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

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                       RESEARCH REPORTING  SERIES
Research reports of the Office of Research  and Development, U.S.
Environmental Protection Agency,  have been  grouped  into 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.

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                                  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

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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) SystemScaled Plant 	 364
        18.4   Theoretical Design of Air Activated Sludge Systems	373
        18.5   Air Activated SludgeNitrification-Denitrification
                  (AASND) System 	 381
        18.6   High Purity Oxygen Activated Sludge (HPOAS)  System	386
        18.7   Activated Trickling FilterHigh 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 sectionA	    110
 7-2    SRC dissolving sectionB	    Ill
 7-3    SRC hydrogen production by gasification of filter residueA .  .    118
 7-4    SRC hydrogen production by gasification of filter residueB .  .    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

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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 sludgenitrification-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 filterhigh 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)

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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

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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

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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 50C  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

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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 SludgeNitrification-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

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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

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                               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 350-450F.
     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.

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                                                                   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.

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                                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.

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                                                        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.)

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 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.)

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                                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.

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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

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                          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,

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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 1150F and 1400F.
The minimum residence time of coal at the desired temperature level of 1400F
to 1600F 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 700F and
1100F, 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

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                    FEED COAL
    DRIVE
GRATE
DRIVE
 STEAM*
 OXYGEN
                                      SCRUBBING
                                      COOLER
                                          =>
                                          GAS
                     x^F WATER JACKET
                                      3
     Figure 4-3.  Lurgi pressure gasifier ,
               16

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     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
1800F.  In the pilot plant water will be added to the leaving gases which
will vaporize and cool the stream to about 900F.  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 650F.
     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 2800F 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 1700F


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
400-500F.  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 1480-1520F 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 1840-1860F.  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 2050F and
the gasifier at about 1800F.
     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 2000F  to
1500-1600F.  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 1850F 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
                          3538
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 1000F 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 1000F.  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     Na2c3 +  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
 400F with a solution  saturated with sodium bicarbonate at 100F 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 100F,  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 3300-3500P.  The endo-
thermic carbon-steam reactions reduce the exit temperature to around 2700F.
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 95F.  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 450C 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
                                                                                       800F
                                                    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.

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         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

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(This page left blank intentionally)
                    62b

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                              REFERENCES SECTION 4
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 2.    Rudolph,  P. F. H., "The LURGI Process, The Route to SNG from Coal,"
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 3.    Lurgi Mineraloltechnik GmbH,  "Clean Fuel Gas from Coal," Fuel Technology
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 8.    Western Gasification Company, "Amended Application for Certificate of
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10.    Batelle Columbus Laboratories, "Detailed Environmental Analysis Concerning
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                                     63

-------
11.  El  Paso Natural Gas Company, "Second Supplement to Application of El
     Paso Natural Gas Company for a Certificate of Public Convenience and
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12.  Milios, Paul, "Water Reuse at El Paso Company's Proposed Burnham I
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13.  Hatten, J. L., "Pipeline Quality Gas from Coal," Presented at the ASME
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14.  Gibson, C. R., Mammons, G. A. and Cameron, D. S,, "Environmental
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15.  Wyoming Coal Gas Co. and Rochelle Co., "Applicants Environmental
     Assessment for a Proposed Coal Gasification Project," October, 1974.

16.  Michigan-Wisconsin Pipeline Co.  and American Natural Gas Coal Gasifica-
     tion Co.,  "Application for Certificates of Public Convenience and
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17.  Hegarty,  W. P. and Moody, B. E., "Coal Gasification: Evaluating the
     BIGAS SNG Process," Chem. Eng^ Progr. 69, No. 3, 37-42, 1973.

18.  Air Products and Chemicals, Inc., "Engineering Study and Technical
     Evaluation of the Bituminous Coal Research, Inc. Two-Stage Super
     Pressure Gasification Process,"  OCR R&D Report No. 60, 1971.

19.  Grace, R.  H., "Development of the BIGAS Process," IGT Symposium on
     Clean Fuels from Coal, Institute of Gas Technology, Chicago, Illinois,
     Sept. 10-14,  1973.

20.  Grace, R.  J., Brant, V. L. and Kliewer, V. D., "Design of BIGAS Pilot
     Plant," Proceedingsof Fifth Synthetic Pipeline Gas Symposium, AGA, OCR
     and IGU,  Chicago.

