United States
Environmental Protection
Agency
Industrial Environmental Research
Laboratory
Research Triangle Park NC 2771 1
EPA-600/7-79-133
June 1979
rxEPA
Wastewater Treatment
in Coal Conversion
Interagency
Energy/Environment
R&D Program Report
-------
RESEARCH REPORTING SERIES
Research reports of the Office of Research and Development, U.S. Environmental
Protection Agency, have been grouped into nine series. These nine broad cate-
gories were established to facilitate further development and application of en-
vironmental technology. Elimination of traditional grouping was consciously
planned to foster technology transfer and a maximum interface in related fields.
The nine 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
8. "Special" Reports
9. Miscellaneous Reports
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 sys-
tems. The goal of the Program is to assure the rapid development of domestic
energy supplies in an environmentally-compatible manner by providing the nec-
essary environmental data and control technology. Investigations include analy-
ses of the transport of energy-related pollutants and their health and ecological
effects; assessments of, and development of, control technologies for energy
systems; and integrated assessments of a wide-range of energy-related environ-
mental issues.
EPA REVIEW NOTICE
This report has been reviewed by the participating Federal Agencies, and approved
for publication. Approval does not 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 recommendation for use.
This document is available to the public through the National Technical Informa-
tion Service, Springfield, Virginia 22161.
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EPA-600/7-79-133
June 1979
Wastewater Treatment
in Coal Conversion
by
R.E. Hicks, D.J. Goldstein, F.B. Seufert, and I.W. Wei
Water Purification Associates
238 Main Street
Cambridge, Massachusetts 02142
Contract No. 68-03-2207
Program Element No, EHE623A
EPA Project Officer: William J. Rhodes *"
Industrial Environmental Research Laboratory
Office of Energy, Minerals, and Industry
Research Triangle Park, NC 27711
Prepared for
U.S. ENVIRONMENTAL PROTECTION AGENCY
Office of Research and Development
Washington, DC 20460
-------
ABSTRACT
Water Purification Associates has contracted with the U.S. Environ-
mental Protection Agency to prepare a manual describing water treatment
control technology specific to fuel conversion plant sites in the western
United States. The information contained in this report may become part of
an EPA Standards of Practice Manual.
Most plants converting coal to other fuels use a large quantity of clean
water in the form of steam and put out a large quantity of dirty water that is
condensed when the products from the coal reactor are cooled. The treatment
of the foul condensate is the subject of this report.
Each aspect of water treatment is discussed separately in this report.
The procedures for removing phenolic compounds are discussed in Section 3 and
they include distillation, extraction and adsorption. We have included design
equations, step by step design procedures, and the calculations for a typical
unit. The physical data that are required for design also have been provided.
Section 4 describing ammonia separation and recovery includes design equations
together with the physical data. Illustrative calculations have been given to
show how the design procedure is used. In Section 5, biological treatment, we
have provided the design procedures showing how to destroy organic contamina-
tion including phenol in the condensate. Cooling tower control is the subject
of Section 6. An economical use of the foul condensate is treating it for
makeup to a plant's cooling system. Finally, Section 7 contains a review of
established procedures for sizing plant equipment.
11
-------
CONTENTS
Abstract ii
List of Figures vii
List of Tables x
Conversion of American to International (SI) Units xii .;,
1. SUMMARY 1
2. INTRODUCTION 2
References - Section 2 3
3. PHENOL REMOVAL PROCESSES 4
3.1 Phenol Removal by Distillation 7
3.1.1 Examples: Design of Phenol Distillation System . 17
3.2 Phenol Recovery by Solvent Extraction 21
3.2.1 Solvent Regeneration 30
3.2.2 Solvent Recovery 34
3.2.3 Example: Design of Solvent Extraction Plant . . 40
3.3 Phenol Removal by Adsorption 52
3.3.1 Adsorption Bed Design 52
3.3.2 Resin Generation 58
3.3.3 Example: Design of Resin Adsorption Plant ... 62
References - Section 3 67
4. AMMONIA SEPARATION AND RECOVERY 71
4.1 Introduction 71
4.2 Ammonia Recovery Process 72
4.3 Stripping Tower Design 77
4.3.1 Vapor-liquid Equilibria for HO - NH - CO -
H^S - C,.Hr OH System 82
Z 6 D
4.3.2 Design Characteristics 89
4.3.3 Example: Design of an Ammonia Stripper 94
4.4 Concentrated Ammonia Fractionation 99
4.4.1 Vapor-liquid Equilibria for the Ammonia-water
: System 102
4.4.2 Enthalpy Correlations for the Ammonia-water
System 107
4.4.3 Design Characteristics 107
4.4.4 Example: Design of an Ammonia-water Fractionator 110
111
-------
4.5 Ammonia Recovery Using the Phosam-W Process 117
4.6 Ammonia Recovery as Ammonium Sulfate 122
4.7 Ammonia Recovery by the All-Distillation Process 122
References - Section 4 130
5. BIOLOGICAL TREATMENT 132
5.1 Air Activated Sludge (AAS) Process 134
5.1.1 The Biokinetic Model 135
5.1.2 Values of the Kinetic Constants 136
5.1.3 The Flow Scheme 139
5.1.4 Relationships for the Mean Cell Residence Time .... 140
5.1.5 Sludge Recycle 143
5.1.6 Aeration Basin Volume 144
5.1.7 Summary of Aeration Basin Equations 144
5.1.8 Example 1. Aeration Basin Volume for Moderate Strength
Wastewater 146
5.1.9 Example 2. Aeration Basin Volume for High Strength
Wastewater 148
5.1.10 Multistage Aeration Basins 148
5.2 Solid-liquid Separation 153
5.2.1 Clarification 153
5.2.2 Theory of Thickening 154
5.2.3 Continuation of Summary of Design Equations 160
5.2.4 Values of the Thickening Constants 161
5.2.5 Sludge Dewatering 161
5.2.6 Continuation of Examples 162
5.3 Aeration 169
5.3.1 Oxygen Requirement I69
5.3.2 Values for Oxygen Requirement 171
5.3.3 Aeration Equipment and Aeration Basin Shape 171
5.3.4 Examples 176
5.4 Operation 177
5.4.1 Variables Under the Operator's Control 179
5.4.2 Startup 182
5.4.3 Operation and Control 183
5.4.4 Nutrients and Activated Sludge Seed 184
5.4.5 Determination of Improved Process Constants 189
iv
-------
5.5 High Purity Oxygen Activated Sludge (HPOAS) Process 194
5.5.1 The Biokinetic Model 194
5.5.2 Solid-liquid Separation 196
5.5.3 Oxygen Transfer 197
5.5.4 Cooling of Mixed Liquor 197
5.5.5 A Model for HPOAS 198
5.5.6 Summary of Design Procedures 205
5.5.7 Example 1. Moderate S jength Wastewater 207
5.5.8 Example 2. High Strength Wastewater 211
5.5.9 Use of the Model 219
5.6 Trickling Filters 220
5.6.1 Design of Trickling Filters 221
5.6.2 Example 1. Moderate Strength Wastewater 222
5.6.3 Example 2. Trickle Filter on Recycle, with a High
Strength Wastewater 222
5.7 /_ aerobic Biologic Destruction 226
b.7.1 Design of Experiments on Anaerobic Fermenation .... 227
5.8 The Place of Biological Oxidation 233
5.8.1 Cost 233
5.8.2 Quality of Effluent 236
References - Section 5 237
6. COOLING TOWER CONTROL 241
6.1 Introduction 241
6.2 Consumption of Cooling Water 242
6.3 Scale Prevention 243
6.3.1 Solubility of Calcium Phosphate 247
6.4 Biological Control 253
6.5 Suggested Limits on Quality of Makeup 254
6.6 Fouling Control 258
6.7 Corrosion Control 261
6.8 Slowdown 263
References - Section 6 265
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7. EQUIPMENT SIZING AND COSTING ................... 267
7.1 Introduction ........................ 267
7.2 Storage or Surge Vessels .................. 2^7
7.2.1 Surge Vessels .................... 257
7.2.2 Reflux Drum ..................... 267
7.3 Heat Exchanger Sizing .................... 268
7.4 Distillation Tower Sizing .................. 2^8
7.5 Cost Data .......................... 269
References - Section 7 and Additional Bibliography
VI
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FIGURES page
2-1. Simplified treatment schemes for treatment of fuel conversion
plant wastewaters. (a) Phenolic process condensate, (b) High
sulfur containing condensate but with medium to low concentra-
tion of organics 4
3-1. Phenol removal and recovery by distillation 8
3-2. Mutual solubility of the phenol-water system at 1 atmosphere . 9
3-3. Azeotrope data for phenol-water system 14
3-4. Phenol-water vapor-liquid data on water side of azeotrope.
Curves calculated using Van Laar equations with Equations (17)
and (18) 15
3-5. Phenol-water vapor-liquid data on phenol side of azeotrope.
Curves calculated using Van Laar equations with Equations (17)
and (18) 16
3-6. Steam rates and theoretical stages in water tower for phenol
removal by distillation 22
3-7. Phenosolvan process 23
3-8. Idealized solvent extraction 23
3-9. Ideal counter-current liquid-liquid extraction 25
3-10. The extraction equation 28
3-11. Possible arrangement of solvent regeneration section for phenol
extraction 31
3-12. Theoretical stages for solvent regeneration fractionation. . . 33
3-13. Solvent recovery section of phenosolvan process. ....... 35
3-14. Theoretical stages for water stripping and gas scrubbing in
solvent recovery section 39
3-15. Phenol removal by resin adsorption 51
3-16. Effect of flow rate on resin capacity, (a) Instantaneous leak-
age, and (b) mean leakage, as a function of volume treated.
Amberlite XAD-4 resin: 500 mg/1 phenol in water stream .... 54
3-17. Effect of influent concentration on resin capacity. Based on
data from Ref. 20 and 21, flow rate =0.5 gpm/ft for zero or
low leakage 55
VI1
-------
3-18. Bed flow rate (residence time) for minimum overall cost in
phenol separation by resin adsorption 57
3-19. Theoretical stages and steam rate for methanol regeneration 60
4-1. First stage in ammonia recovery, stripping without reflux. 74
4-2. First stage in ammonia recovery, refluxed strippers ... 75
4-3. Stripping tower nomenclature 76
4-4. Ammonia stripping from Lurgi type condensate in unrefluxed
columns 90
4-5. Ammonia stripping from Hygas type condensate in refluxed
columns 91
4-6. Phenol removal in refluxed and unrefluxed columns .... 92
4-7. Concentrated ammonia still 100
4-8. Dependence of ammonia vapor-liquid equilibrium on operating
pressure 101
4-9. Ammonia concentration in vapor 105
4-10. Determination of ammonia concentration in liquid 106
4-11. Liquid enthalpies of saturated ammonia water 108
4-12. Vapor enthalpies of saturated ammonia-water 109
4-13. Pressure-temperature relation for saturated ammonia . . . Ill
4-14. Steam required for ammonia distillation 112
4-15. Effect of feed concentration on steam rate for ammonia
distillation 113
4-16. Direct distillation of process condensate 114
4-17. Phosam-W process for ammonia separation 118
4-18. Volatility of H S relative to NH in H S-NH -H2O systems . 125
4-19. Chevron all distillation process for recovery of H S and
ammonia from sour waters 126
4-20. Volatility of CO relative to NH in CO -NH H O systems . 127
£ J £• JL £,
Vlll
-------
4-21. Separation of CO_ and NH by distillation using water wash
and reflux. Stripper specifications: 99% CO , 5% NH removal. 128
4-22. Cost of holding down ammonia when stripping CO . Condenser
cooling water costs, condenser capital costs and costs of
additional steam for ammonia fraction only 129
5-1. Air activated sludge model 134
5-2. Substrate utilization vs. substrate concentration 137
5-3. Sludge growth rate vs. substrate utilization rate 138
5-4. Typical settling velocity vs. solids concentration 157
5-5. Typical curve of solids flux vs. solids concentration 158
5-6. Flow diagram for plant: Example 1 - moderate strength waste-
water 163
5-7. Flow diagram for plant: Example 2 - High strength wastewater. . 165
5-8. Simplified process diagram for air activated sludge, Example 2. 166
5-9. Aeration basins for air activated sludge, Example 2 178
5-10. Sludge settling curve 193
5-11. Schematic diagram of UNOX system with surface aerators .... 195
5-12. Schematic diagram of stage bioreactor for HPOAS 199
5-13. Schematic flow diagram of high purity oxygen actived sludge
(HPOAS) system 209
5-14. Configuration of HPOAS system, Example 1 210
5-15. Schematic flow diagram of high purity oxygen activated sludge
(HPOAS) system, Example 2 213
5-16. Configuration of Step 1 of HPOAS system 214
5-17. Configuration of Step 2 of HPOAS system 215
5-18. Cooler schematic 217
5-19. Activated trickling filter-high purity oxygen activated sludge
system for Hygas plants 224
5-20. Schematic of continuously mixed batch anaerobic digester . . . 230
6-1. Pour Is of water evaporated per 1000 Btu of heat transferred in
a wtt cooling tower 245
6-2. Control of solids from bio-oxidation in cooling loop 259
ix
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TABLES
Page
3-1. Analysis of phenols in tar liquor and oil liquor at
Westfield Works 6
3-2. Approximate cost of phenol removal by solvent extraction • 50
4-1. Composition of feed to stripper 89
4-2. Summary of stage calculations for ammonia system 98
4-3. Calculated steam requirements for water stripping and
ammonia fractionation 120
5-1. Biokinetic constants 136
5-2. Analyses of waters for biokinetic data 137
5-3. Calculations on an air activated sludge plant. Example 1. . 147
5-4. Calculations on an air activated sludge plant, Example 2. . 149
5-5. Calculations on a multistage air activated sludge plant for
Example 1 151
5-6. Second stage calculations for Example No. 2 using biokinetic
set No. 3, Synthane wastewater 152
5-7. Sampling and instrumentation points 167
5-8. Sources of analytical procedures 168
5-9. Activated sludge process control parameters 180
5-10. Relationship between process controls and expected responses
of activated sludge process 185
5-11. Suggested actions for various conditions of process
parameters 186
5-12. Representative elemental composition of dry bacteria
protoplasm , 187
5-13. Hypothetical comparison of trace nutrient composition vs.
indicated bacterial requirement for Synthane condensate and
coke plant ammonia still feed 187
5-14. Required nutrients for activated sludge • . . 188
5-15. Liquid and gas phase sources and sinks 203
-------
5-16. Design of HPOAS system. Example 1 208
5-17. Design of HPOAS system. Example 2 212
5-18. Trickling filter calculations for Example 1 223
5-19. Trickling filter calculations for Example 2 225
5-20. Anaerobic systems studied 227
5-21. Optimum and toxic levels for anaerobic fermentation . . 232
5-22. Cost of activated sludge treatment 234
6-1. Approximate material and energy information on a 65%
efficient coal gasification plant 244
6-2. Equilibrium concentrations of calcium for various total
orthphosphate 252
6-3. Equilibrium concentrations of calcium 252
6-4. Experience with bio-oxidation of phenolic refinery waste-
water in a cooling tower 255
6-5. Calculated effluent from Phenosolvan process 256
6-7. Probable composition of makeup to cooling tower in Lurgi
plant 257
6-8. Suggested limits on makeup to a cooling tower 257
7-1. Equipment cost parameters 270
xi
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CONVERSION OF AMERICAN TO INTERNATIONAL SYSTEM (SI)'
To convert from to
UNITS
multiply by
LENGTH
AREA
VOLUME
MASS
WEIGHT RATE OF FLOW
VOL. RATE OF FLOW
ENERGY
POWER
SPECIFIC ENERGY
PRESSURE
WATER FOR ENERGY
HEAT RATE
ft
fe2
acres
ft3
gallons
Ib
tons
103 Ib/hr
tons/day
galloris/mln
gallons /min
10 gallons/day
Btu
kw-hrs
H.P-
kw
106 Btu/hr
Btu/lb
psia
gal/106 Btu
Btu/kw-hr
meters
meters
meters
3
meters
meters
kilograms
megagrams
kg/sec
kg/sec
meters /sec
3
millimeters /sec
meters /sec
kilo joule
(= Newton x meter)
mega joules
Joules/sec
Joules /sec
kilo joules/sec
kilo joules/kg
kilopascall ~
(= kilonewton/m )
m /megajoule
Joules/kw-sec
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
r\
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.
Xll
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1. SUMMARY
Water Purification Associates has contracted with the U.S. Environmental
Protection Agency to prepare a manual describing water treatment control
technology specific to fuel conversion plant sites in the western United
States. The information contained in this report may become part of an EPA
Standards of Practice Manual.
Most plants converting coal to ot1 ir fuels use a large quantity of clean
water in the form of steam and put o1 - a large quantity of dirty water that is
condensed when the products from the coal reactor are cooled. The treatment
of the foul condensate is the subject of this report. The condensate has
contacted coal and crude gas and contains volatile products of coal pyrolisis
and hydrogenation, principally phenols (mixed), other carbonaceous molecules
and ammonia.
Each aspect of water treatment is discussed separately in this report.
The procedures for removing phenolic compounds are discussed in Section 3 and
they inr'ude distillation, extraction and adsorption. We have included design
equations, step by step design procedures, and the calculations for a typical
unit. The physical data that are required for design also have been provided.
Section 4 describing ammonia separation and recovery includes design equations
together with the physical data. Illustrative calculations have been given to
show how the design procedure is used. In Section 5, biological treatment, we
have provided the design procedures showing how to destroy organic contamina-
tion including phenol in the condensate. Cooling tower control is the subject
of Section 6. An economical use of the foul condensate is treating it for
makeup to a plant's cooling system. The level of treatment required and the
effect on the cooling system of contaminated makeup are discussed. Finally,
Section 7 contains a review of established procedures for sizing plant
equipment.
-------
2. INTRODUCTION
Water Purification Associates has contracted with the U.S. Environmental
Protection Agency to prepare a manual describing water treatment control
technology specific to fuel conversion plant sites in the western United States.
The information contained in this report may become part of an EPA Standards of
Practice Manual.
Most plants converting coal to other fuels use a large quantity of clean
water in the form of steam and put out a large quantity of dirty water that is
condensed when the products from the coal reactor are cooled. The condensed
water contains coal decomposition products including mixed phenols and other
organic compounds, ammonia, hydrogen sulfide and carbon dioxide. The treatment
of the foul condensate is the subject of this report.
We have not included the treatment of source waters, whether they are
reasonable quality surface or poor quality ground waters, because these methods
follow standard design procedures that are not unique to fuel conversion plants.
The treatment of wastewaters other than the foul condensate also is not
included; many are similar to those from coal fired power stations and not
unique to fuel conversion plants. The condensate from a methanation reactor is
clean enough to be reused as boiler makeup after a minor polishing. Domestic
sewage can be treated adequately by conventional biological methods or commercial
package units. How storm runoff is treated will depend on its source and the
general cleanliness of the plant.
The selection of a process and the degree of treatment is influenced partly
by the local discharge regulations, and the costs for energy and source water.
The main factor, however, is the contaminants that are found in the wastewater
and they are determined by the type of fuel and the fuel conversion process.
Depending on the coa.1 rank, condensate from gasification plants is generally
cleaner the higher the temperature and the longer it remains in the gasifier.
Nevertheless, all condensates contain varying degrees of ammonia, acid gases,
organics (particularly phenols), suspended solids and oils.
To design any wastewater treatment an analysis of the condensate is needed.
The analysis is not always straightforward because of the high concentrations of
certain contaminants and the difficulty in preserving the water samples. We
recommend a study of the techniques used by Luthy .
-------
The first step in treating the condensate usually is the removal of
solids and oils. Subsequent processing may follow the steps in Figure 2-1 or
the sequence may be changed so that streams for reuse are taken after any
step. For example, ammonia recovery alone may be sufficient treatment for
water that is used in coal slurrying.
Each aspect of water treatment is discussed separately in this report.
The procedures for removing phenolic compounds are discussed in Section 3 and
they include distillation, extraction and adsorption. We have included design
equations, step by step design procedures, and the calculations for a typical
unit. The physical data that are required for design also have been provided.
Section 4 describing ammonia separation and recovery includes design
equations together with the physical data. Illustrative calculations have
been given to show how the design procedure is used.
In Section 5, biological treatment, we have provided the design procedures
showing how to destroy organic contamination including phenol in the condensate.
The section includes:
air activated sludge process
oxygen activated sludge process
- trickling filters
- anaerobic fermentation.
Cooling tower control is the subject of Section 6. An economical use of
the foul condensate is treating it for makeup to a plant's cooling system.
Using treated condensate has two advantages:
- The degree of treatment that is needed for cooling tower makeup is
less than the treatment that would be needed to discharge into a waterway.
(This statement is substantiated in Section 6).
- The use of source water can be kept to a minimum which is useful in
the water-short western states where withdrawals are generally restricted.
Finally, Section 7 contains a review of established procedures for sizing
plant equipment. Within each section we have used examples that best illustrate
a particular design procedure. Because each section stands independently, the
same example is not necessarily given in each section.
REFERENCE
1. Luthy, R.G., "Manual of Methods: Preservation and Analysis of Coal
Gasification Wastewaters." Energy Research and Development Administra-
tion (now Department of Energy). Report FE-2496-8, 1977.
-------
HIGH NHy C0? BOD, COQ
LOW HZS.
^-
PHENOL
REMOVAL
AMMONIA
REMOVAL
COOLING
TOWER
CONTROL
(a)
HIGH NH , H S.
MODERATE BOD, coo.
LDWcoz
AMMONIA
REMOVAL
RESIDUAL
H2S
CLEANUP
COOLING
TOWER
CONTROL
BIOLOGICAL
TREATMENT
(b)
Figure 2.1. Simplified treatment schemes for treatment of fuel conversion
plant wastewaters. (a) Phenolic process condensate; (b) High sulfur
containing condensate but with medium to low concentrations of organics
-------
3. PHENOL REMOVAL PROCESSES
The wastewaters that are recovered from the quench and scrubbing units
during coal conversion contain dissolved organic chemicals, acid gases and
ammonia. The primary organic constituents are phenols including mono, di and
trihydric members in proportions that depend on the type of gasifier and the
temperature of the water/gas contact. One of the major water management
problems in coal conversion processer _s the lack of adequate quantitative
2
data to facilitate the selection of an optimum phenol removal strategy .
Most experience to date in the removal of phenols has been in petroleum
refining and coke oven plants.
An analysis of phenols that are found in the wastewaters from a Lurgi
plant is shown in Table 3-1. The first high temperature quench produces the
tar liquor with a 52% monohydric phenol content; the oily condensate is
produced at a lower temperature and contains 89% monohydric phenols. Total
phenol concentrations may vary from a negligible value in the relatively clean
condens .-e from a Bigas gasifier, for example, to 8,000 mg/1 in the waste-
waters from a dissolving section of a Solvent Refined Coal (SRC) plant.
Phenols exert a biological oxygen demand of about 2.0 Ib 0 /lb phenol,
(the theoretical oxygen demand is 2.38 lb 0 /lb phenol), and must be removed
before discharge. However, if the wastewater will be used as makeup to a
cooling tower, probably up to 100 mg/1 of phenol may remain (see Section 6) .
The alternatives for treating phenol contaminated water generally fit one
of the following processes :
- Recovery (including distillation, solvent extraction, crystallization,
activated carbon adsorption with solvent regeneration and adsorption on
synthetic polymers with solvent regeneration ) where the phenols are concen-
trated for reuse or sale,
Degradation (including bio-oxidation, incineration, ozonation, contract
disposal and activated carbon adsorption with thermal regeneration) with no
attempt to recover the phenols,
Recovery that is supplemented by degradation.
-------
TABLE 3-1 ANALYSIS OF PHENOLS IN TAR LIQUOR
AND OIL LIQUOR AT WESTFIELD WORKS
Phenols (total)
Monohydric Phenols
Phenol
O-Cresol
M-Cresol
P-Cresol
Total Xylenols
Other Phenols
Catechol
3-Methyl
4-Methyl Catechol
3 : 5 Dimethyl Catechol
3 : 6 Dimethyl Catechol
Resorcinol
5-Methyl Resorcinol
4-Methyl Resorcinol
2 : 4 Dimethyl Resorcinol
Monohydric Phenols as Percentage
of Total Phenols
CONCENTRATION, mg/1
Tar Liquor
3,570
1,843
1,260
155
170
160
100
555
394
385
trace
45
272
40
36
trace
Oil Liquor
5,100
4,560
3,100
343
422
302
393
190
80
110
trace
trace
176
64
—
trace
52%
89%
The economics of the processes depend on the total phenol load, Ib/day,
12
and the phenol concentration in the wastewater . For example, because the
cost of solvent extraction is independent of the influent concentration, this
method becomes attractive as the influent concentration (or the amount of
saleable byproduct) increases. Solvent extraction is capital intensive,
however, and costs increase quickly when high phenol removal efficiencies are
-------
required. Thus, it is unlikely that solvent extraction would be used alone
when the treated effluent will be discharged. Adsorption with solvent
regeneration, using many special adsorbents , may be economical for the high
efficiency removal of intermediate concentrations and useful as a backup to
solvent extraction, or by itself.
Biological oxidation is economical when the influent concentration is
below about 1,000 mg/1 phenol and may be required to destroy organics that are
left behind by extraction or adsorption. Ozonation and activated carbon
adsorption with thermal regeneration are prohibitively expensive for all but
the lowest concentrations and are unsuitable for fuel conversion wastewaters.
Incineration and contract disposal are usually impractical.
A discussion of the recovery processes of distillation, solvent extraction
and adsorption follows.
3.1 Phenol Removal by Distillation
There is an old process for the separation of phenol and water by
distillation when no other components are present. Upon examination we find
this process to be too expensive for use on process condensate and so we have
not modified it to include the presence of ammonia and acid gases. However,
the competitive process of solvent extraction and adsorption both require
small units to separate phenol and water. The need for these small units will
be made clear in the following sections. In this section a design procedure
is given for the separation of water and phenol by distillation when no other
components are present.
The phenol-water system forms an azeotrope at about 9.6% phenol by mass
(98.01 mole %). Because wastewaters from coal conversion plants contain less
than about 0.8% phenol, by a single step distillation we produce a virtually
phenol-free bottoms, and an overhead approaching the azeotrope composition.
The overhead consists of up to 10% of the feed and requires additional treat-
ment to recover its water and phenol content. The recovery is done by cooling
the azeotropic mixture which causes it to separate into a phenol rich and
phenol lean fraction. The lean fraction is refluxed to the first distillation
column, while the rich fraction is distilled to produce nearly pure phenol
(Figure 3-1).
-------
03
FW XF
V xW,n
LW xW,n
N
n+1
f«
1
-Ru» XL |
*
WATER
TOWER
si
foS
BW' XW,B
TREATED WATER
, Rp, XR
1 *
PHENOL
TOVJER
STEAM
f^
N
n+1
t «
1
FP
V Y
V YP,n
LP' xP,n
RECOVERED PHENOL
,, BP' XP,B
Figure 3-1. Phenol removal and recovery by distillation.
-------
0.99
O.98
O.S7
0.9«
XL, PHENOL LEAH PHASE
z
o
CJ
0.75
0.7
O.65
AZEOTROPE COMPOSITION
AT 1 ATMOSPHERE
XR, PHENOL RICH PHASE
J L
20
30
40
50
60
TEMPERATURE °C
Figure 3-2. ' utual solubility of the phenol-water system at 1 atmosphere.
-------
Design Procedure
The usual assumptions of constant molal overflow etc. are applicable and
24
standard textbook methods may be used . Feed composition, decanter tempera-
ture and the required bottoms composition for each tower are specified. The
decantor temperature fixes the lean and rich phase compositions as shown in
Figure 3-2. The decantor temperature must be below 34°C and is usually about
20°C = 68°F. Decantor compositions may either be obtained directly from the
graph or calculated from expressions obtained using curve fitting techniques.
The parabolic curves of Figure 3-2 are well correlated by:
XD = 0.868 - (542 - 8.56.t°C)1/2/100 (la)
JK.
1/2
x = 0.976 + (0.90 - 0.019.t°C) ' /100 (lb)
Ll
Where XR and x are the mole fraction water in the phenol rich and phenol lean
fractions.
If the concentration of phenol in the feed is less than in the lean
fraction, the feed enters the water tower and is designated F . If the
phenol concentration exceeds that in the rich fraction, the feed (F ) enters
the phenol tower. Intermediate concentrations are fed directly into the
decantor (F ) . For a feed rate F at concentration x the water tower bottoms
rate is:
BW=F(Xf - XP,B)/(XW,B ""P
which is derived from the overall total material balance giving the
phenol bottoms rate
Bp = F - Bw (3a)
and the overall water material balance
B • Xw __ — F . x — "- B_ • x
P P,B f W W,B (3b)
10
-------
The vapor (steam) rate V in each column must be sufficient to achieve the
desired separation. If the vapor rate is too low, a "pinched" zone will
develop either at the feed or top regions of the column. At the pinched
condition the vapor leaving the region will be in equilibrium with the liquid
entering that region. Applying this condition to the material balance for the
region of interest leads to the following equations for the minimum vapor
rate:
Water Tower
Pinch e feed: V = Bp (xp/fi - xf)/(yf - xf) (4)
Pinch @ top: V = ( (F + F )(x - x ) + B (x - x )
mm | L) r Jj r c o te )
(5)
Phenol Tower
Pinch @ feed: Vmin = BW (XB w - xf)/(yf - xf) - (6)
{>
(Fn + FWJ (XTJ ~ Xf) + BW(XT» H ~ XR^ /(VP ~ XT?}
jj w K r w o fYt I
(7)
These equations are material balances. For example, Equation(4) follows
from material balances taken from above the feed tray of the water tower and
including the phenol tower, namely
V . . y, = B_ . x_ _ + L, . x_
mm f P P,B w F
and
V . = B_ + L
mrn P w
The limiting value for each tower is used to specify the actual vapor
rate V for that tower. The vapor rate is related to the reflux rate R by
R = V + B - F (8)
and is used to calculate the corresponding overhead vapor composition. For
the water tower:
VN = (Vf + VL - VB,W)/VW (9)
11
-------
Enough information is now available for the tray calculations that are
conveniently started at the bottom of each tower. Below the feed a material
balance yields:
x = (Bx + Vy )/(R + F) (10)
o n
The feed is introduced on the tray. where the calculated x exceeds x . Above
the feed a material balance yields :
Xn+l = (BXB + Vyn - Fxf)/R (11>
The equilibrium vapor composition y that is needed for each tray can be found
using the method outlined below. The top tray is where the equilibrium vapor
composition reaches the value given by Equation (9) .
Phenol-water Vapor- liquid Equilibria
18 2 9— 3 2
Published equilibrium data ' for the phenol-water system were
correlated according to the van Laar equations :
log ZA = A/[l + AxA/(BxB)]2 (12)
and
log ZB = B/[l + BxB/(AxA)]2 (13)
where Zft = yftP/(P* xft) , ZB = yfiP/(P* XB) (14)
P = total system pressure,
D* =
vapor pressure
x, = mole fraction water
A
x = mole fraction phenol
The coefficients A and B in the van Laar equations can be found using:
A = log ZA [1 + XB log ZB/(XA log Zft) ] (15)
and B = log ZB [1 + xft log ZA/(XB log ZB) ] (16)
12
-------
These coefficients were estimated using azeotrope composition data
(Figure 3-3) . The resulting calculated vapor- liquid curves were compared with
experimental data. In general the coefficients evaluated at the azeotrope did
not adequately represent the data in the regions x ->• 1 and x -»• 0 because the
calculated values of the coefficients are very sensitive to small errors in
the azeotrope temperature and compositi' i. After some adjustment it was found
that the van Laar coefficients that best represent the data could be correlated
by:
A (water) = 0.324 - 0.00052t°C 55
-------
O
o.
CD
O
CO
t/o
•z.
o
C_J
160
140
120
100
80
60
40
0.97
SYMBOL
(2)
(3)
REFERENCE
32
29
31
18
0*98
MOLE FRACTION WATER AT AZEOTROPE
0*99
Figure 3-3. Azeotrope data for phenol-water system.
14
-------
0.07
CL97
0.88 0.99
MOLE FRACTION WATER IN LIQUID, x
vo
Figure 3-4. Phenol-water vapor-liquid data on water side of azeotrope.
Curves calculated using van Laar equations with Equations (17) & (18)
15
-------
SYMBOL REFERENCE
0 31
0,2 0.3 0.4
MOLE FRACTION WATER IN LIQUID
Figure 3-5. Phenol-water vapor-liquid data on phenol side of azeotrope.
Curves calculated using van Laar equations with Equations (17) & (18) .
16
-------
If not, a new temperature is selected and the process is repeated until
Equation (21) is satisfied.
3.1.1 Example; Design of Phenol Distillation System
To check our equations the phenol distillation example used in Reference
24 was selected. The feed contains 1 mole % phenol (52,200 ppm) and will be
split into a 99.99% pure phenol stream and a 99.999% pure water stream. The
decantor is cooled to 68°F and the ref" .x streams are heated to their equili-
brium temperatures before they enter -he columns.
(a) Stream Specification. We express concentrations as mole fraction
water: x = 0.99, x = 0.99999, x = 0.0001. The feed composition
X W f B P , B
falls on the water side of the azeotrope so F = 100, F = F = O. The
W LJC
pressure at the top of the columns is taken to be 15 psia.
(b) Decantor Streams. Using t = 68°F (20°C) the compositions of the
lean and rich reflux streams are calculated from:
Eq. (la): x = 0.6754
Eq. (Ib): XT = 0.9835
L
(c) Bottoms Split. Equation (2) is used to calculate the bottoms rate
from the two towers:
BW = 99.00
B = 1.00
(d) Equilibrium Vapor Concentrations. The equilibrium vapor concentra-
tion of the feed, and the lean and rich decantor streams is needed to estimate
the minimum vapor rates. To find the equilibrium vapor concentration of the
feed we assumed that the feed will enter the tower at 15 psia (775.7 mm Hg.).
First we estimate the equilibrium temperature; try t = 220°F (104.4°C).
From Eq. (17): A = 0.270; from Eq. (18) B = 1.573.
From Eq. (19): p* = 48.6 mm Hg; from Eq. (20): p* = 889.2 mm Hg.
2
Using Eqs. (12) & (13): Z = 1.002, Z_, = 25.294
A B
Using Eg (14): Y = 1.161, Yn = 0.016
A B
17
-------
Because Y + Y > 1, we must try a lower temperature. A plot of temperature
against the sum of the vapor mole fractions is useful in finding the correct
equilibrium temperature.
On trying tf = 212.7°F (100.4°C) we get
A = 0.272, B = 1.588,
p* , = 40.45, pu = 771.5,
^phenol HO
Z = 1.002, Z^ = 26.065,
A B
YA = 0.9865, YB = 0.0135
Y + YB = 1.000.
The vapor composition in equilibrium with the feed is therefore Y* = 0.9865
at the boiling point of the feed which is 100.4°C.
Similarly we find that Y* = 0.9773 at 214.5°F and Y* = 0.9822 at 212.7°F,
R L
(e) Minimum Vapor. At the feed tray on the water tower using Equation
(4):
Vmin = 1 x (0-0001 - 0.99)7(0.9865 - 0.99) = 283 moles/100 moles feed
At
At the top of the water tower using Equation (5):
Vmin = 1 X (0-0001 ~ 0.9835)/(0.9822 - 0.9835) = 756
At the top of the phenol tower using Equation (7) :
Vmin = I100 X (0'6754 ~ °'99) + " x (0.99999 - 0.6754U /(0.9773 - .6754)
= 2.234
If we use a steam rate of 4/3 of the minimum,
and
VTT = 756 x 4/3 = 1,008 moles/100 moles feed
w
V = 2.234 x 4/3 = 2.979 moles/100 moles feed
18
-------
Reflux Rate and Overhead Composition. The reflux rates to the two
towers are taken from Equation (8) :
1^ = 1,008 + 99 - 100 - 1,007
R = 2.979 + 1 = 3.979
The composition of the vapors leaving the top trays is taken from Equation
(9):
Y = (100 x 0.99 + 1007 x 0.9835 - 99 x 0. 99999) /1008 = 0.9825
W,N
Y = (4 x 0.6754 - 1 x 0. OOOD/2. 979 = 0.9068
P,N
(g) Water Tower Stages. Stage 1 (reboiler) ^ = xfi = 0.99999. If the
pressure drop per stage is taken to be 0.25 psi and for an estimated 16
stages, the pressure in the reboiler will be 15 + (16-1) x 0.25 = 18.75 psia.
Using the equilibrium calculation method, we find y = 0.99992 and t =
224. 5°F. The liquid composition of the next stage is taken from Equation
(10):
x = (99 x 0.99999 + 1008 x 0.99992)/(1007 + 100) = 0.99993
The procedure
The procedure is repeated for each stage. The feed is introduced on the tray
where x falls below x , and the top tray is reached when
above) . The results are shown below:
< VN (see f
Stage
1
2
3
4
5
6
7
8
9
10
11
Since x < x ,
12
13
14
n
n
0.99999
0.99993
0.99980
0.99961
0.99935
0.99888
0.99799
0.99651
0.99432
0.99158
0.98892
stage 11 is
0.98650
0.98473
0.98374
n
0.99992
0.99978
0.99957
0.99929
0.99877
0.99779
0.99617
0.99376
0.99075
0.98783
0.98553
the feed stage
0.98376
0.98276
0.98219
t°F
224.5
223.8
223.1
222.4
221.6
220.9
220.1
219.3
218.4
217.6
216.8
and Equation (11)
216.0
215.2
214.4
P psia
18.75
18.50
18.25
18.00
17.75
17.50
17.25
17.00
16.75
16.5
16.25
is used for x
n+1
16.00
15.75
15.50
y is less than y (0.9825) so stage 14 is the top stage.
19
-------
1x10
1.286x10
0.0161
0.1694
0.6211
1.685xlO~
2.14x10
0.2263
0.8298
0.9736
366.5
364.4
349.5
274.9
219.1
16.75
16.50
16.25
16.00
15.75
(h) Phenol Tower Stages. Stage 1 (reboiler) xx = *B = 0.0001. For
an estimated 8 plates the pressure in the reboiler is 15 + (8-1) x 0.25 =
16.75 psia. Using the equilibrium calculation method, we find y^ = 1.685x10
at 366.5°F. The liquid composition of stage 2 is taken from Equation (10):
x = (1 x 0.001 + 2.979 x 1.685 x 10~3)/(3.979) - 1.286 x 10~
The results for the other stages are shown below:
Stage X Y t°P P psia
1
2
3
4
5
y,. is greater than y (0.9068) so stage 5 is the top stage.
A water tower of 14 theoretical stages, and a phenol tower of 5
theoretical stages yield the desired separation. The graphical method of
Reference 24 gave 15 and 6 1/2 stages respectively. The steam rate required to
achieve the desired vapor rate in each tower is calculated from a heat balance.
The latent heat of water at 18.75 psia is 962 Btu/lb, and "of phenol at 16.75
psia is about 205 Btu/lb. If, for example, we use 300 psig steam superheated
to 600°F with a latent heat of 921 Btu/lb. , then the steam rate per 100 Ib mole
feed for the water column = 1,008 x 962/921 = 1,053 Ib mole steam and for the
phenol column = 2.98 x 205 x 94/(18 x 921) = 3.5 Ib mole steam. Sizing the
columns, reboilers and condensers is discussed in Section 7.
Design Characteristics
Ca) Phenol Tower
The phenol concentrations in coal conversion condensates are on the water
side of the azeotrope, and the main feed is to the water tower. The reflux
from the decantor to the phenol tower enters at a rate of about 1% of the main
feed rate, and at a concentration of 70-75% phenol by mass.
Steam requirements for the coal conversion condensates are about 0.2 Ibs
per 100 Ib feed to the water tower, and about 4 theoretical stages will
generally produce a 99.9% phenol product. Consequently, capital and operating
costs for the phenol tower are low.
20
-------
(b) Water Tower
The pinch condition for our test calculations based on typical coal
conversion condensates invariably occurred at the feed plate so the minimum
steam rate for this tower is given by Equation (4) . Although they are not as
high as those in the example, the minimum steam rates are still excessive at
about 100-150 lbs/100 Ibs feed. This is illustrated in Figure 3-6; note also
that steam duties increase as the phenc load increases.
Some reduction in steam duty wil" be obtained when the required phenol
removal efficiency is lowered. However, at 80% removal efficiency the minimum
steam rate is about 84 Ib steam/100 Ib feed. Further reductions in. steam may
be obtained by operating at elevated pressures, or by decreasing the decanter
temperature. These reductions are not significant and incur additional
capital .and operating costs.
Steam rates for ammonia stripping are at least an order of magnitude lower
than required for phenol removal, so even if ammonia separation could be
achieved •' n the phenol distillation system, overall steam rates would still be
excessive. At $2.50 per 10 Btu, the cost for the steam to the water tower
amounts to $25 to $30 per 10 gal condensate treated, and it is clear that an
alternative means of breaking the phenol water azeotrope must be used.
3.2 Phenol Recovery by Solvent Extraction
In Figure 3-8 we show an idealized solvent extraction system. The feed
water is contacted counter-currently with liquid solvent in an extraction unit
so that the solute (phenol) is selectively transferred to the solvent or
extract phase. The solvent must not mix with water. One measure of the
quality of a solvent is the "distribution ratio," k, defined by
_ concentration of phenol in solvent
concentration of phenol in water
Solvents that have been used in the past include benzene and kerosene. For
benzene k ~ 2.3 which is quite low. Two modern solvents are isopropyl ether
(k ~ 45} and normal butyl acetate (k Z 70).
Part of the complexity of extraction processes is the large number of
Q
possible solvents that should be evaluated. Mixed solvents can be used .
Published information on distribution coefficients refers to phenol. We know
of no tests c actual coal conversion wastewaters to determine what other
materials are extracted by any particular solvent. And yet information such
21
-------
24
22
20
in
us
e>
a:
o
LU
18
12
10
952 PHENOL REMOVAL
FEED CONCENTRATION
k6,000 ppm PHENOL
150
200
300
STEAM RATE, lb/100 1b feed
Figure 3-6. Steam rates and theoretical stages in water
tower for phenol removal by distillation.
22
-------
DIRTY WATER
PHENOL
Figure 3-7. Phenosolvan process. (References 34 & 35)
DIRTY WATER
TREATED WATER ,
Figure 3-8. Idealized solvent extraction.
23
-------
as the distribution ratio for BOD is what determines whether biological
treatment is or is not required subsequent to extraction.
After extraction residual solvent must usually be removed from the water
(no solvent is completely insoluble in water) and something must be done with
the solution of phenol and other materials in the solvent. In one design
water is stripped with a gas and the gas is then scrubbed with crude phenol to
recover the solvent. Figure 3-8 is more simple than an actual design and it
gives some idea of the complexity of a liquid- liquid extraction system.
Another way of recovering solvent is to extract the first solvent with a
second solvent ' . The second solvent must have a particularly low solubility
in water and be a good solvent for the first solvent. However, the second
solvent need not be good for phenol. :
If the solvent is a cheap fuel, such as kerosene or light oil produced by
coal conversion, it may simply be burnt with all of the extracted organic
matter. More commonly, however, the solvent and phenol are separated by dis-
tillation. This means that two additional attributes of a good solvent are a
high volatility relative to phenol and a low latent heat of vaporization.
The energy consumed in solvent extraction is mostly used to distill
solvent and is approximately the product of the solvent flow rate and the
solvent latent heat of evaporation. The solvent flow rate depends on the
distribution ratio and the design.
Consider a simplified, ideal, counter- cur rent extraction (Figure 3-9) .
Complete immiscibility of solvent and water is assumed,: and equilibrium in
each theoretical stage is assumed so that:
Yn
3T=k (23)
n
Material balances for the solute (phenol) give
Stage 1
S(yl " Yo) = s(kxi ~ V0) =
sk
,. = — x, + x, - sy
2 w 1 1 Jo
24
-------
Teed
W
WATER
NJ
Ul
w
s =
y =
x =
water feed rate, 10 Ib/hr
solvent feed rate, 10 Ib/hr
phenol concentration in solvent, ppm
phenol concentration in water, ppm
Figure 3-9. Ideal counter-current liquid-liquid extraction.
-------
Stage 2
s(y2 - y1) = sk(x2 - x1)
skx.
skx
= W(X3 -
i
= W (X3 - ^ -- Xl
8k
skx
3 w \ w
_
J
skx.
w
Xl - Syo
\
' 2 /
sk \ sk f sk . ;
— x, + — x, + x, - sy ' — + 1 '
w;l wl 1 osw
>
\
If y , the concentration of phenol in the feed solvent, is negligible,
one obtains
Stage n
sk
w
\n
+ ...+— + 1
w
so
, x,
sk f
w x.
n+1
+ —
and, by subtraction :
-J. .\T) -!
1 *-l
W
w
w
(24)
or
log
w
log
' w
n + 1
(25)
Equation (25) 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 3-10 is a graphical representation of
Equation (25).
26
-------
On Figure 3-10, the lines have been labelled "percent removal."
xf - xx
Percent removal = 100
Xf
For 99.9% removal: Xf
~
99% removal: *f_ _
~
95% removal: _f_ _
=
80% removal: f _
xi~
In Figure 3-10 it is shown that at any given solvent-to-water rate and
desired percent removal, the number of stages will decrease as the solvent rate
is increased. With fewer stages the extraction column and its cost will be
smaller; however, all of the solvent has to be evaporated to separate it from
the phenol, and increasing the solvent rate increases the cost of the still and
of the energy. Thus, each design requires an economic optimization that must
be repeated for each solvent and for each percent removal.
There is probably a lower limit to the solvent rate which tells us some-
thing about the energy consumption. Good contact is required between the water
and the solvent. Contacting means mixing followed by settling in many stages
while the solvent and water flow counter-currently. Yet another attribute of a
good solvent is that its density differs appreciably from that of water so that
the water can settle. If the wastewater is contaminated with emulsifying
agents such as detergents, the whole process may not work at all. Good contact
and good separation cannot be achieved if the solvent-to-water ratio, s/w, is
too low. As a reasonable example, let us suppose that the minimum s/w is 0.1,
no matter how large the distribution ratio, k. If the energy consumed is
190 Btu/lb solvent, then a solvent-to-water ratio of 0.1 means 158 x 10
Btu/10 gallons of water. This will apply only if a good solvent is available.
