United States Industrial Environmental Research EPA-600/7-78-186c
Environmental Protection Laboratory September 1978
Agency Research Triangle Park NC 27711
Environmental Assessment
Data Base for High-Btu
Gasification Technology:
Volume III.
Appendices D, E, and F
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.
REVIEW NOTICE
This report has been reviewed by the participating Federal Agencies, and approved
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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-600/7-78-186C
September 1978
Environmental Assessment Data
Base for High-Btu
Gasification Technology:
Volume III. Appendices D, E, and F
by
M. Ghassemi, K. Crawford, and S. Quinlivan
TRW Environmental Engineering Division
One Space Park
Redondo Beach, California 90278
Contract No. 68-02-2635
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
-------
CONTENTS
Page
APPENDIX D - AIR POLLUTION CONTROL D-l
Hydrogen Sulfide Control Module
Claus Process D-2
Stretford Process (See Acid Gas Removal Module,
Appendix B)
Giammarco-Vetrocoke Process (See Acid Gas Removal
Module, Appendix B)
Tail Gas Treatment Module
SCOT Process D-13
Beavon Process D-20
IFP Process D-30
Sulfreen Process D-40
Cleanair Process D-46
Sulfur Oxides Control Module
Wellman-Lord Process D-51
Chiyoda Thoroughbred 101 Process D-61
Shell Copper Oxide Process D-70
Lime-Limestone Slurry Scrubbing Process D-77
Double Alkali Process D-92
Magnesium Oxide Scrubbing Process D-108
Particulate Control Module
Fabric Filtration Process D-l23
Electrostatic Precipitation Process D-l30
Venturi Scrubbing Process D-l36
Cyclones D-142
* • •
in
-------
CONTENTS (Continued)
Hydrocarbon and Carbon Monoxide Control Module
Thermal Oxidation Process .... D-148
Catalytic Oxidation Process D~154
Activated Carbon Adsorption Process (See
Methanation Guard Module, Appendix B)
APPENDIX E - WATER POLLUTION CONTROL E-l
Oil and Suspended Solids Removal Module
Gravity Separation Process (API Separators) E-2
Flotation Process E-10
Filtration Process E-18
Coagulation-Flocculation Process E-24
Dissolved Gases Removal Module
Steam Stripping Process E-36
USS Phosam W Process E-45
Chtvron WWT Process E-52
Dissolved/Particulate Organics Removal Module
Biological Oxidation Process E-60
Evaporation/Retention Pond Process E-77
Chemical Oxidation Process E-80
Ph®ws@lvan Process E-93
Activated Carbon Adsorption Process E-l00
Sludge Treatment Module
Gravity Thickening Process E-l18
Centrifugation Process E-l23
Vacuum Filtration Process . . E-132
Drying Beds E_140
Emulsion Breaking Process E-145
APPENDIX F - SOLID WASTE MANAGEMENT c i
••••«... r — I
Incineration Process ..... c 9
r—c.
Land Disposal Process p_g
Chemical Fixation/Encapsulation Process p_18
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APPENDIX D
AIR POLLUTION CONTROL
Hydrogen Sulfide Control Module
Claus
Stretford (see Acid Gas Removal Module, Appendix B)
Giammarco-Vetrocoke (see Acid Gas Removal Module, Appendix B)
D-l
-------
CLAUS PROCESS
1.0 General Information
1.1 Operating Principles - The catalytic oxidation of H2S, in an acid
gas stream, to elemental sulfur and the recovery of the sulfur. The
catalyst used is either bauxite or alumina in the form of pellets or
balls.
1.2 Development Status - Commercially available.
1.3 Licensor/Developer - The Ralph M. Parsons Co.
100 W. Walnut Street
Pasadena, CA 91124
1.4 Commercial Applications'1^ - There are approximately 170 Claus
plants in the United States used in a wide variety of industries
C\} *
including natural gas and coke production^ . One application of
the Claus process to coal conversion gas purification is in South
Africa^12).
2.0 Process Information^2'12^
2.1 Flow Diagram - There are three basic forms of the Claus Process:
"split-stream," "straight-through," and the "sulfur burning" mode.
The "split-stream" process is used when the C02 concentration ex-
ceeds 30% (volume); the "straight-through" process is generally used
when the feed gas stream contains less than 30% (volume) C0?. The
"sulfur burning" mode is employed where low H2S levels (555-10%) are
to be treated.
In most coal conversion processes the acid gas stream produced as a
result of acid gas treatment will contain CO- in excess of 30%
*For specific information on plant locations; cnifu*. nvn,i,,~+4
atlng'daus pUnts and cables wjlc^eslgn5^^ pUnt'see
pUntsee
D-2
-------
(volume); therefore, the "split-stream" Claus process would be
applicable.* In the "split-stream" process (Figure D-l) the acid
gas, Stream 1, enters the system through a knockout vessel (where
entrained liquids are removed) and is then split into two streams
(4 and 5). Stream 5 enters a sulfur burner where the H2S is oxi-
dized to SCL using a stoichiometricquantity of air. Hot gases enter
a reaction furnace; enough residence time is provided for the Claus
reaction to reach equilibrium. The gas is then passed through a
waste heat boiler and a condenser (where the elemental sulfur pro-
duced is removed) and then it is combined with Stream 4. The com-
bined stream is then reheated and sent to the first catalytic con-
verter for further Claus reaction. A plant may operate with any
number of catalytic converters, depending on the desired sulfur
recovery efficienciy.
The "straight-through" system (Figure D-2) is similar to the above
v
system with the following exceptions: upon existing the knock-out
drum, the entire volume of gas is sent to a sulfur burner where it
is oxidized under free-flame conditions with a stoichiometric quan-
tity of air; it then passes through the reaction furnace, waste heat
boiler, first condenser, reheater, and converter.
The "sulfur burning" mode is similar to the "split stream" mode
except that liquid sulfur is injected into the combustion chamber
to supply SCL for the Claus reaction.
2.2 Equipment - Reaction furnace, sulfur condensers, reheaters, cata-
lytic converters, waste heat boilers.
2.3 Feed Stream Requirements - Claus plants can be designed to operate
at various temperatures and pressures, and with a wide variation of
(3)
feed stream compositions/ '
'•'The 30% maximum CC>2 level for straight-through operation can be extended by
the use of preheat, Hydrocarbons in feed also influence straight-through
applicability, since they may limit the bypassing of gas directly to the
converters.
D-3
-------
mi 5 -T
jr «i "i
x*x
1
§ "I
s *
? T
6J
If LEGEND:
r
1. ACID GAS FEED
2. COMBUSTION AIR
3. TAIL GAS
J
J
(CONVERTER : 1 J
k
— REACTION
FURNACE 8 REHEAT : 1
—
-------
a
SULFUR BURNER
LEGEND:
1. ACID GAS FEED
2. COMBUSTION AIR
3. TAIL GAS
4. CONDENSATE
5. BOILER FEED WATER
6. LOW PRESSURE STEAM
7. HIGH PRESSURE STEAM
8. LIQUID SULFUR
9, SPENT CATALYST
(LOCATION NOT KNOWN)
CONVERTER : 1
w
L
CONDENSER NO. 1
CONVERTER:n
CONDENSER NO. h
AIR BLOWER
CONDENSER NO.n + 1
SULFUR PIT
NOTE: SEE TABLE D-2 FOR COMPOSITION OF STREAMS 1 AND 3.
Figure D-2. Straight-Through Claus Process
(3)
-------
concentration is the most important parameter in Clans
design and operation. Gases with I^S concentrations from less
than 10 vol. % to greater than 90 vol. % can be handled by vari-
ous Claus plant designs(4J2).
• "Standard" conditions have been defined as the following^ ':
- H2S content 90% by volume
- Hydrocarbon content 2% by volume as ethane
- Temperature: 311 °K (100°F)
- Pressure: 0.14 MPa (6 psig)
2.4 Operating Parameters - Operating temperatures will vary as a func-
tion of feed stream conditions and plant design. Pressures are
usually low (below 0.17 MPa or 25 psia).
2.5 Process Efficiency and Reliability - For operating conditions defined
as "standard" in Section 2.3, a "typical" plant (3-stage Claus) is
capable of 97% sulfur recovery^ '. Efficiency decreases as the cata-
(12)
lyst becomes partially deactivated^ .
No information available which would indicate special maintenance
problems or unusual hazardous conditions created by the process.
Principal problems result from frequent shutdown periods from lack
of feed or from upsets caused by operating problems in upstream
^
units
2.6 Raw Material Requirements
• Catalyst makeup: half life is at least two to three years
(12)
2.7 Utility Requirements^ ' - Utility requirements will vary. For gas
stream containing 40% H2$ and 60% C02, typical requirements are as
fol1ows:
• Boiler feed water: 6.25 I/kg of sulfur (0,75 gal/lb)
• Electricity: 0.088 kwh/kg of sulfur (0.05 kwh/lb)
3.0 Process Advantages
• Commercially proven process for bulk H9S removal; process is
known and used extensively L
t Produces high purity salable sulfur
D-6
-------
• Produces steam
• Process design can be readily altered to accommodate a wide range of
feed gas conditions (5). Such process modifications may consist of use
of a multizone combustion chamber and control of flow rate, tempera-
ture and combustion air.
4.0 Process Limitations
t The carryover of high molecular weight hydrocarbons can cause deacti-
vation of the catalyst because such compounds can adsorb on the cata-
lyst and eventually chart6).
• Low molecular weight hydrocarbons in feed can cause increased furnace
temperatures and dilution of reactive sulfur compounds which decrease
conversion efficiency. Carbon oxides formed by hydrocarbon combustion
can increase COS and CS^ formation in the Claus furnace.
• Catalyst plugging problems can occur when NH3 concentration exceeds
500 ppmv in combination with C02 concentrations greater than 30% „
(vol)(3). If C02 is low in feed gas, higher levels of ammonia can,
by design modifications, be handled (up to 18%) (12).
• Excessive hydrocarbons in feed can lead to elevated operating tempera-
tures which can cause accelerated aging of the catalyst^6).
• The presence of HCN in the acid gas can lead to excessive equipment
corrosion and catalyst deactivation via formation of thiocyanates(' ) .
t The presence of various contaminants in the acid gas feed (e.g., NH3,
HzO, C02> hydrocarbons) can lower the sulfur removal ability of the
Claus process and increase the size of the plant required due to
larger volumetric flow rates(°).
t COS and C$2, if present in the feed, are not usually converted to
and then to elemental sulfur in Claus plants using standard catalysts.
Some COS and C$2 are actually formed in Claus plants when feeds high
in CO and C02 are processed.
5.0 Process Economics^ '
• The cost of a Claus plant varies as a function of two major parameters:
the percent of H2S in the acid gas feed; and the daily capacity of
sulfur production.*
*If ammonia containing acid gas is burned, the amount becomes a factor in
plant size and cost.
D-7
-------
• The approximate costs as estimated in 1973 are as follows:
Mole % H2S in Claus Plant Investment Sulfur Production
Acid Gas Feed (102 tonne/day plant size) Cost per Tonne
15 $1,400,000 $14
50 $1,000,000 $11
90 $ 900,000 $ 9
Daily Sulfur Claus Plant Investment
Production Capacity (assumes H2S concentration Sulfur Production
(Tonne) in feed at 50%) Cost per Tonne
10 $ 300,000 $26
102 $1,000,000 $11
1020 $4,300,000 $ 8
6.0 Input Streams (see Figure 1)
• Acid gas stream: (Stream 1); see Tables D-l and D-2
7.0 Discharge Streams (see Figure D-l)
« Tail-Gas: (Stream 3); see Tables D-l and D-2
• Condensate (Stream 6): no data available
• Spent Catalyst (Stream 11): no composition/properties data available!
see Section 2.6 for makeup requirements.
8.0 Data Gaps and Limitations
• Process applicability to coal conversion process gas purification ™<
terns not entirely established. purification sys-
• Definition of the maximum allowable concentrations nf v»y.in ? ± •
nants in the feed gas; e a NH rns re J various contarm-
aceous matter. 9" 3' cos> CS2, trace metals, HCN, carbon-
9 The effects that various contaminants (trace mPtaic «^
matter, etc. have on the process and thl ?*• * ' carbonaceous
taminants in the system the ultlmate fate of such con-
9.0 Related Programs
' ' tUred ^ ^ deSl9" °f the ««•» Pilot punt at
rrr: - -
D-8
-------
TABLE D-l. SPLIT FLOW MODE CLAUS FEED AND TAIL GAS DATA(9)*
Component
COS
H2S
so2
co2
N2
Cl
C2
C3
C4
Temperature
Pressure
Flow (wet basis)
Feed Stream
Stream 1
Mole %
19.72
78.68
0.56
0.66
0.12
0.08
0.18
313°K (105°F)
0.16 MPa
(8.1 psig)
518000 Nm3/d
(19,272 mcf/d)
Tail Gas Stream
Stream 3
Mole %
0.09
0.26
0.10
65.04
34.34
0.19
0.03
;;;;
805°K (990°F)
0.10 MPa
(0.1 psig)
1,073,000 Nm3/d
(37,920 mcf/d
*Datawere selected to represent Claus performance on low H2S, high C02 gases
which would be encountered in coal gasification applications.
D-9
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TABLE D-2. STRAIGHT THROUGH CLAUS FEED AND TAIL GAS DATA
(8)
Acid Gas Feed
Stream 1
Components Mole %
H2S
co2
wll n
t\ -
C2H6.
NH3
H20
>•
90.1
3.6
0.8
0.4
0
5.1
Tail Gas*
Stream 3
Composition Mole %
N2;
co2
H20
H2S
so2
S6 + S8
Entrained Liquid
S°
62
1.4
35
0.30
0.43
0.02
0.13
*Data given are after incineration.
D-1Q
-------
REFERENCES
1, Beers9 W. D., Characterization of Claus Plant Emissions. NTIS9 PB 220-376
for the U.S. EPA9 April 1973.
2. Meisen, A., Bennett, H. A.9 Consider All Claus Reactions, Hydrocarbon
Processing., November 1974.
3. Chute, A. E.9 Tailor Sulfur Plants to Unusual Conditions. Hydrocarbon
Processing, April 1977.
4. Dravo Corporation, Handbook of Gasifiers and Gas Treatment Systems.
ERDA FE-1772-11, Washington, D.C., February 1976.
5. Maddox, R. N., Gas and Liquid Sweetening. Campbell Petroleum Series, 1974.
6. Pearson, M. J., Developments in Claus Catalysts, Hydrocarbon Processing,
February 1973.
7. Homberg, 0. A., Singleton, A. H., Performance and Problems of Claus Plant
Operation on Coke Oven Acid Gases. Journal of the Air Pollution Control
Association, Volume 25, No. 4, April 1975.
8. Goar, B. G., Impure Feeds Cause Claus Plant Problems. Hydrocarbon Proc-
cessing, July 1974.
9. Draft Standards Support and Environmental Impact Statement - Volume I:
Proposed Standards of Performance for Lurgi Coal Gasification Plants, U.S.
EPA November 1976.
10. Norman, W. S., There Are Ways to Smoother Operation of Sulfur Plants. The
Oil and Gas Journal, 15 November 1976.
11. Raymont, M. E. D.9 Role of Hydrogen in Claus Plants. Hydrocarbon Proc-
cessing, May 1975.
12. Information provided to TRW by C« L. Black of Ralph M. Parsons Co.,
June 20, 1978.
13. Information provided to TRW by the Institute of Gas Technology, May 1978.
D-ll
-------
Tail Gas Treatment Module
SCOT
Beavon
IFF
Sulfreen
Cleanair
D-12
-------
SCOT (SHELL CLAUS OFF-GAS TREATMENT) PROCESS
1,0 General Information
(1,2)
1.1 Operating Principles - The purification of Claus plant tail gas by
the catalytic reduction of sulfur species to H2S followed by the
removal and recovery of the H?S in an alkanolamine scrubbing system.
A reducing gas (e,g, hydrogen) is used as the reductant and cobalt/
molybdate catalyst is used),
1,2 Development Status - Commercially available.
1.3 Licensor/Developer - Shell Development Company
One Shell Plaza
P. 0, Box 2463
Houston, Texas 77001
1.4 Commercial Applications^ ' - Primary application is for Claus plant
tail gas treatment; fourteen plants are licensed and operating (see
Figure D-3), and approximately 20 others are in various stages of
planning, design, and construction. No known application to coal
conversion type processes have been reported.
2.0 Process Information
2.1 Flow Diagranr * ' (see Figure D-3) - Claus plant tail gas. Stream 1,
is heated, then sent to a catalytic reactor where the sulfur species
are converted to H?S, This H?S stream. Stream 35 is then cooled and
sent to an alkanolamine gas treating system typically containing
diisopropanolamine. The rich amine solution, Stream 5, is sent to
an amine regeneration unit, and cleaned gas is sent to the Claus
plant incinerator.
2.2 Equipment - Conventional catalytic reactor, cooler, absorber, and
stripper.
D-13
-------
13
1
hr
t
—fi
A
o
I- DC
< ILJ
J
HEATER
LEGEND:
1. GLAUS TAIL GAS
2. SCOT TAIL GAS
3. H2S-RICH GAS (HOT)
4. H2S-RICH GAS (COOL)
5. RICH AMINE SOLUTION
6. COOLING TOWER WATER
7. REDUCING GAS (H2)
8. LEAN AMINE SOLUTION
9. FUEL GAS
10. AIR
11. L. P. STREAM
12. COOLING AIR OR WATER
13. FLUE GAS
14. CONDENSATE
CC
LLJ
m
cr
CO
f""RICH~""J
AMINE |
( j
I SOUR WATER |
"J STRIPPER I
SOUR WATER STRIPPER AND RICH AMINE
REGENERATOR ARE SYSTEMS ALREADY
EXISTING IN REFINERY; HENCE, FOR THE
SAKE OF SIMPLICITY, THEY ARE NOT
SHOWN HERE.
Figure D-3. SCOT Process
-------
2.3 Feed Stream Requirements
Temperature: 400°K - 430°K (260°F - 320°F)
Pressure: 0.13 MPa (19 psia)
2.4 Operating Parameters
2.4.1 Catalytic Reactor
Temperature: 573°K (572°F)^3^
Pressure: -0.13 MPa (-19 psia)
2.4.2 Gas Treatment, Amine Absorber
Temperature: 310°K - 320°K (100°F - 120°F)
Pressure: approximately atmospheric
(3 6^
2.5 Process Efficiency and Reliability^ ' ' - In situations where the
Claus tail gas sulfur content is about 9000 ppm (as $02), typical of
a Claus unit with 94% sulfur recovery, the SCOT system can reduce
the sulfur level in the gas to less than 250 ppm (as S02).
Maintenance is reportedly low, stream factor high.
2.6 Raw Material Requirements^ '
Catalyst: coablt molybdate based, three or more years lifetime
Diisopropanolamine: replacement for mechanical losses only
2.7 Utility Requirements (for a 100 tonne/day Claus plant^ ')
• Electricity: 140 kwh/hr
• Fuel Gas: 1.224 Nm3/min (45,600 scfm), based on 9000 kcal/m3
(1012 Btu/ft3)
§
Cooling Water: (6.7°C, 129F rise): 82 I/sec (1300 gpm)
• Steam (3.4 atm, sat): 1,162 kg/hr (2560 Ib/hr) net. Steam is
produced in the catalytic reactor (2,588 kg/hr) and consumed in
the amine regenerator (3,750 kg/hr).
(3 4}
3.0 Process Advantages^ * '
t Utilizes standard sulfur recovery equipment.
• Easily adapted to existing Claus plants.
D-15
-------
• Process produces some of its steam requirements.
t Process can adapt to variations in feed stream composition.
t Can be integrated with bulk acid gas removal unit (e.g., AOIP for
Claus feed upgrading and tail gas cleanup).
4.0 Process Limitations
^1 '2'
• Requires some type of fuel gas to supply heat and a reducing gas for
the catalytic reaction.
• The SCOT system is utilized for the treatment of Claus plant tail gas;
hence, if sulfur recovery is conducted by means other than the Claus
process, SCOT system may be an inappropriate choice for tail gas
treatment.
• Like other catalytic processes, the efficiency of conversion of COS
and C$2 to ^S is decreased when high levels of C02 are present in
Claus plant tail gas.
5.0 Process Economics^ '
For capital and operating costs (1972 dollars) for the various sized SCOT
units, see Table D-3.
6.0 Input Stream
• Feed gas stream, Claus Tail Gas, Stream 1, see Table D-4.
• Hydrogen stream, Stream 7: 9.5 kg/hr (21 Ib/hr) pure hydrogen re-
quired for 100 tonne/day Claus plant(5J.
• Fuel gas, Stream 9: 1,224 Nm3/min (45,660 scfm), based on 9000 kcal/
m3 (1012 Btu/ft3) for 100 tonne/day Claus plant($).
t Catalyst makeup: typically three or more years lifetime.
• Amine makeup: depends primarily on mechanical losses,
7.0 Discharge Streams
• Tail gas from process, Stream 2, see Table D-4.
• Condensate, Stream 14: Slightly acidic, HoS and
- °-44 - °-63 1/sec (7-i§
• Spent catalyst: ?
D-16
-------
TABLE D-3. CAPITAL AND OPERATING COSTS FOR VARIOUS SIZED SCOT UNITS (IN 1972 DOLLARS)1
I
->4
Total capital .investment,
US $ x 106T
Operating costs, $/stream day
(333 stream days/annum):
Direct costs
Capital charge (17% on
equipment capital)
Totals
Add-On SCOT Unit
Capacity of Claus unit,
100 200 1,000
0.9 1.6 3.6
270 460 1,880
450 770 1,680
720 1,230 3,560
Integrated SCOT Unit5
ton of S
100
0.7
270
370
640
intake/sd
200 1,000
1.2 2.8
460 1,880
570 1,280
1,030 3,160
*The capital investment for the add-on SCOT unit corresponds to about 100% of the
capital investment of the preceding Claus unit. For the integrated SCOT unit it
is about 75%.
^Basis: West Europe; for the USA these figures should be increased by 10%.
Add-on: SCOT unit with gas blower, separate alkanolamine regeneration facilities
and separate sour water stripper.
Integrated: SCOT unit fully integrated but bearing a share of the costs for
combined amine regeneration facilities and sour water stripper. There is no gas
blower but the costs of pressure increase in upstream units has been added.
-------
TABLE D-4. TYPICAL GAS STREAM COMPOSITIOH FOR SCOT PROCESS^1'
Components
H9S
so2
S~ vapor and mist
COS
cs2
CO
co2
H20
N,
H2 :
Temperature
Pressure
Claus Tail-Gas 1
to SCOT
vol %
0.85
0.42
0.05
0.05
5(JUI lei i I-UQO
to Atmosphere
vol %
0.03
-— --
10 ppm
0.04 1 Ppm
0.22 3.05
2.37
33.10
61.30
1.60
413°K (284°F)
0.15 MPa (22 psi)
<0.3
7.00
88.96
0.96
313°K (105°F)
0.1 MPa (14.7 psi)
D-18
-------
8.0 Data Gaps and Limitations
• No information is available which would indicate applicability to coal
conversion processes (e.g., the performance with high C02 levels in
feed).
• The effect that various contaminants (NH3, carbonaceous matter, trace
metals, etc.) have on the process, and the ultimate fate of such con-
taminants in the system are unknown.
9.0 Related Programs
No information available.
REFERENCES
1. Gas Processing Handbook, Hydrocarbon Processing, April 1975.
2. Maddox, R. N., Gas and Liquid Sweetening, Campbell Petroleum Series, 1974.
3. Naber, J. E., J. A. Wesselingh, et al, New Shell Process Treats Claus Off-
Gas, Chemical Engineering Progress, December 1973.
4. Beers, W. D., Characterization of Claus Plant Emissions, USEPA, NTIS No.
PB-220 376, April 1973.
5. Battelle Columbus Laboratories, Characterization of Sulfur Recovery from
Refinery Fuel Gas, USEPA, NTIS No. PB-239-777, June 1974.
6. Information provided to TRW by J. M. Duncan of Shell Development Company,
December 8, 1977.
D-19
-------
BEAVON PROCESS
1.0 General Information
1.1 Operating Principle^1' - The purification of sulfur plant tail gases
by the catalytic conversion of sulfur species to H2S followed by
recovery of the H?S as elemental sulfur in a Stretford unit. Fuel
gas is used to supply heat and to produce a reducing gas for the
catalytic reduction; cobalt molybdate is the catalyst employed.
1.2 Development Status - Commercially available.
1.3 Licensor/Developer - Union Oil Company
P. 0. Box 218
Brea, California 92621
1.4 Commercial Applications - Approximately 30 Beavon units are in opera-
tion. They are used primarily for Claus unit tail gas treatment^
and are presented in Table D-5.
2.0 Process Information
(2)
2.1 Flow Diagram (see Figure D-4)v - Tail gas from the sulfur plant,
Stream 1, is mixed and combustion products and fed to a reactor con-
taining cobalt molybdate catalyst. In the reactor sulfur species
are converted to H2S. The H2S rich gas, Stream 5, flows to a con-
denser where it is cooled and then sent to a Stretford unit for con-
version of H2S to sulfur^3'.
(1 2 3)
2.2 Equipment ' ' - Conventional burner, catalytic reactor, coolers,
absorber, oxidation tank, surge tank.
2.3 Feed Stream Requirements
Temperature: typically 390°K-420°K (250°F-300°F)
Pressure: 0.116-0.122 MPa (17-18 psia)
D-20
-------
TABLE D-5. BEAVON SULFUR REMOVAL PROCESS (BSRP) COMMERCIAL INSTALLATIONS
(7)
Customer and Location
Atlantic Richfield Company
Philadelphia, Pennsylvania
Cities Service Oil Company
Lake Charles, Louisiana
The Dow Chemical Company
Freeport, Texas
Exxon Company, U.S.A.
Baton Rouge, Louisiana
Exxon Company, U.S.A.
Bay town, Texas
Exxon Company, U.S.A.
Bay town, Texas
Exxon Company, U.S.A.
Bayway, New Jersey
General Sekiyu Seisei
Sakai, Japan
Getty Oil Company
Delaware City, Delaware
Hess Oil Virgin Islands Corp.
St. Croix, Virgin Islands
Hess Oil Virgin Islands Corp.
St. Croix, Virgin Islands
Marathon Oil Company
Garyville, Louisiana
Mobil Oil Corp.
Paulsboro, New Jersey
Mobil Oil Corp.
Torrance, California
Claus
Capacity
(LTPD)
140
307
450
300
300
1116
300
150
342
300
320
232
270
200
BSRP
Units
1
3*
2
1
1
2
1
1
1
1
1
2*
2
2
*Employs 1 Stretford unit only.
D-21
-------
TABLE D-5. Continued
Customer and Location
Ninon Ryutan Kogyo K.K.
Tsurusaki , Japan
Texaco, Inc.
Long Beach, California
Toa Oil Company, Ltd.
Kawasaki, Japan
Union Oil Company of California
Chicago, Illinois
Union Oil Company of California
Los Angeles, California
Union Oil Company of California
Rodeo, California
Wintershall AG
Lingen, Germany
Claus
Capacity
(LTPD)
180
350
320
300
200
245
75
BSRP
Units
1
1
2
2*
2
3
1+
*Employs 1 Stretford unit only.
fEmploys Selectox process.
D-22
-------
50 PSIG STEAM
no
CO
1 REDUCING
GENERATOR
LEGEND:
1. FEED GAS
2. TREATED GAS
3. AIR
4. FUEL GAS
5. HOT GAS C<
6. COOL GAS CONTAINING H2S
7. SULFUR
8. SOUR WATER
9. SPERT CATALYST
Q
cc cc
O "J
U_ 00
I- oc
ill O
CC CO
SCO
<
UIKUULAI IIMlj &ULU 1 IUIM j
* *
THREE-STAGE
COOLING
TOWER
Figure -D-4. Beavon Process
-------
2.4 Operating Parameters
2.4.1 Reactor
Temperature: 644°K (700°F)(3)
Pressure: 0.1 MPa (1
2.4.2 Condenser
Temperature: 310°K (100°F)(6)
Pressure: 0.1 MPa (1 atm)(6)
in -j\
2.5 Process Efficiency and Reliability^ ' '
t In refinery applications where the Claus tail gas contains about
4% equivalent H2S, the tail gas from the Beavon process will con-
tain less than 40 ppm equivalent S02* (COS constituting major
portion and h^S will be less than 1 ppm).
• Process involves basic refinery technology and is generally insen
sitive to feed stream upset conditions.
• No unusual maintenance or hazardous conditions are reported.
2.6 Raw Material Requirements
• Fuel gas: 37000 Nm3/day per tonne (1.25 MSCFD/ton) of parent
sulfur plant capacity (2).
• Stretford Solution Alkali: 0.013 to 0.06 l/sec (0.21 to 1.0 gpm)
for 100 tonne per day Claus plant(6>7).
2.7 Utility Requirements
• Power: 70 kwh per tonne (64 kwh/ton) of sulfur in the tail
• Fuel gas: no data available
• Cooling water: 22.7 I/sec (360 gpm) for 100 tonne per day Claus
plantlo;. (Air cooling can be used.)
3.0 Process Advantages'1 '5^
• Recovers organic sulfur compounds and S02 as elemental sulfur.
• Can utilize existing Stretford plant, if available.
*Union guarantees 100 ppmv for refinery applications,
D-24
-------
inrt^! !*: ™nStrUCted of carbon steel with certain items
ing treated with epoxy coating.
• Process is basically insensitive to variations in feed stream
compositions.
• Process produces approximately 80 kg/hr (175 Ib/hr) of 0.43 MPa
(65 psia) steam per ton of sulfur in tail gas(7).
4.0 Process Limitations^2'3^
• High fixed cost of facility including royalty fees.
• Requires some type of fuel gas to supply heat and to produce a reducing
gas for the catalytic reaction.
• Like other catalytic processes, efficiency of conversion of COS and CSo
to HeS is decreased if high levels of CO? are present in Claus plant
tail gas.
5.0 Process Economics'2'
• Costs as reported in 1972 are as follows:
- Fixed costs including royalties
1% Sulfur Equivalent in Feed Gas
Parent Sulfur Plant Capacity Investment
tonne (long ton) per day $ Million
1.11 (1) 0.69
11.1 (10) 1.40
111 (100) 5.80*
4% Sulfur Equivalent in Feed Gas
Parent Sulfur Plant Capacity Investment
tonne (long ton) per day $ Million
1.11 (1) 0.61
11.1 (10) 1.20
111 (100) 3.55*
- Operating costs^ ': approximately $40 per long ton sulfur in
tail gas per day.
*Multiple hydrogenation and Stretford trains.
^Multiple Stretford trains.
D-25
-------
In -j\
6.0 Input Streams v " '
• Feed gas stream, Claus tail gas (Stream 1) see Table D-6.
t Fuel gas (Stream 4): sufficient to heat Stream 1 from 400°K to 620°K
(270°F to 650°F)
• Air (Stream 3): 80%-90% of stoichiometric requirements for fuel gas
• Chemical and catalyst makeup: ADA, vanadium, and caustic soda
lr\
7.0 Intermediate Streams^ '
• Reactor offgas (Stream 5)-see Table D-7.
• Condenser offgas (Stream 6)~see Table D-7.
8.0 Discharge Stream
• Tail gas from process (Stream 2)-see Table D-6.
• Sour water (Stream 8): pH - slightly acidic; h^S and C02 - dissolved
to about 50 ppm each(6). in refineries sour water is recycled to
existing sour water strippers.
9.0 Data Gaps and Limitations
Data gaps exist in the following areas:
• Process applicability to coal conversion process gas purification sys-
tems has not been established, particularly for processing high C09
Claus tail gases. f i
• Characterization of gaseous and liquid feed and discharge streams for
refinery applications (temperature, pressure, composition, etc.).
• The effect that various contaminants (NH3, HCN, carbonaceous matter,
trace metals, etc.) have on the process, and the ultimate fate of such
contaminants in the system.
10.0 Related Programs: No data available.
*Nature of the chemicals not given.
"'"Union has indicated that Beavon systems can be guaranteed to achieve 250
ppmv total sulfur in coal gasification applications^).
D-26
-------
TABLE D-6. TYPICAL BEAVON GAS STREAM COMPOSITION IN
REFINERY APPLICATIONS^)
Components
H2S
so2
s
COS
cs£
co2
H20 vapor
N2
H2
CO
Claus Tail Gas
to Beavon
2.0%
1.0%
0.7%
0.3%
0.3%
10%
26%
56%
2.5%
1.0%
Beavon
Tail Gas
0.0%
0%
0.0%
<250 ppm*
0.0%
14%
5%
80.8%
varies
0.2%
*Union guarantees 100 ppmv; typically 40 ppmv is
attained(7).
D-27
-------
TABLE D-7. TYPICAL COMPOSITION OF BEAVON INTERMEDIATE GAS STREAMS*(6'7)
Components
(vol %)
H2S
so2
S
COS
cs2
co2
H20 vapor
N2
H,
CO
HC (MW=30)
Temperature
Pressure
Claus Tail Gas
to Beavon
0.85
0.42
0.05
0.05
0.04
2.37
33.10
61.30
1.60
i
0.22
—
j 413°K (284°F)
0.126 MPa (18.5 psia)
Reactor Off gas j
1.54
0.00
0.00
40 ppm
2 ppm
3.18
32.30
62.50
0.21
0.20
0.06
673° K (752°F)
0.1 MPa (14.7 psia)
Condenser Offgas
2.13
0.00
0.00
40 ppm
2 ppm
4.39
6.45
86.36
0.29
0.28
.. 0.08
311°K (100°F)
0.1 MPa (14.7 psia)
*Based on the Claus tail gas composition given above for a 100 tonne per day
Claus plant.
D-28
-------
REFERENCES
1. Gas Processing Handbook, Hydrocarbon Processing, April 1975.
2. Beers, U. D., Characterization of Claus Plant Emissions, USEPA, NTIS No.
PB-220-376, April 1973.
3. Maddox, R. N., Gas and Liquid Sweetening, Campbell Petroleum Series, 1974.
4. Beavon, David K., Add-on Process Slashes Claus Tail Gas Pollution, Chemical
Engineering 78 (28), 1971.
5. New Beavon Process Takes Sulfur-Bearing Compounds from Tail Gas, Oil and
Gas Journal, 70 (6), 1972.
6. Battelle Columbus Laboratories, Characterization of Sulfur Recovery from
Refinery Fuel Gas, U.S. EPA, NTIS No. PB-239-777, June 1974.
7. Information provided to TRW by G. E. Tilley of Union Oil Company, June
14, 1978.
8. Letter, G. L. Tilley, Union Oil Research, to C. B. Sedman of Emission
Standards and Engineering Division, Office of Air Quality Planning and
Standards of EPA, January 2, 1976.
D-29
-------
INSTITUT FRANCAIS DU PETROLE (I.P.P.) PROCESS
1.0 General Information
1.1 Operating Principle(1) - The removal of sulfur compounds from Claus
tail gas by catalyticly reacting the H,,S with S02 (the basic Claus
reaction: 2H2S + S02 = 3S + 2H20) in a solvent. The solvent is
generally an alkaline earth metal salt of a carboxylic acid.
1.2 Development Status - Commercially available.
1.3 Licensor/Developer - Institut Francais du Petrole
1 et 4, av. de Bois-Preau
92-Rueil-Malaison
(Hauts-de-Seine) France
1.4 Commercial Applications^ - Claus plant tail gas treatment; approxi-
mately 25 plants in operation or in various stages of planning, de-
sign or construction. Operating plants are located throughout the
world. Table D-8 gives some specific information on four plants in
Japan and a demonstration plant in Canada.
2.0 Process Information
2.1 Flow Diagram (see Figures D-5 and D-6)*
• Figure D-5 illustrates the IFP-1 flow diagram. The Claus tail
gas, Stream 1, is injected into a packed tower, counter-currently
contacting the solvent containing catalyst. Sulfur, Stream 3,
is formed, collected and removed from the tower, and the treated
gas, Stream 2, is sent to an incinerator where the remaining sul-
fur compounds (H2S, COS, C$2) are converted to S02>
*There are two IFP processes: one process, IFP-1, removes H2S and S02 from
Claus tail gas to a S02 equivalent level of 1500 to 2000 ppm; the other or
fefei r^srsiooHir s°2 f™ciaus ^ «•• * ^ i^i
D-3Q
-------
TABLE D-8. I.P.P. PROCESS PLANT LOCATIONS AND APPLICATION
(3)
Plant Owner
Delta Engineering
Corp.
Nippon Petroleum
Refining Company
Idemitsu Oil Co.
Kyokutoh Oil Co.
Showa Oil Co.
Location
Lone Pine
Alberta
Negishi
Japan
Japan
Japan
Japan
Application
Demonstration
Plant
Cleaning tail
gas from 3-
stage Claus
plant
Cleaning tail
gas from 3-
stage Claus
plant
Cleaning tail
gas from 3-
stage Claus
plant
Cleaning tail
gas from 3-
stage Claus
plant
Through-put
Nm3/D (MMSCFD)
21,500 (0.8)
699,400 (26)
592,000 (22)
406,400 (16)
113,000 (4.2)
Sulfur Recovery
Rate (%)
80 - 85
85
85
90
85
-------
CO
PO
LEGEND:
1. Glaus Tail Gas
2. Treated Gas to Incinerator
3. Liquid Sulfur
4. Solvent Recycle
5. Steam Condensate
6. Steam for Start-up
7. Catalyst and Solvent Makeup
Figure D-5. Institut Francais du Petrole Process (IFP-1)
-------
o
CO
co
A. THERMAL CATALYTIC INCINERATOR
B. INCINERATOR
C. AMMONIA SCRUBBER
D. SULFITE EVAPORATOR/SO2 GENERATOR
E. SULFATE REDUCER
F. CATALYTIC REACTOR
LEGEND:
1.
2.
3.
4.
5.
6.
7.
8.
9.
CLAUSTAIL GAS
IFF TAIL GAS (TO ATMOS.)
PURE LIQUID SULFUR
FUEL GAS
NHg MAKEUP
SOLVENT MAKEUP
TAIL GAS PRIOR TO
INCINERATION
AMMONIACAL BRINE
H2S CONTAINING GAS
10. NH3 RECYCLE
11. SO2 CONTAINING GAS
12. SO4= CONTAIN ING GAS
13. SO2/NH3 GAS
Figure D-6. Institut Francais du Petrole Process (IFP-2)
-------
t Figure D-6 illustrates the IFP-2 flow diagram. The Claus tail
gas, Stream 1, is incinerated, then scrubbed with an aqueous
ammonia solution. The scrubbed gas, Stream 7, is incinerated
prior to release to the atmosphere. Ammonical brine, Stream 8,
is piped to a sulfite evaporator/S02 generator and then to a sul
fate reducer unit. S02/NH3 streams, produced in the sulfite
evaporator/S02 generator and the sulfate reducer, are combined
with a supplementary H2S stream (Stream 9) and the combined
stream (Stream 13) is sent to a catalytic reactor for Claus re-
action. The product liquid sulfur, Stream 3, is piped away and
ammonia, Stream 10, is recycled.
2.2 Equipment - Conventional absorbers, evaporators, catalytic reactor,
scrubbers, incinerator, and thermal catalytic incinerator.
- 285°F) maximum for IFP-1 with-
2.3 Feed Stream/Requirements
Temperature: 400°K - 415°K (265°F
out cooling(16).
Pressure: ?
Others: ?
2.4 Operating Parameters
2.4.1 Scrubber - Temperature: ?
Pressure: ?
Solvent loading: ?
Other: ?
2.4.2 Sulfite Evaporator/S02 Generator - Temperature: ?
Pressure: ?
Solvent loading: ?
Other: ?
2.4.3 Catalytic Reactor - Temperature: 393°K - 403°K (248°F
266°F)(6).
Pressure: ?
Solvent/catalyst loading: ?
Other: ?
2.5 Process Efficiency and Reliability - IFP-2 process is reported to ho
capable of removing sulfur species in Claus tail gas to 500 ppm or
less as S02. IFP-1 is capable of removing sulfur species in n
D-34
-------
tail gas to 1500 to 2000 ppm as SQ . No information is available
as to the reliability of the process.
2.6 Raw Material Requirements - No information is available as to the
quantity of makeup ammonia, polyalkaline glycol and catalyst
requirements.*
2.7 Utility Requirements'*"
• Electricity: 35 kwh/hr for IFP-1 process applied to a 100 tonne/
day Claus plant(°).
t Fuel gas: ?
t Water: ?
• Others : ?