21.  Graboski,  M.  S.  and Diehl, E. K., "Design and Operation of the BCR
     Fluidized-Bed Methanation," Proceedings of Fifth Synthetic Pipeline
     Gas Symposium, AGA, OCR and International Gas Union, Chicago, Illinois,

22.  Grace, R.  J.  and Diehl, E. K., "Environmental Aspects of the BIGAS
     Process,"  Symposium Proceedings; Environmental Aspects of Fuel
     Conversion Technology (May 1974, St. Louis, Missouri), Environmental
     Protection Agency EPA-650/2-74-118, 1974.

23.  Hendrickson,  T.  A., ed.,  "Synthetic Fuels Data Handbook," Cameron
     Engineers, Inc.,  1315 S.  Clarkson St., Denver, Colorado 10821, 1975.


                                   64

-------
24.  Air Products and Chemicals, Inc, "Feasibility Study of a Coal Slurry
     Feeding System for High. Pressure Gasifiers," OCR RSD Report No. 68-
     Final Report, 1971.

25.  U.S. Department of the Interior, "Bituminous Coal Research Two-Stage
     Super Pressure Gasification Process 250-Million-SCFD High~Btu Gas Plant,"
     Report 75-7, Bureau of Mines, Morgantown, W. Virginia, 1974.

26.  Fink C. E., "The CO2-ACCEPTOR Process," IGT Symposium on Clean Fuels from
     Coal, Institute of Gas Technology, Chicago, Illinois, Sept. 10-14, 1973.

27.  Fink, C. E., Sudbury, J. D. and Curran, G. P., "CO2-ACCEPTOR Gasification
     Process," AIChE 77th National Meeting, Pittsburgh, Pa., June 3-7f 1974.

28.  Clancey, J. T., Marwig, U. D. and Lindahl, D. R., "Pipeline Gas from
     Lignite Gasification.  Current Commercial Economics," OCR R&D Report
     No. 16 - Interim Report No. 4, 1969.

29.  Fink, C., Curran, G. and Sudbury, J., "CC^-ACCEPTOR Process Pilot Plant -
     1973," presented at Fifth Synthetic Pipeline Gas Symposium, AGA, OCR, and
     IGU, Chicago, Oct. 29-31, 1973.

30.  Fink C., Curran, G. and Sudbury, J., "C02-ACCEPTOR Process Pilot Plant -
     1974, Rapid City, South Dakota," Presented at AGA/OCR Sixth Synthetic
     Pipeline Gas Symposium, Chicago, Oct. 29, 1974.

31.  Corder, C. and Goldberger, W. M., "Design, Installation, and Operation
     of a 25-Ton-A-Day Coal-Gasification Process Development Unit For The
     Agglomerating Burner-Gasification Process," prepared for Office of Coal
     Research, Dept. of the Interior, Washington, D.C. R&D Report No. 97,
     Interim 1, Sept. 1972 - Sept. 1974.

32.  Banchik, I. N., "The WINKLER Process for the Production of Low Btu
     Gas from Coal," IGT Symposium on Clean Fuels from Coal, Institute of Gas
     Technology, Chicago, Illinois, Sept. 10-14, 1973.

33.  Davy Powergas Inc., "Preliminary Technical Data for WINKLER Gasification
     Process," Lakeland, Florida.

34.  Davy Powergas Inc., "Power Gas from Coal via the WINKLER Process,"
     Lakeland, Florida, 1974.

35.  Lewis, P. S., Liberature, A. J. and McGee, J. P., "Strongly Caking Coal
     Gasified in a Stirred-Bed Producer," Bureau of Mines Report of Investiga-
     tion 7644, 1972.

36.  Rahfuse, R. V., Goff, G. B. and Liberature, A. J., "Noncaking Coal
     Gasified in a Stirred-Bed Producer'," Bureau of Mines Technical Progress
     Report 77, March 1974.

37.  Gilmore, D. W. and Coates, N. H., "Behaviour of Caking Coals in Fixed-Bed
     Gasifiers," U.S. E.R.D.A., Morgantown Energy Research Center, W. Va.

                                      65

-------
 38.   Gillmore,  D. W,  and  Liberature,  A.  J.,  "Pressurized,  Stirred,  Fixed-Bed
      Gasification," Presented  at  EPA  Symposium on  Environmental  Aspects  of
      Fuel  Conversion  Technology,  Hollywood,  Fla.,  December  1975,

 39.   M. W.  Kellogg Company,  "Commercial  Potential  for  the  Kellogg Coal
      Gasification Process,"  OCR R&D Report No.  38  -  Final  Report, 1967.