If k gets as low as 2.3 (for benzene), economic optimization will result
in s/w >Q.l. In a preliminary study a ratio about 0.7 was found to be
27
-------
CO
UJ
CD
«a:
LO
o:
CD
cc.
UJ
CD
S K
W
Figure 3-10. The extraction equation.
28
-------
3 3
optimum. In this case the energy consumption is about 1,000 x 10 Btu/10
gallons. The optimum for good solvents (k >20) will be found in a following
example.
Sizing Extraction Equipment
Extraction equipment consists of a mixing or contacting section, and a
settling or phase separation section. The size of each section is determined
by the mean residence time required for th<= respective function.
In the contacting stage the two imm ocible liquids are agitated to produce
a dispersion of droplets and a continuous phase, between which mass transfer
occurs. The residence time is determined by the mass transfer rate, which in
turn is governed by the transport properties of the species in solution and the
characteristics of the droplets. The droplet characteristics are functions of
the geometry, size and materials of construction of the mixer, the intensity of
agitation and the phase fraction and fluid properties including viscosities,
densities, surface energies and concentration of dissolved solutes . The
residence t^jie in the settler is governed essentially by the rate of droplet
coalescence. Considerable work to predict droplet breakage and coalescence
rates has been published ' , but even trace impurities can have a profound
effect on the process. Because most researchers have worked with extremely
pure systems, their results cannot be applied to an industrial situation.
The volumes of the mixing and settling sections can be calculated from the
residence time t in each section:
V ft3 = 2.7 x 10~4w (1 + s/w) t min (26)
for w in Ib/hr wastewater feed.
If residence time information is not available, values of 5 minutes for
the mixer and 20 minutes for the settler may be used.
Designs may be based on horizontal mixer settlers as used and supplied by
14
Lurgi . These are probably convenient for pilot scale and preliminary
designs, but in some instances devices using less floor space may be desirable.
The Scheibel or rotating disc type columns may be used. Designs can be
obtained from Chem Pro
29
-------
3.2.1 Solvent Regeneration
In Figure 3-11 we have shown a possible arrangement for a solvent re-
generation section. The solvent is distilled off the solvent-phenol-water
mixture, condensed and recycled to the extraction stage. Reflux and reboil
steam are used. Vapor-liquid equilibria data for the ternary system which are
required for distillation column design may be found using numerical procedures
36
for multicomponent systems
While ternary vapor-liquid data have been published for some systems, they
18
are usually restricted to specific temperature and/or pressure conditions
If binary data are available, they may be used to generate binary coefficients
for the van Laar or Wilson equations , which in turn may be used for estimat-
9 A oc
ing coefficients for the ternary system ' . Experimental data for at least
two temperatures are useful to estimate the temperature dependence of the
coefficients, and to check whether the chosen system of equations is applicable.
The procedure is tedious and involves a trial calculation for composition and
temperature on each plate, and a mass balance on each column. We attempted the
calculations using the van Laar two suffix equations for the ternary vapor-
liquid system but were unsuccessful, apparently because of the high nonideality
of the butyl acetate-water subsystem.
Approximately the system may be treated as a binary mixture. In estimat-
ing the number of stages for separation, the binary system with the smallest
relative volatility should be chosen. When the solvent is butyl acetate,
18
equilibria data for butyl acetate-phenol should be used . We assumed that all
of the water in the ternary mixture would leave the column with the phenol, so
that our design for the phenol column would be safe. The configuration for the
solvent regeneration is consequently different from the one usually shown for
the Lurgi Phenosolvan system, but it is adequate for estimating purposes.
Vapor-liquid equilibria for butyl acetate-phenol
18
The van Laar coefficients at 44.4°C are
A (butyl acetate) = -0.75
and B (phenol) = -0.95 (27a)
At other than 44.4°C the van Laar coefficients can be determined by
assuming the product of the van Laar coefficient and the absolute temperature
to be constant. Thus:
A (butyl acetate at t°C) = -0.75 x (44.4 + 273)/(t + 273)
B (phenol at t°C) = -0.95 x (44.4 + 273)/(t + 273)
30 (27b)
-------
SOLVENT/PHENOL FROM EXTRACTION SECTION
SOLVENT
RECOVERY
SECTION
PHENOL/SOLVENT fROH GAS SCRUBBER
PHENOL TO GAS SCRUBBER
SOLVENT
TOWER
f OECANTOR \
RECYCLE TO
EXTRACTION SECTION
PHENOL
TOWER
PHENOL PRODUCT
Figure 3-11. Possible arrangement of solvent regeneration section for phenol extraction.
-------
These values are used in the van Laar Equations (12) & (13) to determine
vapor or liquid equilibrium compositions. The vapor pressures may be evaluated
from Equation (19) for phenol, and for butyl acetate from:
Up to 130'C l°*10P*butyl acetate = 5'183 ' 504'41/<93 + t<>C) (28a)
Above 130-C Iog10p*butyl acetate = 7.528 - 1859/(273 + t°C) (28b)
«
The calculation for equilibrium composition is described in Section 3-1.
Tower design follows textbook methods for binary mixtures ' and is
illustrated by the example in Section 3.2.3.
Design for minimum cost
Prior to designing the solvent extraction plant, it is useful to know
under what conditions the individual sections should operate for minimum cost.
Rapid methods for estimating optimum process variables are given in Reference
38. Results of some detailed calculations are presented here for the solvent
regeneration system, and in the next section for the solvent recovery system.
Although they apply only to the stated flow rate and cost values, they do
serve as a general guide.
The costs of the solvent regeneration section are made up of the solvent
fractionator and the phenol tower, and their associated heat exchangers, steam
and cooling requirements. The cost of the phenol section is small and will be
considered to be constant. The cost of the solvent tower, however, will
depend on the solvent rate, the steam rate and the purity of the bottoms going
to the phenol still.
The solvent fractionation section was costed on the basis of a 2,000 ppm
phenol stream for a range of the above variables. We took steam at $2.50/10
Btu, cooling water at 12C/10 gal circulated, and a tray efficiency of 50% in
the distillation column. It was found that steam and cooling costs dominate
and minimum costs are for tall columns using steam rates close to the minimum.
As shown in Figure 3-12, the minimum stripping steam rate is about 14 Ib steam
per 100 Ibs feed, or 15 Ibs steam per 100 Ibs butyl acetate. Assuming maximum
energy conservation by heat interchange between process streams using butyl
acetate, the minimum cost for solvent regeneration is about $(6 s/w) per 1,000
gals wastewater extracted. Steam and cooling account for about 85% of the
total costs.
32
-------
LO
UJ
CD
to
cc.
o
cer
UJ
en
20
15
1O
-v1
FRACTIONATION
EFFICIENCY
10
15
20
25
3O
STEAM RATE lbs/100 Ib feed
Figure 3-12. Theoretical stages for solvent regeneration fractionation
33
-------
Because solvent recovery costs (next section) are virtually independent
of the solvent-to-liquid flow rate, the cost for solvent regeneration can be
used together with the cost for the extraction equipment to find an optimum
solvent flow rate for the desired phenol removal efficiency. We found this
optimum solvent flow to be s/w =0.3 for 99.9% phenol removal, and s/w =0.1
for 80% phenol removal, based on a distribution coefficient k = 10 to 20. The
corresponding costs (extraction equipment and solvent regeneration) are about
$2.0 and $1.0 per 1,000 gals treated respectively. One set of cost data is
included in the example given in Section 3.2.3.
3.2.2 Solvent Recovery
The water leaving the extraction section will contain solvent at its
solubility concentration or higher if adequate settling was not achieved in
the settling unit. Butyl acetate has a solubility of about 0.7 lb/100 Ib
water, which is equivalent to $22 per 1,000 gals. Equivalent values for
isopropyl ether are 1.2 lb/100 Ib water and $16 per 1,000 gals. Thus, solvent
must be recovered from the water stream for economic reasons as well as to
prevent problems in downstream treatment and reuse.
In the Phenosolvan process solvent is stripped from the water with gas,
washed from the gas with some of the wastewater prior to extraction, and then
polished by washing with crude phenol. We have considered only washing with
phenol in the simplified flow sheet (Figure 3-13). The following design
procedure for determining the number of stages in the stripping and scrubbing
towers will be used to find a cost optimized design for the solvent recovery
section using a 2,000 gpm 6,000 ppm phenol wastewater stream when using
nitrogen as the stripping medium.
Design Procedure
A material balance around the lower section of the stripping or scrubbing
tower, up to plate n, yields:
Xn+l " G(yn ~ V/L + Xl (29)
34
-------
Ul
Ul
^J E-6A
PHENOL/SOLVENT
TO SOLVENT REGENERATION
E6-B
o
BLEED FROM PHENOL PRODUCT
STREAM
P-4
UASTFWATFR STREAM FROM SOLVENT EXTRACTION
GAS/SOLVENT
SCRUBBER
n+1
n
GAS
GAS
BLOWER
P-5
III
STRIPPER
n+1
n
DEPHENOLIZED EFFLUENT
Figure 3-13. Solvent recovery section of Phenosolvan process.
-------
where x mole ratio solvent in liquid
y = mole ratio solvent in gas
G = solvent free gas rate, moles/hr, assumed constant
L = solvent free liquid rate, moles/hr, assumed constant
n = stage number, counting from bottom.
In designing the stripper, the value of x , which is the solvent concen-
tration in the treated effluent, is specified and calculations proceed up the
tower until x , > x_, which is the solvent concentration in the feed. Tray n
n+1 f
is then the feed tray and the top of the tower. For the scrubber, the value of
x , which is the solvent concentration in the phenol leaving the tower, is
found from an overall material balance. Calculations then proceed up the tower
using Equation (29) until the solvent has been removed from the gas.
A value for L/G, the liquid to gas mole ratio, must be specified when
using Equation (29). In each tower there will be some limiting L/G value
beyond which the desired separation will not be achieved. This limiting L/G
value may be calculated from an overall material balance:
L (xf - X;L) = G (yn - yQ) (30)
by substituting y = y* for the stripper,
and x = x* for the scrubber. The asterisk denotes an equilibrium
value. Equilibrium concentrations for use with these equations may be obtained
from published data, if available; if butyl acetate is solvent, the following
procedure may be used:
Stripper vapor-liquid equilibria for butyl acetate-water
The feed to the stripper will contain about 0.7 Ib butyl acetate and
probably less than 0.04 Ib phenols per 100 Ib water. We will treat the system
as a binary butyl acetate-water mixture in the presence of an inert saturated
gas with xf = (0.7/100)x(18/116) = 0.0011. Correlation of the data of Weller18
for 112°F in the concentration range x = 0 to x = 0.0011 mole ratio butyl
acetate show that the partial pressure of the solvent is:
PD, = 7.68.106x2 + 19.44.103 x, in mm Hg (31)
BA
36
-------
x is the liquid phase concentration, moles butyl acetate per mole water. The
mole ratio in the gas phase can be calculated from
y = PBA/(P - PBA> (32)
where P is the total pressure in the tower.
Scrubber vapor-liquid equilibria for butyl acetate-phenol
Butyl acetate is completely miscibl* with phenol and concentrations of up
to 70% butyl acetate by mass may be reached in the scrubber. Correlation of
1 8
the data of Weller for 112°F give the partial pressure of the solvent as
p = 9.5 x/(l + x) x < 0.2 (33a)
13A
and
p = 34.5 / x/(l + x) I 1'76 0.2 < x < 10 (33b)
BA ^ j
Here x is moles butyl acetate per mole phenol in the liquid phase. The
corresponding mole ratio in the gas phase is taken from Equation (32).
Design Characteristics
The limiting L/G rates in the stripper and scrubber are found from
Equation (30) and the appropriate vapor-liquid equilibrium equation for
acetate. For the stripper, x = 0.0011, y * = 0.0415 (Equations 31 & 32) so
(G/F) . . . = (0.0415 - 0)/(0.0011 - 0) = 0.026
mm, stripper '
For the scrubber if x, = o and y = 0.0415, then x * >10 (Equations 33b & 32)
f o o o
so
(P/G) . ,, <(0 - 0.0415)/(0 - 10) <0.004
mm, scrubber
where F = water feed rate to stripper, Ib moles/hr
P = phenol feed rate to scrubber, Ib moles/hr. The ratio for the
scrubber probably falls below the limiting value for good contacting in a
packed column, and this should be checked for a final design.
37
-------
The number of theoretical stages decreases as the flow rates are increased
above the minimum (Figure 3-14). The number of stages also decreases as the
specified removal efficiency decreases. The value of butyl acetate' remaining
in the water when stripping to 90, 99 and 99.9% efficiency is about $2.2,
$0.22 and $0.02 respectively. Provided there are no other constraints on the
solvent concentration remaining in the water, the stripping efficiency will be
determined by system economics. The efficiency of the scrubber will generally
be between 99.9 and 99.99%; the number of stages does not change significantly
in this range.
Design for Minimum Cost
The solvent recovery section is treated as an independent system and
costed for various conditions on the basis of a 2,000 gpm, 6,000 ppm phenol
stream after removing 95% of the phenols by extraction. Butyl acetate is the
solvent.
The capital cost of the stripper depends on the solvent concentration in
the feed, the stripper efficiency and the gas rate. The feed concentration is
fixed, and the other two parameters were varied over the range shown in
Figure 3-14. The operating costs are determined by the power required to pump
the gas and the value of the solvent that is not recovered. A cost of 38C/lb
butyl acetate was used.
The cost of the scrubber is dependent essentially on the phenol flow rate
and the volume of gas to be treated. The cost of the two heat exchangers
(Figure 3-13) was included in the costs; returning the phenol stream to the
solvent regeneration section was not included.
In figuring capital costs for the towers, tray columns at 50% stage
efficiency were assumed. We do not anticipate costs for packed columns to
differ significantly. The contribution of capital cost to the annual total
cost for solvent recovery was assessed at 19%; 15% for amortization and 4% for
maintenance.
The minimum cost for solvent recovery is 30C per 1,000 gallons. This was
obtained with a stripper recovery efficiency of 99.9%, for a gas-to-water
ratio (G/F) of 0.05 to 0.055 in the stripper and for a phenol-to-gas ratio
(P/G) of 0.05 in the scrubber. The total cost is divided: 56% for stripping,
8% for solvent losses, 32% for gas blowing and 4% for scrubbing. Costs are
not strongly dependent on (G/F) and (P/G). As (G/F) is increased, scrubber
costs decrease, but gas blowing costs increase. Scrubbing costs make only a
small contribution to the total costs so dependence on (P/G) is weak. A (P/G)
38
-------
2 4 6 8 10
GAS RATE TO STRIPPER, moles gas/100 moles feed water.
20
15
in
5
1O
MINIMUM
SCRUBBING
EFFICIENCY
99.90%
2 4 6 8 10
PHENOL RATE TO SCRUBBER, moles phenol/100 moles qas.
Figure 3-14.
Theoretical stages for water stripping and gas scrubbing
in solvent recovery section.
39
-------
value of 0.1 requires a phenol circulation rate of two to three times the
product phenol rate, and it will increase the cost of the phenol tower in the
regeneration section. However, phenol throughput is relatively low and costs
of the phenol tower will not significantly affect the total plant costs.
At 99.9% solvent recovery, the treated water will contain 7 mg/1 butyl
acetate which will be stripped in the ammonia recovery section. This low
concentration is not expected to cause problems either in the sulfur recovery
section or in the ammonia fractionator where organics are removed by addition
of caustic soda. Thus the optimized costs for the solvent recovery section
may be assumed satisfactory guides in designing solvent extraction plants.
3.2.3 Example; Design of Solvent Extraction Plant
In this design our source is a 2,000 gpm or 10 Ib/hr wastewater stream
containing 6,000 mg/1 phenols. We have chosen butyl acetate as the solvent
and we will design for high phenol removal efficiency of 99.9%. Based on
earlier findings, we select a solvent rate of s/w = 0.3, or 0.3 x 10 Ib/hr
butyl acetate. The extract phase leaving the extraction equipment will contain
dissolved water, most of which will be recycled to the extraction equipment
from the solvent regeneration section. In the worst case, all this water will
be recycled from the decanter on the phenol still and will contain phenol at a
concentration just below the azeotrope. We will design the extraction system
for this worst case.
Water solubility in butyl acetate = 2.7% by mass.
Total water recycled = 0.027 x 0.3 x 106 = 8,100 Ib/hr
If the decantor temperature is 68°F (20°C) , this water will have a mole
fraction phenol of (1 - 0.9832) or 0.0168 as calculated using Equation (Ib).
The phenol recycle is:
(0.0168) moles phenol 8,100 moles water
(1-0.0168) moles water X ~~TI hr = 7 39 ^ moles phenol/hr
or 723 Ib phenol/hr.
The feed concentration to the extraction system consequently increases
from 6,000 ppm to
xf = 10 x (6,000 + 723)/(106 + 8,100) = 6,670 ppm
40
-------
and for 99.9% removal from the original feed, the concentration of the treated
water must be
x = (1 - 0.999) x 6,000 = 6 ppm
The distribution coefficient for butyl acetate-phenol is about 65, but in view
of the other constituents present in coal conversion condensates and the high
overall removal we require, a value of k 10 is used. Experimentally deter-
mined k values should be used if availavj.e. Using the above values for x , x ,
s/w and k, we find from Equation (25) that 6 extraction stages are required.
The actual number will equal the calculated number because we have used a
conservative k value, and will use mixer-settlers that probably operate at
close to equilibrium conditions.
Volume and Cost of Mixer-settlers
Using residence times of 5 minutes for the mixer and 20 minutes for the
settler froir. Equation (26) :
V . = 2.7 x 10~4 x 106 (1+0.3) x 5 = 1,755 ft3 and
mixer
V _ = 1,755 x 20/5 = 7,020 ft
settler
The vessel is 12 ft high with 27 ft sides and contains a 12 x 12 mixing parti-
tion. The volume of metal for the six vessels may be used to estimate the cost
of the unit. A typical mixer-settler arrangement is given in Reference 39. If
we use 1/2" metal plate for sides, top and bottom, and arrange the vessels in
two back-to-back trains of three vessels each, about 600 ft or 300,000 Ib of
metal are required. Assuming a cost factor of $6 per Ib, and an installation
factor of 3.75 allowing for field fabrication, mixing motors, pumps, instrumen-
tation, etc., the total installed cost will be about $6.8 x 10 .
Solvent Regeneration (Figure 3-11)
The solvent still is designed to handle the full solvent and phenol load.
Solvent rate = 0.3 x (10 + 8,100) = 302 x 10 Ib/hr or 2,600 Ib
mole/hr
Phenolrate = 6,000 + 723 = 6,723 Ib/hr or 72 Ib mole/hr
Water -ate = 0.027 x 302 x 103 = 8,100 Ib/hr or 450 Ib mole/hr
The main feed rate is the sum of the components = 3,122 Ib moles/hr.
41
-------
There will be secondary feed from the solvent recovery section consisting of
phenol and butyl acetate. This feed does not significantly alter the design
of the solvent fractionator because the phenol is relatively nonvolatile and
will be removed directly with the bottoms and the butyl acetate has been
accounted for in determining the main feed rate. The extra phenol must,
however, be included in the design of the phenol still.
Solvent Fractionator (Figure 3-11)
The number of stages may be estimated on the basis of a binary phenol-
butyl acetate mixture. Using the above feed rates we obtain
x = 2,600/(2,600 + 72 + 450) = 0.833
We require a virtually phenol-free solvent recycle, say less than 0.02 mole %
impurities, and specify x = 0.9998. Since any butyl acetate that is not
removed in the solvent still can be recovered from the phenol still decanter,
a recovery efficiency of E = 95% is adequate. With a feed of F = 100 moles/hr,
the overhead rate D is:
Dx = EFx
so
D = (0.95)(100)(0.833)/0.9998 = 79.15 moles/hr
and by material balance the bottoms rate is
B = 100 - 79.15 = 20.85 moles/hr
Calculation of equilibrium vapor compositions follows the procedure in the
phenol tower design example, but using the van Laar coefficients A and B from
Equation (27) . The equilibrium temperature of the feed was found to be
277.5°F (136.4*C) at a pressure of 15 psia (775.7 mm Hg) . At this temperature
A = -0.75 x (44.4 + 273)/(136.4 + 273) = 0.5815 and
B = -0.95 x (44.4 + 273)/(136.4 + 273) = -0.7365
42
-------
From
Equation (28b) : p* , = 970.7 nun Hg
* butyl acetate
Equation (19): p* , = 178.4 mm Hg
phenol
Equation (12): z = 0.9466; Equation (13): Z = 0.3401
A 15
Equation (14): yft = 0.9466 x 970.9 x 0.833/775.7 = 0.9869
and yo = 0.3401 x 178.4 x (1-0.833)/775.7 = 0.0131
3
Since y + y = 1.0 the temper.ature is correct and y* = y = 0.9869.
A B • £ A
The minimum reflux ratio for the desired fractionation is:
Rmin = (xo - yf)/(*f - V
= (0.9998 - 0.9869)/(0.9869 - 0.833) = 0.086
In Section 3.2.1 we found that it is economical to keep the steam rate low.
From Figure 3-12 a steam rate of 15 lb/100 Ib feed or one mole steam per mole
feed seems reasonable. The vapor rate for a feed of 100 moles/hr is, there-
fore, V = 100 moles/hr and the reflux ratio is:
R = (V/0) - 1 = (100/79.15) -1 = 0.26
This is adequately above the calculated R . value.
^ J mm
Stage calculations can be started at the bottom of the tower. The bottoms
concentration is given by x^ = (Fxf - Dx)/B = (1-E) Fx_/B = 0.200. For a
pressure drop of 0.25 psi per stage and an estimated 11 stages, the reboiler
pressure is 15+(11-1) x .25 = 17.5 psia. After finding the equilibrium vapor
compositions in the usual way, we can determine the liquid composition on the
next tray by material balance. Below the feed use
Xn+l = (BXB + V*n)/(L + F)
The feed is introduced when X > x_ following which use
n I
Xn+l - (BXB + Vyn ' FXf)/L
where L, the liquid rate at the top of the column is given by V-D or
100 - 79.15 = 20.85 in this example. The stage calculations are listed below:
43
-------
Ppsia
1 0.200 0.2822 365.1 17.50
2 0.268 0.4053 360.1 17.25
3 0.370 0.5902 350.1 17.00
4 0.523 0.8145 329.5 16.75
5 0.708 0.9527 300.7 16.50
6 0.823 0.9847 284.4 16.25
7 0.847 0.9889 280.2 16.00
Plate 7 is the feed tray
8 0.947 0.9978 269.3 15.75
9 0.990 0.9997 264.9 15.50
10 0.999 0.9999 263.2. 15.25
The vapor leaving plate 10 exceeds the specification of
x = 0.9998 and so plate 10 is the top tray.
o
Utility Requirements and Costs
The reboil steam used for the design was 1 mole steam per mole feed or 12
x 3,122 = 56.2 x 1Q3 Ib/hr.
If it is not possible to preheat the feed using waste heat from another
process stream, extra steam will be needed. Butyl acetate has a specific heat
of 0.46 Btu/16°F so this heat load is approximately
3,122 lb mole BA/hr x 116 Ib BA/lb mole x 0.46 x (280-100) Btu/lb
= 30 x 106 Btu/hr
Thus the total heat load is about 86 x 10 Btu/hr and at a cost of $2.50
per 10 Btu, the steam cost is $1.80/10 gal wastewater treated. The total
condenser heat load is the latent heat of the vapor plus the sensible heat to
subcool the product. Butyl acetate has a latent heat of 133 Btu/lb so heat
duty is:
3,122 x 116 lb BA/hr x ) 133 Btu/lb + 0.79 lb product/lb feed x 0.46(263-100)
| 133 Btu/lb + 0.79 lb product/lb feed x 0.46(263-100)1
69.6 x 10b Btu/hr
Assuming a 25°F temperature rise and a cost of 12
-------
Equipment Sizes and Costs
The tower diameter can be found using the procedure given in Section 7.
The liquid rate below the feed is 120.85 x 3,122/100 = 3,773 Ib mole/hr. Using
a molecular weight of 100 and a liquid density of 50 Ib/ft (6.0 Ibs/gal), this
becomes 3773 x 100/(6.0 x 60) = 1,048 gpm. The vapor density at 300°F and 16
3 3
psia is 0.196 Ib/ft . The vapor rate is 3,122 Ib moles/hr or 442.5 ft /s
1/2
giving a vapor loading of 442.5 x (0.196/50) = 27.7. A column diameter of
about 11" 6" is required. The height of a 10 stage column with an assumed
stage efficiency of 50%, 2 ft per plate and a 15 ft skirt, is about 60 ft. The
volume of metal is 88,900 cubic inches or a metal mass of 25,200 Ib. Assuming
a cost factor of $10/lb and an installation factor of 3.75, the tower cost is
$0.94 x 10 . The reboiler condenser and reflux drum may be sized and costed as
shown in Section 7 and will probably raise the cost to $1.06 x 10 .
Phenol Still
The phenol still is designed to handle the full phenol (72 Ib mole/hr) and
water (450 l.h mole/hr) load to the solvent still plus the phenol required for
gas scrubbing in the solvent recovery section. We found in Section 3.2.2 that
the optimum liquid-to-gas mole ratios in the water stripper and gas scrubber
are about 18 and 0.05 respectively. As the water flow to the stripper is 10
4 4
Ib/hr (5.56 x 10 Ib mole/hr), the phenol rate to the scrubber is 5.56 x 10 x
0.05/18 = 150 Ib moles/hr phenol. The total phenol rate to the phenol still
is, therefore, (72 + 150) = 222 Ib mole/hr and the total feed rate is 222 + 450
= 672 Ib mole/hr. The water mole fraction in the feed is x = 450/672 = 0.67.
Because this composition is similar to the rich phase concentration, it can
feed into the top of the phenol tower. The rich phase from the decanter is
recycled to the tower and the lean phase is recycled to the mixer-settlers. We
have designed for a bottoms concentration of x = 0.0001 following the procedure
B
given in Section 3.1. The stage calculations are listed below:
Stage
1
2
3
4
5
x
0.0001
0.0013
0.0156
0.1636
0.6077
y
0.0017
0.0212
0.2216
0.8229
0.9728
t°F
366.5
364.4
349.0
275.4
219.3
P psia
16.75
16.5
16.25
16.0
15.75
Plate 5 is the top and feed plate
45
-------
Utility Requirements and Costs
The minimum vapor rate is 69.5 moles per 100 moles feed. We will use 90
moles per 100 moles feed or (90x661/100) = 595 Ib mole/hr. The steam rate
that will produce this vapor rate is equivalent to:
595 Ib mole/hr x 94 Ib/lb mole x 205 Btu/lb = 11.5 x 10 Btu/hr
There is an additional heat load to heat the reflux from the decanter temperature
(68°F) to the column temperature (224°F). The reflux rate is 21.9 moles/100
moles feed or 21.9 x 661/100) = 144.8 Ib mole/hr and the reflux mixture has a
specific heat of about 23 Btu/lb mole °F. The extra heat load is:
144.8 Ib moles x 23 Btu/lb mole °F x (224-68) °F = 0.5 x 10 Btu/hr
The energy cost at $2.5/106 Btu is 2.5 x (0.5 + 11.5) = $30.1/hr or $0.25/103
gals wastewater treated. The cooling load on the condenser is about equal to
the total heat load of 12.0 x 10 Btu/hr, and cooling water costs at 12C/10
gals circulated, 25°F temperature rise, will be about $0.06/10 gal wastewater
treated.
Equipment Size and Costs
Using the methods described in Section 7, we obtain a liquid rate of 150
gpm and a vapor loading of 4.6 requiring a column diameter of 4 ft. With a
plate efficiency of 50%, the 5 stage 10 plate column will have a height of 40
ft including a 15 ft skirt. The mass of metal is 3,700 Ib, which at $10/lb
and an installation factor of 3.75 gives an installed cost of $0.14 x 106.
The heat exchangers may be sized as suggested in Section 7 and this will
increase the total cost to about $0.2 x 10 .
Solvent Recovery (Figure 3-13)
Stripper. Butyl acetate has a solubility of 0.7 lb/100 Ib water, so the
concentration of the extracted wastewater fed to the stripper is
xf = (0.7/116)/(100/18) = 0.0011
To recover 99.9% of this dissolved solvent, the bottoms concentration is:
x^ = (1 - 0.999) x 0.0011 = 1.1 x 10~6
46
-------
A liquid to gas mole ratio (L/G) = 18 is used because we have shown that it
is economical. We assume that the incoming gas is virtually solvent free, so
y = 0. The stripper should be operated at low pressure; 18 psia (930.9 mm
Hg) is specified for the bottom stage in this example. A pressure drop of 0.1
psi per stage is equivalent to 0.7" HO per foot of column height, based on a
50% stage efficiency and 2 ft of packing per actual stage (i.e. 4 ft per
theoretical stage). This pressure drop is a conservative value for the pack-
ing described in Section 7.
Stage Calculations. The vapor concentration in equilibrium with the liquid is
found first.
From Equation (31): p = 7.68.10 x(1.1.10~6)2 + 19.44.103xl.1. K>6
Dt\.
= 0.0214 mm Hg
From Equation (32): y^ = 0.0214 (930.9 - 0.0214) = 22.99 x 10~6
The liquid concentration on the next stage is:
From Equation (29): x2 = (22.99 x 10~6 - 0)/18 + 1.1 x 10~6 = 2.38 x 10~6
The stage calculations continue until the liquid concentration equals
or exceeds the feed concentration. This stage is then the top of the stripper.
Stage
1
Bottom
2
3
Etc.
20
21
22
x
l.lxlO~6
2.38xlO~6
-6
3.88x10
-4
5.16x10
-4
8. 18x10
1.45xlO~3
y
22.99xlO~6
44.97xlO~6
-6
82.00x10
-2
1.47x10
-2
2.61x10
(Exceeds x
P psia
18.00
17.9
17.8
16.10
16.00
- 1.1x10"
Top
At 4 ft per theoretical stage, a packing height of 88 ft is required and the
total column height will be about 100 ft.
47
-------
Scrubber. If the gas leaving the stripper is in equilibrium with the feed
water, it will have a vapor concentration calculated of y = 0.041 from Equa-
tions (31) and (32). We assume that the phenol scrubbing liquid is virtually
solvent-free when it enters the scrubber, i.e. x = 0. In practice the phenol
may be recycled with a bleed stream being taken to the phenol still. In this
case x / 0 and must be calculated by material balance. If we use a liquid-to-
gas mole ratio (L/G) = 0.05 and all the butyl acetate is absorbed into the
phenol, then the phenol concentration leaving the scrubber is
x = y /(L/G) = 0.041/0.05 = 0.82
1 o
If the total gas pressure drop in the scrubber system is 5 psi (2 psi for the
stripper, 1 psi for the scrubber and 2 psi line losses) , then the pressure at
the bottom of the scrubber must be 15 + 5 = 20 psia (1,034 mm Hg) .
Stage Calculations. The equilibrium vapor concentration is:
From Equation (33b) : PBA = 34.5 j 0.82(1+0.82) J 1*76 = 8.480 mm Hg
From Equation (32) . y.^ = 8.48/(1034 - 8.48) = 8.27 x 10~3
The liquid concentration on the next stage is:
From Equation (29): x2 = (8.27xlO~ - 0.04D/0.05 + 0.82 = 0.165
The stage calculations continue until the vapor is virtually free of solvent,
about yM : 10~8.
H
Stage
1
Bottom
2
3
Etc.
8
9
x
0.820
0.165
2.62xlO~2
5.88xlO~6
1.12xlO~6
y
8.27xlO~3
1.31xlO~3
2.37x!0"4
5.60x!0"8
1.07xlO~8
P psia
20.0
19.9
19.8
19.3
19.2
At 4 ft per theoretical stage, a packing height of 36 ft is required
and the total column height will be about 50 ft.
48
-------
Utility Requirements and Costs. The major energy consumption is for circulat-
ing the gas. The horsepower required to adiabatically compress the gas from 15
to 20 psia is:
WT
HP = 0.00078
V ->]
Mne
where
W = gas mass flow rate
= (106)/18) Ib moles H2Ox( _/18) Ib moles gas/lb mole H^
x 28 Ib/lb mole gas
= 8.64 x 104 Ib/hr
T = inlet temperature = 560°R
M = molecular weight of gas = 28 for nitrogen
n = 0.286 for adiabatic compression
e = overall efficiency =0.7
The ] jrsepower is 577 or 430 kw. If power costs 3£/kw hr, then the com-
pression costs are $12.9/hr or $0.11/10 gal wastewater treated. The installed
cost of the compressor could amount to $10 .
Cooling water and liquid pumping costs may be determined following methods
given in Section 7.- but they are negligible relative to compression costs.
Equipment Costs. The diameter and cost of the packed columns are found in
Section 7. The 100 ft stripper has a liquid rate of 10 Ib/hr and a gas rate
4
of 8.64 x 10 Ib/hr and requires about a 7' diameter column costing $190,000.
4
The 60 ft scrubber has a liquid rate of 1.45 x 10 Ib/hr and the same gas rate,
requiring a 5"6" diameter column costing $80,000. The installed cost of the
two columns will be about $10 .
Summary of the Costs for Solvent Extraction
These costs, summarized in Table 3-2, provide an idea of the costs of sol-
vent extraction in commercial sized coal conversion plants. They are useful for
comparison with phenol removal by adsorption. We emphasize that approximate
"rule of thumb" cost estimating has been used. In the preparation of a budget
equipment costs supplied by vendors will be more accurate. Significant savings
can be realized when integrating the plant into a coal conversion complex
49
-------
if energy conservation by heat interchange between process streams and waste
heat boilers are used where feasible. Air coolers have higher capital costs
than water cooled exchangers, but their overall costs may be lower for cooling
at temperatures above 140°F and this should be checked.
Equipment costs in the table are the capital charges amortized at 15% per
year and include an additional 4% maintenance charge. Labor and chemical
costs are not included.
TABLE 3-2. APPROXIMATE COST OF PHENOL REMOVAL BY SOLVENT EXTRACTION3
$/103 galb % of Total
Extraction System
Equipment 1.36 30
Solvent Regeneration
Equipment: Solvent still 0.21 5
Phenol still 0.04 1
Utilities: Solvent still 2.13 46
Phenol still 0.32 7
Solvent Recovery
Equipment (including compressor) 0.4 9
Utilities 0.11 2
4.57 100
The costs may be partly offset by the market or energy value of the phenol
that is recovered.
a) These costs are to serve as a guide only. See text.
b) 8,000 stream hours/year.
50
-------
MAKEUP
FEED
Figure 3-15. Phenol removal by resin adsorption.
-------
3.3 Phenol Removal by Adsorption
Dissolved organics may be removed from aqueous streams by adsorption onto
porous resins. The aqueous stream is passed through a bed of high surface area
resin which adsorbs the organic pollutants and retains them by van der Waal
forces. The binding energy is generally less than that for activated carbon
adsorption and resin regeneration may be achieved easily by washing with sol-
vent. The organic may be recovered for reuse or sale, and the solvent can be
recycled. Alternatively, the regenerant stream may be destroyed by combustion,
so recovering the heating value of the regenerant and the organic pollutant.
This treatment has not been tested on coal conversion wastewaters.
The Rohm and Haas Company has developed Amberlite polymeric resins that
19 20 23
are effective in adsorbing phenols from wastewater ' ' . The resins report-
21
edly adsorb higher molecular weight phenols, some of which are not easily
handled in solvent extraction or biological systems, as well as or better than
C H OH. Adsorption efficiency is not impaired by the presence of inorganic
6 5
salts or ammonia, although pH values above -v 8.2 may be a problem. Furthermore,
cost estimates have shown that phenol recovery by adsorption can be profitable
under certain conditions '
A typical process flow sheet is shown in Figure 3-15. The system consists
of an adsorption section, a solvent regeneration section, and a still for
drying the extracted phenols. In addition, a stripping/scrubbing step, not
shown, may be used to remove small quantities of solvent regenerant from the
treated water. While solvent regeneration consumes steam, it also allows for
the recovery of a phenol byproduct. Alternatively, the solvent could be burnt
along with its phenol content, so eliminating the need for steam and the two
distillation columns shown in Figure 3-15. For this to be economical, solvent
costs should not be much higher than the fuel (coal) costs on an energy basis.
3.3.1 Adsorption Bed Design
The bed volume for a given application is determined by the bed loading,
Ibs phenol/ft resin, at the desired leakage or phenol removal efficiency. Bed
loadings for Rohm and Haas Amberlite XAD-4 resin are shown in Figure 3-16(a)
from Reference 22. Integrating the curves yields the mean leakage as a func-
tion of the volume of water treated as shown in Figure 3-16(b). In general,
52
-------
when designing adsorption beds, the greater the permissible leakage for a given
flow, the greater the bed loading, and hence the longer the cycle or loading
time. However, the mean leakage curves shown on Figure 3-16(b) are close to
vertical and overall costs are not strongly dependent on the selected leakage.
It is customary to design for zero or very low (1 mg/1) leakage. For low
leakage:
Bed loading (500 mg/1 influent) = 1 j2 - 1.73 q + 0.435 q2 lb/ft3
where q = bed flow rate, gpm/rt (^ inverse residence time)
(34)
The bed loading increases with increasing effluent concentration as shown
in Figure 3-17. Loadings at concentrations between 500 mg/1 and 8,000 mg/1 can
be estimated by multiplying the value in Equation (34) by
Bed loading factor = 0.03 C ' (35)
w" ^re C = influent concentration, mg/1
Equations (34) and (35) are derived from available data for phenol and
they may not be directly applicable to the range of phenols found in coal
conversion condensates. They show that at a bed flow rate of 0.5 gpm/ft , the
bed loading is 1.2 Ib phenol/ft resin at an influent concentration of 500 mg/1
and 5.9 lb/ft for an 8,000 mg/1 feed. Bed costs would increase by a factor of
^ 3 only for the 16 fold increase in concentration.
When the resin bed is loaded, it must be taken off line and regenerated.
A second bed must then be available. Several beds are usually used in parallel
so that only a small percentage of the total bed volume is off line at a time.
Alternatively, it is possible to store the influent stream during the two-hour
regeneration rather than provide additional bed volume. This is economical
when the storage vessel costs less than the additional bed volume required for
the regeneration step.
The fraction of the total bed volume that is loaded in two hours is:
v = (2/t ) V ft3 (36)
e
53
-------
(a)
•CD VOLWMC3
0.5
Eed loading, Ib phenol/ft3
1.5 2.0
By integration of
Fig. Ha)
30 40
Bed Volumes Treated
60
Figure 3-16. Effect of flow rate on resin capacity.
(a) Instantaneous leakage, and (b) mean leakage, as a function .of volume
treated. Amberlite XAD-4 resin; 500 mg/1 phenol in water stream.
54
-------
n
o
<
CL.
<
o
oo
UJ
o:
10
5.0
2.0
1.0
0.5
200
o x
4
m
X
X
X
X
0 X'
I I I I
500
1,000 2,000
5,000 10,000
INFLUENT CONCENTRATION, mg/1 PHENOL
Figure 3-17. Effect of influent concentration on resin capacity.
Based on data from References 20 and 21 , flow rate = 0.5 gpm/ft
for zero or low leakage.
-------
For continuity this must be the bed volume undergoing regeneration at any
time. In Equation (36) V is the total resin volume required and is given by
V = £ ft3, (37)
q
Q is the influent flow rate, gpm.
t is the loading time and is given by
e
t = V (bed loading)/(phenol rate Ib/hr) hrs. (38)
e
The total number of beds, n, of volume v required 4.3:
n = 1 + V/v (39)
The extra bed is for regeneration.
If we do not provide the extra bed for regeneration, a storage basin of
volume, s, is needed where:
S = Q gpm x 120 min/(7.5 gal/ft3) ft (40)
This volume of liquid will be added to the main stream during the loading
cycle increasing the flow rate from Q to Q(l + 2/t ) gpm, and requiring a
proportional increase in the total bed volume V. If the cost of n-1 beds each
of volume v(l + 2/t ) plus the cost of the storage basin is less than the cost
of n beds (and resin) of volume v, the storage route will be cheaper. Storage
is generally cheaper for low feed rate applications.
The choice of bed volume is an economic one. Large beds and therefore
high residence times mean higher bed loadings and lower regeneration costs.
However, bed and resin costs are high. While decreasing the bed volume reduces
capital costs, it increases steam and cooling water requirements. An optimum
must be found. In Figure 3-18 is shown an optimum range of flows and phenol
concentrations. We have included bed (and storage basin in the low flow case)
capital costs and resins costs ($229/ft3), as well as steam ($2/106 Btu) and
cooling water C12C/10 gal circulated) costs for methanol regeneration. A
steam rate of double the theoretical minimum was used to allow for the need to
56
-------
0.2 0.4 0.6
FLOU RATE THROUGH BED, qpm/ft
O. 8
3
1.0
Figure 3-18.
Bed flow rate (residence time) for minimum overall cost in
phenol separation by resin adsorption.
57
-------
reflux so as to achieve the desired separation. Cost of solvent losses, and
capital requirements for solvent regeneration are not included in Figure 3-18,
because they do not affect the flow rate value for minimum cost. The optimum
flow rate falls in the range 0.3-0.5 gpm/ft , depending on phenol concentration
and throughput.
Bed depths should be at least three feet to ensure effective resin
utilization. Pressure drops are about 1/2 psi per foot of depth. Pumping
costs are less than 0.5C/10 gal.
3.3.2 Resin Regeneration
Solvents such as methanol or acetone may be used to absorb phenol from
the resin. About three bed volumes of solvent are required with 1-1/2 bed
21
volumes being used twice; the net usage is 1-1/2 bed volumes . Water is
entrained in the solvent during regeneration and is used also to displace
solvent from the bed following regeneration. Water normally amounts to between
25% to 30% of the final solvent stream. The solvent is recovered by steam
fractionation, and the water/phenol bottoms are treated to recover phenol.
Solvent is expensive and losses must be minimized. About 0.13 Ib solvent
per cubic foot of resin per regeneration cycle is leached from the bed during
21
the loading cycle . Most of this solvent appears in the first part of the
loading cycle. Solvent losses occur also in the fractionator. Solvent
leaving with the bottoms will probably end up in the treated effluent following
the super loading cycle described in the next section.
Methanol costs about 7C/lb and has a latent heat of 480 Btu/lb. Acetone
is more expensive at 19C/lb but has a lower latent heat of 225 Btu/lb. If
total solvent losses are about 1%, overall costs (steam plus losses) are
similar for the two solvents. If solvent losses exceed 1%, it is cheaper to
use methanol. Typically, the feed to the solvent fractionator will have a
solvent:wateriphenol mole ratio of about 25:15:1 and may be treated as a
binary solvent-water mixture. Standard textbook design procedures may
be used together with vapor-liquid equilibrium data for the selected solvent.
58
-------
Vapor-liquid Equilibrium for Methanol-water System
Data at several pressures for both acetone and methanol are available in
25 26
the literature ' . The methanol data correlates well according to the
27 28
Wilson equation , using Wilson coefficients :
and
W = (v /v )/ exp (103.3/T°K) (41)
m w m
W = (v /v )/ exp (242.7/T°K) (42)
w m w
where the subscript m is for methanol, w for water. The molar volumes may be
obtained for any temperature using:
v = 64.51 - 0.197T + 0.387 x 10~3 T2 (43)
m
and
v = 22.89 - 0.036T + 0.686 x 10 6 T2 (44)
w
In use, a temperature is assumed, the molar volumes and Wilson parameters
calculated using the above equations, and the activity coefficients calculated
using the Wilson equations given in Reference 27. If the sum of the vapor mole
.fractions is not unity, another temperature is selected and the calculations
repeated. The procedure, which is similar to that for the van Laar equation
described in Section 3.2, is illustrated in the example given at the end of
this section.
The number of stages for a given separation is a function of the steam
rate as shown in Figure 3-19 for production of a 99% pure methanol overhead.
In practice the recovery efficiency must be high because any methanol in the
bottoms will probably reach the main wastewater stream via the superload
recycle. On the other hand, the overhead purity need not be high as a small
amount of water in the methanol will not adversely affect resin regeneration.
Minimum steam is about 65 lb/100 lb feed, or about 88 lb/100 Ib methanol. The
theoretical steam for evaporation is about 50 lb/100 lb methanol; the additional
steam required for fractionation provides reflux. Lowe reflux ratios can be
used when a less pure overhead is acceptable. The actual steam rate/column
size used will depend on their costs and the corresponding solvent loss. It
turns out that minimum overall costs can be realized for high methanol recovery
efficiencies (>99.9%) with tall columns at near minimum steam rate. At these
59
-------
99.95 METHANOL RECOVERY
OVERHEAD = 99Z CHjOH
SO 75 10O
STEAH RATE lb/100 Ib feed
125
150
Figure 3-19. Theoretical stages and steam rate for methanol regeneration.
60
-------
conditions the costs for regeneration including capital charges, steam and
cooling water costs, as well as solvent losses amount to $1.30 per 1,000
gals. The solvent loss of about 0.1% is only in the solvent fractionator.
Losses from mixing and leaching from the resin bed may be more significant.
A total methanol loss of 1% would add an additional $0.4/10 gals to the
operating costs.
Phenol Distillation
The bottoms from the solvent fractionation column contain about 30%
phenol by mass. Removal of all of the phenol is expensive because the phenol-
water system forms an azeotrope of 9.4% phenol by mass at atmospheric pressure,
and 4.4% phenol at 5 psia. It is possible to break the azeotrope by a double
24
distillation/decanting system , or by solvent extraction. If we choose
atmospheric distillation, we must recycle a 9.4% phenol solution back to the
adsorption bed. If we use the more expensive vacuum distillation route, we
recycle 4.4% phenol; if we break the azeotrope, there is no need to recycle
the water fraction of the regenerant stream. We have found that because the
bed capacity increases with increasing influent concentration (Figure 3-17),
recycling a concentrated stream of up to 10% phenol after blending it with
the wastewater feed does not require additional bed capacity. In fact initial
blending is not necessary. In the patented Rohm and Haas superloading process ,
the concentrated stream is passed directly through the resin bed at the end
of the loading cycle, and then blended with the influent stream to another
bed. Alternatively, the superloading water may be kept separate and used for
solvent displacement. Since solvent not recovered in the fractionation tower
will inevitably end up in the superloading water, use of the superloading
water for solvent displacement might help reduce solvent losses.
