V. and Stenzel, R. A., "Reuse of Wastewaters—Possibilities and
Problems," presented at the Workshop on Industrial Process Design for
Pollution Control, AIChE, New Orleans, October 1974.
461
-------
APPENDIX 1
ANALYSES OF WASTEWATER SAMPLES
Although analysis was not a major part of this study, we were
fortunate to have received three sets of samples from the Institute of Gas
Technology's Hygas pilot plant (two derived from Montana lignite and one
derived from Illinois coal) and two sets of samples from Pittsburgh and
Midway's Solvent Refined Coal pilot plant. Our analytical procedures
follow*. The analytical results are given in Section 14. As is emphasized
there, further analyses are essential for an economic design of a water
treatment plant. When making our analyses we ran into some troubles and
found some still unresolved inconsistencies which we have recorded here in
hope of giving some assistance to future analytical programs.
Our samples were shipped to Cambridge, Massachusetts for analysis.
One sample, shipped unpreserved from Fort Lewis, Washington, contained 0.5%
phenol when shipped and only 0.0065% phenol when we analyzed it. Even with
waters as foul as these, preservation of a special sample for phenol analy-
sis (using copper surfate and phosphoric acidl) is necessary.
We had some trouble with total organic carbon (TOO and chemical
oxygen demand (COD) analyses. First we found that two different samples
(both from Montana lignite), shipped unpreserved from the Hygas plant, gave
Sample 1 Sample 2
COD (mg/1) 13,586 29,993
TOC (mg/1) 3,936 5,270
COD/TOC (calculated) 3.5 5.7
Sulfide not analyzed none detected
We have trouble justifying the high ratio COD/TOC in Sample 2. Either the
TOC measurement is too low, which can occur if volatile carbonaceous mole-
cules such as benzene are lost during sparging to remove inorganic carbon,
or the COD is too high. If COD is too high, we are not sure what non-
carbonaceous reducing compounds are present. An unpreserved sample from
the Solvent Refined Coal plant gave
*Analyses were done by Dr. Judith Clausen.
462
-------
TOC (mg/1) 6,600
COD (mg/1) 43,655
COD/TOC 6.6
Sulfur 10,500
but reduced forms of sulfur could account for the high COD.
When a sample is preserved for shipment by reducing the pH to less
than 2 with sulfuric acid2, then an extra sample must be separately pre-
served and shipped for sulfide analysis2 and inorganic carbon and alkalinity
cannot be determined. We received a preserved sample from the Solvent
Refined Coal plant on which phenol, sulfur and nitrogen analyses had been
performed, at the plant, before acidifying for shipment. We found
COD, mg/1 as received 31,718
COD, mg/1 sulfide removed 25,200
TOC, mg/1 7,390
Sulfide was removed by pH adjustment followed by addition of lead carbonate.
Since the sample before preservation contained 16,200 mg/1 sulfur, clearly
some of the sulfide survived acid preservation. Sulfide removal reduced the
COD/TOC ratio from 4.3 to 3.4.
Being disturbed by possibly high COD measurements, we tried to find
whether the unusually large quantities of ammonia present could affect the
COD analysis. We prepared a solution of 2.5510 grams potassium acid
phthalate in 100 ml water (theoretical COD = 30,000 mg/1) and measured a
COD of 29,676. To the same solution we added 5.00 grams ammonium carbonate
and measured a COD of 29,472. However, we then took the ammoniacal solution,
made it alkaline and bubbled nitrogen through to remove ammonia. After two
hours bubbling the COD measured 33,263 and after three hours of bubbling
the COD measured >38,860. We have no explanation for results obtained when
we tried removing ammonia.
Finally, samples were shipped from the Hygas plant (from Illinois coal),
one unpreserved and one split into three with part preserved with H2S04
(for COD, TOC and nitrogen), part preserved for phenol and part preserved
for sulfide. We found
463
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Nitrites - N
Nitrates - N
Ammonia - N
Sulfide
Sulfate
Phenol
TOC
Total Carbon
BOD - 5 days
BOD - 20 days
COD - as received
COD - ammonia removed
COD - ammonia & sulfide removed
Preserved
Samples
(mg/1)
None detected
None detected
7,200
540.0
355.0
273.4
702
1,225
2,999
Unpreserved
Samples
(mg/1)
None detected
None detected
7,200
490.0
340.6
265.4
3,573
4,722
2,682
3,016
3,445
3,025
2,686
Nonpreservation of the sample caused no loss of ammonia and some loss of
sulfide and phenol. We cannot explain the effect of acid preservation on
TOC and total carbon; the COD/TOC ratio in the preserved sample, being 4.3,
is possible, but the ratio of 1 in the unpreserved sample seems quite
impossible.
In closing, we point out that all of our BOD measurements were made
on samples shipped unrefrigerated.
Analytical Procedures
General.