 40.   Cover, A.  E., Schreiner,  W.  C.,  Skaperdas,  G. T.,  "The  Kellogg Coal
      Gasification Process,"  paper presented  at American Chemical Society
      Meeting, Washington, D. C.  (1971) September.   (Division of  Fuel Chemistry
      preprint Vol. 15, No. 3,  page 1.)

 41.   Cover, A.  E., Schreiner,  W.  C. and  Skaperdas, G.  T.,  "The Kellogg
      Coal  Gasification Process Single Vessel Operation," IGT Symposium on
      Clean Fuels from Coal,  Institute of Gas Technology, Chicago, Illinois,
      Sept.  10-14, 1973.

 42.   Cover, A.  E., Schreiner,  W.  C. and  Skaperdas, G.  T.,  "Coal  Gasification:
      Kellogg's  Coal Gasification  Process," Chem. Eng.  Progr. 69, No. 3,
      31-36, 1973.

 43.   Fraley, L. D. and Kumar,  C,  A., "Application of MOLTEN  SALT Gasification
      to Combined Cycles," Presented at IGT Clean Fuels  from  Coal -  II
      Sympoj5ium, Institute of Gas  Technology, Chicago,  111.,  June 1975.

 44.   Hebden, D., "High Pressure Gasification Under Slagging  Conditions,"
     presented  at 7th Synthetic Pipeline Gas Symposium, Chicago,1975.

 45.  Farnsworth, J.  F., Leonard, H. F., Mitsak, D. M. and Wintrell,   R.,
      "K-T:  KOPPERS-Commercially Proven Coal  and Multiple-Fuel Gasifier,"
     AISE Annual Convention,  Philadelphia, Pa., April 22-24, 1974.

 46.  Farnsworth, J.  F., Mitsak, D. M., Leonard, H. F. and Wintrell,   R.,
      "Production of Gas from Coal by the KOPPERS-TOTZEK Process," IGT
     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

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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

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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

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     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

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hydrogasification, hot gases at about 1700F 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
1200F.  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

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     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

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      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

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      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
C2
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 400F 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 750F to adjust the ratio of hydrogen to CO for the downstream
methanation reaction.  The shifted gas is cooled to 100F 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,  1000F)
                  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 230F.
The excess sensible heat in coal is contributed to the coal slurry heater in
raising the recycle oil to 200F.
     (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

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                                   SECTION 6
                               SYNTHANE PROCESS
 6-l  INTRODUCTION AND SUMMARY OF RESULTS

      The Synthane process for the production of high-Btu pipeline gas is
 x                                      19
 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 750-800P
 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 1800F.  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 750P 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 1700F to about 800F.

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
                                                                                        330f
                                     DOWTHERM
                                     STEAM HEAT
                                     EXCHANGE*
          GAS
                CHAR
                       HMOPSIA

1
A
X
                                                                                             -0J
WATER
METHANATO*
1

7S2F
_ 71F
VOOVY,*-,
_:
'
GAS-GAS
HEAT
EXCHANGER

1*
                                                                                         WOfSIA
                                                         WAHU
                                                                93SPSIA
                                                                          -*-c.
                                                                                                              CW-
                                                                                                              tiorsiA
                                                                                                                IOOF
                                                                                                          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 220F 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

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  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

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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  550-800F.
The coal-solvent slurry is pumped together with hydrogen to 1700-2500 psi
and heated to about 850F.  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 350-450F, 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 sectionA.

-------
SRC
SOLUTION
FROM
SECTION A
                               I
FUEL
                        FILTER RESIDUE
                                                        STEAM \j
                                                         CW
                                                                 c
                                                                                                                    VENT
                                                                                                         cw
                                                                                  CW
                                                                        1
BFW
j450F
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  sectionB.

-------
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 750F (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 residueA.

-------
VO
              FROM
              SECTION A
                                STEAM
                       600F
STEAM
                          SHIFT
                        CONVERSION
                         REACTOR
                                         BFW
                                7SOF
         SHIFT
       INVERSION
        REACTOR
                                                    750F
                                                         STEAM
                                                         \^Jl8QF*\       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  residueB.