The Rohm and Haas superloading scheme has been included in Figure 3-15.
The solvent fractionator bottoms are first decanted since the phenol rich
portion - about 30%-50% of the bottoms - is distilled in the phenol tower at
atmospheric pressure. The overheads and the lean fraction from the decantor
contain about 10% phenol. They are combined and used to superload the resin
prior to regeneration.
The design of the phenol tower follows the method given in Section 3.1.
61
-------
3.3.3 Example; Design of Resin Adsorption Plant
The plant will be based on the 2,000 gpm wastewater stream containing
6,000 mg/1 phenol that as used for the solvent extraction example (Section
3.2.3).
1. Resin Bed Design
(a) Select optimum bed flow rate. In Figure 3-18 we indicate a rate
of 0.3 gpm/ft ; a value of 0.27 gpm/ft gives a convenient loading time of 8
hrs and will therefore be used.
(b) Calculate bed loading. From Equations (34) and (35) with q =
0.27 and C = 6,000 mg/1:
Bed loading = 6.34 Ib phenol/ft resin
(c) Calculate bed volume. The total resin volume required is given
by Equation (37) :
V = 2,000 gpm/(0.27 gpm/ft ) = 7,408 ft
We will use 7,500 ft .
(d) Calculate loading time. The phenol rate to the resin beds is
2,000 gpm x 8.33 Ib/gal x 6,000 x 10~6 Ib/lb = 100 Ib/min, or 6,000 Ib/hr.
Substituting in Equation (38) gives the loading time:
t = 7,500(ft3) x 6.34 (Ib/ft3)/6,000 (Ib/hr) : 8 hours
(e) Number of beds and costs. From Equation (36) the bed volume
undergoing regeneration at any time is:
V = (2/8) x 7,500 = 1,875 ft3
The selected bed volume should either be 1,875 ft or some whole number
fraction of 1,875 ft . A volume of about 600 ft is convenient; we will
3 3
select one third of 1,875 ft or 625 ft . From Equation (39) the total number
of beds of this volume required is:
n = 3 (1 + 7,500/1,875) = 15 beds
of which three will be undergoing regeneration at any time. The cost of the
skid mounted beds can be estimated from:
62
-------
Bed cost = $4,500 (Bed volume, ft3)0'55
This equation gives a cost of $2.33 x 10 for the 15 beds. Resin cost, at
$225/ft , adds an additional $2.11 x 106.
2. Methanol Distillation
(a) Feed composition to the methanol column. The volume of methanol
used is 1.5 bed volumes per cycle, or 1.5 x 7,500 ft /8 hrs = 1,406 ft /hr.
At a density of 49.3 Ib/ft this amounts to 69.3 x 103 Ib/hr or 2.17 x 103 Ib
mole/hr. The water rate is taken to be 25% by volume of the methanol rate,
that is 351.5 ft /hr, or 21.9 x 103 Ib/hr, or 1.22 x 103 Ib mole/hr. The
phenol rate is the phenol input with the wastewater, plus the phenol recycled
in the superload cycle. Assuming all the water entering the methanol tower
is returned to the resin beds with the lean stream (9.4% phenol by mass) from
the phenol tower decanter (see Figure 3-15), the phenol in the recycle stream
is .094 x 21.9 x 10 = 2,060 Ib/hr. The total phenol throughput is therefore
2,060 + 6,000 = 8,060 Ib/hr or 85.7 Ib mole/hr.
The total feed rate is 99.3 x 10 Ib/hr or 3.48 x 10 Ib mole/hr, and
the mole fraction methanol in the feed is xf = 2.17/3.48 = 0.624.
(b) Specify tower operating conditions. Any methanol not recovered
in the methanol tower will be returned to the resin beds in the superload
stream and may end up in the treated wastewater. Because there is no solvent
recovery system proposed for the treated wastewater, a high methanol recovery
of 99.9% is specified for the methanol distillation. However, the overhead
need not be pure methanol and 95 mole % is probably adequate. We specify x
= 0.95.
The minimum reflux rate is calculated from
Rmin = (XD - yf)/(yf - V
and requires a knowledge of yf, the vapor composition in equilibrium with the
feed. Equilibrium vapor composition calculations proceed as in the phenol
tower design example (Section 3.1.1), but here we use the Wilson equation
parameters W and W given in Equations (41) and (42).
c n w
As before, temperatures are assumed and the resulting mole fraction water
and methanol summed. If the sum exceeds 1.0, a lower temperature is selected;
63
-------
if the sum is less than 1.0, the temperature is increased. At a. temperature
of 160.2°F (71.2°C, 344.1°K) and a total pressure of 15 psi (775 mm Hg) , we
get:
V , Molar volume of methanol (Eq 43) = 42.55 cm /g mole
n 3
V , Molar volume of water (Eq 44) = 18.63 cm / g mole
w
W , Wilson coefficient for methanol (Eq 41) = 0.324
n
W , Wilson coefficient for water (Eq 43) = 1.129
w
Substituting in the Wilson equation given in Reference 27, we get the
activity coefficient for methanol to be 1.065, and that for water to be 1.357.
The vapor pressure of water is calculated from Equation (20) to be 246 mm Hg.
The vapor pressure of methanol is given by
log.,, p* _ . =7.879- 1473/(230 + t°C)
^10 ^methanol
which gives a value of 974 mm Hg. Substitution in the activity coefficient
equations (Equation 14) gives
y ., = 1.065 x 974 x .624/775 = 0.835
•methanol
y . 1.357 x 246 x (1 - .624)/775 = .162
water
The minimum reflux rates can now be calculated :
Rmin = ('95 ~ -835)/(.835 - .624) = 0.545
A value of R=l will be used for this example.
(c) Calculate liquid and vapor flow rates. Based on 100 Ib moles/hr
feed, we get by material balance:
Product (distillate) rate:
DxD = EFxf or D = .995 x 100 x .624/.9S = 65.'36 Ib moles/hr
Bottoms rate:
B = 100 - D = 100 - 65.36 = 34.64 Ib moles/hr
Vapor rate:
V = (R+1)D = 2 x 65.36 = 130.72 Ib moles/hr
64
-------
Liquid rate above feed:
L = V - D = 130.72 - 65.36 = 65.36 Ib mole/hr
(<3) Stagewise calculations. The stagewise calculations are made
as for the butyl acetate distillation example in Section 3.2.3. The results
are summarized below.
st
1
2
3
4
5
6
7
8
9
10
:age x
_3
(bottom) 1.84x10
1.20xlO~2
-2
6.96x10
0.272
0.507
0.620
0.663
The feed is
0.756
0.839
0.910
y
_2
1.46x10
8.71xlO~2
0.342
0.636
0.780
0.834
0.853
introduced on Plate
0.894
0.930
0.961
t°F
221
217
202
181
170
165
163
7
159
156
153
P psia
18
17.7
17.4
17.1
16.8
16.5
16.2
15.9
15.6
15.3
The vapor leaving plate 10 exceeds the product specification; plate 10
is therefore the top plate.
(e) Utility requirements and costs. The steam rate is calculated as
in the example given in Section 3.2.3. A rate of 82 x.10 Ib steam/hr is re-
quired for reboil and preheat. At $2.5/10 Btu, the steam-cost is $1.8/10 gal
wastewater treated. The cooling water required to condense the reflux and
subcool the product is about 5,700 gpm, which at a cost of 12C/10 gal cooling
water circulated amounts to $0.34/10 gal wastewater treated.
(f) Equipment size and cost. Using the equations given in Section 7
and following procedures outlined in previous examples, a tower diameter
of 9 ft and a height of 60 ft is obtained at an installed cost of $642 x 10 .
Cost of the total condenser and reboiler will increase the total cost of the
methanol regeneration to about $1.1 x 10 .
65
-------
3. Phenol Distillation
(a) Feed Composition. From the methanol tower calculations we have
Phenol rate = 8,060 Ib/hr = 85.7 Ib mole/hr
Water rate = 21,900 Ib/hr = 1,220 Ib mole/hr
Total feed = 29,960 Ib/hr = 1,306 Ib mole/hr
Water mole fraction in feed = 1,220/1,306 = 0.934
(b) Tower design. Design procedures for phenol distillation columns
were presented in Section 3.1.1 and 3.2.2 and will not be repeated here. Five
theoretical stages are required resulting in a 2'6" diameter 49 ft tall column
with an installed cost of $90 x 10 . Heat exchangers will boost the equipment
cost to $170 x 10 . Steam and cooling water costs are relatively small,
amounting to $0.09 and $0.02 for each 1,000 gallons wastewater treated.
4. Summary of Costs for Resin Adsorption
The costs estimated in this example are summarized in Table 3-3 and are
intended primarily for comparison with costs given in Table 3-2 for solvent
extraction. They are not intended to represent an absolute cost for phenol
removal. The tables show that the cost of phenol removal by both solvent
extraction and resin adsorption fall in the range $4 to $5 per 1,000 gallons
wastewater treated, the slightly higher cost for solvent extraction being
attributable to the solvent recovery system.
A study of the distribution of costs for the two systems can indicate
where improved design may lead to significant cost savings. For instance, both
systems have high (^50%) utility (energy) costs, and application of heat
40
recovery techniques or selection of solvent having lower latent heats should
be considered.
The solvent extraction system has a high capital cost which can be reduced
by decreasing the number of extraction stages. A reduction from six to two
stages will reduce capital costs by 67% and it will now be economical to reduce
the solvent/water flow ratio resulting in a proportional reduction in the
utility costs for solvent regeneration. Of course, the penalty for the reduc-
tion in cost is a reduction in phenol removal efficiency as shown by Equation
(25). However, the solvent extraction system may be followed by a resin
adsorption system. Because the phenol load on the resin is now considerably
reduced, regeneration frequency and consequently utility costs are reduced.
66
-------
We have found that overall costs for a hybrid system may, for some wastewaters,
be lower than the costs for either single system accomplishing the same job.
Apart from economical advantages, the hybrid system may result in a cleaner
effluent because pollutants not removed in the one system may be removed in the
other. Further, a single solvent recovery system will serve to remove both
solvents from the treated water. A small fractionation system may be added to
separate the solvents and reduce the costs for solvent losses shown in Table
3-3.
TABLE 3-3. APPROXIMATE COST OF PHENOL REMOVAL BY RESIN ADSORPTION
$/103 gal % of Total
Adsorption System
Equipment 0.46 11
Resin, initial load 0.32 8
Resin, replacement (5 yr life) 0.43 10
Solvent Regeneration
Equipment 0.22 5
Utilities 2.14 52
Methanol makeup (1% loss, 7C/lb) 0.40 10
Phenol Distillation
Equipment 0.03 1
Utilities 0.11 33
Total Cost 4.11 100
REFERENCES - Section 3
1. Sinor, J.E., Ed., "Evaluation of Background Data Relating to New Source
Performance Standards for Lurgi Gasification," Interagency Energy-
Environment Research and Development Program Report, Environmental
Protection Agency, EPA-600/7-77-057, June 1977.
2. Goldstein, D.J., and Probstein, R.F., "Water Requirements for an
Integrated SNG Plant and Mine Operation," Environmental Protection
Technology Series, EPA - 600/2-76-149, June 1976, Symposium Proceedings:
Environmental Aspects of Fuel Conversion Technology, II, Dec. 1975,
Hollywood, Florida.
3. Earhart, J.P., Won, K.W., Wong, H.U., Prausnitz, J.M. and King, C.J.,
"Waste Recovery: Recovery of Organic Pollutants via Solvent Extraction,"
Chem. Eng. Prog., 73, (No. 5) 67, 1977.
4. Earhart, J.P., Won, K.W., Wong, H.U., Prausnitz, J.M. and King, C.J.,
"Removal of Phenolics from Industrial Wastewater by Dual-Solvent Extrac-
tion," Chem. Eng. Prog., 73, (No. 5) 67, 1977.
67
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5. Beychok, H.R. , Aqueous Wastes from Petroleum and Petro-chemical Plants,
John Wiley & Sons, London 1967.
6. Kiezyk, P.R. and Mackay, D. , "Wastewater Treatment by Solvent Extraction,"
Can. J_. of_ Chem. Eng. , 49, 747, Dec. 1971.
7. Lauer, F.C., Littlewood, E.J. and Butler, J.J., "Solvent Extraction
Process for Phenols Recovery from Coke Plant Aqueous Waste," Chem. Pro.
Equipment Corp., Iron and Steel Engineer, Reprint.
8. Medir, M. and MacKay. D. , "Extraction of Phenol from Water with Mixed
Solvents," Can. J. of Chem. Eng. , 53, 274, 1975.
9. Priestley, James J. , "Treatment of Effluent Liquors from Gasworks," Gas
J., 260. 532, 594, 659, 665, 1949.
10. Zogorski, J.S. and Faust, S.D., "Removing Phenols via Activated Carbon,"
Chem. Eng. Prog. , 73, 65, 1977.
11. Fox, C.R., "Remove and Recover Phenol," Hydrocarbon Processing, 54, 109,
July 1975.
12. Winter, T.H., Fox, R.D. and Himmelstein, K.S., "Economic Evaluation of
Phenolic Waste Treatment Systems," presented at 1973 International Water
Conference, Institute of Engineers of Western Pennsylvania.
13. Lauer, F.C., Littlewood, E.J. and Butler, J.J., "Solvent Extraction
Process for Phenols Recovery from Coke Plant Aqueous Waste," Iron and
Steel Engineer. Presented at Eastern States Blast Furnace and Coke Oven
Assoc. meeting, Pittsburgh, Pa., February 1969.
14. Pollaert, T. , Manager Coal Gasification, American Lurgi Corporation.
Private communication, Dec. 22, 1977.
15. Verhoff, F.H., Ross, S.L. and Curl, R.L., "Breakage and Coalescence
Processes in an Agitated Dispersion Experimental System and Data Reduc-
tion," Ind. Eng. Chem. Fundanu , 16, 371, 1977.
16. Ross, S.L. , Verhoff, F.H. and Curl, R.L. , Ibid. "2. Measurement and Interpre-
tation of Mixing Experiments," Ind. Eng. Chem. Fundam., 17, 101, 1978.
17. Hanson, C. ed. , Recent Advances in Liquid-liquid Extraction, Pergamon
Press, Elmsford, NY, 1975.
18. Weller, R. , Schuberth, H. and Leibnitz, E. , "The Vapor Liquid Phase
Equilibria for the System Phenol/n-Butylacetate/Water at 44.4°C," Journal
ftir Praktische Chemie, 21, 234, 1963.
19. Fox, C.R., "Remove and Recover Phenol," Hydrocarbon Processing, 54, 109,
July 1975.
20. Rohm and Haas Company, "Process for the Recovery of Phenolics from Aqueous
Streams," Form 20B, August 1977.
68
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21. Fox, C.R., Rohm and Haas Company, private communication, March 1978.
22. Simpson, R.M., "The Separation of Organic Chemicals from Water," paper
presented at the Third International Symposium of the Institute of
Advanced Sanitation Research, April 1972.
23. U.S. Patent 3,979,287 (Rohm and Haas Company).
24. Robinson, C.S. and Gilliland,E. R. , Elements of Fractional Distillation,
McGraw-Hill, 1950.
25. Othmer, D.F. and Benenati, R.F., "Composition of Vapors from Boiling
Binary Solutions," Ind. Eng. Chem., 37, 299, 1945.
26. Othmer, D.F. and Morley, J.R., "Composition of Vapors from Boiling Binary
Solutions," Ind. Eng. Chem., 38, 751, 1946.
27. Perry, R.H., Chilton, C.H., eds., Chemical Engineers' Handbook, 5th
edition, 13-8, 1973.
28. Holmes, M.J. and Van Winkle, M., "Prediction of Ternary Vapor Liquid
Equilibria from Binary Data," Ind. Eng. Chem., 62(1), 21, 1970.
29. Horsley, L.H., "Azeotrope Data III," Advances in Chemistry Series 116.
American Chemical Society, Washington, D.C., 1973.
30. Sims, H.D., "Relative Volatility of Binary Mixtures at Low Pressures,"
MIT ScD Chem. Eng., 1933.
31. Rhodes, F.H., Wells, J.H. and Murray, G. W., "Vapor Composition Relation-
ships in the Systems Phenol-water and Phenol-cresol," Ind. Eng. Chem.
17(11), 1199, 1925.
32. Bogart, M.J.P. and Brunjes, A.S., "Distillation of Phenolic Brines," Chem.
Eng. Prog. 44(2), 95, 1948.
-* **- •
33. Goldstein, D.J. and Yung, D., "Water Conservation and Pollution Control in
Coal Conversion Processes," U.S. Environmental Protection Agency Report
EPA/600/7-77-065, 1977. NTIS catalog PB-269-568/2WE.
34. 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.
35. El Paso Natural Gas Co. , "Second Supplement to Application for a Certifi-
cate of Public Convenience and Necessity," Docket CP 73-313, October 1973.
36. Van Winkle, M., Distillation, McGraw-Hill, 1967.
37. Heck, R.P-, "Munitions Plant Uses Adsorption in Wastewater Treatment,"
Industrial jastes, 35, March/April 1978.
69
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38. Probstein, R.F. and Gold, H., "Water in Synthetic Fuel Production,"
MIT Press, Cambridge, Mass., 1978.
39. Wurm, H.J., "Phenol Removal from Coke Oven Condensates by the Phenosolvan
Process," in German, Gluckauf 104(12) 517, 1968.
40. Goosens, G. , "Principles Govern Energy Recovery." Hydrocarbon Processing,
57(8) 133, 1978.
70
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4. AMMONIA SEPARATION AND RECOVERY
4. 1 Introduction
Ammonia is formed in coal conversion reactors by the combination of
nitrogen and hydrogen. Nitrogen is introduced to the reactors in combined
form with the coal which may contain up to 2% nitrogen by mass. The propor-
tion of this nitrogen that is converted to ammonia will depend on reactor
conditions; up to 60% conversion has been reported for the Syn thane process
Large concentrations of atmospheric nitrogen are present in airblown
gasifiers and may react
(1)
The rate of reaction (1) is slow, but coal constituents could conceivably act
as catalysts. Equilibrium is shifted to the right by decreasing temperature
and increasing pressure.
Gaseous ammonia leaving the reactors is absorbed in the downstream quench
step and removed with the dirty condensate waters. In low temperature gasifiers,
e.g. Hygas, Lurgi, the resulting concentration of ammonia in the process
condensate may range from 3 , 600 mg/1 to 14 , 000 mg/1 or more . These concentra-
tions are considerably higher than those found in common wastewaters (for
example, ammonia concentrations in raw sewage are of the order of 35 mg/1 and
far in excess of those permissible for effluent discharge. High temperature
gasifiers such as the Texaco and Koppers Totzek units are relatively clean
burning and their process condensates may not need treatment for ammonia
removal .
Since current trends are to reuse water and to minimize or possibly
eliminate its discharge, we will briefly review the consequences of the
presence of ammonia in the process condensate.
A major reuse proposed for process condensate is cooling. Ammonia in the
cooling tower circuit provides nutrients for biological growth, leads to
inactivation of zinc corrosion inhibitors, and consumes chlorine. It acceler-
ates the corrosion of copper, so copper tubing should not be used. Ammonia is
stripped from the water as it cascades down the cooling tower, and will there-
fore be present in the cooling tower plume. The threshold odor limit for
2
ammonia has been stated to be 21.4 ppm . If all the ammonia entering a
cooling tower is stripped, then the ammonia concentration in the plume will be
about one twenty-fifth that of the ammonia concentration in the makeup
71
-------
water (cooling tower operating with liquid to gas flow ratio of two to one, 2%
of circulating water evaporated per cycle, about 10 cycles of concentration).
The ammonia concentration in the makeup water must, therefore, be limited to
below about 500 ppm to avoid exceeding the threshold concentration for odor.
The ammonia concentration in the drift will be essentially the same as in the
circulating water. Recommended values for ammonia concentration in the
3 12
circulating cooling water of 75 ppm and 10 ppm have been reported. A full
discussion of cooling water control is given in Section 6.
In some plants process condensate is recycled back to the process for
slurrying the coal. Although a very clean water is probably not required,
ammonia and other constituents may have to be removed to prevent their buildup
in the recycle stream. As in the cooling water qase, 100% ammonia removal is
not necessary; it is sufficient merely to maintain an acceptable concentration
level. This level must be determined.
Phenol is removed either by solvent extraction, resin adsorption, or by
biological oxidation. Where the biological process is used, ammonia concen-
4
trations have to be reduced to below about 1,800 ppm to prevent inhibition of
the oxidation reactions. However, some ammonia is essential to provide
nutrients; normally a phenol to N ratio of 14 to 1 is used.
In the case of solvent extraction, ammonia may or may not interfere,
depending on the solvent used. On the other hand, if phenol is present in the
ammonia stripper, some fraction of it will be stripped and may cause problems
in downstream processing. For example, if the phenol is removed with the acid
gases, it can accelerate deactivation of catalysts used in sulfur recovery,
e.g. Claus processes. It is possible to hold down the phenols by refluxing
the ammonia stripper, but this is expensive and would not normally be considered
unless reflux was required for other reasons. Removal of phenol upstream of
ammonia recovery may, therefore, be preferred. However, depending on its
concentration, phenol removal by biological oxidation, which must follow the
ammonia recovery step, may be cheaper. The final system choice will depend on
the overall economics which must be determined for each' specific case.
4.2 Ammonia Recovery Process
There are two major considerations. First, we do not wish to go from a
water pollution to an air pollution or other disposal problem; second, ammonia
72
-------
removal is costly and we should try to recover our costs. Both factors dictate
that we recover ammonia in a form suitable for sale, preferably as anhydrous
ammonia or possibly as ammonium sulfate. In addition to ammonia, process
condensates contain phenols and the volatile acid gases CO and H S. While
removal of ammonia from water by stripping is straightforward, separation of
ammonia from carbon dioxide and hydrogen sulfide to obtain saleable ammonia is
not straightforward; this defines the problem.
While there is published experience and design procedures for the treat-
ment of sour waters , that is, waters containing ammonia and H S, we have
found no published procedures relevant to the stripping of water containing
ammonia, phenols, CO and H S. There are proprietary design procedures for
18
strippers treating aqueous feeds containing up to 30 component species and
they require the use of large digital computers. Our objective was to develop
a relatively rapid calculational procedure suitable for desk top checking of
demonstration plant designs. While we could not retain the versatility and
detailed ac .uracy of the large programs, our procedure does meet the require-
ments of determining, within design accuracy, the utility requirements and
stripper size for a range of process variables. To our knowledge, such
procedures have not previously been published and we have described our design
procedures here in some detail.
Two main possibilities exist for the stripping of the gases from the
dirty condensate. Ammonia and the acid gases may be stripped off simultan-
eously with ammonia separated and recovered from the vapor. Or the acid gases
may be stripped off with the ammonia held down; 'the ammonia is then concentrated
in a separate stage. Ammonia may be held down by a water wash or a combination
of wash and reflux. The stripper configurations that might arise are shown in
Figure 4-1 (non-refluxed strippers) and Figure 4-2 (refluxed strippers).
Reflux is used to concentrate the ammonia in the overheads when stripping a
dilute (less than about 500 ppm NH ) feed, or it may be used to hold back
phenol, or in combination with a water wash to hold back ammonia.
It is useful to describe the design procedures for the stripping stage in
the ammonia recovery schemes in the most general terms. In Section 4.3
73
-------
Water +
NH,, and
Volatile acids
BOTTOMS
Water +
some NH,
TO COOLING TOWER
(A) STRIPPING
DISTILLATE (vapor)
Water +
NH3, and
Volatile acids
To Phosam-W,
Ammonium sulfate
or other process
for separation and
recovery of ammonia
STEAM
WASH
DISTILLATE
Water +
FEED
Water + NH3, and "
Volatile acids
BOTTOMS rf
*^^^ ^S
Volatile acids (C02 + H2
TO DESULFURIZATION
STFAM
V/9t"r +
NH., |TO AMMONIA CONCENTRATION
(B) STRIPPING WITH WASH TO HOLD DOWN AMMONIA
Figure 4-1. First stage in ammonia recovery; stripping without reflux.
74
-------
(A) STRIPPING WITH REFLUX TO INCREASE AMMONIA CONCENTRATION IN OVERHEAD
CW
PARTIAL
CONDENSE
•^DISTILLATE (vapor)
Water + ammonia, volatile acids,
and some phenol
FEED ;
Water +
NH
volatile acids,
and phenols
BOTTOMS
Water + -<-
phenols and
some NH3 )TO BIOTREATMENTl
STEAM
WASH
Water
FEED
Water + MH3,
volatile acids,
and phenols
BOTTOMS
Water + NH
To Phosam-W
or ammonium sulfate,
or other process for
separation and recovery
of ammonia from acids
(B) STRIPPING WITH WASH TO HOLD DOWN AMMONIA
CW
DISTILLATE (liquid)
Water, and volatile
acids (C02, H2S)
fro OESULFURIZATIOH)
STEAM
and
3
phenols
TO AMMONIA/PHENOL SEPARATION AMD AflMONIA
CONCENTRATION
Figure 4-2. First stage in ammonia recovery; refluxed strippers
75
-------
WASH WATER
W moles/hr
DISTILLATE y]
moles/hr
(Partial Condenser)
DISTILLATE, x
r- D moles/hr
CCIIIY (Total Condenser)
Lr LUA
C moles/hr
FEED, x.
F moles/hr
BOTTOMS, Xg
B moles/hr
-^ \/ f^j / / /
REBOIL STEAM
SRmoles/hr
1 1
I vn
1 Ln-] t 1
» yn+i
lLn \ I
///////
///////
Tray n-1
Tray n
Tray n+1
LIVE STEAM
<; mnlp<;/hr
"T>~f*~~f~~7*~7^)k* ^ ^~
/ / /// n
'//////<
< ( / ' if Fl fl^H UATFR
*^ r LMon , WttliLK
A moles/hr
Figure 4-3. stripping tower nomenclature.
76
-------
a set of design equations incorporating all possible streams: feed, wash,
reflux, hot water flash, etc. is developed. Those streams not required for a
specific design are simply assigned zero values. Because of its generality,
the procedure may be used for ammonia fractionation as well as stripping, and
is applicable to the case where phenol is present and must be held down. The
one limitation is that the aqueous ammonia product must not exceed 30% by
mass. The production of more concentrated ammonia requires special treatment
*
to account for departures from ideality with respect to both vapor-liquid
equilibrium and heat effects. This is done in Section 4.4.
4. 3 Stripping Tower Design
A typical stripping tower is shown in Figure 4-3. The condenser may be
either a partial or total condenser, or may not be used. If a partial conden-
ser is used, it is considered to be the first equilibrium stage in the column.
In this case, and for the case of no condenser, material balance above the
feed yields
W + V , = Vn + L ...
n+1 1 n (2)
which, for any component becomes
V.y. + L x
_ 1 1 n n ,_.
Yn+l ~ Vn + L - W
1 n
x and y are mole fractions; the nomenclature is given on Figure 4-3.
Below the feed the equations become
W + V , + F = V, + L (4)
n+1 1 n
and
V,y, + L x -Fx_
11 n n f , ,-N
(->)
V, + L - W - F
1 n
77
-------
In the case of the total condenser, we must replace V and y by D and x
respectively. However, the distillate or liquid concentration from the total
condenser is identical to the vapor concentration entering the condenser which
is y , the vapor concentration leaving the top plate.
In general the material balances become
W + V , = X, + L - F (6)
n+1 1 n
and
L x -Fx_
n n f
n+l ~ X, + L -W-F
1 n
where
X = D for total condenser
X = V for partial or no condenser
F = 0 above the feed plate.
We have chosen to treat the feed as a saturated liquid.
To being the calculations some conditions must be specified. We will
specify the steam, flash and wash rates, the pressure at the top of the
tower, and the type of condenser and reflux rates. If we are stripping, we
must specify the amount of ammonia to be removed; if deacidifying, that is,
removing only the acid gases, the amount of CO- to be removed must be
specified.
We could not develop, from published vapor-liquid equilibrium data, a
means of determining the complete overhead vapor composition from a knowledge
of the pressure, temperature and the water vapor and ammonia (or CO7) concen-
trations. It is therefore necessary to estimate removal efficiencies of three
of the four components (the fourth component removal efficiency is specified)
and to check the estimates by mass balance. Generally CO is stripped in
excess of 99% for 90% NH3 removal. H S may or may not lead NH stripping. If
CO2 is present, the H2S is held down and under some conditions may concentrate
in the tower until the CO is stripped. Relative removal efficiencies are
determined by their relative initial concentrations, temperature (pressure),
78
-------
steam rate, and whether or not reflux is used. The amount of water vapor
leaving with the stripped gases is determined by heat balance.
It is assumed that the concentrations are everywhere sufficiently low
that the enthalpies of the streams are those of pure water, and that the molal
overflow rates based on the water component are equal. This assumption is
probably not applicable to the vapor leaving the top tray, so in doing the
Heat balance only the water vapor component of this stream is considered.
The temperature everywhere is taken to be that equivalent to the water
vapor pressure
. _ 7036.3
" 384 (8)
14.49 - InP
v
where P is the vapor pressure of water, psia.
This equation, based on a Cox type plot, has an accuracy of better than
in the range 5O - t°F - 400.
The overall balances for the tower, in the case of no condenser, are
Material: F + ST + A + W = B + v, (9)
L J.
Thermal:
FHf + SR(Hv,B - HifB} + SLHv,B
(10)
where H and H. are the molar vapor and liquid enthalpies of water at the
v,n £,n
conditions of stream n.
79
-------
For the case of a partial condenser,
Material: F + S, + A + W = B+V, (11)
Thermal: FHf + SR(HV/B
BH£,B + VlHv,lYl,H20 + C(Hv - Ei 1) (12)
For the case of a total condenser,
Material: F + S+ W + A = B + D (13)
Thermal: FHf
= BHn „ + V,H ny, n - CHn , (14)
A,B 1 Vjl-'ljH 0 1,1
For the general case, these equations may be written:
Material: F + S + W + A = B + X (15)
F(H.= - H. „) + (S + S ) (H - Hn ) + W(Hn , - Hn )
Thermal: V, = —^ ^ RH L HV'B *il *ii ^B_
1 Yl,H20Hv,l - HJl,B + X2
(16)
Where V is the vapor from the top equilibrium stage,
X = R = 0 for no condenser
X = R(H - H ) for partial condenser
^ V , X J6 , J.
X = :r-7-(H _ - H ) for total condenser
£. K+ X J6 , 13 J6 , J.
(2
and R,the reflux ratic, is defined R = — (17)
Xl
80
-------
The feed must be introduced at the top if there is no condenser (R=0)
and no wash (W=0) . Otherwise, it may be introduced at a point where it im-
proves the desired separation. For stripping, this will be where it increases
the phenol to ammonia liquid concentration ratio; for deacidifying it will be
where it increases the ammonia to CO concentration ratio. Both the above
feei and below feed material balance equations, Equations (6) and (7) must
be used until the feed plate is found.
The feed enthalpy is that for the top plate when the feed is at the top
(R and W = 0) . Otherwise the feed enthalpy is not known in advance but may be
estimated as being close to that of the top plate for stripping, or close to
that of the reboiler if deacidifying.
The plate-by-plate calculations are begun by assuming a top temperature,
p
calculating the water vapor concentration y = v using Equation (8) ,
2 P
and then determining the top vapor rate using Equation (16) . The temperature
assumption is now checked by calculating the vapor composition from the top
equilibri -in stage using the specified removal efficiencies n.:
n . Fx _ .
The correct temperature is that for which Zy. = 1.
Once the top vapor is known, the equilibrium liquid composition is
determined using the vapor-liquid equilibrium relationships presented in
Section 4.3.1. The composition of the vapor from the next plate can then be
calculated using Equation (7) for each component. The pressure on each
successive plate is taken to increase by 0.5 psi, and the temperature of each
plate is calculated using Equation (8) . The calculations are repeated until
the liquid composition satisfies one of the specified removal efficiencies.
When stripping, the ammonia concentration is checked; when deacidifying, the
CO concentration is checked.
81
-------
4.3.1 vapor-liquid Equilibria for HO - NH - CO2 - H2S -C^H^OU System.
Van Krevelen et al have published vapor-liquid equilibrium for the HO -
NH - CO - H-S system; we have been unable to find data for this system in
the presence of phenols. It is sometimes reported that phenol, being an acid,
19
ties up the ammonia . However, phenol has an lonization constant nearly
three orders of magnitude less than that for H S , so we have made the
assumption that the presence of phenol does not significantly affect van
Krevelen's data.
Apart from phenol, there are eight species to consider: NH , NH ,
NH C00~, CO , HCO , CO ~, H S and HS . The total ammonia concentration A,
the total CO concentration C, and the total H S concentration S, in
moles/litre are given by:
A = (NH3) + (NH4+) + (NH2COO~) (19)
C = (C02) + (HC03) + (C03) + (NH2COO~) (20)
S = (H2S) + (HS~) (21)
Generally, (CO ) and (H S) are negligible.
^ £•
The relationships given by van Krevelen governing the concentration of
the individual components are
The ionic balance :
(NH4+) = (HC03~) + 2(C03=) + (NH2COO~) + (HS~) (22)
The ionic equilibrium relationships :
(HCO)
(NH2COO~)
K = - — (24)
(NH3) (HC03 )
82
-------
_ (NH ) (CO )
K3 ~ - ^— (25)
(NH3)(HC03 )
(NH )(HS~)
and Henry's law for ammonia:
(NH )
-- ^ Hg (27)
The "constants" K, to K. and H vary with system temperature and, in
1 4 NH
some cases, composition. Using data presented by van Krevelen, we obtain
Antoine type correlations:
Kl - ex* toF + 1605.5 -'8.64 InC ' 75'63 > (28)
\t°F + 557.5
- 6.704 (29)
-19.96
a +0.089S+a C
= 10 (31)
where a = - 0.168 (t°C)°'627
1 147
and a = 0.0029 (t°C) '
0.025(NH3)
(32)
83
-------
where H = exp - 16.54 ) (33)
0 T°K
K , K , H as determined above are to be used with partial pressures
1 4 NH
expressed in mm Hg.
Van Krevelen developed a method for calculating the equilibrium vapor
composition from a known liquid. In our work we start off with a specified
vapor at the top of the stripping column and require the equilibrium
liquid composition. This may be done as follows:
1) The free ammonia concentration is first determined from a knowledge
of p and Equations (27) and (32) . Some iteration is required because the
Henry coefficient is slightly dependent on (NH3).
2) A value for (HCO ) is then assumed and used to calculate (NH COO )
O £•
from Equation (24).
3) C may now be calculated using Equation (20); as a first approximation
(CO ) is taken to be zero. A better estimation of (CO ) is obtained using
•J J
Equation (25) with Equation (30) after finding the (NIK) concentration from
Equation (23) and Equation (28) .
4) Step 3 is repeated until the calculated (CO ) value stabilizes.
5a) (HS ) is calculated from the ionic balance Equation (22) and used
to calculate p using Equation (26) and Equation (31) . Steps 2-5 are
2
repeated until the assumed (HCO ) results in the known p
2
5b) In the case p is zero, (HS ) is zero and the assumed (HCO_ ) value
us 3
may be checked against one calculated from Equation (22)
The procedure is not suitable for systems containing H S in the absence of
CO . In these cases use the following procedure:
1) The free ammonia concentration is determined.
2) A value of K is calculated from Equation (31) with C = 0. As a
first try, S = 0.
3) A better value for S is obtained using Equation (26). Note that
for this system (NH ) = (HS~) = S.
4) Steps 2 and 3 are repeated until the value of S stabilizes. This
fixes (NH ) and hence A.
84
-------
The above calculations are correct if fixed ammonia is not present; that
is, if there is no mineral acidity and so no NH Cl or (NH ) SO . If fixed
ammonia is present, the calculations are approximately correct for that part
of the ammonia which is free.
If phenol is present, its vapor pressure must be accounted for. Available
data for phenol water systems (see Section 3) suggest that in the concen-
tration range of interest the vapor-liquid equilibria for the phenol-water
system is only slightly pressure dependent and may be adequately represented
by:
y = 1.67 x 15 - P psia - 60 (34)
It is not known to what extent the presence of the ammonia and dissolved acid
gases affect this equilibrium relationship. Equation (34) should nevertheless
be used until information for the multicomponent system becomes available.
Note that the condensates under consideration have phenol concentrations of
less than "" % by mass so the phenol-water azeotrope (10% phenol by mass) is not
a problem.
Example of Equilibrium Calculations
The liquid composition A = 1.189, C = 0.41 and S = 0.189 presented in
Table 8 of van Krevelen's paper is used as a basis of the exemplary
calculations. All concentrations are in g moles/1. Van Krevelen's procedure
is used to calculate the equilibrium vapor composition:
1. Assume a value for the bicarbonate composition. As a first guess,
use a value about 40% of the total CO concentration; try (HCO ) = 0.17.
2. Calculate (NH3). From Equations (19), (20) and (21):
(NH ) = (HCO~) + A - 2C - S
= 0.17 + 1.189 - 2 x 0.41 - 0.189
= 0.35
3. Calculate (CO~). Using Equations (20), (22), (25), and assuming
(CO ) Z 0 and that (CO ) is negligible relative to C:
£ -3
(HCO~)/(C + S)
85
-------
From Equation (30) the value of K at 20°C (68°F) is 0.140. Substituting:
(CO ) = 0.140 x 0.35 x 0.17/(0.41 + 0.189)
= 0.014
4. Calculate (NH*). From Equations (20) and (22)
(NH*) = C + S + (CCf)
= 0.41 + 0.189 + 0.014
= 0.613
5. Calculate (NH COO ) using Equations (24) and (29) . From Equation
(29) the value of the equilibrium constant is K = 3.395. Substituting in
Equation (24):
COO ) = K2 (NH3)(HC03)
= 3.395 x 0.35 x 0.17
= 0.202
6. Check the ammonia balance using Equation (19):
A = (NH3) + (NH*) + (NH2COO~) (19)
= 0.35 + 0.613 + 0.202
= 1.165
But we know A = 1.187. Since the calculated ammonia balance is too low,
we must increase the assumed bicarbonate concentration. Repeating the calcu-
lations: 1. Assume (HCO ) = 0.178
2. (NH3) = 0.358
3. (C0~) = 0.015
4. (NH4) = 0.614
5. (NH2COO~) = 0.217
86
-------
The ammonia balance now gives
A = 0.358 + 0.614 + 0.217
= 1.189
which coincides with the given liquid composition. The gas composition can
now be calculated:
7. Determine the ammonia partial pressure. From Equation (33) : H =
0.0929, and from Equation (32): H = 0.0948. Substituting in Equation (27)
NH_
p ' = 0.358/0.0948 = 3.776 mm Hg
8. Determine the CO partial pressure. From Equation (28): K = 0.148.
Substituting in Equation (23) :
Pco = (HCO~)
= (0.178) (0.614)7(0.148) (0.358)
= 2.066 mm Hg
9. Determine the H S partial pressure. From Equation (31): K = 0.090.
Substituting in Equation (26) :
PH s = (NH+)(S)/(K4)(NH3)
2 = (0.614) (0.189)/(0.090) (0.358)
= 3.602 mm Hg
As an exercise in determining equilibrium liquid compositions from a
given vapor, the three partial pressures determined above will be used.
1. Calculate the free ammonia concentration. From Equation (33) : H =
0.0929. Assuming H = H , Equation (27) gives:
3 ° i
/
3 ?MS '
4tf ^
= 3.776 x 0.0929 ± 0.351 f','4 ——'—•
Hfi!.^
This value of (NH ) is used in Equation (32) to calculate a more correct
value of H to get H = 0.0948, which in turn gives (NH ) = 0.358.
Further iterations are not necessary.
87
-------
2. Calculate (NH COO~). A value of (HCO ) is first assumed; we select
(HCO~) = 0.178. Substituting in Equation (24) :
(NH2COO ) = K2 (NH3)(HC03)
= 3.395 x 0.358 x 0.178
= 0.216
3. Calculate the total CO2 concentration.
C = (C02) + (HCO~) + (C03) + NH2COO~)
= 0 + 0.178 + (CO3) + 0.216
i.e. C = 0.394 + (CO3)
To calculate (CO~) use Equation (25) , but must first estimate K and
(NH ) - To calculate K from Equation (28) we require a value for C which is
not yet accurately determined. As a first assumption use C = 0.394. This
gives K. = 0.146 so that from Equation (23):
(NH*) = K^ (NH ) p /(HCO )
2
= 0.146 x 0.358 x 2.066/0.178
= 0.605
Substituting in Equation (25) :
(C03)
0.140 x 0.358 x 0.178/0.605
0.015
4. This value of (CO ) is used to calculate a corrected value of C and
the process is repeated until the calculated C concentration stabilizes. The
first two iterations are:
Iteration
1
2
C
0.394
0.409
Ki
0.146
0.148
(NH*)
0.605
0.613
(COJ
0.015
0.015
The value of C stabilizes after only two iterations.
88
-------
5. Calculate the (HS~) concentration. From Equation (22):
(HS ) = (NH ) - (HCO ) - 2 (CO ) - (NH COO )
= (0.613) - (0.178) - 2(0.015) - (0.216)
= 0.189
Thi? partial pressure of H S is now calculated using Equation (26) . From
Equation (31) K = 0.090 so:
p = (NH*) (HS~)/K (NH )
n b 4 33
= (0.613) (.189)/(0.090) (0.358)
= 3.593
The calculated H S partial pressure agrees closely with the given value
of 3.602 so that the HCO concentration assumed in Step 2 must be correct.
Had the calculated p been too low, the assumed HCO_ concentration would
H S 3
have to be increased and the calculations repeated. The calculated liquid
concentratio: A = (NH ) +
) + (NH2COO~) = 1.187, C = 0.403 and S = 0.189
agree closely with the liquid composition assumed originally.
4.3.2 Design Characteristics
The variation of stripper performance with steam rate, reflux, and wash
is shown in Figures 4-4 through 4-6. The feed compositions used for the
calculations are shown in Table 4-1. Composition "A" is typical of a Lurgi
process condensate after solvent extraction and has a relatively high ammonia
concentration; composition "B" is typical of a Hygas process condensate with a
relatively low ammonia content but high phenol loading.
TABLE 4-1. COMPOSITION OF FEED TO STRIPPERS
Component
NH3
co2
H2S
Phenols
A. Typical Lurgi Condensate
After Solvent Extraction
mg/1
17,000
33,000
100
300
Mole Fraction
0.0180
0.0135
0.000053
0.000057
0.9684
B. Typical Hygas
Condensate
mg/1 Mole Fraction
4,370 0.00466
7,600 0.00312
4,000 0.00076
0.9914
89
-------
8 -
I 4
1.0
2.0
~1 I
3.0 4.0
Ibs STEAM/gal FEED
1
0.1
0.2 0.3 0.4
STRIPPING STEAM Ibs/lb FEED
0.5
Figure 4-4.
Ammonia stripping from Lurgi type condensate
in unrefluxed columns.
90
-------
nr
2.0
-------
i
70
60
50
40
30
20
10
1.0
2.0
3.0 4.0
Ibs STEAM/{{GAL FEED) (R + 1)|
90 < T} ". < 99 ; R = 0
NH.J
O HNH = 92% , 0 < R < 10
I I'
0 0.1 0.2 0.3 0.4 0.5
STRIPPING STEAM Ibs STEAM/{(R + l)(lbs FEED)}
Figure 4-6. Phenol removal in refluxed and unrefluxed columns,
92
-------
The relation between steam rate, number of theoretical plates and
the stripping efficiency of ammonia from the Lurgi type feed water is
shown in Figure 4-4 for columns operating at 25 psia. These data are
for unrefluxed columns. The steam rate has to be approximately doubled
to increase the ammonia removal from 90% to 99% for a fixed column
height. Column heights cannot be significantly decreased by increasing
the steam rate above ^ 40 Ibs. per 100 Ibs feed (% 3 Ibs steam/gallon
feed); nor can significant steam savings be effected by increasing
column heights above 8 to 10 theoretical stages.
Results for the Hygas type feed are shown in Figure 4-5 for 92%
ammonia removal. The steam requirement increases with increasing reflux
as do the cooling water requirements of the overhead partial condenser.
Reflux, which results in increasing the ammonia concentration and decreasing
the phenol concentration in the vapor, will be used only if it results
in some downstream savings. For instance, USS Engineers and Consultants,
Inc. (UEC) :inds that refluxing is beneficial in the Phosam-W process
(Section 4.5) if the ammonia concentration in the stripper feed is less
than about 6,000 mg/1 . Refluxing increases the ammonia concentration
and lowers the temperature in the ammonium phosphate absorber. This
improves absorber performance resulting in both a decrease of absorber
size and reduced steam requirements for absorbant regeneration.
The amount of phenol removed in the stripper is dependent on the steam
rate and the amount of reflux as shown in Figure 4-6. Phenol may be held down
by providing adequate reflux but this is expensive in steam. In the Phosam-W
process some of the stripped phenol leaves the system with the absorber over-
heads while the remainder is carried over to the ammonia fractionator where
10
it is removed.by the addition of NaOH. UEC states that as the phenol concen-
tration in the absorber vapor increases, an increased in the NaOH consumption
must be anticipated and this affects the operating cost of the plant slightly.
If the phenol in the feed water exceeds 300 to 500 mg/1 and if it is required
to minimize NaOH consumption, reflux would be used but this results in an
increase in both capital and operating costs. We do not expect reflux to be
often used to hold down phenol.