2.8 Miscellaneous - No information available which indicates special
maintenance problems or unusual hazardous conditions created by the
process.
3.0 Process Advantages
• Solvent and catalyst are readily available at a low cost^ .
t Produces high quality sulfur.
• Low foaming tendency of solvent.
(3 4)
• Minimum solvent loss due to its low vapor pressure^ * .
t Catalyst is highly
(3 4)
• Both solvent and catalyst are chemically and thermally stable ' .
(5)
• Carbon steel can be used throughout process v '.
• The total H2S + S02, concentration in the feed gas has little effect
on investment cost '^'.
*The solvent and catalyst makeup costs for an IFP-1 process are reported to be
approximately $350 per day for a 1274-tonne (1400-ton) per day Claus plant^2).
Further, the solvent and catalyst makeup cost for an IFP-2 process is approxi-
mately $5 per day for a 228-tonne (250-ton) per day Claus plant(2).
"fThe utility cost for an IFP^-1 process is reported to be approximately $30 per
day for a 1274-tonne (1400-ton) per day Claus plant^'. Also, the utility
cost for an IFP-2 process is approximately $70 per day for a 228-tonne (250-
ton) per day Claus plant(2).
D-35
-------
• IFF process can be made up of a combination of remote unit locations
and central plant to optimize capital investment*^'.
0 Investment is small in comparison with the cost of the Claus plant.
(5)
• Does not create any water pollution^ .
4.0 Process Limitations
• For optimum operating conditions, the H2S/S02 ratio in the feed to the
catalytic reactor should be maintained in the range of 2.0 to 2A(6>,
(5)
• COS and CSp, if present, are not removed in the catalytic reactorv '.
• Tail gas must be incinerated prior to release to the atmosphere via
stacks(3).
• No commercial applications reported for the process other than Claus
tail gas cleanup.
5.0 Process Economics
• The overall cost of a 182-tonne/day (200-ton/day) sulfur plant is
approximately $2.00 per 1000 Nm3 ($53.00/MMSCF) of tail gas treated
6.0 Input Streams
• Feed stream, Claus tail gas, Stream 1, see Table D-9.
• NH, makeup, Stream 5, Figure D-6: ?
• Fuel gas, Stream 4, Figure D-6: ?
• Solvent catalyst makeup, Stream 6, Figure D-6: ?
7.0 Discharge Streams
• IFP tail gas prior to incineration, Stream 2, Figure D-5, see Table D-9.
• Production sulfur, Stream 3, see Table D-9.
• IFP tail gas prior to incineration, Stream 2, Figure D-5: and after
incineration, see Table D-10.
8.0 Data Gaps and Limitations
• Input and discharge stream data supplied above is for IFP-1 orocpss- no
stream information was available for the IFP-2 process! Process' no
*This is a desirable feature in applications where Claus plants are located at
different locations in a major facility or in several near-by pllnts
D-36
-------
TABLE D-9. THREE APPLICATIONS OF THE IFM PROCESS FOR
''• TREATING CLAUS TAIL GAS ('4)*
Stream Composition/
Operating Conditions
Claus Tail Gas
Composition
Stream 1 , Mol e % :
H2S
so2
s
H20
N2, C02, Misc.
Temperature, °K (°F)
Pressure, MPa (psig)
Sulfur Recovery
H2S + S02 reaction, %
Production Rate
(Stream 3), kg/hr
(Ib/hr)
Treated gas to
incinerator, Stream 7,
ppm of H2S + S02
Treating Tail Gas After
One-Stage Claus
1.48
0.74
1.26
28.58
67.94
400 (260)
0.10 (0.50)
95
112.3 (247)
1100
2-Stage Claus
0.59
0.29
0.14
24.96
69.02
400 (260)
0.10 (0.50)
90
36.8 (81)
900
3-Stage Claus
0.34
0.17
0.13
30.25
69.11
400 (260)
0.10 (0.50)
80
19.5 (43)
1000
*Refer to Figure D-5.
D-37
-------
TABLE D-10 TYPICAL COMPOSITION OF GAS STREAMS FOR THE IFP-1
PROCESS FOR TREATING CLAUS TAIL GAS*(6)
Components
(vol X)
H,S
£
so2
s
COS
cs2
CO
co2
H,
H2S
N,
°2
Temperature
Pressure
Claus Tail Gas
to IFP
0.85
0.42
0.05
0.05
0.04
0.22
2.37
1.60
33.10
61.30
—
413°K (285°F)
0.126 MPa (18.5 psia)
After Catalytic
Reactor
0.085
0.042
0.040
0.040
0.075
0.219
2.376
1.607
33.990
61.545
—
3920K (246°F)
0.1 MPa (14.7 psia)
After Incinerator
--
0.212
— •-
—
--
--
4.483
--
30.502
64.299
0.504
923°K (1200°F)
0.1 MPa (14.7 psia)
*Based on the tail gas composition
plant.
given above for a 100-tonne per day Claus
iD-38
-------
• Data gaps exist in the following areas:
- Applicability of the process to coal conversion processes; e.g.,
efficiency, reliability, feed stream requirements.
- Characterization of gaseous and liquid streams (e.g., purified gas,
feed gas) for the IFP-2 process in commercial refinery gas treating
application.
- Definition of the maximum allowable concentrations of various con-
taminants in the feed gas; e.g., COS, CS?, trace metals, carbonace-
ous matter.
- The effect that various contaminants (trace metals, carbonaceous
matter, COS, CS£, HCN) have on the process, and the ultimate fate
of such contaminants in the system. '
9.0 Related Programs
No data available.
REFERENCES
1. Maddox, R. N., Gas and Liquid Sweetening, Campbell Petroleum Series, 1974.
2. Gas Processing Handbook,Hydrocarbon Processing, April 1975.
3. Beer, W. D., Characterization of Claus Plant Emissions, U.S. EPA, NTIS
PB-220-376, April 1973.
4. Barthel, Y., Y. Bistri, et al, Treat Claus Tail Gas, Hydrocarbon Processing,
May 1971.
5. Bonnifay, P., R. Dutrian, et al, Partial and Total Sulfur Recovery, Chemical
Engineering Progress, Vol. 68, No. 8, August 1972.
6. Battelle Columbus Laboratories, Characterization of Sulfur Recovery from
Refinery Fuel Gas, U.S. EPA, NTIS No. PB-239-777, June 1974.
D-39
-------
SULFREEN PROCESS
1.0 General Information
1.1 Operating Principles^ '2'3'4'5) - S02 and H2S in Claus tail gas are
removed by further promotion of the Claus reaction on a catalytic
surface. This process was designed specifically for treatment of
tail gases of Claus type sulfur plants, either in gas processing
plants or refineries. The reaction is carried out in a solid batch
reactor utilizing an activated alumina catalyst. Adsorbed sulfur is
desorbed with hot Claus tail gas (circulating in a closed loop sys-
tem) from which the sulfur vapor is removed in a condenser-coalescer.
1.2 Licensor/Developer - Developed by Lurgi Gesellshaft for Warme and
Chemotechnic (Lurgi) of West Germany and Societe Nationale des
Petroles D'Aquitaine (SNPA) of France. In addition to Lurgi and
SNPA, the process is licensed to:
Ralph M. Parsons Company, Pasadena, California
Fluor Engineers and Constructors, Inc., Irvine, Calif.
Ford, Bacon & Davis, Dallas, Texas
Partec Lavalin, Inc., Calgary, Alberta, Canada
1.3 Commercial Status Application - Commercially available. Sixteen
commercial scale plants treating sulfur plant tail gas have been
constructed.
2.0 Process Information
2.1 Flow Diagram (see Figure D-7(1'4))- Claus tail gas (Stream 1) is
introduced in parallel into a battery of catalytic reactors where the
Claus reaction is carried out at 130°C to 160°C (265°F to 320°F).
This temperature is lower than that used in a Claus process, and the
formation of elemental sulfur is favored. A battery of six reactors
is shown in Figure D-7; four are utilized for sulfur adsorption, one
D-40
-------
SULFUR PRODUCING
REGENERATION
COOLING
o
-pa
LL J J
CONDENSER COALESCER
1. SULFUR PLANT TAIL GAS
2. PURIFIED GAS
3. COOLING GAS
4. REGENERATION GAS
5. COOLING AND REGENERATION GAS
6. LIQUID SULFUR
STEAM
Figure D-7. Sulfreen Process Flow Diagram
-------
for catalyst regeneration, and one for cooling after regeneration.
A Sulfreen process may consist of only three catalytic reactors, two
in adsorption and one in desorption, depending on tail gas composi-
tion and flow rate and economic considerations. Desorption of sulfur
from the catalyst is achieved by heating a regeneration gas (Stream
4), usually tail gas from the Glaus unit, to 320°C (608°F) and circu-
lating it through the catalyst bed, thereby vaporizing the adsorbed
sulfur. The vaporized sulfur is condensed and removed from the re-
generation gas in the condenser-coal escer. The condenser-coal escer
reduces the regeneration gas temperature to about 120°C (248°F). The
cooled gas is utilized to reduce the catalyst bed to a temperature
suitable for adsorption after the regeneration process is completed.
The process operates continusouly and the reactors are sequenced
between the adsorption and desorption processes.
2.2 Equipment^ ' ' - All equipment can be constructed of carbon steel if
provision is made to maintain temperatures above the water dew point
to avoid corrosion. However, stainless steel may be used for cata-
lytic reactors and a portion of the regeneration circuit.
(1 4)
2.3 Catalyst Lifev ' ' - Activated carbon or alumina catalyst is expected
to have a life of approximately 4 to 6 years.
2.4 Process Efficiency - Capable of 80% to 85% removal of sulfur from
tail gas( ' . The Sulfreen unit operating at LACQ, France is 75%
efficient '. When using ordinary alumina catalyst, 80% sulfur
removal is obtained, with combined H2S and S02 concentration of
2000-2500 ppm in the treated gas stream. Overall sulfur removal
efficiencies for Claus and Sulfreen are in the range of 98.5% to 99%.
A promoted activated alumina has been developed by Lurgi and SNPA
to prevent poisoning by sulphation and aging. Overall sulfur
removal efficiency for Claus and Sulfreen of 99.5% is obtained
with combined H2S and S02 concentration of 1000-1200 ppm. Very
good conversion of COS is also obtained.
D-42
-------
2.5 Chemical Requirements(1) - (Based on 100 tonne per day Claus unit)
• Catalyst alumina: 11.0 kg/hr (24.3 Ib/hr)
• Nitrogen: 44.9 Nm3/hr (1675 scf/hr)
2.6 Utility Requirements"' - (Based on 100 tonne per day Claus unit and
alumina catalyst)
t Steam produced in process (0.48 MPa, 70 psia saturated):
735 kg/hr (1620 Ib/hr)
• Electricity: 124 kwh/hr
t Fuel Gas: 60.9 Nm3/hr (2272 scf/hr)
• Boiler feed water: 0.189 I/sec (3.0 gpm)
3.0 Process Advantages
• Process does not have a major liquid waste stream' '.
• Produces high quality sulfur - 99.9% pure'3'.
• Alternating catalytic reactors between adsorption and desorption
permits continuous operation(2).
4.0 Process Limitations
t C$2 is not appreciably removed^ '.
stoichiometric H2S:Su?
Claus tail,gas, which necessitates careful control of Claus unit
.(2)
cu ;>pci_ i i i i~a i ly i \ji v-> i au.a ia i i yuc
f-3 /n
5.0 Process Economics
t Optimum performance requires a stoichiometric h^SOg ratio of 2:1 in
operations(2).
• Process designed specifically for Claus tail gas'
.(3,4)
A Sulfreen plant processing 110 MM SCFD of tail gas from a one million
long ton per day sulfur plant was constructed for three million dollars.
Operation and maintenance costs varied from $150,000 to $180,000 per
year (1969 dollars). A Sulfreen plant processing 220,000 Nm2/hr (197 MM
SCFD) of tail gas from a 2,200 ton per day sulfur plant was constructed
for 3.2 million dollars (Ram River Stage II, 1973 dollars).
6.0 Input Streams
6.1 Claus tail gas, prior to incinerator (Stream 1), see Table D-ll.
D-43
-------
TABLE D-ll. TYPICAL GAS STREAM COMPOSITION FOR SULFREEN PROCESS
(0
Components
(VOL %)
H2S
so2
S Vapor and
Mist
COS
cs2
CO
co2
H2
H20
N2
°2
Temperature
Pressure
Claus Tail Gas Prior
to Incineration
0.85
0.42
0.05
0.05
0.04
0.22
2.37
1.60
33.10
61.30
-
140°C (284°F)
0.126 MPa (18.5 psia)
Purified Gas
0.18
0.085
0.013
0.051
0.04
0.222
2.39
1.62
33.44
61.93
-
140°C (284°F)
0.1 MPa (14.7 psia)
Incinerated
Sulfreen Tail
Gas
-
3,385 ppn
—
-
-
-
2.9
-
28.93
67.23
0.61
650°C (1202°F)
0.1 MPa (14.7 psia)
*Based on a 100 tonne per day Claus plant
NOTE: This stream data is considered to be out of
the only data available at this time.
However, it is
D-44
-------
7.0 Discharge Streams
7.1 Purified gas (Stream 2), see Table D-ll.
7.2 Incinerated Sulfreen tail gas, see Table D-ll.
8.0 Data Gaps and Limitations
• No information which would indicate applicability to coal conversion
processes.
REFERENCES
1. Battelle Columbus Laboratories, Characterization of Sulfur Recovery from
Refinery Fuel Gas, USEPA, NTIS.No. PB-239-777, June 1974.
2. Riesenfeld, F. C., and Kohl, A. L., Gas Purification, Second Edition, Gulf
Publishing Co., Houston, Texas, 1974.
3. Maddox, R. N., Gas and Liquid Sweetening, Campbell Petroleum Series, 1974.
4. Morin, M. M., and Philardeau, T. M., Sulfreen Process Experiences at Ram
River (Alberta, Canada), CNGPA Meeting, June 9, 1976, Calgary, Canada.
5. Grancher, P., Recent Advances in Claus Techniques for Sulfur Recovery from
Acid Gases, International Sulfur Symposium, October 25, 1977, Calgary,
Canada.
6. Information provided to TRW by Y. M. Philardeau of the Aquitaine Company
of Canada Ltd., June 6, 1978.
D-45
-------
CLEANAIR PROCESS
1.0 General Information
1.1 Operating Principles' ' - The purification of sulfur plant tail
gases by catalytic conversions to H2S, followed by a continuation
of the Claus reaction and a Stretford unit where h^S is recovered
as elemental sulfur. A fixed-bed reactor contains both a reduction
(possibly Co/Mo) and a hydrolysis (unknown) catalyst. The Claus
reaction is promoted in a packed reactor with an unknown proprietary
chemical solution.
1.2 Licensor/Developer - J. F. Pritchard Company
4625 Roanoke Parkway
Kansas City, Missouri 64112
1.3 Commercial Status/Application - Commercially available. Three units
built in the U.S. remained on start-up status due to recurrent
operating problems; two other units have been constructed in the
U.S.S.R.
2.0 Process Information
2.1 Flow Diagram (See Figure D-Sr1'2'3' - A limited amount of informa-
tion is available pertaining to specific details of the Cleanair
process. The J. F. Pritchard Company apparently is reluctant to
divulge specific process information. The process can be installed
in three stages: the first stage converts S02 to sulfur; the second
stage removes H2S in a Stretford process; and the third stage con-
verts COS and CS2 to H2S. The Stretford process offgas may be:
incinerated in a typical Claus incinerator, converting residual
H2S to S02 and CO to a>2; or discharged directly into the atmosphere.
(2)
2.2 Catalyst Lifev ' - Catalyst life will generally vary from 2 to
5 years depending on plant operation and feed characteristics. High
C02 concentrations will shorten catalyst life.
D-46
-------
STAGE III
STAGE
o
-Pi
v
Y REACTION
/\ TOWER
CLAUSTAIUGAS
QUENCH REACTION SOLUTION
REACTION SOLUTION
H2S RICH STREAM
5. PURGE STEAM TO COOLING TOWER
6. STEAM
AIR
TREATED GAS TO INCINERATOR
LIQUID SULFUR TO STORAGE
Figure D-8. Cleanair Process Flow Diagram
-------
2 3 Process Efficiency - Plant effluent is normally guaranteed to contain
(2)
less than 250 to 300 ppm of equivalent S02V '.
2.4 Utility Requirements^2'3^ - (Claus unit capacities in long tons per
day: Case 1 = 50; Case 2 = 150; Case 3 = 500)
• Electricity: Case 1 - 200 kw
Case 2 - 580 kw
Case 3-1900 kw
• Fuel Gas (8000 Kcal/m3, 900 Btu/scf):
Case 1 - 13.4 Nm3/hr (500 scfh)
Case 2-40 Nm3/hr (1500 scfh)
Case 3-121 Nm3/hr (4500 scfh)
• Cooling Water (27.7°C, 80°F; 8.3°C, 15°F rise):
Case 1 - 30. I/sec (475 gpm)
Case 2 - 88. I/sec (1400 gpm)
Case 3-287. I/sec (4550 gpm)
• Steam (0.44 MPa, 50 psig saturated):
Case 1 - 181. kg/hr (400 Ib/hr)
Case 2 - 544. kg/hr (1200 Ib/hr)
Case 3 - 1814. kg/hr (4000 Ib/hr)
C\ ?\
3.0 Process Advantagesv ' '
• Produces high quality sulfur
t Can be adapted and retrofitted to existing Claus plants
• Provides flexibility in handling varying amounts of sulfur con-
stituents (may vary threefold)
• H2S : S02 ratio in the tail gas can vary up to 8:1 without affectina
efficiency a
• Potentially capable of very low sulfur emissions.
D-48
-------
4.0 Process Limitations^
• Operational difficulties have been encountered
• High cost
5.0 Process Economics^2'3) - (1972 dollars)
• Capital investment: Case 1 - $ 925,000
Case 2 - $1,400,000
Case 3 - $2,200,000
• Annual operating and maintenance: Case 1 - $203,700
Case 2 - $332,500
Case 3 - $624,500
6.0 Input Streams
6.1 Feed Gas (Stream 1) - no data available
7.0 Waste Streams
7.1 Treated Gas (Stream 8) - no data available
7.2 Sour water stream purged from depurator (Stream 5) -no data available
8.0 Data Gaps and Limitations
t Disclosure of technical details of the Cleanair process requires com-
pletion of a secrecy agreement. Therefore, detailed stream data and
process information are not available.
• No information which would indicate applicability to coal conversion
processes.
9.0 Related Programs - none known
REFERENCES
1. Battelle Columbus Laboratories, Characterization of Sulfur Recovery from
Refinery Fuel Gas, U.S. EPA, NTIS No. PB-239-777, June 1974-
2. Beers, W. D., Characterization of Claus Plant Emissions, U.S. EPA, NTIS No.
PB-220-376, April 1973.
3. Maddox, R. N., Gas and Liquid Sweetening, Campbell Petroleum Series, 1974.
D-49
-------
Sulfur Oxides Control Module
Well man-Lord
Chiyoda Thoroughbred 101
Shell Copper Oxide
Lime/Limestone Slurry Scrubbing
Dual Alkali
Magnesium Oxide Scrubbing
D-50
-------
WELLMAN-LORD PROCESS
1.0 General Information
1.1 Operating Principles - Absorption of sulfur dioxide in a concen-
trated sodium sulfite solution followed by recovery of sulfur di-
oxide gas and solution regeneration.
1.2 Developmental Status - Commercially available.
1.3 Licensor/Developer - Developed by Davy-Powergas, Inc. (Lakeland,
Florida), formerly Wellman-Lord, Inc.
1.4 Commercial1Applications^1*2'3'9' - To date the commercial applica-
tions have been primarily for desulfurization of flue gas from
fossil fuel-fired boilers. Systems in service include 14 oil-fired
boilers in Japan and a 115-MW demonstration plant at the Mitchell
Station of the Northern Indiana Public Service Company (the only
application of process in coal-fired electric utility service).
The most recent installation is on two 335-MW coal fired boilers in
start-up operation at Public Service Co. of New Mexico at Farmington.
The process would be applicable to the control of S02 emissions from
other types of industrial operations and non-ferrous smelting, sul-
furic acid and Claus plants. Twenty-five commercial installations
are in current operation worldwide (all applications, including
those for Claus plant and sulfuric acid plants). Possible applica-
tions in a commercial coal gasification facility may be in connec-
tion with support operations such as utility boilers and sulfur
recovery (Claus plant tail gas treatment).
2.0 Process Information
2.1 Flow Diagram - see Figure D-9
• Process Description - The process can be viewed as composed of
four major processing steps - flue gas pretreatment, S02
D-51
-------
SO2 GAS SCRUBBING
AND ABSORPTION
01
r\>
rx
EVAPORATOR-
CRYSTAL LI ZER 11
-^Sl
A
L_
SODA ASH
DISSOLVING
TANK
1. RAW GAS
2. VENTURI MAKEUP WATER
3. STEAM
4. SODA ASH OR CAUSTIC
5. CLEANED GAS
6. PRETREATMENT SCRUBBER SLOWDOWN
7. COOLING AIR/WATER
8. SULFATE PURGE STREAM
9. COMPENSATE (THERMAL)
10. CONCENTRATED S02 STREAM
11. MAKEUP WATER
Figure D-9. Wellman-Lord Process Flow Sheet
-------
absorption, absorbent regeneration and purge treatment. See
Table D-12 for brief descriptions of these processing steps.
2.2 Equipment^2'4^ - see Table D-12.
2.3 Feed Stream Requirements
Pressure: Slightly above standard pressure.
Temperature: Mormally designed to receive gas into the prescrubber
at less than 533°K (500°F) and to receive saturated gas into the
absorber at 311°K-339°K (100°F-150°F). Flue gas from utility
boilers are usually somewhat less than 533°K.
Loading: Process can handle SCL concentrations well over
10,000 ppm.
Contaminant and Other Limitations^ ': The system is very sensitive
to the buildup of contaminants (sulfate, thiosulfate, and flyash).
For applications to Claus plant tail gases it may be necessary to
incinerate the gas to destroy hLS, COS and CSp prior to SO^ absorp-
tion, since these constituents tend to form thiosulfates which do
not regenerate. The resultant sulfate levels are controlled at a
level of about 5 wt % in the absorber feed by continuously purging
sulfate at a rate equivalent to approximately 5%-10% of the
absorbed sulfur value.
2.4 Operating Parameters
• Absorption
Temperature: 310°K-340°K (100°F-150°F)(4)
Pressure: Close to 0.10 MPa (1 atm).
Loading: to 10,000 ppm SOg.
• Regeneration
Temperature: 369°K (205°F)(9)
(Q\
Pressure: 0.068 MPa (10 psiar
D-53
-------
TABLE D-12. WELLMAN-LORD PROCESS DESCRIPTION AND EQUIPMENT (SEE FIGURE D-9)
Processing Steps
Description
Equipment Used
Flue gas pretreatment
S09 absorption
p
in
Absorbent regeneration
Purge treatment
Removal of residual fly ash by scrubbing,' gas
is also cooled and saturated.
Absorption of SO? in a concentrated solution
of sodium sulfite and bisulfite; sodium sul-
fite reacts with SO? to form more sodium bi-
sulfite; side reactions include oxidation of
sodium sulfite forming non-regenerabTe sodium
sulfate.
Thermal treatment to release the absorbed S02
(some crystalling sodium sulfite precipitates
out during treatment); the S02-bearing stream
is partially condensed to remove water (which
is recycled to the dissolution tank); the
concentrated SOg-bearing stream can be pro-
cessed to produce elemental sulfur, sulfuric
acid, or liquid S02- Soda ash or caustic
soda is added to the dissolution tank as the
make-up chemical.
Purging of nonreactive/nonregenerable sodium
species (sulfate, thiosulfate) from the
system. A slip stream is treated and a
crystalline product containing primarily
sodium sulfate is produced. Details of the
operation are not known.
Low pressure-drop (10-15 cm
H20) venturi scrubbers,
followed by an entrainment
separator.
Conventional multi-stage
(commonly 3 to 5 stage) (e.g.,
tray tower). The largest
absorption unit will handle
the flue gas from 150-200 MW
boilers. A large capacity
surge tank installed between
the absorber and the
regenerator.
Forced-circulation evaporator/
crystal!izer of either
sing!e-or-multi pie-effect
design; single-effects usually
used for systems smaller than
150 MW. Stainless steel
piping also recommended for
all 5.1 cm (2 in.) and smaller
pipes in solution service.
Not known.
-------
2.5 Process Efficiency and Reliability - SCL removal efficiency has
been proven at >90% for S02 concentrations up to 20,000 ppnr9'.
Reliability in terms of on-stream time has been >97% for all
installations^4'9^.
The system successfully fulfilled performance acceptance test
requirements at NIPSCO, a coal -fired utility, on Sept. 15, 1977.
S09 removal was 91%, particulate emissions were 0.072 g/10 cal
t- c
(0.04 lb/10 Btu), sulfur product purity was 99.9% and sodium
( 2} *
carbonate makeup was 0.26 kg/kg sulfur removal v.
2.6 Raw Materials Requirements
Basis - Performance of oil-fired systems in Japan using
Wellman-Lord process.
Sodium makeup: 5%-20% of . absorbed sulfur value. A purge
containing 10% of the absorbed sulfur value tied up as sodium
salts corresponds to a sodium makeup equivalent to 0.25 kg
NaOH or 0.33 kg Na2C03 per kg S absorbed^.
Basis - NIPSCO coal -fired boiler acceptance test results.
Sodium carbonate: equivalent to 9.4% of S02 absorbed.
2.7 Utility Requirements
Steam: 15 kg/kg S02absorbedv ' . Another source reports 5-10 kg
steam/kg S02 evaporated^. 25,455-29,318 kg/hr (56,000-64,500
Ib/hr) usage reported for NIPSCO system' '. This is equivalent to
14.5-16.5 kg steam/kg S02 absorbed.
Electricity: 3.0 kcal/Nm3 flue gas
Process Water: 0.055 1/Nm3 flue gas (0.0004 gal/scf) including
(9)
prescrubberv '.
Electricity: 3.72 kcal/Nm3 (0.0070 kw/scfm)^4^.
Electricity: 3.0 kcal/Nm3 flue gas (0.0056 kw/scfm)^
Calculated from data reported for NIPSCO acceptance tests performance.
•(•Calculated from data in Reference 2 based on acceptance test data at NIPSCO
plant.
D-55
-------
o tq\
Cooling Water: 1.1 liters/Mm (0.008 gal/scf)v ' .
3 3
Reheater Flue Gas, 1010 Btu/scf (9000 kcal/m ): 215 Nm /hr
(7600 scf/hr) based on 100 tonne per day Claus plant for reheat
to 588°K (600°F).
3.0 Process Advantages
(4)
• A concentrated S02 stream containing up to 90% S02 can be produced
• Can remove in excess of 95% of S02 from streams containing as much
as 20,000 ppm(9).
• By installing large surge capacity, absorption and regeneration
sections of, plant can operate independently, thus enhancing its
reliability^.
• Low scale potential in the scrubber system; no potential of calcium
scaling(4).
• Ability to separate the scrubber system operation from the regenera-
tion section, which allows the use of a centrally located regeneration
facility serving a number of different scrubbers™).
• Considerable operating experience has been obtained with oil-fired
boilers, sulfuric acid plants, and Claus plants in addition to the
present coal -fired utility at NIPSCOU).
• The sulfuric acid plant size requirements are relatively small
due to the high concentration of recovered SO,^7'.
t Low liquid-to-gas ratios are required in the scrubber^ .
4.0 Process Limitations
• Sensitivity of the system operation to the buildup of contaminants.
The system requires a prescrubber for feed gases containing high
particulate loading. The liquid bleed from the prescrubber has a
low pH (1.5-2.0) and must be neutralized prior to being discharged
• Some oxidation of sulfite to nonreactive sulfate if high sulfur
trioxide or high oxygen levels exist in the feed
Small quantities of nonreactive sodium species such as sodium
sulfate, thiosulfate, (formed from H2S, COS, CS2 in gas) must be
purged from the system and replaced by caustic or soda ash. This
creates a handling and disposal problem™).
D-56
-------
• The process operates near the solubility limit of sodium sulfite
in designs where a prescrubber is not needed. If the S02 level in the
feed gas drops suddenly, less of the more soluble sodium bisulfite
would be formed and sodium sulfite precipitation could occur locally
in the scrubber as the gas is cooled(4).
t High steam usage: around 6-10 kg steam per kg SC>2 absorbed^ ,
depending on the application.
• The evaporator system must be maintained free of solids' '.
• Problems have arisen in the past from pitting corrosion of
evaporati on tubes(7).
(Q\
5.0 Process Economics^ ;
A Wellman-Lord unit handling a gas containing 5820 kg/hr (12,800 Ibs/hr)
SOp is estimated to cost about 16 million dollars (1978). A unit handling
580 kg/hr (1200 Ibs/hr) S02 is estimated to cost about 2.3 million
dollars (1978). These estimates do. not include facilities for
incineration or S0? compression.
6.0 Input Streams
6.1 Raw gas (Stream 1): See Table D-13.
6.2 Steam (Stream 3): 0.3 MPa saturated (30 psia)
6.3 Soda ash or caustic (Stream 4); 3.7 tonnes/day for a 100 tonne/
day Claus plant^
6.4 Venturi makeup water (Stream 2): Not required for Claus plant,
can use cooling water; no operating data available.
6.5 Cooling water (Stream 7): No operating data available, designs
usually specify a 14°K (25°F) temperature rise.
6.6 Makeup water (Stream 11): No data available.
7.0 Discharge Streams
7.1 Reheater Exhaust Gas (Stream 5) - See Table D-13.
7.2 Concentrated S02 Stream (Stream 10)^ - 85 vol % S02, 15% H20.
7.3 Pretreatment Scrubber Slowdown (Stream 6) - No data available.
D-57
-------
TABLE D-13. TYPICAL COMPOSITION OF GAS STREAMS ENTERING AND LEAVING.100 TONNE PER DAY
REFINERY CLAUS PLANT WITH WELLMAN-LORD TAIL GAS PROCESS(
Composition
(Vol 30
so2
co2
H20
N2
°2
Temperature
Pressure
Claus
Incinerated
Exhaust
1.08
4.23
26.57
66.68
1.44
650°C
0.1 MPa
(14.7 psia)
Quench Outlet
(Absorber Inlet)
Stream 1
1.34
5.26
8.76
82.85
1.79
43°C
0.1 MPa
(14.7 psia)
Absorber Outlet
Stream 8
250 ppm
5.33
8.88
83.98
1.81
43°C
0.1 MPa
(14.7 psia)
Reheater Exhaust
to Atmosphere
Stream 15
215 ppm
5.66
9.80
82.43
2.11
316°C
0.1 MPa
(14.7 psia)
en
oo
-------
TABLE D-14. APPROXIMATE COMPOSITION OF WELLMAN-LORD PURGE
STREAM FROM CLAUS PLANT APPLICATION^8'9)*
Component
Wt %
TDS:
26
5
Na2S03
14
Water
74
*No centrifuge in the system; water added to
dissolve all solids.
tActual amount in solution unknown, but is
estimated to be about 1% by weight.
7.4 Purge Stream (Stream 8) - See Table D-14 for Claus plant application,
composition for fossil fuel boiler will be different.
7.5 Heat Exchanger Condensate (Stream 9) - No data available.
8.0 Data Gaps and Limitations
Several limitations exist in Well man-Lord process operating data;
these include:
t Lack of stream characterizations for most streams due to the
proprietary nature of the process.
t Actual operating data are limited for commercial installations.
t Data are lacking on the most optimized version of the process which
would operate with double-effect evaporators, and convert purged
salts to a final, solid by-product or waste material.
D-59
-------
9.0 Related Programs
EPA has contracted for an independent analysis of the full-scale
Wellman-Lord process at NIPSCO. The objectives of the test program
are to: assess the technical and economic feasibility of the process;
determine the applicability and control capability of the process;
determine the magnitude and characteristics of the liquid and solid
waste streams; and investigate performance with respect to varying
inlet flue gas conditions. An interim report on the testing/analytical
results is expected to be published in late 1978 or early 1979.
REFERENCES
1. Delgado, F. F. Recent Operating Experience of the Wellman-Lord FGD
Process on a Coal-Fired Boiler. Davy Powergas Inc., Lakeland,
FL 33803, 7 pp.
2. Link, W. F. and W. H. Ponder. Status Report on the Wellman-Lord/Allied
Chemical FGD Plant at Northern Indiana Public Service Co.'s Dean H.
Mitchell Station. Presented at Fourth FGD Symposium, EPA, Hollywood,
Florida, Nov. 8-11, 1977, 18 pp.
3. Boyer, H. A. and R. I. Pedroso. Sulfur Recovered from S02 Emissions at
NIPSCO's Dean H. Mitchell Station. Presented at Fourth FGD Symposium,
EPA, Hollywood, Florida, Nov. 8-11, 1977, 18 pp.
4. Kittrell, J. R. and N. Godley. Impact of SOX Emissions Control on
Petroleum Refining Industry. Vol II, Appendix L. EPA-600/2-76-161b, U.S.
EPA, Research Triangle Park, N. C., June 1976, 300 pp.
5. Maddox, R. N. Gas and Liquid Sweetening, 2nd ed. Campbell Petroleum
Series, Norman, Oklahoma, 1974, 300 pp.
6. Davis, John C. S02 Absorbed from Tail Gas with Sodium Sulfite. Chemical
Engineering, Nov. 29, p. 45-46, 1971. .
7. The Status of. Flue,f Gas Desulfurization Applications ,in the U.S.:, A
Technological Assessment. The Federal Power Commission, Bureau of
Power, July 1977.
8. Genco, J. M. and S. S. .Tarn, Characterization of Sulfur Recovery from
Refinery Fuel Gas, Battelle-Columbus Laboratories, EPA, NTIS: PB-239-777,
June 1974.
9. Information provided to TRW by L. H. Grieves of Davy Powergas,
June 16, 1978.
D-60
-------
CHIYODA THOROUGHBRED 101 PROCESS
1.0 General Information
M p o\
1.1 Operating Principlesv ' '°' - Purification of boiler flue gas or
incinerated Glaus tail gas, utilizing a dilute sulfuric acid solution
with a catalyst to absorb/oxidize S02- Gypsum is produced as an
end product. The scrubbing liquor used to absorb S0? is sent to an
oxidizer where residual sulfurous acid is oxidized to sulfuric acid.
Sulfuric acid from the oxidizer is neutralized with limestone to
crystallize and separate gypsum. The absorber offgas is reheated
and discharged to the atmosphere.
1.2 Developmental Status^ ' - Commercially available and fully tested
both in the U.S. and Japan. A new process modification that will
remove oxides of nitrogen has been piloted^ .
1.3 Licensor/Developer - Chiyoda Chemical Engineering and Construction
Co., Ltd.
Chiyoda International Corp.
1300 Park Place Building
1200 6th Avenue
Seattle, Washington 98101
(206) 624-9350
1.4 Commercial Applications^ ' - The Chiyoda Thoroughbred 101 process
has been applied to three Claus sulfur plants, eight industrial
boilers and one industrial incinerator in Japan, as of mid-197411 '. In
the U. S., Chiyoda International has tested a 23-MW prototype unit
on a coal fired utility boiler at Plant Scholz Station of Gulf Power
Co. in Sneads, Florida'5'. Presently, there are fifteen Chiyoda
(4)
installations in Japanv .
D-61
-------
2.0 Process Information
2.1 Simplified Flow Diagram (see Figure D-IO)"'2'3' - Incinerated Claus
flue gas (Stream 1) is scrubbed with recirculated water for removal
of particulate matter and cooling to approximately 328°K (131 F).
Particulates scrubbed from the flue gas are filtered from the
scrubbing water before returned to the prescrubber. The SO,, con-
tained in the flue gas is absorbed by dilute (2%-3%) sulfuric acid
in the absorber at 323°K to 343°K (120°F to 160°F). Absorber vent
gas is reheated by direct combustion of fuel to avoid steam plume
formation from the stack. Sulfurous acid formed in the absorber is
reacted with oxygen from the air in the oxidizer to produce sulfuric
acid in the presence of soluble sulfate catalyst. Sulfuric acid
produced in the oxidizer (Stream 5) is neutralized with limestone,
or other calcium compound, in the crystallizer, thus producing gypsum.
Gypsum crystals are separated by a centrifuge and dry gypsum (5%
to 20% moisture content) is conveyed to storage (Stream 2). Catalyst
makeup is added to the mother liquor tank before the liquor is
recycled to the absorber (Stream 7). Some purging (Stream 10) of
liquor may be required to minimize the level of solubles in the
system. The purge rate is determined by the rate at which solubles
enter the system via flue gas particulate matter or corrosion.
2.2 Equipment - Venturi prescrubber, stainless steel absorber (packed
column), Chevron type mist eliminator, oil or gas fired reheater,
bubbling column oxidizer. Dilute sulfuric acid storage tank,
limestone silo and slurry vessel, precipitator/crystal!izer reactor
clarifier; centrifuge and fly ash thickener.
2.3 Feed Stream Requirements^ * '
t Pressure: No experience above one atmosphere.
• Temperature (Flue Gas): typically 427°K-478°K (310°F-400°F),
no actual restriction^
D-62
-------
CHIYODA THOROUGHBRED 101 PROCESS
.0 General Information
( 1 23)
1.1 Operating Principles^ ' 9 ' - Purification of boiler flue gas or
incinerated Claus tail gas, utilizing a dilute sulfuric acid solution
with a catalyst to absorb/oxidize S0?. Gypsum is produced as an
end product. The scrubbing liquor used to absorb S02 is sent to an
oxidizer where residual sulfurous acid is oxidized to sulfuric acid.
Sulfuric acid from the oxidizer is neutralized with limestone to
crystallize and separate gypsum. The absorber offgas is reheated
and discharged to the atmosphere.
1.2 Developmental Status^ ' - Commercially available and fully tested
both in the U.S. and Japan. A new process modification that will
remove oxides of nitrogen has been piloted^ ' .
1.3 Licensor/Developer - Chiyoda Chemical Engineering and Construction
Co., Ltd.
Chiyoda International Corp.
1300 Park Place Building
1200 6th Avenue
Seattle, Washington 98101
(206) 624-9350
1.4 Commercial Applications^ ' - The Chiyoda Thoroughbred 101 process
has been applied to three Claus sulfur plants, eight industrial
boilers and one industrial incinerator in Japan, as of mid-1974^ '„ In
the U. S., Chiyoda International has tested a 23-MW prototype unit
on a coal fired utility boiler at Plant Scholz Station of Gulf Power
(5)
Co. in Sneads, Florida^ . Presently, there are fifteen Chiyoda
(A)
installations in Japan^ .
D-61
-------
2.0 Process Information
( ~\ 1 Q \
2.1 Simplified Flow Diagram (see Figure D-lOr ' ' - Incinerated Claus
flue gas (Stream 1) is scrubbed with recirculated water for removal
of particulate matter and cooling to approximately 328 K (131 F).
Particulates scrubbed from the flue gas are filtered from the
scrubbing water before returned to the prescrubber. The SO,, con-
tained in the flue gas is absorbed by dilute (2%-3%] sulfuric acid
in the absorber at 323°K to 343°K (120°F to 160°F). Absorber vent
gas is reheated by direct combustion of fuel to avoid steam plume
formation from the stack. Sulfurous acid formed in the absorber is
reacted with oxygen from the air in the oxidizer to produce sulfuric
acid in the presence of soluble sulfate catalyst. Sulfuric acid
produced in the oxidizer (Stream 5) is neutralized with limestone,
or other calcium compound, in the crystallizer, thus producing gypsum.
Gypsum crystals are separated by a centrifuge and dry gypsum (5%
to 20% moisture content) is conveyed to storage (Stream 2). Catalyst
makeup is added to the mother liquor tank before the liquor is
recycled to the absorber (Stream 7). Some purging (Stream 10) of
liquor may be required to minimize the level of solubles in the
system. The purge rate is determined by the rate at which solubles
enter the system via flue gas particulate matter or corrosion.
2.2 Equipment - Venturi prescrubber, stainless steel absorber (packed
column), Chevron type mist eliminator, oil or gas fired reheater,
bubbling column oxidizer. Dilute sulfuric acid storage tank,
limestone silo and slurry vessel, precipitator/crystallizer reactor
clarifier; centrifuge and fly ash thickener.
2.3 Feed Stream Requirements^ ' '
t Pressure: No experience above one atmosphere.
t Temperature (Flue Gas): typically 427°K-478°K (310°F-400°F),
no actual restriction.