-------
        TABLE 7-7.   STREAMS IN THE PRODUCTION OF HYDROGEN BY
Stream
No.
10 02
11 HO
12 Raw gas
CO
C2
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

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         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

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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

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   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

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               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 50F 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 170F.  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 100F.  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 100F
drop in temperature.
     (4)  The vacuum preheater is designed to heat the vacuum tower feed
through 100F 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 390F  in a waste  heat
 recovery  boiler.
      (6)  The  SRC  is  assumed  recovered  as a  liquid at 550F.

 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 220F.
      (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
     EPA50/2-74-009-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

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                                  SECTION 8
                          OTHER PROCESS WATER NEEDS
81  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 300F 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
100F.
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 1,500F 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 1,500F.  '       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

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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 10F 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
125F.  However , to best match published water consumption the conditions
assumed are 120F (49C) 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

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            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

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               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

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                                 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

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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

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                              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

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              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

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                 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

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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

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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.

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       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 1000F superheat.  However, the efficiency is not important-
ly increased if steam is produced above about 1000F.  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 1000F 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

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                                                            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 1950F inlet temperature and a pressure ratio of 10:1, producing
          a gas turbine exit temperature of about 1000F.  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
          2400F at a pressure ratio of 16:1 is anticipated.  The correspond-
          ing steam conditions would be 1500 psig at the 1000F level.
               Ultimately a low-Btu gas turbine inlet temperature of 3000F
          at 16:1 pressure ratio is anticipated.  The corresponding steam
          system would operate at 2400 psig 1000F.
     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)
  1950F (Current)
  10/1 P-ratio                900F/1250 psig           11,200 (31%)
  2400F
  16/1 P-ratio               1000F/1500 psig             9750 (36%)
  3000F
  16/1 P-ratio               1000F/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

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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

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            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

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     (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 260F 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 900F and ash leaves at 500F.
    (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 1200F.  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

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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.

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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

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turbine and air compressor, and a cooling air flow.  The temperature of the
combustion products is held to 1950F.
    (17)  To calculate point 18, the stack temperature is taken to be 300F.
    (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

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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

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                  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

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        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

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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
                   58
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

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   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

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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

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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

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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

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      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

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  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

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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 140F 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 140F 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 300F, 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 140F 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 TF.  The
ambient air temperature is 90F.  Cooling water is available at 80F and may
be heated to 105F.  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 = 130F.  The air temperature rise  (see
Table 10-1) is
                                 T       \
                           -=   -  - 90 J  =  40.25    (where T = 130F)

     The high temperature difference, h^ = 280 -  (90+t) = 169.75   (where
T = 130F) .
     The low temperature difference  A2 = T - 90 = 40  (when T = 130F) .
     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

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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

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     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

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     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

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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

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      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 140F  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

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:.
 -
                    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

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the feed stream to the condenser is a mixture of water vapor and nonconden-
sible carbon dioxide at about 230F.  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 140F 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 100F to
keep the water in the system in balance.  By air cooling to 130F and then
wet cooling to 100F, 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 130F contains about
0.18 moles H O/mole CC>2 and the stream saturated at 100F 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

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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 134F,
respectively.  Above 134F or below 115F, efficiency falls.
     The heat rates required when the condenser temperature is in the range
of 115 to 134F 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 115F to a value of 12,200 Btu/kw-hr at the condenser
temperature of 134F.
     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 134F 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 115F to a value of
8,700 Btu/kw-hr at the condenser temperature of 134F.  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

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   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

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   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

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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 134F.  The
condenser is designed to have a maximum temperature of 134F 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
115F.  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 115F.
     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