93
-------
A comparison was made between the theoretical results presented in this
section and existing sour water (CO free) stripper performance data given in
References 5 and 11. The wide range in actual stripper performance makes
direct comparison difficult. However, the smoothed characteristic curves
presented in Reference 11 show that on increasing the stripping steam rate
from 0.5 to 2 Ibs steam per gallon, feed removal efficiencies for phenol in-
creased from 20 to 48% and for ammonia from 80 to 99% in unrefluxed strippers.
Our calculations for the same conditions yield phenol efficiencies from 15 to
44%, and virtually identical results for ammonia. Similar agreement was found
for refluxed strippers.
4.3.3 Example; Design of an Ammonia Stripper
The number of theoretical stages required to strip 99% of the ammonia
from the condensate A shown in Table 4-1 will be determined. A high steam
rate of 20 Ib steam per 100 Ib feed will be used for this example and a
column pressure of 25 psia at the top will be used.
1. Estimate overhead composition.
At this high ammonia removal efficiency, virtually all the CO will be
stripped. The H S concentration is relatively low, and in the presence of the
large amount of CO it may not be stripped as much as the ammonia. An
indication of the phenol removal may be obtained from Figure 4-6. The follow-
ing removal efficiencies are selected to begin the calculations:
Component Removal Efficiency %
NH3 99 (specified)
CO2 100 (assumed)
H_S 89 (assumed)
Phenol 36.8 (assumed) --^-
2. Calculate top vapor rate.
A temperature is assumed at the top of the column which is slightly lower
than the boiling point of pure water at the operating pressure; this is 232°F.
At this temperature the molar enthalpies for water, obtained from steam tables,
are:
Liquid water 200.25 x 18 = 3,605 Btu/lb mole
Saturated vapor 1157.7 x 18 = 20,837 Btu/lb mole
94
-------
The vapor pressure of water at 232°F is 21.57 psia so the mole fraction
water in the vapor at the top of the column is, approximately,
Y1 = 21.57/25.0 = 0.8628
According to Figure 4-4, about four theoretical stages will be required. The
pressure at the reboiler (taking 0.5 psi per theoretical plate) will be 25 +
(4-1) x 0.5 = 26.5 psia corresponding to a water boiling point of 243°F. At
this temperature the water enthalpies are
Liquid water 211.4 x 18 = 3,805 Btu/lb mol
Saturated vapor 1161.5 x 18 = 20,907 Btu/lb mol
We now have sufficient information to use Equation 16, which may be written
for this case:
F(H«i ~ Hp R) + sp (HV r, ~ Ho tJ
„ _ •*•!• *? i& K v to * /o
= 100(3,605-3,805) + 20(20,907-3,805)
0.8628 x 20,837 - 3,805
= 22.72
3. Calculate composition of vapor leaving top plate.
Applying Equation (18) : y = (Efficiency %) x /V :
NH3: y1 = (99) (0.018) /22. 72 = 0.0784
C°2; Yl = (1°0) <°-0135>/22-72 = 0.0594
H2S: y1 = (89) (0. 000053)/22. 7 = 0.000207
phenol: ?1 = (38.8) (0. 000057)/22. 7 = 0.000097
Previously determined for water: y = 0.828
Total Ey = 1.001
The mole fractions sum to 1.0 indicating that the assumption of 232°F at
the top is correct. If the sum exceeds 1.0, a lower temperature must be
tried; if the sum is less than 1.0, a higher temperature must be tried.
95
-------
4. Calculate the equilibrium liquid composition falling from the top
plate.
An example of the determination of the equilibrium NH , CO and H S
liquid concentrations was given in Section 4.3.1 and will not be repeated
here. The phenol concentration in the liquid is estimated using Equation
(34):
x = y/1.67 = 0.000097/1.67 = 0.000058
The complete composition of the liquid on plate 1 is:
Component Mole Fraction in Liquid
NH3 0.00729
C02 0.00107
H2S 0.000017
Phenol 0.000058
Water 0.99157
(by difference)
Sum 1.0
S. Calculate the vapor rate from the second plate.
A heat balance around each plate is required to exactly determine the flow
rates frcm each plate. However, we are dealing with dilute mixtures and the
liquid rate may be estimated with sufficient accuracy by assuming a constant
molal overflow for water. This is equivalent to stating (Lx ) = constant.
2
For the feed Lx^ = 100 x 0.9684 = 96.84. The liquid rate leaving plate 1 is
L - 96.84/0.99157 = 97.66
Using Equation (6) the vapor rate from the plate below is calculated by
material balance:
V2 " Vl + Ll - F
= 22.72 + 97.66 - 100 = 20.38
96
-------
6. Calculate composition of vapor leaving the second plate.
Equation (7) may be written
VnYn + L x - Fx^
11 n n f
yn+l ~ V, + L - F
1 n
= (22.72)(0.0784) + (97.66)(0.00729) - (100)(.0180)
3! Y2 ~ 22.72 + 97.66 - 100
C02:
H2S:
= 0.03401
y = 0.00511
y = 0.000052
phenol:
Water, by
difference:
y = 0.00010
= 0.9607
The temperature of this plate is assumed to be equivalent to the water
vapor pressure. The total pressure on plate 2 is 25.5 psia. The partial
pressure of water is 25.5 x 0.9607 = 24.50. The temperature is 239.3°F. The
liquid composition is in equilibrium with the vapor at this temperature and
the plate pressure of 25.5 psia is now calculated as in Step 4. Steps 4 to 6
are repeated until the required ammonia removal efficiency has been attained.
The results are summarized in Table 4-2.
The table shows that the required ammonia removal efficiency of 99% is
reached by plate 4 and that the CO , H S and phenol removal efficiencies
correspond satisfactorily with the efficiencies originally assumed in Step 1.
In cases where a poor initial estimate of a removal efficiency has been made,
an improved value in between the assumed and calculated values should be
selected and the stage calculations repeated.
97
-------
TABLE 4-2. SUMMARY OF STAGE CALCULATIONS FOR AMMONIA STRIPPER
Mole Fraction
Removal Efficiency*
Stage Component
1 NH_,
3
co2
H^S
2
Phenol
H O
2
2 NH.,
3
co_
2
H S
2
Phenol
H O
2
3 NH_,
3
CO
2
H S
2
Phenol
H.,0
2
4 NH3
co2
H-.S
2
Phenol
H O
2
In Vapor
0.0784
0.0594
0.000207
0.000097
0.8628
0.0340
0.0051
0.000052
0.00010
0.9607
0.0109
0.00053
0.000029
0.000086
0.9884
0.00289
0.000047
0.000029
0.000072
0.9970
In Liquid
0.00729
0.00107
0.000017
0.000058
0.9916
0.00248
0.000110
0.000012
0.000056
0.9973
0.00076
0.000009
0.000012
0.000052
0.9992
0.000202
^ 0
0.000008
0.000043
0.9997
%
68.7
93.9
75.2
21
_
89.4
99.4
82.5
24
_
96.7
99.9
82.5
29.5
—
99.1
^ 100
88.3
41.4
—
Calculated using Efficiency = 1 - (Bx )/(Fx ) with
B = 100 - V , 100 - 22.72 = 77.28 n
98
-------
4.4 Concentrated Ammonia Fractionation
There are only two components: ammonia and water. However, elementary
textbook calculations are inadequate because the modified
Henry law for ammonia as used by van Krevelen at low concentrations is not
valid above a vapor concentration of about 30% by mass and we wish to produce
99.9% ammonia. No suitable correlations were found in the literature. In
addition, the molal heat of vaporization of ammonia is approximately
half that of water and ammonia has a significant heat of solution. These
heat effects could be neglected in the dilute system but must be accounted
for in concentrated solutions by doing an enthalpy balance around each plate.
In making concentrated ammonia, we will generally require a liquid
product and so the tower will be fitted with a total condenser. A typical
tower is shown in Figure 4.7. A material balance for tray n is;
V . = L +D-F (35)
n+1 n
which for the ammonia component becomes
V ,, y ,. = L x + Ox,, - Fx. (36)
n+1 n+1 n n D f
Equations (35) and (36) yield:
L D(XD " yn+l> + F(yn+l - V (37)
Yn+l - Xn
The enthalpy balance for plate n is:
V , H n + L H. . = V, H , + L H. - FH_ _ (38)
n+1 v,n+l o £,1 1 v,l n £,n £,f '
L
On substituting R = —, V = L + D = D(R+1), and Equation (35) into the
enthalpy balance:
H - RH0- - H _,, )
v,l H, I v,n+lj
D { (R+1)H . - RH. , - H . f + F(H . - HB _)
, = \ v>l ^fl v,n+lj v,n+l i,f (39)
n H , - H.
v,n+l SL,n
99
-------
TOP VAPOR, y
V, moles/hr
1
FEED, xf
F moles/hr
BOTTOMS
'B
B moles/hr
.1 ,
H"
REFLUX, x
moles/hr
o
Tray n-1
Tray n
Tray n+1
TOTAL CONDENSER
DISTILLATE, X
D moles/
fr
D
D _ o
R " ~~
STEAM
rJLs
Figure 4-7. Concentrated ammonia still.
100
-------
1.0
0.8 -
DC
O
Q.
0.6 _
o
O
0.4 _
0.2
250 psia
0.2 0.4 0.6 0:8
AMMONIA CONCENTRATION IN LIQUID, x
Figure 4-8. Dependence of ammonia vapor-liquid equilibrium on operating pressure.
-------
The value of F is zero when Equations (35) to (39) are used above the feed.
<
The feed tray is that for which x - xf.
The tray-by-tray calculations proceed as follows:
Specify the distillate composition x , the condenser pressure P, the reflux
ratio R, and the ammonia removal efficiency n- F°r the total condenser, XD =
y = x . Calculate:
flow rate of distillate: D = n Fxf/xD
top vapor rate: V^ = D(R+1)
bottoms flow rate: B = F - D
bottoms concentration: x = (1 - n)Fxf/B
Specifying the composition and pressure of the condenser fixes the temperature
which is found from Equation (46).
After determining the liquid composition of the top tray from vapor-
liquid equilibrium correlations, the only remaining unknown for the top tray
is the liquid rate. The following procedure may be used. The pressure of the
next tray, n+1 say, is set at
P . = P + 0.5 (40)
n+1 n
A temperature for this tray is assumed, which fixes the vapor composition for
the binary system, using Equation (46). The value of y , so obtained is used
n+1
in the material balance, from Equation (37), to calculate the required liquid
rate L . The temperature assumption is checked using the heat balance from Eq«-
uation (39) . The process is repeated until the heat and material balances cor-
respond. This procedure fixes the liquid rate for tray n, and the temperature
and vapor composition for tray n+1. The vapor rate for tray n+1 is calculated
from Equation (35) and the liquid composition from the vapor-liquid correlations.
The calculations continue until the obtained liquid composition meets the
bottoms specification x . This tray is the reboiler.
4.4.1. Vapor-liquid Equilibria for the Ammonia-water System
There are two degrees of freedom for the binary sy; tern. The vapor
composition is fixed and may be calculated once the liquid composition and
102
-------
pressure (or temperature) is stated:
y = f(x, P)
(41)
This is the usual way of presenting vapor-liquid equilibria. The form of
relationship for the ammonia-water system is shown in Figure 4-8.
For our calculations it was useful to have a slightly different relation-
Ship: y = f (t, P) (42)
x = f(t, P) (43)
A knowledge of pressure and temperature can be used to calculate both liquid
and vapor composition. Alternatively, a knowledge of pressure and any one
composition can be used to determine the temperature and the other composition.
8 *3
Published ' vapor-liquid equilibrium data for the ammonia-water system were
used to obtain the correlations.
(a) Vapor composition
If P is the vapor pressure of pure water, P is the total pressure
and x and y are the mole fractions ammonia in the liquid and vapor phase
respectively, then Raoult's law may be written:
(1-y) = Pv(l-x)/P (44)
P
Y
Since the ratio — is unique for any temperature and pressure, both the liquid
and vapor compositions must be a unique function of this ratio. The vapor
composition may be represented by;
(1-y) = f(Pv/P) (45)
where the function f includes deviations from Raoult's law. Figure 4-9 is a
plot of (.1-y) vs. P /P. Over most of the concentration range for which data
are available, a least squares fit of the variables used in Equation (45) gave:
(1 - y) = 1.03 (pyp)1'27 , 0.2 < y < 0.97 (46a)
This equation is not satisfactory in the ranges y -> 1 and y -> 0 and two other
equations have to be introduced. In the high concentration range the data are
correlated by:
(1-y) = 7.85 (P /P)1'75 , y > 0.97 (46b)
103
-------
Data for the low concentration range were not available. However, Henry's
Law, previously used in the ammonia stripper calculations, may be applied in
this concentration range. Henry's Law as given in Equation (27) may be
rewritten in terms of mole fractions and psia.
H is the Henry's Law coefficient as given by Equation (33). Substituting for
o
x in Equation (44) yields:
y = (1 - P /P)/(l -HP /I. 07)
The denominator in this equation varies from 0.95 at 150°F to 0.90 at 500°F
and may be considered constant over the temperature range corresponding to the
dilute (lower) end of our f ractionator . We also found that to match the
available equilibrium data in the higher concentration range, and to ensure
the correct limit y -*• 0 as (P /P) -»• 0, the denominator should be taken to be
one. With sor 3 approximation we propose for the filute ammonia solutions:
(1 - y) = PV/P , y < 0.2 (46c)
Note that Equation (46c) follows directly from Equation (44) as x -> 0.
(b) Liquid composition
The most satisfactory method for determining the liquid composition over
the full concentration range was based on an Antoine type relationship:
ID
In P = A - t°F + 384~ (P in psia) (47a>
Both A and B are dependent on x.
Values of A and B were calculated for each tabulated x value using two
extreme temperature data points. The calculated values, shown plotted in
Figure 4-10, were found to be correlated by:
A = 14.49 - 1.90 { sin (130x) }1//2
(47b)
B = 7036 - 3502 { sin (100x)}°'6 (47c)
where the sine is to be taken of the numerical value in degress, not radians.
104
-------
0.6 —
0.1 -
0.2 0.3 0.4 0.5 0.6
Figure 4-9. Ammonia concentration in vapor.
105
-------
6 _
5 _
4 _
o CALCULATED FROM TABULATED DATA
0.6
.B = 7036 - 3502 {sin TOO x i
12
o CALCULATED FROM TABULATED DATA
A = 14.49 - 1.90 {sin 130 x
0.5
0.2 0.4 0.6 0.8
MOLE FRACTION AMMONIA IN THE LIQUID, x
1.0
Figure 4-10. Determination of ammonia concentration in liquid.
106
-------
Note that Equation (47) reduces to Equation (8), the vapor pressure expression
for pure water, when x = 0. The liquid composition for a known pressure and
temperature must be determined by trial.
Although highly empirical, Equation (47) yields liquid compositions
within 5% of those presented in Reference 9. Differences of up to 10% exist
between the data in Reference 8 and Reference 9. For very low liquid concen-
trations, X < 0.02, with known vapor composition, it is advisable to use the
Henry's law expression, Equation (27).
4.4.2. Enthalpy Correlations for the Ammonia-water System
Enthalpy correlations are needed to determine the temperature gradient in
the concentrated ammonia still. In Reference 9 liquid and vapor enthalpy data
are tabulated against temperature and liquid composition, and against pressure
and liquid composition. The liquid composition may be eliminated to obtain
enthalpy as a function of temperature, with pressure the parameter as shown in
Figures 4-11 and 4-12.
The liquid enthalpy is not strongly dependent on pressure over most of
the concentration range, and a least squares fit to all the data (with a
correlation coefficient of 0.97) gave:
H = 1.25t - 138 Btu/lb
JO
100 < t°F < 400; 25
-------
400
o
i
-------
400
350
300
250
ce.
•f
ce
UJ
200
150
100
50
psia
• 25
• 50
O 100
• 150
a 200
A 250
v 300
JL
1
500
600
700 800 900 1000
SATURATED VAPOR ENTHALPY, Btu/lb
1100
1200
Figure 4-12. Vapor enthalpies of saturated ammonia-water.
109
-------
then we have to operate at elevated pressures. Cooling waters are typically
available at 80°-90°F and to keep condenser surface areas within practical
limits, a cooling driving force of about 20°F is required which implies
condensing temperatures around 105°F. The saturation temperature-pressure
curve for 100 %NH3 shown in Figure 4-13 suggests that the distillation column
should operate between 200 and 250 psia.
The steam rate and number of theoretical plates to produce 99.9% amnonia
are shown in Figure 4-14. Increasing the steam rate above ^ 0.6 Ibs/lb feed
does not reduce the number of plates required, while increasing the number of
plates above 8 to 10 does not result in significant steam savings. Having to
operate at 225 psia rather than 150 psia increases the steam required by 25%.
An increase in the feed concentration increases the steam requirements per
unit feed, particularly for towers designed to operate at low steam rates.
This is shown in Figure 4-15 for feed concentrations in the range 20 to 40%
ammonia. For a tower with 10 theoretical stages, doubling the ammonia feed
concentration increases the steam per unit of feed by about 33%. However,
doubling the feed concentration means a 50% reduction in feed rate for the same
product rate, which results in a net saving of 33% in steam. Steam rates for
concentrating ammonia from a typical process condensate concentration of 4,000
mg/1 are shown in Figure 4-16. A steam rate of 0.2 Ibs/lb feed is required to
produce 30% aqueous ammonia at 25 psia in a 7 stage tower, or 99.9% anhydrous
ammonia at 225 psia in an 11 stage tower. In comparing these steam rates with
those for the more concentrated feed (0.4-0.6 Ibs steam/lb feed), the volume of
feed required for the same product volume is increased by a factor of about 50
to 100.
4.4.4 Example; Design of an Ammonia-Water Fractionator
Ninety-nine percent of the ammonia in a 30 mole % feed solution will be
recovered as a 99.9% pure product. A relatively high reflux ratio of R = 2 is
specified and a total condenser pressure of 225 psia is selected.
1. Determine overall material balance.
(a) The product rate D = n FX /x
= 99 x 0.3/0.999
= 29.73
110
-------
140
120
X- ioo
i 80
Si 60
40
20
50
100
I
I
150 200
PRESSURE, psia
250
300
350
Figure 4-13. Pressure-temperature relation for saturated ammonia.
Ill
-------
11
10
,/> 9
UJ
J—
3
CC
o
FEED CONCENTRATION xf
0.3 MOLES NH3/MOLE SOLUTION
99S RECOVERY AS 99.95 NH,
psia —J
_L
JL
0.2 0.4 0.6
STEAM RATE, Ibs/lb FEED
0.8
1.0
Figure 4-14. Steam required for ammonia distillation.
112
-------
11
10
FEED CONCENTRATION
xf MOLES NH3/MOLE SOLUTION
99X RECOVERY OF 99.91 NH,
JL
JL
0.2 0.4 0.6 0.8
STEAM RATE, Ibs/lb FEED
1.0
Figure 4-15.
Effect of feed concentration on steam rate for
ammonia distillation.
113
-------
10
•a:
_l
a.
ce
o
o
ee.
>t
PRODUCTION OF 99.9% NH
PRODUCTION OF 30% NH
FEED CONCENTRATION
xf = 0.0042 MOLES NH-j/MOLE
SOLUTION
(4,000 mg NH3/litre}
J_
JL
0.2
0.4 0.6 0.8
STEAM RATE Ibs/lb FEED
1.0
1.2
Figure 4-16. Direct distillation of process condensate.
114
-------
Vapor enthalpy, plate2 from Equation (49): HV = 669.1 Btu/lb
Substitution into Equation (39) yields the liquid rate from the top plate L =
46.6.
The water vapor pressure and ammonia vapor concentration are calculated:
Water vapor pressure at 257.4°F, from Equation (8) = 33.8 psia
Mole fraction ammonia in vapor. Equation (46a) = 0.879
and substituted in Equation (37) to get a second estimate of the liquid rate
L = 46.3. The correspondence between the two calculated liquid rates indi-
cates that the assumed temperature for stage 2, t = 257.4°F, is correct. In
cases where liquid rate determined by material balance is low relative to that
determined by the enthalpy balance, a higher temperature must be assumed.
Calculation of the liquid rate completes the information required for the
first plate and we now proceed to the next plate.
6. Calculations for plate (n+1).
The temperature and vapor composition for each lower plate are found when
determining the liquid rate for the previous plate. The vapor rate for plate
(n+1) follows from Equation (35):
V = L + D - F
n+1 n
or, for plate 2:
V2 = 46.4 + 29.73 - 0 (F = 0 above feed)
= 76.1
Steps 3 to 6 are repeated for successive stages until the liquid compositions
are equal to or less than the calculated bottoms compositions x = 0.0043.
The results are:
Stage
1
2
3
4
5
X
0.802
0.260
0.066
0.011
0.0015
y
0.999
0.879
0.458
0.113
0.016
V
89.1
76.1
90.7
94.0
97.1
L
46.4
161.0
164.3
167.4
(70.3)
T°F
111.6
257.4
346.3
382.1
391.0
Ppsia
225.5
226.0
226.5
227.0
227.5
116
-------
The liquid composition on stage 2 is just less than that in the feed, so
the feed is introduced on stage 2. By stage 5, the liquid composition has
fallen below the calculated x value, so stage 5 is the reboiler and the
liquid rate is the bottoms rate calculated in Step 1.
7. Calculate the steam rate.
The steam required for fractionating is equivalent to the vapor leaving
the reboiler. Since the molecular weight of the feed is 0.3 x 17 + 0.7 x 18 =
17.7, the steam rate on a mass basis is
Steam rate = -(97.1 x 18)/(100 x 17.7) = 0.99 Ib steam/lb feed
4.5 Ammonia Recovery Using the Phosam-W Process
The Phosam-W process is a proprietary process of USS Engineers and
Consultants Incorporated (UEC), a wholly-owned subidiary of U. S. Steel
Corporation. The design procedure is quite complex and certain proprietary
information is necessary; designs are made to order by UEC. We have given a
brief description of the process and some of the interrelating effects involved
in making a good design. The description shows why a complex computer program
is used for designs and why rules of thumb are inadequate.
An exemplary process flow diagram is shown in Figure 4-17. There are four
major pieces of equipment: (1) a water stripper in which volatile contami-
nants, ammonia and acid gases (carbon dioxide and hydrogen sulfide) are driven
out of the wastewater; the wastewater passes on for subsequent treatment or
use; (2) an absorber in which ammonia is absorbed out of the vapor into a
solution; the acid gases pass on for further treatment or disposal; (3) a
Phosam stripper in which the absorbing solution is regenerated for reuse and
aqua ammonia is recovered; and (.4) an ammonia still in which aqua ammonia is
concentrated to anhydrous ammonia for sale or use.
(1) The Water Stripper
The water stripper is not proprietary and design procedures were
given in Section 4.3.
It is usually economical to keep the stripper between 20 and 40 actual
trays. In these systems with a low vapor rate compared to the liquid rate,
tray efficiencies of between 25 and 50% are common. The higher efficiency is
Only realized w th well-designed trays operated at 75 to 80% of flooding. If
throughput is decreased, efficiency will fall. Twenty to forty actual trays
117
-------
oo
FEED
STRIPPER
. FRACTIQNATOR
Aqua Ammonia
E-5
cu
WASTEWATER
C£NDENSATE
Figure 4-17. Phosam-W process for ammonia separation.
-------
is in the range of 7-15 theoretical stages. Referring to Figure 4-4 we see
that for 99% ammonia removal about 10 Ibs steam per 100 Ibs feed are required.
In practice, steam rates of about 11-13 Ibs per 100 Ibs feed are reportedly
close to an economical optimum. Increasing the steam rate above this value
without reflux increases the water vapor rate to the downstream absorber and
may increase the cost of ammonia recovery. Ninety-nine percent ammonia
removal is not always necessary for coal conversion plants since certain
amounts of ammonia are tolerable or required in downstream processes.
Integration of the water stripper into the rest of the system serves to
conserve energy. Wastewater is preheated partly by exchange with stripped
water (Exchanger 8 in Figure '4-17) and partly by exchange with lean Phosam
solution CExchanger 1). If the ammonia concentration in the feed water is
low, there will be a lot of water to preheat with not much Phosam solution to
supply the heat. The rate of circulation of Phosam solution depends on the
rate of feed of ammonia, not water. In this case steam preheat may be needed.
The water stripper may have a boiler or use live steam; it makes no
difference to the overhead vapor. If live steam is used, the condensate
enters the stream of stripped wastewater and is not recovered. If a reboiler
is used, the steam must be at a higher pressure to provide enough temperature
difference across the reboiler. A typical reboiler temperature is 290°F and
a typical steam pressure for the reboiler is 64 psia.
Some of the water stripper steam duty may be provided by the reboil
condensate of the two downstream towers, as well as bottoms from the ammonia
fractionator. These streams are saturated at 200-300 psig, and ^ 16% will
flash at the water stripper conditions. The bottoms stream should be kept
separate from the reboil steam and may be flashed directly into the water
stripper as shown in Figure 4-17. The steam savings realized by utilizing
reject heat from the fractionator is shown in Table 4-3. The steam saved by
using condensate from the Phosam stripper reboiler is not included in Table 4
-3. A small quantity of caustic, added to the ammonia fractionator to control
carbonate buildup, is introduced into the water stripper with the fractionator
bottoms, and serves to free an equivalent amount of fixed ammonia.
119
-------
TABLE 4-3. CALCULATED STEAM REQUIREMENTS FOR WATER STRIPPING AND AMMONIA
FRACTIONATION. (BASED ON 100 LBS FEED AND 11 LBS STEAM TO THE WATER STRIPPER)
Ammonia Concentration Fractionator Flash*(Fron.
Stripper
Feed
%
0.6
1.2
Fractionator
Feed
%
10
20
10
20
Feed
Ib
6
3
12
6
Steam
(10 Stages)
Ib
1.74
0.99
2.90
1.98
Bottoms
Ib
5.4
2.7
9.7
4.9
Fractionator
Bottoms and Reboiler)
% Stripper Steam
10.0
5.1
17.6
9.6
*UEC states that flash from fractionator bottoms provides from 7 to 11% of
water stripper steam requirements, depending on water stripper feed NH
concentration.
(2) The Absorber
The Phosam solution is phosphoric acid. A lean solution contains
close to but more than one NH per PO ~ and a rich solution contains close to
but less tha: two NH per PO ~. The solution is used inside the range of
monoammonium phosphate with a little diammonium, and mostly but not 100%
diammonium phosphate. The concentration of the solution is such that crystals
do not form (probably below 120 g PO /I) . Absorption requires a low tempera-
ture because the vapor pressure of ammonia over the solution decreases with the
temperature. Absorption temperatures are usually kept below 220°F. This can
be achieved by operating unrefluxed water strippers at a maximum of about 25
psia. The pressure drop for the vapor passing through the absorber is low and
the acid gases leave at above atmospheric pressure.
Absorption is exothermic. The heat evolved per absorption of unit q
quantity of solution (the differential heat of absorption) decreases as the
ratio of NH4 : P04 in solution increases. Since the absorption equilibrium is
temperature dependent, the calculations are necessarily complex.
Cooling the absorber is necessary. Solution is p.ulled out of the
absorber, cooled in Exchanger 7, and returned. (Exchanger 7 can be an air
cooler.) Since the return is under pressure from the pump, it is sprayed into
the absorber and the bottom section of the absorber is a spray. Above the
spray section trays are used.
120
-------
Cooling usually causes water to condense out of the vapor. This will
always happen if the water stripper has no reflux so the water content of the
vapor is high. Condensed water enters the circulating Phosam solution and
passes to the Phosam stripper where it leaves in the aqua ammonia. Thus this
cooling effects the concentration of the aqua ammonia to the ammonia still and
henco the heat supply to the ammonia still. Optimization requires a balance
here.
(3) The Phosam Stripper
To reverse the absorption reaction the pressure is raised, the
temperature is raised, steam is passed through the solution to strip it, and
heat is supplied because the desorption is endothermic.
Excessive steam for stripping is required if the pressure is below about
200 psig and 200-250 psig is usually used. If a reboiler is used, steam is
usually required at 300 psig or above. Live steam can be used but this
dilutes the aqua ammonia. If the water stripper has a reflux so that little
water is condensed from the vapor, then it is often economical to use live
steam in the Phosam stripper (except that the condensate goes to the wastewater
stream). If the water stripper has no reflux, excessive dilution will occur if
more than 50% of the steam to the Phosam stripper is live. Aqua ammonia is
best recovered at not less than 10% NH and cannot usually be recovered at
greater than 20% NH . However, this range means a possible variation of two-
fold in the feed to the ammonia still and has a considerable effect on the
system economics. Thus, the use of live steam to the Phosam stripper requires
careful evaluation.
Aqua ammonia is condensed in Exchanger 3 against rich solution and cooling
water. Aqua ammonia is not refluxed to the Phosam stripper. The Phosam
stripper feed temperature can be controlled to condense more or less vapor from
the top tray. This affects the aqua ammonia feed strength and the plant water
balance.
(4) The Ammonia Still
Design procedures were discussed in Section 4.4 where it was shown
that steam rates are slightly dependent on the concentration of the aqua
ammonia, but will generally be in the range 0.2-0.4 Ib steam per Ib feed. If
a reboiler is not used, the steam condensed passes to the wastewater after
recovering resi ual heat by flashing in the water stripper.
121
-------
The ammonia still will plug up if tramp oils are not released and if
caustic soda is not added to dissolve carried over carbon dioxide. These
important procedures are proprietary to UEC.
4.6 Ammonia Recovery as Ammonium Sulfate
Ammonia is removed from the ammoniacal liquor in a water stripper and
bubbled through dilute (approximately 60%) sulfuric acid to form ammonium
sulfate crystals. The crystals, which may contain some tars and free acid,
are suitable for use as fertilizer. Process flow sheets for ammonium sulfate
14
production are described in the literature
Modern plants must be designed so that the acid gases, which are not
dissolved in the sulfuric acid, do not become a pollution problem. If dilute
sulfuric acid is not a process by-product, then from considerations of both
transport and materials of construction it might be more economical to use
concentrated acid.
The economy of the ammonium sulfate process depends on the market prices
of ammonia and ammonium sulfate. Depending on the quality of the ammonium
sulfate, its market value (FOB) will be from about 30% to 50% of that of
ammonia. However, the ammonium sulfate market is variable, shipping costs are
high, and the quality of the product will probably be such that the antici-
pated market value is at the lower end of this range. This is little more than
the ammonia content of the salt, so the costs of sulfuric acid will not be
recovered. Recovery of ammonia wastes as anhydrous ammonia is feasible, and
the projected ammonia market appears to be stable ; therefore, the ammonium
sulfate route is not recommended.
4.7 Ammonia Recovery by the All-Distillation Process
This process consists of two distillation columns. In the first, the acid
gases are stripped with steam while the ammonia is held in solution. In the
second column, the ammonia is concentrated; either a 30% aqueous ammonia, or
anhydrous ammonia may be produced.
Details of the all-distillation system, known as the CLL (Chemie-Linde-
Lurgi) process are proprietary. Lurgi has proposed this system for their El
Paso and Wesco plants , and claims that it is economically more attractive
than other processes for ammonia recovery in large plants. It is not known how
122
-------
the ammonia is held in solution in the CLL process. One method would be to
provide a wash stream to the top of the stripper (see Figure 4-16). The wash
water would exit with the tower bottoms and form part of the feed to the
ammonia concentrator, so increasing the steam requirements of this unit.
Another method would be to fix the ammonia with acid. The fixed ammonia must
then be released by addition of alkali; milk of lime is usually used. The
addition of chemicals means increased costs and an increased TDS of the waste-
water, and may lead to carbonate plugging of the ammonia concentrator. It is
possible that a combination of wash and chemical additives are used in the
proprietary process. • If chemical additives are not used, the separation will
depend on the relative volatilities of the gases.
The volatility of H S relative to NH in a system containing only H S, NH
•^ J £• J
and HO is shown in Figure 4-18. To increase the relative volatility, and
thereby facilitate the separation of H S from NH , the temperature and/or the
ft -J
H S concentration must be increased. Of course the H S concentration drops as
£* £•
stripping proceeds, and high temperature (pressure) operation is therefore
selected. Chevron indicates operation at 100 psig, but also indicates
recycle of what appears to be concentrated ammonia to the H S stripper. (See
Figure 4-19.) This is not understood since we know from Figure 4-18 that
increasing the ammonia concentration decreases the volatility of H S relative
to NH .
The relative volatility of CO and NH in systems containing CO , NH and
HO is shown in Figure 4-20. The relative volatility of CO is an order of
£t £
magnitude greater than that of H S at low temperatures, and increases rapidly
with temperature. CO is therefore considerably easier to separate from NH .
£ J
The effective relative volatilities may be increased by using wash and/or
reflux. Wash will hold down the less volatile component in the overhead as
will reflux. We found that a total condenser should be used to provide the
reflux since this increases the concentration of the more volatile component
and also increases its relative volatility (Figures 4-18 and 4-20).
We investigated stripping 99% of the CO while holding down 95% of the
ammonia in an H S free Lurgi type process condensate (Table 4-1) , using wash
rates of 25% to 100% of the feed rate and the reflux necessary to achieve the
specified separation. Some results are shown in Figure 4-21. Several combina
tions of steam rate and reflux yield the desired separation for each wash rate
123
-------
selected; however, we did not search all possibilities and these results may
not be optimum.
The cost of the wash and reflux can be approximated as follows: the
ammonia fractionator requires high pressure steam at about 0.2 Ib/lb feed; at
a steam cost of $2.5 per 10 Ib, additional steam due to the use of wash costs
7 fi
$0.5 per 10 Ib wash used. About 10 Btu must be removed in the condenser for
each 103 Ib reflux; if cooling water costs 12C per 10 gals circulated, reflux
costs $0.6 per 10 Ib reflux. In addition, the capital cost of the condenser,
amortized at 19%/yr, amounts to $0.03 per 10 Ib reflux. These costs are shown
in Figure 4-22 and suggest that the optimum operation will be with small wash
rates and at high pressure. The small differences in steam rate in the acid
gas stripper, and the dependence of the cost of this stripper on pressure are
not included in Figure 4-22, but they do not significantly affect these
findings.
In Figure 4-22 we see that in addition to the usual costs for stripping
and ammonia i/'actionation, an additional $1 to $2 per 10 gallons is required
to separate tne acid gas and ammonia by this method. By comparison, the value
of ammonia in a condensate containing 17,000 mg is nearly $8 per 10 gals.
We emphasize that our findings here are based on our theoretical stripper
design procedure (Section 4.3); they may not be achievable in practice and
they do not reflect any proven commercial process.
124
-------
12345
TOTAL AMMONIA/TOTAL H.,S IN SOLUTION, HOLE/MOLE
Figure 4-18. Volatility of H2S relative to NH3 in H2S-NH3-H2O systems,
125
-------
so
en
HYDROGEN
SULFIDE
STRIPPER
RECYCLE
1
I
X,
1
1
*/
,DEAPATFn
CONOENSATE
AMMONIA
STRIPPER
^
Q
^~
X_
/^-
E
DEGASSED SOUR WATER FROM STORAGE TANKS
HYDROGEN SULFIDE PRODUCT 50 PPM (WT) AMMONIA MAXIMUM
( D
RECYCLE
AMMONIA PRODUCT
TO FEED 5 ppM (HT) HYDROGEN SULFIDE MAXIMUM
rSTRIPPED WATER PRODUCT TO PROCESS UNITS 50 PPM (WT) AMMONIA b PPM (WT)
HYDROGEN SULFIDE MAXIMUM
Figure 4-19. Chevron all distillation process for recovery of H S and ammonia from sour waters.
-------
5,000
2,000
1,000
z 500
200
100
50
20
atm)
J_
0123456
TOTAL AMMONIA/TOTAL C02 IN SOLUTION, MOLE/MOLE
Figure 4-20. Volatility of CO2 relative to NH3 in CO2-NH3-H2O systems
127
-------
0 -25 —
0-25 0 .SO 0.75
WASH Ib/lb FEED
1 .0
Figure 4-21. Separation of CO2 and NH3 by distillation using water
wash and reflux. Stripper specifications: 99% CC>2/ 5% NH3 removal.
128
-------
CO
CD
oo
o
60 psia
100 psia
200 psia
0.25 0.5 0.75
WASH Ib/lb FEED
1 .0
Figure 4-22. Cost of holding down ammonia when stripping CO2-
Condenser cooling water costs, condenser capital costs (19%/yr),
and costs of additional steam for ammonia fraction only. Reflux
rates required as per Figure 4-21.
129
-------
REFERENCES Section 4
1. Bolton, R.L., and Klein, L., Sewage Treatment; Basic Principles and
Trends, Ann Arbor Science Publishers, Mich., 1971.
2 Arthur D. Little, Inc., "Research on Chemical Odors," prepared for the
Manufacturing Chemists Association, October 1968.
3. Kleusener, J. , Heist, J. , and van Note, R.H., "A Demonstration of Waste-
water Treatment for Reuse in Cooling Towers at Fifteen Cycles of Concen-
tration," AIChE Water Reuse Conference, Chicago, 111., May 4-8, 1975.
4. Wei, I.W., and Goldstein, D.J., "Biological Treatment of Coal Conversion
Condensates," presented at Third Symposium on Environmental Aspects of
Fuel Conversion Technology, Hollywood, Fla., September 13, 1977.
5. Beychok, M.R., Aqueous Wastes from Petroleum and Petrochemical Plants,
John Wiley & Sons, 1967.
6. van Krevelen, D.W., Hoftijzer, P.J., and Huntjens, F.J. , "Composition and
Vapor Pressures of Aqueous Solutions of Ammonia, Carbon Dioxide and
Hydrogen Sulfide," Recueil Trav. Chim. , 68_ 191-216, 1949.
7. Water Purification Associates, "Water Conservation and Pollution Control
in Coal Conversion Processes," Environmental Protection Agency, EPA
600/7-77-065, June 1977. NTIS catalog PB269-568/ZWE.
8. Perry, R.H. , and Chilton, C.H., eds., Chemical Engineers Handbook,
Fifth edition, McGraw-Hill, 1973.
9. Macriss, R.A., et. al. , "Physical and Thermodynamic Properties of Ammonia
Water Mixtures." Institute of Gas Technology, Research Bulletin No. 34,
September 1964.
10. Rice, R.D., USS Engineers and Consultants, Inc., Private communication.
11. American Petroleum Institute, "Sour Water Stripping Survey Evaluation,"
Publication No. 927, June 1973.
12. Morresi, A.C., and Cheremsinoff, P.N., "Cooling Water and Boiler Possi-
bilities for Wastewater Reuse," Industrial Wastes, 24,33, Mar/Apr. 1978.
13. Klett, R.J., "Treat Sour Water for Profit," Hydrocarbon Processing, 97,
October 1972.
14. Fluor Corporation, Ltd., "Ammonium Sulfate," Petroleum Refinery, 38, 222,
November 1959. See also Petroleum Refinery, 219, November 1957.
15. Anon., "Key Chemicals: Ammonia," Chem S Eng News, 9, January 31, 1977.
16. Beychok, M.R., "Coal Gasification and the PhenoscIvan Process", 168th
National Meeting, American Chemical Society, Atlantic City, N.J., Div.
of Fuel Chemistry Preprints, Vol. 19, Number 5, September 8-13, 1974.
130
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17. Strakey.- J.P., Forney, A.J. , Haynes, W.P. , and Plants, K.D., "Effluent
Treatment and its Cost for the Synthane Coal-to-S.N.G. Process," 168th
National Meeting, American Chemical Society.- Atlantic City, N.J. , Div.
of Fuel Chemistry preprints, Vol. 19, number 5, September 8-13, 1974.
18. Melin, G.A., Niedzwiechi, J.L., and Goldstein, A.M., "Optimum Design of
Sour Water Strippers," Chem. Eng. Prog. 71(6), 78-82, 1975.
19. Sinor, J.E., ed. "Evaluation of Background Data Relating to New
Source Performance Standards for Lurgi Gasification," Interagency Energy-
Environment Research and Development Program Report, Environmental
Protection Agency, EPA-600/7-77-057, June 1977.
131
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5. BIOLOGICAL TREATMENT
Biological oxidation is the process in which living organisms, mostly
bacteria, cause the reaction of dissolved oxygen with dissolved and sus-
pended organic molecules to make more organisms and small residual molecules,
mostly carbon dioxide. The result is water cleaned of most of its organic
content and a settleable, separable sludge of dead and living bacteria.
Very approximately, less than half the weight of BOD (biochemical oxygen
demand) is converted to weight of sludge.
In the absence of much carbonaceous matter bacteria will flourish
which can oxidize ammonia to nitrate. This process is nitrification. The
water is now passed to a separate vessel in which oxygen is withheld.
Different bacteria will use the oxygen in nitrate leaving the nitrogen in
the molecular form; this is denitrification. Selected carbonaceous materials,
usually methanol, must be added to a denitrification reactor. Nitrification
and denitrification are expensive and are probably not suitable for removing
the massive quantities of ammonia found in wastewaters from fuel conversion
plants.
In the absence of oxygen bacteria will convert organic matter to more
bacteria and methane. This is anaerobic treatment and it has the particular
advantage of not requiring a large energy input for oxygenation or aeration.
However, evidence is lacking that anaerobic treatment will function smoothly
on wastes from coal conversion.
Not all organic molecules are susceptible to biological oxidation;
those that are not are biorefractory. One measure of the content of biore-
fractory molecules in water is the ratio of chemical oxygen demand to
biochemical oxygen demand (COD/BOD). If the ratio is high, there are many
molecules that are chemically oxidizable and not biologically oxidizable.
When the ratio is greater than about two, the water is considered to have a
large amount of biorefractory content. All available evidence is that
wastewaters in coal conversion are predominantly biodegradable and that
biological oxidation is a functional treatment. This does not mean that it
is economical at all flow rates and concentrations.
132
-------
Many biodegradable molecules are toxic to bacteria in high concentra-
tions. In particular phenol, which is rapidly biodegradable at low concen-
trations, must usually be kept below 500 mg/1 in the reaction tank.
Ammonia in high concentration is also toxic, but it will be removed before
biological treatment.
Biological oxidation requires the presence of bacteria, organic matter
and dissolved oxygen. For most wastes the concentration of bacteria yielded
by a once-through reactor is not high enough to obtain on adequate rate of
reaction. It is usual, therefore, to separate the sludge from the slurry
leaving the oxidation tank and return much of it to the tank; the balance
of the sludge is disposed of. This procedure maintains a high bacterial
concentration in the reactor at all times and is called the activated
sludge process. Another way to obtain high concentration of bacteria is to
grow the bacteria on a solid surface and to pass the wastewater in a thin
film over the surface which is the principle of a trickle filter.
The bi^logical agents in municipal sewage plants are simply soil
bacteria. More efficient will be the sludge taken from a plant treating
coke oven wastewater which is phenolic. These bacteria are acclimated to a
phenolic waste and will grow more quickly than unacclimated bacteria.
.26,27
23
Acclimated sludge is available . Bacteria deliberately mutated for acclima-
tion to phenol are also available and seem to be even more efficient
The oxidant can be air or oxygen. In municipal sewage plants air is
the most common oxidant. It is necessary to stir the bioreactor to keep
the sludge in suspension. If the stirrers are located so that they also
pull air into a vortex in the tank, the oxidant addition is obtained for
only a modest increase in energy consumption over that needed for stirring.
While municipal sewage may have only 200 mg BOD/1, the wastes from coal
conversion plants will have up to 20,000 mg BOD/1. Thus, the use of simple
surface aerators will consume a lot of energy. Surface aerators may need
0.4 to 0.45 kw-hrs/lb BOD removed. For 15,000 mg BOD/1 removed this trans-
lates to 50-56 kw-hr/thousand gallons, therefore, it may easily be more
economical to use oxygen instead of air. Most coal conversion plants are
making oxygen on site in any case.
For background information the reader is referred to References 28,
29, 34 and 45. The use of oxygen is described in References 30 and 31; a
133
-------
description of anaerobic treatment is given in Reference 32; a combination
of trickle filters and activated sludge are described in Reference 33.
5.1 Air Activated Sludge (AAS) Process
The activated sludge process is widely used, mostly in municipal
sewage treatment plants. On Figure 5-1. is shown a scheme appropriate for
SOLIDS SEPARATOR
(clarifier)
Q, S
^ BIOREACTOR
(aeration basin)
Xr.
RECYCLED
SLUDGE
DISPOSAL
SLUDGE
DENATURING
Q = flow rate of liqniid waste to be treated biologically,
volume/time;
q = flow rate of recycled sludge, volume/time;
w = flow rate of wasted sludge, volume/time;
So = influent substrate concentration, mass/volume;
Si = effluent substrate concentration, mass/volume;
X = microbial mass concentration, mass/volume;
X = microbial mass concentration in the clarified overflow from
the solids separator, mass/volume;
*r = microbial mass concentration in the underflow from the solids
separator, mass/volume;
P » power requirement for aeration, energy/time;
PX = excess microorganisms production rate, mass/time.
V = volume of aeration basin
Figure 5-1. Air activated sludge material.
134
-------
foul process condensates. The feed will have to be from a storage pond or
tank to keep changes in the flow and concentration to the bioreactors as
slow as possible. Phosphorus, an essential nutrient for the bacteria, will
have to be added. The bioreactor is one or more stirred basins arranged in
series and the concentration of organic waste in the reactor is close to
the concentration in the water leaving the reactor. By carefully distribut-
ing the concentrated feed, the toxicity of high concentrations of phenol
can be avoided.
Design equations and two examples will be given for this process.
5.1.1 The Biokinetic Model. The biokinetic model developed by
2 28
Lawrence and Mccarty ' has been widely used for the activated sludge process
and will be used here. It is primarily based on a relationship between
substrate utilization and microbial growth, which can be formulated in
two basic equations. The first equation is an empirically developed
relationship between the net growth rate of microbes and the rate of
substrate utilization:
= Y -- k X (1)
dt Y dt V U)
where X = concentration of microorganisms
t = time
Y = growth yield coefficient; weight of microorganisms
produced per weight of substrate removed
S = concentration of substrate or degradable contaminant
k = microorganism decay coefficient, time
If Equation (1) is divided by X, we obtain:
3.dX_YdS
X dt ~ X dt " *d (2)
In Equation (2) each term has the dimension (time ) and compatible units
must be used. The lefthand side of Equation (2), which is the rate of
increase of concentration of microorganisms per unit concentration, may
also be written 1/6 , where 0 is called the mean cell residence time, or
c c
sludge age. The first term on the righthand side of Equation (2) includes
135
-------
-i AC
the quantity — —, which is the rate of decrease of concentration of substrate
per unit concentration of microorganisms, often described as F/M ratio.