Most of the procedures used were those from "Standard
(Reference 1) which are usually the same as those of the EPA
Page references and specific detail, where helpful, are given
However, the basis of the method has not been copied
Methods"
(Reference 2).
in the following list.
from the reference; detail has been given only when a choice was made.
The procedures used for water analyses at the trials of American coals in
the Lurgi gasifier at Westfield Scotland (Reference 4) became available
near the end of our work; we tried their method for chloride analysis and
suggest that this reference be studied by someone embarking on analytical
work.
A suggested list of procedures specific to coal conversion studies
was published towards the end of our work (Reference 3). We find little
difference between these recommendations and our procedures, and have tried
to note those differences which we did find.
464
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1 9
Biochemical Oxygen Demand. Standard Methods , page 489; EPA ,
#00310, page 11. Sample dilutions used were 0.5%, 0.2%, 0.1%, 0.02% and
0.01%. Dissolved oxygen was determined using the modified Winkler (azide)
titration as described in Standard Methods2, page 474.
Chemical Oxygen Demand. Standard Methods-'-, page 495; EPA2, #00340,
page 20.
Total Carbon and Total Organic Carbon. Measured using a Dohrmann
Analyzer Model DC-50. This machine measures total carbon directly and
total organic carbon on a sample acidified with hydrochloric acid.
Inorganic carbon is obtained by difference.
Alkalinity (total). Standard Methods1, page 52; EPA2, #00410, page 3,
Chloride. The method of titration with silver nitrate (Standard
Methods2, page 96) proved unworkable because the high concentration of
organic matter caused the solution to darken with time. Sulfides also
interfere. On a sample from the Hygas plant (using Illinois coal), we
tested the following methods with the following results:
1) Reference 4, Method L.12, page 296. Hydrogen peroxide is used to
decompose the organic matter in the sample. Chloride is titrated with
silver nitrate using potassium chromate as an indicator. 10 ml of water
is diluted to 150 ml and boiled for 15 minutes with 10 ml of 30% (by
volume) hydrogen peroxide. A second 10 ml of 30% H202 is then added, and
boiling continued for another 5 minutes. The standard argentometric
analysis did not give a satisfactory endpoint, so excess silver nitrate
was added and the solution back titrated with thiocyanate using ferric
ammonium sulfate as indicator.
2) As No. 1, except two 5 ml portions of 30% by volume hydrogen
peroxide were used.
3) Argentometric method for chloride, Standard Methods1, page 96.
The sample was not pretreated. 5 ml aliquot was diluted to 100 ml. Sul-
fide and thiosulfate interfere with the titration. Hydrogen peroxide pre-
treatment removes the interference.
4) Dry-ashing Method for preparing food samples for analysis from
Reference 5. 100 ml of sample is evaporated to dryness and then ashed in
a muffle furnace at 525°C. The residue is taken up in dilute nitric acid
and Method 3 used for titration. The dry ashing easily removes organic
matter and sulfides. It may be necessary to evaporate samples larger than
100 ml to insure good accuracy as the method requires 20 mg of chloride
for 1% accuracy.
465
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Results:
Chloride, ppm Cl.
Method 4 is particularly easy and seemed to give a satisfactory result.
Cyanide. Standard Methods1, page 399; sulfide is precipitated with
lead carbonate before distilling the sample. Cyanide is determined color-
imetrically using pyridine-pyrazolone reagent.
Ammonia. Standard Methods1, page 224; sample distilled into boric
acid and titrated with sulfuric acid.
Total Kjeldahl Nitrogen. Standard Methods1, page 224; EPA2, #00625,
page 175.
Nitrite. Standard Methods1, page 240; EPA2, #00615, page 215.
Nitrate. Standard Methods1, page 234.
Phenol. Standard Methods1, page 502 (Distillation) and page 507
(Direct Photometric Method); EPA2, #32730, page 241.
Sample distilled and the aqueous solution reacted with 4-aminoanti-
pyrene,
Under ordinary conditions the Standard Methods procedures does provide
for the removal of hydrogen sulfide by the addition of phosphoric acid
(pH less than 4) and copper sulfate to the sample. We do not know whether
all the sulfide is removed from the high sulfide containing samples using
the quantity of copper sulfate recommended in Standard Methods. For
precise and accurate analyses, this aspect of the phenol analysis should
be investigated.
Sulfide. Standard Methods1, page 551; EPA2, #00746, page 284.
• Sulfate. Gravimetric as BaS04, Standard Methods1, page 331; EPA2,
#009471, page 283.
Metals. Calcium, magnesium, sodium, potassium, manganese, iron,
aluminum, barium and silica were determined by atomic absorption spectros-
copy. 100 ml of sample are taken to dryness on a hot plate. The residue
is treated with 5 ml HC1 and 5 ml HNO3 and heated to dryness. This treat-
ment is repeated three times. The residue is then quantitatively trans-
ferred to a 100 milliliter volumetric flask with deionized distilled water.