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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 134F.
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 115F, it is taken to be 115F 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 115F.  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  134F.  At the sum-
mer  design  condition  the  condenser  is designed to have a  temperature of 134F.
As the  ambient temperature drops  during  the year, the condenser  cools  even-
tually  to a temperature of  115F.   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  115F  during these colder months  (below 115F,  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
                 rv*      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) = 99F, TWH  (hot  water temperature)  = 124F and the
condenser temperature TC = 134F.   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, 99F cold water temperature and 25F  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
 115F,  some  cooling  tower  fans are shut off to hold  the condenser at  115F.
 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 115F.
     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  - 96F; condenser temperature T  =  134F;  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 
-------
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   = 71F.  The calculation is trial-
                                      -A\-
and-error.  For a first try assume T  = 115F 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   = 79F.
Since the monthly average temperature is 8F 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  = 134F, T   = 99F, and T   = 124F.  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 60F wet bulb tempera-
ture (the design wet bulb), 25F range and 39F 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 71F and  the wet bulb temperature WBT is
                              \\-*
55F.  The calculation is trial-and-error.  If  the cold water temperature to
the condenser is assumed to be 98F, the warm water leaving the condenser,
based on Equations  (4) and (5) in Section 10-7, is 123F and the condenser
temperature is 133F with a cooling load of 8,680 Btu/hr.  Now with the
ambient wet bulb temperature at 55F 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 98F, 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 
-------
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 115P.  For example,'in month 1 the cold and hot water temperatures of
the condenser are 82 and 106F respectively in order to keep the condenser
at 115F.  With a wet bulb temperature of 20F, the temperature of the cool-
ing water leaving the tower is 77F.  A bypass of the tower, therefore, is
necessary in order to maintain the temperature of the return cooling water
at 82F.
     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 months5, 6, 9 and 10the 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 - 96F               WET CONDENSER AREA - 0.540 ft_
               DESIGN WBT - 60F               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 95F 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 95F 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 140F
followed by wet cooling to 95F.  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 95F 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.

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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 95F 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 95F 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

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         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

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         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

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                                       $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

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       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

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                                 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

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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

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       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
Hage.  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

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                                 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

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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

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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

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            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

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     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

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           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 77F and it is only necessary to quench the
ash to about 200F.  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

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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

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                            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

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                                 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

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     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

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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

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                   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

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                   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

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                   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

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                   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

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                   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

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                   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

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                   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

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                   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

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                   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

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                   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

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                          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

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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~
S42
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
S4

"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

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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

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 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

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             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

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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  122F
 (50C) and 800 ppm total dissolved solids.
      It must be emphasized that Table 14-16 gives equilibrium solubilities.
                                     274

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              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

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               TABLE 14-16.  SOLUBILITY OF CaCO  AT 50C 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

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                                      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 200C, 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

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   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

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            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

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                            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 BoilersA
     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  200C 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-UpA 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 (212F) and the reac-
tion is carried out under pressure to prevent boiling.  The pressure is
                                     285

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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 280C
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

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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

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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

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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

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                                                     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 80P  (25C) 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

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      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

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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.

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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

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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-660/2-75020, 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 WatersPossibili-
      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

-------
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

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                                 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

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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, DCis stated.  The reflux or wash water0is stated.  The  feedF,
FN, FCis given.   The  tower bottomsW, WN, WCare  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
                           TK  =  352 + 1.5PH Q  (psia)
                                 for  18 < P    < 22
                                            H2
                           TK  -  0.555(tF + 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

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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

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                    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 300F (422K) 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

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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

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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

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                           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

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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

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     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

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                     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

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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

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  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

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 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

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 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

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                                 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 FilterHigh Purity Oxygen Activated Sludge
             (ATF-HPOAS) System
          Air Activated SludgeNitrification-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

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    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) SYSTEMSCALED 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

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       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 phenolwhich 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'
4.3xlOgols./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

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     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













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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 SLUDGENITRIPICATION-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
                                                    N3~-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

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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

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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

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                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   80F
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

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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 95-100F 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 1F 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
1F.  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 85F.  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 97F and
to reduce its temperature to 80F 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 80F 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

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                                           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 1F
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 15F,
from 80F to 95F.  Assume that the cooling tower will lower the temperature
of the recycled mixed liquor from 95F to 80F.
     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 95to 80F 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

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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 95to 80F.
     (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

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                                          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 filterhigh 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 1F 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
                                      (T20)
      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 = 30C = 86F
          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 1F rise for 200 mg BOD/1
removed, so the temperature rise in Step 1 is 60F 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 60F
must only heat (R+3) million gallons through 15F, 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 = 1.41C
            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

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      (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

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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

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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

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                            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 4C
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

-------
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

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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

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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

ng/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
S4
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
P4
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.
    S4~
    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^
*^27
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
Si2
P4
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
                               26
 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 AUON,,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.

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                                                 $/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

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                            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 WastewatersPossibilities  and
    Problems," presented at the Workshop on Industrial  Process  Design for
    Pollution Control,  AIChE,  New Orleans, October 1974.
                                    461

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                               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

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                          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 525C.  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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