The quantity — — = F/M is a function of the concentration of substrate
X ut
and the second equation of the Lawrence and McCarty model assumes the function:
1 dS kS
X dt K + S
s
(3)
The coefficients k and K in Equation (3) may be evaluated by plotting
S
1 f3C
— — versus S, and the values of Y and k in Equation (2) can be determined
X dt i dX 1 dS
by plotting versus — —. Typical plots prepared by us are shown in Fig-
X dt X dt « ~
ures 5-2 and 5-3 using data from the AAS at the Bethlehem Steel/Coke Plant
Other possible values for k, K , Y and k will now be discussed.
5.1.2 Values of the Kinetic Constants. Reports are available on
three wastewaters of interest. The first is coke oven liquor investigated
at Bethlehem Steel . Biokinetic constants were not derived in Reference 23,
but we have Blotted the data from Reference 23 in Figures 5-2 and 5-3.
The data and the plots both use phenol as the substrate. To convert to
BOD we used 2 Ib BOD = 1 Ib phenol which is the usual experience (the
theoretical oxygen demand is 2.38 Ib/lb phenol). This gives the constants
shown as Set No. 1 in Table 5-1. The constants shown as Set No. 2 were derived
by us from experimental results on an H-Coal wastewater and the constants
shown as Set No. 3 are for the Synthane process using a Montana Rosebud
37 38
coal ' . The sets of constants are quite different.
TABLE 5-1 BIOKINETIC CONSTANTS
k, g soluble BOD/g MLVSS day
K , mg BOD/1
S
Y, g VSS/g BOD processed
k , day
d
Temperature °C
*Estimated on the basis of 1 g TOC ~ 2 g BOD
136
No. 1
Coke Oven
Liquor
1.8
0.34
0.2
0.17
27-32
No. 2
H-Coal
Wastewater
^ 0.3
^ 30
0.48
0.03
23
No. 3
Syn thane
Wastewater
•^ l*
•^ 266*
0.37
0.033
-------
1.0 I—
0.8
X dt ' 0.6
Ib phenol
Ib MLSS-day
0.4
0.2
k = 0,9
Ib phenol
Ib MLSS.day
K =0.17 mg/1
s
0.2
0.4
0.6
S, mg/1 phenol
Figure 5-2. Substrate utilization vs. substrate con-
centration (based on data from Ref.23).
137
-------
0.2
0.1 —
X dt
0.1 —
0.2 »—
Ib sludge
Ib phenol processed
Figure 5-3. Sludge growth rate vs. substrate utiliza-
tion rate (based on data from Ref.23).
138
-------
In Table 5-2 analyses from H-Coal waste, from Synthane waste, and a
typical coke oven liquor are shown. The coke oven liquor analysis is not
from Bethlehem Steel because it was not given in Reference 23. The
analysis of the Synthane wastewater is from a parallel publication.
TABLE 5-2 ANALYSES OF WATERS FOR BIOKINETIC DATA
39 25 38
Typical Coke H-Coal Synthane
Oven Liquor Wastewater Wastewater
BOD 4,000 - 6,000 1,900 - 2,600
COD 7,000 3,100 - 4,200 5,700
Phenol 2,000 750 - 1,500 1,200
In the examples we will show the effect that these different sets of
constants have on the size of the aeration basin. It is preferable that
the water from a given plant be tested. Lacking tests, Set No. 3 can be
used for gas plants and Set No. 2 for liquifaction plants.
5.1.3 The Flow Scheme. Activated sludge processes include the two
major components shown in Figure 5-1, the bioreactor followed by the
solids separator. These components should be considered an integral
system in terms of treatment performance, design and operation of the system.
Certain features of the flow scheme in Figure 5-1 affect the design
equations and require discussion before these equations are given. The
first feature is that the removal of waste is assumed to be accomplished
in the bioreactor only, and the solids separator is assumed to have no
effect on the effluent quality S . These assumptions may well be valid
if the waste concentration is measured as specific soluble compounds such
as phenol. However, if certain gross parameters like total BOD or TOC
(dissolved plus suspended forms) are used as the measure of efficiency,
the effect of the solids separator will have to be considered. Efficiency
is defined as:
100(S - SJ
E - s° i (4)
o
where S = influent waste concentration, mass/volume
S = effluent waste concentration, mass/volume
139
-------
The concentration of biological solids in the effluent is a function of their
settling properties and the design and operation of the solids separator.
Any release of biological solids in the effluent would tend to reduce E
as measured by gross parameters. For coal conversion wastewaters it is
assumed that the major pollutants will be soluble organics and that the
release of biological solids will be minimized and insignificant. Therefore
the conditions depicted in Figure 5-1 will be applied in the following
discussions.
The remaining features of the flow scheme are:
The bioreactor is completely mixed
There are no active biological solids in the raw waste entering the
bioreactor
The total active biological solids in the system equals the biological
mass in the bioreactor.
The design and control equations developed in the following sections
are specific ;o the flow diagram of Figure 5-1.
5.1.4 Relationships for the Mean Cell Residence Time. The best
parameter for the design and operation of an activated sludge process is
probably the mean cell residence time 9 , which is defined as:
XT
c (AX/At)T
in which X = total active microbial mass in the treatment system, mass;
and (AX/At) = total quantity of active microbial mass withdrawn daily,
including those solids purposely wasted as well as those lost in the
effluent, mass/time. In essence, 9 is a measure of the average retention
time of active microoorganisms in the treatment system, and we will show
that a desired treatment efficiency, E, can be obtained by controlling 9 .
c
Furthermore, 0 is readily measurable and easily controlled in the
C ',
operation of an activated sludge process.
One particular value of 9 , 0™, is of major concern in the design
and operation of activated sludge process. 9™ represents the minimum 9
c c
at which a complete failure of the biological process will occur. Below
9 , the organisms are removed from the system at a rate greater than their
140
-------
synthesis rate so that eventually no organisms will be left in the system.
6 , the 6 value to be used for design, must be significantly greater
c c
than 6 . The ratio (6 /Q ) gives the safety factor for the system. The
c c c
evaluation of 6 and 6 will now be made for the flow scheme in Fiqure 5-1.
c c
For this system, Q , as defined in Equation (5) is:
e = vx
c wX + QX (6)
r e
in which the various symbols are defined in Figure 5-1. Usually the value of
X may be assumed to be negligible, and Equation (5) then leads to:
6 ~ ^2L
c w X (7)
r
A mat-rial balance for microbial mass around the entire AAS system
gives:
(net rate of change of microbial mass) = (growth rate)
- (washout rate)
that is,
(ff) (V) = [Y (f) - kdX] V - (wXr + QXe) (8)
Under steady state, (—) will be equal to zero, and substituting for
(wX + QX ) from Equation (6) will lead to a relationship between 6 and
substrate utilization rate:
1 Y dt _
e x kd O)
141
-------
Substituting for — from Equation (3) results in a relationship between 9c and
effluent substrate concentration S :
YkS
1 - k, (10)
- ,
9 K + S, d
c s I
Vkd+ 1/ec> V1 + kd9c) (ID
si Yk - k - i/e e (Yk - k ) - i
dec d
Equation (9) may also be written as:
- = Y (F/M) - k., (12)
9 d
c
in which F/M is food to microorganism ratio usually in Ibs substrate per Ib
biological solids per day.
Equation (11) shows how control of 9 controls the effluent concentration
C
5 and that as 9 is reduced, S increases, and if 9 is made small enough,
then S1 will equal S and the process will fail. Thus the value of 9 may be
obtained by setting S = S in Equation (10):
YkS
7m- = F-TT - kd (13)
9 so
c
when S » K , a limiting value of 9 is:
os ^ c
(8™) lim. = (Yk - kd)-1 (14)
Given the values of Y, k, K , k and S , then 6™ may be calculated by
s d o c
using Equation (13) or Equation (14). The calculated value of 9™ represents
c
the lower bound of 9 allowable for operation. The actual value of 9
c c
preferably should be more than twice 9 if adequate control is to be achieved.
142
-------
There is a second lower limit to 6 . The sludge residence time must be
c
more than the hydraulic residence time or there will be no recycle of sludge
and no control. This limit is important and it is further discussed with
numerical examples.
In fact, there is an upper bound to Q . The settling characteristics of
activated sludge get worse when 6 is greater than about fifteen
days . Although it might appear theoretically possible to use very large
values of 0 and to compensate by using large solids separators, this usually
C
leads to inoperable solids separators and is not practiced. 0 is usually
c
controlled somewhere between 20 and 15 days. However, the upper bound of
28
Q may be somewhat higher than 15 days for pure oxygen activated sludge ,
probably because of the improved settling properties of the biological floes.
The preferred limit of 15 days is based on municipal, not industrial,
experience. In fact, one set of biokinetic constants that might apply to
coal conversion waste force mean cell residence times of up to 70 days or
more.
5.1.5 Sludge Recycle. A material balance for microbial mass around the
solids separator gives:
(Q -I- q + w)X = (q + w)X + QX (15)
Assuming X to be negligible, that is,perfect solids capture by the separator,
and letting p- = r, the recycle ratio, Equation (15) gives:
= x ( - «• W/Q )
r V + r + w/Q; (16)
or, , w I r _
0 V X
r = ^—-J *- = (using Equation 7)
A
_£_ _ i
X X
X
_£ ,
X X (17)
Equations (16) and (17) can also be derived by a material balance for
microbial mass around the bioractor alone.
143
-------
5.1.6 Aeration Basin Volume. Under steady state conditions the rate of
~ " dS_
dt'
jq C
substrate utilization by microorganisms, —, may be related to the actual
substrate removal as follows:
as o
dt V
Incorporating Equation (18) into Equation (9) leads to:
l_+k = : d_c = Y
e d e x v
c c
and
YQQ (S - S )
XV = XC+ ° (19)
d c
In an operating system, Equation (19) is a relationship between 0 and X.
c
For the designer, Equation (19) allows determination of the aeration basin
volume if X is chosen.
The microbial concentration X is usually expressed as mixed liquor
volatile suspended solids (MLVSS) in mg/1. The value of X cannot be chosen
arbitrarily since there are practical constraints due to the requirements for
oxygen transfer and liquid-solid separation. For air activated sludge systems,
X is generally within the range of 1,500 to 5,000 mg/1. The upper bound of X
is higher for high purity oxygen systems.
5.1.7 Summary of Aeration Basin Equations.
To design a process the designer is given the influent concentration, S ,
o
and the feed rate, g. Preferably there exist reasonable estimates of the
biokinetic constants k, k , K and Y. The minimum value of the mean cell
d s
residence time can be calculated frOm Equation (13) :
, YkS
1 _ o
- k
,,m K + S d
G so
c
144
-------
For a series of values of the mean cell residence time, 0 , the following
c
can tie calculated; the effluent concentration from Equation (11) :
K (1 + k,0 )
_ s d c
O _ ~~
1 6 (Yk - k,) -1
c a
the aeration basin volume for a series of chosen values of MLSS from Equation
XV =
the recycle ratio, for which an experimentally determined value of recycle
concentration is required, from Equation (17):
e«c
X
X~
Xr
the flow rate of waste sludge, w, and the waste sludge production rate, P ,
X
from Equations (7) and (20):
_ V_ X_
W ~ 0 X
c r
P = XV
x ——
c (20)
If biokinetic data are lacking, they can be estimated from experience on
related wastes. In an extreme case of lack of information, the aeration basin
must be sized on some such criterion as hydraulic residence time, but this is
unreliable.
The designer has two variables at his disposal, the mean cell residence
time, 0 , and the mixed liquor suspended solids concentration, X. 0 must be
C *•*
greater than the minimum value, 0™, must be greater than the hydraulic resi-
dence time, and must be long enough to give an acceptable effluent concen-
trating S . The chose value of X may affect the recycle sludge concentration,
X , and therefore the recycle rate. Both 0 and X affect the aeration basin
r' • * c
145
-------
volume and the clarifier volume. The final choice, preferably made on
cost, requires an understanding of the clarifier. We return to this
summary in the following section on solid-liquid separation.
5.1.8 Example 1. Aeration Basin Volume for Moderate Strength Wastewater
Calculations were made for a demonstration size (50 x 10 scf/day)
pipeline gas plant having a wastewater flow of 0.3 x 10 gal/day with
2,500 mg BOD/1 (which could be a Hygas plant using an eastern coal). The
sequence of equations listed above were used with the three sets of
biokinetic coefficients from Table 5-1. The recycle sludge concentration
was taken from Equation (37), X = 3.4X. The results are shown in Table 5-3.
If the biokinetic constants for coke oven liquor, Set No. 1, were correct,
a low value of effluent BOD, S , could be obtained for modest values of mean
cell residence time (for example, 6 = 10 days) in an aeration vessel of
about 0.2 x 10 gallons.
Calculations based on Set No. 2, for H-Coal wastewater, are presented
next. In changing constants we have more than doubled the production rate
of sludge (Y increased from 0.2 to 0.48) and very much decreased its decay
rate (k,decreased from 0.17 to 0.02). At the same time the rate of oxidation
d
of BOD has decreased (k decreased from 1.8 to 0.3). The lower oxidation
rate will mean a larger aeration basin and the higher sludge production
rate will mean that the waste sludge flow rate, w, must be large or
operation must be at large sludge concentrations, X.
Because this is a gas plant, we suggest using biokinetic Set No. 3,
for Synthane wastewater. The aeration basin would be 0.8 x 10 gals
(107,000 ft ). The system could be operated at a mixed liquor suspended
solids concentration, X, between about 3,600 and 4,000 mg/1 and a mean cell
residence time, 9 , between about 12 and 20 days. The recycle pipes and
c
pumps should be designed for a ratio of 0.38 (0.38 x 0.3 x 1Q6 = 114,000
gals/day = 79 gals/min). The waste sludge pipes and pumps should be designed
for 0.026 x 10 gals/day = 18 gals/min. The dry weight of waste sludge
is equal to
wx = |*
r 6
c
The maximum expected rate is about
0.8 (10 gals) x 8.33 Ib/gal x 4,000 ppm v 12 days = 2,200 Ib/day
146
-------
TABLE 5-3 CALCULATIONS ON AN AIR ACTIVATED SLUDGE PLANT, EXAMPLE 1
FEED 0.3 X 10 GAL/DAY AT 2,500 MG BOD/L.
Biokinctic Constants Set No. 1, Coke Oven Liquor, em - 5.3 days
9c
(days)
10
12
15
Biokinetic
ec
(days)
15
20
25
Biokinetic
e
c
(days)
10
12
15
20
25
Sl
(mg/1)
1.02
0.81
0.65
Constants Set
Sl
(mg/1)
61
38
28
Constants Set
Sl
(mg/1)
149
122
98
77
65
V
X=2,000
0.28
0.30
0.32
(106 gals)
3 , 000 4 ,
0.19 0
0.20 0
0.21 0
No. 2, H-Coal, 6m »
c
V
X-=2,000
1.82
2.22
2.54
(106 gals)
3 , 000 4 ,
1.21 0
1.48 1
1.69 1
No. 3, Synthane, e <
V
X*2,000
0.98
1.13
1.34
1.62
1.85
(106 gals)
000 X-2,000
.14 0.39
.15 0.39
.16 0.40
8.9 days
000 X-2,000
.91 0.30
.11 0.31
.27 0.32
•3.3 days
3,000 4,000 X-2,000
0.65 0
0.76 0
0.89 0
1.08 0
1.23 0
.49 0.32
.57 0.32
.67 0.33
.81 0.34
.93 0.34
r
3,000
0.40
0.40
0.40
r
3,000
0.34
0.34
0.35
r
3,000
0.35
0.35
0.36
0.36
0.37
4,000
0.40
0.41
0.41
4,000
0.36
0.36
0.37
•
4,000
0.37
0.37
0.37
0.38
0.38
w (10 gals/day)
0.0082 0.0054 0.0041
0.0073 0.0048 0.0036
0.0062 0.0041 0.0031
w (10 gals/day)
X-2,000 3,000 4,000
0.036 0.024 0.018
0.033 0.022 0.016
0.020 0.030 0.015
w (10 gals/day)
X-2,000 3,000 4,000
0.029 0.019 0.014
0.028 0.019 0.014
0.026 0.017 0.013
0.024 0.016 0.012
0.022 0.014 0.011
147
-------
5.1.9 Example 2. Aeration Basin Volume for High Strength Wastewater
These calculations were made fisr a demonstration size (50 x 10
scf/day) pipeline gas plant having a wastewater flow of 0.17 x 10 gals/day
at 18,000 mg BOD/1 (which could be a Hygas plant using a Montana coal).
In Table 5-4 we have shown the hydraulic liquid residence time calculated
from
t = V/(Q + w + rq) (21)
This is a very concentrated waste and alternative procedures to
remove organic contamination, such as solvent extraction or resin absorption,
should be considered. With the slower kinetics of Set Nos. 2 and 3, large
hydraulic residence times are needed to degrade the waste. Attempts to
design for moderate mean cell residence times result mathematically in
negative recycles when the cell residence time is less than the hydraulic
residence time.
Assuming again that Set No. 3 for Synthane wastewater is taken to be
correct, the aeration basin would be 4.6 x 10 gallons (615,000 ft ) and
operation could be for X between 3,000 and 4,000 mg/1, 0 between 20 and
35 days with a 19-day hydraulic residence time. This means a large basin
and we suggest that alternatives to biotreatment be considered. The
recycle ratio may reach 0.2; recycle flow up to 34,000 gals/day = 24 gals/
min. The waste sludge rate may reach 70,000 gals/day = 49 gals/min; the
dry waste sludge rate may reach 7,700 Ib/day.
5.1.10 Multistage Aeration Basins
An aeration basin may or may not be well mixed. In a well mixed
basin the concentrations are those of the effluent. In the feed water if
the contaminant concentrations exceed the toxic level, then good mixing
is required. For other situations an approach to plug flow or the use of
multistages can be advantageous. The rate of disappearance of contaminant
substrate was given by Equation (3) :
1 dS kS
X dt K +S
s
When S is much greater than K , the reaction rate is independent of S.
S
This is a zero order reaction and the rate is the same whether the reactor
148
-------
TABLE 5-4 CALCULATION ON AN AIR ACTIVATED SLUDGE PLANT, EXAMPLE 2
Feed O.17 K 10 gal/day at 18,000 mg BOD/1.
Biokinetics Set No. 1, Coke Oven Liquor, e « 5-9 days
en
(days)
10
15
20
Sl
(mg/1)
1.02
0.65
0.53
6
V(10
X-2,000
1.13
1.29
1.39
gals)
3,000
0.76
0.86
0.93
Biokinestics Set No. 2, H-Coal, I
e
c
(days)
15
20
35
60
70
Sl
(mg/1)
61
38
21
14
13
Biokinetics Set
e
c
(days)
15
20
25
35
45
lmg/11
98
77
65
53
47
V(106
X=2,000
7.57
9.16
12.52
15.72
16.57
gals)
3,000
5.04
6.11
8.35
10.48
11.04
No. 3, Synthane,
VI 106
X=2,000
5.65
6.79
7.72
9.16
10.2
gals)
3,000
3.77
4.53
5.15
6. 11
6.81
4,000 X=2,000
0.57 0.22
0.65 0.26
0.70 0.30
a" - 9.1 days
4,000 X-2,000
3.79 -0.46
4.58 -0.376
6.26 -0.20
7.86 -0.04
8.29 +0.007
~m ^ « ,
8 =-3.0 days
4,OOO X=2,000
2.82 -0.23
3.40 -0.17
3.86 -0.117
4.58 -0.04
5.12 +0.02
r
3,000
0.29
0.32
0.34
TC
3,000
-0.17
-0.11
+0.004
+0.11
+0.14
r
3,000
-0.017
+0.025
+0.061
+0.11
0.15
4,000
0.32
0.35
0.36
4,000
-0.02
+0.02
+0.11
+0.19
+0.21
4,000
+0.09
+0.12
+0.15
+0.19
0.27
v(106
X-2,000
0.033
0.025
0.020
w(106
X-2,000
0.15
0.13
0.11
0.08
0.07
w(106
X-2,000
0.11
0.11
0.09
0.08
0.07
gals/day)
3,000 4,
0.022 0.
0.017 0.
0.014 0.
gals/day)
3,000 4,
0.10 0
0.09 0
0.07 0
0.05 0
0.05 0
gals/day)
Liquid Residence
Time (days)
000
017
013
010
X=2,000 3
4.7
5.4
5.8
,000 4,000
3.1 2.4
3.6 2.7
3.8 2.9
Liquid Residence
Time (days)
000
.07
.07
.05
.04
.03
X=2,000 3
31
38
52
65
69
,000 4,000
21 16
28 19
35 26
44 33
46 34
Liquid Residence
Time (days)
3,000 4,000
0.07 0
0.07 0
0.06 0
0.05 0
0.04 0
.06
.05
.05
.04
.03
X=2,000 3
23
28
32
38
42
, 000 4 , 000
16 12
19 14
21 16
25 19
28 21
-------
is plug flow or well mixed. The principle is discussed more fully in
standard texts such as Reference 40. If S is comparable to Kg, then the
reaction rate depends on S and will be higher when S is larger. A plug
flow reactor will accomplish the same conversion in a smaller total
volume than a well mixed reactor. Particularly for the biokinetics of
the Synthane wastewater, K is quite large. Some calculations for
Example 1 are given in Table 5-5 for multistages. GC must be increased
from stage to stage to reduce the effluent BOD, S^ For low effluent
concentrations there is a savings in using two stages.
In at least one published case, it has been experimentally demon-
strated that the configuration of tanks-in-series can achieve better
performance for activated sludge systems in terms of process stability.
reliability and efficiency. These process improvements are attributed to
the change in flow pattern and fluid regime provided by the serial staging
„• 44
configuration
This does not necessarily mean that two stages can or should be used.
As discussed in Section 6, quite a high BOD is acceptable in cooling water
makeup. Of course this increases the trouble and expense of operating the
cooling system but it decreases the expense of the biological treatment.
Also, it is not always possible to use low values of & to obtain low
conversions in the first stage. Q must be larger than the hydraulic resi-
dence time or a negative recycle of sludge must be used. This means
recycling clean water which is possible, but wasteful. In the second
example with the concentrated wastewater, calculations for multistages
could not be made.
In practice, the aeration basin in Example 1 should be so shaped and
mixed as to encourage plug flow. In Example 2, plug flow must be discouraged
by the shape and mixing pattern of the aeration basin and the entry and
exit points of the water. The wastewater is so concentrated that it is
certainly toxic and must be diluted by mixing in the aeration basin.
Unless we are quite sure that the biokinetic constants have been accurately
measured on the wastewater in question, a second stage does give protection
and might be useful in Example 2. A sample calculation for a second stage is
shown in Table 5-6. A basin of about 0.1 x 10 gallons = 13,400 ft3 would
suffice, which is 2% of the first stage basin. The recycle ratio is 0.41;
48 gals/min recycle. The waste sludge rate is 0.4 gals/min.
150
-------
TABLE 5-5 CALCULATIONS ON A MULTISTAGE AIR ACTIVATED
SLUDGE PLANT FOR EXAMPLE 1
Feed: 0.3 x 10 gals/day at 2,500 mg/1 BOD
Biokinetic constants, Set No. 3, Synthane, 6 =3.3 days
c
MLSS: 3,000 mg/1
(106 gals)
1.233
0.405
0.125
0.530
Three stages
5 2500 452 0.325
12 452 122 0.104
25 122 65 0.024
0.458
e
c
(days)
One stage
25
Two stages
6
25
S
o
(mg/1)
2500
2500
312
Sl
mg/i
65
312
65
151
-------
TABLE 5-6 SECOND STAGE CALCULATION FOR EXAMPLE NO. 2 USING
BIOKINETIC SET NO. 3, SYNTHANE WASTEWATER
Inlet flow
Inlet BOD,
MLSS
e
c
(days)
20
35
70
, Q = 0.17 x 10
S = 300 mg/1
, X = 3,000 mg/1
S
(mg/1) V
77
53
39
gals/day
(106 gals)
0.056
0.084
0.116
r w (10 gals/day)
0.41 0.0008
0.41 0.0007
0.41 0.0005
Liquid Residence
Time (days)
0.23
0.35
0.48
to
-------
5.2. Solid-Liquid Separation
The solid-liquid separator, usually called a clarifier, is a critical
and integral part of the activated sludge process. The separator should be
designed to satisfy two functions: 1) clarification, which is the produc-
tion of an overflow relatively free of suspended solids; and 2) thickening,
which is the production of an underflow that contains, in sufficiently high
concentration, the active biological solids to be returned to the bioreactor.
5.2.1 Clarification
The requirement of clarification has long been the sole basis of
clarifier design, and the design criteria are in the form of a hydraulic
overflow rate in gallons per day per square feet. The procedure for
determining the unit area required for sedimentation involves the measure-
ment of the settling rate of the sludge in question, using the laboratory
settling rate procedure. The settling rate expressed in feet per hour can
be converted to an overflow rate by means of the following equation:
OR = (SR ft/hr) x 7.45 gal/ft3 x 24 hr/day (22)
2
where, OR = overflow rate (gallons/day/ft )
SR = settling rate (ft/hr)
This overflow rate based on the laboratory settling rate thus determines
the maximum overflow rate which can be expected to produce a reasoriably
clarified effluent. Hydraulic considerations and short-circuiting introduce
inefficiencies in the final clarifier which should be accounted for in the
design. An adjusted overflow rate can be calculated which will make
allowances for these inefficiencies.
ORD = K X OR (23)
2
where, ORD = design overflow rate (gpd/ft )
2
OR = laboratory overflow rate (gpd/ft )
K = a dimensionless correction factor
The magnitude of the constant K is a function of the tank configuration,
the inlet and outlet design, and the hydraulic efficiency of the clarifiers.
153
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Although exact measurements of K. are not available, the magnitude of this
factor ranges from about 0.5 to 0.8 for the various final settlers which
are commercially available.
5.2.2 Theory of Thickening
The relationship between thickening and the clarifier design is more
complex and was not fully recognized until the early 1970's. (A complete
summary to 1972 is given in Reference 24) . In Reference 24 it was found
that the settling characteristics of municipal plants sludges varied from
plant to plant and also varied with time in any one plant. (This variation
with time is because municipal plants deal with a wastewater whose charac-
teristics vary with time). Thus, an understanding of thickening requires
experiments on the specific wastewater. The following theoretical dis-
cussion is intended to help correlate experiments, however, since we know
of no experiments on coal conversion wastewater, the theory cannot be used
for design at this moment.
Today, final separators must be designed for an arbitrary overflow
rate and the concentration of the thickened sludge must be arbitrarily
assumed. If the size of the separator required to give the desired degree
of thickening exceeds the size required for clarification, the thickening
function controls the sizing of the separator. Conversely, if the clarifi-
cation requirement is larger, the clarification controls. The hydraulic
overflow rate should be determined for both clarification and thickening,
and the larger clarifier should be used.
The theory of solids separation for thickening has been well docu-
mented . The rate of solids transport to the bottom of the separator,
M, in pounds of solids per unit time, is directly proportional to the
cross-sectional area of the separator, A, and the rate at which solids
reach the bottom of the separator per unit area; that is, the solids flux
G, commonly in pounds per square foot per day:
G = M/A (24)
Generally there will be a limiting solids flux, G_ , whose value depends
L
on the solids settling characteristics and other parameters, and on a
solids loading rate, Mg, being applied to the separator. The goal of
154
-------
design is to provide an adequate area A so that the applied solids flux:
= jL = (Q + q + w)X = Q(l + r + w/Q)X
a L A A A (25)
If the area is inadequate and the applied solids flux, G , becomes
3i
larger than G , then the solids tend to build up and are eventually lost
L
through the overflow. This will cause the effluent water quality to
deteriorate. On the other hand, if a more than adequate area is provided,
then G is less than'G,. and the functions of thickening and clarification
a L
may be simultaneously satisfied. If anaerobic denitrification can be
avoided in the separator by a quick withdrawal of the concentrated sludge,
an overdesign of the separator should have no detrimental impact except for
its higher cost.
To evaluate G , a general expression for G is needed. Qualitatively
L
the solids are transported to the bottom of the separator by two mechanisms:
gravitational subsidence, and the bulk downward transport due to sludge
withdrawal from the bottom of the separator. The total solids flux at
which solids of concentration X. pass downward in the separator is:
G = X.V. + X.U (26)
in which V. is the settling velocity of the solids at concentration X., and
U is the average downward velocity due to the removal of sludge at concen-
tration X , the solids concentration being returned to the bioreactor. The
value of U depends on the flow rate of recycled and wasted sludge, q + w,
and A, so:
(27)
The value of V. is a function of the solids concentration, C, , and a
commonly accepted expression for V. is:
V± = aC."n (28)
155
-------
in which a and n are empirical coefficients. Larger values of a represent
sludges settling well. Larger values of n represent sludges with settling
properties very sensitive to changes in C (n is greater than one) . If no
inert solids are applied to the separator, then C. is equal to X. , the
biological solids concentration or MLSS. Since coal gasification wastes
contain primarily soluble organics and little inert solids:
V± = aX± (29)
A typical relation between V. and X. is shown in Figure 5-4.
Substitution of Equations (27) and (28) into Equation (26) gives:
G- ax.- + x. (30)
1 A 1
A typical graphical presentation of Equation (30) is shown in Figure 5-5.
Equation (30) is not applicable for very small values of X. because the
11 \
first term, aX. , goes to infinity when n > 1. On Figure 5-5 the
actual curves are shown with dashed lines for small X. . The total flux is
shown as the sum of its components, X.V. and X.U, and there will be a
minimum which acts as the limiting solids flux, G , with a corresponding
L
solids concentration of X_ . The lower solids flux when X. is less than X
ij i m
is of no practical concern because the MLSS in the influent to the solids
separator is usually larger than X and less than x • To reach the
m ii
returned solids concentration, X , the solids concentration in the separator
will have to pass through X^ and, thus, G becomes the maximum solids flux
which can be transported through the bottom of the solids separator.
In Figure 5-5 we find that X corresponds to the value of X when:
L i
3G „ „ 3 -G
•r^- = 0 and > 0
3Xi 3X±2
156
-------
100 i—
U-l
o
r-l
0)
cn
c
•H
0)
CT>
"S
•H
-p
•H
C
H
10
0,1
V. = aX.
i i
-n
0.001
0.01
0.1
x., Initial Solids Concentration, Ib/lb
i
Figure 5-4. Typical settling velocity vs. solids concentration.
157
-------
o
to
SOLID CONCENTRATION, X.
Figure 5-5. Typical curve of solids flux vs. solids concentration.
158
-------
Differentiation of Equation (30) gives:
|f-= a(l-n)X.-n
and by setting the righthand side of Equation (31) equal to zero, X can be
J_i
determined as:
w
Differentiation of Equation (31) gives:
32G an(n-l)
<33)
As long as n > 1:
3X.
i
and Equation (32) gives the X_ corresponding to G_ . Substituting Equation
L Li
(32) into Equation (30) :
f , ..""I 1/n . n . ,rQ + w . (n-l)/n
= a(n-l) (—H-^i-.—, (34)
The hydraulic overflow rate required to satisfy the thickening function
can be obtained by equating G from Equations 034) and (25) to obtain:
Li
159
-------
n
n ,
. n-1
a (n-1) n-1 (r + w/Q)
A Xn
or, since (Q + q + w)X = (q + w)Xr
that is (Q + rQ + w)X = (rQ + w)Xr
£ _ a (n-1)
A n
A Xrn(r+w/Q)
5.2.3 Continuation of Summary of Design Equations
Equation (35) or (36) can be used to find the thickener area, A. In
these equations the terms, a and n, are experimentally determined constants
in the settling Equation (29) . The mixed liquor suspended solids concen-
tration, X, is set by design. The feed rate, Q, is given. The recycle
ratio, r, and the wasted sludge rate, w, are from the equations for the
aeration basin previously given. The return sludge concentration, X , is
experimentally determined; it cannot be found from Equations (35) and (36)
because they are not independent equations.
Adequate clarification of the effluent water is also required. If the
area of the clarifier determined from the equations above does not give
satisfactory water, then the area must be made bigger according to the
clarification requirement.
Equation (35) or (36) must, of course, be used with the design equations
for the aeration basin previously given; in particular with Equation (17)
which relates the recycle ratio, r, to the recycle concentration, X . If
the dependence of X on Q/A is known, it will be necessary to iterate
between Equations (17) and (35) until a consistent set of values for X , r
and Q/A is found.
At this point it begins to be possible to estimate the cost of the
system for chosen values of 9 and X. The cost of the aeration basin is
not so easy to determine because it is not just a function of volume, but
160
-------
depends also on shape, which is controlled by choice of aeration equipment
(discussed in the next section). At first a simplified cost estimate
dependent only on volume will be used. The cost of the clarifier is mostly
a function of the area. In general, a larger X will mean a smaller
aeration basin (but not a smaller oxygen requirement) and a larger clarifier.
A minimum cost value for X can, in theory, be found. We must say that
experimental information to do such a perfect design is seldom available.
5.2.4 Values of the Thickening Constants
We have found no published results for thickening sludges from the
biological oxidation of coal conversion plant wastewater. One piece of
23
information has been reported for coke oven liquor , namely
X
= 3.4 (37)
This particular point was measured at X = 2,600 mg/1. Until better infor-
mation is available, Equation (37) should be used for air systems. For
oxygen, Union Carbide has suggested that X = 20,000 mg/1.
Also, until settling rates have been measured, an arbitrary overflow
rate should be used to find the area of the solids separator. The coke
23 2
oven liquor was clarified satisfactorily at 685 gal/ (day) (ft )- For an
2 25
H-Coal wastewater a design overflow of 560 gal/(day) (ft ) has been used
For municipal plants the Ten States Standards call for an overflow rate of
2
600-800 gal/ (day) (ft ) , but this may be too high for coal conversion
waste. Until better information is available, we suggest sizing final
2
clarifiers for safety at 400-500 gal/ (day) (ft ). If an intermediate
2
clarifier is. used, it can be sized at 600-800 gal/ (day) (ft ). These rates
are applicable whether air or oxygen is used.
5.2.5 Sludge Dewatering
Sludge dewatering is a big job in these plants. Available experience
is all on municipal waste and often when the sludge is partly primary and
partly secondary. In our example the sludge is entirely secondary and need
161
-------
not have the same dewatering characteristics. For many municipal secondary
sludges the dewatering rate on a vacuum filter is dependent on feed concen-
tration and the use of a dissolved air flotation thickener which can
multiply the sludge concentration about 2 to 2.5 times is justified. We do
not know if a thickener is justified for this sludge.
Filters require attention. One has the choice of using a small
filter, a small sludge holding basin, a lot of operating labor and some
standby equipment or, on the other hand, a large filter, a large sludge
holding basin, a little operating labor and no standby equipment. A
comparison is given in the following example. We suggest sizing the filter
2 2
for 3 Ib dry weight/(ft ) (hr) = 72 Ib dry weight/(ft ) (day) .
5.2.6 Continuation of Examples
Example 1. A flow diagram is given in Figure 5-6. The clarifier will
2 2
be sized for 424 gals/(day) (ft ) and have a 30 ft diameter and 707 ft
area. The dry sludge production rate may reach 2,200 Ib/day. Assuming
that a filter is operated 7 hrs/day every day of the week at 3 Ib dry
2 2
weight/(ft ) (hr) , the active filter area must be 105 ft . A rotary filter
of 6 ft diameter and 8 ft long has adequate active area. A 24-hour sludge
holding basin must be provided. The liquid volume is 0.026x10 gallons =
3,500 ft . To avoid having the sludge set-up in the basin, stirring is
necessary. The recycle flow passing through the basin may supply adequate
stirring. A clarifier mechanism is often used.
Example 2. This flow diagram is given in Figure 5-7. Two stages are
shown as discussed under Multistage Aeration Basins. The first clarifier
2 2
will be 18 ft diameter, 254 ft , 668 gals/(day)(ft ). The second clarifier
will be 22 ft diameter, 380 ft2, 447 gals/(day)(ft2).
The dry sludge rate is large and can reach 7.700 Ib/day. The follow-
ing options are typical of those that can be considered:
162
-------
M
NUTRIENTS
Q = 0.3 >k A
FOR POSSIBLE DILUTION AT START-UP
Flows in 106 Gal/day
AERATION BASIN
107,000 ft3 liquid
CLARIFIER
TREATED WATER
SLUDGE HOLDING BASIN
Figure 5-6- Flow diagram for Plant Example 1. - Moderate strength wastewater.
-------
Filter operated (hrs/day) 7 15 24
(days/week) 7 7 7
2
Active filter area required (ft ) 367 171 107
Installed area including
standby (ft2) 372 302 216
Possible filter configuration*
diameter (ft) 10 66
length (ft) 12 86
no. of filters 1 22
Sludge holding basin capacity
hours 24 24 12
ft3 9,360 9,360 4,680
Sludge holding basin size
side x side x depth(ft)** 34x34x10 34x34x10 24x24x10
*From Ametek catalog
**Includes 2 ft. free board
Two shift operation does not seem worthwhile and the choice is between
one shift and three shift. For this plant we suggest one shift operation.
A commercial scale plant would require three shift operation. Operation
for five days a week was not considered.
A simplified diagram is shown in Figure 5-8. Gravity flow through the
plant is assumed. A transfer pump, P6, can move sludge back and forth
between the stages for wasting and for start-up. Some sampling and instrumen-
tation points are shown in Figure 5-8 and Table 5-7. These are more
analyses than usual on demonstration size plant and they are intended to
provide experience for use on a larger plant. Standard sources of
analytical procedures are given in Table 5-8.
164
-------
FOR POSSIBLE DILUTION AT START-UP
NUTRIENTS
CLARIFIER
AERATION BASIN
615,000 ft3
SLUDGE HOLDING BASIN
Flows in 106 Gal/day
Figure 5-7. Flow diagram for Plant Example 2. - High strength wastewater.
-------
NUTRIENT
CLARIFIER TREATED WATER
V j
VACUUM SLUDGE
FILTER
I -A_IV-^_
FILTRATE
VV Sample 4 Instrument Points
P8
Figure 5-8. Simplified process diagram for air activated sludge. Example 2.
-------
TABLE 5-7 SAMPLING AND INSTRUMENTATION POINTS
10
Flow
Phenol
BOD
COD
TOG
SS
Temp.
PH
DO
Temp.
PH
SS
VSS
Flow
DO
SS
Temp.
Phenol
BOD
COD
TOC
SS
Flow
DO
SS
Temp.
Phenol
BOD
COD
TOC
SS
Flow Flow Flow & Flow
SS SS Totalizer SS
VSS SS VSS
-------
TABLE 5-8 SOURCES OF ANALYTICAL PROCEDURES
Analysis
pH
Alkalinity
Acidity
Specific Conductance
Chloride
Total Dissolved Solids
(Filterable)
Suspended Solids
(Non-filterable)
Dissolved Oxygen
BOD
COD
TOC
Phenol
Ammonia Nitrogen
TKN
Nitrite - N
Nitrate - N
Phosphorous
Calcium
Magnesium
Silica
Sulfate
Turbidity
Color
Residual Chlorine
Coliform
Temperature
Standard
Methods
14th. Ed.
460-465
278-282
273-277
71-75
302-309
92-93
94
440-454
543-550
550-557
532-534
574-592
407-418
437-440
434-436
418-433
466-483
185-191
221-223
484-492
493-498
113-139
64-71
309-345
913-942
ASTM
178-187
111-119
111-119
102-125
270-275
188-197
188-197
372-377
-
473-475
469-472
553-568
237-241
577-579
366-371
363-365
388-407
160-163
160-163
401-407
429-435
222-228
-
276-289
_
EPA42
239
3-8
1-2
275-276
29-34
266-267
268-272
51-68
11-12
20-28
236-238
241-248
159-174
175-196
215-216
197-214
249-265
19,103-104
114-115
274
277-283
295
36-39
35
_
286
168
-------
5.3 Aeration
Oxygen must be supplied to maintain the biological oxidation. To be
used the oxygen must be dissolved. A dissolved oxygen level of 1.5 to 2
mg/1 should be maintained in the aeration basin. We begin by discussing
the oxygen required and then the equipment that can be used to cause oxygen
from air to dissolve at the desired rate.
5.3.1 Oxygen Requirement. The oxygen requirement, pounds of oxygen
required per pound of BOD removed, is a function of F/M and COD/BOD ratios.
The effect of COD/BOD ratio may be of particular interest in these waters
since the fate of COD in the biological treatment of coal conversion conden-
sates is unknown at present. Ideally, this oxygen requirement should be
determined in pilot studies by one of the following methods: measurement of
oxygen utilization rate (OUR in mg/l/hr) of the biomass, or total COD
balance.
The OUR measurements should be made two or three times per day on the
mixed liqv>r from the bioreactor. The measurement involves the use of a
probe to uonitor the dissolved oxygen concentration to depletion in a mixed
liquor sample. The slope of the straight line section of the curve of
dissolved oxygen versus time is taken as the OUR.
Although the OUR measurement is simple to make, the result may be
inaccurate because the dissolved oxygen measurement using a probe is not
always exact. The inaccuracy may occur particularly with high BOD waters.
To make a more precise determination of the oxygen requirement, the COD
balance may be used, which is more complex.
The method of total COD balance involves a complete inventory of COD
around the biotreatment system, including the bioreactor and solids-liquid
separator. The oxygen requirement can be so determined since COD is
measured and calculated in oxygen equivalents. For a general situation the
COD fluxes around the system can be represented in the following diagram:
COD,.
COD .
out
COD.
accum.
CODwasted
169
-------
where:
COD. = total COD content of the influent to the bioreactor,
in excluding the recycled sludge;
COD = total COD content of the effluent from the solids-liquid
Ou separator;
COD = total COD content of the wasted sludge;
wasted
COD = total COD gained due to MLSS and clarifier solids inventory
accum.
increases.
The following equation then represents a COD balance around the system:
COD - COD - COD , - COD = oxygen requirement
in out wasted accum.
Theoretically, the COD in this equation will drop out if a true
accum.
study state can be maintained in the system. However, the MLSS and clarifier
solids blanket tend to vary. In actual operation, the total COD will be
determined on daily composite samples of influent, effluent, waste sludge and
mixed liquor; thus COD. , COD and COD , can be calculated. The total
in out wasted
COD determination of mixed liquor and measurement of MLVSS leads to a ratio
between COD and MLVSS, which can be used to evaluate COD . The COD
accum.
gained due to clarifier solids increase may be monitored by measuring the
thickness of the sludge blanket. One technique to convert the thickness
increase to a COD value is to assume half of the thickness increase as having
the solids concentration of the return or wasted sludge, and the other half
as having the solids concentration of the mixed liquor. The conversion of
solids to COD will be achieved by using the COD/MLVSS ratio previously
described.
Given the design flow, BOD feed concentration and required BOD removal,
the oxygen requirement in pounds of oxygen required per day can be calculated
using the ratio between pounds of oxygen required per pound of COD removed,
determined above and the corresponding BOD removal.
170
-------
5.3.2 Values for Oxygen Requirement
We know of no oxygen uptake rates or COD balances done on coal conver-
sion wastewater. When surface aerators were used in the treatment of coke
oven liquor , it was found that 18.2 Ib phenol were removed per day per
horsepower. This is about 43 Ib BOD removed/(day)(hp) which compares
closely with typical values in the literature of 45-50 Ib BOD/(day)(hp).
Pending additional information, this value must be used if surface aerators
are used. For diffused aeration, a true oxygen requirement is needed.
Assuming that a surface aerator transfers between 1.8 and 2.6 Ib O /hp-
hr, the figure of 43 Ib BOD/(day)(hp) translates to between 1 and 1.4 Ib
0 /Ib BOD removed. Pending additional information, a figure of 1.3 Ib
0 /Ib BOD should be used.
5.3.3 Aeration Equipment and Aeration Basin Shape
A variety of equipment is available to mix air and water and to
encourage the dissolution of oxygen. A useful classification of the
equipment ' includes
1. Surface aerators; These are rotating stirrers near the
surface of the water designed either to throw water up in a spray, or to
pull air down into a vortex.
2. Turbine mixers; These are rotating stirrers set near the
bottom of the water with compressed air released from a pipe just below
the stirrer.
3. Diffusion systems; These are spargers set on pipes at the
bottom of the aeration basin. If porous pipes are used, fine gas bubbles
are obtained. Coarse gas bubbles result when individual spargers are
used. The individual spargers are like mushroom caps over the top of open
vertical pipes. Individual static mixers are also used and these draw
water into the airstream before it is released. In all diffusion systems
the only moving machinery is the air compressor which is not in the water.
At this point in the design, the designer knows a required aeration
basin volume and a required rate of supply of oxygen. For any chosen
piece of equipment supplying oxygen, the energy used and the capital cost
will depend on the shape of the aeration basin; on ...ts length, width and
depth; whether partition walls are used, and whether tank baffles are
used. The basin shape and the aeration equipment must be matched to
171
-------
result in a system with the lowest cost. The value of the number Ib
O /(day) (ft3 of basin) is not constant for the industry and this means
that there is no one preferred type of aeration equipment. Furthermore,
all equipment of one type (for example, all surface aerators) is not
designed to operate with maximum efficiency in the same basin. Some
surface aerators are designed for optimum efficiency in wide, shallow
basins; others for deeper, narrower basins with wall baffles. The lengthy
process of matching equipment, basin and oxygen requirements cannot be
avoided. Manufacturers of aeration equipment are highly competitive and
will provide test results together with recommendations of the basin shape
preferred with their equipment. Taking into account the cost of the
basin, as well as the cost of the aeration equipment, the designer can
choose the least costly overall system.