D-62
-------
REHEATER
—e
LIMESTONE
HOPPER
CENTRIFUGE
C JMOTHER
V 1 / LIQUOR
• TANK
- ' —
10
LEGEND:
1. INCINERATED CLAUS TAIL GAS, SOX RICH
2. GYPSUM TO STORAGE
3. VENT GAS
4. PRESCRUBBER MAKEUP WATER
5. SULFURIC ACID RICH SOLUTION
6. NEUTRALIZED SOLUTION
7. MOTHER LIQUOR
8. AIR
9. STEAM
10. MOTHER LIQUOR PURGE STREAM
11. FILTER SOLIDS
12. CATALYST SOLUTION
13. CALCIUM SALT
14. PRESCRUBBER OFFGAS
15. S
15. ABSORBER OFFGAS
Figure D-10. Chiyoda Process
D-63
-------
• Loading: Normally designed for about 2000 ppm SO? inlet
concentration, but can be designed for any typical utilitytlue
gas concentrations resulting from coal combustion. Maximum
loading to date is 11,000 ppm S02-
• Other: Absorbent chloride concentrations shall not exceed 200i ppm
to prevent pitting and corrosion in the stainless steel vessels.
2.4 Operating Parameters '
• Absorption System
- Temperature: 322°K (120°F) - Normal recirculating liquid
stream temperature.
- Pressure: Atmospheric
- Sulfuric Acid Concentration: Maintained at about
2% by weight.
o
• Particulate Loading: Inlet 0.2 g/Nm (0.1 gr/scf)
Outlet 0.02 g/Nm3 (0.01 gr/scf)
2.5 Process Efficiency and Reliability^ ' ' - Process efficiency is
dependent upon liquid-to-gas ratio used in the absorber, which
in turn determines the absorber packing height. Typical process
efficiency is about 95%, but efficiencies of over 99% have
been achieved. The capability for removing HpS, COS, CSp, HCN
and other possible species from coal gasification is not known.
Some problems which have occurred at the Gulf demonstration
facility required minor process redesign. Several plants are opera-
ting in Japan with greater than 99% reliability'8'.
2.6 Raw Material Requirements
• Catalyst: ferric sulfate solution can be any value up to
saturation^)
t Calcium Salt: 21.1 tonne/day (23.3 ton/day) based on use at
limestone (90% purity) for a 100 tonne/day Claus plant.
• Air: Quantity not known.
D-64
-------
2.7 Utility Requirements^ '' - (Based on 100 tonne per day Claus
plant).
t Electricity: 425 kwh/hr
• Fuel gas (9,000 kcal/m3, 1012 Btu/scf)
(4,256 Nm3/min. 158,750 scfm)
t Cooling water (5.6°K, 10°F Rise): 64.5 I/sec (1020 gpm)
• Steam (3.2 MPa, 470 psia saturated): 2470 kg/hr (5446 Ib/hr)
3.0 Process Advantages
• Continuous Stable Operation - No slurry is used in the
absorption and oxidation processes, thereby avoiding any scaling
or clogging problems!2).
e Special chemicals and utilities are not required
• Gypsum produced is of sufficiently good quality for use in wallboard^ '.
Gypsum shows good mechanical stability, not requiring stabilization
for landfill ing.
• Simple process flow results.in operational flexibility and lower
construction/operation cost^' '.
4.0 Process Limitations
t If the gypsum produced is not marketable, it must be transported
to a landfill(').
• Process offgas may have to be reheated before discharging into the
atmosphere depending on stack requirements(^>2).
• Relatively large packed absorber size required.
• Chloride levels in the absorbent solution must be controlled
below 200 ppm.
• Since the process requires the handling of sulfuric acid solutions,
special corrosion resistant metals are required^'/.
• Special corrosion resistant alloys are required (e.g., 316 LS.S for
lim'ngs
D-65
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5.0 Process Economics^ - (1972 dollars, Japanese yen basis*)
The following costs are based on a Chiyoda process applied to a boiler,
utilizing 2.7% sulfur fuel oil.
Design conditions:
Power generating capacity, MW 250 800
Flue gas volume, Nm3/hr (scfln) 750,000 2,400,000
(441,400) (1,413,000)
S02 in flue gas, ppm 1500 1500
Flue gas temperature, °K(°F) 413 (284) 413 (284)
Desulfurization rate, % 90+ 90+
Economics:
Capital investment $4,970,000 $11,850,000
Annual fixed cost (18% of $ 894,600 $ 2,133,000
capital)
Annual operating and $1,810,600 $ 4,779,800
maintenance
Overhead (12% of 0 & M) $ 109,900 $ 317,600
6.0 Input Streams
• Catalyst (Stream 12): See Section 2.6
t Calcium salt (Stream 13): See Section 2.6
• Make-up Water (Stream 4): See Section 2.6
7.0 Intermediate Streams^ '
• SO rich gas (Stream 1): see Table D-15
/\
t Prescrubber offgas (Stream 14), see Table D-15
• Absorber offgas (Stream 15), see Table D-15
dollar fl'9ures are low-
D-66
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TABLE D-15. TYPICAL CHIYODA PROCESS GAS STREAM COMPOSITIONS
'(1)
Components
(Vol %)
so2
co2
H20
"a
°2
Temperature
Pressure
Incinerated Claus Tail
Gas (Stream 1)
1.08
4.23
26.57
66.68
1:44
650°C (1200°F)
0.1 MPa 14.7 Psia)
Prescrubber
Off gas (Stream 14)
1.242
4.872
15.404
76.822
1.66
55°C (131°F)
0.1 MPa
(14.7 Psia)
Absorber
Off gas (Stream 15)
O.lOt
4.928
15.582
77.822
1.679
Of\
f\ / "I O T v r~ \
oo C (131 F)
0.1 MPa
(14.7 Psia)
Reheater
Vent Gas (Stream 3)
0.0865
5.31
15.571
77.061
1.972
316°C (600°F)
0.1 MPa
(14.7 Psia)
o
en
*Based on a 100-tonne per day Claus plant.
tSO? levels below 100 ppm can be achieved in coal fired boiler aoplicationv .
-------
8.0 Waste Streams"'
t Mother liquor purge stream (Stream 10) - Glaus tail gas (containing
about 33% water) is cooled, thereby condensing water, which must be
removed from the system. For a 100 tonne per day Claus unit, 0.148
I/sec (2.34 gpm) is purged with the following composition (in weight
percent):
H20 97.0%
H2S04 0.8%
MgO 2.2%
Fe2(SO.)3 (catalyst) trace
t Gypsum - (Stream 2) - moisture content of 5 to 20 percent. Gypsum
quality is dependent upon the impurities in the limestone feed.
• Reheater Vent gas (Stream 3) see Table D-15.
• Filtered solids (Stream 11) composition dependent on the character-
istics of the input gas.
9.0 Data Gaps and Limitations
• Cost data for Claus plant application are not available.
t Quantity of catalyst required is not known.
0 Detailed characterization data not available for all input,
intermediate and waste streams.
10.0 Related Studies: Not known.
D-68
-------
REFERENCES
1. Battelle Columbus Laboratories, Characterization of Sulfur Recovery
from Refinery Fuel Gas, U.S. EPA, NTIS No. PB-239-777, June 1974.
2. Beers, W. D., Characterization of Claus Plant Emissions, USEPA, NTIS No.
PB-220-376, April 1973.
3. Maddox. R. N., Gas and Liquid Sweetening, Campbell Petroleum
Series, 1974-
4. Siddiqi, A. A. and 0. W. Tenini, FGD - A Viable Alternative. Hydrocarbon
Processing, Houston, Texas, October 1977, pp.104-110.
5. DaRan, R. B., R. A. Edwards, and R. E- Rush, Interim Report on Chiyoda
Thoroughbred 101 Coal Application Plant at Gulf Power's Scholz Plant.
Presented in: Proceedings from Symposium on FGD, Vol. II, EPA-600/
2-76-136b, New Orleans, LA., May 1976, pp. 761-783.
6. Ando, J., Status of S02 and NOx Removal Systems in Japan, presented
at: Fourth FGD Symposium, EPA, Hollywood, Florida, November 8-11,
1977.
(
7. The Status of Flue Gas Desulfurization in the U. S.: A Technological
Assessment, The Federal Power Commission, Bureau of Power, July
1977.
8. Information provided to TRW by R. B. Dakan of Chiyoda International
Corp., March 10, 1978.
D-69
-------
SHELL COPPER OXIDE PROCESS
1.0 General Information
1.1 Operating Principles - A concentrated S02 gas stream is produced by
reaction ("adsorption") of sulfur oxides with CuO, followed by in-
situ regeneration using a reducing gas at approximately the same
temperature as S02 adsorption. The concentrated S02 stream is sent
to a Claus plant for sulfur recovery.
1.2 Development Status - Commercially available for oil refineries.
Pilot plant testing in coal-burning power plant.
1.3 Licensor/Developer - Developed by Shell International Petroleum,
The Hauge, Netherlands. Licensor:
Universal Oil Products Company
Des Plains, 111.
M 2}
1.4 Commercial Applications^ ' ' - One unit is currently in operation
at the Showa Yokkaichi Sekiyu in Japan, with a capacity of 2.8 x
105 Nm3/D (103 MMSCFD). The pilot plant at Tampa Electric
Company's (TECO) Big Bend Station has a capacity of 55.6 x 103 Nm3/D
(2.0 MMSCFD).
2.0 Process Information
2.1 Flow Diagram Pilot Plant^ ' (see Figure D-ll) - A raw gas stream is
heated to about 644°K (700°F) by heat exchange with the treated gas
followed by a trim burner for temperature control. The gas enters
the fixed-bed reactor containing CuO on alumina where it is "adsorbed"
D-70
-------
UUNGSTROM
EXCHANGER
ELECTRIC
SUPERHEATER
LEGEND:
1. Flue Gas
2. Steam
3. Hydrogen
4. Hydrogen Vent
5. Product Gas
6. Regeneration Off-gas
Figure D-ll.
Shell Flue Gas Desulfurization Process
for TECO Pilot Plant(2)
D-71
-------
on the dry bed. The principal reaction between SO,, in the flue
gas and copper activated alumina absorbent is:
S02 + 1/2 02 + CuO -> CuSO
The bulk of the accepted sulfur is released during the regeneration
cycle. Hydrogen is used for regeneration:
CuS04 + 2H2 -> Cu + S02 + 2H20
During the initial stages of adsorption, the Cu produced during
regeneration is oxidized:
Cu + 1/2 02 •*• CuO
Any CUpS present in the regenerated acceptor is oxidized:
Cu2S + 2 1/202 -»• CuO + CuS04
During regeneration the reactor is isolated from flue gas by flapper
valves and the gas flow is bypassed. Both treated gas and
regenerated gas are sent to the stack.
2.2 Flow Diagram - Yokkaichi Plantv ' (see Figure D-12) - Flue gas from
an oil-fired boiler at 673°K (752°F) containing 1300 ppm S02 flows
into one of the two adsorber reactors. Approximately 90% of the S0?
is absorbed as described in Section 2.1. After 90 minutes, the
flue gas is introduced into the second absorber, while the first
absorber is regenerated. The two adsorbers are alternated between
the acceptor and regeneration stages to allow continuous operation.
Because the SO^ released from regeneration will vary from nil to a
maximum every 90 minutes, an absorber-stripper system is utilized to
produce a constant flow to the Claus plant. About 99.5% of the SO
is absorbed (in water, under pressure) and then removed from the
absorber liquor in the stripper column.
D-72
-------
a
co
WASTE HEAT BOILER
Legend:
1. Flue Gas
2. Treated Flue Gas
3. Regeneration Gas (Hydrogen)
4. Absorber Offgas
5. Excess Stripper Water
6. Low Pressure Steam
7. S02 to Claus Plant
Figure D-12. Shell CuO Process (Yokkaichi Plant)
(9)
-------
2.3 By-Product - SCU rich gas. In a commercial plant, SO,, can be reduced
to elemental sulfur.
2.4 Equipment - Reactor fixed bed adsorber. Flue gas flows through
open channels alongside the acceptor. Designed by UOP.
2.5 Operating Parameters
• Adsorption temperature: 644°K-700°K (700°F-800°F)
t Regeneration temperature: 644°K-700°K (700°F-800°F)
2.6 Process Efficiency and Reliability^ ': 90% S02 removal efficiency.
No data for process reliability. The Yokkaichi plant experienced
the following operational problems: quench column corrosion,
sticking of the hydrogen line valve and plugging of waste heat
boiler tubes. These problems have been solved.
2.7 Raw Material Requirements - Hydrogen v ': 0.19 to 0.20 kg/kg of
S recovered.
2.8 Utility Requirements - ?
3.0 Process Advantages
• Dry process - handling of waste slurries not required.
• Acceptance and regeneration occur at the same temperature obviating any
heating or cooling of adsorption beds.
• Continuous processing can be achieved by using two units in alternating
acceptor c>nd regenerator modes.
• Process could be expanded for NOX removal by ammonia injection into
acceptor bed. CuO and CuS04 as catalysts for reduction of NOX
to nitrogen gas.
4.0 Disadvantages
• Equipment costs are high.
• A hydrogen source is needed for regeneration.
• Stripper requires steam, i.e., high energy inputs.
• The steam used in regeneration results in an acidic wastewater
stream. (In the pilot plant, the steam is vented to the stack
with the regeneration off-gas.)
D-74
-------
5.0 Process Economics
Capital cost of the Yokkaichi plant was $3.3 million (1974 dollars),
excluding the hydrogen plant^ .
6.0 Input Streams
• Flue Gas - Yokkaichi planter 673°K (752°F), 1300 ppm S02-
7.0 Intermediate Streams
No data reported.
8.0 Discharge Streams
(g\
• Excess Stripper Waterv ;: Contains 20 to 40 ppm (wt) of sulfur.
• Treated Flue Gas: 130 ppm S02 (based on 90% adsorption efficiency)
9.0 Data Gaps and Limitations
Comprehensive data are not published for either the pilot plant or
commercial facility.
10.0 Related Programs
None.
D-75
-------
REFERENCES
1. Kittrell, J. R. and Nigal Godley, Impact of SOX Emissions Control on
Petroleum Refining Industry, Vol. II, Appendix L,pp. 68-79, EPA 600/
2-76-161b, June 1976.
2. Anneson, A. D., F. M. Nooy, et al, The Shell FGD Process: Pilot Plant
Experience at Tampa Electric, paper presented at Fourth Symposium on Flue
Gas Desulfurization, EPA, November 1977.
3. Conser, E., Anderson, F., New Tool Combats S02 Emissions, Oil and Gas
Journal, pp. 81-86 (October 29, 1973).
4. Ploeg, J. E. G., Akagi, et al, How Shell's Flue Gas Desulfurization
Unit has Worked in Japan, Pet. Int. 14(7}, pp. 50-58, July 1974.
5. Pohlenz, J. B., The Shell Flue Gas Desulfurization Process, Flue Gas
Desulfurization Symposium, Atlanta, GA, November 4-7, 1974.
6. Vicari, F. A., J. B. Pohlenz, Energy Requirements for the Shell FGD
Process, Flue Gas Desulfurization Symposium, New Orleans, LA, March 8-11,
1976.
7. Dry Process for S09 Removal Due Test, Oil and Gas Journal, 67-70
(August 1972). L
8. Dry Scrubbing of Utility Emissions, Environmental Science and Technology,
9(8), 712-713 (August 1975).
9. Ando, J., et al, SO? Abatement for Stationary Sources in Japan,
EPA 600/2-76-013a, January 1976.
D-76
-------
LIME/LIMESTONE SLURRY SCRUBBING PROCESSES
1.0 General Information
1.1 Operating Principles - Sulfur dioxide absorption in a lime or lime-
stone slurry. The spent slurry is discharged to a settling pond
or thickener with the return of the clarified liquid to the scrubber
circuit.
1.2 Developmental Status - Commercially available.
1.3 Licensors/Developers^ ' ' - The engineering design of slurry
scrubbing systems for commercial installations is offered by a
number of companies including Babcock and Wilcox, Chemico, Com-
bustion Engineering, Peabody Engineering, Research Cottrell,
Universal Oil Products, and Zurn Air Systems.
1.4 Commercial Applications^ ' - There are currently 37 flue gas
desulfurization (FGD) units (7,441 MW total capacity) operating
in the United States on utility and industrial boilers, with lime -
or limestone scrubbing systems accounting for 82% of total operating
capacity. When units currently under construction or in the planning
stage are added to the present capacity, a total capacity of
50,419 megawatts (131 units) is projected of which 64% (by MH
capacity) will utilize lime or limestone slurries. In Japan there
are 333 operational FGD installations, 64 of which are lime/
limestone systems.
With the exceptions of the few applications to tail gases from
sulfuric acid and Claus plants, all .existing slurry scrubbing sys-
tems in the U.S. are applied to boiler gas streams. Although no
application to coal gasification currently exists, possible applica-
tions in a commercial gasification plant may be in connection with
support operations such as utility boiler and sulfur recovery.
D-77
-------
2.0 Process Information
2.1 Flow Diagram (see Figure D-13)
• Process Description^ '^ - The raw gas is treated in a venturi
scrubber for the removal of residual particulates and some S02
(up to 30%). Additional particulates and the bulk of the remain-
ing S02 are removed in an absorption tower where a slurry of lime
or limestone (generally 6Z-155K) is circulated. The integrated
scrubber/process absorber shown can achieve up to 95% S02 removal.
The slurry effluents from the venturi scrubber and absorber are
channeled into separate "reaction tanks" where the process reac-
tions are allowed to approach equilibrium. The overall reactions
are:
(a) with lime: Ca(OH)2 + S02, ^ CaSOg • 1/2 H20 + 1/2 H20
CaS03 + 1/2 02 ^CaS04
(b) with limestone: CaC03 + S02 + 1/2 H20 — CaS03*1/2 H20
CaS03 + 1/2 02~ CaS04
Most of the reaction tank slurries (containing precipitated
reaction products) are recirculated to the scrubber and absorber.
A slurry bleed stream is sent to a thickener for processing
and disposal. The process may also be designed to route the
slurry bleed stream directly to a disposal pond.
Residual parti cul ate removal can be achieved by utilizing a high
efficiency electrostatic precipitator or wet scrubber upstream
of the absorption tower.
2.2 Equipment - Conventional absorption towers (usually spray packed
towers), marble bed absorbers, venturi scrubbers, turbulent bed
absorber, stirred reaction tanks, and thickeners.
In addition to conventional absorbers, a number of newly developed
variations are now in service, such as the venturi rod scrubber,
and eggcrate (polygrid packed absorber) scrubber.
D-78
-------
•REMOVAL OF RESIDUAL PARTICULATES CAN B6
ACCOMPLISHED USING A HIGH EFFICIENCY
ELECTROSTATIC PRECIPITATOR OR FLOODED
DISC SCRUBBER,
0
LEGEND:
1. INLET GAS
2. THICKENER WASTE SLUDGE
3. THICKENER OVERFLOW
4. OUTLET GAS
5. POND RETURN WATER
6. LIME OR LIMESTONE
7. MAKEI.T WATER
8. VENTURI HOLD TANK EFFLUENT
9. REACTION TANK BLEED
10. VENTURI RECYCLE
11. ABSORBER RECYCLE
12. LIMESTONE SLURRY
Figure D-13. Wet lime/limestone process
(2)
D-79
-------
2.3 Feed Stream/Requirements^ '
• Temperature: 408°K to 443°K (274°F-338°F) typically
• Pressure: atmospheric
• Loading: 400 to 5000 ppm S02; 0.01-9.2 g/Nm3 (0.0041-4.0 gr/scf)
particulatesO4)
• Other: system design favors gases produced from low chloride
coals (when coal is burned).
2.4 Operating Parameters^3' '
• Venturi Scribber
- Temperature: flue gas temperature in venturi section,
flue gas saturation temperature in the separation section
- Pressure: atmospheric
- Pressure drop: 30-230 mm H20 (1.2-9.1 in. H20)
3
- Solution circulation rate: 2-10 liters/Mm
(14-71 gal/1000 scf)
• Absorber^3'14)
- Temperature: flue gas saturation temperature
- Pressure drop: 35-505 mm hLO (1.4-19.9 in. FLO)
3
- Solution circulation rate: 5-12 liters/Nm
(35.5-85 gal/1000 scf)
- Slurry concentration: 5%-15%
2.5 Process Efficiency and Reliability - S00 removal efficiencies up to
(2)
95% can be expectedv ', although generally efficiencies of 7Q%-90%
are reported based on utility firing of high sulfur coal. For low
sulfur coals, at least 50% of the sulfur dioxide can be removed^.
Fly ash removal efficiencies of 98% are typical for the integrated
scrubbing-absorption system( '. Process reliability may be a weak
point in slurry-based systems. Availabilities have been reported
D-80
-------
from 70%(5'6) up to 93% for the Kansas City Power and Light,
LaCygne Unit No. 1*( '. In Japan, numbers up to 100% have been
reported^3'.
2.6 Raw Material Requirements
• Lime/limestone: Stoichiometries of 1.0-1.05 moles of CaO/mole
S02 absorbed are reported in Japan(3). Limestone Stoichiometries
of 1.0-1.5 are required for the Kansas Power and Liaht
installation04).
t Fixation Agents: Patented materials are available for fixating
the sludge. One such material is Calcilox by Dravo(8).
• Air(3'4}9'10'14): 350%-600% of stoichiometric quantities used
in many Japanese slurry scrubbing systems to oxidize calcium
sulfite sludge to gypsum by-product. Also, forced oxidation
to gypsum is used increasingly in American systems to enhance
solids settling, dewatering and storage properties and requires
150%-400% of stiochiometric quantitiesO2*).
• Make-up Water: The quantity of water required for closed loop
operation is determined by evaporation losses and the amount
discharged with the waste sludge streamO4/. 390 l/min (103 gpm)
required for Kansas Power and Light installation.
2.7 Utility Requirements
• Steam: Usually none; may be used to reheat outlet flue gas.
• Fuel (for reheat): determined by flue gas flow rate and
reheat required.
• Electricity'3': 1.2% to 2.1% of station power generated.
t Total Energy^ ': 2%-5% of total station generation and includes
pump, fan and reheat energy requirements.
2.8 Miscellaneous'3^ - Scaling and corrosion problems can result from
improper design and operation, and are often controlled during
scheduled shutdowns to minimize unscheduled downtime.
*Availability reflects the percentage of time when the boiler is operating
and the scrubber system is available for operation. Slurry scrubbers in the
U.S. generally have greater down times than utility boilers.
D-81
-------
(2)
3.0 Process Advantages
• The basic process is fairly simple and very few process steps are
involved.
• The capital and operating costs are relatively low. Reserves of
absorbent materials are abundant in the United States.
• S02 removal efficiencies are generally high.
t The two-stage treatment of flue gases allows for the removal of
both S02 and the residual particulates.
• The lime/limestone process is the most commonly used S02 control
method by utilities exclusive of low-sulfur fuel. Commercial
installations have been operating for more than four years.
to)
4.0 Process Limitationsv '
• Large quantities of waste sludge require processing and disposal in
an environmentally acceptable manner.
• If not designed carefully or operated attentively, lime/limestone
systems have a tendency towards chemical scaling, plugging, and
erosion. These problems can frequently halt operation of the
system.
• Oxidation of sulfite to sulfate 'increases the tendency towards
serious scaling. Excess air, high pH, fly ash, residence time in the
reaction tank, and the presence of N0£ in the flue gas are suspected
to be factors which contribute to oxidation. Scaling can be reduced
by forcing the oxidation completely to gypsum.
• Efficiency of S02 removal decreases with decreasing sulfur content
of fuel.
5.0 Process Economics
Commonwealth Edison reported the cost of the 160 MW retrofitted lime-
stone scrubbing installation at $95/kw (1972 dollars) at its Will
County (Illinois) Station. In addition, an expenditure of $13/kw is
required for sludge treatment and disposal. Operating costs were
estimated at 2.8 mils/1000 kcal ($0.70/MM Btu), coal fired at 6Q% load
factor or 7.3 mils/kwh, including 2.1 mils/kwh for sludge treatment
and disposal^ '.
D-82
-------
79.9
88.4
61.1
68.4
44.9
51.4
TVA updated cost estimates for lime and limestone srubbing processes
designed to remove 90% of S02 from utility gas fired with 3.5% sulfur coal
follow. On-site sludge disposal and 7,000 hrs/yr operation are assumed
for new facilities. Sludge fixation costs are excluded. Cost
basis: mid-1977^.
Capital Investment, $/kw
200 MW 500 MW 1000 MW
Lime
Limestone
Operating Costs, Mils/kwh
200 MW 500 MW 1000 MW
Lime 4.54 3.65 2.94
Limestone 4.20 3.41 2.74
Sludge fixation would add 151-20% to the annual operating cost .
6.0 Input Streams
0 Inlet gas (Stream 1) - see Table D-16.
• Lime or limestone (Stream 6) - see Section 2.6.
• Make-up water (Stream 7) - see Section 2.6.
7.0 Intermediate Streams
• Limestone slurry (Stream 12).
• Venturi hold tank bleed (Stream 8) - see Table D-17.
• Reaction tank bleed (Stream 9) - see Table D-18.
• Venturi recycle (Stream 10) - same composition as Intermediate
Stream 8 - see Tables D-17 and D-19.
t Absorber recycle (Stream 11) - same composition as Intermediate
Stream 9 - see Tables D-18 and D-20.
D-83
-------
TABLE D-16. PROPERTIES OF FEED GAS TO LIME/LIMESTONE S02 SCRUBBERS (STREAM 1)
Temperature
Flow Rate
Particulate
so2
Particulate Ash
Analysis
P2°5
Si02
FeO
A1203
CaO
MgO
so3
Na20
TiO
Other
Kentucky Utility Green River ^ '
Lime Scrubber
422°K (300°F)
611,280 m3/hr (360,000 ACFM)
382,466 Nm3/hr (238,000
SCFM dry)
5.3 g/Nm3 dry (2.2 gr/SCF
dry)
49.4 kg/min (108.9 Ib/min)
45,239 liters/min (11,968 gpm)
—
— —
—
::.
—
—
—
Kansas Power and Light
Lawrence No. 4(10)
Limestone Scrubber
411°K (280°F)
684,294 m3/hr
(408,000 ACFM)
7.25 g/Nm3 dry
(3 gr/SCF dry)
748 ppm
--
— -
--
_ _
—
—
—
Kansas City Power and Light
La Cygne Unit No. id 2)
Limestone Scrubber
411°K (280°F)
4,686,480 m3/hr
(2,760,000 ACFM
total , 7 scrubbers)
17.8 g/1000 kcal
(9.9 Ib/MM Btu)
5000 - 5700 ppm
0.15
46.1
19.2
14.1
6.9
1.0
7.9
2.5
0.6
1.0
0.7
o
00
-------
TABLE D-17.
VENTURI SCRUBBER SLURRY SLOWDOWN AT KANSAS POWER
AND LIGHT LAWRENCE NO. 4(10) - LIMESTONE
SCRUBBER (STREAM 8)
Parameter
Flow Rate
Solids Cone*
Dissolved Ions*
Ca++
so3=
so4=
Solids Composition*
CaS03 • 0.5 H20
CaS04 • 2 H20
CaCOo
CaSO. Relative Saturation*1"
Value
3037 kg/hr (6695 Ib/hr)
473 liter/min (125 gpm)
9-11%
876 ppm
137 ppm
106 ppm
2,340 ppm
2.41 wt %
11.57 wt 5
5.85 wt %
1.45
*At 100% limestone feed stoichiometry.
"""•Relative saturation of 1.45 indicates a calcium sulfate supersaturation
of 45% under certain conditions.
D-85
-------
TABLE D-18. REACTION TANK SLOWDOWN AT KANSAS POWER AND LIGHT
LAWRENCE NO. 4U°) - LIMESTONE SCRUBBER
(STREAM 9)
Flow Rate
Solids Cone*
Dissolved Species*
Ca++
Mg++
so3=
so4=
Solids Composition*
CaS03 • 0.5 H20
CaS04 • 2 H20
CaOL
CaS04 Relative Saturation*t
957 kg/hr (2110 Ib/hr)
151.4 liters/min (40 gpm)
5 - 7%
715 ppm
127 ppm
23 ppm
2,064 ppm
0.20 wt %
19.25 wt %
21.52 wt %
1.22
*At 100% limestone feed stoichiometry.
"^Relative saturation of 1.22 indicates a calcium sulfate supersaturation
of 22% under certain conditions.
D-86
-------
TABLE D-19. VENTURI SCRUBBER RECYCLE LIME/LIMESTONE SCRUBBER (STREAM 10)
Kansas Power and Light
Lawrence No. 400)
Limestone Scrubber
Kentucky Utility
Green River (7)
Lime Scrubber
Flow Rate
Ca(OH),
CaxSOx
Total
Solution
Circulation Rate
13,626 liters/min
(3,600 gpm)
2.67 Uters/nT
(20 gal/1000 ACF)
44,663 liters/min
(11,800 gpm)
49.4 kg/min
(109 Ib/min)
3348 kg/min
(7,380 Ib/min)
*At 100% limestone feed stoichiometry.
TABLE D-20. ABSORBER RECYCLE LIMESTONE SCRUBBER (STREAM 11)
Kansas Power and Light
Lawrence No. 4(1)
Limestone Scrubber
Northern States Power
Sherburne County (4)
Limestone Scrubber
Flow Rate
L/G
PH
Dissolved Calcium
Dissolved Magnesium
Dissolved Sulfate
Dissolved Sulfite
Solid Calcium
Solid Magnesium
Solid Sulfate
Solid Sulfite
20,061 liters/min
(5,300 gpm)
4.0 liters/m3
(30 gal/1000 ACF)
5 - 5.5
500-700 ppm
1,500-2,500 ppm
8,000-15,000 ppm
0
1 - 10% (wt)
0.5 - 1.5%
15 - 20%
0
D-87
-------
8.0 Discharge Streams
t Thickener waste sludge (Stream 2)
Kentucky Utility - Green River ' Lime Scrubber
Flow Rate
HJD: 863 liters/min (228 gpm)
Ca(OH)2: 4.1 kg/min (9.0 Ib/min)
CaSO : 86.2 kg/min (190 Ib/min)
A
• Outlet gas (Stream 4) - see Table D-21.
• Thickener overflow (Stream 3) - no data available.
0 Pond return water (Stream 5) - see Table D-22.
9.0 Data Gaps and Limitations
Although numerous articles have been published describing operation
of lime and limestone slurry scrubbing processes, full stream character-
ization data are usually unavailable for a given full-scale operating
plant. Notably, gas composition data, liquid stream comprehensive
trace element analysis data, flue gas reheat fuel requirements and
realistic capital and operating costs are unavailable.
10.0 Related Programs
To minimize the chemical limitations of lime and limestone systems,
efforts have been made to improve the process by the use of magnesium
additives. Research conducted at the bench-scale and pilot plant level
has stimulated further work on prototype facilities (EPA/TVA Alkali
Scrubbing Test Facility, Shawnee Station), and at the demonstration '
level (EPA Scrubber/Sludge Evaluation Program, Paddys Run Station, ;
Louisville Gas and Electric). Two proprietary absorbents have been
developed, one by Dravo (thiosorbic lime), the other by Pullman Kellogg
(catalytic limestone)^ •
D-88
-------
TABLE D-21. OUTLET GAS - LIME/LIMESTONE SCRUBBER (STREAM 4)
Temperature
Flow Rate
Parti cul ate
so2
H20
Opaci ty
Kentucky Utility
Green River (7)
Lime Scrubber
320°K (116°F)
514,494 m3/hr
(303,000 ACFM saturated)
382,466 Nm3/hr
(238,000 SCFM dry)
0.106 g/Nm3 dry
(0.044 gr/SCF dry)
9.9 kg/min (21.8 Ib/min)
636 liters/min (168 gpm)
—
Kansas Power and Light
Lawrence No. 4(10)
Limestone Scrubber
336°K (144°F)
616,374 m3/hr
(363,000 ACFM)
0.053-0.094 g/Nm3 dry
(0.022-0.039 gr/SCF dry)
200 ppm
--
2.5 - 7.5%
TABLE D-22. POND RETURN WATER - LIMESTONE SCRUBBER (STREAM 5)
Kansas City Power and Light
La Cygne Unit No. 1(12)
Limestone Scrubber
Calcium
Magnesium
Sodium
Potassium
Bicarbonate
Chi ori de
SuTfate
Sulfite
Silica
PH
Conductivity
696 ppm
48 ppm
22 ppm
23 ppm
36.6 ppm
177.8 ppm
1627 ppm
Not detected
20.6 ppm
7.0
4380 micromhos
D-89
-------
Two pilot plants, sponsored by the EPA, are actively involved in forced
oxidation test programs to enhance solids settling properties, to decrease
sludge disposal land requirements, and to improve the quality of recycled
water. They are: TVA/EPA Alkali Scrubbing Test Facility, Shawnee No. 10;
(9)
and EPA/IERL pilot plant, Research Triangle Parkv '.
Two major suppliers are involved in chemical fixation of scrubber
wastes: The Dravo Corporation and U Conversion Systems
REFERENCES
1. Siddiqi, A. A. and J. W. Tenini. FGD-A Viable Alternative. Hydrocarbon
Processing, Houston, Texas, Oct. 1977, pp 104-110.
2. The Status of Flue Gas Desulfurization Applications in the U.S.: A
Technical Assessment. The Federal Power Commission Bureau of Power,
July 1977.
3. Ando, Jumpei. Status of S02 and NOX Removal Systems in Japan.
Presented at Seventh FGD Symposium, EPA, Hollywood, Florida,
November 1977, 21 pp.
4. Kruger, R. J. Experience with Limestone Scrubbing - Sherburne County
Generating Plant Northern States Power Co. Presented at Fourth FGD
Symposium, EPA, Hollywood, Florida, November 8-11, 1977. 27 pp.
5. Stober, W. G. Operational Status and Performance of the Commonwealth
Edison Will County Limestone Scrubber. Presented in Proceedings from
Symposium on FGD, Vol. I, EPA-600-2-76-136a,."New Orleans, La., May
1976, pp. 219-248.
6. Knight, G. R. and S. L. Pernic R., Jr. Duquesne Light Co. Elrama
and Phillips Power Stations Lime Scrubbing Facilities. Presented
in Proceedings from Symposium on FGD, Vol. I, EPA-600-2-76-136a,
New Orleans, La., May 1976, pp 205-218.
7. Beard, J. B. Scrubber Experience at the Kentucky Utilities Co.
Green River Power Station. Presented at Fourth FGD Symposium, EPA,
Hollywood, Florida, November 1977, 8 pp.
8. Workman, K. H. Operating Experience - Bruce Mansfield Plant Flue Gas
Desulfurization System. Presented at Fourth FGD Symposium, EPA,
Hollywood, Florida, November 1977, 5 pp.
9. Laseke, B. A. and T. W. Devitt. Status of Flue Gas Desulfurization
Systems in the United States. Presented at Fourth FGD Symposium,
EPA, Hollywood, Florida, November 1977. 35 pp.
D-90
-------
10. Green, K. and 0. R. Martin, Conversion of the Lawrence No. 4 Flue Gas
Desulfurization System. Presented at Fourth FGD Symposium, EPA,
Hollywood, Florida, November 1977, 21 pp.
11. Ring, T. A. and J. M. Fox. Stack Gas Cleanup Progress. Hydrocarbon
Processing, 119-121, October 1974.
12. McDaniel, C. F. La Cygne Stations No. 1 Wet Scrubber Operating
Experience. Presented in Proceedings from Symposium on FGD, Vol. I,
EPA-600-2-76-136a, New Orleans, La., May 1976, pp 355-372.
13. LaSeke, Bernard A. Jr., PEDCo Environmental, Inc., EPA Utility FGD
Survey: December 1977-January 1978.
14. Information provided to TRW by the technical staff of EPA's Industrial
Environmental Research Laboratory (RTP), June 1978.
D-9,1
-------
DUAL ALKALI PROCESS
1.0 General Information
1.1 Operating Principles - Sulfur dioxide removed by scrubbing in a
liquid-vapor absorption tower using a clear, concentrated sodium
sulfite absorbent solution to form sodium bisulfite. The sodium
bisulfite solution is reacted with lime in a separate vessel to
precipitate calcium salts and regenerate sodium sulfite which is
returned to the scrubber.
1.2 Developmental Status - Commercially available although unproven on
a commercial scale. Systems have been demonstrated on utility and
industrial coal-fired boilers up to a maximum capacity of 32 MW
in the U.S. The first full-scale application is presently under
construction at Louisville Gas and Electric.
1.3 Licensor/Developer - There are various developers of the basic
double alkali process, each utilizing their own patented ideas.
The developers offering the most fully developed commercial processes
are FMC, Envirotech and Arthur D. Little/Combustion Equipment
Associates.
FMC Corporation
Environmental Equipment Division
1800 FMC Drive West
Itasca, Illinois 60143
Envirotech Corporation
Eimco BSP 669
W. 2nd South
Salt Lake City, Utah 84110
Arthur D. Little/Combustion Equipment Associates, Inc.*
555 Madison Avenue
New York, N. Y. 10022
*Licensed under Combustion Equipment Associates.
D-92
-------
1.4 Commercial Applications - A dilute mode, 32 MW systems was started
up in March 1974 and tested by General Motors/Koch on coal-fired
6 M boilers in Parma, Ohicr1*2'.
Arthur D. Little (ADL) and Combustion Equipment Associates (CEA)
successfully completed testing of a 20 MW prototype at Gulf Power
Company's Scholz Station in Sneads, Florida, a coal-fired facility.
A concentrated mode system was started up in February 1975^'4^.
ADL/CEA presently has under construction a 277 MW demonstration sys-
tem on coal-fired boilers at Louisville Gas and Electric Power Co.
The system is scheduled for service in 1979^3'5'.
A 250 MW system by FMC and a 575 MW system by Buell/Envirotech are
in the planning stages. The FMC system is scheduled for service in
1979 at Southern Indiana Gas and Electric Co. FMC pilot plants have
operated on stoker boilers, sulfuric acid plants, and on a dual-
fired bark oil/coal boiler. In addition, an FMC industrial scale
system has operated successfully on a chemical kiln at Modesto,
California(3'6'7'8).
Full-scale utility applications of dual alkali systems ranging in
(9)
size from 150 to 450 MW are in operation in Japanv '. Kureha-
Kawaski, Showa Denro and Tsukishima have a number of systems operat-
ing on oil-fired boiler flue gases as listed below^ ' :
Number of
Process Supplier Absorbent By-Product Units
Kureha-Kawasaki Sodium sulfite, Gypsum 3
limestone
Showa Denro-Ebara Sodium sulfite, Gypsum 16
limestone
Tsukishing Sodium sulfite, Gypsum 4
lime
2.0 Process Information
2.1 Flow Diagram - See Figure D-14.
t Process Description(4'7): The process consists^ three major
sections: absorption, regeneration and dewatenng. Hot fjue_gas
(Stream 1) enters the absorber countercurrent to a clear liquid
D-93
-------
s.^
If
ABSORBER
J I
MIXING
TANK
1 1
REGENERATION
REACTOR
SYSTEM
1. FLUE GAS FEED
2. SCRUBBER EFFLUENT
3. SCRUBBED FLUE GAS
4. LIME SLURRY MAKEUP WATER
5. LIME
6. FILTER WASH WATER
7. FILTER CAKE
8. SODA ASH
9. MAKEUP WATER
10. PURGE STREAM
11. SCRUBBED ABSORBENT FEED '
12. REACTOR SLURRY
13. THICKENER UNDERFLOW
14. THICKENER OVERFLOW
15. FILTRATE
Figure D-14. Dual Alkali Scrubbing with Lime Regeneration
(ID
-------
fresh absorber feed solution consists of a mixture of sodium
suime' S0d1um "
S02 is absorbed by the active sodium constituents,* converting them
to sodium bisulfite (Reactions 1, 2 and 3 below). Some oxidation of
sodium sulfite to inactive sulfate also occurs in the scrubber
(Reaction 4):
Absorber Reacti
(1) Na2C03
(2) NaOH +
(3) Na2S03
(4) Na9SO,
ons
+ 2S02
so2 ->
+ so2 •
+ 0.5 i
NaHS03
2NaHS03
2NaHS03
A bleed stream from the scrubber hold tank (Stream 2) is pumped to a
regeneration reactor where lime is added and calcium salts are pre-
cipitated by the following reactions:
Regeneration Reactions
(5) Ca(OH)2 + 2NaHS03 ->• CaS03 • 0.5H2Oi + Na2S03 + 1.5 H20
(6) Ca(OH)2 + Na2S03 + 0.5H20 -> CaS03 • 0.5H2(K + 2NaOH
(7) Ca(OH)2 + Na2S04 + 2H20 -*• CaS04 • 2H2Oi + 2NaOH
Sodium sulfate produced in the absorber is precipitated according
to Reaction 7.