Thus prepared, the samples were anlayzed using the methods described in the
Jarrell-Ash Atomic Absorption Methods Manual6; these methods parallel
those described in Reference 2. Potassium, sodium, calcium, magnesium,
manganese and iron were determined using an air-acetylene flame whereas
466
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aluminum, barium and silica were determined using an air-nitrous oxide
flame. To surpress ionization, suitable amounts (0.1-1%) of potassium
or lanthanum were added to standards and samples of aluminum, barium and
calcium; 0.1% cesium was added for potassium determinations.
Emission spectrographic surveys are based on 10 drops (0.5 ml) of
solution. Each drop is individually dried with a sun lamp in the elec-
trode crater.
REFERENCES
1. "Standard Methods for the Examination of Water and Wastewater," 13th
Edition, American Public Health Association, Washington, DC 1971.
2. "Methods for Chemical Analysis of Water and Wastes," EPA-625/6-74-003,
U.S. Environmental Protection Agency, Washington, DC, 1974.
3. Kalfadelis, C.D., et al, "Evaluation of Pollution Control in Fossil
Fuel Conversion Processes - Analytical Test Plan," pp. 141-150,
EPA-670/2-74-009-1, U.S. Environmental Protection Agency, Research
Triangle Park, N.C., October 1975.
4. Woodall-Duckham, Ltd., "Trials of American Coals in a Lurgi Gasifier
at Westfield, Scotland," pp. 276-302, NTIS Catalog FE-103, November
1974.
5. "Official Methods of Analysis of the Association of Official Analyti-
cal Chemists," Washington, DC, p. 296, 1970.
6. Jarrell-Ash Atomic Absorption Methods Manual, Jarrell-Ash Division,
Fisher Scientific Corporation, Waltham, Massachusetts.
467
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TECHNICAL REPORT DATA
(Please read Instructions on the reverse before completing)
I. REPORT NO.
EPA-600/7-77-065
2.
3. RECIPIENT'S ACCESSION NO.
4. TITLE AND SUBTITLE
WATER CONSERVATION AND POLLUTION CONTROL
IN COAL CONVERSION PROCESSES
S. REPORT DATE
June 1977
6. PERFORMING ORGANIZATION CODE
7. AUTWOR(S)
D. J. Goldstein and David Yung
8. PERFORMING ORGANIZATION REPORT NO.
9. PERFORMING ORGANIZATION NAME AND ADDRESS
Water Purification Associates
238 Main Street
Cambridge, Massachusetts 02142
10. PROGRAM ELEMENT NO.
EHE623
11. CONTRACT/GRANT NO.
68-03-2207
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
3. TYPE OF REPORT AND PI
Final; 6/75-12/76
14. SPONSORING AGENCY CODE
EPA/600/13
15. SUPPLEMENTARY NOTES E pA project officer Mark J. Stutsman is no longer with IERL-RTP;
for details, contact W. J. Rhodes, Mail Drop 61, 919/549-8411 Ext 2851.
16. ABSTRACT
The report gives results of a study to determine water consumption and
environmental impacts of coal conversion processes in Western states. Part 1 gives
brief descriptions and process water requirements for nine conversion processes.
Detailed designs and analyses are given for the Hygas, Synthane, and Solvent Refined
Coal (SRC) processes, and for Lurgi combined-cycle power generation. At three pro-
posed sites (in North Dakota, New Mexico, and Wyoming), complete water require-
ments and effluents, including all mining and related off-site uses, are given for the
power, Hygas, and SRC plants. The Synthane process is analyzed only at the Wyoming
site. Part 2 gives analyses of influent and effluent waters, with examples for study.
For the three selected plants at the North Dakota site, source water of good quality is
assumed to be cheap and available. For three plants in New Mexico, source water
will be available but is brackish. All the plants at the Wyoming site receive sewage
from a satellite town; additional fresh water is available but is assumed to be
expensive. The plants, being net consumers of water, are designed for no discharge
of water. For each process at each site (10 cases) an integrated water treatment
plant block flow diagram is given with approximate costs and energy requirements.
17.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b.lDENTIFIERS/OPEN ENDED TERMS C. COSATI Field/Group
Pollution
Water Conservation
Coal
Coal Preparation
Coal Gasification
Electric Power Generation
Mining
Water Treatment
Energy
Pollution Control
Coal Conversion
Western Coal
Water Treatment Plants
13B
02C
08G,21D
081
13H
10A
18. DISTRIBUTION STATEMENT
Unlimited
19, SECURITY CLASS (This Report)
Unclassified
21. NO. OF PAGES
483
20. SECURITY CLASS (Thispage)
Unclassified
22. PRICE
EPA Form 2220-1 (9-73)
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