Mixing is an important factor in the design and operation of aeration
systems. The purpose of mixing is to uniformly disperse dissolved oxygen
and activated sludge throughout the basin, and mixing requirements have
been related to the power density in aerator horsepower per unit volume of
aeration basin. It has been reported that 0.006 to 0.01 hp/1,000 gal is
required to maintain uniform dissolved oxygen, while the minimum hp/1,000
gal to prevent any bottom deposition of activated solids increases with
increasing aerator spacing. To maintain a minimum bottom velocity of 0.4-
0.5 ft/sec, at least 0.02-0.03 hp/1,000 gal is needed at a distance of 14
ft from the aerator, and 0.05-0.06 hp/1,000 gal at a distance of 35 ft.
These minimum power densities may be used as a design guide when no pilot
plant data are available with the specific surface aerators to be used. No
such general experience has been published on diffusion aerators and the
manufacturers must be asked. However, we foresee th-i the energy needed
to dissolve oxygen into coal conversion wastewaters will so far exceed the
minimum energy for mixing that mixing will not usually be a limitation.
The following section offers some general opinions on aeration equipment
by type, suggests the kind of information that most manufacturers can
supply, and shows how to use this information.
Surface Aerators
This type of aerator is often the cheapest, both to buy and ~,.o run.
Surface aerators also may be the noisiest type, the noise coming 1 -am the
172
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movement of water. Upflow, or spray type, surface aerators give the most
cooling of the water and are the most prone to form aerosols. Maintenance,
while not complex, does involve going out on the water. (If the water
level must vary in the aeration basin, then pier-mounted surface aerators
and turbines should not be used.)
Surface aerators are made for an upflow or a downflow of water. They
are made to float or to be mounted on concrete piers. Each type is most
efficient in a particular shape of basin. The performance of surface
aerators is rated in terms of their oxygenation efficiency as pounds of
oxygen per horsepower-hour at standard conditions. Standard conditions
refer to a test temperature of 20°C, zero initial dissolved oxygen concentra-
tion, and tap water as the test liquid. In addition, oxygenation efficiencies
reported by the manufacturers are usually determined by testing a single
aerator rather than multiple aerators as usually used in actual treatment.
Also, tests are made at a fixed depth, commonly in a basin 10 to 12 feet
deep; t? ^s depth frequently gives the lowest excavation cost.
To correct the results of the test under standard conditions, the
28
following equation may be used :
D/-» _ *-»
" - a (38)
where,
N = Ib O /hp-hr transferred under field conditions
N = Ib O /hp-hr transferred in water at 20°C and zero dissolved oxygen
8 = salinity-surface tension correction factor, usually 0.95
C = oxygen-saturation concentration for waste at given temperature
and altitude
C - operating oxygen concentration
L
T = temperature, °C
a = c ^gen-transfer correction factor for waste, usually 0.8 to 0.85
:>r wastewater
173
-------
Use of Equation (38) requires information better known to the plant
designer than to the aeration equipment manufacturer. The temperature
chosen will usually be a hot, summer temperature when oxygen solubility is
minimal. C applies at the chosen temperature. The constants $ and a
wait
can be found from tests or taken from experience.
Equation (38) says nothing about the basin shape and each manufacturer
should be asked to state the preferred shape. It is particularly important
that multiple aerators in the same basin not be so close to each other
that they interfere and lose efficiency .
Turbine Mixers
Test results and nomographs for designing have been given in Refer-
ence 1. Turbine mixers can be useful when a high oxygen rate is required
in a small basin, or if very deep basins are to be used; turbines give
better mixing than any other type of aerator in such basins.
Diffused Aeration
Fine bubble diffusers can cause the dissolution of about 25% more air
than coarse bubble diffusers, but they are easily clogged and maintenance
costs can often overcome their inherent advantage. Coarse bubble diffusers
(_and probably static mixers although less experience is available) have
the particular advantage of ease of maintenance. The only moving machinery
is the air compressor which is usually inside a building and not on the
water. Diffused aeration is particularly useful if dissolved oxygen is to
be controlled when the feed is varying . Piping procedures and a descrip-
tion of compressors is fully given in Reference 28. A series of tests on
static mixers has been published . Piping is straightforward. Pipes
should be large enough that frictional losses are less than one-fifth of
the total pressure loss; the main pressure loss equals the depth. Another
procedure frequently found to be economical is to allow 2 psi frictional
pressure drop. The main header can be sized for about 1,000 ft/sec and
the secondary lines for a lower velocity.
Placement of the diffusers and the surface shape of the water affect
not only the efficiency of absorbtion, but also the type of mixing that
results in the basin. Diffusers are often placed to one side to produce a
rolling motion in the mass of water in the basin. The mixing pattern, and
therefore placement of the diffusers, can determine whether the basin
174
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approaches closer to plug flow or to complete mixing. Plug flow is advanta-
geous with fairly dilute feeds, but good mixing is required with strong
feeds to prevent toxic contaminants from killing the system.
An extensive series of tests have shown that placement of the diffusers
and basin width affect the absorbtion efficiency . In these cases a
narrow basin and a midwidth position are preferred; however, this may not
be true for other manufacturers.
There is discussion on just how efficient a diffusion system will be
and how this is affected by depth. Certainly the efficiency of a diffusion
aeration system increases with increasing depth. The limit is where the
basin is so deep that the rising air bubbles cease giving adequate mixing.
The deeper the basin, the larger the fraction of the compressed air which
dissolves. In a deep basin less air needs to be compressed. The actual
relationship depends on the diffuser and can be supplied by the manufacturer.
For Kenics static mixers a rule of thumb is:
Percent of compressed air which dissolves = (depth in feet) - 2.5
(39)
The deeper the tank, the higher is the pressure required from the
compressor. The work done by the compressor at low pressures when the
compressibility factor equals one is given by :
HP = 0.00078
WT
Mne
P2 \n
P
-1
1
(40)
where W = massflow, Ib/hr
T = inlet temperature, °R
M = molecular weight = 28.8 for air
n = approx. 0.286 for air in adiabatic compression
approx. 0.408 for air in polytropic compression
e = efficiency =0.7
P,, P- = inlet and outlet absolute pressures.
175
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Adiabatic compression is probable and the horsepower is very close whether
adiabatic or polytropic compression is assumed. (Reference 28 gives the
formula for adiabatic compression.) If the inlet temperature is taken to
be 80°F = 540°R (which is necessary for design, but high for a year-round
average), then for air:
HP = 0.073 W
0.286
(41)
Equations (39) and (41) were used to determine the theoretical horsepower
to dissolve 1,000 Ib O /nr at various depths. The pressure was taken to be
(1.2 x depth). Air contains 23 wt% oxygen. The results are:
Depth (ft)
10
20
30
40
Fraction of
Air Which
Dissolves
0.075
0.175
0.275
0.375
Air to be
Compressed
(Ib/hr)
57,800
24,800
15,700
11,600
Pressure
Drop (psi)
5.2
10.4
15.6
20.9
Compressor
hp
382
300
264
244
These are calculated power requirements without allowance for inefficiencies.
The .real power requirements for surface aerators probably would fall in
the range 400-500 hp. The capital cost of compressors is very dependent
on the pressure ratio because the design changes if a higher pressure
outlet is required. Diffusion systems will tend to cost more than surface
aerators to install but less to run. Diffusion systems should be considered
in most situations.
5.3.4 Examples
A definitive study has not been made of aeration in the two examples
under consideration. The cost of excavation enters into the economics and
was not known. Surface aerators have been discussed with only one manu-
facturer, Bird Machinery, South Walpole, Mass., and the following is a
possibility, not a recommendation.
176
-------
Example 1. The horsepower for surface aeration is
2,500 Ib BOD 6 lb water J_ (hp) (day)
6 . x (0.3 x 8.33 x 10 ) . x ^ - 145 hp
10 lb water
Two 75 hp aerators can be used. The aeration basin which has 107,000
ft liquid capacity can be 20 ft deep x 52 ft wide x 103 ft long. This is
quite deep and draft tubes will be needed to ensure mixing near the bottom.
Some advantage of partial plug flow c ^n be taken by having the inflow and
outflow on the short sides.
Example 2. The horsepower for surface aeration is:
18,000 lb BOD .. ,_ . _,„ ,^6. lb water 1 (hp day) „„„ ,_
'— x 0.17 x 8.33 x 10 ) — x -rr— ,7. ** = 593 hp
..6 ., . day 43 lb BOD c
10 lb water *
Good mixing is required. Six 100 hp aerators can be installed in a
basin 20 ft deep x 143 ft wide x 215 ft long. Inflow and outflow should
be on the long side. Possible arrangements for both basins are shown in
Figure 5-9.
In both examples deep basins have been chosen to suit the particular
aerators under consideration; other possibilities should be sought.
5.4 Operation
Of all the treatments likely to be used for treating wastewater from
coal conversion, biological treatment appears to have the least firm
experimental basis for design. Therefore, a summary of operating techniques
follows with some emphasis on how to obtain better information for future
designs.
177
-------
143'
A *L
STEP 1.
215'-
t
^
fx,
< >
X
X J
X '
x^
X
/
\ *
1
T T T t
WATER DEPTH 20 ft
STEP 2.
40'
40'
X
WATER DEPTH 12 ft
lOOhp AERATORS
-INLET PIPE
OUTFLOW BOTH SIDES
30hp AERATOR
Figure 5-9. Aeration basins for air activated sludge, Example 2.
178
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5.4.1 Variables Under the Operator's Control
Activated sludge processes are affected by many process variables.
Once the plant is built, some of these are fixed. Fixei parameters include
the aeration basin volume, the aerators, the clarifier size, and the
wastewater flow and concentration. Parameters over which the operator has
some control are . nown in Table 5-9.
Parameter 1; Oxygen Utilization Rate (OUR)
Because oxygen usage is directly related to the removal of biodegradable
organics, the OUR in the aeration basin serves as a measure of the bio-
chemical activity in the system. The value of OUR will generally range
from 10 to 70 mg/l/hr. Excessively low OUR indicates low biochemical
activity, which may be due to non-biodegradability of the pollutant or
lack of active biota in the aeration basins. Too high an OUR, on the
other hand, might lead to too low a level of dissolved oxygen in the mixed
liquor, and oxygen transfer may be limited. The amount of biological
solids may be controlled by adjusting the return sludge flow rate, and OUR
may thus be controlled to a certain extent. The dissolved oxygen concentra-
tion in the mixed liquor should be maintained at no less than 2 mg/1.
In a high purity oxygen activated sludge system we must determine the
coefficients related to gas transfer. These include the Henry constants,
Ht, and the overall gas transfer coefficient, Ka, as shown in Section
5.5. The values of H. for various gas constituents in pure water can be
found from standard references., but the value of 1C a will have to be
determined in the actual operation because it is highly dependent on the
actual design and operation.
Under a steady state condition, the Equations of Table 5-15 may be
rearranged to yield:
a'
kX C.
4n
K + C.
s 4n
• - b'X
O
-v1 (c
n
" - C )
1~ 1 ^"1 „'
,n-l l,n
K a =
(HP - C )
1 l,n l,n
In the actual operation of a demonstration plant, all terms on the righthand
side of this Equation may be determined and the ILa can be calculated.
179
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TABLE 5-9 ACTIVATED SLUDGE PROCESS CONTROL PARAMETERS
Parameter
1. Oxygen Utilization Rate
Variability
2. Sludge Wasti-ig
3. Return Sludge Flow Rate
4. Depth to Clarifler Sludge
Blanket
5. Sludge Settling
Characteristic*
Some
within design
range
7 - 10 ft
Some (SVI)
50 to 100)
Control
Some
Some
Comment
Control may be
by adjusting return sludge
flow rate.
Control is based on the value
of Kludge residence time which
depends on the BOO removal and
solids separation to be achieved
in final clariflers.
Hourly adjustment to be considered
on the basis of sludge settling
characteristics and final clarifier
depth of sludge blanket.
Control through adjustment of
return sludge flow rate and
sludge wanting.
Control through adjustment of
sludge wasting.
To determine the actual oxygen requirement in terms of pounds of oxygen
required per pound of BOD removed, the COD balance technique discussed in
Section 3 is also applicable to the High Purity Oxygen Activated Sludge System.
Parameter 2; Sludge Wasting
Sludge wasting is one of the most important parameters over which the
operator has complete control. Its control is based on the value of 0 ,
c
which has been defined by Equations C.5) , (6) and (7). Equation (7) is
VX
0 =
c
wX
From this Equation the sludge wasting rate, w, can be calculated when the
values of V. X, X and 0 are known. V is fixed by design. X and X have
re r
to be actually measured by the operator. The value of 0 usually ranges
C
between 3 and 40 days, although a single © value is used for design
C
calculations. This design 0 value serves as a guideline for startup and
C
after several months of operation, the operator will be able to determine
what the optimum 0 range should be to produce an effluent of the desired
quality. The operator should also try to waste sludge as uniformly as possi-
ble, that is, to avoid wasting in surges so that the potential for unstable
conditions may be .minimized.
180
-------
Parameter 3: Return Sludge Flow Rate
The purpose of returning sludge from the clarifier to the aeration basin
is to sustain a desirable MLSS level to react with the biodegradable organics.
If the return sludge flow rate, represented by q in the design equations, is
too low, the following undesirable conditions may develop:
There may be insufficient MLSS to effectively react with the
incoming BOD in the aeration basin, causing a deterioration in
effluent quality.
The larger accumulation of sludge in the clarifier may create a
thicker sludge blanket at the clarifier bottom, which may allow
more solids to escape in the overflow.
The relationship between MLSS, i.e. X, and q or r (recycle ratio) was
previously derived through a mass balance in Equation (16) :
= X +
r w
1 + r + Q
—
With constant X and Q, Equation (16) shows that as r increases, X will
also increase, although not proportionally.
Too high a return sludge rate should also be avoided for the following
reasons :
As r increases, there will be less sludge consolidation in the
clarifier, and therefore X will decrease in Equation (15) .
As X increases, the sludge settling velocity, V, will decrease in
accordance with Equation (29) :
V. = axTn
i i
Eventually there will be a limiting solids flux, preventing any further
increase of solids transport from clarifier to aeration basin by increasing
the recycle ratio r. Qualitatively this has been shown in Figure 5-5, which
can be quantitatively determined by the operator.
Although the purpose of adjusting the return sludge rate is
to shift the biological solids from the clarifier to the aeration basin or
vice versa, limits mus t be observed at both high and low ends. The recycle
ratio r gene lly ranges from 15 to 100%.
181
-------
Parameter 4: Depth to Sludge Blanket in the Clarifier
The depth of sludge blanket provides the operator with an indication of
the sludge settling characteristics and the need for adjusting return sludge
rate and/or sludge wasting.
This parameter is determined by measuring the depth from the water
surface to the top of the sludge blanket using a sludge blanket depth finder
or equivalent devices and measurement should be done at least once every
eight-hour shift and can be made by inserting a transparent tube, closing the
top end, lifting out the full tube and measuring it.
Parameter 5; Sludge Settling Characteristics
Sludge volume index (SVI) may be used as a measure of the sludge settling
characteristics. This involves the standard settling test which is usually
conducted with the aeration basin mixed liquor in a one-liter graduated
cylinder. After 30 minutes of settling, the volume of sludge is measured, and
SVI in ml per gram can be calculated as follows:
SVKml/gm) = Volume of sludge (ml/1)
MLSS (mg/1) '
Sources for the required tests are given in Table 5-8.
In general, sludges with an SVI between 50 and 100 have good settling
characteristics. As SVI increases, the settleability becomes poorer. Common
control measures include adjusting 9 and/or chemical treatments, such as the
c
addition of flocculating agents.
5.4.2 Startup
Before the actual startup, all structural, mechanical and electrical
components should be carefully checked, especially the following items:
The drive mechanisms of the mechanical aerators and clarifier
scrapers and skimmers
All pumps
All gates and valves
All effluent weirs in the aeration basins and final clarifiers
should be level.
182
-------
The following sequence of steps should be followed to start:
(1) Fill the aeration basins with the bacteria seed and river water.
(Sources of bacteria will be discussed later.)
(2) Turn on all mechanical aerators.
(3) Slowly start the inflow of wastewater into the aeration basins,
and the aeration basin effluent will start to fill the clarifier.
The purpose is to build up sludge. The quickest measurement
is the oxygen uptake rate and we want to obtain a steady increase in
OUR. Let in wastewater at such a rate that OUR increases. We also
want a steady increase in MLVSS, but this takes longer to determine.
If the increase is not found, the first possibility is that the
sludge is not acclimatized, that is, the waste is toxic. In this
event slow down the feed. The effluent quality is not controlled
during start-up and if it is unsatisfactory, the effluent must be
returned to the equalization pond.
(4) When the clarifier is about 3/4 full, start the sludge collection
mechanisms and the recycle sludge pumps. Set the pumps at 100%
recycle to return all settled solids in the final clarifier to the
aeration basins to build up the activated sludge concentration, X.
(5) Steps (1) through (4) may be applied to Stage 2 of the activated
sludge process except that the inflow of wastewater in Step (3) will
be replaced by the overflow from the Stage 1 clarifier.
(6) At intervals, probably every four hours, check the dissolved oxygen
(DO) level in the aeration basins to ensure a minimum of 2 mg/1.
The activated sludge will take several weeks to develop in the aeration
basins. Sludge recycle rate should be gradually reduced as the mixed liquor
volatile suspended solids, X, increases. Sludge wasting should be started
when X exceeds 4,000 mg/1. Adjustments of return sludge rate, q, and sludge
wasting rate, w, will have to be made to maintain a steady state.
5.4.3 Operation and Control
The activated sludge process is designed to be run continuously and semi-
automatically. The equalization basin effluent wil_ be uniformly fed to
the aeration basins and mixed with the returned sludge. After contact in the
183
-------
aeration basin, the mixed liquor then flows by gravity to the clarifier.
The clarified effluent overflows to any downstream processes, and the
settled downflow will be pumped to a sludge holding tank. Part of the sludge
will be wasted from the holding tank to maintain QC and the remainder will be
pumped back to the aeration basins to react with the wastewater influent.
The process controls that the operator can regulate are sludge return
rate, q, and sludge wasting rate, w. The relationships between the controls
and expected responses of the activated sludge process are detailed in
Table 5-10. Based on an understanding of these relationships, a list of
suggested actions to cope with various conditions of process parameters is
given in Table 5-11. The same principles apply to the operation and
control of both stages of the activated sludge process.
5.4.4 Nutrients and Activated Sludge Seed
Phosphorus has been considered the only mineral nutrient supplement
required for the activated sludge. However, a number of studies have
advocated the addition of other mineral nutrients. In a study of the
applicability of coke plant water treatment technology to coal gasification,
20
Parsons and Nolde based the mineral requirements on the elemental composi-
tion of typical dry bacterial protoplasm, which is shown in Table 5-12. A
comparison between the approximate quantities of nutrients present in
Synthane coal gasification condensate and coke plant ammonia still feed,
and the bacterial composition from Table 5-12 is made in Table 13. The com-
parison shows that Synthane condensate is more deficient in phosphorous,
potasium, iron and magnesium than the ammonia still feed. However, this does
not represent a universal conclusion because the condensate composition
depends on the type of coal feed and the upstream treatment processes.
In a recent bench-scale pilot study of the biotreatment of a weak
ammonia liquor from a coke plant which had been pretreated by multi-stage
21
flash evaporation , Adams supplemented various mineral nutrients to the raw
wastes in the ratios shown in Table 5-14. The ratios are based on a litera-
ture review of the nutrients required for activated sludge and the relative
amounts of minerals that are added are related to the BOD of the wastewater.
In their experimental study they found that filamentous sludges were
produced in the absence of the nutrients shown in Table 5-14.
184
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TABLE 5-10 RELATIONSHIP BETWEEN PROCESS CONTROLS AND EXPECTED
RESPONSES OP ACTIVATED SLUDGE PROCESS
Case
Process Control
Parameters
Return
Sludge
Flow Rate, q
Sludge
Wasting
Rate, W
Expected
Process Responses
Sludge Settling
Characteristics
Solids
Mass
Wasted, P
x
Solids
Retention
Time, 9
1.
oo
Ul
Increase
Decrease
Constant
Constant
Constant
Increases
1) Sludge Volume Less
Increases
2) X Decreases
r
1) Sludge Volume More
Decreases
2) X Increases
r
No Effect Initially, More
then Follows Case 2
Increases
Decreases
Decreases
Constant
Decreases
No Effect Initially
then Follows Case 1
Less
Increases
-------
TABLE 5-11 SUGGESTED ACTIONS FOR VARIOUS CONDITIONS OF PROCESS PARAMETERS
Process Parameter
Dissolved Oxygen
(DO) in Aeration
Basins
Mixed Liquor
Volatile Suspended
Solids (MLVSS)
Depth to Sludge
Blanket
Condition
(1) Satisfactory
(2.O-4.0 mg/1)
(2) Too high
(>4.0 rag/1)
(3) Too low
(<2.0 mg/1)
(1) Satisfactory (3,000-
4,000 mg/1)
(2) Too high
(>4,000 mg/1)
(3) Too low
(< 3.0OO ng/1)
(1) Satisfactory
(7-10 ft)
(2) Too large
{> 10 ft)
(3) Too small
(< 7 ft)
Suggested Action
NONE
Reduce aerator speed from
high to low
(1) Increase aerator speed
from low to high
(2) Measure oxygen uptake rate
(OUR) on other sections of
the basin to determine if
primary effluent feed to all
sections should be modified
NONE
Increase sludge wasting slowly
Decrease sludge wasting slowly
NONE
Increase Return Sludge Rate or
decrease Sludge Hasting Rate
Increase Sludge Waiting Rate or
decrease Return Sludge Rate
186
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TABLE 5-12 REPRESENTATIVE ELEMENTAL COMPOSITION
OF DRY BACTERIAL PROTOPLASM
(From Reference 20)
c
H
O
N
P
S
Wt. %
50
5.8
27
12
2.5
0.7
Na
K
Ca
Mg
Fe
Wt.
0.
0.
0.
0.
0.
%
7
5
7
5
1
TABLE 5-13 HYPOTHETICAL COMPARISON OF TRACE NUTRIENT COMPOSITION
VS. INDICATED BACTERIAL REQUIREMENT FOR SYNTHANE CONDENSATE
AND COKE PLANT AMMONIA STILL FEED (From Reference 20)
Indicated
Requirement
lb/1,000 tons
Synthane Condensate
Indicated
Present
lb/1,000 tons
Coke Plant Ammonia
Still Feed
Indicated Present
lb/1,000 tons
Ca
Fe
K
Mg
Na
P
10
2.8
14
14
20
70
5.
0.
0.
1.
19
0.
1
28
78
6
12
5.
3.
16
3.
95
57
9
2
5
187
-------
TABLE 5-14 REQUIRED NUTRIENTS FOR ACTIVATED SLUDGE (From Reference 21)
gm/30 1 mg/1
Manganous Sulfate 0.831 27.7
Manganous Sulfate: H2O 0.930 31.0
Cupric Chloride 0.969 32.3
Sub. Copper Sulfate: 5H O 1.670 55.6
Zinc Sulfate 1-20 40.0
ZnS04:7H20 2.14 71.3
Molybdic Anhydride 1.93 64.3
Sub.Ammonium Molybdate
Selenium Oxide 6(10~ ) .2(10~ )
Magnesium Sulfate 45.15 1,505.0
Magnesium Sulfate:7H O 92.4 3,080.0
Sub. Magnesium Chloride
Cobalt Chloride 1.1 36.6
Cobalt Chloride:6H O 2.01 67.0
Calcium Chloride 51.2 1,706.6
Sodium Carbonate 14.46 482.0
Sub. Sodium Bicarbonate 28.92 964.0
Sodium Chloride 0.395 13.1
Sub. Sodium Sulfate
Potassium Chloride 25.9 863.3
Ferric Chloride 107.0 3,566.6
Ferric Chloride:6HO 178.0 5,933.3
Amount/Day (liters) = BOD in mg/1 x 10~ x Flow in I/Day
188
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The addition of nutrients in the ratios shown in Table 5-12 reduced the
SVI from 250 - 350 to 100 - 200 range.
Although the above data will serve as a general guideline, the exact
quantities of mineral nutrients required will have to be determined on a
case-by-case basis.
The mineral nutrient requirements of activated sludge should be first
studied by laboratory experimental procedures involving several culture
22
transfers in the determination of BOD rates . The results of these
experiments should then be tested on a continuous-flow basis with particular
emphasis on the settleability of the generated sludge as measured by SVI.
The activated sludge seed used to start up the plant should be obtained
from a source which is free from pathogenic organisms. This is because
the treated water may be used as cooling tower makeup and every precaution
should be taken to prevent or minimize the potential spread of any harmful
organism, in cooling tower drift or blowdown. Potential sources for the
activated sludge seed include commerical biological cultures such as
mutated Pseudomonas sp., or the activated sludge successfully treating a
similar industrial wastewater such as coke plant wastes, in which no
contamination by pathogens has occurred.
5.4.5 Determination of Improved Process Constants
Published process constants on related waters vary over a wide range.
This uncertainty makes it necessary for the constants to be determined in
laboratories and demonstration plants so that any future plants can be
designed better. Detail on laboratory determinations will be found in
Reference 45. Demonstration plants should be designed with flexibility,
and the determinations of process constants should be made.
A. Biokinetic Constants, k, K , Y, k
These four biokinetic constants are evaluated graphically as shown on
Figures 5-2 and 5-3, which are graphs of Equations (3) and (2). All values
plotted should be taken under steady state corresponding to specific
values of 6 as calculated by Equation (6). The general procedures are:
c
1) After startup select an appropriate value of X, say 3,000
mg/1, and adjust the sludge wasting rate so that a steady state can be
maintained.
189
-------
2) Measure SQ, S^, 0, w, X, X^, X£
ids
3) Calculate - —= —
wX P
!_ dX_ _ r _ x
X dt ~ VX XV
0 = VX
c —r, r~
4) Increase X by reducing w gradually.
5) Repeat steps 2), 3) and 4) until at least five sets of data are
obtained.
6) Plot — — vs. S as in Figure 5-2 and evaluate k and K .
X dt s
7) Plot vs. — -r- as in Figure 5-3 and evaluate Y and k .
X dt X dt a
It is also desirable to find the maximum X which can be maintained for
satisfactory operation, i.e., without exhausting dissolved oxygen in the mixed
liquor with aerators operating at full capacity, and without excessive
deterioration of effluent quality.
The F/M ratios, if desired, may be calculated from 9 by using Equation
C
(12). The substrate S may represent any of the following parameters: phenols,
BOD, COD or TOC. It is recommended that all of these parameters be measured
when any S is to be determined.
The effects of trace nutrients, if any, may be evaluated by comparing
values of biokinetic constants with nutrients added and those without.
B. Oxygen Requirements
The oxygen requirements may be determined by OUR measurements or total
COD balance as explained. Because of the high BOD strength of coal conversion
wastewaters, total COD balance would probably give the most accurate results.
The OUR measurements are fast and simple, and may serve to quickly identify
any potential trouble in the bioreactor. For instance, a low OUR indicates
slow bio-oxidation which may be due to the presence of toxic substances in the
feed.
190
-------
The oxygen requirements determined in accordance with the procedures
given are usually in pounds of 0 required per day. To present such data in a
more meaningful way, the total oxygen requirements may be divided into two
major components as in the following equation:
i i
Ib O /day = a (Ib BOD removed/day) + b (Ib MLVSS)
The coefficient, a , represents the oxidation of BOD in the wastewater. The
coefficient, b , represents the endogenous respiration rate. The amount of
BOD5 removed per day may be calculated from Equation (18) , and the amount of
MLVSS is equal to VX.
The total oxygen requirements will be a function of the BOD/COD ratio.
If the wastewater feed varies in BOD/COD, extra care must be taken to ensure
that total oxygen requirements are determined under steady state conditions
when the BOD/COD ratio is constant.
C. Solids Settling Experiments
The settling characteristics of activated sludge have been discussed and
the relationship between the settling velocity, V., of the solids at concentra-
tion, X. , has been described by Equation (29):
V. = a X.~n
i i
One objective of solids settling experiments is to evaluate coefficients a and
n from multiple batch settling tests as described below:
1) The mixed liquor from the sludge holding basin is diluted with
treated water to obtain various solids concentrations within the range of
concentrations expected to be maintained in the aeration basin.
2) Pour the mixed liquors into standard one-liter graduate cylinders,
each with a different solids concentration, X.. Each test cylinder should be
equipped with a variable-speed stirrer.
3) Mix the solids gently in each cylinder to obtain homogeneity, then
stop mixing and allow solids to settle under quiescent conditions.
4) A solids-liquor interface will form and its position vs. settling
time is recorded.
191
-------
5) Plot the position of interface vs. settling time as shown in
Figure 5-10.
6) The slope of the straight line portion of the settling curve in
Figure 5-10 is the settling velocity, V. , of the sludge at an initial
solids concentration of X. . This settling velocity is also referred to as
zone settling velocity (ZSV).
7) Plot V. vs. X. on a log-log graph paper as shown in Figure 5-4
and evaluate a and n from the straight line obtained.
8) At the completion of settling tests, siphon off the clear super-
natant and determine the solids concentration in the remaining sludge
blanket, which may be a good measure of X in the full scale clarifier.
192
-------
1000
i.
OJ
•o
3C
CD
- £ 500
IO
3
•a
-------
5.5 High Purity Oxygen Activated Sludge (HPOAS) Process
HOP AS represents a modification of AAS with major changes in the
following:
1. High purity oxygen is used in place of ordinary air to supply the
oxygen needed to oxidize the organics. Because of the higher partial
pressure of oxygen in the gas phase, the dissolved oxygen can be main-
tained at a higher concentration in HPOAS than in AAS system.
2. Staged covered reactors are used to facilitate the transfer of
oxygen from the gas to the liquid phase. The reactors in stages represent
a configuration between the completely mixed and the plug flow reactor.
If the treatment reaction follows a nonzero order kinetics, this config-
uration should require less volume than the completely mixed to achieve
equivalent treatment efficiency. However, any toxic substance in the feed
would be distributed in a smaller volume and may, therefore, have more
detrimental impact.
3. The mixing and oxygenation capacity is gradually reduced from the
initial to final stages. The high-purity oxygen and wastewater are usually
concurrently fed into the staged reactors, and a typical example is the
UNOX system shown in Figure 5-11. In this arrangement, the higher oxygen
demand in the initial stages will be satisfied with higher oxygen supply.
As the biotreatment intensity decreases along the sequence of stages, the
required mixing as well as oxygenation capacity are also reduced. The
reduced mixing intensity is highly conducive to the agglomeration of
biological floes and the subsequent solid-liquid separation.
Several commercial HPOAS systems are presently available, and the
UNOX system marketed by Union Carbide Corp. appears to be the most widely
waste
12,30
used for the biotreatment of industrial wastes . A recent overview of
the UNOX process can be found elsewhere
5.5.1 The Biokinetic Model. In general, it has been observed13'14
that HPOAS seems to be more efficient than AAS in removing BOD under
higher organic loadings, and also tends to produce less biomass for sludge
disposal. These advantages of HPOAS are not to be attributed to any
metabolic changes in the microorganisms indigenous to the activated sludge
process. Biochemical analysis of sludge samples from both AAS and HPOAS
systems showed no difference in all constituents in terms of their
percentages of volatile suspended solids.
194
-------
AERATION
TANK COVER
CONTROL
VALVE
OXYGEN
FEED GAS
WASTE
LIQUOR
FEED
RECYCLE
SLUDGE
Ul
EXHAUST
GAS
MIXED LIQUOR
EFFLUENT TO
CLARIFIER
Figure 5-11. Schematic diagram of UNOX system with surface aerator (3 stages shown).
-------
A widely accepted hypothesis accounts for the difference as different
rates of penetration of substrate and oxygen into the biological floes
The ideal situation would be that both substrate and oxygen penetrate to
an equivalent level into biological floes. However, at high organic
loadings, the oxygen penetration may not be commensurate with the substrate
penetration and the kinetics of biotreatment becomes limited by the oxygen
penetration. Since the dissolved oxygen of the mixed liquor can be main-
tained at a higher level in HPOAS, the driving force for oxygen penetration
is larger, and the oxygen-limiting situation may be reduced or eliminated.
Under a more aerobic environment, the production of biomass by the floes
will be less.
Since no biological difference exists between the microorganisms in
the AAS and those in the HPOAS system, the same principles of biokinetic
modeling may be applied. The equations describing the AAS process also
describe the HPOAS if the reactor configuration is the same. The effect
of high dissolved oxygen concentration in the mixed liquor of HPOAS will
be reflected in the values of coefficients: k, K , Y, k . The major
effects of using staged closed reactors will be discussed below.
5.5.2 Solid-Liquid Separation. The design of solid-liquid separation
for HPOAS is identical with that for AAS, i.e. the separator should be
designed to satisfy the requirements for both clarification and thickening.
However, since the HPOAS system usually maintains a higher level of MLVSS
in the bioreactor, the ability to deliver a high concentration of recycled
sludge from the solid separator to the bioreactor will be especially
important i.e. the design of the solid separator is more likely to be
controlled by the thickening than by the clarifying function. Those
equations and discussions developed for solid-liquid separation will be
particularly important for the design of a HPOAS system.
12 13
It is commonly accepted ' that the settling characteristics of
activated sludge improve significantly as the dissolved oxygen increases
from 2 to 5 mg/1 or more. This improvement will be reflected in the value
of the coefficients a and n used in the settling equation. However, the
general procedures for using such information in the design should be the
same as that described.
196
-------
5.5.3 Oxygen Transfer. The principles and discussions already given
are generally applicable to HPOAS except for certain additional complica-
tions primarily due to the reactor configuration used. The use of staged
closed reactors for HPOAS introduces the following conditions different
from those for AAS:
1. While the closed reactor is desirable in maintaining relatively
higher partial pressure of oxygen in the gas space inside each reactor, it
also tends to collect any gases which may be stripped out of the liquid
phase as a result of the mixing action. These gases include nitrogen
(N ), carbon dioxide (CO ) and any volatile organics originally present in
the wastewater. The N content in the wastewater is due to its solubility,
but the CO will be mostly from biochemical oxidation.
2. The compositions of both gas and liquid phases will change from
stage to stage in a system depicted in Figure 5-11. Since the oxygen is
fed co-currently with the wastewater, the partial pressure of oxygen will
be reduced" from stage to stage because of the oxygen demand of wastewater
and the diluting effect of the stripped gases.
Although the overall objective of oxygen transfer remains the match-
ing of oxygen supply and demand, this task becomes complicated as the
objective should be achieved in each stage under dynamic conditions.
Obviously the rational approach would be to develop a mathematical model
describing the mass transfer of key constituents in both gas and liquid
phases so that the condition of each stage can be readily computed. This
modeling approach will be further developed.
The collection of any volatile organics in the gas phase may facili-
tate the treatment of gaseous emissions from an HPOAS plant should it
become necessary. This should be considered as an advantage of the HPOAS
system.
5.5.4 Cooling of Mixed Liquor. One of the complications introduced
by enclosing the reactors is the rise of water temperature due to the
oxidation of biodegradable organics. Assuming 100% biodegradability, the
oxidation of phenol can lead to about 1°F temperature rise per 165 mg/1
BOD removed.
197
-------
Biological agents are known to be temperature sensitive. Mesophilic
bacteria, which constitute the bulk of activated sludge, do not operate
effectively of the temperature rises beyond about 104°F (40°C) . In the
aerobic biological treatment of coke plant wastes, it has been recom-
mended16 that the water temperature be 95-100°F throughout the year. When
a high level of BOD removal is to be achieved by HPOAS, the rise in water
temperature may become excessive and a means of cooling becomes necessary.
One possibility is to control the temperature by internally recycling the
wastewater stream (mixed liquor) from the last stage, back to the first
stage. This has the possible additional advantage of diluting potentially
toxic materials in the concentrated wastewater.
5.5.5 A Model for HPOAS. As mentioned, a rational approach to
matching the oxygen supply with demand may be to develop a mathematical
model to account for the material balance of key constituents in both gas
and liquid phases of each stage of the reactors. Such a model has been
developed by Mueller et al for a municipal wastewater plant and successfully
used to predict changes in operation . In this section basic equations
will be developed and the general approach for coal conversion wastewaters
using the kinetic equations developed for AAS will be described.
The key constituents in the liquid phase are dissolved oxygen (O ) ,
carbon dioxide (CO2) , nitgrogen (N2) , BOD and pH. In the gas phase the
parameters of concern are the partial pressures of O_, CO , N . In addi-
^ £* £
tion, since a net gas transfer occurs from the gas to liquid phase, the
gas flow rate is also of concern. Thus there are, in total, nine key
parameters which are interrelated and whose magnitudes are to be deter-
17 18
mined for every stage of reactor '
The principle of mass conservation will be applied to a reactor
configuration shown in Figure 5-12. Wastewater and oxygen are co-
currently applied to a series of n stages of bioreactors, and the outflow
of liquid and gas from the (n-l)th stage is the inflow to the n'th stage.
The liquid and gas phases of each stage are assumed to be completely
mixed. The concentration of the i'th liquid constituent in the n'th stage
is denoted by (^ n» and the partial pressure of the i'th gas constituent
in the n'th stage by P.
198
-------
GAS
PHASE
vo
VD
AERATION TANK STAGES
SECONDARY
CLARIFIER
Figure 5-12: Schematic diagram of staged bioreactors for HPOAS
(From Reference 17)
-------
The general mass conservation equation is:
Change = Inflow - Outflow + Sources - Sinks (42)
The inflow and outflow may be represented by identical formulations for any
constituent in the liquid or gas phase while sources and sinks will vary from
constituent to constituent depending on the interactions among constituents.
In the liquid phase the concentration of a dissolved constituent is governed
by: dc-
(C. - C, _) + St _ (43)
i ,n _ Q_ . _ _1
"«~ V
n
where
V = volume of the n'th stage liquid phase
n
Q = liquid flow rate, including the influent
plus recycle flows = Q + q
S. = sum of sources and sinks of the i'th constituents
' in the n'th stage liquid phase
t = time
In the gas phase the mass conservation equations are of the form:
M± V9 dP. M.
RT~ dt~~ = FT (Gn-l Pi,n-l ~ Gn Pi,n) + Vn Si,n (44)
where
Mi = molecular weight of the i'th gas constituent
R = universal gas constant
T = absolute temperature of the gas phase
g
V = volume of the n'th gas stage
Gn_1 = gas flow rate from the (n-l)th- to n'th stage
g
S. = sum of the sources and sinks of the i'th gas
constituent in the n'th stage.
In writing Equations (43) and (44) it has been assumed that each volume, V1
g n
and V , is constant in time, and that the ideal gas law applies for each gas
constituent.
200
-------
The interactions among these constituents, which constitute various
sources and sinks, are described in the following paragraphs:
1. Gas transfer between gas and liquid phases.
As oxygen is introduced into a bioreactor, three transfer mechanisms
occur: (1) O^ is transferred from the gas to liquid phase; (2) N originally
present in the liquid is initially transferred to the gas phase; and (3) CO
produced by the biochemical reaction is transferred from the liquid to gas
phase.
The rate of gas transfer across the gas-liquid interface is proportional
to both the driving force (the difference between the dissolved gas saturation
and the actual concentration) and the overall gas transfer coefficient, K a.
L
For a dissolved gas in water, its transfer rate can be expressed as:
aT-V
where
C. = actual concentration of the dissolved gas, i
saturation concentration of the dissolved gas, i,
which can be related to its j
the gas phase by Henry's Law
' which can be related to its partial pressure in
C . = H.P. (46)
v D'
where
H. = Henry's constant for gas i.
2. Consumption due to biochemical oxidation.
Biochemical oxidation leads to a co
oxygen requirements may be expressed as:
Biochemical oxidation leads to a consumption of O and BOD. The total
Ib O /day = a1(Ib BOD removed/day) + b1(Ib MLVSS) (47)
201
-------
The coefficients a' and b1 will be used to formulate the sinks for C>2 in the
liquid phase. The consumption of BOD has been described in the form:
ds _ KXS (48)
dt K + S
s
The terms in Equation (48) have been defined in the section on air activated
sludge.
3. Production due to biochemical oxidation.
While 0 is consumed during biochemical oxidation, CO is produced. The
moles of CO produced per mole O consumed is called the respiratory quotient
(RQ) and its value may vary with loading conditions and can be experimentally
determined. The production of CO may be expressed in the same fashion as
Equation (47):
Ib CO2/day = a"(lb BOD removed/day) + b"(lb MLVSS) (49)
where
(50>
The various sources and sinks for key constituents in the liquid and gas
phases are summarized in Table 5-15.
As shown in Table 5-15, the only interaction with N in the liquid phase
is the interphase gas transfer; no other reaction is assumed to take place.
Also no BOD transfer from the liquid to gas phase has been assumed. Because
no reaction occurs in the gas phase, the only source of each species is that
through interphase gas transfer.
202
-------
TABLE 5-15 LIQUID AND GAS PHASE SOURCES AND SINKS
Species
Designation
Sources and Sinks
Interphase
Gas Transfer
BOD
Oxidation
Bacterial
Respiration
(a) Liquid Phase, S.
°2
co2
N2
BOD
Cl,n
C2,n
C3,n
°4,n
+K_a(H P - C )
L 1 l,n l,n
+V(H2P2,n - C2,n>
+V(H3P3,n - C, )
3,n
kXC,
, 4,n
" K + C,,
s 4,n
la" 4/n
K + C.
s 4,n
kXC.
4,n
K + C.
s 4,n
-b'X
+b"X
(b) Gas Phase, S?
CO,
2,n
3,n
-K a(H P - C )
L 1 l,n l,n
-KLa(H2P2,n * C2,n>
-K a(H P - C )
L 3 3,n 3 ,n
203
-------
As a typical example, the mass balance equation for dissolved oxygen in
the n'th stage reactor is:
dc,
(C - C ) + K_a(H P - C ) (51)
i l,n-l l,n L 11,n ±,n
n
— kXC_ I
. •
>'X
Solving for the dissolved oxygen concentration in the n'th stage at steady
state yields :
kXC
aH.P. - a' — — - — + b'X
V
n
~|
I
+ K a (52)
V
n
Other equations for the concentrations of other species can be obtained
similarly.
In addition to the seven equations indicated in Table 5-15, two more
equations are needed to complete the equation set. One is for the evaluation
of pH. Assuming that the only significant source of alkalinity is carbonaceous
alkalinity, and that the pH is in the neutral range, then all the alkalinity
occurs substantially as bicarbonate alkalinity. For the first dissociation
reaction of carbonic acid, the following equations can be obtained:
_ (H+) (HCO~) ->- 4-
Kl ~ 3
So
(H CO ) (CO ) C_
£ J «£ 2 ,n
pH = pK + log (Alk) - log (C, ) (53)
n i 2. ,n
The final equation is the continuity condition in the gas phase. The gas
flow rates, G , may vary from stage to stage, so a continuity condition is
204
-------
required to calculate the gas flows. In the present practice of staged
covered reactors for HPOAS, negligible head loss usually occurs in the gas
flow from stage to stage; the total pressure in the gas phase of each
stage is, therefore, a constant, P and equal to the sum of the partial
pressure of the individual gases present, i.e.
P, + P~ + P =P = P
l,n 2,n 3,n Tn T (54)
For the UNOX system, oxygen gas enters the first stage at a slight pressure
(1-4 inches water column gauge) and flows from stage to stage with enough
velocity to prevent gas backmixing
Equations (53), (54) and those partially described in Table 1-1 make
up a total of nine equations, five for the liquid phase and four for the
gas phase, to solve for the nine key parameters discussed previously.
Given all the necessary coefficients in these equations, the steady state
condition of the key parameters can be determined for each stage. In
general these calculations may be done for the following purposes:
1. To determine the number of stages required for a certain treatment
efficiency and the capacity of each stage;
2. To match the supply with the demand of oxygen in each stage and
thus determine the gas flow rate and the capacity of the aerators;
3. To select the optimal combination of various design components,
e.g. the economic tradeoff between increasing gas flow and increasing
aerator power to obtain a desirable mixed liquor dissolved oxygen concen-
tration.
The use of this model for design requires considerably more information
than has been published. Some calculations have been made for our examples
and they are given later. To make the actual designs for our examples we
have relied on the past experience of Union Carbide Corporation treating
similar wastewaters using their UNOX system. We are much indebted to
Union Carbide.
5.5.6 Summary of Design Procedures
The design of a High Purity Oxygen Activated Sludge system using past
experience ii ludes sizing the oxygenation basins and clarifiers and
determining the oxygen and electric power requirements. The procedures
are outlined in the following paragraphs.
205
-------
Oxygenation Basins
1. Determine BOD (mg/1) applied to bioreactors.
2. Determine F/M (applied) 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) _ ,„.
RTQ (hrs) ~ 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, d, in the oxygenation basin, A = V/d.
6. Determine the total number of stages to provide the area, each stage
being a square in cross section.
7. Design the overall configuration of the oxygenation basins, including
number of trains and number of stages in each train.
It should be noted that the use of F/M ratio can actually be related to
the biokinetic model discussed, as shown in the following equation:
- = Y (F/M) (BOD removal %) - k
0 Q
c
Such correlation of different design techniques has been discussed by
19
Sherrard
Recirculation and Cooling of Mixed Liquor
1. Calculate the BOD removed in mg/1
2. Calculate the temperature rise of mixed liquor on the basis of 1°F
temperature rise per 165 mg/1 BOD removed.
3. Assume the temperature of the influent wastewater to be 90°F, and that
the temperature of mixed liquor is not to exceed 104°F. Calculate the heat
removal rate, in Btu/hour for example, to be achieved by a heat exchanger or
equivalent.
4. Determine the required area of heat exchanger on the basis of heat
removal rate, wet bulb temperature and heat exchange coefficient.
206
-------
Clarifiers
1. Assume the overflow rates for Steps 1 and 2 to be 700 and 400 gpd/sq
ft respectively.
2. Calculate the total surface area by dividing the hydraulic flow rate
to the clarifiers by the overflow rate.
3. Determine the diameter of the clarifiers by considering the number of
clarifiers to be preferably more than one.