Slurry from the reactor (Stream 12) is pumped to a thickener where
calcium salts are concentrated (Stream 13) and pumped to a rotary
vacuum filter. A waste cake is produced (Stream 7), and some of the
soluble sodium in the cake is washed back into the system (Stream 6).
Filtrate (Stream 15) is combined with clear overflow solution from
the thickener (Stream 14) and returned to the scrubber.
*Active sodium refers to sodium derived principally from NaOH, Na^SOs,
Na?C03 and NaHC03, as opposed to inactive forms derived from NaCl or NagS
When active Na+ is less than 0.15M, the system is considered to be in the
dilute mode.
D-95
-------
In some versions of the dual alkali process an additional liquid purge
stream (Stream 10) is required to control sulfate and chloride
buildup.
2.2 Equipment^'8'12^ - All equipment is conventional and includes
venturi scrubbers - dual throat and variable throat types - sieve
tray towers, FMC patented disc contactor towers, packed towers,
agitated reactors, thickener with circulating rake, rotary vacuum
filters, oil or steam type reheaters and demisters.
2.3 Field Stream Requirements
t Temperature^'1 2): 408°K to 561°K (275°F-550°F)
• Pressure: atmospheric
• S02 Loading (dry basis)
- Concentrated mode*(8'12): 1800-8000 ppm
- Dilute mode^13': 250-400 ppm, up to 1500 ppm
• Particulate Loading
- With venturi : 7.25 g/Nm3 dry (3 gr/scf dry)
- Without venturi^8': 0.048 g/Nm3 dry (0.02 gr/scf dry)
Contaminant Limitations
- Oxygen (dry basis)
Concentrated mode^ ' ': 6.5/6-7.6% maximum
Dilute mode: No upper limit; system favors high oxygen
concentration.
- Chloride ion^ ' ': Harmful to equipment, not process. Suc-
cessful prototype operation demonstrated with 0.5-0.10 weight
'% Cl coal , dry basis.
Concentrated mode dual alkali systems are not normally designed for flue gas
streams with S02 concentrations less than about 1500 ppm or 02 concentrations
greater than about 7% because of the resulting high oxidation rates of sul-
fite to sulfate in the scrubber. Dilute mode systems are normally designed
for low sulfur coals (S02 < 1500 ppm) because these systems regenerate sodium
sulfate and make gypsum more efficiently.
'^Assuming an electrostatic precipitator is used.
D-96
-------
- Potassium ion' ': ?
- Fluoride ion- ': ?
2.4 Operating Parameters
• Absorption System
- Temperature : Flue gas saturation temperature; for boiler
flue gas, 328<>K (13QOF) is typical
- Pressure: Atmospheric
- Loading: Dilute mode(13) - L/G of 2.67 liters/m3
(20 gal/1000 ACF)
Concentrated mode^8'12^ - L/G of 1.82 liters/m3
(13.6 gal/1000 ACF), dual throat venturi; 0.67 -
1.34 liters/nP (5-10 gal/1000 ACF), tray tower
- pH(8'12): 4.8 - 7.0
0 Regneration System
- Temperature: Same as absorber
- Pressure: Atmsopheric
- FMC Single Reactor (Concentrated Mode)
nU\'/. Q C
pn . o.o
(12)
Residence timev ': 5 minutes
- ADL/CEA Two Reactor System (Concentrated Mode)
11.0 - 12.5
Residence time' ': 5 minutes - Reactor 1
35 minutes - Reactor 2
2.5 Process Efficiency and Reliability^8'14' - Concentrated, mode dual
alkali systems have demonstrated S02 removal capabilities over 99%
for typical flue gas streams from coal-fired utilities. Lime uti-
lization has ranged from 95%-100% based on one mole Ca(OH)2/mole S02
removed.
D-97
-------
A 3-MW prototype of the FMC process has had 94% availability for the
initial year of operation, and a 50-MW commercial system started up
in October 1975 and also had a high availability since that time.
A 20-MW ADL/CEA prototype logged 78% availability over a 17 month
test period.
2.6 Raw Material Requirement
• Makeup Chemicals(8>9'12'13)
Na«CO,: 0.98-1.0 moles Ca/mole sulfur removed - concentrated
mode; 1.0-1.1 Ca/S - dilute mode, generally. 1.4 to
1.65 Ca/mole sulfur dilite mode system actual operat-
ing data.
« Makeup Water: Depends heavily on type of system. System with
separate particulate scrubber requires additional water. System
for liquid purge for sulfate control (FMC) uses more water than
system without sulfate purge.
• Air: Used only in dilute mode systems to oxidize all sodium
sulfate to sulfate prior to regeneration by sparging.
2.7 Utility Requirements
t Steam: None typically, but can be used to reheat flue gas in
lieu of fuel oil.
(4)
t Fuel Oilv ': 1.4% - 2.1% of energy input to a power generation
station.
t Electricity^4'6'8^: 1.3% - 2% of utility station generation for
concentrated mode, tray tower system removing 95% of S02 produced
from burning 3-4 wt % sulfur coal with no particulate removal.
2.5% - 3.0% of station generation for system removing both S0?
and particulates. ^
3.0 Process Advantages' '
• Capital and operating costs are relatively low. The process utilizes
conventional chemical processing equipment and materials required are
commonly used and available.
• Very high S02 removal efficiencies can be obtained. Tray towers can
be designed for insertion or removal of extra trays to adjust to
changing performance requirements.
t The soluble product in the absorber minimizes solids buildup and
erosion problems.
D-98
-------
• Process can simultaneously remove particulates and SO .
• A low L/G ratio is featured by the scrubber.
t Corrosion and erosion problems are minor compared to those in wet
lime/limestone processes.
• High fly ash contents can be tolerated in the system.
4.0 Process Limitations
(15)
• Large quantities of waste calcium sulfite and calcium sulfate salts
containing soluble sodium must be disposed of.
• Design complexities must be introduced to deal with the following
problems:
1) Excessive purge of Na2S04 produced as a result of oxidation (the
ADL/CEA process does not require a purge stream when high sulfur
coal is burned and less than 40% excess air is used in combustion
of coal).
2) Clean scrubbing liquor saturated with calcium sulfate. Excessively
high levels of calcium sulfate could lead to scaling problems.
(Concentrated mode systems do not scale as long as there is proper
pH control.)
• Requires some makeup to replenish sodium losses.
• Problems of pitting and corrosion due to chloride buildup. Special
coatings and linings and/or higher grade alloys must be used.
• Generates predominantly calcium sulfite solids - a material that does
not occur naturally in nature, is thixotropic and has a high potential
COD.
5.0 Process Economics
• General Motors Dilute Mode System^1^: $3.5 million (1974) capital
investment for "first-of-a-kind" industrial boiler system equivalent
to 32 MW.
Unit capital cost: $88/kw
Operating costs: Not available
• FMC Concentrated Mode(12): Estimated cost for 150-MW utility boiler
system is $6 million (1974). Unit capital cost: $40/kw
• EPA Estimates(9): Based on concentrated mode, dual alkali system,
coal-fired power plant flue gas, 90% S02 removal, s^a ash makeup,
new 200 MW system, 80% load factor, throwaway CaS03/CaS04 salts
D-99
-------
Unit capital cost: $50-60/kw (1974 dollars)
Operating cost: 2.5-3 mils/knh
Actual costs are not available for full scale applications.
6.0 Input Streams
• Flue Gas (Stream 1): see Table D-23.
• Lime Slurry Makeup Water (Stream 4): Quantity depends on particular
system water balance as determined by inlet flue gas temperature,
sulfur dioxide concentration, waste cake moisture and wash require-
ments, sulfate and chloride purge requirements, demisters and pump .
seals.
• Lime (Stream 5): see Section 2.6.
0 Filter Wash Water (Stream 6): Flow rate depends on system design
requirements. Wash ratios of 3:1 or less are normal(').
t Soda Ash Makeup (Stream 8): see Section 2.6.
• System Makeup Water (Stream 9): Rate depends on system with balance.
7.0 Intermediate Streams
• Scrubber Effluent (Stream 2): see Table D-24.
t Scrubber Regenerated Absorbent Feed (Stream 11): see Table D-25.
• Reactor Slurry (Stream 12): see Table D-26.
• Thickener Underflow Slurry (Stream 13):
ADL/CEA-Gulf Power
' Prototype^)
Solids Concentration less than 30 wt %
• Thickener Overflow (Stream 14): see Table D-27.
• Filtrate (Stream 15): No data available.
8.0 Discharge Streams
• Flue Gas Outlet (Stream 3): see Table D-28.
• Filter Cake (Stream 7): see Table D-29.
• Liquid Purge (Stream 10): The only commercially available double
alkali process in the U.S. requiring liquid purge is the FMC process
in some applications. Characteristics are the same as for Inter-
mediate Stream 11. See Table 0^25. No specific data available.
D-100
-------
TABLE D-23. CHARACTERISTICS OF FLUE GAS FEED TO DUAL ALKALI PROCESS
Parameter/
Concentration
Temperature
Flow Rate
so2
Parti cul ate
Composition
°2
N2
co2
H2
so2
General Motors -.Parma, Ohio
Demonstrator (12)
450°K (350°F) - two large
boilers
561 °K (550°F) - two small
boi 1 ers
101,880 m3/hr (60,000 acfm)
per large boiler
84,900 m3/hr (50,000 acfm-)
per small boiler
900-1600 ppm, range
1200-1300 ppm, average
0.725 g/Nm3 (0.3 gr/scf)
—
--
ADL/CEA
Gulf Power
Prototype (8)
4080K (275°F)
127,350 m3/hr
(75,000 acfm)
1800-3800 ppm,
dry
0.048 g/Nm3
dry (0.02 gr/
scf dry) with
precipitator
energized
6 . 5% max . ,
dry
:
FMC Pilot
Plant(2)
478°K (400°F)
4,377 m3/hr
(2,578 acfm)
3363 ppm
5.8 g/Nm3
(2.4 gr/scf)
7.6%
76.3%
11.4%
4.7%
D-101
-------
TABLE D-24. CHARACTERISTICS OF SCRUBBER EFFLUENT IN DUAL ALKALI PROCESS
Parameter/
Constituent
Temperature
PH
Chloride ion (Cl~)
Potassium ion (K+)
Fluoride ion (F~)
Total sodium (Na+)
Other non-Na, K,
Ca metals
Active alkali
Sulfate (S04=)
Hydroxide (OH-)
Bisulfite (HS03~)
Calcium (Ca++)
General Motors
Parma, Ohio
Demonstrator (12)
—
5.5 - 6.0
--
--
--
--
—
—
0.35M
Trace
0.03M
300 - 400 ppm
ADL/CEA
Gulf Power
Prototype (4»8)
--
4.8 - 6.0
13,000 ppm, max.
300 - 1300 ppm
70 ppm, max.
--
<1 ppm each
--
--
--
__
--
FMC Pilot.
Plant^7'12'
328°K (130°F)
6 - 7
__
--
--
>2M
—
>0.5M
—
--
--
—
TABLE D-25. CHARACTERISTICS OF REGENERATED ABSORBENT IN DUAL ALKALI PROCESS
General Motors - Parma, Ohio
Parameter/Constituent Demonstrator^)
PH
Sulfate (S04=)
Hydroxide (OH")
Bisulfite (HSQg
Calcium (Ca++)
9.0
0.1M
Trace
300 - 400 ppm
D-102
-------
TABLE D-26. CHARACTERISTICS OF REACTOR SLURRY IN DUAL ALKALI PROCESS
Parameter/Cons ti tuent
ADL/CEA-Gulf
Power
Prototype (4,8)
FMC Pilot
No. Reactors in Series
Residence Time
• Reactor 1
• Reactor 2
pH
Solids Concentration
Flow Rate
3-5 min.
30-40 min.
11.0 to 12.5
Up to 5%
700 liters/min
(185 gpm)
5 min.
Not applicable
8.5
TABLE D-27. CHARACTERISTICS OF THICKENER OVERFLOW IN DUAL ALKALI PROCESS
Parameter/Consti tuent
ADL/CEA-Gulf Power
Prototype
pH
Active Sodium (Na+)
Sulfate (S04=)
Chloride ion (Cl~)
Calcium ion (Ca++)
11 - 12.5
0.2 - 0.6M
0.6 - 1.05M
4000 - 5000 ppm
50 - 200 ppm
D-103
-------
TABLE D-28. CHARACTERISTICS OF FLUE GAS IN DUAL ALKALI PROCESS
Parameter/
Constituent
Temperature
Flow Rate
so2
Solids
Entrainment
Liquid
Entrainment
Sodium
Entrainment
General Motors
Parma, Ohio
Demonstrator02)
Saturation temp.
--
20 - 200 ppm
--
"
ADL/CEA
Gulf Power
Prototype^)
Saturation temp.
0.0085 g/Nm3
(0.0035 gr/scf dry)
0.060 g/Nm3
(0.025 gr/scf dry)
0.0048 g/Nm3
(0.002 gr/scf dry)
FMC Pilot Plant'12'
Saturation temp.
343 ppm
—
"
D-104
-------
TABLE D-29. CHARACTERISTICS OF FILTER CAKE IN DUAL ALKALI PROCESS
Parameter/
Constituent
General Motors-Parma, Ohio
Demonstratorv12)
FMC Pilot Plant
(12)
ADL/CEA
Gulf Power Prototype
CJ
Ca(OH)2
Fly Ash
Solubles as Na^SO.
CaSO
A
Moisture
Wash Efficiency
CaS03- 0.5 H20
CaC03
Insoluble Solids
Soluble Solids
Total Solids
CaS04/CaSOx
(molar)
10-20%, dry basis
1-2%, dry basis
4-5%, dry basis
Remainder of dry cake
50%
20% reduction of
solubles
-48%
13.95%, wet basis
1.18%, wet basis
35.80%, wet basis
90% reduction
(2 displacement
washes)
47.93%, wet basis
1.14%, wet basis
-63%
up to 90% reduction (2-3)
displacement washes)
50%, wet washed cake
3-5%, dry washed cake
12%, dry cake, unwashed
45 - 60%
0.10 - 0.25
-------
9.0 Data Gaps and Limitations
Complete data sets are not available for concentrated mode operation of
double alkali processes, the most commercially viable mode. Only par-
tial data are available from the 20 MW prototype system at Gulf Power,
the largest successful, closed loop, dual alkali system operated on a
coal-fired utility in the U.S. to date. In most cases, neither total
gas analyses nor extensive trace component analyses have been conducted.
Commercial data are not available for concentrated mode, closed loop
operation. The concentrated process cannot operate closed loop on
streams containing either a high oxygen content or a low sulfur
dioxide concentration.
10.0 Related Programs
Arthur D. Little has an $800,000 contract from EPA to study waste salt
fixation, disposal and utilization. Both ADL/CEA and FMC are construct-
ing their first full-scale, commercial, utility-based processes.
REFERENCES
>.
1. Tuttle, John. Summary Report on S02 Control Systems for Industrial
Combustion and Process Sources, Vol. I, Industrial Boilers, U. S.
Environmental Protection Agency, Research Triangle Park, N.D.,
December 1977, 147 pp.
2. Siddigi, A. A. and J. W. Tenini. FGD-A Viable Alternative, Hydrocarbon
Processing, Houston, Texas, October 1977, pp 104-110.
3. Laseke, B. A. and R. W. Dewitt. Status of Flue Gas Desulfurization
Systems in the United States, presented at Seventh FGD Symposium, EPA,
Hollywood, Florida, November 1977, 35 pp.
4. Rush, R. E. and R. A. Edwards. Operational Experience with Three 20 MW
Prototype Flue Gas Desulfurization Processes at Gulf Power Company's
Scholz Electric Generating Station, Electric Power Research Institute
Report Summary, Fourth Quarter, 1977, 82 pp.
5. Van Ness, R. P., Louisville Gas and Electric Company Scrubber
Experiences and Plans, presented at Seventh FGD Symposium of the
Environmental Protection Agency, Hollywood, Florida, November 1977,
10 pp.
6. Chemical Processing, Chem-Trends, August 1977.
D-106
-------
7. FMC Corp. Environmental Equipment Division Capabilities Statement
Sulfur Dioxide .Control Systems, Technical Progres Report 100 "
Itasca, Illinois, March 1976, 44 pp. H '
8. LaMantia, C. R. and R. R. Lunt, et al. Operatina FxnpriPnrp
Dual Alkali Prototype System at Gulf PoweSiKer! ? s'er e ,~I
on FGD> VOL '• »*-™
9. Kaplan, N. Introduction to Double Alkali FGD Technology, presented in
Proceedings from Symposium on FGD, Vol. I, EPA-600-2-76-136a New
Orleans, La., May 1976, pp 387-421.
10. Ando, Jumpei. Status of S02 and NOX Removal Systems in Japan, presented
at Seventh FGD Symposium, EPA, Hollywood, Florida, November 1977
21 PP.
11. Kittrell, J. R. and N. Godley. Impact of SOX Emissions Control on
Petroleum Refining Industry, Vol. II, EPA-600/2-76-161b, U. S.
Environmental Protection Agency, Research Triangle Park, N.C.,
June 1976, 300 pp.
12. Kaplan, N. An Overview of Double Alkali Processes for Flue Gas
Desulfunzation, presented in Proceedings from the EPA Symposium on Flue
Gas Desulfurization, Atlanta, Georgia, November 1974, 65 pp.
13. Cornell, C. G. and D. A. Dahlstrom. Sulfur Dioxide Removal in a Double-
Alkali Plant, Chemical Engineering Progress, 69(12): 47-53, 1973.
14. Legatski, L. K. , K. E. Johnson, and L. Y. Lee. The FMC Concentrated
Double Alkali Process, presented in Proceedings from Symposium on FGD,
Vol. I, EPA-600-2-76-136a, New Orleans, La., May 1976, pp. 471-502.
15. The Status of Flue Gas Desulfurization Applications in the U. S.: A
Technological Assessment, The Federal Power Commission, Bureau of
Power, July 1977.
16. Princoitta, F. EPA Presentation on Status of Flue Gas Desulfurization
Technology, presented at National Power Plant Hearings, October 1973,
66 pp.
D-107
-------
MAGNESIUM OXIDE PROCESS
1.0 General Information
1.1 Operating Principles^ - Magnesium oxide slurry absorption of
sulfur dioxide from flue gas, after particulate removal, in a wet
scrubber. The aqueous slurry is centrifuged; spent magnesium
solids are recovered and calcined at elevated temperature with
coke to regenerate magnesium oxide crystals. An SO^-rich gas
stream is produced during regeneration.
1.2 Development Status - Currently available and commercially tested
in the U.S. and Japarr .
1.3 Licensor/Developer '
In the U.S.: United Engineers and Contractors, Inc.
Philadelphia, Pa. (Mag-Ox process)
Chemico (Chemical Construction Co.) and
Basic Chemicals
In Japan: Onahama - Tsu Rishing
Mitsui Mining
Chemico-Mitsui
1.4 Commercial Applications^1'2'3'
• Chemico: Boston Edison Company operated a magnesium oxide
scrubbing system at its Mystic Station in Everett, Mass.
Chemico Air Pollution Control Company installed the system on
Unit No. 6, a 150 MW oil-fired boiler unit. The unit was oper-
ated from April 1972 through June 1974. The regeneration facil-
ity for the Mag-Ox unit was located in Rumford, R.I.
Potomac Electric Power Company operated a magnesium oxide
scrubbing system at its Dickerson Station in Dickerson,
Maryland on Unit No. 3, a 190 MW coal-fired boiler unit. Only
50 percent of the gas, or 95 MW equivalent, was processed through
the scrubber. The system was operated from September 1973 to
D-108
-------
(S™ J^5' .The generation facility utilized, the same as
Boston Edison's, was located in Rumford, R.I.
• ""Ited Engineers: A demonstration program is in progress at
t days tone No 1 boiler of Philadelphia Electric Company. Pending
the outcome of the scheduled one-year test program, full-scale
application of the process may result at this station and at
Cromby. The present system is 120 MW in capacity. United
Engineers has a total of four Mag-Ox plants in operation in the
United States on 846 MW total capacity.
In Japan, both Onahama-Tsu Rishing and Mitsui Mining have one
plant each using the magnesium oxide process and producing sul-
furic acid as the byproduct. Chemi co-Mi tsui has one plant using
magnesium oxide as absorbent and producing sulfur.
2.0 Process Information
2.1 Flow Diagram - see Figures D-15, D-16 and D-17.
(2 3)
• Process Descriptionv * ' - In the United Engineers version
of the Mag-Ox Process, hot flue gas enters the parti cul ate
scrubber where it is contacted with water, removing the majority
of parti cul ate matter. Most of the hydrogen chloride, a variable
fraction of the sulfur trioxide and a minor amount of the sulfur
dioxide contained in the flue gas are absorbed in the particulate
scrubbing liquor. Caustic soda is added to control solution
pH and prevent corrosion. A liquid blowdown stream is taken off
the particulate scrubber at a rate sufficient to prevent exces-
sive chloride buildup. This stream is neutralized and sent to
the station ash settling basin.
The flue gas from the particulate scrubber is contacted with an
aqueous slurry of magnesium sulfite to remove better than 90%
of the sulfur dioxide in the flue gas:
MgO + S02 + H20 = MgS03 • nH20, n= 3 or 6
Insoluble magnesium sulfite is converted to soluble magnesium
bisulfite:
MgS03 + S02 + H20 •> Mg(HS03)
?
In the scrubber surge tank slaked magnesium oxide is added to
Ihe circulating scr&ber liquor, converting bisulfite back to
sulfite:
Mg(HS03)2 + MgO -> 2 MgS03 + H20
D-109
-------
PARTICULATE
SCRUBBER
SURGE TANK
I
o
PARTICULATE
SCRUBBER
S02 SCRUBBER
SURGE TANK
LEGEND:*
1. PRESCRUBBER INLET GAS
2. MAKEUP MgO
3 REGENERATED Mfl TO SCRUBBER
4. UNDERSIZED Mg(OH)3 SLURRY
5. PARTICULATE SCRUBBER MAKEUP WATER
6. PART'.CULATE SCRUBBER PURGE
7. SCRUBBER OUTLET GAS
8. ABSORBER BLEED STREAM. RECYCLE
9. PARTICULATE SCRUBBER
10. ABSORBER INLET GAS
11. ABSORBER DOWNCDMER
12. CYCLONE DUST
13. ABSORBER RECYCLE
14. MAKEUP WATER
29. CAUSTIC SODA
THE SPECIFIC STREAM NUMBERING SYSTEMS CONFORM TO THOSE
USED IN FIGURES D-16 AND D-17
Figure D-15. Magnesium Oxide Scrubber System
(3)
-------
MOTHER
LIQUOR
TANK
4.
8.
12.
15.
UNDERSIZED Mg (OH)2 RECYCLE
MgSO-j SLURRY
CYCLONE DUST
AIR TO DRYEP
19.
20.
21.
FUEL OIL TO DRYER
THICKENER UNDERFLOW
THICKENER OVERFLOW
CENTRIFUGED SOLIDS
CENTRATE
DRYER PRODUCT
THE SPECIFIC STREAM NUMBERING SYSTEMS CONFORM TO THOSE
USED IN FIGURES D-15 AND D-17
Figure D-16. Magnesium Oxide Process - MgS03 Recovery System
(3)
-------
LEGEND
THE SPECIFIC STREAM NUMBERING SYSTEMS CONFORM TO THOSE
USED IN FIGURES D-15 AND D-16
I
FLUIDIZED
BED
REACTOR
21. DRYER PRODUCT
22. CALCINER FUEL OIL
23, CALCINER AIR
3. REGENERATED MgO
24. VENTURI COOLING TOWER SLOWDOWN
25. VENTURI COOLING TOWER MAKEUP WATER
26. ENRICHED SO2 STREAM
27. CALCINER PRODUCT
28. CYCLONE PRODUCT GAS
COMBUSTION
AIR BLOWER
VENTURI WET
SCRUBBER AND
GAS COOLING
TOWER
Figure D-17. MgO Regeneration Plant
(3)
-------
A bleed stream from the scrubber slurry oroceprk tn a th,vi,
centrifuge and rotary kiln where niagnesiu', s l?te i r ™ e
a dry product. Some of the sulflte is oxidized to sulfate:
MgS03 + 1/2 02 + MgS04
The dry crystals are calcined in a fluid bed reactor by the
following reaction: *
MgS03 ^ MgO + S02 t
The regenerated magnesium oxide crystals are separated from the
S02-rich gas stream. The gas stream is scrubbed and cooled prior
to transfer to a suitable conversion facility.
The Mag-Ox process developed by Chemico differs somewhat from the
United Engineers process. Aside stream of the magnesium sulfite
(hydrate) is separated and sent to a centrifuge. Also, carbon
(coke) is added to the feed stream to the calciner to improve
sulfate regeneration:
MgS04 + 0.5 C A* MgO + 0.5 C02 + S02 t
2.2 Equipment - Most of the equipment in the Mag-Ox process is conven-
tional, such as the variable throat venturi, scrubber, thickener,
surge tanks, centrifuges, cyclones, cooling towers/wet scrubber and
the dryer. Some of the equipment must be altered to handle magne-
sium slurries and its associated problems. One item of equipment,
the fluidized bed reactor, is specifically designed for this use.
2.3 Feed Stream/Requirements
t Temperature(3'4) - 400°K-440°K (250°F-330°F) typically , from
utility boiler
• Pressure - near atmospheric
t Loading(5) - 0.06 to 4.5 gr/SCF (0.145 to 10.9 g/Ha?) dry par-
ticulate acceptable to first stage particulate scrubber^ Systems
have been designed for 1850 ppm S02 range but can be designed for
much higher loadings.
t Contaminant Limitation - ?
D-113
-------
2.4 Operating Parameters
§ Particulate Scrubbing Step
- Temperature(3): 326°K-328°K (125°F-130°F)
- Pressure drop^:* liquid 254-305 mm H20 (10-12 in. H20)
- pH(3): 2.8 to 3.1
(3)
- Slurry concentrationv . 2%
- L/G^:1" 0.65 liters/m3 (4.85 gal/1000 ACF outlet)
- Turndown: ?
• Absorption Step
- Temperature (slurry)^: 320°K (120°F)
- Pressure drop^3M 254 mm FLO (10 in. water total),
5 in. per stage)
- phT ': about 6.3
- Slurry concentration^ . 5-10 wt % magnesium sulfite
- L/G^: 6.5 liters/m3 (48.5 gal/1000 ACF out)
- Turndown: ?
• Drying Step
- Temperature^: 500°K (450°F) off-gas
- Mass throughput rate: ?
t Regeneration Step
- Temperature^3^: 1230°K (1750°F)
- Mass throughput rate: ?
*Pressure drop is for variable thraot venturi scrubber or Environeering
Ventri-Rod unit. Both types are in service .
tL/G is liquid-to-gas ratio.
fPressure drop is for two-stage Environeering Ventri-Rod unit.
D-114
-------
2.5 Process Efficiency and Reliability^3'6^ - Based on the magnesium
oxide scrubbing system installed at the Eddystone Station of
Philadelphia Electric Co. in mid-1971, the cumulative availability
of 32% has been disappointing. Many operating difficulties have
been encountered despite high particulate removals (9756-98%) and
sulfur dioxide removal (greater than 95%). The longest continuous
run has been 140 hours. Particulate removal efficiency is based on
an outlet stream from the 93% efficient mechanical/electrostatic
precipitator system. Also, most of the hydrogen chloride and some
of the sulfur trioxide are removed in the particulate scrubber.
A Chemico-Mitsui plant at Idemitsu in Japan reports operability
of 100% for a system operating an oil burner and Claus furnace gases.
2.6 Raw Material Requirements
t MgO Makeup^5' - 7% replacement per year.
• Coke - ?
t Caustic Soda - ?
• Lime - ?
2.7 Utility Requirements
• Electricity^5'6^ - About 2% to 3% of station generating capacity.
Pump and fan power usage at a Japanese plant treating Glaus fur-
nace and oil burner gases is 11.7 kwh/Nm3 (0.31 kwh/scf).
• Process Water
- Particulate scrubber: ?
- Absorber: ?
• Air
- Dryer: ?
- Calciner: ?
• Fuel Oil
- Dryer: ?
- Calciner: ?
D-115
-------
- Total: 9.8 liters/mwh (2.6 gal/MWH) of No. 2 fuel oil. How
ever, No. 6 oil can be substituted^), 311°K (100°F) reheat
required at Philadelphia Electric's Eddystone plant(^).
2.8 Miscellaneous
• Oxidation^3' - Sulfite oxidation has been reported at around
15% (during drying).
3.0 Process Advantages
• Minor disposal problems since MgO is regenerated and sulfur is recovered
in a usable form.
• Sulfur can be recovered as high grade sulfuric acid or as elemental
sulfur.
• Regeneration can be carried out at a remote site, thus permitting use
of a central regeneration facility serving several FGD units.
• Only minor plugging and scaling problems encountered in the scrubber.
(5)
4.0 Process Limitations^ '
t Energy requirements are relatively high.
• Past demonstrations at Boston Edison and Potomac Electric's Dickerson
Station exhibited relatively low reliability.
• Difficulties with the centrifuged MgS03 • nH^O cake (r\~3 or 6) because
the trihydrate and hexahydrate crystals have different handling
properties.
t Corrosion problems in the slurry-handling systems.
• Regenerated and makeup MgO is required in a pulverized form.
• The scrubber requires a high liquid-to-gas ratio.
• Fly ash must be kept out of the regeneration system, thus necessitating
extensive particulate removal prior to gas processing.
5.0 Process Economics
A 1973 EPA study provided the following costs^ ':
Capital investment: $33-66/kw
Operating costs: 1.6-3.0 mils/kwh - no credit for sulfur recovery
1.4-2.8 mils/kwh - with credit for sulfur recovery
D-116
-------
In another more recent study, estimated costs of providing 1420 MW's
(gross) of scrubber service, excluding acid plant, for the Potomac
Electric Power Company's Dickerson Station is $106/kw (1975 dollars).
Operating cost was estimated at 5.0 mils/kwh^.
Total capital cost for Philadelphia Electric's Eddystone plant is
estimated to be about $130/kw, including particulate scrubbing and a
sulfuric acid facility. Also included is about a $20/kw retrofit charge.
The operating and maintenance cost of the Eddystone scrubber is estimated
at about 2.3 mils/kwh excluding any credit for by-product sulfuric acid.
If credit is taken for by-product acid, then the O&M cost drops to
2.0 mils/kwh. Use of a more normal 283°K (50°F) reheat instead of the
311°K (100°F) used at Eddystone would reduce the cost an additional 10%'3'.
6.0 Input Streams (see Figures D-15, D-16 and D-17)
• Prescrubber inlet gas (Stream 1) - see Section 2.3.
Potomac Electric- Philadelphia Electric-
Dickerson^' Eddystonel3/
Temperature: 400°K (250°F) 420°K-440°K (300°F-330°F)
Flow Rate: 500,000 m3/hr (295,000 acfm) 545,000m3/hr (321,000 acfin)
1700 ppm, 3% sulfur coal
Particulate: 0.145 g/NrrT (0.06 gr/scf
dry) with precipitator
operating
10.8 g/Nm3 (4.5 gr/scf dry)
without precipitator
Derived from 2.1
coal, dry
sulfur
90.7 kg/hr (200 Ib/hr)
•
•
•
t
•
•
•
Makeup MgO (Stream 2) - ?
Particulate Scrubber Makeup Water (Stream 5) - ?
Absorber Makeup Water (Stream 14) - ?
Dryer Air (Stream 15) - ?
Dryer Fuel Oil (Stream 16) - ?
Calciner Fuel Oil (Stream 22) - ?
Calciner Air (Stream 23) - ?
D-117
-------
• Venturi Cooling Tower Makeup Water (Stream 25) - ?
• Caustic Soda (Stream 29) - ?
7.0 Intermediate Streams (see Figures D-15, D-16 and D-17)
• Absorber Bleed Stream (Stream 8)
t
•
Philadelphia Electric-
Eddystone(3)
pH:
concentration
-6.3
5 - 10 wt %
Sulfate concentration: 2,000 - 5,000 ppm
Particulate Scrubber Recycle (Stream 9) - see Discharge Stream 6
Absorber Inlet Gas (Stream 10)
Potomac Electric-
Dickerson(4)
Temperature: 320°K (120°F)
Flow Rate:
Particulate:
Absorber Downcomer (Stream 11) - ?
Absorber Recycle (Stream 13) - ?
pH:
Flow Rate:
Cyclone Dust (dryer) (Stream 12)
Temperature:
Philadelphia Electric-
Eddystone(3) _
330°K (130°F)
455,000 m3/hr(268,000 acfin)
14 kg/hr (30 Ib/hr)
Philadelphia Electric-
Eddystone(3)
5.8 - 6.8
50,658 liters/min
(14,000 gpm)
Potomac Electric - Pickersorr '
500°K (450°F)
Undersized Mg(OH)2 Slurry (Stream 4) - ?
Thickener Underflow (Stream 18) - ?
D-118
-------
• Centrifuged Solids (Stream 19) - ?
• Centrate (Stream 20) - ?
t Dryer Product (Stream 21) - ?
• Regenerated MgO (Stream 3) - ?
• Calciner Product (Stream 27) - ?
• Cyclone Product Gas (Calciner) (Stream 28) - ?
8.0 Discharge Streams (see Figures D-15, D-16, and D-17)
• Particulate Scrubber Purge (Stream 6)
Philadelphia Electric -
Eddystone(3)
Temperature: 325°K-328°K (125°F-130°F)
PH: 2.8 - 3.1
Slurry concentration: 2 wt %
Chloride concentration: <1000 ppm
t Scrubber Outlet Gas (Stream 7)
Philadelphia Electric -
Eddystone(3)
Temperature: 325°K-328°K (125°F-130°F)
Flow Rate: 469,000 m3/hr (276,000
ACFM)
Particulate: <0.18 g/1000 kcal
(<0.1 Ib/MM Btu);
13.6 kg/hr (30 Ib/hr)
t Venturi Cool ing Tower Slowdown (Stream 24) - ?
• Enriched S02 (Stream 26) - ?
9.0 Data Gaps and Limitations
Incomplete composition data are currently available for most
streams in the Magnesium Oxide (Mag-Ox) Scrubbing Processes. Data per-
taining to feed stream pressure, contaminant limitations, scrubber turn-
down ratio, are missing. Dryer gas velocity, mass throughput rate,
D-119
-------
turndown ratio and fuel requirement are missing. Calciner pressure
drop, superficial gas velocity, and fluidized bed height data are missing.
Other missing data include material requirements for coke, caustic soda
and lime, (when required), and process water usage. At present, there
are no full scale commercial Mag-Ox Scrubber systems in operation in the
U.S. To date, units have only been tested on utility flue gases. How-
ever, the process is regenerative and thus is closely tied into an
elemental sulfur or sulfuric acid plant. At a coal gasification facility
it would be possible to use a central sulfur plant for the Mag-Ox plant
and for acid-gas treatment enriched gas. Particulate scrubber, absorber
and system makeup and air flow rates required for dryer and calciner
operations are unknown.
10.0 Related Programs
None known in the U.S. Commercial scale processes are in operation
in Japan.
D-120
-------
REFERENCES
1. Siddigi, A.A. and J.W. Tenini. FGD-A Viable Alternative, Hydrocarbon
Processing, Houston, Texas, October 1977, pp. 104-110.
2. Kittrell, J.R. and N. Godley. Impact of SOX Emissions Control on
Petroleum Refining Industry. Vol. II, EPA-600/2-76-161b, U.S. EPA,
Research Triangle Park, N.C., June 1976, 300 p.
3. Gille, J.A. and J.S. MacKenzie. Philadelphia Electric's Experience
with Magnesium Oxide Scrubbing, presented at Fourth FGD Symposium, EPA,
Hollywood, Florida, November 8-11, 1977, 15 pp.
4. Taylor, R.B. and D. Erdman, Summary of Operations of the Chemico-Basic
MgO FGD System at the PEPCO Dickerson Generating Station, presented in
Proceedings from Symposium on Flue Gas Desulfurization, Vol. II,
EPA-600-2-76-136b, New Orleans, La., May 1976, pp 735-758.
5. The Status of Flue Gas Desulfurization in the U.S.: A Technological
Assessment, The Federal Power Commission, Bureau of Power, July 1977.
6. Ando, Juniper. Status of S02 and NOX Removal Systems in Japan, presented
at Seventh Flue Gas Desulfurization Symposium, EPA, Hollywood, Florida,
November 1977, 21 p.
7. Princiotta, F. EPA Presentation on Status of Flue Gas Desulfurization
Technology, presented at National Power Plant Hearings, October 1973,
66 p.
D-121
-------
Particulate Control Module
Fabric Filtration
Electrostatic Precipitation
Venturi Scrubbing
Cyclones
D-122
-------
FABRIC FILTRATION PROCESS
1.0 General Information
1.1 Operating Principles - Physical removal of participates from a gas
stream by impaction, interception, diffusion and/or electrostatic
attraction^ '.
1.2 Developmental Status - Commercially available.
1.3 Licensor/Developer - Many companies manufacture fabric filtration
systems, each system incorporating certain proprietary features.
A complete listing of manufacturers are presented in technical and
trade journals (e.g., Ref. 2).
1.4 Commercial Applications - Fabric filtration has been applied to coal
fired boilers, coal loading and transport facilities, coal proc-
essing^ ', and a large number of miscellaneous industrial gas clean-
ing operations.
2.0 Process Information
2.1 Flow Diagram - See Figure D-18.
• Process Description - Fabric filters are a series of tubular or
envelope shaped bags contained in a structure called a baghouse.
The filters or bags can be constructed of a variety of fibers,
depending on design requirements. Table D-30 presents a limited
selection of fabrics with their relative costs and properties for
various applications^ A variety of methods (both continuous
and intermittent) are used for bag cleaning/
*
*Bag cleaning methods include use of shakers, high velocity reverse air flow
sonic energy and ring type spargers. These can be employed on both continuous
and intermittent bases.
D-123
-------
,(CLEANING PULSE)
/ L
BAG SUPPORT
DUST HOPPER
LEGEND:
1. Dirty Gas
2. High Pressure Air (for bag cleaning)
3. Cleaned Gas
4. Collected Dust
Figure D-18. Continuous Cleaning Bag Filter
D-124
-------
TABLE D-30. FILTER FABRIC PROPERTIES
(5)
Fabric
Cotton
Wool
Nylon
Dacron
Orion
Creslan
Dyne!
Polypropylene
Teflon
Fiberglas
Fi 1 tron
Nomex
Operating
Temperature*
(°F)
180
200
200
275
260
250
160
200
450
500
270
375
Acid Resistance
Poor
Very good
Fair
Good in most acids,
cone. H2S04 partially
dissolves fabric
Good to excellent
Good
Excellent
Excellent
Excellent
Fair to good
Good to excellent
Fair
Alkali Resistance
Very Good
Poor
Excellent
Good in weak alkali,
fair in strong
Fair to good
Good
Excellent
Excellent
Excellent
Fair to good
Good
Excellent at low
temperature
Flex and
Abrasion
Very Good
Fair to Good
Excellent
Very good
Good
Good to very good
Fair to good
Excellent
Fair
Fair
Good to very good
Excellent
o
ro
Continuous operating temperature as recommended by the IGC.
-------
2.2 Equipment - Bags, structure and bag cleaning device. All equipment
is usually supplied by the manufacturer.
2.3 Feed Stream Requirements*
Temperature: to 700°K (800°F)
Pressure: None
Gas composition: Wet, corrosive, explosive or oily gases are not
well suited for treatment in a baghouse. However,
design and operation modifications can be employed
which will make treatment of most gases possible,
although in some cases very expensive.
2.4 Operating Parameters - As above except:
Pressure: Pressure drop through the filter ranges from about
0.5 KPa (2 in. H20) to 2.4 KPa (10 in. HgO).
2.5 Process Efficiency and Reliabilityf
• Efficiency^
Particle Mean Control
Diameter Efficiency
0.25 98.5% - 99.7+%
0.50 98.7% - 99.5+%
0.75 99.1% - 99.5+%
1.00 99.0% - 99.5+%
Above 100 99.5+%
• Reliability - Many years of use have proven that well designed,
operated and maintained fabric filters can provide trouble-free
service in many varied industrial applications.