4. Select the depth of clarifiers, usually in the order of 15 ft.
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 Ib 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 Ibs of oxygen 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).
5.5.7 Example 1. Moderate Strength Wastewater.
A feed of 0.3 x 10 gals/day at 2,500 mg/1 BOD means an application of
6,255 Ib BOD/day. This and other design information is tabulated in Table 5-
16. The chosen F/M ratio is 0.8. The chosen MLVSS is 4,000 mg/1. The
hydraulic residence time from Equation (55) is:
(2'500)(24) = 18. 75 hours
(4,000)(0.8)
The aeration basin volume is 0.23 x 10 gals = 31,200 ft . This is shown in
Figures 5-13 and 5-14. A single train, two-step basin has been chosen.
The temperature rise is
2,500 mg BOD/1 = 150F
165
and cooling is not required.
207
-------
TABLE 5-16. DESIGN OF THE HPOAS SYSTEM,
EXAMPLE Ia
Design Basis
Flow, 106 gal/day 0.3 (1,671 ft3/hr)
Phenols, mg/1 300-500
BOD , Ibs/day 6,255
BOD5, mg/1 2,500
COD, mg/1 3,000
COD/BOD5 !-2
Influent suspended solids, mg/1 Negligible
Influent wastewater temperature, °F 90°F
pH Adjusted as required
Nutrients Phosphorous to be added
Required effluent BOD, mg/1 45
System Design
Retention time, hrs (based on feed flow) 19
MLSS, mg/1 5,000
MLVSS, mg/1 4,000
Sludge recycle rate, %Q 30
F/M ratio, Ibs BOD5/(lb MLVSS)(day) 0.8
Average DO level, mg/1 5-0
Oxygen supplied, tons/day 5
Average oxygen utilization efficiency, % 75
Secondary clarifier overflow rate, gal/(day)(ft2) 400
Recycle suspended solids concentration, wt % 2.0
Effluent soluble BOD5/ mg/1 45
Utility Requirements Oxygenation Motors
Operating Energy:
Brake HP 58
Connected Load:
Nameplate HP 80
Nameplate HP (stagewise) 40/40
information supplied by Union Carbide on the basis of assumptions
provided by WPA.
Used as basis for determining oxygen requirement.
208
-------
NUTRIENTS
to
a
US_ED_FOR_ _START-UP
CLARIFIER
SLUDGE
DISPOSAL
97,370
SLUDGE
HOLDING
BASIN
ALL FLOWS IN GAL/DAY
Figure 5-13. Schematic flow diagram of high purity oxygen activated sludge (HPOAS) system
Example 1.
-------
10'
39.5'
t
>_
•
40 hp
— rh—
-LjJ-
40,hp
rb
T
*-
to
I-1
o
SLUDGE HOLDING
BASIN
39.51
Figure 5-14. Configuration of HPOAS system, Example 1.
-------
2
The clarifier is sized for 400 gals/(day)(ft ).
The oxygen requirement is assumed to be 1.22 Ib O /lb BOD removed, so the
oxygen required is:
(1.22)(6,255 lb BOD/day) = 7,631 Ib/day
With an assumed utilization efficiency of 75%, the oxygen supplied is:
10,175 Ib/day = 5 tons/day
The oxygenation motor brake horsepower is based on 7.3 lb O /hp-hr;
200 kw-hrs/ton oxygen supplied.
The sludge growth rate, based on UNOX experience, is 0.22 lb TSS/lb BOD
removed, or 1,376 lb TSS/day. If the clarifier yields an underflow of 2%
solids, the wasting rate is 67,430 lb water/day or 7,370 gallons/day. The 24-
hour sludge holding basin has a capacity of 7,370 gallons = 985 ft . At
2
3 lb dry weight/(ft )(hr), the active filter area for operation 7 hrs/day is
1,376 lb sludge/(21 lb/ft2) = 66 ft2.
5.5.8 Example 2. High Strength Wastewater.
For this example the flow is 0.17 x 10 gal/day at 18,000 mg/1 BOD. Two
steps have been used, each giving 95% removal of BOD. In each step a MLVSS
concentration of 4,000 mg/1 is used. In Step 1, with the higher BOD loading,
F/M has been taken as 0.5; in Step 2 F/M has been taken as 0.3. The design
basis is listed in Table 5-17.
The Aeration Basins
From Equation (55) the hydraulic residence time in Step 1 is:
(18,000)(24)
(4,000)(0.5) =
and in Step 2 is 18 hours.
6 36
The aeration basin volumes are 1.5 x 10 gal = 202,000 ft and 0.13 x 10
gals = 17,000 ft . These volumes are shown on Figuies 5-15 through 5-17.
Single train basins have been chosen; four stages in Step 1 and two stages in
Step 2.
211
-------
TABLE 5-17. DESIGN OF THE HPOAS SYSTEM
EXAMPLE HO. 2*
Design Basis
Flow, 106 gal/day 0.17 (947 ft3/hr)
Phenols, mg/1 3.000 - 5,000
BODj, Ibs/day 25,520
BODS, ng/1 18,000
COO, Kg/I 25,000 - 30,000
COD/BOD 5 1.5
Influent suspended solids, ag/1 Negligible
Influent waste water temperature, *T 90T
pH Adjusted as required
nutrients Phosphorous to be added
Required effluent BOD, mg/l 45
System Design Step 1 Step 2
Plow, CUO6 gal/day) 0.17 0.17
Influent temperature *F 90 101
Retention tone, hrs (based on feed flow) 216 IB
HLSS, ng/1 5,000 5,000
HLVSS, ng/1 4,000 4,000
Sludge recycle rate, %Q 14 30
Recycle for cooling, %Q 640
F/H ratio. Iba BODs/(lb HLVSS) (day) 0.5 0.3
JUrerage DO level, ng/1 5.0 5.0
Oxygen supplied, tons/day 31 2
Average oxygen utilization efficiency, % 70 65
Secondary clarif ler overflow rate, gal/ (day) (ft2) 700 400
Recycle suspended solids concentration, wt % 2.0 2.0
effluent soluble BOO^. ag/1 900 45
Utility Requirements Oxygenation Motors
SteP l steP 2
Operating Energy:
Brake HP 311 20 331
Connected Load:
Naneplate HP 385 30 415
Nemeplate HP (stagewise) 12S/100/10O/60 15/15
information supplied by Union Carbide on the basis of assumptions
provided by WPA.
Used as basis for determining oxygen requirement.
212
-------
5. 3x 10° Btu/hr
NUTRIENTS
STEP 2
CLARIFIER
Flows in 10 gal/day
SLUDGE HOLDING
BASIN
SLUDGE
DISPOSAL
Figure 5-15. Schematic flow diagram of high purity oxygen activated sludge (HPOAS) system,
Example 2.
-------
15'
-2-
INFLUENT
58'
10
58'
125 hp
58'
100 hp
58'
100 hp
58'
60 hp
58'
CLARIFIER
SLUDGE
HOLDING BASIN
Figure ;5-'i6- Configuration of Step 1 of HPOAS system, Example 2.
-------
21
to
M
(J\
EFFLUENT 32.5^
FROM STEP 1 *
k
^
i
J£. D
15 hp
3L
15
-E
.D
?
[CLARI
I 24 f
>
^
8'
TO SLUDGE HOLDING
BASIN
32.5'
Figure 5-17 . Configuration of Step 2 of HPOAS system, , Example 2.
-------
Cooling Load
The expected temperature rise in step 1 is:
(18,000 - 900) rog/1 BOD _
165
so cooling is required. The permitted temperature rise is from a feed of 87°F
to a maximum of 101°F; that is 14°F. The heat load on the cooler is:
(0.17 x 106 gal/day) x (0.347 (lb/hr)/(gal/day)) x (104 - 14°F)
= 5.3 x 106 Btu/hr
The recycle rate for cooling is:
5.3 x 106/14 = 379,000 Ib /hr = 758 gal/min
which is 640% of the plant feed rate.
Cooler Design
The cooler is designed for
- 400,000 lb/hr
- range 101° to 87°F
- high suspended solids
- design ambient wet bulb temperature, 77°F
It is not likely that circulating cooling water from a cooling tower will
be cold enough for a countercurrent heat exchanger. In any event, it is
probably convenient to be independent of cooling water because the cooler
will be near the waste treatment facility. A wet surface cooler will be
used. Among the manufacturers of suitable equipment axe Baltimore Air
Coil and Niagara Blower Company. Figure 5-18 is a typical configuration.
Consider first a Baltimore Air Coil industrial fluid cooler. The speci-
fied equipment is Model 200-3 or Model 300-1 as may be determined from BAG
Bulletin S 404/1-0. The circulated flow is passed through cooling tubes,the
outside of which are kept flooded with water. A forced current of air is
also passed over the tubes. The water used for flooding is picked up from
a pan below the cooling tubes and resprayed over the tubes. A small
amount of makeup is added as the flooding water evaporates. At design,
flooding water will approach the wet bulb temperature of 77°F. The
design temperature difference is the log mean of (87-77) and (101-77) and
equals 16.0°F; the load is 400,000 x 14 = 5.6 x 10 Btu/hr; the coefficient
in this type of cooler is about 130-140 Btu/(ft2) (hr) (°F) so the area
required is 2,500 to 2,700 ft . The tubes in this equipment do not
216
-------
AIR DISCHARGE
A A
FLUID OUT
AIR FLOW CONTROL
WATER DISTRIBUTION WEIRS
/r/l\/lvl\/!\/!\./ivlvivivi\/lviv!v!vi'
WATER CASCADE
AIR
FLUID IN
WATER PUMP
Figure 5-18 . Cooler schematic.
217
-------
have fins. Dry tubes with forced air will have a heat transfer coefficient of
not more than 15 Btu/(hr)(ft2)(°F) and even on the coldest winter days the
flooding water is not turned off. Control of temperature is by dampers in the
air flow.
Makeup to the flooding water pan is from the final clarifier effluent of
the biological treatment. Lack of fins makes the tubes not too susceptible to
becoming coated with dirt and scale on the outsides; the outside of the tubes
is quite easy to clean. The blowdown from the flooding water loop probably
need not exceed 50% of the evaporation rate which is 5,600 Ib/hr (= 11.2 gpm)
in the summer and less in winter. The blowdown can be dumped into the bio-
treatment clarifier.
Niagara Blower makes similar equipment, but the cooling tubes are finned
2
on the outside. The dry tube coefficient is about 100 Btu/(hr)(ft )(°F) and
the flooding water will be turned off for about nine months of the year.
Temperature control will be mostly by control of air flow, but the blower
energy will be higher when flooding water is turned off.
Either piece of equipment will suffice.
The Clarifiers
The first clarifier is sized for 668 gal/(day)(ft ) and the second for
376 gal/(day)(ft2).
Oxygen Requirement
The oxygen required depends on the values of F/M and the ratio COD/BOD.
For Step 1, 1.8 Ib O /lb BOD has been assumed and for Step 2, a value of 2.1
Ib 0 /lb BOD. These are quite high values. The oxygen required is:
Step 1. (0.059 x 10 lb feed/hr) x (17,100 x 10~6 lb BOD/lb feed
x (1.8 lb 02/lb BOD)/0.7 efficiency
= 2,590 Ib/hr = 31 tons/day
Step 2. (0.059) (855) (2.D/0.65 = 163 Ib/hr = 2 tons/day
The average oxygenation motor efficiency is 8.2 lb O supplied/hp-hr.
Sludge
The total sludge production is about
(24,244 lb BOD removed/day) x (0.22 lb sludge/lb BOD) = 5,300 Ib/day
218
-------
The total waste sludge water rate at 2% solids is 0.26 x 10 Ib/day =
0.031 x 10 gal/day. The waste sludge rate is 0.029 x 10 gal/day in Step
1 and 0.002 x 10 gal/day in Step 2. The sludge holding basin for 24
3 2
hours is 31,000 gallons = 4,100 ft . The active filter area is 252 ft
for operation 7 hrs/day.
5.5.9 Use of the Model
Some calculations were made on the model described in Section 5.5.5 to
find if the two examples could be described. There is not enough information
to use the model properly, but it appears to promise a method of correlating
an operating plant. At this time the model requires too much data to be
useful for design.
The constants used were as follows:
For both examples we used the biokinetic constants used for air activated
sludge:
k = 1 g soluble BOD/(g.MLVSS)(hr)
k, = 0.033 hr'1
d
Y = 0.37 g.MLVSS/g.BOD removed
K = 266 mg.BOD/1
The Henry's law constants were:
0 ; H = (1.779 x 1Q6)/(7,241 t°'545)
ft D QRR
C02; H = (2.446 x 10 )/(64.58 t )
N2; H = (1.557 x 106)/(27,300 t°'363)
where t is in C, c is in mg/1, and P is in atm.
The respiratory quotient from Equation (50) was taken as 1.
The oxygen requirements were taken from the examples:
Example 1: a1 = 1.22, b1 = 0.0042
Example 2: a1 = 1.80, b1 = 0.0042
The recycle and sludge conditions also were taken from the examples.
We found that in Example 1 a possible simulation was obtained with K a
(the gas transfer rate in Equation (45)) equal to 14 in Step 1 and 6 in Step
2. The oxygen transfer rate depends on the gas p^'ase concentration and on
the aerator horsepower. The model gave negative oxygen concentrations in the
liquid in Step 1 because the equations permit it. Removals of BOD were higher
than expected. Modification of the model to show possible oxygen starvation
may help. However, we used the plant design to find gas transfer rates for a
219
-------
given set of biokinetic constants which may not be correct. If the biokinetic
constants are wrong, the model is probably so complex that it cannot be
used to find them. The biokinetic constants must be determined independently.
In Example 2 we modeled the first stage only using:
Step 1: lea = 44
Step 2: Ka = 31
Step 3: 1C a = 25
Step 4: K a = 12
The model gave very high BOD removals and sludge concentrations not over
500 mg/1. Other values were simulated. It seems that the kinetics were
too fast to really simulate the design of Example 2.
5.6 Trickling Filters
A trickling filter is a method of obtaining a high concentration of
bacteria by allowing them to grow on the solid surface of grid-like packing
while wastewater trickles in a thin film over the surface. An important
cost in activated sludge biological oxidation is the cost of supplying oxygen.
In a trickling filter, water and not air, is pumped and the energy required
is much lower than in an activated sludge process. The use of dual biological
processes (using a combination of trickling filter and activated sludge) for
industrial wastewater treatment is not new . Success in the treatment of
wastewaters from organic chemical manufacturing, petrochemical refining and
meat processing industrials has been reported ' . In most of the reported
cases the water contaminants of primary concern have been phenols and BOD,and
they were successfully removed by combining the desirable attributes of a
trickling filter and an activated sludge process into the most economical
treatment system
Trickling filters and experience with them are classified according
to the hydraulic loading and the organic loading . Of interest in the
treatment of coal conversion wastewaters are:
Hydraulic Loading '• Organic Loading
gal/(min)(ft2) Ib BOD/(day)(103ft3)
Super rate filters with
synthetic media 0.25 to 1.5* up to 300
Roughing filters 1.0 to 3.0 100 plus
*not
*not including recirculation
220
-------
In general trickling filters will not give as clean an effluent as an
activated sludge process. With high organic loadings, the fraction of the
BOD removed can be quite low; as low as 40%. With wastewaters from coal
conversion, the proper place for a trickling filter is as a roughing filter
ahead of the activated sludge plant. However, the use of a trickling filter
on wastewaters as concentrated as those from coal conversion has not been
reported and must be regarded as experimental. Although design procedures are
given below, they are based on correlations of trickling filter performance
on quite different wastes. The designs will serve to produce an experimental
trickling filter, not to build equipment that is certain to perform to
specification.
5.6.1 Design of Trickling Filters
Several different correlations have been used for trickling filters . The
formula most applicable for our purposes is the modified Germain formula :
L / \
e / K0© \ ,^r^
-= exp (-— (56)
\ J /
where
L = BOD of the effluent, mg/1
L = BOD of the influent, mg/1
3.
K = treatability factor
D = fill depth, ft
(T-20)
6 = temperature factor, = 1.035 , where T = water temperature in °C
2
J = water flux through the filter, gal/(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, that is, its specific surface area and geometry,
etc., which can affect the contact time between water and biological mass
on the surface of the medium.
Equation (56) is a correlating formula, not a design formula. The term
D/J has the dimensions of time. Equation (56) relates the fractional removal
of BOD to the contact time on the filter. There is no expression for the
effect of feed concentration on the fractional removal in a given filter.
221
-------
The treatability factor, K, lumps together the concentration and nature of
the waste. K must be determined experimentally. Experience is that K
generally lies in the range 0.01 to 0.045 for wastewaters like those from
coal conversion. The factor, n, is generally close to 0.5.
5.6.2 Example 1. Moderate Strength Wastewater
O
Example 1 is a wastewater from a demonstration plant of 0.3 x 10
gal/day 2,500 mg/1 BOD. In Table 5-18 some calculations are shown for
different values of K and for n = 0. 5. The temperature is taken to be
30°C; 6 = 1.41. The range of possible values of K is wide enough to make
it impossible to suggest a design. Experiment is required to evaluate K.
If even half of the influent BOD is removed, the aerators in a follow-
ing air activated sludge plant can be reduced by 72 hp, saving 0.47 x 10
kw-hrs/yr, or about $14,000/yr. Assuming that it is economical to invest
five times this saving, that is $70,000 in a trickle filter, it is clear
that a trickle filter should be investigated. The cost is about $10/ft .
5.6.3 Example 2. Trickle Filter on Recycle with a High Strength Wastewater
Our second example is high strength wastewater having 18,000 mg BOD/1.
We have not considered a trickling filter for this water since the concen-
tration is so high as to be, most probably, toxic. However, in designing
a high purity oxygen activated sludge plant for this water, it was found
to be necessary to recycle 640% of the feed flow for cooling. The recycle
contains about 5,000 mg/1 MLSS but this will most probably not render it un-
fit to feed to a trickling filter. Rather than using the flow circuit shown
in Figure 5-15, the circuit shown in Figure 5-19 can be used. The recycle
is mixed with the feed and applied to a trickling filter. The mixed feed to
the trickling filter is 1.26 x 10 gals/day at 3,200 mg BOD/1. For n=0.5 and
6=1.41 some calculations are shown in Table 5-19. The filter should probably
be 30,000 ft to hold the organic loading down to the level of experience.
This trickling filter must also supply cooling. Packing used in a
trickling filter is also usable in a cooling tower and a forced draft
cooling tower was designed for this service . A generous volume of packing
was 14' x 14' x 5' high = 980 ft . A trickling filter thirty times this size
is most likely to provide adequate cooling without a forced draft. It will
be necessary to support the fill above ground level to provide space for air
to enter at the bottom. It will also be necessary to provide louvers so
that air entry can be reduced in cold weather to avoid excessive cooling.
222
-------
TABLE 5-18 TRICKLING FILTER CALCULATIONS FOR EXAMPLE 1
0.3 x 106 gal/day at 2,500 mg BOD/1; 6,250 Ib BOD/day
K)
N)
Filter
Diam.
(ft)
11.2
13.8
16.0
13.8
16.0
19.5
16.9
19.5
23.9
21.9
25.2
30.9
Area
(ft2)
100
150
200
150
200
300
225
300
450
375
500
750
Depth
(ft)
30
20
15
40
30
20
40
30
20
40
30
20
Vol.
(ft3)
3,000
3,000
3,000
6,000
6,000
6,000
9,000
9,000
9,000
15,000
15,000
15,000
Organic Loading
Ib BOD/ (day) (103ft3)
2,083
2,083
2,083
1,040
1,040
1,040
694
694
694
417
417
417
Hydraulic
Loading
(gpm/ft2)
2.1
1.4
1.0
1.4
1.0
0.69
0.93
0.69
0.46
0.56
0.42
0.27
Percent Removal
1 -
K=0.045
73%
66
60
88
84
78
93
90
85
97
95
91
- L /L
e a
K=0.02
44%
38
25
61
57
49
89
64
56
78
73
66
K=0.01
25%
21
13
38
34
29
44
40
34
53
48
42
-------
NUTRIENT
NJ
it*
FROM EQUALIZATION
COOLING TOWER/TRICKLE FILTEF
STEP 1
CLARIFIER
STEP 2
STEP 2 HPOAS CLARIFIER
SLUDGE HOLDING CLARIFIER
TO SLUDGE DISPOSAL
Figure 5-19- Activated trickling filter-high purity oxygen activated
sludge system (ATF-HPOAS) for Hygas plant.
-------
TABLE 5-19 TRICKLING FILTER CALCULATIONS OR EXAMPLE 2
1.26 x 10 gal/day at 3,200 mg BOD/1; 33,600 Ib BOD/day
Filter Organic Loading
Diam. Area Depth Vol.
(ft) (ft2) (ft) (ft3) Ib BOD/(day) (103ft3)
30.9 750 40 30,000 1,120
35.7 1,000 30 30,000 1,120
43.7 1,500 40 60,000 560
50.5 2,000 30 60,000 560
10
to
in
Hydraulic Percent Removal
Loading 1 - L /L
e a
(gpm/ft2) K=0.045 K=0.02 K=0.01
1.2 90% 64% 40%
0.88 87 59 36
0.58 96 77 52
0.44 94 72 47
-------
5.7 Anaerobic Biological Destruction
In the anaerobic biological treatment processes, organic constituents
of the wastewater are converted to methane and carbon dioxide. General
explanations are given in References 32 and 39. The major 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 BODL/ while the corresponding
figure for the aerobic process may be as high as 0.4. Sludge production var-
ies with the nature of organic constituents and the mean cell residence
time, 0c.
(2) The problem of biological sludge disposal and the requirement of
a nutrient, phosphorus in this case, will be significantly reduced.
(3) Elimination of aeration results in significant savings in energy
and operating costs.
(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.
47
Studies have shown that petrochemical wastewaters may be successfully
treated by anaerobic lagoons in a large-scale pilot plant. Anaerobic
filters and contact digesters (i.e. anaerobic activated sludge) provided
satisfactory removals when treating dilute wastes in bench-scale studies,
but neither unit provided satisfactory removals when tested on a semi-
pilot scale with actual wastes. The difference in performance between the
lagoons and the other two systems was due to the ability of the lagoons to
metabolize volatile acid intermediates. The characteristics of these
three anaerobic systems are shown in Table 5-20.
Only limited data are available on the inhibitory effect of certain chemi-
48
cals found in petrochemical industrial wastewaters. A study showed that for
an active methane-producing, unacclimated, domestic digester sludge the 50 per-
cent inhibitory concentration of phenol (based on gas production) was found to
be larger than 1,000 mg/1 under substrate-limiting conditions, 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 conditions with a phenol concentration
226
-------
TABLE 5-20. ANAEROBIC SYSTEMS STUDIED
(From Reference 47)
Submerged Filter
Contact Digester
Open Lagoon
Description
Flow pattern
Biosolids level
Metabolic pathways
Retention time
Gas collection
Temperature control
Rock or gravel
packed column
Plug flow
High biomass
through attached
growths
Fermentation and
anaerobic
respiration
Completely mixed
vessel
Backmixed
High biomass
through settling
and return
Fermentation and
anaerobic
respiration
Basin with con-
siderable
stratification
Some wind and
wave mixing ,
thermal turnovers
Low suspended
solids , bottom
sludge layer
Fermentation ,
anaerobic
respiration ,
sulfur oxidation,
photosynthesis ,
some aerobic
respiration
1 to 3 days 1 to 10 days
Normally collected Collected
Not normally
practiced
Usually practiced
10 to 100 days
Gas is released,
although a plastic
covering with peri-
pheral collection
tiles is possible
Unfeasible unless
covered and
insulated
227
-------
of about 500 mg/1, the methane-forming activity of the acclimated biomass
was found to be about 95 percent of that for the noninhibited.
In a series of packed-bed anaerobic reactors, the effect of four mixed
inhibitors, including phenol, formaldehyde, acrylonitrile and ethyl acrylate,
48
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.
48
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
considerable difficulty would likely be experienced in the startup of a
full-scale anaerobic process treating an inhibitory waste.
49
A more recent study on the anaerobic treatment of oil shale retort
water reached the following conclusions:
(1) The retort water studied had to be pretreated to remove toxic
constituents (ammonia and sulfide) and to add nutrients in which it was
deficient (calcium, magnesium and phosphorus) before it could be successfully
treated by anaerobic fermentation. Pretreatment included pH adjustment,
air stripping and skimming, and nutrient addition.
(2) A digested sludge from a conventional municipal sewage treatment
plant was successfully acclimated to the retort water studied.
(3) A major fraction of the organics in the retort water studied was
stabilized by conversion to CH. and CO by anaerobic fermentation. BOD
4 2. 5
and COD removal efficiencies were 76-80%. Within the limits of experimental
error, the same removal rate was obtained for both BOD and COD.
(4) The effluent from anaerobic fermentation of the retort water
studied (BOD5:530 - 580 mg/1) might be suitable for treatment by conventional
aerobic processes.
(5) The growth of the methane formers, which stabilize the organics,
is nutrient limited in the retort water studied.
(6) Pretreatment of the retort water removed 49% of the BOD . This was
probably due to the reduction in solubility of high molecular weight fatty
acids at neutral pHs; they precipitate out of solution and do not exert a BOD.
228
-------
(7) During anaerobic fermahtation a major component removed from the
retort water studied was fatty acids.
(8) The long hydraulic residence time used in this study (about 50 days)
would not be used in practice. Cell recycle, which increases the cell resi-
dence and decreases the hydraulic residence time, should be to achieve hydrau-
lic residence times on the order of 2 - 3 days. In this study all recycle
was achieved by removing a sample of the digester mixed liquor, centrifuging
it at 2,500 rpm for 3 min, and returning the centrate (biological mass) to
the digester.
Based on the findings of these recent studies, the following recommenda
tionsmay be offered for coal gasification wastewaters:
(1) Experiments on anaerobic fermentation should be conducted with
continuously mixed reactors with long mean cell residence time, 6c, of
49
probably more than 50 days
(2) Pretreatment procedures, such as those described in other parts
of this report, will be required to reduce the concentration of any toxic
constituents below the threshold level.
(3) The nutrient requirements of anaerobic fermentation should be
established and any nutrient deficiency should be corrected by adding
appropriate chemical(s).
(4) It is unlikely that anaerobic fermentation alone can achieve the
treatment objective satisfactorily. After the technical feasibility of
anaerobic fermentation is established, then the possibility of following
anaerobic fermentation by aerobic biotreatment should be evaluated.
5.7.1 Design of Experiments on Anaerobic Fermentation
In general/ References 47 and 49 are excellent guides for conducting
experiments on anaerobic fermentation. Preliminary design of such experiments
is described in the following paragraphs.
A. Reactor Configuration
The reactor should be continuously mixed and heated to maintain a
temperature of about 35°C. It should also have provisions for feeding the
wastes, withdrawing effluent or samples and collecting and measuring the
gaseous products as shown in Figure 5-20. The capacity of the reactor
depends on th hydraulic residence time and the volumetric feed rate.
229
-------
EFFLUENT
or
SAMPLE
GASES
(CH4 & C02 etc/)
MIXER
FEED
C
ii
OJ
on
o
o
Figure 5-20. Schematic of continuously mixed batch anaerobic digester.
230
-------
B. Toxicants and Nutrients
For the same chemical there may be an optimal and toxic concentration as
shown in Table 5-21. Furthermore, the toxic levels may vary considerably
from the listed values because of synergistic or antagonistic effects of
other chemical species . However, the values in Table 5-21 may be used to
identify potential toxicity, required pretreatment, and nutrient deficiency.
C. Acclimation
If a digested sludge from a conventional municipal sewage treatment
plant is used to seed the reactor, the microbes in the digested sludge may be
acclimated to coal conversion wastes by incremental additions of coal conver-
sion wastewater to the municipal sludge feed. After each incremental addition,
the digester must be monitored until steady state conditions are attained. The
parameters used as indicators of a steady state should include volatile sus-
pended solids, volatile acids, and volume and composition of digester gas.
D. Cell Recycle
When it is desired to maintain a short hydraulic residence time and a
long cell residence time, the separation and recycle of the active biomass,
especially the methane former, will become necessary. Digester failures
have frequently been attributed to a population imbalance between acid
formers and methane formers. Methane formers are generally.more sensitive
to environmental conditions and usually grow more slowly than acid formers.
The population imbalance usually means a gradual washout of methane formers
and eventually a digester failure.
Although it is highly desirable to selectively separate and recycle
methane formers, all of the past experience shows that this is not practical.
The failures of anaerobic filters and anaerobic activated sludge in treating
47
petrochemical wastewaters might be due to the inability of these systems
to selectively separate methane formers from the rest of solids.
The successful cell recycle reported in Reference 44 indicates that
centrifugation may be a good technique to separate methane formers from
the liquid phase, but may not necessarily be a preferred technique to
separate methane formers from acid formers. The following separation
techniques are recommended for testing:
- Centrifugation
- Polymer flocculation followed by sedimentation
- Flotation using C02
231
-------
TABLE 5-21 OPTIMUM AND TOXIC LEVELS FOR ANAEROBIC FERMENTATION
(From Reference 49)
Parameter Optimum Level (a) Toxic Level (b)
(mg/1) (mg/1)
Sodium
Potassium
Ammonium (as NH )
Calcium
Magnesium
Soluble Sulfide (as S)
Nitrogen
Phosphorus
230
390
180
200
120
(c)
(d)
(d)
>4,600
>3,900
>1,350
>2,000
>1,200
> 200
-
—
(a) Maximum efficiency from an anaerobic treatment system can be obtained by
maintaining the major ions as close to their optimum values as possible.
(b) Toxic levels may vary considerably from the values shown due to
synergistic or antagonistic effects of other ions.
(c) At levels lower than 200 mg-S/1, sulfides can have a beneficial effect
by precipitating certain toxic heavy metals, e.g., Cu, Zn, Ni, Fe.
(d) Optimum levels of nitrogen and phosphorus have been demonstrated to depend
on the growth rate and concentration of organisms present in the system.
They are ~ 11% of cell volatile solids weight for N and 2% of cell
volatile solids weight for P.
232
-------
E. Test Parameters
Test parameters may be broadly classified into two types:
1. Operational parameters: BOD, TOC, volatile acids, alkalinity
volatile and total suspended solids, volume and composition of digester
gas.
2. Toxicity and nutrient parameters: ammonia, sulfide, phosphorus,
calcium, magnesium and others.
It has been reported in Reference 49 that a toxicity and a nutrient
deficiency problem could be differentiated by comparing total gas production
and gas composition. When toxicity was the cause of failure, the gas
composition changed significantly; CH decreased and CO and trace gases
increased. When the system was nutrient limited, total gas production
decreased while gas composition remained constant and comparable to that
observed in a properly operating digester.
F. Data Analyses
The 'Behavior of the anaerobic fermentation process may be described by
the same biokinetic model used for the air activated sludge process. In
other words, its behavior may be summarized by the value of four coeffi-
cients: k-, K , Y and k,. The same data analyses should be undertaken for
s d
anaerobic fermentation as for air activated sludge except that oxygen
balance does not apply to the anaerobic process. Values of the four
coefficients have been reported in the literature primarily for the
anaerobic treatment of municipal sludges.
5.8 The Place of Biological Oxidation
Biological oxidation can be considered to be in competition with
solvent extraction and adsorption as a method of removing organic contami-
nation from wastewater. The choice of process will depend on the cost and
on the quality of the effluent water.
5.8.1 Cost
Cost estimates for our two examples of activated sludge treatment and
a related third example are shown in Table 5-22. The cost of an air
activated sludge plant is always higher than the cost of an oxygen acti-
vated sludge plant. This is because the slow kinetics with air necessitate
much larger basins than are needed for oxygen and the expense of pure
oxygen is noj high enough to offset it. The cost of the aeration basin is
233
-------
TABLE 5-22 COST OF ACTIVATED SLUDGE TREATMENT
Flow (10 gal/day)
BOD (mg/1)
Capital investment, $10
Capital related charges,
$/103gal
Oxygen, $/10 gal
Other charges
Total charges, $/10 gal
$/100 Ib BOD
Example 1
Air Oxygen
0.3
Example 2
Air Oxygen
0.17
0.3 0.3 0.17 0.17 0.41
2,500 2,500 18,000 18,000 9,000
1.5 0.8 6.6 3.2 5.6
Example 3
Air Oxygen
0.41
9,000
3.4
3
0
3
18
.3
0
.6
.9
.8
1.
0.
0.
2.
9.
7
12
18
00
68
25
2
27
18
.04
0
.72
.76
.46
12.
1.
2.
15.
10.
1
42
21
73
46
7
1
8
11
.83
0
.10
.83
.73
4.
0.
0.
6.
8.
74
64
87
25
32
234
-------
over 70% of the capital cost of an air activated sludge plant. It is also
apparent that the correlation for cost is as $/100 Ib BOD removed. All the
oxygen plants lie close to $10/100 Ib BOD. For municipal plants with
intake waters having about 200 mg BOD/1 it is not true that the cost is
proportional to the rate of removal of BOD. This proportionally only holds
true for higher concentrations of feed water.
39
In an earlier study we also found cost to be proportional to the
rate of removal of BOD. In that study we used a trickling filter plus
faster kinestics and less oxygen consumption than in this report and found
$2.50/100 Ib BOD removed. The cost, $2.50, is too low. The cost in
Table 5-22 is not necessarily correct and needs reevaluation. What does
seem to be correct is that cost is proportional to the rate of removal of
BOD and that oxygen is preferable to air.
In other parts of this report we show that the cost of solvent extrac-
tion is proportional to the rate of water flow and nearly independent of the
concentration. Thus, solvent extraction should be considered first for
higher concentrations, as in our Example 2. Activated sludge should be
considered for lower strength wastewaters and for effluents from solvent
extraction. Probably, but we are not certain, solvent extraction is prefer-
able aboye 2,000 to 4,000 mg BOD/1 and activated sludge below 2,000 to
4,000 mg BOD/1.
The cost of solvent extraction increases as the fraction of contami-
nant removed increases. Solvent extraction is optimum when high fractions
removed (perhaps greater than 80%) are not required. In many situations we
expect that solvent extraction will be sollowed by adsorption or biological
oxidation.
To prevent toxicity most wastewaters must be diluted for biological
oxidation to occur. In the usual activated sludge plant this means that the
effluent must be at least .dilute enough not to be toxic. Furthermore, as
has been shown in the examples, sludge residence times large enough to be
about twice the minimum and large enough to require a positive sludge recycle
will usually result in destruction of more than 90% of the influent BOD.
We expect biological treatments to be operated with an effluent BOD in the
range 300 to 50 mg/1, usually at the lower level.
235
-------
The correlation of cost for an absorbtion process is not as simple
as for solvent extraction and activated sludge, neither is the design as
reliable. The cost depends on both the water flow rate and on the influent
concentration. It does seem, however, that low effluent concentrations
can be obtained without incurring a high cost. The economical range of
concentrations for using adsorption is probably close to the same range
for biological oxidation.
Both solvent extraction and adsorption, but not biological oxidation,
allow for recovery of the organic matter which may have some value, if only
as a fuel.
5.8.2 Quality of Effluent
Biological oxidation removes only biodegradable compounds. However,
some additional removal of even nonbiodegradable chemicals occurs by
adsorption onto the sludge. It is the experience at several coal conver-
sion pilot plants that wastewaters are satisfactorily biodegradable.
Phosphorous, as a nutrient, must be added as biological treatment and an
unknown, small, concentration will be left in the effluent. Ammonia as a
nutrient will be left over as well, but this is not usually different from
the competitive processes.
Since biological treatment converts soluble organic contaminants partly
to a sludge, the effluent will have suspended solids in the range 30-100 mg/1.
The quality of water effluent from solvent extraction and adsorption
is not well known. Probably solvent extraction is best for removing simple
phenol (CgH OHl, less good for cresols and resinol, and least good for
fatty acids: and two ring phenols. Adsorption may be quite good for removing
higher molecular weight phenols; the evidence is not conclusive.
236
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REFERENCES - Section 5
1. Winter, R.L., Heimark, E.L. and Uhl, V.W., "Submerged Turbine Aeration
Design," Philadelphia Mixers Corporation, King of Prussia, Pa., 19406.
2. Lawrence, A. W. and McCarty, P. L., "Unified Basis for Biological Treatment
Design and Operation," Journal of the Sanitary Engineering Division,
American Society of Civil Engineers, 96, 757, 1970.
3. Schmidt, F.L., Wren, J.D. and Redmen, D.J., "The Effect of Tank Dimensions
and Diffuser Placement on Oxygen Transfer," J. Water Pollution Control
Federation, 50 (No. 7) 1950-1967, July 1978.
4. Physicochemical Processes for Water Quality Control, edited by W. J. Weber,
Jr., Wiley-Interscience, 1972.
5. Price, K.'S. , Conway, R.A. and Cheely, A.H. , "Surface Aerator Interactions,"
Journal of the Environmental Engineering Division, ASCE, 283-300, June'73.
6. Eckenfelder, W.W. and Ford, D.L., "Engineering Aspects of Surface Aeration
De~sign," Proceedings of 22nd Purdue Industrial Waste Conference, 279-291,
196 .
7. Dick, R.I., "Role of Activated Sludge Final Settling Tanks," Journal of the
Sanitary Engineering Division, American Society of Civil Engineers, 96,
SA2, 423-436, April 1970.
8. Dick, R.I., "Gravity Thickening of Waste Sludges," Proceedings Filtration
Society, Filtration and Separation, 9^(2), 177-183, March/April 1972.
9. Dick, R.I. and Young, K.W., "Analysis of Thickening Performance of Final
Settling Tanks," Proceedings of the 27th Industrial Waste Conference,
33-54, 1972.
10. Mynhier, M.D., and Grady, C.P. Jr., "Design Graphs for Activated
Sludge Process," Journal of the Environmental Engineering Division, ASCE
Vol. 101, No. EE5, Proc. Paper 11659, 829-846, October 1975.
11. "Prospects Strong for Wastewater Oxygenation," Chemical and Engineering
News, 17-18, March 28, 1977, p. 17-18.
12. Vaseleski, R.C.,"The UNOX Process-Effective Wastewater Treatment Practice,"
presented at 70th Annual AIChE Meeting, New York, November 13-17, 1977.
13. Eckenfelder, w.W., Jr. et al, "Oxygen Activated Sludge-Considerations for
Industrial Applications," presented at 70th Annual AIChE Meeting, New York
November 13-17, 1977.
14. D'Antoni, J.M. & Steimle, S.E., "Effects of Dissolved Oxygen in the
OxygenaJion Activated Sludge Process," presented at 70th Annual AIChE
Meeting New York, November 13-17, 1977.
237
-------
15. Benefield, L.D., Randall, C. W., and King, P.H. , "The Effect of High
Purity Oxygen on the Activated Sludge Process," Jour. Water Poll.
Control Fed., 49 269-279, 1977.
16. Adams, C.E., Jr., Stein, R.M., and Eckenfelder, W.W., Jr., "Treatment of
Two Coke Plant Wastewaters to Meet EPA Effluent Criteria," 27th Proceed-
ings of Purdue Industrial Waste Conference, 864-880, May 1974.
17. Mueller, J.A., Mulligan, T.J., and DiToro, D.M., "Gas Transfer Kinetics
of Pure Oxygen System," Journal of the Environmental Engineering
Division, ASCE, Vol. 99, No. EE3, 269-282, June 1973.
18. Watkins, J.P-, Mulligan, T.J., and Shema, J., "Pure Oxygen and Conven-
tional Air Activated Sludge Treatment - Pilot Plant Evaluation for a
Pulp and Paper Mill Waste," Proceedings of the 31st Purdue Industrial
Waste Conference, May 1976.
19. Sherrard, J.H., "Significant Correlation Exists Between Design
Techniques," Water & Wastes Engineering, 50-52, January 1978.
20. Parsons, W.A., and Nolde, W., "Applicability of Coke Plant Water
Treatment Technology to Coal Gasification," Proceedings of the Third
EPA Symposium on Environmental Aspects of Coal Conversion Technology,
Hollywood, Fla., September 1977.
21. 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 Purdue Industrial Waste Conference, 617-630, May 1974.
22. Orford, H.E., Rand, M.C., and Gellman, I., "A Single Dilution Technique
for BOD Studies," Sewage and Industrial Wastes 25, ^., 284-289, (1953).
23. Kostenbader, P.D. and Flecksteiner, J.W., "Biological Oxidation of Coke
Plant Weak Ammonia Liquor," Journal Water Pollution Control Federation,
_4JL(2) , 199-207, February 1969.
24. Rex Chainbelt, Inc., "A Mathematical Model of a Final Clarifier,"
Environmental Protection Agency Report 17090 FJW 02/72, February 1972.
25. Reap, E.J., et al, "Wastewater Characteristics and Treatment Technology
for the Liguification of Coal Using H-Coal Process," Proceedings of the
32nd Purdue Industrial Waste Conference, May 1977.
26. 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 Environmental
Aspects of Fuel Conversion Technology II, Hollywood, Fla., December 15,
1975, Environmental Protection Agency, Research Triangle Park, N.C.,
EPA-600/276149.
238
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27. Zitrides, T.G., "Using Customized Bugs for Biological Waste Treatment,"
Plant Engineering, 117-119, June 23, 1977.
28. Metcalf & Eddy, Inc., Wastewater Engineering, McGraw-Hill, 1972.
29. Eckenfelder, W.W., Jr. , Water Quality Engineering for Practicing
Engineers, Barnes & Noble, 1970.
30. McWhirter, J.R. , ed., The Use of High-Purity Oxygen in the Activated
Sludge Process, 2 volumes, CRC Press, 1978.
31. Oxygen Activated Sludge Wastewater Treatment Systems, EPA Technology
Transfer, August 1973.
32. McCarty, P.L., "Anaerobic Treatment of Soluble Wastes," Advances in Water
Quality Improvement, E.F. Gloyna and W.W. Eckenfelder, Jr., eds., 336-
352, University of Texas Press, Austin, Tex., 1968.
33. Bryan, Edward H., "Two-stage Biological Treatment Industrial Experience,"
Proceeding of Eleventh Southern Municipal & Industrial Waste Conference,
North Carolina State University, 1962.
34. Adams, C.E. and Eckenfelder, W.W., eds., Process Design Techniques for
Industrial Waste Treatment, AWARE, Inc., Enviro Press, Nashville,
Term. , 1974.
35. Gilbert, R.G. and Libby, D., "Field Testing for Oxygen Transfer and
Mixing in Static Mixer Aeration Systems," proceedings of the 32nd
Purdue Industrial Waste Conference, May 1977.
36. Neerber, R.F. , "Compressor Selection for the Chemical Process Industries,"
Chemical Engineering, 78-93, January 20, 1975.
37. Neufeld, R.D., Drummond, C.J. and Johnson, G.E., "Biokinetics of
Activated Sludge Treatment of Synthane Fluidized Bed Gasification
Wastewaters," 175-186, Preprints of American Chemical Society, Division
of Fuel Chemistry, Vol. 23, .No. 2 for Meeting of March 12-17, 1978.
38. Johnson, G.E. et al, "Treatability Studies of Condensate Water from
Synthane Coal Gasification," Pittsburgh Energy Research Center, Depart-
ment of Energy, Report PERC/RI-77/13.
39. Goldstein, D.J. and Yung, D., "Water Conservation and Pollution Control
in Coal Conversion Process," EPA Report 600/7-77-05, Research Triangle
Park, N.C., June 1977. NTIS Catalog No. PB 269-568/2WE.
40. Leyenspiel, O., Chemical Reaction Engineering, Jchn Wiley & Sons, 1972.
239
-------
41. Standard Methods for the Examination of Water and Wastewater, 14th ed.
Amer. Pub. Health Assn., Washington, D. C. (1976).
42. Methods for Chemical Analysis of Water and Wastes, U.S. EPA Technology
Transfer (1974).
43. Annual Book of ASTM Standards (1976), Part 31 Water, American Society for
Testing and Materials, Philadelphia, Pa.
44. Lau, C.M., "Staging Aeration for High Efficiency Treatment of Aromatic
Acids Plant Wastewater," Proceedings of the 32nd Purdue Industrial Waste
Conference, May 1977.
45. Eckenfelder, W.W. and Ford, D.L., Water Pollution Control, Experimental
Procedures for Process Design, Jenkins Publishing Co., Austin, Texas,
1970.
46. Smith, R.M., "Some Systems for the Biological Oxidation of Phenol-Bearing
Waste Waters," Biotechnology and Bioengineering, J5, 275-286, 1963.
47. "Anaerobic Treatment of Synthetic Organic Wastes," Water Pollution Control
Research Series 12020 DIS pi/72, Environmental Protection Agency.
January 1972.
48. "Identification and Control of Petrochemical Pollutants Inhibitory to
Anaerobic Processes," Environmental Protection Technology Series,
EPA-R2-73-194, April 1973.
49. Ossio, E. A., Fox, J.P., Thomas, J.F., and Poulson, R.E., "Anaerobic
Fermentation of Simulated In Situ Oil Shale Retort Water," Preprints of
Papers presented at Division of Fuel Chemistry, 175th American Chemical
Society Meeting, Anaheim, Vol. 23, No. 2., 202-213, March 12-17. 1978.
50. Kugelman, I.J. and Chin, K.K., "Toxicity. Synergism and Antagonism in
Anaerobic Waste Treatment Processes," in Advances in Chemistry Series
No. 105, American Chemical Society, 1971.
51. Wastewater Treatment Plant Design, Water Pollution Control Federation,
Washington, D.C. 20037, 1977.
52. Water Purification Associates, work under DOE Contract EF-77-C-01-2635
to be published as "Conceptual Designs for Water Treatment in
Demonstration Plants".
53. Flanagan, M.J. and Bracken, B.D., "Design Procedures for Dissolved Oxygen
Control of Activated Sludge Processes", U.S. EPA Report 600/2-77-032,
Cincinnati, Ohio, June 1977.
240
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6. COOLING TOWER CONTROL
6.1 Introduction
Cooling towers in coal conversion plants have uncommon control problems
because they will usually be fed with a mixture of source water and treated
process condensate. Treated process condensate may have a lower salt content
than the river water which decreases the possibility of scale formation by
slightly soluble salts. On the other hand, condensate will usually have a
higher organic content (phenol and COD) and a higher ammonia content than is
common in cooling tower makeup. In this section we will estimate the
consumption of cooling water and describe methods of dealing -with the parti-
cular contaminants in the mixed feed.
One approach is to pretreat the process condensate to very low levels of
ammonia and COD. Pretreatment techniques are the subject of most of this
report so the cost of this approach can be, as it should be, estimated. If
the process condensate is vigorously pretreated, then those cooling tower
control problems unique to coal conversion disappear. In this case cooling
tower control is no different from usual practice. Our purpose in this
section is to show how a cooling tower might be controlled when the makeup
water has organic contamination and to suggest upper limits to acceptable
concentrations of contamination.
As found by most authors, it is useful to classify the problems of
cooling tower control under four headings:
Scale formation
A wet cooling tower is an evaporator, and salts dissolved in the makeup
water concentrate, often to the point of precipitation. The precipitate tends
to adhere to heat transfer surfaces forming a hard scale.
Fouling
Not only may the makeup water contain silt and BOD which may be converted
to sludge, but the circulating water in its passage through the tower scrubs
dust out of the air. Circulating water thus contains an increasing amount of
suspended matter, which will settle out in stagnant spots in the pipes and
heat exchangers.
Corrosion
The well-oxygenated circulating water is very corrosive to heat transfer
surfaces.
241
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Biological Control
Circulating cooling water is warm and well-oxygenated and is an ideal
habitat for microbial growth. The water is seldom sterile when fed to the
system and, in any case, receives a steady supply of airborne growth.
Untreated cooling systems are subject to fungal rot of the wooden parts of
the tower, bacterial corrosion of iron and bacterial production of sulfide,
algae growth in the sunlit portions of the tower, and suspended sloughed-
off growth that can lodge in the system and block the flow. Biocidal
chemicals must be added to control growth.
Before beginning the description of cooling tower control, the consump-
tion of cooling water will be estimated.
6.2 Consumption of Cooling Water
To introduce the terminology of a cooling tower, the material balance
can be written as
Makeup = Slowdown + Evaporation
By definition
Cycles of Concentration = (Makeup)/(Slowdown)
we also write
Evaporation = (Cooling Load)/(Heat dissipation rate),
so
so
M(l - 1/X) = Qc/q (1)
where M = makeup rate, Ib/hr
X = cycles of concentration
Q = cooling load, Btu/hr
q = heat dissipation rate, Btu/lb evaporated
The Cooling Load
This is the total rate of energy loss to the atmosphere through the
cooling tower from the whole plant. If the plant is fully designed, the
cooling load will be known. An accurate preliminary estimate can be
obtained by considering the major points of cooling as described in
References 1-4. If time for an estimate cannot be taken, we suggest
assuming that the conversion plant is 65% efficient and that 25% of the
242
-------
unrecovered heat is dissipated by evaporative cooling. The cooling load
will then be
(0.25)(1-0.65)(energy in total coal feed)
or
(energy in product fuel)(0.25)(0.35)/(0.65)
The Cycles of Concentration
This is decided upon in the usual way by comparing the cost of blow-
down disposal with the cost of water treatment. When the cooling tower
makeup is contaminated, blowdown to the river without treatment may not be
suitable. In most coal conversion plants we foresee the need to limit the
blowdown quantity so that it can be disposed of conveniently, for example,
by using it to quench and wet the ash. This may eliminate blowdown treat-
ment, but will usually require alternative treatment to permit high cycles
of concentration without the formation of scale. We will return to the
question of blowdown disposal in a later section.
Heat Dissipation Rate
In a cooling tower the air is heated at the same time as the water is
cooled by evaporative and convective heat transfer. It is not correct to
take the heat dissipation rate to be the latent heat of evaporation of
water (.1000 Btu/lb). A suitable estimating number is 1,400 Btu/lb evapo-
rated. A sample of a rough estimate using the quantities discussed so far
is given in Table 6-1.
A more accurate estimate of the heat dissipation rate requires climatic
information. Given this information cooling tower calculations can be made
month by month and the annual average heat dissipation rate determined.
Cooling tower calculations are lengthy and the loss in accuracy will be
5,6
less than 3% if the results of the calculations made by Leung and Moore
are used. Their results are shown on a copy of their graph as Figure 6-1.
For any period of time, such as a month or a quarter of the year, for which
the average wet bulb temperature and average relative humidity are known,
the average heat dissipation rate can be read from Figure 6-1.
6.3 Scale Prevention
Scaling is prevented by controlling the concentration of species which
form part of slightly soluble salts, particularly salts whose solubility
decreases with an increase in temperature. These salts tend to precipitate
243
-------
TABLE 6-1. APPROXIMATE MATERIAL AND ENERGY INFORMATION ON A
65% EFFICIENT COAL GASIFICATION PLANT
Product:
Coal:
Unrecovered energy:
Fraction lost to
cooling water:
Evaporation rate:
Coal ash:
Water lost with
coal ash:
Concentration in
cooling tower:
Makeup:
10 Btu as gas
106/0.65 = 1.54 x 106 Btu
Approximately 154 Ib
0.54 x 10 Btu
25%
1400 Btu dissipated per pound of
water evaporated
96 Ib water evaporated
Approximately 15 Ib
35% moisture in wet solids
8 Ib water
13 fold
104 Ib
244
-------
80
75
70
65
60
55
50
Elevation: Sea Level
40
35
30
25
0
Til
ILL
TLUJ\
LL
LL
LL
I
I
0.7
0.8
0.9
EVAPORATED WATER, LBS/10J BTU
1.0
Figure 6-1. Pounds of water evaporated per 1000 Btu of heat
transferred in a wet cooling tower.
245
-------
on the surface of the heat exchanger forming hard, adherent scale. The
most common scales are calcium carbonate and calcium sulfate with addi-
tional problems from silica, magnesium silicates, and calcium phosphates.
it _^__j. — 2—
The species that must be controlled are Ca , Mg , SiC>2. CC>3 and PC>4 .
Phosphate deserves special mention because it is not a common component of
cooling tower makeup water. It is present in coal conversions plants
because (1) it was added to an upstream biological treatment plant as a
nutrient and cannot all be consumed in the treatment, and (2) it is needed
as a nutrient to encourage biological oxidation in the cooling towers. The
solubility of calcium phosphate is given in Section 6.3.1.
Most makeup waters contain carbon dioxide in excess of the concentra-
tion in equilibrium with the atmosphere. Quite high alkalinities occur,
nearly all as bicarbonate (HCO ) . In the cooling tower there is a tendency
for CO to be driven off converting some of the alkalinity to carbonate
(CO ) and causing precipitation. This is common and is usually prevented
by adding sulfuric acid which replaces some of the alkalinity with equiva-
lent sulfate. Scale by calcium sulfate is sometimes prevented by
softening the makeup water. Lime and soda ash are added to precipitate
calcium. This has the added advantage of removing some carbonate alka-
linity as well. If desired, the softener can be operated to precipitate
magnesium as hydroxide. Silica quite readily adsorbs onto the surface of
magnesium hydroxide and is partly removed with the magnesium. This is
advantageous. Phosphate can be removed by the addition of lime.
In addition to procedures for removal of the precipitating species,
scale formation can be prevented even when precipitation is occurring by
adding chemicals which inhibit crystal growth. Some of the most useful
modern compounds to prevent scale are polymers of acrylates and acrylamide.
Acidification, lime treatment, and the use of antiscalants are not
specific to coal conversion and are dealt with in other literature. Designs
are not given in this report. The final point to be made is that no matter
which treatment is used, dissolved salts will concentrate and some blowdown
is essential to control their concentration in the circulating water.
There is always some drift of water droplets from a cooling tower and this
is a sort of accidental blowdown. When the circulating cooling water is
246
-------
very concentrated, the drift may damage foliage and land where it settles.
For a high concentration cooling system, the best modern drift eliminators
will be used, slowdown cannot acceptably be prevented by allowing a high
drift.
6.3.1 Solubility of Calcium Phosphate
Since phosphate is a contaminant not commonly found that will be found
in most coal conversion plants, it is worth discussing its control. The
solubility of calcium phosphate decreases as the pH increases. Calcium
phosphate can be deliberately precipitated from a water stream by adding
lime. Alternatively, precipitation can be prevented in circulating cooling
water by lowering the pH. Solubility calculations suggest that the precipi-
tation of calcium phosphate should be a serious problem. Experience
24
suggests that it is not a problem . A pseudosolubility product has been
successfully used in water treatment and the incorporation of this
pseudosolubility product into a simplified solubility theory gives an
equation which is probably adequate for determining the proper pH to prevent
precipitation. This equation is derived below.
Simplified Theory
The least soluble salt is hydroxyapatite, Ca (PO ) OH. Tricalcium
phosphate, Ca (PO ) , is also only very slightly soluble. If fluoride is
O ™i ^
present fluoroapatite, Ca (PO) F, which is the least soluble of all, may be
found in the precipitate. Phosphate exists in solution in the forms H PO ,
H PO , HPO and PO , in proportions dependent on pH. A method of deter-
2444 ^5
mining permissible concentrations has been given by McCoy for tricalcium
phosphate. A much more complex procedure has been used by DeBoice and
25
Thomas who accounted for the presence of carbonate species plus twelve
complexes. This complexity is not justified here because of the inaccur-
acies of the solubility products. McCoy used a graphical approximation
which we have not thought to be necessary today when a calculator is
readily available.
It is important to note, as stated by McCoy, that the following calcu-
lation predicts equilibrium values. With increasing alkalinity the rate of
crystal growth is slowed and even at pH 8, several hours may elapse before
any precipitation is formed. Furthermore, with suspending agents present the
precipitate does not necessarily form a hard scale. Because of this the calcu-
lations have been approximated by making all activity coefficients equal to one.
247
-------
Following McCoy take
K = 7.5 x 10 3 = [H+] x [H2P04 ]/[H3PO4] (2)
K2 = 6.2 x 10~8 = [H+] x [HP04=]/[H2P04~] J3)
K3 = 4.8 x 10~13 = [H+] x [P04=]/[HP04=] (4)
Analysis gives tot-al orthophosphate which is
[o-PO ] = [H3P04] + [H2P04 ] + [HP04 ] + [P04 ] (5)
To develop the required relationships we will need to express both [PO ] and
[HPO ~~] in terms of [H ] and [o-PO ] only.
From Equations (2), (3) and (4)
[HP04 ] = [H] x [P04~]/K3 (6)
= [H+]2 x [P0~]/KK (7)
[H3P04] = [H+]3 x [P04 l/^K^ (8)
If Equations (6) , (7) and (8) are substituted into Equation (5) and the
equation is rearranged
[P04 ] = [o-P04] x K^K^/F (9)
where
F = [H+]3 + Jfil[E+]2 + K^tH"1"] + K K K (10)
., -3pH . -2pH -pH
=10 * + K. X 10 * + K K X 10 * + K,K K (11)
248
-------
If the numerical values for K , K2 and K from Equations (2), (3) and (4),
are put into Equation (10)
p = 10 — + 7.5 x 10 J ^" + 4.65 x 10 •LU'pn + 2.23 x 10~" (12).
Also, Equations (9) and (6) give
[HP04~] = [o-P04] x K K2 x [H+]/F (13)
Equation (13) is used below to determine the solubility of hydroxyapatite.
Now consider the solubility of tricalcium phosphate. McCoy ~ gives
[Ca++]3 x [P045]2 = Kscap = 5 x 10~3° (14)
so,
1.5 Log [Ca++] + Log [PO4=] = - 14.651 (15)
In Equation (15)-, [PO ~] can be taken from Equation (9) and numerical values
can be substituted for K , K and K to obtain
1.5 Log [Ca++] + Log [o-PC>4] - Log F = 7.00 (16)
Finally, it is convenient to change the units of concentration. Let CaH
(meaning calcium hardness) be the concentration of calcium in mg/1 as calcium
carbonate, then
105[Ca++] = CaH
Let o-PO be total phosphate in mg/1, then
0.95 x 105[o-PO4] = o-P04
In the new units Equation (16) becomes
1.5 Log CaH + Log o-PO - Log F = 19.478 (17)
249
-------
Equation (17) , with Equation (12) , is an equation relating the concentra-
tion of calcium, the concentration of orthophosphate and the pH at equilibrium.
If for any water the lefthand side of Equation (17) is greater than the
righthand side, then the water is supersaturated. If the lefthand side is
less than the righthand side, the water is undersaturated.
As an example, the permissible calcium hardness at equilibrium is given
on Table 6-2 for various orthophosphate concentrations. The results for
10 mg/1 are close to McCoy's results.
Now consider the solubility of hydroxyapatite. Truesdall and Jones
give the equation
Ca (PO )3OH + 3H O«-5Ca + 3HPC>4 + 4OH
for which
Log K =5 LoglCa**] + 3 Log[HPO. ] + 4 Log K + 4 pH
S. Ap 4 W
= -29.233 - 8'996TX 1Q3 (18)
—59
For an example consider K at 30°C when it has the value 1.19 x 10
-14 = p
Take KW = 10 . If [HPC>4 ] from Equation (13) is substituted into Equation
(16) , if numerical values for K and K are
tration are changed as before, one obtains
(16), if numerical values for K and K are taken,and if the units of concen-
1.667 Log CaH + Log o-PO - Log F + pH/3 = 21.670 (19)
Calculations from Equation (19) are also presented on Table 6-2 . They are
quite close to the results for tricalcium phosphate.
As a final example, Chen and others in their tabulation use a
solubility limit expressed as
(Ca^mg/l) x (o-PO mg/1)2 = 0.06
250
-------
or, in the units of this section,
(CaH)3 x (o-P04)2 = 0.94
If this is substituted for Equation (14) then Equation (17) becomes
1.5 Log CaH' + Log o-PO - Log F = 19.137
which makes little difference. However, Chen and others use the solubility
product at all pH which seems to limit all waters to the concentration
permissible at a pH of about 10.
A More Practical Equation
In the circulating cooling water of many coal conversion plants we can
expect about 10 mg/1 phosphate and 600-800 mg/1 calcium as CaCO . The
solubility calculations given above suggest that a suspension of calcium
phosphate will usually be present. Experience is that phosphate scale will
not be troublesome, but most plants will require the addition of some
suspending or antiscalant chemical.
25
For the purposes of water treatment, DeBoice and Thomas have
successfully used a pseudosolubility product
[Ca+V x [P0=]2 = 10~25'46 = 3.467 x ID*26 (20)
This is very different from Equation (14). DeBoice and Thomas also used
= 7.2 x 10~3, K = 6.3 x 10~8 and K = 4.5 x 10~13
which are close enough to Equations (2) through (4) .
If Equation (20) is used instead of Equation (14)
the righthand side of Equation (15) = -12.73
the righthand side of Equation (16) = £.921
251
-------
TABLE 6- 2. EQUILIBRIUM CONCENTRATIONS OF CALCIUM
(in mg CaCO /I) FOR VARIOUS TOTAL ORTHOPHOSPHATE
CALCULATED FROM EQUATIONS (17) AND (19)
Equilibrium Calcium Concentration
for tricalcium phosphate for hydroxyapatite
orthophosphate (mg/1) 10
E5.
5.0
6.0
6.5
7.0
7.5
8.0
9.0
Log F -
-12.122
-14.
-15.
-15.
-16.
-17.
-18.
099
045
915
653
268
325
1.73xl04
830
194
51
16
6
1
3
3.85xl04
1850
434
114
37
14
3
1
8.02xl04
3850
902
237
76
30
6
10
1.34xl04
552
119
28
8
3
0.4
3
2.76xl04
1140
244
58
17
6
0.8
1
5.34x10
2200
472
113
32
11
1.
4
6
TABLE 6-3. EQUILIBRIUM CONCENTRATIONS OF CALCIUM
CALCULATED FROM EQUATION (21)
10 mg/1 Orthophosphate
pH Calcium as mg/1 CaCO
5.0 330,000
6.0 16,000
6.5 3/000
7.0 980
7.5 310
8.0 120
9.0 24
252
-------
and Equation (17) becomes
1.5 Log CaH + Log o - PO4 - Log F = 21.399 (21)
In Equation (21), F is given by Equation (12).
Calculations from Equation (21) are given in Table 6-3. Experience is
lacking, but Equation (21) seems to give answers that are safe to use.
Equation (21) is suggested for a control equation.
6.4 Biological Control
If treated wastewater is added to the cooling system, the control of
biological growth can become a serious problem since the water contains nutri-
ents. Relevant experience obtained in the petroleum industry has shown that,
within limits, organic contamination is permissable in cooling water makeup.
In a cooling tower several changes occur in the circulating water.
The first change is that ammonia is vaporized with water. Ammonia does not
concentrate in a cooling tower and will be reduced in concentration if a
large eno-.gh fraction (approximately more than 85%) of the makeup water is
evaporated rather than blown down. This is easily calculated from vapor
pressure data and has been demonstrated .
It has been completely demonstrated that refinery phenolic wastewaters
can be used in a: cooling tower and that bio-oxidation of phenol will occur
with very high removal efficiencies. This has been practiced at the Sun
Q Q
Oil refinery in Toledo, Ohio , at the Mobil Oil refinery in East Chicago
and at other refineries according to a brief note by Maguire of Betz
Laboratories . In Table 6-4 we have tried to bring together some of the
pertinent information.
It must be remembered that "phenol" in a refinery sour water probably
contains a higher fraction of C,H OH than will be found in the coal conver-
8
sion wastewater, so the analogy is less than perfect. Mohler and Clere
identified several species of bacilli and cocci in their tower; five species
utilized 200 ppm phenol and two of the five grew in a 1,000 ppm phenol/
mineral salts medium. The bacteria were mesophilic with optimum growth in
9
the range 60-100°F. Hart reports that his towers operated at 80-86°F.
Q
Necessary conditions for successful bio-oxidation are low sulfide
10 8
(below 2 ppm is suggested ) , and only slow, small excursions in pH .
References 8 ind 9 report a pH range of 7.8 to about 8.3. Hydrogen sulfide
in coal conversion wastewater is lower than in stripped refinery sour water.
253
-------
Chlorination to prevent slime was used at Mobil , but had to be carefully
controlled to maintain bio-oxidation. Maguire reports similarly and suggests
the use of acrolein or other non-oxidizing biocides to prevent the formation of
toxic chlorinated phenols. We recommend not using chlorine for coal conversion
wastewaters because of the possible release of chlorinated aromatic compounds.
Most of the suppliers of proprietary chemical mixtures can supply biocides
other than chlorine. However, experience with these chemicals in a cooling
tower which is intended to be biologically active has not been reported. We
can make no firm recommendation.
Corrosion of steel has been low at both Mobil and Sun Oil. Low corrosion
is also experienced when treated sawage is used in a. cooling tower
6.5 Suggested Limits on Quality of Makeup
Before setting quality limits of organic contamination in cooling system
makeup, an example is useful. The Lurgi process plant designed by El Paso to
make 250 x 10 srf/day pipeline gas in New Mexico will require cooling water
makeup of about 1,500 x 10 Ib/hr of which 50% will be treated process conden-
sate. The quality of treated process condensate is not precisely known, but two
12 13 ~
estimates have been made ' . They are reproduced in Tables 6-5 and 6-6-
When treated condensate is diluted with an equal volume of San Juan River
water, a probable composition of the mixture is that shown in Table 6-7. With
the addition of acid and antiscale chemicals, this water can be concentrated
14
eightfold and circulated at a pH of about 7.5 without scale formation . It
has been assumed by the designers, and we agree, that the organic contamination
is not bad and that the tower will operate. The most important foreseeable
problem is a high suspended solids as discussed below.
Of all the plant designs we have seen or made, the El Paso plant has the
highest organic contamination in the makeup water. We have used this to
suggest upper limits of permissable concentrations in makeup. The limits
shown in Table 6-8 apply to the mixed makeup; they are high and will require
good control by the operators, but they are not necessarily the maximum.
Please recognize that the limits are rather arbitrary.
The limit on ammonia requires discussion.
Ammonia is driven off in a cooling tower and the most important limit to
the permissable concentration in the makeup water is the formation of noxious
fumes. Our calculations show that the odor threshold for ammonia in the
cooling tower plume will not be exceeded with ammonia concentrations of up to
254
-------
TABLE 6-4. EXPERIENCE WITH BIO-OXIDATION OF PHENOLIC
REFINERY WASTEWATER IN A COOLING TOWER
Mobil Oil Summary
9
Refinery from Ref.10
Makeup concentration of phenol ppm 13 to 70 " 14 (a) 33 to 110
Fraction removed (%) >99.4% 92-97% avg > 90%
Tower loading:
gpm/ft2 3.3 to 3.4
Ib phenol/day 102 to 840 100
Ib phenol/Ift3)(day) - 68(a)
Notes:
(a) Estimate by us from information in the publication.
TABLE 6-5. 'CALCULATED EFFLUENT FROM PHENOSOLVAN PROCESS
(Reference 13 )
mg/1
Monohydric phenol 27
Polyhydric phenol 432
Other organics 1,834
BOD 3,151
Basis;
Assumed composition of crude phenol:
85V monohydric phenol
15% polyhydric phenol
5% other organics
Assumed extraction recoveries:
99.5\ for monohydric phenols
60.0% for polyhydric phenols
15.0% for other organics
BOD- factors:
monohydric phenols • 1.7
polyhydric phenols « 1.9
other organics >0.7
255
-------
TABLE 6-6. PHENOSOLVAN PLANT PERFORMANCE
SASOL FACILITY (From Ref.12 )
(For combined clean and contaminated gas liquor stream)
COMPONENT
Phenols
Sodium
Ammonia (free)
Ammonia (fixed)
Suspended Tar & Oil
CN
Total S
Fatty Acids as C_H O
CO,
Concentration, ppm
COMPONENT
Phenols (Steam volatile)
Phenols (Bound)
Fatty Acids as C^R^p
Ammonia as Nitrogen
Hydrogen Sulfide
CN
Fluoride
Chloride
Calcium (As Ca)
Iron (As Fe)
Orthophosphate
Total Dissolved Solids
Suspended Solids
COD
pH
EFFLUENT
Concentration, ppm
1
60-160
560
215
12
1
56 gm/1
25
18
. 1 mg/1
2.5
875
21
1,126
8.4
256
-------
TABLE 6-7. PROBABLE COMPOSITION OF MAKEUP TO COOLING TOWER IN
LURGI PLANT (50% from river & 50% treated effluent)
mg/1
Phenols 100
BOD 1,000
COD 2,000
Ammonia as N 100
Total P 1
HCO 80
Ca 37
Mg 5
Na 22
Cl 17
SO, 69
TABLE 6-8- SUGGESTED LIMITS ON MAKEUP TO A COOLING TOWER
mg/1
Phenols 100
BOD 1,000
COD >2,000
Ammonia as N 100
HS 2
257
-------
500 mg/1 so 100 mg/1 (shown in Table 6-8) should be no problem. Because of
ammonia, copper should not be used as a material of construction, and
chlorination which is not planned will be reduced in efficacy.
6.6 Fouling Control
To prevent fouling by settled suspended solids in the circulating
cooling water, it is necessary to keep the suspended solids below a concen-
tration somewhere between 100 and 300 mg/1. When biological activity is
going on, biomass can settle on and with other suspended solids and cause
serious fouling. Thus, the lower limit on suspended solids is recommended.
Demonstrations with a refinery wastewater have shown that about 50%
of the organic carbon and most of the phenolics are removed in the cooling
tower circuit. Thus, it is reasonable to expect that at least half of the
BOD in the makeup will be oxidized in the cooling tower so that the sludge
production rate will be about 0.2 Ib sludge per Ib BOD entering the system.
Water having a BOD of 1,000 mg/1 (as suggested in Table 6-8) has an equiva-
lent sludge concentration of 200 mg/1 and it is apparent that some form of
solids separation is required. Of course the water shown in Table 6-8 has
a high BOD. In many plants the BOD of process condensate will be reduced to
about 50 mg/1 by, for example, an activated sludge treatment. In these
cases the sludge formed by bio-oxidation in the cooling tower, about 10
mg/1, is of little importance and the most important consideration is to
ensure that water leaves the final clarifier of the bio-treatment plant
with a low suspended solids, with filtering if necessary.
If the makeup does contain a high BOD resulting in the formation of a
high concentration of sludge, the cooling tower basin must not be relied
upon as a settling basin. Water is usually distributed over a cooling
tower packing at more than 3 gpm/ft and settling of biomass usually occurs
2
at about 0.5 gpm/ft . The basin of a cooling tower is bigger than the
packed area of the tower, but it is not that much bigger. Mohler and
18
Clere have found it worthwhile to skim the basin.
The recommended scheme is shown in Figure 6-2 and follows the concepts
8,18
of Mohler and Clere . They found that most of the BOD removal (about
90%) occurs in a cooling tower/cooling loop system very quickly, so concen-
trations of 1.1 cycles will suffice. They also found that 110 ppm sus-
pende'd solids in the circulating cooling water had no adverse effects on
heat transfer provided that the heat exchangers were backwashed systematically.
258
-------
Makeup
TOWER No. 1
1.1 CYCLES
FILTER
-IAIN
COOLING TOWER
HEAT EXCHANGERS
Supernatant
Back flush
~l
SETTLING TANK
I
Sludge
Figure 6-2. Control of solids from bio-oxidation in cooling loop.
259
-------
One section of the cooling tower should be isolated. This section
should be connected to heat exchangers which are easy to backwash. It
should receive the full makeup and operate at about 1.1 cycles of concentra-
tion. The heat exchangers supply heat for biological activity and minimize
the investment in cooling towers. The blowdown is filtered. Enough blowdown
is returned to the basin to hold suspended solids to about 100 ppm, which
is required for good filtration and to avoid fouling the heat exchangers.
The rest of the filtrate serves as makeup for the main cooling loop.
The recommended filter is upflow multimedia with the large particles
on the bottom. The filtration rate will be about 5 gpm/ft at pressure
drops probably between 15 and 60 psi with about 85% removal of solids. A
2
loading of about 3 Ib solids/ft can be expected so flushing is required
once every 12 hours.
The recommended flush is also upf low at a slow water rate with a large
injection of compressed air. The a-JT causes the bed to rise and a large
space is needed to prevent loss of sand in the flush water. Free space
above the bed should equal the bed depth. An alternative procedure is a
short air-water scrub followed by 5 to 8 minutes of a high water rate (12
gpai); a second air scrub and fast water flush will usually be required to
complete the cleaning. Compressed «ir injection should be stopped for a
few minutes before the water rate is raised otherwise the rapidly rising
air bubbles will upset the bed. About 4% of the water fed to the filter
will be in the flush. The backflush should be allowed to settle for one or
two hours. The supernatant can be returned to the cooling basin and the
sludge disposed of.
A possible disposal process for the sludge is to dewater it with flue
gas scrubber sludge. The filtrate becomes makeup to the flue gas scrubber.
A filter is suggested rather than a clarifier. A clarifier can handle
only 0.5 to 0.75 gpm/ft and must be much larger than a filter. It is not,
therefore, necessarily cheaper. A clarifier can handle a higher suspended
solids in the feed obviating recycle of clarified water to the number 1
cooling loop. However, it must be demonstrated that a higher suspended
solids does not harm heat transfer. The overflow from a clarifier may not
be as clear as from a filter and if about 10 cycles of concentration are
planned in the main cooling loop, the makeup should have less than 20 ppm
260
-------
suspended solids. A clarifier will need flocculating chemicals. Mohler
and Clere found a low usefulness for chemicals in their filter system.
6.7 Corrosion Control
The use of chromium or much zinc to control corrosion should be
avoided. It is probable that nonmetallic corrosion preventing agents will
be adequate. In most coal conversion plants the surface contacted by
cooling water is not particularly hot, and it is on hot surfaces that
nonmetallic corrosion inhibitors perform least well. Process streams are
usually cooled to below 300°F in an air cooler; the water cooler follows
the air cooler in series. Much of the cooling load may be condensers for
steam turbines and here the temperature difference is not more than about
20°F. The hottest heat exchangers will be interstage coolers on gas compres-
sors and here the use of corrosion resistant metals might be considered.
Special metals will anyway be used for oxygen and hydrogen service and they
might be ronsidered for air compressors. Note that because of the presence
of ammonia, copper alloys should not be used.
All the cooling towers should be operated just at the point where
calcium carbonate precipitates. This allows a thin film of scale to protect
the heat exchangers. Determination of the desirable pH involves tedious
and lengthy calculations and some shortcuts have been given. Kunz, Yen and
Hess have measured the pH and the alkalinity in forty successfully
operating cooling towers and find them related by the curve:
pH = 4.4 + 1.6 Log1Q (Alk) (22)
with 90% confidence limits - 0.75 pH units. In this equation alkalinity is
expressed in mg/1 as CaCO . Calgon Corporation supplies a slide rule for
calculating the amount of acid to be added for desired cycles of concentra-
tion. The curve of Kunz and others is close to the curve recommended by
Calgon.
There has been enough study on nonchromium corrosion control agents
that a recommendation can be made.
Krisher20 has conducted a series of tests and makes certain points.
He successfully used polyphosphates and organic phosphonates. The trade-
off between me re fouling and less corrosion is more prounounced than with
261
-------
chromate inhibitors and very careful control is needed. We recommend
monitoring both scale formation (in one or two selected heat exchangers, or
in a test section) and corrosion (by use of a meter), and adjusting the pH
control regularly as needed.
Krisher points out that phosphorus is a nutrient but since our towers
are intended to be biologically active, this is an advantage. Krisher
suggests 50-100 ppm inhibitor, which will be expensive, but, as far as we
can see, unavoidable.
Kumer and Fairfax have tested several nonchromium corrosion inhibi-
tors and rated them for anticorrosion effectiveness, scale formation and
cost (1971 costs). They prefer an amino-phosphonate + polyacrylamide system.
22
Additional tests have been reported by Breske , whose two preferred non-
metal formulations were (1) a proprietary mix of phosphonate, polymer and
copper inhibitor, and (2). a mixture of aminomethylene phosphonate, hydroxy-
ethylidene diphosphonate, benzotriazole and hexametaphosphate. Benzotriazole
is a copper inhibitor and may not be needed. Because of the presence of
ammonia, copper should be avoided.
Vukasovich and Robitaille studied the use of sodium morybdenate. Their
best formulation was (mg/1 in the circulating water) , 5 mg/1 Na MoO , 5 mg/1 -
hydroxyethylidene-l,l-diphosphonic acid, 5 mg/1 2-mercaptobenzothiazole,
2 mg/1 ZnSO . The MET is a copper inhibitor and, as mentioned, may
not be needed.
There is enough agreement among the authors quoted that we suggest the
following be tried for corrosion inhibition (mg/1 in the circulating water;
the suppliers mentioned are not necessarily the only suppliers)
hydroxyethylidene diphosphonate, 5 mg/1
(available from Monsanto)
sodium molybdate, 5 mg/1
(available from Climax)
If copper inhibition is required,
benzotriazole, 1 mg/1 or mercaptobenzothiazole, 5 mg/1
(available from Sherwin Williams)
If inhibition is inadequate,
zinc sulfate, 2 mg/1 or
sodium gluconate (Ref.23), 10 mg/1
In addition, an aerylamide-aerylate polymer will be added as antiscalant
in an amount that differs for each plant.
262
-------
6.8 Slowdown
Firm recommendations cannot be given for the disposal of cooling tower
blowdown. The chosen method will depend on the local situation and regula-
tions, the local source water quality, and on the coal conversion process
and coal rank as these effect the quantity and quality of condensate fed
to the cooling towers.
Blowdown will contain all the salts that were in the makeup. Traces
of metals derived from coal and not natural to rivers may be present, but
there is no information available. Probably the salts are natural to
rivers and, depending on local preference, could be disposed of in the
river, in the mine mixed with coal ash, or in a pond (lined or unlined).
If concentration is required before disposal, evaporation using vapor
compression evaporators will be used. If a concentrated waste is required,
the cooling towers will be operated at the maximum safe concentration and
the blowdown will be saturated. Concentration procedures intolerant of
precipitation such as reverse osmosis, will, in the usual case, not be
useful. Organic contamination may easily distill over during evaporation,
decreasing the value of the distilled water. Dollar credit for this water,
which is usually used as boiler feed, cannot be taken if an absorption
cleanup step is required.
Blowdown will have some, but very little, contamination added to
control corrosion and biofouling. This is controllable, and harmful or
long-lived chemicals need not be used.
Blowdown will usually contain organic contaminants whose nature is
unknown. We expect that these contaminants will have a high ratio of COD/BOD
because most of the biodegradable material will have degraded. If concentra-
tion or destruction of the organic contamination is required, very general
procedures should be considered until something is known about the nature of the
the contaminants. Wet oxidation is the direct combustion of organic matter
in water by air under pressure and it has a wide applicability. It is very
expensive; the cost is proportional to the throughput and independent of the
concentration so it should only be considered for highly contaminated waters
where treatment is positively essential. Wet oxidation is described in
263
-------
References 3 and 26. Adsorption, as described in this report, should also
be considered if treatment is required. A suitable carbon or synthetic
absorbent must be found. The recovered waste will probably be burnt.
Adsorption has the advantage that the cost may not be strongly dependent on
the effluent concentration so a clean water can be obtained.
264
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REFERENCES Section 6
1. Probstein, R.F. and Gold, H., Water on Synthetic Fuel Production,
Chapter 4, MIT Press, 1978.
2. Gold H. et al, "Water Requirements for Steam-Electric Power Generation
and Synthetic Fuel Plants in the Western United States," EPA report
#600/7-77-037, 1977.
3. Goldstein, D.J. and Gold, H., "Water Conservation and Pollution Control
in Coal Conversion Processes," EPA Report #600/7-77-065, 1977,
N.T.I.S. Catalog PB-269-568/2WE.
4. Gold, H. and Goldstein, D.J., "Water Related Environmental Effects in
Fuel Conversion", Vols. 1 & 2, EPA, Research triangle Park, Report
600/7-78-197a, b, October 1978.
5. Leung, P. and Moore, R.E., "Water Consumption Determination for Steam
Power Plant Cooling Towers: A Heat-and-Mass Balance Method," Paper 69-
69-WA/Pwr-3, American Society of Mechanical Engineers, New York,
Winter Annual Meeting, November 1969.
6. Leung, P., "Evaporative and Dry-Type Cooling Towers and Their Applica-
tion to Utility Systems," Water Management by the Electric Power Industry,
(E.F. Gloyna, H.H. Woodson, and H.R. Drew, eds.), 106-116, Center for
Research in Water Resources, the University of Texas at Austin, 1975.
7. Grutsch, J.F. and Griffin, R. W., "Water Reuse Studies by the Petroleum
Industry," p_aper 5e delivered at 85th National Meeting of A.I.Ch.E.,
Philadelphia, June 4-8, 1978.
8. Mohler, E.F., Jr., and Clere, L.T., "Bio-oxidation Process Saves HO,
Hydrocarbon Processing, 84-88, October 1973; also, same authors,
"Development of Extensive Water Reuse and Bio-oxidation in a Large Oil
Refinery," delivered at the National Conference on Complete Water
Reuse, Washington, D.C,, 1973.
9. Hart, J.A., "Wastewater Recycle or Reuse in Refinery Cooling Towers,"
The Oil and Gas Journal, 92-96, June 11, 1973.
10. Maguire, W. F., "Reuse of Sour Water Stripper Bottoms," Hydrocarbon
Processing, 151-152, September 1975.
11. Donahue, J.M. and Nathan, C.C., "Unusual Problems in Cooling Water
Treatment," Chem. Engineering Progress T±, (No. 7), 88-89, July 1975.
12. Sinor, J.E., ed., Cameron Engineers, "Evaluation of Background Data
Relating to New Source Performance Standards for Lurgi Gasification,"
EPA Report 600/7-77-057, Research Triangle Park, June 1977.
13. Beychok, J'.R. , "Coal Gasification and the Phenosolvan Process,"
Symposiuir on Processing of Phenolic Aqueous Wastes, 85-89, American
Chemical Society, Div. of Fuel Chemistry preprints, Vol. 19, No. 5,
September 1974.
265
-------
14. Goldstein, D.J., Hicks, R.E. and Liang, L. , "Conceptual Designs for
Water Treatment in Demonstration Plants," ongoing work under DOE
contract EF-77-C-01-2635 by Water Purification Associates.
15. McCoy, J. W. , The Chemical Treatment of_ Cooling Water, Chemical
Publishing Company, New York, 1974.
16. Truesdall, A. H., and Jones, B. F. , "WATEQ, A Computer Program for
Calculating Chemical Equilibria of Natural Waters," J. Research U.S.
Geological Survey 2 (No. 2), 233-48, March-April 1974.
17. Chen, Y. S., Petrillo, J. L. , and Kaylor, F. B., "Optimal Water Reuse
in Recirculating Cooling Water Systems for Steam Electric-Generating
Stations," 528-541 in Proceedings of_ the Second National Conference
on Complete Water Reuse, AIChE, 1975.
18. Mohler, E.F. and Clere, L.T. , "Removing Colloid Solids via Upflow
Filtration," Chem. Eng. Progress, 74-82, April 1977.
19. Kunz, R.G., Yen, A.F. and Hess, T.C., "Cooling Water Calculations",
Chem. Eng. 84(16), 61-71, August 1, 1977. Also same authors, "Basic
Cooling Water Calculations, General Algorithm for Cooling Water
Chemistry," paper 41b presented at 85th National AIChE Meeting,
Philadelphia, PA, June 1978.
20. Krisher, A.S., "low Toxicity Cooling Water Inhibitors - How-They Stack
Dp," Chemical Engineering, 115-116, February 13, 1978.
21. Kumer, J. and Fairfax, J.P., "Rating Alternatives to Chromates in
Cooling Water Treatment." Chemical Engineering, 111-112, April 26,
1976.
22. Breske, T.C., "Testing and Field Experience with Non-heavy Metal
Corrosion Inhibitors," Materials Performance, 17-24, February 1977.
23. Vukasovich, M.S. and Robitaille, D.R., "Corrosion Inhibition by Sodium
Molybdate," Proceedings of the Climax Second International Conference
on the Chemistry and Uses of Molybdenum, Oxford, England, September
1976; distributed by Climax Molybdenum Company.
24. Hoppe, T.C., "Secondary Effluent without Phosphate Removal Used for
Cooling Water Makeup," Water and Sewage Works, 62-65, February 1976.
25. DeBoice, J.N. and Thomas, J.F., "Chemical Treatment for Phosphate
Control," J. Water Pollution Control Federation - 47_(9) , 2246-2255,
September 1975.
26. Water Purification Associates and Process Research, Inc., "Innovative
Technologies for Water Pollution Control," National Commission on
Water quality, Report 75/13, December 1975, N.T.I.S. Catalog PB 247 390.
266
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7. EQUIPMENT SIZING AND COSTING
7.1 Introduction
The equations for making preliminary size and cost estimates for the
equipment required for the water treatment procedures described in the
preceding sections are presented in this section in handbook style. The
costing and sizing procedures described below are to be used to compare
alternative water treatment schemes, and are not to be used for detailed
design purposes.
7.2 Storage or Surge Vessels
The optimum length/diameter (L/D) ratio for a vessel of a given volume
is dependent on the design pressure, the maximum allowable working stress
of vessel shell material, the joint efficiency, and the corrosion allowance.
For carbon steel vessels with design pressures under 200 psig, the savings
realized by using L/D ratios greater than 2:1 are negligible. However, use
of alloy steel shell materials at high design pressures strongly favors L/D
ratios of 5:1 or greater if plot space permits.
7.2.1 Surge Vessels
The amount of liquid holding capacity required by a given surge vessel
is usually determined by the amount of time required by the plant operator
to intervene to prevent a serious upset in the downstream processing units.
Liquid holdup times of 3 to 5 minutes are standard practice.
7.2.2 Reflux Drum
The volume required for a reflux drum is governed by both vapor and
liquid handling capacity. In most circumstances the liquid handling
capacity, described in Section 7.2.1, will dictate the vessel volume. For
\
systems with partial condensers; the vapor handling capacity must be
checked to determine if the cross-sectional area available for vapor flow
is such that the vapor velocity (V) will not exceed-:
0.5
V = 0.15
r°i - D i
_i v
D
L v J
ft/sec. (1)
The available cross-sectional area is defined as the full cross-section for
vertical vessels, or the transverse cross-section between the maximum
liquid level and the top of the vessel for horizontal vessels.
267
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7.3 Heat Exchanger Sizing
The required heat transfer surface area (A) is a function of the quantity
of heat to be transferred (Q), the log-mean temperature difference (AT. ) and
the overall heat transfer coefficient (D ). Equation 2 states the relation-
o
ship between the above parameters:
A- ——, ft" (2)
The heat transfer surface area can be estimated from using the
following procedure: first, the heat duty (Q) and the log-mean temperature
difference (AT, ) are determined from the system mass and energy balances.
1m
Then, an overall heat transfer coefficient U is selected from tables such
o
as those given in Reference 1. The estimated heat transfer area is then
calculated from Equation (2) .
7.4 Distillation Tower Sizing
The size of a distillation column is determined by the allowable vapor
and liquid flow velocities, and the number of stages necessary to accomplish
the required mass transfer. The column diameter is a function of the vapor
and liquids loads. The vapor load for a given stage is defined as:
load = CFS I D^/XDj^ - DV>
0.5
(3)
where:
CFS = vapor rate, ft /sec
3
D = vapor density, Ib/ft"
Dj^ - liquid density, Ib/ft3
The liquid load is simply the liquid rate leaving stage n in gallons per
minutes (GPM). Once the irapor and liquid loads have been determined/ using
methods described in previous sections of this manual, the column diameter can
be estimated from Equation (4):
268
-------
D =
- 0.5
V load (A + B-GPM) + C-GPM
"J
ft
where:
B
Tray towers
Packed towers
3.5
3.8
4 x 10 4
2 x 10~4
3 x 10 2
7 x 10~3
4.5
0.9
(4)
7.5 Cost Data
The cost data used in preparing the cost estimates in this manual are
presented in this section in equation form:
n
Cost = k x (size parameter) , $
(5)
Table 7-1 lists the parameters for Equation (5).
After the major equipment cost has been determined, the total installed
cost is estimated by multiplying the equipment cost by a Lang factor of
3.75 as shown below in Equation (6).
Total installed cost = 3.75 (major equipment cost), $ (6)
269
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TABLE 7-1. EQUIPMENT COST PARAMETERS
10
~j
o
Equipment
Distillation towers
Surge and storage vessels
Heat exchangers, shell
and tube
Heat exchangers, reboilers
Material
Size Parameter
c.s
ss
c.s.
ss
cs
ss
cs
995
2786
73
204
44
350
292
weight, Ib
weight , Ib
weight, Ib
weight, Ib
2
area, ft
2
area, ft
2
area, ft
0.44
0.44
0.65
0.65
0.84
0.69
0.65
-------
REFERENCES - Section 7, AND ADDITIONAL BIBLIOGRAPHY
1. "Platecoil for Heat Transfer," Catalog No. 5-63, Tranter, Inc., Wichita
Falls, Texas.
2. "Ballast Tray Design Manual," Bulletin No. 4900-3rd ed. , Glitsch, Inc.,
Dallas, Texas.
3. "Koch Flexitray Design Manual," Bulletin 960, Koch Engineering Co., Inc.,
Wichita, Kansas.
4. "Packed Towers," Norton Chemical Process Products, Akron, Ohio.
5. "Cost Engineers' Notebook," American Association of Cost Engineers,
Morgantown, West Virginia.
271
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TECHNICAL REPORT DATA
(Please read Instructions on the reverse before completing)
1. REPORT NO.
EPA-600/7-79-133
3. RECIPIENT'S ACCESSION NO.
4. TITLE AND SUBTITLE
Wastewater Treatment in Coal Conversion
5. REPORT DATE
June 1979
6. PERFORMING ORGANIZATION CODE
7. AUTHOR(S)
R.E.Hicks, D.J.Goldstein, F.B.Seufert, and
I.W.Wei
8. PERFORMING ORGANIZATION REPORT NO.
9. PERFORMING ORGANIZATION NAME AND ADDRESS
Water Purification Associates
238 Main Street
Cambridge, Massachusetts 02142
10. PROGRAM ELEMENT NO.
E HE 62 3 A
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
Final; 10/76 - 1/79
14. SPONSORING AGENCY CODE
EPA/600/13
15. SUPPLEMENTARY NOTES IERL-RTP project officer is William J. Rhodes, Mail Drop 61,
919/541-2851.
16. ABSTRACT
The report describes water treatment control technology specific to fuel
conversion plant sites in the western U.S. Most plants converting coal to other fuels
use a large quantity of clean water (as steam) and put out a large quantity of dirty
water that is condensed when the products from the coal reactor are cooled. Treat-
ment of this foul condensate is the subject of this report. The report discusses each
aspect of water treatment separately. Procedures for removing phenolic compounds
are discussed: they include distillation, extraction, and adsorption. The report in-
cludes design equations, step-by-step design procedures, and calculations for a
typical unit. It also provides physical data that are required for design. The discus-
sion of ammonia separation and recovery includes design equations and physical data
Illustrative calculations show how the design procedure is used. For biological
treatment, the design procedures show how to destroy organic contamination (inclu-
ding phenol) in the condensate. It also discusses cooling tower control. An econom-
ical use of the foul condensate is treating it for makeup to a plant's cooling system.
The report also reviews established procedures for sizing plant equipment.
17.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b.IDENTIFIERS/OPEN ENDED TERMS
COS AT I Field/Group
Pollution
Coal Gasification
Waste Water
Water Treatment
Phenols
Ammonia
Bioassay
Cooling Towers
Organic Compounds
Pollution Control
Stationary Sources
Coal Conversion
Biological Treatment
Organic Contamination
13B
13H
07C
07B
06A
13A,07A
8. DISTRIBUTION STATEMENT
Release to Public
19. SECURITY CLASS (ThisReport)
Unclassified
21. NO. OF PAGES
284
20. SECURITY CLASS (Thispage)
Unclassified
22. PRICE
EPA Form 2220-1 (9-73)
272
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