2.6 Raw Material Requirements - None.
*Requirements are dependent on fabric used; the figures given represent the
maximums which could be handled by commercially available bags.
tReliability and efficiency will be dependent on the specific fabric filtration
design and on gas stream characteristics such as: chemical composition, site
distribution, water vapor content, temperature, etc. Values given report
results of a range of applications.
D-126
-------
2.7 Utility Requirements
Electri
method.
• Electricity: Dependent on gas flow rate, pressure, and cleaning
• High pressure (for bag cleaning): Dependent on above factors.
Not used in all designs.
2.8 Miscellaneous - Maintenance needs are dependent on the fabric selected
and the nature of the gas being handled.
3.0 Process Advantages
• High efficiency on fine particulates.
• Low energy requirements.
• Can be adapted to a wide range of gas stream characteristics.
• Proven system; considerable experience has been acquired in a wide
range of applications.
4.0 Process Limitations
• Requires large structures for high volume flows.
• Bag replacement required.
• Cannot handle explosive, corrosive or wet gas mixtures without special
design considerations (e.g., control of temperature, use of proper
fabrics, and selection of proper cleaning methods).
• Not suitable for operation at relatively high temperatures (generally
above 550); temperatures limited, dependent on fabric selected.
5.0 Process Economics* - Installed costs for fabric filters have been
reported to vary from $50.76 to $80.26 per actual cubic meter per minute
($1.48 to $2.34 per acfmr • The combined operating and maintenance
costs are reported as $12.36 to $17.65 per actual cubic meter per minute
($0.35 to $0.50 per acfm) on an annualized basis '
6.0 Input Streams - Basis for stream compositions (input and discharge) is
the same unit as in Section 5.0 (Process Economics), operating on a
coal fired industrial boiler.
*The costs presented are in early 1974 dollars and for a 39,100 Nm /hr
(70,000 acfm) unit operating at 395°K (250°F) using NomexK felt bags.
D-127
-------
6.1 Feed Gas Stream No. l - see Table D-31.
6.2 High Pressure Air Stream No. 2^ r High pressure airj varies with
manufacturer.
TABLE D-31. INPUT AND DISCHARGE STREAM CHARACTERISTICS FOR BAGHOUSE
FILTRATION^7'
Particle Mean
Diameter, ym
79.5
6
4
2.8
<0.9
Inlet Loading
mg/scf
4.221
2.292
1.482
1.254
3.2557
Outlet Loading
mg/scf
0.0068
0.0060
0.003
0.003?
0.0221
Overall Efficiency
Removal
Efficiency, %
99.84
99.74
99.78
99.69
99.31
99.55
7.0 Discharge Streams
(7)
7.1 Cleaned Gas Stream No. 3V ' - see Table D-31.
7.2 Collected Dust Stream No. 4 - Dry collected dust quantity (rate)
dependent on particle loading and removal efficiency.
8.0 Data Gaps/Limitations
Extensive performance data are available for fabric filtration appli-
cations to a variety of industrial gas cleaning operations. Evaluation
of expected performance of the system in applications to coal gasifica-
tion plant gas streams (e.g., coal and ash lockhopper vent gases) requires
data on detailed characteristics of the gases to be treated. Such data
which include gas temperature, particle size distribution and chemical
characteristics of the gas stream including the particulates, are
generally either not available or are incomplete.
9.0 Related Programs
Acurex/Aerotherm Corporation and Westinghouse are both presently
involved in research programs to develop high temperature and pressure
fabric filtration devices^ '.
D-128
-------
REFERENCES
1. Simon, H., Baghouses, Air Pollution Engineering Manual; AP-40, p. 106.
2. Product Guide, Journal of the Air Pollution Control Association,
Vol. 27, No. 3 (1977).
3. Lear Siegler Inc., Installations of LUHR System Bag Houses , LUHR/PL/
002A/9/74, September 1974.
4. Hesketh, H.E., Understanding and Controlling Air Pollution, Ann Arbor
Science Publishers, 1973, p. 341.
5. Reigel, S.A. and R.P. Bundy, Why the Swing to Baghouses, Power, January
1977, pp. 68-73.
6. Turner, J.H., Extending Fabric Filter Capabilities, JAPCA, 24:1182, 1974.
7. McKenna, J.D., et al, Performance and Cost Comparisons Between Fabric
Filters and Alternate Particulate Control Techniques, JPACA, 24:1144, 1974.
8. Bush, J.R., Future Need and Impact on the Particulate Control Equipment
Industry Due to Synthetic Fuels, paper presented at the 3rd Symposium
on Environmental Aspects of Fuel Conversion Technology, Hollywood, Florida,
September 1977.
D-129
-------
ELECTROSTATIC PRECIPITATION PROCESS
1.0 General Information
1.1 Operating Principle - The removal of participates from a gas stream
by imposing an electrical charge and collecting the charged particles
on oppositely charged collector plates.
1.2 Development Status - Commercially available.
1.3 Licensor/Developer - Many companies are involved in the commercial
production of electrostatic collection devices. Complete listing
can be obtained in various trade and technical journals (e.g.,
Ref. 1).
1.4 Commercial Applications - Electrostatic precipitators have been
used to control particulate emissions in a wide range of industrial
applications including: electrical power generation, cement making,
(2}
steel making, smelting and the pulp and paper industriesv '.
2.0 Process Information
2.1 Flow Diagram - see Figure D-19.
2.2 Equipment - Support structure, electrodes, collection bin, power
supply (rectifier) and controls.
2.3 Feed Stream Requirements
t Temperature: precipitators have been applied at temperatures
from ambient to about 700°K (800°F)(2,3).
• Pressure: precipitators have been applied on only a limited
basis above atmospheric pressure. It has been reported that
pilot and full scale tests have been successful at up to
55
D-130
-------
COLLECTION
ELECTRODES
CORONA WIRES
GAS OUT
•GAS IN
GAS OUT
GAS IN
COLLECTED PARTICULATE
SIDE VIEW
Figure D-19. Typical Electrostatic Precipitator
D-13T
-------
• Nature of participate matter: The electrical resistivity*
of particles should be between 109 and 1010 ohm-cm for optimum
operation. The addition of certain additives (such as $63) can
act to correct resistivity problems.
2.4 Operating Parameters - Precipitators have been operated in the
above ranges; considerable testing will need to be done for
operation at higher temperatures and/or pressures.
2.5 Process Efficiency and Reliability^ ' - Efficiency is dependent on
gas stream and particle parameters. Values given are for a range of
applications operating on both the hot and cold side of coal fired
boilers.
Particle Mean Collection Efficiency Range
Diameter, ym %
0.1 90 - 99.4
0.5 90 - 98.7
1.0 95 - 99.6
above 5.0 98 - 99.9
The system is generally reliable if operated near design conditions.
Changes in particle size distribution, particle resistivity, flow
rate and/or temperature can cause severe changes in performance.
2.6 Raw Material Requirements - None.
2.7 Utility Requirements
• Electricity: for corona power plus the pressure drop through
unit.
2.8 Miscellaneous - Maintenance needs are relatively high as compared
to other particulate collection devices. This is due mainly to
the complexity of the system and the precipitators inability to
J
accept changes in process parameters
*Electrical resistivity is a function of temperature as well as particle
composition, therefore tests are usually performed on the actual gas stream
before detailed design of the precipitator is performed.
D-132
-------
3.0 Process Advantages^ '
• Highly effective collection
• High efficiency on small particles
• Low energy consumption (pressure drop is less than 25% of a fabric
filter)
• Easy expandability
4.0 Process Disadvantages' '
• High capital expenditures and space requirements
• Sensitivity to process upsets or changes
• Relatively high maintenance costs
(3 7}
5.0 Process Economics^' '
Capital and operating costs for electrostatic precipitators vary
widely depending on the application. The range for installed costs has
been $7 to $17 per actual m3/hr ($4 to $10 per acfm). Operating costs
range from $0.08 to $0.37 per actual m3/hr ($0.05 to $0.22 per acfm).
Based on these figures, the annualized costs would be about $0.85 to
3
$3.40 per actual m /hr ($0.50 to $2.00 per acfm). These figures are
based on 1974 dollars.
6.0 Input Streams - see Table D-32.
7.0 Discharge Streams
7.1 Gas Stream - see Table D-32
7.2 Particulates Collected - depends upon application
8.0 Data Gaps and Limitations
Electrostatic precipitators are used widely in numerous, and a range
of industrial applications and considerable design and operating exper-
ience exist for existing applications. The design requirements for the
application of the system to the purification of gases from coal conver-
sion processes have not been evaluated. Such an evaluation would require
D-133
-------
TABLE D-32. PERFORMANCE DATA - AN ELECTROSTATIC PRECIPITATOR
(6)
Figures given as base on actual field data for one installation.
The location of the installation, collection conditions, tempera-
ture, etc. are not reported.
Particle Diameter
ytn
0.2
0.4
0.6
1.0
2
4
6
8
10
Removal Efficiency
%
95
94
95.5
97.3
99
99.4
99.4
99.5
99.5
data on the characteristics of the specific gas streams to be treated
(e.g., particle size and size distribution, temperature, particle
resistivity and presence of other constituents in the gas).
9.0 Related Programs
Research-Cottrell is presently involved in research on high tem-
perature and pressure precipitators for use on coal conversion processes.
D-134
-------
REFERENCES
1. Product Guide, JAPCA, Vol. 27, No. 3, 1977.
2. Simon, H. Electrical Precipitators, Air Pollution Engineering Manual,
AP-40, p. 135.
3. Walker, A.B., Hot-Side Precipitators, JAPCA, 25:143, 1975.
4. Rao, A.K., et al, Particulate Removal From Gas Streams at High
Temperature/High Pressure, EPA 600/2-75-020, August 1975.
5. McCain, J.D., et al, Results of Field Measurements of Industrial
Particulate Sources and Electrostatic Precipitator Performance, JAPCA
25:117, 1977.
6 . Lasater, R.C. and Hopkins, J.H., Removing Particulates from Stack Gases,
Chemical Engineering, October 17, 1977.
7. McKenna, J.D. et al, Performance and Cost Comparisons Between Fabric
Filters and Alternate Particulate Control Techniques, JAPCA, 24:1144,
1974.
8. Bush, J.R., Future Need and Impact on the Particulate Control Equipment
Industry Due to Synthetic Fuels, presented at the 3rd Symposium on
Environmental Aspects of Fuel Conversion Technology, Hollywood, Florida,
September 1977.
D-135
-------
VENTURI SCRUBBING PROCESS
1.0 General Information
1.1 Operating Principle - Physical removal of particulates from a gas
stream by the inertia! impactions of the particles with diffused
scrubbant droplets.
1.2 Developmental Status - Commercially available.
1.3 Licensor/Developer - Many companies manufacture vsnturi scrubbing
systems; each system incorporates certain proprietary features.
A complete listing of manufacturers is presented in technical
and trade journals (e.g., Ref. 1 and 2).
1.4 Commercial Applications - Venturi scrubbers have been applied to
foundry cupolas, blast furnaces, lime kilns and a large number
(3}
of miscellaneous industrial gas cleaning operations^ .
2.0 Process Information
2.1 Flow Diagram - see Figure D-20.
• Process Description - Particulate-laden gas stream (Stream 1)
enters scrubber housing, passes through a venturi section through
which low pressure water flows. The water is atomized by using
some of the energy from the gas stream. Particles in the
gas stream are moving faster than the atomized water; these
particles are captured by the atomized water droplets by inertia!
impaction. The particulate-laden water is separated from the
cleaned gas stream and sent to a water treatment facility and
the cleaned gas is sent to further processing or ducted to a
stack for release to the atmosphere.
2.2 Equipment - The scrubber housing and appropriate fittings and con-
nections are part of the basic scrubber design; and are usually
supplied by the manufacturer of the scrubbing device.
2.3 Feed Stream Requirements - Venturi scrubbing devices are applicable
over wide ranges of temperature, pressure, and gas compositions;
D-136
-------
VENTURI SECTION
L SCRUBBING SECTION
CO
-J
GAS/WATER SEPARATOR SECTION
LEGEND;
1.
2,
3.
4.
Dirty Gas Steam
Scrubbed Water
Cleaned Gas Stream
Particulate Laden Scrubber
Hater
Figure D-20. Venturi Scrubber
-------
however, these devices must be designed to meet the specific needs
of each application.
2.4 Operating Parameters (see Section 2.3 above) - The pressure drop
across the venturi scrubbers varies depending upon the application;
generally, the pressure drop is in the 2.4 kPa to 24 kPa range
(10 in. H20 to 100 in. HgO).
2.5 Process Efficiency and Reliability*
(4)
• Efficiencyv '
Particle Mean Diameter Control Efficiency
0.25u 60% to 92.5%
0.50y 85% to 97.2%
0.75U 92% to 99%
l.OOu 95% to 99.6%
(3 5)
• Reliabilityv * ' ~ Venturi scrubbers have been used for the
removal of particulates from gaseous streams for over thirty years.
When properly designed, operated and maintained, they can provide
relatively trouble free service in a variety of industrial
applications.
2.6 Raw Material Requirements
• Water treatment chemicals (e.g., for pH adjustment; chemical
and design dependent on gas and raw water characteristics)
2.7 Utility Requirements
Water: quantity dependent on the particulate loading in the gas
stream and whether provisions are made for treatment and recycling
of the spent water.
Electricity: requirements dependent on the system pressure drop
and water circulation ratio.
2.8 Miscellaneous - Care must be taken to correct pH adjustment to the
scrubbing water to avoid excessive corrosion and/or plugging
problems^ '.
*Reliability and efficiency are dependent on a number of parameters, primarily
relating to the characteristics and changes in the characteristics of the gas
stream (e.g., chemical composition, flow rate and particle size and distri-
bution) and the scrubbing water (e.g., pH)(3,4,5)
D-138
-------
3.0 Process Advantages^3'4'5'6)
t Effective performance over a wide loading range.
• Practically no re-entrainment of the particulates.
4.0 Process Limitation^4'5'6)
• High energy costs
• Difficulty of disposing of wet sludge
• Corrosion problems
t Possible visible moist plumes
• Scrubbing water is potential source of water pollution necessitating
extensive water treatment facilities.
5.0 Process Economics - Installed cost
Installed costs for venturi scrubbers have been reported to vary from
$80.00 to $170.00 per normal cubic meter per min ($2.11 to $4.51 per
acfm).* The operating costs (including the maintenance costs) are reported
as $30.00 to $38.00 per normal cubic meter per min (10.76 - 10.96 per
acfm) on an annulized basis^ '.
6.0 Input Streams
• Input gas stream (Stream 1), see Table D-33.
0 Scrubbing water (Stream 2), see Table D-33.
7.0 Discharge Stream
• Purified gas stream (Stream 3), see Table D-33.
• Particulate-laden water, Stream 4: (rate and composition dependent
on particulate loading in the inlet gas stream, scrubber efficiency,
and input water characteristics)
3
*The costs presented are in early 1974 dollars and for a 39,100 Nm/hr
(70,000 acfm) unit operating at 395°K (250°F) with scrubber effidennes
of 97% and 99%, respectively.
D-139
-------
TABLE D-33. SOME TYPICAL APPLICATIONS OF VENTURI SCRUBBERS WITH APPROPRIATE GAS AND
LIQUID STREAM INFORMATION(3)
Application
Black Liquor
recovery boiler
Fly Ash Sinter
Furnace
Blast Furnace
Foundry Cupola
Saturated Gas
Flow Rate
Nm-Vm
(acfm)
5,000
(185,000)
1,600
(60,000)
6,100
(225,000)
1,100
(41 ,400)
Temperature
°K
(°F)
354
(177)
314
(105)
330
(135)
327
(130)
Throat
Gas Velocity
m/s
(ft/sec)
61
(200)
32
(105)
110
(360)
107
(350)
Liquid
Flow Rate
Stream 2
gpm/mi n
10500
(2775)
1360
(360)
12000
(3150)
2650
(700)
Dust Concentration
gm/Nm3 (gr/scf)
In
Stream 1
6.7
(3)
1.36
(0.61)
8.9
(4)
17.9
(8)
Out
Stream 3
0.11
(0.05)
0.11
(0.05)
0.01
(0.005)
0.07
(0.03)
Pressure Drop
kPa
(inches H20)
8.4
(35)
1.2
(5)
14.4
(60)
13.2
(55)
I
o
-------
8.0 Data Gaps and Limitations
Extensive performance data are available for venturi scrubber
applications to a variety of industrial gas cleaning operations.
Evaluation of expected performance of the system in applications to coal
gasification plant gas streams (e.g., coal and ash lock hopper vent
gases) requires data on chemical characteristics of the gas and the
particle size distribution of the particulates to be removed.
REFERENCES
1. Produce Guide, Journal of the Air Pollution Control Association, Vol. 27,
No. 3, 1977.
2. Environmental Control Issue, Control Equipment, Environmental Science and
Technology, October 1977.
3. Hesketh, H.E., Fine Particle Collection Efficiency Related to Pressure
Drop, Scrubbant and Particle Properties and Contact Mechanisms; Journal
of the Air Pollution Control Association, Vol. 24, No. 10, 1974.
4. McKenna, J.D., and J.C. Mycock, et al, Performance and Cost Comparison
between Fabric Filters and Alternate Particulate Control Techniques;
JAPCA, Vol. 24, No. 12, 1974.
5. Ekman, F.O. and Johnstone, H.F., Collection of Aerosols in a Venturi
Scrubber; Industrial and Engineering Chemistry, Vol. 43, No. 6, June 1951.
6. Striner, B.A. and Thompson, R.J., Wet Scrubber Experience for Steel
Mill Applications, JAPCA, Vol. 27, No. 11, 1977.
D-141
-------
CYCLONES
1.0 General Information
1.1 Operating Principles - Removal of participates from a gas stream
by the action of centrifugal forces.
1.2 Developmental Status - Commercially available.
1.3 Licensor/Developer - Many companies manufacture cyclone collectors.
Listings of manufacturers are contained in various trade and tech-
nical journals (e.g., Ref. 1).
1.4 Commercial Application - Widely used in the chemical process
industry for removal of particulates from gaseous streams; also used
as a final collection device before gas discharge where particles
are large or loading light. Frequently used as a precleaner before
more efficient control devices.
2.0 Process Information
2.1 Flow Diagram - see Figure D-21
• Process Description - Cyclones operate by imparting a centrifugal
force to the gas stream. The circular shape and tangential
entrance change the gas flow pattern to a vortex, spiraling it
downward. The inertia of the particles carries them to the cyclone
wall and down the sides to a collector section from which they are
removed(3).
2.2 Equipment - Cyclones can be either single or multiple; multiple units
can be arranged in either series or parallel (see Figure D-21).
• Construction - May be constructed from any suitable material.
Primary considerations are: temperature, pressure, abrasiveness
and corrosive tendencies of the gas or particulate matter.
D-142
-------
,' H, •
£-'•>' "VJT75.,,.
i-~' *$XJ»
1 -^ V^x/
Single Tube
Multiple Tube
Parallel
Arrangement
Multiple Tube
Series
Arrangement
Figure D-21.
Cyclone - Illustrating Single and Multiple Tube Arrangements
for Parallel or Series Operation(2)
D-143
-------
2.3 Feed Stream Requirements
• Temperature - Relatively unaffected by temperature^ ; construction
materials and methods must take temperature into account.
(4)
t Pressure - Relatively unaffected by pressurev ; construction
materials and methods must take pressure into account.
t Particle Size and Composition - Cyclones are generally limited to
applications where particle size is about 5 ym. Multi-cyclone
units with high pressure drop have been applied for removal of
particles down to 3 ym.
2.4 Process Efficiency and Reliability^4' - Efficiency is dependent
on: particle size and density, loading, inlet velocity, cyclone
dimensions and the gas density and viscosity. The general range of
efficiencies are:
Particle Size, ym Range of Efficiency, %
5 ym 50 to 80
5 to 20 ym 80 to 95
Reliability is high due to a simple system with no moving parts.
Some problems with removal of collected particulates (in coal gasifi-
cation gasifiers) has been reported^ .
2.5 Raw Material Requirements - none
2.6 Utility Requirements
• Electricity - To overcome the pressure drop across the system
(to power fan or blower).
2.7 Miscellaneous - Maintenance requirements are very low due to a
simple system with no moving parts.
3.0 Process Advantages
• Mechanically simple
• Highly reliable
t Relatively small space requirements
• Can handle hot, high pressure gas streams with little change in
efficiency
D-144
-------
t Low cost
• Low energy consumption
4.0 Process.Limitations
• Inability to collect small particles (below 5 ym)
• Large gas flows require multiple units
5.0 Process Economics
Due to the wide variety of applications, sizes, materials of construc-
tion and types of cyclones in use, no generalized cost figures are
available.
6.0 Input Streams
Due to the many variables involved in determining cyclone performance,
a generalized figure is presented (Figure D-22). Efficiency is plotted
against particle size for various pressure drops. The figure shows the
rapid deterioration of efficiency as particle size decreases. The figure
also shows the increase in efficiency with increased pressure drop.
7.0 Discharge Streams - see Figure D-22.
8.0 Data Gaps and Limitations
The performance of cyclones can be accurately predicted once particu-
late and gas stream parameters are known. Characterization, with
respect to these parameters, of coal gasification streams is needed to
determine applicability and performance for specific cases.
9.0 Related Programs - Not applicable.
D-145
-------
n. wg
in. wg
3.0 in. wg
10 15 20
Particle diameter, microns
Figure D-22. Cyclone Efficiency vs. Particle Size and Pressure Drop
(2)
REFERENCES
1. Product Guide, JAPCA, Vol. 27, No. 3, 1977.
2. Walker, A.B., Operating Principles of Air Pollution Control Equipment:
Guidelines for Their Appl ication, Research-Cottrell.
3. Hesketh, H.E. Understanding and Controlling Air Pollution, Ann Arbor
Science Publishers Inc., 1973.
4. Lasater, R.C. and Hopkins, J.H., Removing Particulates from Stack
Gases, Chemical Engineering, October 17, 1977.
5. Haynes, 'W.P., et al, Synthane Process Update, Mid-77, presented at the
4th Annual Conference on Coal Gasification, Liauefaction and Conversion
to Electricity, University of Pittsburg, August 2-4, 1977.
D-146
-------
Hydrocarbon and Carbon Monoxide Control Module
Thermal Oxidation
Catalytic Oxidation
Activated Carbon Adsorption (see Methanation
Guard Module, Appendix B)
D-147
-------
THERMAL OXIDATION PROCESS
(Direct-Flame Afterburners)
1.0 General Information
1.1 Operating Principle - The oxidation of combustible compounds
(e.g. hydrocarbons, CO, H-S) from many types of industrial waste
gas streams by direct combustion. In practice, thermal oxidizers
are generally used for the destruction of residual combustibles after
bulk of such materials is removed by prior treatment (e.g., by con-
ventional incineration). The typically low concentration of com-
bustibles in such waste gases usually requires that supplemental
fuel be used.
1.2 Development Status - Commercially available.
1.3 Licensor/Developer - Many companies manufacture direct-flame after-
burners; some systems incorporate certain proprietary features. A
complete listing of manufacturers are presented in technical and
trade journals (e.g. Reference 1).
1.4 Commercial Applications - Widely used in various industrial applica-
tions to control odors, smoke, total hydrocarbons and carbon mon-
(2)
oxidev . Potential applications of thermal oxidation in a coal
gasification facility may be in connection with emission control from
lock hoppers, Claus plant and regeneration of process catalysts.
2.0 Process Information
2.1 Flow Diagram - see Figure D-23.
t Process Description: The contaminated gas (Stream 1) enters the
unit, passes through the burner flames in the upstream part of the
unit. The hot gaseous mixture then passes through the remaining
part of the chamber where the combustion process is complete,
prior to being discharged to the atmosphere.
D-148
-------
TEMPERATURE
INDICATOR
IO
TEMPERATURE
CONTROLLER
BURNERS
LEGEND:
1. Inlet Gas Stream
2. Purified Gas Stream
Figure D-23. Direct-Fired Afterburner (Thermal Oxidation)
-------
2.2 Equipment - Conventional refractory- lined chamber, one or more
burners, temperature indicator-controllers.
2.3 Feed Stream Requirements
• Temperature: The inlet gas at any temperature can generally
be handled; however, the lower the temperature, the greater would
be the requirements for preheating and hence the supplementary
fuel requirements.
t Pressure: No limitation.
• Gas Composition: Many waste gas streams containing a wide variety
of combustible materials can be treated with thermal oxidation.
2.4 Operating Parameters
2.4.1 Combustion Chamber*
• Temperature: 839°K to 1005°K (1050°F to 1350°F)t;
1005°K to 1089°K (1350°F to 1500°F)t
• Pressure drop: 1.2 to 4.9 cPa (0.5 in. to 2.0 in. water)
• Residence time: 0.3 to 1.0 second
t Superficial velocity: 1.23 to 15.25 meters/sec
(4 to 5 afps)
2.5 Process Efficiency and Reliability - Through the control of tempera-
ture (supplemental fuel and air addition), different levels of com-
bustion efficiency can be achieved in response to changes in inlet
gas characteristics. In typical applications, 90% reduction of
hydrocarbons can be achieved. Years of operation have proven the
system reliable and efficient^ '.
2.6 Raw Material Requirements - None.
*Basis: 90% elimination of carbonaceous material from gas stream as determined
by the following equation for operating efficiency:
Hydrocarbons In - [Hydrocarbons Out + (CO out - CO in)] irvn
Hydrocarbons In x 10°
tApproximate temperature requirements for hydrocarbon control.
^Approximate temperature requirement of hydrocarbon and CO control.
D-150
-------
2.7 Utility Requirements
~in¥?l re1u1:ei1ients vary as a function of the type of fuel
aS Stream ter"Perature and the amount of combustibles
Table D"34 gives some
2.8 Miscellaneous - No unusual maintenance or hazardous conditions are
reported.
TABLE D-34. SUPPLEMENTAL FUEL REQUIREMENT FOR DIRECT-FIRED OXIDIZER AS A
FUNCTION OF RAW GAS TEMPERATURE AND HYDROCARBON CONCENTRATIONS (AS
HEXANE)*(5)
Natural Gas as Fuel
Cone, of Hydrocarbons, ppm
Raw Gas Temperature, °K (°F)
311°K (100°F)
366°K (200°F)
477°K (400°F)
589°K (600°F)
700°K (800° F)
Oil as Fuel
Cone, of Hydrocarbon, ppm
Raw Gas Temperature, °K (°F)
311°K (100°F)
366°K (200°F)
477°K (400°F)
589°K (600°F)
700° K (800°F)
Volume/hr of fuel to volume/mi n waste gas
0
2.6
-2.42
2.22
1.61
1.22
1000
2.18
2.0
1.62
1.21
0.81
2000
1.73
1.58
1.22
0.81
0.41
3000
1.3
1.18
0.8
0.04
—
3
Liters of oil per hour per 100 Nm per minute
waste gas (GPH/MSCFH waste gas)
0
16.7 (17)
15.7 (16)
13.0 (13.3)
10.8 (11)
8.3 (8.3)
1000
13.7 (14)
12.7 (13)
10.3 (10.5)
8.0 (8.2)
3.2 (3.3)
2000
13.2 (13.5)
9.8 (10)
7.8 (8)
5.4 (5.5)
2.9 (3)
3000
8.5 (8.7)
7.4 (7.5)
4.9 (.5)
2.3 (2.4)
—
*Assumes an oxidizer operating temperature of 1033°K (1400°F).
D-15T
-------
3.0 Process Advantages
• Process has been successfully used for years as an air pollution
control device to reduce hydrocarbons and various other combustible
contaminants from gas streams.
• May be easily installed as a retro-fit to existing installations.
• The performance of the system does not deteriorate with time.
• There is only one control point variable, chamber temperature.
The temperature and hence the process efficiency is readily controlled
by varying the fuel flow rate.
4.0 Process Limitations
(2,3)
• Requires supplemental fuel to raise raw gas temperature ' .
• Will not oxidize/ remove contaminants which are already in oxidized form
(e.g., S02, S03)(2'.
• Oxidation or partial oxidation of contaminants containing halogens
may create extremely hazardous effluent (e.g., hydrogen chloride,
phosgene) (2).
• At lower operating temperatures, only partial oxidation of organics
may be achieved and the decrease in organics concentration may be
accompanied by an increase in the CO level. When the generation of
CO is considered in the determination of the system efficiency,* the
efficiency of a direct- fixed boiler is usually less than 90% at oper-
ating temperatures below 477*K
5.0 Process Economics
Typical costs of a direct flame unit is as follows:
Equipment - $15.00 to $30.00 per 100 Nm3/min ($5.00 to $10.00/scfm)
Fuel - $0.00 to $80.00 per 100 Nm3/min per year ($0.00 to $20.00
per 1000 scfm per year)
6.0 Input Streams
6.1 Feed Gas Stream (see Figure D-23, Stream 1) - This stream generally
contains hydrocarbons, CO, CO^, and possibly small amounts of H^S,
S02» COS and various trace elements. The stream may be a discharge
"r-ff. • . _ Hydrocarbons in - [hydrocarbons out + (CO out - CO in)] inn
cTTiciency Hydrocarbons in x IUU
D-152
-------
stream from sulfur recovery plant tail gas treatment plants or
other air pollution abatement equipment.
6.2 Fuel - Quantities required depend on raw gas temperature, operating
efficiency desired and type of fuel used (see Table D-34.
7.0 Discharge Streams
7.1 Purified Gas Stream (see Figure D-23, Stream 2) - This stream will
contain primarily C02, H20 and N2 with small amounts of SO , NO and
traces of hydrocarbons and particulate matter.
8.0 Data Gaps and Limitations
Extensive performance data are available for direct-fired oxidation
applications to a variety of industrial gas cleaning operations. Evalua-
tion of expected performance of the system in applications to coal gasifi-
cation plant waste gas streams requires data on detailed characterization
of the gas to be treated. Such data which include gas temperature, and
chemical characteristics including a trace element survey, are generally
currently either not available or are incomplete.
9.0 Related Programs - No available data.
REFERENCES
1. Environmental Control Issue, Control Equipment, Environmental Science
and Technology, October 1977.
{. Waid, D.W., Afterburners for Control of Gaseous Hydrocarbons and Odors,
AIChE Symposium Series No. 137, Vol. 70, 1974.
3 Weiss S.M., Direct-Flame Afterburners, Air Pollution Engineering Manual,
2nd Edition (AP-40) U.S. EPA, May 1973.
4. Hesketh, H.E., Understanding and Controlling Air Pollution, Ann Arbor
Science Publishers, Inc., 1972.
5. Rolke, R.W., et al, Afterburner System Study. Report No. EPA-R2-72-062,
NTIS No. PB-212 560, EPA Contract No. EHS-D-71-3, Emeryville, California,
Shell Development Co., August 1972.
D-153
-------
CATALYTIC OXIDATION PROCESS
1.0 General Information
1.1 Operating Principles - The oxidation of combustible compounds
(e.g. hydrocarbons, CO, H2S) in a gas stream by passing the gas
stream through a catalyst bed. The catalysts most commonly used
are precious metals (e.g., platinum and palladium) supported on
various carrier materials (e.g. alumina, nickel). Various other
catalysts (e.g. metallic oxides, copper chromite) can also be used.
In practice, catalytic oxidizers are generally used for the destruc-
tion of residual pollutants in a gas stream, after bulk of such
pollutants are removed by prior treatment (e.g., thermal oxidation
or conventional incineration). Use of supplementary fuel is usually
required.
1.2 Development Status - Commercially available.
1.3 Licensor/Developer - Many companies manufacture catalytic oxidation
systems; most systems generally incorporate certain proprietary
features and catalytic formulation. A listing of major manufacturers
are presented in technical and trade journals (e.g., Reference 1).
1.4 Commercial Applications - Catalytic oxidation units have been applied
to coke ovens, catalytic cracking units, and a large number of
m
miscellaneous industrial gas cleaning operationsv '. Potential appli-
cations of catalytic oxidation in a coal gasification plant may be
in connection with the control of emissions from lockhoppers, Claus
plant and regeneration of process catalysts.
2.0 Process Information
2.1 Flow Diagram - see Figure D-24
• Process Description: The contaminated gas, Stream 1, is blown
into a preheat zone, where it is heated to the required tempera-
ture. The gas then flows through the catalytic element. The
D-154
-------
Figure D-24. Catalytic Oxidation
en
01
TEMPERATURE
PREHEATER BURNERS
TEMPERATURE
CONTROL
AIR
LEGEND:
1. Contaminated Gas
2. Purified Gas
-------
purified gas, Stream 2, exits the unit and can be released
directly to the atmosphere, or sent on for further treatment and
heat recovery(3).
2.2 Equipment - Conventional refractory-lined chamber, one or more
burners, catalyst chamber, and where applicable,heat recovery
(3)
equipment, temperature indicator-controlled .
2.3 Feed Stream Requirements
• Temperature: The inlet gas at any temperature can generally be
handled; however, the lower the inlet gas temperature, the
greater would be the requirements for preheating and hence the
supplementary fuel requirements.
• Pressure: None.
• Gas Composition: Many waste gas streams containing a wide variety
of oxidizable material can be treated by catalytic oxidation; how-
ever, the presence of heavy metals (e.g., mercury, lead, arsenic),
certain sulfur compounds, resin solids and particulates in the
gas stream can seriously effect the performance of the system due
to plugging and/or poisoning of the catalyst^3'.
(3 4)
2.4 Operating Parametersv ' '
2.4.1 Preheat Chamber
t Temperature: 589°K to 8H°K (600°F to 1000°F)
• Pressure: 0.1 to 13.7 MPa (atmospheric to 2000 psig)
2.4.2 Catalytic Chamber
• Temperature: For many applications, the temperature is
111°K (200°F) higher than the preheat chamber due to
exothermic chemical reactions occurring in the catalyst bed.
• Pressure: Approximately the same as in the preheat chamber;
any difference is due to pressure drop across catalyst bed.
2.5 Process Efficiency and Reliability
• The efficiency of a catalytic oxidizer is a function of several
variables including(3):
- operating temperature of the unit
- type of catalyst being used, including surface area, bed
depth, amount of catalyst to volume of gas being treated.
D-156
-------
- nature of contaminants to be oxidized
- uniformity of gas flow through catalyst bed
t Catalytic oxidizing units can be selected and/or designed to
remove 90% or more of the oxidizable materials from most qas
streams U).
• With proper catalyst maintenance and replacement the system
will continuously meet efficiency requirements.
2.6 Raw Material Requirements
• Catalyst makeup: Depending upon the type of catalyst used, gas
stream components, and flow rate, the interval between catalyst
replacement can vary from a few months to as long as two yearsi3/.
2.7 Utility Requirements
• Fuel: Fuel requirements vary as a function of the type of fuel
used, inlet gas stream temperature and the amount of oxidizable
chemicals present in the gas stream. Table D-35 gives some
typical fuel requirements.
2.8 Miscellaneous - Depending on the nature of the contaminants present
in the inlet gas stream, frequent changing of the catalyst may be
required particularly if catalyst poisons (arsenic, mercury, tars,
etc.) are present in the inlet stream. The handling and disposal
of the spent catalyst may require special attention because it may
be contaminated with various toxic substances, such as heavy metals.
3.0 Process Advantages
• Catalytic oxidation is particularly suitable for removing small
amounts of oxidizable contaminants.
• The system can be readily applied (either by incorporation in original
design or by retrofit) to a variety of industrial applications as an
an air pollution abatement device(5).
• Lower supplemental fuel requirements than conventional thermal
oxidizers(S).
D-157
-------
TABLE D-35. SUPPLEMENTAL FUEL REQUIREMENTS FOR CATALYTIC OXIDIZERS AS A
FUNCTION OF RAW GAS TEMPERATURE AND HYDROCARBON CONCENTRATIONS
(AS HEXANE)*(8)
Natural Gas as Fuel
Cone, of Hydrocarbons, ppm
Raw Gas Temperature- °K (°F)
311°K (100°F)
366°K (200°F)
377°K (400°F
589° F (600°F)
Oil as Fuel
Cone, of Hydrocarbons, ppm
Raw Gas Temperature, °K (°F)
311°K (100°F)
366°K (200°F)
477°K (400°F)
589°K (600°F)
Volume/hr fuel to volume/mi n waste gas
0
1.3
1.19
0.90
0.53
1000
0.95
0.80
0.48
--
2000
0.55
0.45
—
—
3000
0.2
--
--
—
3
Liters of oil per hour per 100 Nm per minute
waste gas (gph fuel/Mscfh waste gas)
0
8.3 (8.5)
7.3 (7.4)
5.6 (5.7)
3,4 (3.5)
1000
6.1 (6.2)
5.1 (5.2)
3.0 (3.1)
—
2000
3.7 (3.8)
2.7 (2.8)
--
--
3000
1.4 (1.5)
0.5 (0.5)
--
--
*Assumes a gas exit temperature of 775°K (900°F).
4.0 Process Limitations
t The efficiency of a catalytic oxidation unit deteriorates with use,
due to catalyst degradation caused by age and contaminations).
• The catalyst must be periodically changed; the spent catalyst (which
may contain hazardous substances) must be handled and a suitable means
of disposal must be used(2,4,5)
• The products of the catalytic oxidation process may still constitute
an air pollution problem (e.g., H2S + 02 = H20 + SO )(4).
t Catalysts are susceptible to poisoning and/or fouling by heavy metals,
sulfur oxides and resin solids(3).
D-158
-------
5.0 Process Economics^ '
Typical costs of a catalytic oxidation unit are as follows:
Equipment - $5.00 to $15.00 per 100 Nm3/m ($1.75 to $5.00 per scfm)
Fuel - $0.00 to $18.00 per 100 Nm3/m per year ($0.00 to $4.50/1000
scfm per year)
6.0 Input Streams
6.1 Feed Gas Stream (see Figure D-24, Stream 1). This stream is
generally made up of hydrocarbons, CO, C02> H2§ and possibly small
amounts of H2S, S02, COS, and various trace elements. This stream
may be the tail gas stream from a sulfur recovery plant, a tail gas
treatment plant or other air pollution abatement equipment.
6.2 Catlyst replacement/makeup - Depends on the nature of catalyst and
operating conditions.
7.0 Discharge Streams
7.1 Purified gas stream (see Figure D-24, Stream 2) - This stream will
contain primarily C09, H90 and N9 with small amounts of SO , NO
£. C. L. XX
and traces of hydrocarbons artd particulate matter.
7.2 Spent catalyst - no data available.
8.0 Data Gaps and Limitations
Extensive performance data are available for catalytic oxidation
applications to a variety of industrial gas cleaning operations. Evalua-
tion of expected performance of the system in applications to coal gasifi-
cation plant waste gas streams requires data on detailed characteristics
of the gas to be treated. Such data which include gas temperature, and
chemical characteristics including a trace element survey are generally
either not available or are incomplete.
9.0 Related Programs - No available data.
D-159
-------
REFERENCES
1. Environmental Control Issue, Control Equipment, Environmental Science
and Technology, October 1977.
2. Riesenfeld, F.C., and Kohl, A.L., Gas Purification, 2nd Edition, Gulf
Publishing Company, 1974.
3. Weiss, S.M., Catalytic Afterburners, Air Pollution Engineering Manual,
2nd Edition (AP-40) U.S.E.P.A., May 1973.
4. Hawthorn, R.D., Afterburner Catalysts - Effects of Heat and Mass Transfer
Between Gas and Catalyst Surface, AIChE Symposium Series No. 137, Vol. 70,
1974.
5. Farkas Adalbert, What You Should Know. . .Catalytic Hydrocarbon Oxidation,
Hydrocarbon Processing, July 1970.
6. Miller, M.R. and Wilhoyte, H.J., A Study of Catalyst Support Systems for
Fume-Abatement of Hydrocarbon Solvents, Journal of the Air Pollution Con-
trol Association, Vol. 17, No. 12, December 1967.
7. Hesketh, H.E., Understanding and Controlling Air Pollution, Ann Arbor
Science Publishers, Inc., 1972.
8. Rolke, R.W., et al, Afterburner Systems Study. Report No. EPA-R2-72-062,
NTIS No. PB-212 560, EPA Contract No. EHS-D-71-3, Emeryville, California,
Shell Development Co., August 1972.
D-160
-------
APPENDIX E
WATER POLLUTION CONTROL
Oil and Suspended Solids Removal Module
Gravity Separation (API Separators)
Flotation
Filtration
Coagulation-Flocculation
E-l
-------
GRAVITY SEPARATION (API SEPARATORS) PROCESS
1.0 General Information
1.1 Operating Principles - Free oil is separated from a wastewater by
retaining the wastewater in a basin where the oil globules (having
a lower density than water) rise to the surface under the influence
of gravity and are collected at the surface. Particles denser than
water which may settle to the bottom are collected as bottom sludge.
1.2 Development Status - Oily water separators have been used indus-
trially for decades to recover oil from process waste streams.
1.3 Licensor/Developer - The American Petroleum Institute has developed
guidelines for design of rectangular shaped gravity separators
(known as API separators). API and other types of separators,
especially the smaller parallel plate separators, are offered by a
number of water pollution control equipment suppliers.
1.4 Commercial Applications - Widely employed in refineries for removal
of oil and suspended solids.
2.0 Process Information
Three of the most common types of gravity separators are the API,
circular and parallel plate separators.
• The API Separator [Figure E-l(a)] - The wastewater enters the basin
and passes under the oil retention baffle, then over the diffusion
baffle (to minimize turbulence). As the wastewater travels the
length of the channel, the oil globules move toward the surface and the
heavy particles settle downward. Flight scrappers push the oil which
has reached the surface towards one end and into the slotted pipe for
removal. At the same time the flight scrappers push sludge deposits
on the bottom of the basin to sludge hoppers. Clarified water passec
under the oil retention baffle No. 2 and leaves the unit.
E-2
-------
oil retention baffle
diffusion baffle
(a). API Separator
diffusion baffle
oil retention
baffle
LEGEND:
1 - Influent
2 - Clarified Effluent
3 - Collected Oil
4 - Sludge
(b). Parallel Plate Separator
Figure E-1. Gravity Separators
E-3
-------
• Circular Separators (not shown) - In general, these types of separators
are designed similar to the conventional circular clarifier. In most
designs, influent enters at a central location in the circular tank
and has a peripheral dischargeO). Pilot plant studies indicate that
circular units may be as effective as an API separator^ ' »5;.
0 Parallel Plate (or tube) Separators [Figure E-l(b)] - The wastewater
enters the separator and flows over a weir and through the parallel
plates (or tubes). The plates can either be corrogated or flat. The
oil particles coalesce on the under side of the plates and rise up
to the surface where they are removed. Solids collect on the bottom
of the plates and slide downward towards the sludge hopper for removal.
2.1 Equipment
e Separation tank - Concrete, tile or coated steel. The design
criteria for rectangular API type separators are as follows(2):
- Horizontal velocity (VH) maximum = 0.91 m/min (3 ft/min)
or 15 Vt (Vt* = rate of rise of oil) whichever is smaller.
• Depth = 0.9 m (3 ft) minimum to 2.44 m (8 ft) maximum
• Depth-to-width ratio - 0.3 minimum to 0.5 maximum
• Width = 6 ft minimum to 20 ft maximum
2.2 Feed Stream Requirements - Due to large variation in wastewater
characteristics (e.g., specific gravity of oil, concentration of
settleable solids, temperature of feed stream, presence or absence
of emulsions) design of the separator is "tailored" to the specific
wastewater to be treated (V. determined in laboratory tests)' .
2.3 Operating Parameters - Operating parameters are variable because
each separator is designed to handle a specific influent. The
vertical rate of rise of the oil globules (Vt) ideally will be equal
to the overflow rates ^ '. In most cases, however, turbulence and
short circuiting affect the efficiency of oil removal. Ideally, as
v = 0.0241 (-^ -) where Vt = rate of rise of oil gTobules (0.015 cm in
diameter) in wastewater, in fpm; S and S = specific gravity of water and oil
in wastewater at design temperature, respectively; and N = absolute viscosity
of wastewater at design temperature, in poises.
E-4
-------
long as Vt is greater than the overflow rate oil will be removed
and not carried out in the clarified effluent.
2.4 Process Efficiency and Reliability - Typical efficiencies of
various oil separation units and parallel plate separators are shown
in Tables E-l and E-2. API separators are widely used in the
petroleum refining industry; the system has proved effective and
reliable for separation of oil from wastewaters.
2.5 Raw Material Requirements - None.
2.6 Utility Requirements
y(6
(5)
Electricity^ - 14.9 kwh/1000 gal
3.0 Process Advantages
• General
- Economical^ '
(5)
- Simple Operationv '
« Parallel Plate Separator
- 15%-20% of the installation area of regular separators^ '
(3)
- Removal efficiencies generally higher than for regular separatorsv '
(see Figure E-2)
4.0 Process Limitations
• Removal efficiency decreases as the wastewater temperature drops
• Removes little or no soluble and emulsified oils^ '
• API type separators are designed to affect complete removal of oil
globules with diameters equal to or greater than 0.015 cm (0.006 in.);
smaller particles would be removed only fractionally.
• Peak flow rates may decrease removal efficiency due to rise in the
overflow rates (see Figure E-3)
5.0 Process Economics
Capital cost for a 3,780 1/min (1000 gal/min) capacity gravity
separator for treatment of oily wastewaters depends largely on the
desired effluent oil concentration (see Figure E-4). Operating cost
E-5
-------
TABLE E-l. TYPICAL EFFICIENCIES OF OIL SEPARATION UNITS
(3)
Oil Content
Influent,
mg/1
300
220
108
108
98
100
42
2,000
1,250
1,400
Effluent,
mg/1
40
49
20
50
44
40
20
746
170
270
Oil
% Removed
87
78
82
54
55
60
52
63
87
81
Type of
Separator
Parallel plate
API
Circular
Circular
API
API
API
API
API
API
CCD
% Removed
—
45
—
16
--
—
— :
22
--
—
SS
% Removed
--
--
—
—
—
--
—
33
68
35
TABLE E-2. OIL REMOVAL, TILTED-PLATE SEPARATOR
(3)
Oily Water
Throughput
1/hr (gal/hr)
2100 (8000)
2100 (8000)
2100 (8000)
4200 (16,000)
4800 (16,000)
4800 (18,600)
4800 (18,600)
Influent
Oil
mg/1
150
375
500
500
500
470
700
Effluent
Oil
mg/1
50
66
86
178
190
185
330
Percent
Removal
67
82
83
65
62
67
53
E-6
-------
400r
300
5 200
!00
Tilted-plqte separator
— — Existing API gravity separator
100 200
Influent oil, mg/l
300 400
Figure E-2.
Removal Efficiency of Pilot-Scale Titled-Plate Separator
Compared to Full-Scale API Separator(3)
100
§
-------
2.50
150
125 100 75
Oil in Effluent, ppm
50 25
Figure E-4. Capital Cost vs. Quality of Effluent
E-8
-------
for a 39780 1/min (1000 gpm) capacity gravity separator is $72,750
annually (includes depreciation of plant at 10%).
6.0 Input Streams
6.1 Raw Wastewater (Stream!), Figure E-l - See Table E-1 and E-2.
7.0 Discharge Streams
7.1 Clarified Effluent (Stream 2), Figure E-l - See Tables E-l and E-2.
7.2 Recovered Oil (Stream 3), Figure E-l - No data available.
7.3 Sludge from Settled Solids (Stream 4), Figure E-l - 3.3% to 59.8%
oil (average 22.6%) and 7%-98% oil (average 53%)
8.0 Data Gaps and Limitations
No available data on the use of gravity separation for the treatment
of wastewaters from coal conversion facilities. Also, no available
data covering feed stream requirements and operating parameters.
9.0 Related Programs
Not known.
REFERENCES
1. Petroleum Refining - Development Document for Effluent Limitations
Guidelines and New Source Performance Standards, U.S. EPA Contract
No. 440/l-74-014a,
2. American Petroleum Institute, Manual on Disposal of Refinery Wastes,
Volume on Liquid Wastes, Chapter 9, 1969.
3. Azad, Hardman S. Industrial Waste Management Handbook, McGraw Hill Book
Co., 1976.
4. Eckenfelder, Industrial Water Pollution Control, McGraw Hill Book Co.,
1966.
5. Ford, Davis L., et al , Removal of Oil and Grease from Industrial Waste-
waters, Chemical Engineering Deskbook Issue, October 17, 1977.
6. Thompson, C.S., et al , Cost and Operating Factors for Treatment of Oily
Wastewater, Oil and Gas Journal, November 20, 1972.
7. Jacobs Engineering Co., Assessment of Hazardous Waste Practices in the
Petroleum Refining Industry, NTIS No. PB-259-097, USEPA-SW-1296, June 1976.
E-9
-------
FLOTATION PROCESS
1.0 General Information
1.1 Operating Principle - Separation of solid or liquid particles from
a liquid phase by the addition of a gas (usually air and under
pressure) to the waste stream (or a fraction thereof) and subsequent
release to atmospheric pressure, thereby forming fine bubbles which
adhere to and are trapped in the particle structure reducing the gross
particle density and hence causing the particles to rise to the sur-
face where they can be removed by skimming. Removal efficiency can
be enhanced by the addition of chemical flocculants, particularly
when colloidal or emulsified oils are present^ .
1.2 Development Status - Has been used for decades in industrial and
municipal wastewater treatment.
1.3 Licensor/Developer - Not a patented or proprietary process; flota-
tion systems/equipment are offered by a large number of water pol-
lution control equipment suppliers.
1.4 Commercial Applications - Dissolved air flotation (DAF) is used
by a number of refineries for the removal of oil from wastewaters.
In a few refineries dissolved air flotation is used to clarify
(2\
biologically treated effluentsv '. Flotation has also been used
for treatment of other industrial wastes (e.g., pulp and paper
wastewater) and for sludge thickening. Total flow, split flow
and recycle pressurization are the most common DAF systems used
in refinery applicationsv
E-10
-------
2.0 Process Information (see Figure E-5)
Three most common variations of DAF are total flow, split flow
and recycle flow pressurization (depending on whether the entire raw
wastewater flow or a portion of the raw wastewater flow or treated
effluent is pressurized).
The pressurization is carried out in a pressurized chamber where
a short retention is provided for air dissolution. The pressurized
liquid is discharged directly to the flotation chamber (total flow
pressurization) or mixed with the entire or the remaining portion
of the raw wastewater (recycle flow or split-flow pressurization,
respectively) and then discharged to the flotation chamber. The solids
which float to the surface are skimmed off in the flotation chamber.
• Desirable features of total flow pressurization include
- bubbles are released through entire volume of wastewater
- smaller flotation chamber is required than for recycle
pressurization.
• Some desirable features of split flow and recycle pressurization^ ' ' ':
- requires smaller pressurizing pump and reduces pumping cost
especially if the DAF system is gravity fed(5)
- pump control is easier and can be run at a constant rate
- reduces amount of emulsion that would be formed if all influent
was pressurized
- allows optimum floe formation in portion of feed stream that
bypasses pressurization system.
2.1 Equipment
t pressurization pump
t air injection equipment
t pressurization tank
• pressure regulating device
• flotation chamber; rectangular or circular chambers made of
concrete, tile or coated steel(3)
E-ll
-------
FLOTATION CHAMBER
a. Total Flow Pressurization
»• FLOTATION CHAMBER
b. Split Flow Pressurization
1. WASTE INFLUENT
2. FLOCCULATING AGENT (OPTIONAL)
3. AIR
4. PRESSURIZED STREAM
5. OIL SCUM
6. CLARIFIED EFFLUENT
i
1
V \
1 ^
J2
1
LEGEND
4
/^"\ /
r*Jt*i i F
FLOTATION CHAMBER
RESSURE \ f
6
1-^ i \ CHAMBER 1^ f V
c. Recycle Flow Pressurization
Figure E-5. Dissolved Air Flotation Process
E-12
-------
e skimming equipment - flight scrapers or other design
c chemical feeding equipment (if required).
2.2 Feed Stream Requirements - Because of a large number of factors
which affect flotation efficiency (e.g., temperature, and nature
and concentration of solids), design of flotation systems are
usually "tailored" to specific waste application. The design
criteria (air-to-solid ratio,* solid rise rate, recycle ratio,
etc.) are developed based on laboratory/bench-scale flotation
tests. In general, very high concentrations of separable oil
(greater than about 1000 ppm) and high wastewater temperature
reduce process efficiency^ ' '. (Solubility of air in water
decreases with the rise in temperature.)
2.3 Operating Parameters
9 Stream pressurization
- total flow system(2'4'6'8) -0.31-0.52 MPa (45-75 psia)
- partial flow(3) -0.52-0.62 MPa (75-90 psia)
• Retention time
- pressure retention tank^ ' ' ' - 1-5 minutes
- flotation chamber^1'2'3' - 10-40 minutes
« Flotation chamber overflow rates
- total flow and split flow(3>4) - 81-4-102 1/min/m2
(2.0-2.5 gpm/ft2)
- recycle flow^ - 40.7-61 1/min/m2 (1.0-1.5 gpm/ft )
G Air-to-solids ratio - 0.02-0.06
tion factor (less than 1.0); Sa = concentration Sf air in wastewater at
saturation at wastewater temperature, cm3/!; R = pressurized volume, 1,
P = absolute pressure, atm; Q = waste flow, 1; S = influent suspended
solids concentration, mg/1.
E-13
-------
2.4 Process Efficiency and Reliability - Depending on wastewater
characteristics, unit design/loadings and use of chemical aids,
oil, suspended solids and BOD removal ranges of 50% to 100%, 30% to
(3 4 5)
80%, and 30% to 50% may be expected, respectively^ ' ' '. Properly
designed and operated flotation units are in operation in many
plants; extensive records of trouble-free operation are available
for these units.
2.5 Raw Material Requirements
• Air at 0.38 MPa (55 psia) - 0.035 to 0.07 SCM per 1000
(5 to 10 SCF per 1000 gal) of pressurized wastev3).
• Chemical aids^ ' ' - Alum, ferric salts, and activated silica are
used. Alum and ferric salts are added before or at the pressuri-
zation pump. Activated silica is added downstream of the pressure
release valve. Amount of chemicals used depends on the type and
quantity of the effluent. Common concentrations of alum used in
DAF systems are 100-130 mg/lO).
2.6 Utility Requirements
t Electricity^ - 0.55 kwh/1000 gal
3.0 Process Advantages
• Handles fluctuations in feed rates weir1'3'
• Captured solids are low in volume compared to large volumes of
backwash in filtration(7)
,(5)
(5)
KSX,I i<_ i i v, i u i ill *7 VI l|^pllivj I l/-j*J UIIU Mlli
4.0 Process Limitations
• High temperatures reduce effectiveness^ '
• Does not remove soluble oil '
• Does not effectively remove oil emulsions without the use of
chemicalsO »3).
• Effectiveness is sometimes unpredictable^ '
• Effective in reducing BOD and
• DAF systems are beneficial in stripping H2$ and NH3
E-14
-------
5.0 Process Economics
Capital cost for a 3,780 1/nrin (1000 gal/min) capacity gas flotation
system for treatment of oily wastes is $330,000 (1972)^. Operating
cost for the same system is $80,190 annually (1972{5^.
6.0 Input Streams
6.1 Raw Wastewater (Stream 1), Figure E-5 (e.g., oily water from API
separator, usually 200-1000 mg/1 of free oir4^- (see Table E-3)
6.2 Air (see Section 2.3)
6.3 Chemical flocculation aids, if required (see Section 2.5)
7.0 Discharge Streams
7.1 Clarified effluent (Stream 6), Figure E-5 (see Table E-3)
7.2 Float (Stream 5), Figure E-5 - Percent solids in the float and
characteristics of the float (water content, settleability, etc.)
dependent on raw wastewater characteristics and design/operating
conditions. For many applications, the solids content of the float
is in the 1% to 4% range.
8.0 Data Gaps and Limitations
No data available on the use of flotation for the treatment of
wastewaters from coal conversion facilities.
9.0 Related Programs
Not known.
E-15
-------
TABLE E-3. DISSOLVED AIR FLOTATION - PERFORMANCE DATA
(1)
Influent Oil
mg/1
1930
580
105
68
170
125
100
133
94
638
153
75
61
360
Effluent Oil
mg/1
128
68
26
15
52
30
10
15
13
60
25
13
!
% Removal
93
88
78
75
70*
71
90
89
86
91
83
82
15 75
45 | 87
i
*No chemical additives used in this case; chemical additives
used in all other cases.
E-16
-------
REFERENCES
1. Ford, Davis L., et al, Removal of Oil and Grease from Industrial
Wastewaters, Chemical Engineering Deskbook Issue, October 17, 1977.
2. Petroleum Refining - Development Document for Effluent Limitation
Guidelines and New Source Performance Standards, U.S. EPA Contract
No. EPA-440/l-74-014a.
3. American Petroleum Institute, Manual on Disposal of Refinery Wastes,
Volume on Liquid Wastes, Chapter 9, 1969.
4. Liptak, E.G., Environmental Engineers Handbook, Volume 1 Water Pollution,
Chi 1 ton Book Co., Radnor, Penn., 1974.
5. Thompson, C., et al, Cost and Operating Factors for Treatment of Oily
Wastewater, Oil and Gas Journal, No. 47, p. 53, November 20, 1972.
6. Metcalf and Eddy, Inc., Wastewater Engineering, McGraw-Hill Book Co.,
New York, 1972.
7. Parsons, W.A. and W. Nolde, Abstract Applicability of Coke Plant Water
Treatment Technology to Coal Gasification, paper presented at EPA
Environmental Aspects of Fuel Conversion Technology Symposium, Hollywood,
Florida, September 13-15, 1977.
8. Ross, R.D., Industrial Waste Disposal, Reinhold Book Corp., 1968.
9. Eckenfelder, Industrial Water Pollution Control, McGraw-Hill Book Company,
New York, 1966.
E-17
-------
FILTRATION PROCESS
1.0 General Information
1.1 Operating Principle - Wastewater containing suspended solids is
passed through a bed of granular material, resulting in deposition
of the suspended solids in the bed. When the pressure drop across
the bed becomes excessive.the bed is cleaned by backwashing with
water. In some cases air scouring of the filter bed is implemented
to enhance cleaning*1 '.
1.2 Development Status - Filtration has been practiced for decades in water
treatment plants but only recently has it been used in the treatment
io\
of wastewaters* .
1.3 Licensor/Developer - Not a patented or proprietary process.
1.4 Commercial Applications - Sand filters are employed to polish
domestic water supplies and polish industrial wastewaters. Sand
filters are currently in use at the Lurgi-type gasification faci-
lity at Westfield, Scotland. Hay filters are used in petroleum
refining to remove suspended solids (SS) and adsorb oil. Diato-
maceous earth filters are used to obtain an extremely high quality
effluent. Mixed media filters are used to lengthen the filtering
cycle.
2.0 Process Information (see Figure E-6)
Wastewater enters the filtration unit and slowly percolates through
the filter media (i.e., sand, charcoal, diatomaceous earth, anthracite,
etc.) to the underdrain. As the filtration process proceeds, the
pressure drop across the filter (head loss) increases. (At a head loss
of 1.5 to 2.4 m of water column or whenever breakthrough occurs, the
filtration unit is backwashed to flush the collected SS from the inter-
stices)^ '. Backwashing proceeds by flushing water, and in some cases
E-18
-------
BACKWASH
DRAIN -*•:
HIGH HEAD
AIR
DIFFUSER |-T
BACKWASH
TROUGH
UNDER DRAIN
RAW FEED (1)
;. •• -,r:. • •.; ^1~. *, - *_-" "•., ;*•. : \ • f. •-,<-,
SINGLE OR
MULTIPLE LAYER
FILTER MEDIUM
AIR (2)
-BACKWASH (3)
-EFFLUENT (4)
Figure E-6. Typical Filtration Bed
E-19
-------
air, back through the filter bed via the underdrain. Backwash water is
collected in the backwash trough and exits the system via the backwash
drain.
2.1 Equipment
t Air injection equipment
• Filters media (sand, coal, anthacite, diatomaceous earth,
gravel )-60.9 to 71.4 cm (54 to 36 in.)
• Backwash equipment (pump to inject water)
(4)
t Piping - cast iron or coal-tar enamel-lined welded steelv '
• Concrete filter tank although small steel units may be pur-
chased as a whole or assembled in the field
2.2 Feed Stream Requirements
;tituent Present
in Influent in Influent
Constituent Present Acceptable Concentration^ '
Solids 100 mg/1
Fiber 10-25 mg/1
Particle size 200 y
Oil 25-75 mg/1
2.3 Operating Parameters*
Rate Operating Time
Backwash 84-1218 lpm/m2 5-15 min. ^4'7^
(2-29 gpm/ft2)(4,7,9)t
Air scour 0.6-1.5 Mm3 min/m2 3-10
(2-5 scfm/ft2)(l,4)
Waste Influent 168-210 lpm/m2
(4-5 gpm/ft2)(4,6-)
Rotating 31.5-42 lpm/m2
Surface wash (0.75-1.0 gpm/ft2
*Coal-sand mixed beds. ,
tOr a rate that will expand the filter bed 20%* »'',
E-20
-------
Begin backwash when breakthrough occurs or when head loss becomes
1.5 to 2.4 m (5 to 8 ft) of water column^.
2.4 Process Efficiency and Reliability - 60%-95% removal of suspended
solids can be expected for granular media filtration^1). Removal
efficiencies for other species, can be seen in Table E-4 data
obtained from operation at the Lake Tahoe reclamation plant. Relia-
bility of filtration is good providing that the system is backwashed
at proper intervals before breakthrough occurs' '. Also, oil, fiber
and suspended solids can be handled with good reliability as long as
their respective influent concentrations are kept within certain
limits. Fiber in concentrations greater than 10 to 15 mg/1 will
(5)
cause plugging problemsv . Filters equipped with water backwash
can handle 25 mg/1 of free oil while filtration equipped with
heavy duty air scouring can handle oil in concentrations of 50 to
75 mg/r . At higher oil concentrations the oil will cause plugging
of underdrains and prevent complete backwashing^ . High concen-
trations of suspended solids (above 100 mg/1) will cause reduced
cycle time* .
TABLE E-4. TYPICAL REMOVALS BY MIXED-MEDIA FILTERS FROM WASTEWATER
PRETREATING BY COAGULATION SEDIMENTATION
Substance
Phosphorus
COD
BOD
Suspended Solids
Turbidity
Range (% Removal)
70 - 95
20 - 45
40 - 70
100
60 - 95
E-21
-------
2.5 Raw Material Requirements
t Compressed air - 1.2 Nm3/m2 (4 scfm/ft2)
• Backwash water - clarified effluent (permissible to use effluent
from another filtration unit) 630-840 Ipm/m? (15-20 gpm/ft2) to
achieve 38-40% expansion(lO)
• Surface wash - 3.45 x 6.89 x 105 Pa (50-100 psi) supplied at
31.5 - 42 Ipm/m2 (0.75 - 1.0 gpm/ft2)
« Polymer aids may be added in doses of less than 0.1 mg/1 to
enhance filtration
2.6 Utility Requirements - None (except where pumping of influents is
required).
3.0 Process Advantages
• High removal efficiencies and reliabilities^ '
(4)
• Removes small amounts of suspended solids very well
4.0 Process Disadvantages
• Large volumes of backwash water needed^ '
(8)
t Does not handle shock loads wellv '
5.0 Process Economics
For 1000 gpm capacity plant^ ' (cost at 1967 dollars)
Capital cost per 1000 liters treated - 0.38<£
Operation and maintenance per 1000 liters treated - 1.4<£
6.0 Input Streams
6.1 Raw Wastewater (Stream 1), Figure E-6 - Wastewater is usually
secondary effluent from biologically or chemically clarified
waste streams
6.2 Air (Stream 2), Figure E-6 (compressed air)
6.3 Backwash water (Stream 3), Figure E-6 (clarified effluent from other
filtration unit)
7.0 Discharge Streams
7.1 Clarified effluent (Stream 4), Figure E-6 - 2-4 mg/1 of suspended
solids under good conditions^ '
7.2 Backwash effluent (Stream 5), Figure E-6
E-22
-------
8.0 Data Gaps and Limitations
No data were found on increased process efficiency when adding polymer
aids to filtration. No data available on the quality of backwash streams
9.0 Related Programs - Not known.
REFERENCES
1. Azad, Hardam S. Industrial Wastewater Management Handbook, McGraw-Hill
Book Co., New York, 1976.
2. American Petroleum Institute Manual on Proposal of Refinery Wastes,
Volume on Liquid Wastes, Chapter 9, 1969.
3. Advanced Waste Treatment Research Laboratory, Cincinnati, Ohio. Current
Status of Advanced Waste Treatment Processes, Federal Water Quality
Administration, U.S. Department of the Interior Publication PPB 1101,
July 1, 1970.
4. Culp, Russell L., and Culp, Gordon L. Advanced Wastewater Treatment,
Van Nostrand Reinhold Co., New York, 1971.
5. Liptak, B.G., Environmental Engineers Handbook, Volume 1, Water Pollution,
Chi 1 ton Book Co., Randor, Penn., 1974.
6. Cohen, Jesse M., Solids Removal Processes, Advanced Waste Treatment
Seminar on Removal of Solids and Organics, San Francisco, October 29
and 30, 1970.
7. Burns and Roe, Inc. Process Design Manual for Suspended Solids Removal
for EPA Technology Transfer, October 1971.
8. Metcalf and Eddy, Inc. Wastewater Engineering, McGraw Hill Book Co.,
New York, 1972.
9. Kriessl, James F. Granular Media Filtration of Secondary Effluent
U.S. EPA Advanced Waste Treatment Research, December 13, 1974.
10. Cleasby, J.L. and Baumann, E.R. Backwash of Granular Filters in
Wastewater Filtration, EPA-600/2-77-016, April 1977.
E-23
-------
COAGULATION-FLOCCULATION PROCESS
1.0 General Information
1.1 Operating Principle - Suspended solids and colloidal materials are
removed from wastewaters by the addition of chemical coagulants and
coagulant aids to produce finely divided precipitates or microflocs.
The process of coagulation is followed by flocculation of these
small particles into larger clumps or agglomerates which may be
removed by sedimentation. Coagulation results from (a) neutraliza-
tion of negative surface charges on colloidal particles by positively
charged metallic or polymeric ions used as coagulants (and/or their
hydrolysis products); and (b) the binding and enmeshing actions
of the metal hydroxide gels.
1.2 Developmental Status - Commercially available.
1.3 Licensor/Developer - Coagulation-flocculation treatment systems
and equipment are offered by numerous suppliers (Ref. 1). Chemical
coagulants, coagulant aids and pH adjustment chemicals are available
through various chemical supply houses (e.g., Calgon Corporation,
Betz, etc.).
1.4 Commercial Applications - An oil flocculation system is in use
at the SASOL Lurgi-type coal conversion facility, Sasolburg, So.
Africa^ '. Numerous applications to treatment of oily wastewaters
in petroleum refineries.
E-24
-------
2.0 Process Information
2.1 Flow Diagram (see Figure E-7)
• Process Description - Influent wastewaters (Stream 1) are fed
to a retention/equalization basin where they are mixed to produce
more constant feedwater quantity and quality. The equalized
wastewater (Stream 3) is pumped to the coagulation-flocculation
unit which usually consists of three chambers: a mixing chamber,
a flocculation compartment, and a sedimentation chamber. In the
mixing chamber, the wastewater is flash-mixed with chemical
flocculants, flocculation aids and pH adjustment chemicals by
means of vertical or horizontal mechanical paddles. The wastewater
passes from the mixing chamber to the flocculation compartment
where it is agitated by slowly moving paddles, then flows into the
sedimentation chamber by means of an inlet device which distributes
the waste uniformly throughout the cross-sectional area of the
chamber. Clarified waters (Stream 4) leave the sedimentation
chamber over a weir. Residual sludge (Stream 5) is scraped from
the bottom of the sedimentation chamber and discharged.
(3)
2.2 Equipment^ '
t Coagulation-flocculation unit - Two basic types: (a) the sludge-
blanket unit which combines mixing, flocculation and settling in
a single unit; and (b) the conventional system using a rapid mix
tank, followed by a flocculation tank containing longitudinal
paddles which provide slow mixing; and a conventional settling
tank.
• Retention/equalization basin
• Pumps
• Mechanical paddles
(3 4)
2.3 Feed Stream Requirementsv '
• Flow rate - For ease of operation and constant effluent
quality, the influent rate to the coagulation-flocculation unit
should be as uniform as possible.
e Composition - Composition of the influent waste should be as
uniform as possible to minimize the number of adjustments ot
chemical dosages required.
t pH - Optimum pH for coagulation, as determined by laboratory
tests, should be maintained.
E-25
-------
MECHANICAL
-
RETENTION
EQUIALIZATION
BASIN
® f \
^-^ (to )
>LA
r^\
n
SI
\'
"•'" N
^*
&
e__s
o
W/i"4?
-------
2.4 Operating Parameters(4>5) - Coagulation, flocculation and settling
times vary with the specific waste being treated; optimum conditions
can be determined by lab-scale testing of the wastewaters (e.g., "jar
tests"). The period for flash mixing of coagulation chemicals into
the waste usually varies from 30 seconds to 5 minutes while floccula-
tion times range 5-30 minutes. Retention time in the sedimentation
chamber varies from 2 to several hours. Requirements are determined
by sedimentation tests of the flocculated water. See Table E-5.
2.5 Process Efficiency and Reliability - Efficiency depends upon the
design of the system, coagulant dosage, temperature and on the
characteristics of the wastewaters. The chemical nature of the
wastewater contaminants (e.g., suspended solids, colloidal materials,
etc.) determines their propensity for coagulation. Hydrophobic
contaminants (e.g., clays, inert solids) have no adsorption affinity
for aqueous media, and are readily susceptible to coagulation.
Hydrophilic contaminants (e.g., emulsified oils)tend to adsorb
or absorb water which retards coagulation and flocculation, and
special coagulant aids are often required to achieve effective
coagulation. Coagulation-flocculation processes have been widely
used and proven reliable for treatment of a range of industrial
and municipal wastewaters (e.g., paperboard, laundromat, chemical,
synthetic rubber, and vegetable processing wastes). See Table E-6.
2.6 Raw Materials Requirements
• Coagulation chemicals - Commercial grade alum [Al2(S04)3 • 14H20],
ferric salts, high molecular weight polyelectrolytes, etc.
See Table E-7.
t Coagulation aids - Activated silica, inorganic salts (e.g.,
CaCl2), etc. See Table E-7.
• pH adjustment - Sulfuric acid, sodium hydroxide, lime. See
Table E-7.
2.7 Utility Requirements
t Electricity - Used for driving pumps, mechanical P^dles.
Requirements vary with the specific design and removal efficiency
desired.
E-27
-------
TABLE E-5. COMPARISON OF LIME, ALUM AND FERRIC CHLORIDE TREATMENT*^
Parameter
Chemical Cost, if/kg
U/lb)
Chemical Dosage, mg/1
Polymer, mg/1 @
$2.75/kg
Acid Dosage, mg/1
H2S04 @ 2.2<£/kg (U/lb)
Total Chemical Cost,
tf/1000 liter
U/1000 gal)
Operating pH
2
Rise Rate, pm/m
(;gpm/ft2)
Thickened Sludge, %
Thickener Loading, „
kg/day-m2 (Ib/day-ft/)
Vacuum Filtration Rate,
kg/hr/m2 (lb/hr/ft2)
Ca(OH)2
2.48 (1.25)
350
--
200
1.4 (5.3)
10.2-10.8
30.5 (0.75)
10-15
10.2-15.3 (50-75)
1.6-2.0 (8-10)
Alum1"
7.7 (3.5)
150
0.5
--
1.3 (4.9)
6.5-7.5
20.32 (0.50)
3-5
1.0-2.0 (5-10)
0.4-0.6 (2-3)§
Fea3*
9.9 (4.5)
100
0.5
—
1.2 (4.7)
6.5-7.5
20.32 (0.50)
3-5
1.0-2.0 (5-10)
0.4-0.6 (2-3)§
*Data and calculations shown are based partially on results obtained from a
(25 gpm) pilot plant using South Salt Lake Sewage conducted by Sanitary
Engineering R&D Dept. of Eimco Corp.
tFilter alum - Al2(S04)3'14H20
t40 Percent liquid solution by weight
§ Sludge conditioning chemical required
E-28
-------
TABLE E-6. EFFICIENCY OF COAGULATION-FLOCCULATION TREATMENT OF INDUSTRIAL AND MUNICIPAL WASTEWATERS
(3)
Coagulation Agents Added
Alum
Silica
H2S04
Polyelectrolyte
Cationic Surfactant
Calcium Chloride
Lime (Ib/lb BOD)
Constituent Removal (%)
BOD
SS
Oil and Grease
Fe
PO
A
COD
TSS
pH
Detention Time, hr
Sludge, % Solids
Paperboard
50 mg/1
5 mg/1
--
--
--
--
--
--
95.7-86.6
--
--
--
--
--
--
1.7
2-4
Waste Type
Ball-bearing
800 mg/1
—
450 mg/1
45 mg/1
--
--
--
--
92.6
90.7
91.1
95.9
--
--
10.3 (influent)
7.1 (effluent)
--
--
Laundromat
--
--
--
--
88 mg/1
480 mg/1
--
54.7
--
--
--
43.8
66.6
--
7.1 (influent)
7.7 (effluent)
--
--
Latex Paint
345 mg/1
—
--
--
--
--
--
91.6
--
--
--
--
95.8
82.5
3.5-4.0
(influent)
3
Synthetic
Rubber
100 mg/1
—
—
—
--
--
--
82.3
--
--
--
--
82.4
--
6.7
(influent)
--
—
Vegetable
Processing
--
—
--
--
--
--
0.5
35-70
--
--
—
--
--
—
--
--
rn
i
ro
10
-------
TABLE E-7. CHEMICAL COMPOUNDS USED IN COAGULATION-FLOCCULATION PROCESSES^
Compound
Coagulants
Aluminum sulfate
Sodium aluminate
Ammonium alum
Potash alum
Copperas
Chlorinated
cooperas
Ferric sulfate
Ferric chloride
hydrate
Magnesium oxide
Formula
A12(S04)3-14H20
Na2Al204
A12(S04)3(NH4)2S04-24H2Q
A12(S04)3-K2S
-------
TABLE E-7. Continued
Coagulant Aids
Bentonite
Sodium silicate
pH Adjusters
Lime, hydrates
Soda ash
Caustic soda
Sulfuric acid
--
Na20(Si02)3_25
Ca(OH)2
f!a2C03
NaOH
H2S04
--
40 Be
solution
90 oercent
Ca(OH)2
99 oercent
98 oercent
riaOH
100 cercent
Powder
Solution
Powder
Powder
Flake
Solid
Ground
Solution
Liquid
967 (60)
1372 (86)
393-797 (25-50)
542-829 (34-52)
--
Iron
Steel
Iron
Steel
Rubber
--
--
--
--
Essentially insoluble - fed in
slurry form
oH adjustment and softening
pH adjustment and softening
oH adjustment, softening, oil
removal systems
oH adjustment
m
i
u>
*0ther compounds, for which no information is available, suitable as coagulant aids are activated silica, clay, activated carbon
ca'istici zed starches, and ethyl cellulose.
-------
3.0 Process Advantagesv ' '
• Effective for removal of suspended solids and oils from a variety
of wastewaters, including nonbiodegradable and refractory organics.
t Minimum utility and raw materials requirements.
• Minimum maintenance requirements for most systems.
• Relatively inexpensive method for removal of oil and suspended solids.
4.0 Process Limitations^4'5^
t Process is highly sensitive to fluctuations in wastewater characteris-
tics; wastewater composition must be as uniform as possible.
t Unreacted coagulant chemicals may cause after-flocculation when the
clarified effluent is discharged into receiving waters.
• Large quantities of sludge are generated which may require dewatering
by filtration or centrifugation and disposal, either by incineration
or landfill.
• Iron and aluminum salts form gelatinous hydroxide floes that are
difficult to dewater in many cases.
• Use of iron and aluminum salts add large amounts of ions (chlorides
or sulfates) to the wastewater.
t Polyelectrolytes used alone are ineffective for removal of phosphorous.
5.0 Process Economics
Capital and operating equipment costs depend on the type and size
of the coagulation-flocculation unit used. Chemical costs depend on
the specific chemical.fs) utilized and on the quantity of wastewater to
be treated. The prices of common coagulation chemicals (1978 dollars)
are: $142/tonne ($129/ton) of alum; $27.60/tonne ($25/ton) of lime;
$2977-5513/tonne ($2700-5000/ton) of polyelectrolytes^. See Table E-6.
6.0 Input Streams
6.1 Influent Waste (Stream 1) - Wastewater characteristics will vary,
depending on the source. See Table E-6 for typical industrial/
municipal wastewater characteristics.
6.2 Chemical Flocculants, Flocculation Aids, and pH Adjusters (Stream 2)-
See Section 2.6.
E-32
-------
7.0 Intermediate Streams
7.1 Equalized Wastewater (Stream 3) - Composition will vary, depending
on that of the influent wastes (Stream 1).
8.0 Discharge Streams
8.1 Clarified Effluent (Stream 4) - See Table E-6. Effluent charac-
teristics will vary, depending on the composition of the influent
waste. Effluent will also contain unreacted, excess coagulation
chemicals.
8.2 Sludge (Stream 5) - See Table E-6. Will contain flocculated oil,
suspended solids and colloidal matter.
9.0 Data Gaps and Limitations
Effective use of coagulation-flocculation processes on oily and
colloidal wastewaters from coal gasification operations will depend on
trial-and-error experimentation to determine the most suitable processes
and conditions, since actual operating conditions for these waters are
unknown.
10.0 Related Programs - None known.
E-33
-------
REFERENCES
1, Environmental Control Issue: Control Equipment, Environmental Science
abd Technology, October 1977.
2. Infprmation provided by South African Coal Oil and Gas Corp. Ltd., to
EPA's Industrial Environmental Research Laboratory (Research Triangle
Park), November 1974.
3. W.W. Eckenfelder, Jr., Industrial Water Pollution Control, McGraw-Hill
Book Co., New York, 1966, p. 87-99.
4. Manual on Disposal of Refinery Wastes, Chapter 9, Filtration, Floccula-
tion and Flotation, American Petroleum Institute, Washington, D.C.,
First Edition, 1969, p. 9-1 to 9-20.
5. W.J. Weber, Jr., Physicochemical Processes for Water Quality Control,
Wiley-Interscience Publishers, New York, 1972, p. 63-109.
6. Organic Flocculants Market Set for Big Growth, Chemical and Engineering
News, January 23, 1978, p. 9.
7. Envirotech Municipal Equipment Division, Seminar for Consulting Engineers,
Roger Young Center, Los Angeles, Calif., August 11, 1971, 50 p.
E-34
-------
Dissolved Gases Removal Module
Steam Stripping
USS Phosam W
Chevron WWT
E-35
-------
STEAM STRIPPING PROCESS
1.0 General Information
1.1 Operating Principle - Removal of hydrogen sulfide and/or ammonia
from sour waters by stripping with steam (the most common stripping
medium*), flue gas or an inert gas. The stripping efficiency ca.n
be enhanced by adjustment of the pH of the sour water. Depending
on the system design and operating conditions, other volatile
components such as phenols and cyanides may also be partially
removed during stripping.
1.2 Development Status - Process has been used commercially for several
decades, primarily in petroleum refineries.
1.3 Licensor/Developer - Not a patented or proprietary process.
1.4 Commercial Applications - Widely employed in refineries for removal
of H2S and NH., from sour waters. The Lurgi facility at SASOL, S.A.
employs steam stripping for wastewater treatment and ammonia
recovery^ .
2.0 Process Information
2.1 Flow Diagram (see Figure E-8) - Sour water is fed to the top of the
stripping tower and flows downward countercurrent to steam over
trays or packing. Overhead vapors may be cooled to condense mois-
ture (in a reflex drum) if the off gas is to be treated for sulfur
*Based on a recent survey of refineries conducted by the American Petroleum
Institute (API)U), the vast majority of sour water strippers use steam as the
stripping medium. Use of flue gas, though suitable for ^S removal, yields
poor ammonia removal efficiencies. Only steam stripping is covered in this
data sheet. Licensed processes (Chevron WWT and PHOSAM W) which feature both
stripping and ammonia and sulfur recovery are covered by separate data sheets.
E-36
-------
LEGEND:
1. SOUR WATER FEED
2. STRIPPER VAPOR PRODUCT
3. STEAM
4. REFLUX DRUM VAPOR PRODUCT
(TO SULFUR AND/OR AMMONIA
RECOVERY, OR INCINERATION)
5. REFLUX CONDENSATE
6. STRIPPER BOTTOMS (TO WASTE
TREATMENT OR DISCHARGE)
STRIPPER
(TRAYS OR
PACKING)
c
REFLUX DRUM
Figure E-8. Refluxed Sour Water Stripper
E-37
-------
and/or ammonia recovery. Reflux is not employed when the offgas
is incinerated. Stripper bottoms are commonly cooled by heat
exchange with feed water.
2.2 Equipment^1'
• Stripping Tower - Carbon steel shell and lining. Tower may
contain packing (ceramic) or trays (carbon steel). All towers
constructed since 1970 employ trays according to an API survey.
• Reflux Drum - Linings commonly carbon steel, although stainless
sometimes used to minimize corrosion.
2.3 Feed Stream Requirements - Hydrocarbons in feed can cause fouling
of stripper unit and reflux system, and can create further down-
stream problems if stripper offgas is fed to a sulfur recovery
plant. A surge vessel with skimming facilities and depressurization
equipment is commonly employed ahead of the sour water stripper to
(2)
minimize organics in stripper feed^ '.
Carbonates in sour feed can lead to tower deposits and can
limit hydrogen sulfide removal efficiency. High carbonate waste-
waters are either neutralized prior to stripping or handled by
(4}
means other than strippingv '.
Temperature of feed affects steam consumption in the tower. A
temperature of about 393°K (250°F) is an ideal feed temperature, with
about 338°K (150°F) being a practical minimum^.
2.4 Operating Parameters
/3\
t Temperaturev ': 355°K (180°F) is approximate minimum operating
temperature for reflux drum overhead to inhibit ammonium hydro-
sulfide deposition.
t Tower Bottom Pressure^3'6': 170-400 kPa (17-40 psia)
t Loading: Depends upon efficiency desired, liquid flow rate,
steam/sour water ratio, and nature of packing or trays used
(see Section 2.6 for steam quantities used).
2.5 Process Efficiency and Reliability - With efficient steam stripping
about 99+% H2S removal, 90%-95% NH3 removal, and 50-70% phenol
removal may be realized^ '. Sour water stripping units commonly
show good reliability, although problems with foaming and fouling
E-38
-------
can occur when organics are present in the feed or temperature is
too low in the reflux drum (allowing ammonium hydrosulfide to
precipitate).
2.6 Raw Material Requirements
' narHnn ^L" °'32f^ (?'3 ' 2'7 1bS/9al) ^endi"9 on tower
packing and degree of ammonia removal requiredO). A typical
refinery steam rate is about 0.13 kg/1 (1.0 Ib/gal) feed A
design case for stripping of sour waters from a coal gasification
operation specified 0.18 I/kg (1.5 Ibs/gal )(0.9).
( 2)
• Caustic or Acidv ': Lime or sodium hydroxide may be added to
release feed ammonia from relatively acidic sour waters. Simi-
larly, flue gas, hydrochloric or sulfuric acid may be added to
alkaline sour waters to enhance hydrogen sulfide removal. Quan-
tities needed depend on the buffer capacity of the sour
water feedUO).
2.7 Utility Requirements
t Steam (see Section 2.6)
• Electricity^9^: about 0.92 kwh /10001 (3.5 kwh /1000 gal feed)
• Cooling waterv ' : about 0.6 1/1 feed
3.0 Process Advantages
t High degree of removal of hydrogen sulfide from sour waters.
• Ammonia levels in stripper water are relatively independent of feed
concentration.
• Process is widely used and reliable.
t Equipment can be constructed primarily of carbon steel.
• Low utility requirements.
• Relatively inexpensive operation.
t Vapor product can be processed for ammonia and/or sulfur recovery.
4.0 Process Limitations
t Hydrogen sulfide and ammonia removal efficiencies depend upon the
buffer capacity of the sour feed.
t Overhead vapors from stripper are corrosive and can lead to the
formation of deposits (primarily NH4HS).
f:
E-39
-------
• Organics can cause tower fouling, can affect ammonia and hydrogen
sulfide removal efficiencies, and can carry over to vapor product.
0 Stripper bottoms generally require additional control before
discharge.
5.0 Process Economics
Capital cost for a 4 x 106 I/day (1 MGD) capacity sour water stripping
operation for a coal gasification facility is estimated at about 1 million
dollars (1976r . Operating costs depend largely on steam rate and
energy source for its generation. The operating cost for the above sour
water treatment plant is estimated at about $240,000 in 1975 dollars.
6.0 Input Streams
6.1 Sour Water Feed (Stream 1) - Ammonia levels in refinery sour waters
are commonly in the range of 1000-10,000 mg/1 (see Table E-8 for the
range of feed ammonia levels encountered in an API survey^ '). Hydro-
gen sulfide levels range from 300 to 10,000 mg/l' ' '; phenol levels
(4)
range from 30 to 800 mg/lv '. Table E-8 presents data on the sources
(2)
and composition of sour waters in an example petroleum refinery .
Gas liquor feed at the SASOL Lurgi gasification plant contains 1.0 to
1.2 weight % ammonia.
6.2 Steam (Stream 3) - see Section 2.6 and Table E-9
TABLE E-8. SOURCES AND COMPOSITION OF SOUR WATERS IN AN EXAMPLE
PETROLEUM REFINERYU) - STREAM 1
Stream Source
Fluid Catalytic Cracker
Gas Plant/Sour Crude Unit
HDS* Unit Foul Water
Sulfur Plant Sour Water
Miscellaneous
Total (or average)
Typical Flow Rate
1/min
(gpm)
570 (150)
95 (23)
175 (46)
217 (57)
95 (25)
1,380 (363)
Ammonia
Concentration
(mg/1)
3,800
1,400
300
280
1,000
1,850
Sulfide
Concentration
(mg/1)
11,125
1,706
471
770
4,950
5,350
*Hydrodesulfurization
E-40
-------
TABLE E-9. PERFORMANCE DATA FOR REFINERY SOUR WATER STRIPPERS WITH HIGH AMMONIA
REMOVAL(l) - STREAMS 1, 3 AND 6
m
Tower Media
10 valve trays
8 valve trays
30 sieve trays
30 sieve trays
24 sieve trays
23 sieve trays
52 valve trays
5 glitch trays
20'-3" Raschig rings
10 flex trays
15'-3" Raschig rings
18 trays
20 bubble cap trays
1 2 Socony trays
20 sieve trays
Average Steam Rate
kg/1 (Ibs/gal)
* 1.4 (3.1)
1.0 (2.2)
1.1 (2.4)
2.5 (5.5)
0.6 (1.3)
1.6 (3.5)
--
7.8 (17.2)
1.3 (2.9)
1.8 (4.0)
1.5 (3.3)
--
--
1.2 (2.6)
--
Refluxed Strippers
Average Ammonia
in Feed (mg/1)
2500
1200
1720
430
74
4000
1600
5.410
3550
2000
1400
19,000
2000
32,200
1600
Average Ammonia
in Bottoms (mg/1 )
78
25
68
64
63
100
65
45
--
200
80
80
15
56
25
Minimum Ammonia
in Bottoms (mg/1)
25
--
--
--
--
40
--
19
37
25
7
--
10
--
7
Non-Refluxed Strippers
8 bubble cap trays
6 shower trays
15'-3" Raschig rings
15'-3" saddles
8 valve trays
28 bubble cap trays
5 valve trays
8 flex trays
1.5 (3.3)
0.3 (0.7)
0.6 (1.3)
0.2 (0.4)
1.9 (4.2)
0.8 (1.8)
0.4 (0.9)
2.7 (5.9)
960
1850
1200
2600
5450
2625
215
4400
50
96
65
200
56
10
76
11
30
--
36
34
--
--
—
10
-------
7.0 Intermediate Streams
7.1 Stripper Vapor Product (Stream 2) - no data available (this stream
may be a discharge stream in the case of a non-refluxed stripper).
7.2 Reflux Condensate (Stream 5) - see Table E-10 (applies only to
refluxed strippers).
8.0 Discharge Streams
8.1 Stripper Bottoms (Stream 6) - see Table E-ll and Table E-9.
8.2 Reflux Drum Vapor Product (Stream 4) - Limited data available;
H2S, NH3, and C02 will be present in offgas in approximately the
same ratio as found in feed water (assuming near equal removal
efficiencies). The offgas may contain other volatile organics
(e.g., phenols, light hydrocarbons) and inorganics (e.g., HCN).
At the SASOL Lurgi gasification facility, ammonia stripper column
overhead contains 6% ammonia and 0.1% FLS.
TABLE E-10. COMPOSITION OF SOUR WATER STRIPPER REFLUX CONDENSATE AT
A LARGE REFINERY*0) - STREAM 5
Steam Rate1" - kg/1
(Ibs/gal)
1.20
1.41
1.58
1.76
1.90
2.23
2.87
Reflux Ammonia
(mg/1)
32,600
36,300
42,600
26,300
30,300
10,000
52,300
Reflux H9S
(mg/ir
16,300
30,600
28,600
17,700
8,600
9,200
12,500
*10 actual stages (or trays) in tower.
fSteam rate is only one variable influencing reflux ammonia and hydrogen
sulfide levels; data reflect varying operating conditions.
E-42
-------
TABU
Example
Refinery No.
NH3 (as N)
Sulfide (as H2S)
Phenols
BOD
\j\JiJ
COD
TOC
TSS
Total alkalinity**
Chloroform
extractables
Ca++
Mg++
Si02
cr
so4=
Total P04=
N02"
MO ""
NUo
Cu++
Fe+2 & +3
7 ++
Zn
Specific
conductance
PH
—
1
28
0.01
110
""
--
--
28
88
13
5
0.5
<2
39
16
3
0
.-, i
-------
9.0 Data Gaps and Limitations
Data gaps and limitations relate primarily to the performance of sour
water strippers with regard to minor constituents such as oils, phenolics,
amines, and cyanides. No data for actual applications to coal gasifica-
tion are known.
10.0 Related Programs
No programs are known to be underway or planned which are aimed at
the assessment of sour water stripper performance in coal gasification
applications.
REFERENCES
1. Gantz, R.G., Sour Water Stripper Operations, Hydrocarbon Processing,
May 1975, p. 85-88.
2. Rodriguez, D.G., Sour Water Stripper: Its Design and Application, in
Water-1973 AIChE Symposium Series, No. 136, Vol 70, 1974.
3. Melin, G.A., et al, Optimum Design of Sour Water Stripper, Chemical
Engineering Progress, Vol 71, No. 6, June 1975, p. 78.
4. Walker, G.J., Design Sour Water Strippers Quickly, Hydrocarbon Processing,
June 1969, p. 121-124.
5. Water Conservation and Pollution Control in Coal Conversion Processes,
Water Purification Associates, Draft Report to EPA under Contract No.
68-03-2207, 1977.
6. Hart, J.A., Waste Water Recycled for use in Refinery Cooling Towers,
Oil and Gas Journal, June 11, 1973, p. 92-96.
7. Maguire, W.F., Reuse Sour Water Stripper Bottoms, Hydrocarbon Processing,
September 1975, p. 131-152.
8. Beychok, M.R., Aqueous Wastes from Petroleum and Petrochemical Plants,
John Wiley and Sons, New York, 1967.
9. Bonham, J.W. and Atkins, W.T., Process Comparison Effluent Treatment
Ammonia Separation, ERDA Document No. FE-2240-19, June 1975.
10. Bombergen, D.C. and Smith, J.H. Use Caustic to Remove Fixed Ammonia,
Hydrocarbon Processing, Vol. 56, No. 7, July 1977, p. 157-162.
11. Information provided by South African Coal, Oil and Gas Corp., Ltd. to
EPA's Industrial Environmental Research Laboratory (Research Triangle
Park), November 1974.
E-44
-------
USS PHOSAM W PROCESS*
1.0 General Information
1.1 Operating Principles - Ammonia is absorbed from a gas stream
(usually overhead vapors from a sour water stripper) by counter-
current flow of an ammonium phosphate solution. The ammonia-
enriched solution is subsequently steam stripped at elevated
pressure to release the absorbed ammonia. The resulting water/
ammonia vapor stream is fractionated to produce anhydrous ammonia.
The absorption/regeneration reaction may be represented as follows:
(NH4)1<3 H1>7P04 + 1/2 NH3 ^ (NH4)K8H1-2P04
1.2 Development Status - commercial; several PHOSAM units are currently
in operation on coke oven gases.
1.3 Licensor - USS Engineers and Consultants, Inc. (UEC)
600 Grant Street
Pittsburg, PA 15230
1.4 Commercial Applications^ - The PHOSAM process is used at the
U.S. Steel's Clairton coke facility (Pittsburg, PA) and nine other
PHOSAM plants are operating worldwide; others are in the design
and construction phase. PHOSAM W is licensed to at least one
proposed coal gasification plant.
*The PHOSAM Process is for application to byproduct coke production. The
PHOSAM W refers to the application of the process to other wastewaters and
gases.
E-45
-------
2.0 Process Information
2.1 Flow Diagram (see Figure E-9 for one design of the PHOSAM W process
for treating sour water)* - Sour water enters a steam stripper where
free NHg, H2$, C02 and other acid gases and volatile organics are
stripped. The top of the stripper is an ammonia absorber where
lean ammonium phosphate solution contacts the sour vapors and
absorbs ammonia and small amount of acid gases. Stripped sour water
leaves the bottom of the stripper. Rich ammonium phosphate solution
is purged of acid gases in a contactor and sent to a high pressure
steam stripper for ammonia removal. Lean solution returns to the
absorber, while the ammonia is separated from water by distillation.
Caustic is added to the fractionation system to inhibit acid gas
accumulation in the still. Still bottoms are returned to the sour
water stripper. Reboilers may be used on the columns, if condensate
recovery is required.
2.2 Equipment^3' - Based on a design for 13.7 x 106 I/day (3.6 MGD) sour
water treatment.
• "Superstill" consisting of a sour water stripper and absorber
Height Diameter
Bottom 25 m (84 ft) 4.4 m (14.5 ft)
Top 18 m (61 ft) 2.9 m (9.5 ft)
• PHOSAM Stripper
Height - 21 m (70 ft)
Diameter - 2.9 m (9.5 ft)
• Fractionator
Height - 20 m (65 ft)
Diameter - 1.6 m (5.25 ft)
*Process can also be designed to handle ammonia-containing gas streams, in
which case the sour water stripper is omitted.
E-46
-------
C.W.
LEGEND:
1. SOUR WATER FEED
2. LOW PRESSURE STEAM
3. HIGH PRESSURE STEAM
4. HIGH PRESSURE STEAM
7. RICH AMMONIUM PHOSPHATE SOLUTION
8. LEAN AMMONIUM PHOSPHATE SOLUTION
9. STILL BOTTOMS
10. PRODUCT ANHYDROUS AMMONIA
5. CAUSTIC FEED (LOCATION NOT KNOWN! 1t. PURIFIED GAS
6. MAKEUP PHOSPHORIC ACID (LOCATION 12. STRIPPED SOUR WATER
NOT KNOWN) 13. PHOSPHATE SOLVENT SLOWDOWN
(LOCATION NOW KNOWN)
Figure E-9. USS PHOSAM W Process
(1)
-------
• Heat Exchangers - approximately 5640 m2 (60,000 ft ) of surface
area
PHOSAM W absorber, strippers and fractionator are stainless steel
clad. Sour water stripper is carbon steel.
2.3 Feed Stream Requirements - Tars and pitches can cause fouling
problems in sour water stripper reboiler or bottoms interchanges
if used.
Temperature of feed affects steam consumption in the sour water
(2)
stripper and heat exchanger surface requirements . Feed at its
bubble point is ideal, but any temperature is acceptable.
2.4 Operating Parameters
Temperature^: Absorber - 378°K (220°F)
PHQSAM stripper - 465°K-475°K (380°F-400°F)
Pressure: Absorber3^ - 0.07-0.16 MPa (0-10 psig)
PHQSAM stripper^ - 1.3-1.7 MPa (180-
250 psig)
Solution Circu- Depends on feed ammonia concentration and
lation Rate: pressure of absorber and stripper.
2.5 Process Efficiency and Reliability^ ' - On coke oven gases, an
ammonia removal of up to 99.7% can be obtained at absorption tem-
peratures of 3H°K-333°K (105°F-140°F). Recovered anhydrous ammonia
is 99.99% pure. The PHOSAM process has reportedly been in successful
operation at U.S. Steel's largest coke plant (in Clairton Works,
Pittsburgh, PA) since 1968.
2.6 Raw Material Requirements^ '
H3P04 makeup (as 100% H3P04) - 0.002 kg/kg NH3
NaOH for ammonia fractionator (as 100% NaOH) - 0.003 kg/kg NH_
O
2.7 Utility Requirements^
Steam @ 3.7 MPa (550 psig): 12 kg/kg NH-
0
Steam @ 0.27 MPa (25 psig): 8 kg/kg NH,
-------
Cooling water: 300 I/kg (40 gal/lb) NH,
O
Electric power: 0.066 kwh /kg NH3
3.0 Process Advantages
• Commercially available.
• Process recovers anhydrous ammonia.
• Process is efficient at separating ammonia from H,S and other acid
gases in gaseous or wastewater streams. i
• Process uses a relatively inexpensive and non-hazardous/non-toxic
solvent.
• Process can reduce free ammonia levels in wastewaters to 100-200 mg/1.
4.0 Process Limitations
• Wastewater from the process contains levels of ammonia ("100 ppm) and
phosphate (~7 mg/1) which may require further treatment.
« Moderately high steam and cooling water requirements.
5.0 Process Economics
c a
A PHOSAM W plant handling 13.7 x 10 I/day (3.6 x 10 gal/day) of
sour water is estimated to have a capital cost of 8.2 million 1976
dollars^. Operating costs for such a plant are estimated at about
$1/1000 1 ($4/1000 gals). Sale of ammonia can offset about $0.14/1000 1
($0.55/1000 gals) for each 1000 mg/1 of ammonia in the feed.
6.0 Input Streams
6.1 Sour Water Feed (Stream 1) - No operating data available. The
following sour water feed composition and flow rate were submitted
to USS Engineers and Consultants by C.F. Braun and Co. requesting
design and cost data for coal gasification applications :
Component
Carbonate C02 I3'000
Sulfide (as H2S) 350
HCN 33°
NH3 4'800
E-49
-------
Component mg/1
Phenol 3,500
Flow Rate - 1/min 12,540 (3,300)
(gal/min)
Temperature 360°K (200°F)
6.2 Low Pressure Steam (Stream 2) - see Section 2.7
6.3 High Pressure Steam (Streams 3 and 4) - see Section 2.7
6.4 Caustic Feed (Stream 5) - see Section 2.6
6.5 Make-up Phosphoric Acid - see Section 2.6
(2}
7.0 Intermediate Streams^ '
7.1 Rich Ammonium Phosphate Solution (Stream 7) - Concentration about
40% by weight; salts in solution approximate the formula
(NHj, gH, 2P04> The pH of the rich solution is approximately 6.8.
7.2 Lean Ammonium Phosphate Solution (Stream 8) - Concentration about
40%; salts in solution approximate the formula (NH^)^ ^H, yP^V The
pH of the lean solution is approximately 5.2.
7.3 Fractionator Bottoms (Stream 9) - Ammonia level is approximately
500 mg/r ' . No other composition data available. Bottoms will be
alkaline since caustic is added to enhance ammonia removal and pre-
vent acid gas contamination of product ammonia. This stream is nor-
mally recycled to the sour water stripper.
8.0 Discharge S tr earns
8.1 Product Ammonia (Stream 10) - Anhydrous and approximately 99.99%
pure^ . No trace composition data available.
8.2 Purified Gas (Stream 11) - Equilibrium data are proprietary, but
UEC reports that normal design allow 0.5% to 0.8% of the free
ammonia in the feed water to remain in the purified gas. Acid
gases (e.g., C02, H2S, HCN) are almost completely stripped from
sour feed water and will be present in purified gas in the same
molar ratio as they appeared in the feed to the sour water stripper.
E-50
-------
8.3 Stripped Sour Water (Stream 12) - Free ammonia content is
approximately 100-200 mg/l(1). Acid gases are essentially completely
removed by stripping. Phenols and other organics are only partially
removed. Fixed ammonia salts remain in the stripped water unless
neutralized by alkali addition.
9.0 Data Gaps and Limitations
Due to the proprietary nature of the process, limited data are
available on the properties and flow rates of most streams associated
with the PHOSAM W process.
10.0 Related Programs
C.F. Braun, as the evaluation contractor for the ERDA-AGA program
on high Btu gas from coal, has obtained designs and data for the applica-
tion of PHOSAM W to sour waters likely to be encountered in coal gasifi-
(5)
cationv '. Detailed information about this design are not currently
publicly available, due to the proprietary nature of the process.
REFERENCES
1. Dravo Corp., Handbook of Gasifiers and Gas Treatment Systems, ERDA
document No. FE-1772-11, February 1976.
2. Information provided to TRW by R.D. Rice of USS Engineers and Consultants,
December 27, 1977.
3. Water Purification Associates, Water Conservation and Pollution Control
in Coal Conversion Processes, EPA Report No. 600/7-77-065, 1977.
4. Colaianni, L.J., Coke Oven Offgas Yields Fuel, Chemical Byproducts, Chemical
Engineering, March 29, 1976, p. 82.
5. Bonham, J.W. and Atkins, W.T., Process Comparison Effluent Treatment
Ammonia Separation, ERDA Document No. FE-2240-19, June 1975.
E-51
-------
CHEVRON WWT PROCESS
1.0 General Information
1.1 Operating Principle - Stripping of hydrogen sulfide and ammonia from
sour waters with steam in two separate stages to produce gaseous
streams suitable for sulfur and ammonia recovery.
1.2 Development Status - Commercially available (first commercial unit
was constructed in 1966). Several units are now in operation in
(4)
refineries in California, Texas, Canada, Japan and Kuwait' .
1.3 Licensor/Developer - Chevron Research Company
575 Market Street
San Francisco, Calif.
1.4 Commercial Applications - To date, all commercial applications have
been for the processing of refinery sour waters.
2.0 Process Information
2.0 Flow Diagram (see Figure E-10) - Degassed sour water is fed to a
reboiler stripper column where hydrogen sulfide and carbon dioxide
are stripped overhead. Stripper bottoms with the bulk of the
ammonia are fed to a second reboiler stripper column (operated under
different temperature and pressure) for ammonia stripping. The over-
head from the second stripper is scrubbed with cold aqueous ammonia
to remove traces of H2S, and compressed and condensed to form
anhydrous or aqueous ammonia; hydrogen sulfide rich aqueous ammonia
is recycled to the degasser.
2.2 Equipment - The process employs pressure vessels, distillation
columns, scrubbing towers, and compression equipment. Materials
used in these equipment are not known.
2.3 Feed Stream Requirements - Process incorporates a degasser for the
removal of highly volatile organics. Volatile inorganics (C02 and
HCN) will appear in the hydrogen sulfide stripper overhead. Phenols
E-52
-------
c.w.
CAN.
CO
LEGEND:
1. SOUR WATER FEED
2. STEAM - 1.1 MPa (165 PSIA)
3. STEAM - 0.44 MPa (65 PSIA)
4. DEGASSED FEED
5. HYDROGEN SULFIDE RICH GAS
(TO SULFUR RECOVERY OR
INCINERATION)
6. FLASH GAS
AMMONIA
CONDENSER
7. AMMONIA RICH WATER
8. REFLUX CONDENSATE (TO
STRIPPER OR TO DEGASSERI
9. AMMONIA RICH GAS
10. H2S RICH AMMONIA SOLUTION
11. PRODUCT AMMONIA
17 WASH WATER MAKE-UP
13. STRIPPER BOTTOMS
COMPRESSOR
Figure E-10. Chevron WWT Process Flow Diagram
-------
will appear largely in ammonia stripper bottoms. The process is more
economical when applied to feeds containing high levels of hydrogen
sulfide and ammonia (3 to 5 wt % each'1'). However, Chevon Research
Company has patented a presentation process to handle feeds with low
(2)
H?S and NH., concentrationsv .
2.4 Operating Parameters^ ' - Basis is petroleum refinery design.
Hydrogen Sulfide Stripper
Still bottoms: Temperature: ?
Pressure: ?
Overhead: Temperature: 311°K (110°F)
Pressure: 780 kPa (115 psia)
Ammonia Stripper
Still bottoms: Temperature: 367°K (200°F)
Pressure: 440 kPa (65 psia)
Overhead: Temperature: ?
Pressure: ?
fl 2)
2.5 Process Efficiency and Reliability^ ' ' - Process is capable of pro-
ducing a hydrogen sulfide stream with less than 50 ppm (wt) NH~; an
ammonia stream with less than 50 ppm (wt) H?S; and a stripped water
stream containing less than 50 mg/1 ammonia and 5 mg/1 sulfide. Reli-
ability of the process is reportedly high. (A Chevron WWT plant in
El Segundo, Calif, has operated for several years without a major
shutdown.)
2.6 Raw Material Requirements - No raw materials are required for the
process.
2.7 Utility Requirements^ ' - Based on a design feed for coal gasifica-
tion application (see Section 6.0)
Total Steam: 1.0 MPa (150 psig) with returnable condensate -
0.16 kg/1 (1.31 Ibs/gal) feed
Electric power: 0.04 kwh /I (0.01 kwh /gal)*
Cooling water: 0.48 1/1 feed
*Approximately 1/3 of electric power is for ammonia product compression.
E-54
-------
3.0 Process Advantages
• Commercially available and has been demonstrated to be reliable.
• Can produce either anhydrous or aqueous ammonia.
• Can produce a concentrated hydrogen sulfide stream suitable for sulfur
* "\f \j v 1.1 y •
• Achieves low ammonia and hydrogen sulfide levels (50 and 5 mg/1, respec-
tively) in stripped wastewaterU). H
• Relatively low cooling water requirements^.
4.0 Process Limitations
• Process does not remove phenols (or other low .volatile organics) from
wastewaters.
• Process consumes relatively large amounts of electricity^3'.
• Process economics are highly dependent upon plant size, feed ammonia
and hydrogen sulfide levelsM).
5.0 Process Economics
A study comparing the economics of the Chevron WWT process with the con-
ventional sour water stripping for application to coal conversion waste-
waters has indicated that the capital cost for the Chevron process would
be 3.2 million dollars (1975 dollars) higher than that for a conventional
steam stripper handling 18 x 10 I/day (4.75 mgd) of wastewater^ '. This
same study indicated annual utility costs of 1.35 million dollars (1975
dollars) for the Chevron plant compared to 1.4 million dollars for the
conventional stripper. The ammonia recovered in the above Chevron plant,
however, has an estimated annual sales, value of 4.65 million dollars (1975
dollars) which significantly offsets the higher capital and utilities
(3)
costs of the Chevron plantv '.
Chevron Research Company has reviewed and updated the above cost study and
has estimated that the installed cost of the 18 x 106 I/day (4.75 mgd)
plant would be 11 million 1978 dollars(4). This plant includes a precon-
centration process to produce a suitable feed to Chevron WWT.
E-55
-------
6.0 Input Streams
6.1 Sour Water Feed (Stream 1) - In refinery applications feed water
ammonia concentrations have ranged from 12,000 to 55,000 mg/1,
sulfide concentrations from 25,000 to 55,000 mg/1 as H^ '.
Levels expected in coal gasification wastewaters are generally
much lower. The following feed compositions have been assumed
in a design of a Chevron process for application to coal gasifica-
(3 4)
tion sour waterv ' '.
Constituent Concentration (mg/1)
Carbonate Carbon (as C02) 13,000
Sulfide (as H2S) 230
Cyanide 330
Ammonia 4,800
Phenol 3,500
6.2 Steam (Streams 2 and 3) - see Section 2.7.
6.3 Make-up Wash Water (Stream 12) - No data available.
7.0 Intermediate Streams
7.1 Degassed Feed (Stream 4) - No data available.
7.2 Ammonia Rich Water (Stream 7) - No data available.
7.3 Reflux condensate (Stream 8) - No data available.
7-4 Ammonia Rich Gas (Stream 9) - Contains about 2% (wt) hydrogen
sulfide in refinery applications^ '. No other composition data
available.
8.0 Discharge Streams
8.1 Hydrogen Sulfide Rich Gas (Stream 5) - Product specifications for
this stream are less than 50 ppm (wt) ammonia and less than 5000 pprn
(wt) water vapor '. Depending on the composition of the sour
water feed, the stream may contain C02, HCN, organics, etc. No
actual composition data available on this stream.
E-56
-------
8.2 Flash Gas (Stream 6) - No data available.
8.3 Stripper Bottoms (Stream 13) - Product specifications for this stream
are less than 50 mg/1 ammonia and less than 5 mg/1 sulfide. No
actual composition data available.
A design case for coal gasification wastewater treatment (see
Section 6.1) has specified the following composition^:
Component Concentration (mg/1)
Carbonate Carbon (as COJ 0
Sulfide (as H2S) 0
Cyanide 17
Ammonia 11
Phenol 2900
8.4 Product Ammonia (Stream 11) - Product specifications for this
stream are less than 5 ppm (wt) hydrogen sulfide and less than
1000 ppm water. No other data available.
9.0 Data Gaps and Limitations
Data gaps and limitations for the process relate primarily to the composi-
tion of various process and waste streams. Existing applications of the
Chevron process have been to refinery sour waters which contain very
high levels of ammonia and hydrogen sulfide. Information about process
performance in applications to feeds containing lower levels of ammonia
and hydrogen sulfide (e.g., sour waters expected in coal gasification)
is not publicly available, although Chevron Research has patented a pre-
concentration process to handle dilute feeds such as those encountered
4)
in coal gasification applications" '.
10.0 Related Programs
C. F. Braun, as evaluation contractor for the ERDA-AGA program on high
Btu gas from coal, has obtained designs and data for the application of
Chevron WWT to sour waters likely to be encountered in coal gasifica-
tion^. Detailed information about this design is not currently
publicly available.
E-57
-------
No other programs aimed at the evaluation of the applicability of the
Chevron process to coal gasification or at environmental assessment of
the process are known to be under way or planned.
REFERENCES
1. Annessen, R. J., and Gould, G. D., Sour Water Processing Turns Problem
into Payout, Chemical Engineering, March 22, 1971, p. 67-69.
2. Klett, R. J., Treat Sour Water at a Profit, Hydrocarbon Processing,
October 1972, p. 97-99.
3. Bonham, J. W., and Atkins, W. T., Process Comparison Effluent Treatment
Ammonia Separation, ERDA Document No. FE-2240-19, June 1975.
4. Information provided to TRW by J. D. Knapp of Chevon Research Company,
February 3, 1978.
E-58
-------
Dissolved/Particulate Organics Removal Module
Biological Oxidation
Evaporation/Retention Pond
Chemical Oxidation
Phenosolvan
Activated Carbon Adsorption
E-59
-------
BIOLOGICAL OXIDATION PROCESS
1.0 General Information
1.1 Operating Principle - Use of microorganisms to convert organic
compounds to carbon dioxide, water and other end products. Air or
oxygen is provided for the biological oxidation of organics.*
1.2 Development Status - Commercially available. Numerous units are in
operation throughout the world for municipal waste treatment and for
treatment of industrial wastes, including coal gasification and
petroleum refinery wastes.
1.3 Licensor/Developer - Many biological treatment systems and equipment
are offered by numerous suppliers. Some licensed versions of biologi-
cal processes are patented, such as the UNOX pure oxygen activated
sludge technology (Union Carbide Corporation, So. Charleston, West
Virginia)^ . A complete listing of these systems is available in
the literature (e.g., pollution control editions of ES&T, Pollution
Engineering, Chemical Engineering, etc.).
1.4 Commercial Applications - Coal related applications include:
(a) SASOL Lurgi-type coal conversion facility, Sasolburg, So. Africa -
(2}
trickling filters^ '; (b) HYGAS pilot plant, Chicago, Illinois - waste
(3)
stabilization pondv '; and (c) Bethlehem Steel Co., coke plant,
Bethlehem, Pa. - commercial scale air activated sludge systenr .
2.0 Process Information
2.1 Flow Diagram (see Figure E-ll) - The most widely used biological
treatment systems are: (a) activated sludge (air activated and high
*When air or oxygen are used, the biological oxidation is classified as aerobic
oxidation. In the absence of air or oxygen (anaerobic conditions), the decom-
position of organics is incomplete and results in the production of inter-
mediate organic compounds, methane, sulfide, etc. Except for certain special
applications (e.g., treatment of organic sludge or concentrated waste), aerobic
treatment is the system of choice and is discussed in this data sheet.
E-60
-------
I
BIOLOGICAL TREATMENT
SYSTEM
LEGEND:
1. Influent Waste
2. Nutrients/pH Adjustment Chemicals
3. Air or Oxygen
4. Fugitive Emissions
5. Effluent to Clarifier for Solid
Separation
6. Sludge to Recycle
7. Excess Sludge to Treatment/
Disposal
Figure E-ll. Simplified Schematic of Biological
Oxidation Svstenn5/
E-61
-------
purity oxygen activated); (b) trickling filters; (c) lagoons (waste
stabilization ponds); and (d) oxidation towers.
• Activated sludge: The conventional activated sludge process con-
sists of a biological reactor unit containing a high concentration
of microorganisms. Air or oxygen (Stream 3) is supplied either
by mechanical aeration or by a diffused air system. The treated
waste is sent to a clarifier for solids/liquids separation. A
portion of the settled sludge (Stream 6) is recycled to the bio-
logical reactor to "seed" the raw wastewater; the excess sludge
(Stream 7) is sent to disposal.
t Trickling Filter: This system consists of a filter bed and
wastewater distribution (i.e., sprinkler system) and a sedi-
mentation tank. The filter bed, which is typically 9.3 - 124 m
(3-40 ft) deep, consists of rock or synthetic media to which
microbial films are attached. Most systems employ recirculation
to increase efficiency and minimize shock loadings.
• Lagoons (waste stabilization ponds): These systems consist of
large basins ranging from 3.1 to 37.2 m (10-122 ft) in depth
and are classified as aerobic, anaerobic or facultative. In
aerobic lagoons, air or oxygen is provided through natural sur-
face aeration or by mechanical means (aerated lagoons). In
facultative lagoons, both aerobic and anaerobic waste digestion
occur (anaerobic conditions exist near the bottom of the pond).
• Oxidation towers: Wastewater is used as make-up water for cool-
ing towers. Biological floes become established in the system
(mostly on the surfaces in the cooling tower) and excess floes
are discharged in the cooling tower blowdown.
2.2 Equipment
• Biological reactor unit (i.e., tank, lagoon, filter bed, tower,
etc.)
• Sedimentation tanks for solids/liquids separation (activated
sludge)
• Mechanical aerators (aerated lagoons, activated sludge)
• Compressors and air diffusers (activated sludge)
• Pumps (all systems)
2.3 Feed Stream Requirements
• Temperature: Optimum for aerobic biological treatment systems -
20°C to 35°C (60°F to 80°F).
t Biodegradability: The organics to be removed must be biodegradable.
E-62
-------
.
Loading: Varies with the specific biological system, removal
efficiencies desired, and specific design. A raw waste may
require either dilution or concentration for effective treatment.
Inhibitory constituents: Wastes containing high concentrations
of chemicals toxic to biological systems (e.g., heavy metals,
tars, phenols, ammonia, etc.) cannot be treated effectively
Threshold concentrations reported for phenols, ammonia and
and°2o2o «-^°° t0 ^'°°° mg/l(5'6)' 120° to 200° -ng/l(7,8);
• Equalization: Wide fluctuations in wastewater characteristics
may be detrimental to biological systems. Equalization may be
necessary to achieve uniform waste concentrations.
• pH: Optimum pH is between 6 and 8.
• Nutrients: A BOD:N:P ratio of approximately 100:5:1 is required
for biological treatment.
2.4 Operating Parameters - Considerable data are available on biological
treatment for petroleum refinery and coal conversion wastes. Typi-
cal operating parameters for an activated sludge system for testing
coke plant wastes are given in Table E-12. Selected operating param-
eters and data for a hypothetical design activated sludge system for
a coal conversion plant are presented in Table E-13.
2.5 Process Efficiency and Reliability - Efficiency depends upon the type
and design of process used, and on the nature of the wastewaters.
The chemical nature of the compounds determines their biodegradability.
Di- and polyhydric phenols (found in coal gasification effluents)
are less completely biodegraded than simple phenols. The biodegrada-
bi lities of polyaromatic phenols, and most aromatic and hetero-
cyclic compounds are unknown. Data on biological treatment
efficiencies for coal gasification facilities are presented in
Tables E-14 and E-18. A COD removal efficiency of 91.7% has been
reported for the trickling filters in use for wastewater treatment
at the SASOL Lurgi-type coal conversion facility, Sasolburg, South
Africa(2). Tables E-12 and E-15 present data on coke plant and
petroleum refinery applications, respectively. Biological
processes have been widely used and proven highly reliable treat-
ment of range of industrial and municipal wastewaters.
E-63
-------
TABLE E-12. OPERATING PARAMETERS AND DATA FROM AN ACTIVATED
SLUDGE PILOT PLANT FOR TREATING A PHENOLIC
WASTEWATER FROM A COKE PLANT*(7)
Parameter
Parameter Ranget
Flow rate, 1/min (gpm)
Phenol, mg/1
Dilution water,^ 1/min
(rpm)
Recycle sludge,
1/min (gpm)
Retention time, hr.
Temperature, °C (°F)
pH
Phenol removed, kg/day/kg
sludge in aeration
Sludge growth, kg/kg
phenol removed
Effluent phenol
concentration, mg/1
1.5 - 4.6 (0.40 - 1.21)
3,350 - 3,900
2.0 - 4.5 (7.6 - 17.0)
9.1 - 37.8 (2.4 - 10.0)
1.6 - 4.4
27 - 33 (80
6.9 - 7.7
0.43 - 0.93
0.13 - 0.23
0.2 - 0.8
- 92)
*Range for 7 individual measurements.
Dilution water added to raw waste to lower strength prior
to treatment
E-64
-------
TABLE E-13.
DESIGN AND EXPECTED PERFORMANCE DATA FOR A
HYPOTHETICAL HIGH PURITY OXYGEN ACTIVATED
SLUDGE (HPOAS) SYSTEM FOR A COAL CONVERSION
PLANT (4>
Influent Waste Characteristics
BOD5, mg/1
COD, mg/1
Phenol as CeH5OH in
mg/1
NHg.as N, mg/1
Flow, I/day (gal/day)
Temperature, °C (°F)
Design Parameter or Utility
Reguirementt
Volume of unit, 1 (gal)
Area of Clarifier, m^
(ft2)
Retention Time, hrs.
(based on feed flow)
Sludge recycle rate, %Q,
I/day (gal/day)
Mean biomass loading,
kg BOD5/kg MLVSS-day
Volumetric organic loading,
kg BOD5/103 m3-day
Recycle suspended solids,
wt %
13,000-18,000
25,000-30,000
3,000-5,000
290
3.22 x 106 (0.85 x 106)
26.7 (80)
Step 1
23.0 x 106
(2.62 x 106)
198.1
(2,130)
74
132.3 (35)
0.8
5.84
2.0
Step 2
2.2 x 106
0.57 x 106)
263.7
(2,835)
16
132.3 (35)
0.3
1.35
2.0
(continued)
*Based on Hygas plant using lignite feed. The design assumes that the
biodegradability of the coal conversion wastes are similar to coke
plant wastes.
fTwo units ("steps") in series, each consisting of a HPOAS unit and
a clarifier.
E-65
-------
TABLE E-13. Continued
Design Parameter or Utility
Requirementt (Continued)
Effluent solution BOD.
mg/1 b
Oxygen supplied,
kg/day (tons/day)
Average oxygen utiliza-
tion efficiency, %
Electric power for
aerators, kw-hr
Step 1
900
7.2 x 10
(79.0)
79
613
Step 2
45
4.2 x Iff
(4.6)
80
38
*Based on Hygas plant using lignite feed. The design assumes that
the biodegradability of the coal conversion wastes are similar to
coke plant wastes.
tTwo units ("steps") in series, each consisting of a HPOAS unit and
a clarifier.
E-66
-------
TABLE E-14. ANTICIPATED WASTEWATER COMPOSITIONS AND BIOLOGICAL TREATMENT EFFICIENCIES FOR
COAL CONVERSION EFFLUENTS
Wastewater
Compounds
Phenols
Aromatic Amines
Monoaromatic
hydrocarbons
Thiophenes
Polycyclic
hydrocarbons
Thiocyanate
Cyanide
Sulfide
BOD5
Suspended
Solids
Anticipated Untreated
Effluent (Stream 1)
Concentration Range, mg/1
1,000 - 10,000
100 - 1,000
10 - 100
1 - 10
0.1 - 1
—
—
—
30,000
...
Anticipated Treated
Effluent (Stream 5)
Concentration Range, mg/1
1 - 10
70 - 500
9 - 90
*
0.03 - 0.08
1 - 10
1 - 10
0.01 - 0.3
50-150
60 - 200
Biological Wastewater
Treatment Removal
Efficiency, %
99.9+
30 - 50
40+
—
30 - 80
—
—
—
99.8+
...
Ref.
11
11
n
n
n
12
12
12
12
12
m
cr>
*No data available
-------
TABLE E-15. EFFICIENCY OF BIOLOGICAL TREATMENT FOR PETROLEUM
REFINERY EFFLUENTSU3)
Biological
Treatment
Method
Activated Sludge
Trickling Filters
Waste
Stabilization
Pond (Aerobic)
Aerated Lagoons
Cooling Tower
Oxidation (Air
Stripping)
Percent Removal*
S.
S~ BOD COD Solids Oil Phenols
97-100
—
95-100
88-90
60-85
40-95
75-95
90+
60-85
30-70
30-65
60-85
90+
—
50-80
20-70
40-65
—
50-80
50-90
70-90
95-99+
—
— — _
90-99
99.9
*Thiocyanates are approximately 70% removed by all processes.
2.6 Raw Material Requirements
• Air or oxygen (Stream 3): Varies with the type of biological
treatment, waste loading and removal efficiency; for most sys-
tems, 0.6 to 1.5 kg 02/kg BOD5 removed.
• Nutrients/pH adjustment chemicals (Stream 2): See Section 2.3
• Microorganisms: Some strains of bacteria may be added to improve
removal efficiency (e.g., PHENOBAC - a commercially available
strain of mutated Pseudomonas sp. for removal of phenols).
2.7 Utility Requirements
• Electricity: Used for driving pumps, compressors, etc., and
varies with the specific design and removal efficiency desired.
Power requirements for activated sludge and aerated lagoon sys-
tems are generally between 0.020 to 0.022 hp-hr/lb BOD (0.006 to
0.0074 kw-hr/kg BOD) removed per day.
E-68
-------
3.0 Process Advantage^4^
* ^U?^ TXPfSiV? metnods f°r the removal of biodegradable organics
and low levels of certain reduced inorganics (e.g., cr, SOT, S-, etc.).
• Minimum maintenance requirements for some of the biological systems
(e.g., lagoons, trickling filters).
• Little or no raw materials required except for oxygen and air in the
case of activated sludge and aerated lagoon systems, and possibly
nutrients, and chemicals for pH adjustment.
4.0 Process Limitations^4'15^
• Ineffective for removal of nonbiodegradable and refractory organics
(some of which are present in coal gasification effluents - see
Section 2.5).
• Inapplicable when waste contains intolerably high concentrations of
toxic materials (e.g., heavy metals, toxic organics, etc.).
• Process is highly sensitive to wide fluctuations in wastewater char-
acteristics (e.g., pH, acidity, and organic and hydraulic loadings).
• Some processes (e.g., conventional activated sludge systems trickling
filters), generate sludge requiring further treatment and disposal.
5.0 Process Economics
See Tables E-16 and E-17 for actual and estimated costs.
6.0 Input Streams
6.1 Influent Waste (Stream 1) - Wastewater characteristics vary depend-
ing on the source. See Tables E-13, E-14 and E-19 for coal conver-
sion wastewater characteristics. See Table E-20 for listing of
chemical classes in coal gasification wastes.
6.2 Nutrients/pH Adjustment Chemicals (Stream 2) - see Section 2.6
6.3 Air or Oxygen (Stream 3) - See Section 2.6 and Table E-15.
7.0 Intermediate Streams
7.1 Fugitive Emissions (Stream 4) - NH^ HgS, .mercaptans and other
malodorous organic compounds may be released during routine opera-
tions and especially during upsets. No data are currently avail-
able on quantities and characteristics of such emissions.
E-69
-------
TABLE E-16. ESTIMATED COST OF HYPOTHETICAL DESIGN HIGH PURITY OXYGEN
ACTIVATED SLUDGE SYSTEM FOR A COAL CONVERSION PLANT*(4,16)
Capital Costs
10° $ (1977)
Equalization
Step 1 HPOAS:
Oxygenation Basins
Clarification
Cooling Tower
Pumps for Recirculation
Step 2 HPOAS:
Oxygenation Basins
Clarification
Pumps for Recirculation
Oxygenation Equipment and Related Instrumentation
for Steps 1 and 2t
Installation and Oxygenation Equipment and Related
Instrumentationt
DAF Thickening
Vacuum Filtration
TOTAL
1.19
1.96
0.25
0.08
0.13
0.45
0.29
0.02
3.50
0.32
0.54
0.36
9.09
Operating Costs
10° $/yr (1977)
Amortization and other capital-related items
at 15% of capital/yr
Maintenance:
Concrete work
Machinery
Electricity at 2,470 kw*
Chemicals:
Phosphorous
Oxygen, 295 tons/day at $14.32/ton
TOTAL
TOTAL OPERATING COSTS
1.36
0.05
0.08
0.40
0.32
1.40
3.61
0.95 $/1000 liter
(3.61 $/1000 gal)
*See Table E-13 for waste characteristics, design parameters, and utility
requirements.
tQuotation from Union Carbide.
^Excluding electricity required for oxygen generation.
E-70
-------
TABLE E-17.
YEARLY COSTS FOR TREATMENT OF 1,000 GPM OF OILY
PETROLEUM REFINERY WASTEWATER USING AN ACTIVATED
SLUDGE TREATMENT *™{ Aumitu
Cost
Dollar Value
($000)
Investment (excluding land)
Operating Costs
1. Power ($0.01/hp)
2. Maintenance at 4% plant cost
3. Direct Labor and Overhead
4. Depreciation at 10% plant
5. Insurance and Taxes at 3% plant cost
6. Chemicals
$/1000 liters ($/1000 gal)
1,160
9.6
46.4
40.0
116
34.8
1.0
247.8
0.13
(0.5)
*A11 costs are in 1970 dollars.
8.0 Discharge Streams
8.1 Effluent to Clarifier for Solids Separation (Stream 5) - Section 2.5.
See Table E-18 for effluent data for Synthane wastes.
8.2 Sludge to Recycle (Stream 6) - Approximately 20% of the BOD removed
is discharged as sludge in conventional activated sludge systems.
A portion of the sludge is returned to the biological reactor to
"seed" the raw wastewater. Sludges contain approximately 2%-5%
solids. The chemical composition depends on the influent waste and
nature of treatment. Generally, sludges contain biological floes,
heavy metals, undegraded or organic degradation products and inerts.
No composition data are available for coal gasification wastes.
8.3 Excess Sludge to Treatment/Disposal (Stream 7) - See Section 8.2.
E-71
-------
TABLE E-18. TYPICAL PERFORMANCE OF BIOLOGICAL TREATMENT OF
SYNTHANE WASTES*(25)
Applied F/M, kg TOC/kg
MLVSS-day
TOC, mg/1
Influent
Effluent
TOC Removal , percent
COD, mg/1
Influent
Effluent
COD Removal , percent
Phenol , mg/1
Influent
Effluent
Phenol Removal , percent
MLVSS, mg/1
Run 1
0.7
1960
850
57
5960
2030
64
1205
25
98
2750 ;
Run 2
0.2
i
500
!
150
70
1250
390
69
175
<.!
99
2520
treatment unit was a 7-liter activated sludge bioreactor; hydraulic
retention time = 24 hours. Wastewater was generated in the Synthane
PEDL) at Pittsburgh Energy Research Center.
E-72
-------
TABLE E-19. SUMMARY OF CLASSES OF ORGANIC CONSTITUENTS IN COAL GASIFICATION
RAW GAS QUENCH CONDENSATES (ALL CONCENTRATIONS IN mg/1)
Constituent Class
Monohydric phenols
Dihydric phenols
Polycyclic hydroxy
compounds
Monocyclic n-aromatics
Polycyclic n-aromatics
Aliphatic acids
Others (i.e. ,
benzofurans,
benzofuranols,
benzothiophenols,
hydroxy benzaldehydes ,
benzoic acid)
Synthane
TPR-86U8)
1690-9380
*
90-660
30-580
0-210
--
210-580
(19^
Synthanev ;
5250
— t
40
—
—
730
Lurgi-,2Qx
Westfield^u;
1833-4560
546-1751
—
—
—
--
Synthane^21)
4506
—
66
46
97
--
SASOL^22)
2410
7718
234
--
226
m
i
*Dihydric phenols have been identified in the wastes; however, concentrations have not been
determined.
tlndicates data not available.
-------
9.0 Data Gaps and Limitations
Limited data are available on the characteristics and biotreatability of
wastewaters from commercial coal gasification facilities. Data for pilot
facilities may not be representative of those for large scale operations.
10.0 Related Programs
An experimental program to determine the composition, biodegradability
and biodegradation kinetics of organics in coal conversion wastewaters is
currently underway at the University of North Carolina under an EPA con-
tract^23 . Treatability studies, including an assessment of activated
sludge, trickling filters and anaerobic treatment methods for coal con-
version wastewaters, are currently being conducted by the Environmental
Studies Institute, Carnegie-Mellon University, Pittsburgh, Pennsylvania,
under contract with DOE^ . Biological treatment of Synthane wastewaters
is currently being investigated under PERC sponsorship at the Synthane
pilot plant. Limited data from this study have already been published
and additional results are anticipated^ '.
REFERENCES
1. Hardistz, D. M., and H. E. Bishop, Jr., Wastewater Treatment Experience
at Organic Chemical Plants Using a Pure Oxygen System, in AIChE Symposium
Series, Water - 1976 II. Biological Wastewater Treatment, Vol. 73, 19,
p. 140-144.
2. Information provided by South African Coal,Oil and Gas Corp. Ltd., to
EPA's Industrial Environmental Research Laboratory (Research Triangle
Park), November 1974.
3. Massey, M. R., R. W. Dunlap, et al, Characterization of Effluents from
the Hygas and ^-Acceptor Pilot Plant, Interim Report for the Period
July-September 1976, ERDA Document No. FE-2496-1, November 1976.
4. Wei, I. W., and D. J. Goldstein, Biological Treatment of Coal Conversion
Condensates, presented at Third Symposium on Environmental Aspects of
Fuel Conversion Technology, Hollywood, Florida, Water Purification
Associates, Cambridge, Mass., September 1977, 31 pp.
5. Sawyer, C. N., and P. L. McCarty, Chemistry for Sanitary Engineers,
McGraw-Hill Book Co., 1976.
E-74
-------
6 jsJ&,Et .tfs. ^r ^
Symposium on Environmental Aspectsof Fuel Conversion Technology II! SSlly-
wood, Florida Environmental Protection Agency, Research Triangle Park, No.
Carolina, hPA-oUu/z-67-i49.
7. Kostenbader, P. D. and J. W. Flecksteiner, Biological Oxidation of Coke
?!a?^ e?nJ%Sma Llcluor> Journal Water Pollution Control Federation,
41 (2), 199-207, February 1969.
8.' Baker, J. E. and R. J. Thompson, Biological Removal of Carbon and Nitro-
gen Compounds from Coke Plant Wastes, EPA-R2-73-167, April 1973.
9. Environmental Assessment of the Hygas Process, Report to ERDA from the
Institute of Gas Technology, Chicago, Illinois, NTIS No. FE-2433-8,
May 1977.
10. Environmental Assessment of the Hygas Process, Report to ERDA from the
Institute of Gas Technology, Chicago, Illinois, NTIS No. FE-2433-13,
August 1977.
11. Herbes, S. E., G. R. Southworth and C. W. Gehrs, Organic Contaminants in
Aqueous Coal Conversion Effluents: Environmental Consequences and
Research Priorities, Oak Ridge National Laboratory, Oak Ridge, Tenn.,
CONF-760632, 1976, 18 pp.
12. Parsons, W. A., and W. Nolde, Applicability of Coke Plant Water Treatment
Technology to Coal Gasification, presented at Third Symposium on Environ-
mental Aspects of Fuel Conversion Technology, Hollywood, Florida,
September 1977, 15 pp. A. G. McKee & Co., and McKee-Otto Engineers and
Constructors, Cleveland, Ohio.
13. Development Document for Proposed Effluent Limitations Guidelines and New
Source Performance Standards for Petroleum Refining, U.S. Environmental
Protection Agency, Washington, D. C., December 1973, p. 110.
14. Gloyna, E. F., and D. L. Ford, Petrochemical Effluents Treatment Practices-
Summary, Engineering-Science, Inc., Austin, Texas, PB-192-310, Water Pollu-
tion Control Research Series, February 1970, 98 pp.
15. Azad, H. S., Industrial Wastewater Management Handbook, McGraw-Hill Book
Co., New York, N.Y. 1976, p. 3-17.
16. Goldstein, D. J., and A. Yung, Water Conservation and Pollution Control
in Coal Conversion Processes, Water Purification Associates, Cambridge,
Mass., EPA-600/7-77-065, PB-269-568, June 1977, 482 pp.
17. Thompson, C. S., J. Stock, et al, Cost and Operating Factors for Treatment
of Oily Waste Water, The Oil and Gas Journal, 70(47), pp. 53-56, November
1972.
E-75
-------
18. Forney, A. J., W. P. Haynes, et al, Analysis of Tars, Chars, Gases, and
Water in Effluents from the Synthane Process, U.S. Bureau of Mines Tech-
nical Progress Report 76, Pittsburgh Energy Research Center, Pittsburgh,
Pa. 1975.
19. Schmidt, C. E., A. G. Sharkey, et al, Mass Spectrometric Analysis of Pro-
duct Water from Coal Gasification, U.S. Bureau of Mines Technical Progress
Report 86, Pittsburgh Energy Research Center, Pittsburgh, Pa., 1974,
20. Janes, T. K. and W. J. Rhodes, Industrial Environmental Research Labora-
tory, Environmental Protection Agency, personal communication.
21. Spinola, A. A., Ozonation of Process Wastewaters from the Production of
Synthetic Natural Gas Via Coal Gasification, M. S. Report, Department of
Civil Engineering, University of Pittsburgh, Pa., 1976.
22. Jolley, R. L., W. W. Pitt, et al, Organics in Aqueous Process Streams of
a Coal Conversion Bench-Scale Unit Using the Hydrocarbonization Process:
HPLC and GC/MS Analysis, Environmental Technology Annual Technical Meet-
ing of the Institute of Environmental Sciences, Los Angeles, Calif., 1977.
23. Sincer, P. C., F. K. Pfaender, et al, Composition and Biodegradability of
Organics in Coal Conversion Wastewaters, University of No. Carolina,
Chapel Hill, N. C. September 1977, 31 pp.
24. Massey, M. J., R. W. Dunlap, et al, Environmental Assessment in the ERDA
Coal Gasification Development Program, Carnegie-Mellon University,
Pittsburgh, Pa., March 1977, 154 pp.
25. Johnson, G. E., Neufeld, R. D., et al, Treatability Studies of Condensate
Water from Synthane Coal Gasification, PERC/RI-77/13, Pittsburgh Energy
Research Center, Pittsburgh, Pa., 1977.
E-76
-------
EVAPORATION/RETENTION POND(1'2'3)
1.0 General Information
Evaporation/retention ponds or lagoons are natural or man-made basins
constructed either by digging out a depression on the land or by erecting
dikes. Waste is discharged to the pond and water is allowed to evaporate,
thus reducing the waste volume and making room for additional waste. The
solids or sludge may be removed and landfilled, or the waste may perma-
nently remain at the pond site.
1.1 Applicability - Method is most suitable when large land areas are
available, there is a significant net evaporation rate, and there is
little risk of contaminating groundwater.
1.2 Development Status - Evaporation/settling ponds have been used widely
for the disposal of municipal and a wide variety of industrial wastes,
Ponds are used at the SASOL-Lurgi facility and at all coal gasifica-
tion pilot plants in the U.S. and have been featured in all proposed
commercial scale SNG facilities.
1.3 Operating Parameters - Some operating parameters and design considera-
tions include: available land area, climatic and atmospheric condi-
tions, subsoil permeability and distance to surface/groundwaters,
pond depth and volume. Ponds may require lining with clays, plastic
or other impervious material to prevent groundwater contamination.
2.0 Advantages
• Method is simple and economical to use.
• Wide variations in waste types and loadings are accommodated.
• Minimal maintenance is required.
• Can generate no effluent streams requiring further treatment or
disposal.
0 The clarified water may be suitable for recycling to plant.
E-77
-------
3.0 Disadvantages
• Adequate protection must be provided against surface and groundwater
contamination (e.g., use of liners; diversion of surface runoff, etc.).
• Ponds must be provided with suitable containment mechanisms to prevent
overflow due to rainfall accumulation. Operation is dependent on
climatological conditions. In areas of heavy rainfall, flood protection
equipment may be difficult and expensive to provide.
• Method depends on the availability of adequate land and suitability of
climate.
• Leachate and undesirable odors may be generated, depending on the type
of waste deposited.
• Pond must be monitored for leachate and for erosion control.
t In nonpermanent sites, the deposited solids must be excavated for
ultimate disposal, usually by landfilling.
• Surfaces of ponds used for permanent disposal may require stabilization
to prevent erosion by wind and precipitation.
• Ambient air above the pond may pick up low levels of volatile materials
from the influent waste.
4.0 Process Economics - Depends on the quantity of waste handled, land area
required, the cost of labor and equipment (i.e., drainage pipes, monitor-
ing equipment, etc.). When the climate is suitable and large land areas
available, use of evaporation/retention basins can be the most economic
method for waste disposal.
5.0 Related Programs^ ' - Radian Corporation is about to conduct an EPA-
sponsored program to assess state-of-the-art holding pond design, construc-
tion and management, and to investigate the interactions of chemicals in
coal conversion effluents or clay liners used in holding pond construction.
E-78
-------
REFERENCES
1. Powers, P. W., How to Dispose of Toxic Substances and Industrial Wastes,
Noyes Data Corporation, Park Ridge, N.J., 1976, p. 25.
2. Cavanaugh, E. C., J. D. Colley, et al, Environmental Problem Definition for
Petroleum Refineries, Synthetic Natural Gas Plants, and Liquefied Natural
Gas Plants, Radian Corporation, Austin, Texas, EPA-600/2-75-068, PB-252-245,
November 1975, p. 327.
3. The Cost of Clean Water, U.S. Department of the Interior, Federal Water
Pollution Control Administration, Washington, D.C., 1967.
4. White, I. L., M. A. Chartrock, et al, Work Plan for Completing a Tech-
nology Assessment of Western Energy Resource Development, University of
Oklahoma City, Oklahoma, EPA-600/7-78-012, February 1978, 70 pages.
E-79
-------
CHEMICAL OXIDATION PROCESS
1.0 General Information
1.1 Operating Principle - Use of chemicals (primarily ozone, chlorine,
chlorine dioxide, and oxygen/air)* to oxidize phenols, cyanides,
sulfides, thiocyanates, refractory organics and other wastewater
constituents, and to reduce the COD and BOD of the waste; ozone and
chlorine compounds are also used for water disinfection .
1.2 Development Status - Commercially available. Numerous units in
operation throughout the world for municipal and industrial water
and wastewater treatment.
1.3 Licensor/Developer - Many chemical oxidation systems are offered by
numerous suppliers. Some licensed or patented versions include the
(2}
UV-OX process for organics removalv , the Zimpro unit air oxidation
(3)
system for wastewater and sludge treatment , and processes for air
oxidation of sulfidic (including ammoniacal sulfidic) sour waters
and sulfidic spent caustics from petroleum refineries (e.g., Sulfox
process)^ ' .
1.4 Commercial Applications - Applications include municipal water and
wastewater treatment plants, petroleum refineries, coke plants and
numerous other industries. Although there have been no commercial
or pilot-scale applications to coal gasification wastewaters,
laboratory-scale testing of Hygas pilot plant wastes using the
Zimpro process has been conducted^ '.
*Numerous other chemical oxidants, such as ^2 and MnO/T, have been or are
currently being used in water and wastewater treatment. The use of these
chemicals has been on a small scale, and they are not considered in this data
sheet.
E-80
-------
2.0 Process Information
2.1 Flow Diagram (see Figure E-12) - Influent waste (Stream 1) is reacted
with a chemical oxidant (Stream 2), under controlled conditions
(e.g., temperature, mixing regime, reaction time, etc.) in the reac-
tion unit. Other feed materials (e.g., air, oxygen, etc. -
Stream 3), are supplied as required by the specific process employed.
Treated effluent (Stream 4) and sludge consisting of reaction
products and byproducts (Stream 5) are discharged.
2.2 Equipment
• Chemical oxidation reaction vessel - Design varies with the
specific process utilized. Air oxidation of sulfidic sour
waters is usually carried out in pressure vessels or multi-
stage oxidation towers)(5>8)
• Oxidant source equipment - Varies with the type of process
utilized. For ozone oxidation, 03 generator is required,
including air cleaning and drying equipment when the 03 is
manufactured from air.
• Pumps, heat exchangers, mixers, compressors, driers, etc.,
as required.
2.3 Feed Stream Requirements - Vary with the specific process; many are
applicable over wide ranges of wastewater compositions. Most
chemical oxidation processes are highly pH dependent; the optimum
pH varies with the specific process reactant characterization and
reaction time involved (usually 7 or greater for C12). Other impor-
tant feed variables include temperature (in the case of chlorine,
the waste temperature must be below 316°K (110°F) before C12 is added
to prevent CIO- formation), oxidant concentrations, the presence of
O
inhibitory constituents, and the presence of rate-improving or
(9)
mechanism-directing catalysts' .
2.4 Operating Parameters - Temperature, pressure and reaction times vary
with the specific waste and process utilized. See Table E-20 for
operating parameters of an air oxidation column for sulfidic
petroleum refinery wastewaters. Generally 1-2 hours or less reac-
tion time are required for chlorine oxidations^ '; less than one
hour reaction time is usually required for C102
E-81
-------
CHEMICAL OXIDATION
REACTION UNIT
LEGEND:
1. Influent Waste
2. Chemical Oxidant
3. Raw Materials/Additives
(air, steam, etc)
4. Treated Effluent
5. Sludge
Figure E-12.
Simplified Schematic of Chemical
Oxidation Systems
E-82
-------
TABLE E-20. OPERATING PARAMETERS FOR AIR OXIDATION COLUMN
FOR SULFIDIC PETROLEUM REFINERY
Parameter
Design Value
Air Flow, m /min (ft3/min)
Temperature, °K. (°F)
Pressure (bottom)
psig (atm)
Water Flow, I/day
(gal/day)
Air-Water Ratio (approx.
inlet conditions), m3/m3
Vessel Volume, m3 (ft3)
S in feed, mg/1
S oxidized, tonnes/day
(tons/day)
S~ Oxidation Rate,
kg/hr/m3 (lb/hr/ft3)
Excess Air, %*
37.5 (1,325)
366 (200)
85 (5.8)
7.15 x lof
(1.9 x 10D)
9.7
54.5 (1,925)
8,000
7.6 (8.4)
5.8 (0.36)
100
*Basis for oxidation of sulfide to thiosulfate.
2.5 Process Efficiency and Reliability - Efficiency depends on the type
and design of process used, and on the nature of the wastewaters.
See Tables E-21 through E-23 for efficiencies of: (a) Zimpro wet
air oxidation of Hygas pilot plant wastewaters; (b) ozonation of
chlorinated hydrocarbons in petrochemical wastewaters; (c) coke
plant wastewater oxidation by C12> C102 and 03; and (d) ozonation
of mixed industrial-municipal wastewaters. Efficiencies of
removal of contaminants from oil refinery wastes are shown in
Table E-24-
E-83
-------
TABLE E-21. OPERATING EFFICIENCY FOR ZIMPRO WET AIR OXIDATION OF HYGAS PILOT PLANT WASTEWATERS
(3)
Parameter
Temperature, °K (°F)
Time, min
COD, g/1
% COD Reduction
Total Solids, g/1
Total Ash, g/1
PH
NH3 as N, g/1
TKN, g/1
Total S, g/1
Total Hal ides as Cl , g/1
Phenol , mg/1
Cyanide, mg/1
Thiocynate, mg/1
BOD5, mg/1
Catalyst
Feed
287
(56)
—
13.7
—
1.75
0.36
8.2
3.25
3.25
0.17
0.1
740
0
0
—
—
694-56-1
513
(464)
60
4.9
64.2
1.36
0.34
8.2
2.84
3.04
0.13
0.1
<1.0
0
0
2350
No
V
h
Removal
—
—
64.2
—
22.2
5.5
—
12.6
6.5
23.5
0
<0.13
—
—
—
—
Sample
694-55-1
553
(536)
60
3.3
75.9
1.16
0.37
8.1
2.88
3.01
0.17
0.1
<1.0
0
0
190
No
01
h
Removal
—
—
75.9
—
—
—
11.4
7.4
0
0
0.13
—
—
—
—
694-58-1
553
(536)
60
1.0
92.7
2.46
0.38
8.0
3.07
3.29
0.43
0.1
<1.0
0
0
—
Yes
°/
k
Removal
—
—
92.7
—
—
—
—
5.5
0
0.13
—
—
—
—
00
-------
TABLE E-22. ^REMOVAL^FFICIENCIES FOR OZONATION OF PETROCHEMICAL
Ozone Dosage,
mg/1
994
2,530
2,700
3,920
4,640
5,400
pH Initial
12.2
12.6
7.0
12.6
12.6
12.6
COD,
Raw Waste
mg/1
3,340
3,340
3,340
3,340
3,340
3,340
COD,
Treated Waste,
mg/1
1,410
900
1,460
745
450
413
01
la
Reduction
57.8
73
56.5
77.5
86.5
90.5
*Exact composition of wastewater is unspecified.
TABLE E-23. EFFICIENCIES AND COSTS FOR COKE PLANT WASTEWATER
OXIDATION BY C12, C102, AND
Treatment
Method
Chlorine
Chlorine
dioxide
Ozone
Liquor
Treated
Bio-effluent,
NhL removed
Bio-effluent
Bio-effluent
Cyanide
Concentration,
mg/1
Influent Effluent
10.0 <1.0
4.0-5.2 1.8-3.6
2.0-5.0
Reduction
Efficiency,
%
90+
30-55
Not
Effective
Approximate Cost,*
$/million liters
($/million gal)t
35.9 (9.50)
68.0 (18.00)
_ _ _
*Costs are for chemicals only; plant capital costs are not included.
tin 1969 dollars.
E-85
-------
TABLE E-24. EFFICIENCY OF OZONATION OF OIL REFINERY
WASTEWATERSU4)*
Parameter
Removal Efficiency, %
BOD
COD
Phenol
Sulfide
Suspended Solids
Chloride
Ammonia
Cyanide
Toxicity
50 - 90
50 - 90
80 - 99
80 - 99
Not Applicable
Not Applicable
10 - 30
80 - 99
Reduced
*Wastewaters are secondary effluents from chemical or biological
treatment.
2.6 Raw Materials Requirements
• Chemical oxidant (e.g., oxygen, chlorine, etc.) - Actual quantity
varies with the oxidant used and species being oxidized. Some
oxidants (e.g., ozone, C102) require on-site generation from
chemical elements.
t Chemical additives (e.g., for pH adjustment) - Varies, depending
on the optimum pH for oxidation and the wastewater character-
istics. (In certain applications, chemical oxidants are combined
with activated carbon or other materials which serve as catalysts
in the oxidation process and result in more effective BOD and COD
removal)l°).
2.7 Utility Requirements
• Electricity - Varies with the specific process used. For 0
generation, one kwh is required to generate 150g (0.33 Ib)
• Stam - Required in air oxidations. See Table E-20 for petroleum
refinery application.
E-86
-------
3.0 Process Advantages'6'13'
• Effective for removal of refractory organics not amenable to biological
treatment and wastes containing toxic chemicals.
• Suitable for treatment of certain organics and inorganics present in
municipal and industrial wastewaters (e.g., phenol, sulfide, cr,
SCN-, etc.). Some of these constituents will be components of coal
gasification effluents.
• Some processes (e.g., air oxidation of sulfidic wastewaters and chlori-
nation) are widely used commercial processes for which extensive
operating experience is available.
• Most processes impart no taste or order to the treated water and
wastewater.
• Some processes result in the destruction of microorganisms and dis-
infection of the wastewater being treated.
4.0 Process Limitations^6'14'15^
• Generally applicable to small, concentrated waste streams, usually
where biological oxidation is ineffective (e.g., when the wastes con-
tain toxic chemicals or refractory organics).
0 Effects of chlorination and ozonation not completely understood. Can
result in the production of potentially hazardous substances (e.g.,
chlorinated hydrocarbons when chlorine compounds are used as oxidants).
• 03 and C102 are unstable and require on-site generation. Capital
costs are modest; operating costs may be high.
5.0 Process Economics
Cost of 03 oxidation is 0.53-0.79
-------
on the source. See Tables E-21 and E-25 for coal gasification waste-
water characteristics.
TABLE E-25. PERFORMANCE OF BENCH SCALE OZONE TREATMENT OF SYNTHANE PROCESS
WATER (ALL UNITS ARE mg/1 EXCEPT pH AND OZONE RATE)(20)
Constituent/Parameter
Phenol
Cyanide
Thiocyanate
Ammonia
COD
BOD
TOC
Pyridine and Picolines
Lutidines
Napthalene and Aniline
Toluidines
2-Methyl naphthalene and
Xylidine
2-6-Xylenol
Quinoline
Phenol and 0-cresol
M,p-cresol; 2,3 and
2,5-Xylenol
Methyl quinoline
2,3-Xylenol
3,5-Xylenol; m,p-ethyl phenol
3,4-Xylenol
3-Ethyl-5-Methyl phenol
C_-Phenols
4-Indanol
Indol
PH
Raw
Wastewater
2,320
2.28
418
4,250
17,162
420
5,800
—
—
—
—
—
—
—
—
—
—
—
—
—
—
—
—
—
9.4
Ozone Rate
0.3 mg/ml
2,340
1.31
445
4,200
17,062
399
5,600
17
5
19
8
10
23
7
2,084
1,527
22
44
333
102
54
26
47
55
9.4
1.2 mg/ml
2,225
0.6
450
4,010
16,504
340
3,800
17
6
18
5
8
17
3
1,990
1,438
24
44
330
99
55
27
40
50
9.4
3 mg/ml
820
10.2
240
3,710
10,020
286
—
9
1
3
2
1
6
2
443
308
0
14
77
16
23
21
16
16
8.0
E-88
-------
TABLE E-26.
CAPITAL AND ANNUAL OPERATING COSTS OF
.,.,.,.._-_„, , ' '•• * • w '"' wni_/ un i ; vtviic
MUNICIPAL WASTEWATER TREATMENT PLANTS, IN
THOUSANDS OF 1977 DOLILARS^T) '
Capital Costs
Ozone Generators
Compressors and Driers
Mixers and Pumps
Reactors
Piping and Electrical
Building and Supports
Other
Overhead, Fees and Profit
Total
Operating Costs
Electric Power
Amortization
Ozone Feed
per
40
$ 440
112
98
200
230
160
146
416
$1,302
Ratio (mg ozone generated
liter of feedwater)
75
$ 770
180
98
200
276
170
146
552
$2,392
100
$ 980
212
98
200
289
170
146
628
$2,723
$ 147
128
Oxygen i 37
Operation and Maintenance
Total
41
$ 353
$/1000 gal 9.7
$ 208
170
44
47
$ 469
12.8
$ 245
193
48
50
$ 536
14.7
E-89
-------
6.2 Chemical Oxidant (Stream 2) - See Section 2.6.
6.3 Raw Materials (Stream 3) - See Sections 2.6 and 2.7 and Table E-20.
7.0 Discharge Streams
7.1 Treated Effluent (Stream 4) - Varies with the type of waste treated
and nature of the oxidant. Effluent is likely to contain quinones
(from oxidation of phenol), cyanates (from CN~ oxidation) and thio-
sulfates (from sulfides and mercaptans). May also contain chlorinated
hydrocarbons and aromatics due to incomplete chlorination, or
ozonides and epoxides if ozone was used. See Tables E-21 to E-25
for actual effluent data.
7.2 Sludge (Stream 5) - Some processes generate sludges which may contain,
depending on the nature of the waste, heavy metals, nondegradable
organics, and inerts.
(??)
8.0 Data Gaps and Limitationsv '
Data needed for engineering and cost estimates for a commercial-scale
chemical oxidation facility as an alternative to biological oxidation
for treatment of coal conversion wastewaters are currently unavailable.
For example, the rate and extent of oxidation of polyhydric and substi-
tuted phenolics, such as those found in coal gasification wastes, by
oxidants such as ozone, are not known. Also, the efficiency of ozonation
as a function of pH is not well defined.
9.0 Related Programs
The EPA is currently sponsoring a program in the city of Wyoming, Michigan,
to develop an understanding of the effects of ozonated effluents on the
(231
environmentv '. Other related programs are not known.
E-90
-------
REFERENCES
L Sinn" ?QfiS0luUmh-nSJ1tUtnVManUal °n DisP°sal of Refine^ Wastes, First
Edition, 1969, Washington, D.C., Chapter 11.
2. Zeff, J D.} UV-pX Process for the Effective Removal of Organics in
Wastewaters, in Water-1976: II. Biological Wastewater Treatment, AIChE
Symposium Series, Volume 73, No. 167, 1977, p. 206-220.
3. Water Conservation and Pollution Control in Coal Conversion Processes,
X!!tolJ>!£iflcat1on Associates> Cambridge, Mass., EPA-600/7-77-065, NTIS No.
PB-269-568, June 1977, p. 285-314.
4. Canada Patent No. 601,035.
5. Beychok, M. R., State-of-the-Art Wastewater Treatment, Hydrocarbon Pro-
cessing, December 1971, p. 109-112.
6. Weber, W. 0., Jr., Physiochemical Processes for Water Quality Control,
Wiley-Interscience Publishers, New York, N.Y., 1972, p. 364.
7. Azad, H. S., Industrial Wastewater Management Handbook, McGraw-Hill
Book Co., New York, N.Y., 1976, p. 8-58.
8. American Petroleum Institute, Manual on Disposal of Refinery Wastes,
First Edition, 1969, Washington, D.C., Chapter 15.
9. Watkins, J. P., Controlling Sulfur Compounds in Wastewaters, Chemical
Engineering/Deskbook Issue, Vol. 84 (No. 22), October 17, 1977, p. 61-64.
10. Martin, J. D. and L. D. Levanas, Air Oxidation of Sulfide in Process
Water, Proc. API 36 (III), 313-7, 1956.
11. Gloyna, E. F., and D. L. Ford, The Characteristics and Pollutional Prob-
lems Associated with Petrochemical Wastes, Federal Water Pollution Con-
trol Administration, Ada, Oklahoma, 1970.
12. Kostenbader, P. D. and J. W. Flecksteiner, Biological Oxidation of Coke
Plant Weak Ammonia Liquor, Journal of the Water Pollution Control Federa-
tion, Vol. 41 (No. 2), 199-207, 1969.
13. Mulligan, T. J. and R. D. Fox, Treatment of Industrial Wastewaters,
Chemical Engineering/Deskbook Issue, Volume 83 (No. 22), October 18,
1976, p. 64-66.
H. The Cost of Clean Water, Federal Water Pollution Control Administration,
November 1967, p. 66.
15. Majumdar, S. B., W. H. Ceckler and 0. J. Sproul, A Physical and Mathe-
matical Model of Mass Transfer and Reaction Kinetics of Ozonation, in
Water-1976: I. Physical, Chemical Wastewater Treatment, AIChE Symposium
Series, Volume 73, No. 166, p. 188-205.
E-91
-------
16. Bush, K. E., Refinery Wastewater Treatment and Reuse, Chemical
Engineering, April 12, 1976, p. 113-118.
17. Summary Report, The Advanced Waste Treatment Program, January 1962 through
June 1964, U.S. Public Health Service, Division of Water Supply and Pollu-
tion Control, Washington, D.C., April 1965.
18. Sondak, N. E. and B. F. Dodge, The Oxidation of Cyanide-Bearing Plating
Wastes by Ozone, Plating, Part I, Vol. 48 (No. 2), 173-80 (1961); ibid,
Part 2, No. 3, 280-4.
19. Water Purification Associates, Innovative Technologies for Water Pollu-
tion Abatement, Report No. NCWQ 75/3, National Commission on Water
Quality, Washington, D.C., NTIS No. PB-247-390, December 1975.
20. Wynn, C. S., B. S. Kirk, et al, Pilot Plant for Tertiary Treatment
of Wastewater with Ozone, in Water-1972, AIChE Symposium Series, Vol. 69,
No. 129, 1963, p. 42-60.
21. Anderson, G. L., Ozonation of High Levels of Phenol in Water, in Water-
1976: I. Physical, Chemical Wastewater Treatment, AIChE Symposium
Series, Volume 73, No. 166, 1977, p. 265-271.
22. Information provided to TRW by Dr. R. Johnson, University of No. Carolina,
February 13, 1978.
23. Disinfection of Wastewater, EPA Agency Task Force Report, July 1975.
24. Neufeld, R. D. and Spinola, A. A., Ozonation of Coal Gasification Plant
Wastewater, Environmental Science and Technology, Vol. 12,'No. 4, April
1978.
E-92
-------
PHENOSOLVAN PROCESS
1.0 General Information
1.1 Operating Principles - Extraction of phenols and other organics
from process/waste water using organic solvents.
1.2 Development Status - Commercially available.
1.3 Licensor/Developer - Lurgi Mineralotechnik GmbH
American Lurgi Corporation
377 Rt. 17 South
Hasbrouck Heights, M.J.
1.4 Commercial Applications^1' - Since 1940 over 30 commercial Pheno-
solvan plants have been installed worldwide, including plants at
Sasolburg, South Africa and Kosovo, Yugoslavia which process waste
waters from Lurgi gasification operations.
2.0 Process Information
2.1 Flow Diagranr ' (see Figure E-13) - Filtered phenol containing
wastewater is fed to a mixer-settler where it contacts lean organic
solvent. After solvent-water phase separation, the solvent is sent
to a distillation column for solvent recovery. Lean solvent from
the column returns to the mixer-settler while crude phenol is
fractionated for purification and additional solvent recovery.
The'dephenolized wastewater is stripped of solvent with nitrogen (N^)
gas in a packed tower. Solvent rich N2 gas is then contacted with
scrubbing phenol from the crude phenol stripper to recover most of
the solvent. Phenolic vapors remaining in the N2 gas are then largely
removed via contact with a portion of the feed wastewater. Clean
H9 returns to solvent recovery scrubber and the feed wastewater
proceeds to the mixer settler.
2.2 Equipment - Filter bed (sand or gravel), mixer-settler, distillation
columns, packed towers. All vessels are carbon steel and operate at
low pressure.
E-93
-------
FILTER
BED
I
vo
^ f MIXER- V*—~
II SETTLER |7
SOLVENT
DISTILLA-
TION
COLUMN
CRUDE
PHENOL
STRIPPER
12
' »
14
PHENOL
RECOVERY
SCRUBBER
SOLVENT
RECOVERY
SCRUBBER
I
LEGEND:
1. GAS LIQUOR FEED
2. MAKEUP SOLVENT (LOCATION
NOT KNOWN)
3. LEAN SOLVENT
4. RICH SOLVENT
5. RAFFINATE
6. DEPHENOLIZED GAS LIQUOR
7. CRUDE PHENOL
8. STRIPPED PHENOL
SOLVENT
RECOVERY
SCRUBBER
9. STEAM
10. NITROGEN
11. SOLVENT RICH VAPOR
12. RECYCLE PHENOL
13. PHENOL RICH VAPOR
14. RECYCLE GAS LIQUOR
15. FILTER BACKWASH SLUDGE
Figure E-13. Flow Diagram for Phenosolvan Process
(1)
-------
2.3 Feed Stream Requirements - Incoming wastewater is conmonly filtered
to remove suspended solids. Such materials can cause foaming and
sludge buildup and can reduce process efficiency if not largely re-
moved prior to the solvent extraction step.
Depending on the solvent used, feed water may require cooling to pre-
vent excessive solvent losses. Generally, the extraction step is con-
ducted at about 300°K-345°K (100°F-160°F).
2.4 Operating Parameters
0 Pressure: atmospheric
300
solvent.
(l 9}
9 Temperatureu' ': 300°K-345°K (100°F-160°F) depending on the
* t Loading ': weight flow ratio of wastewater to solvent is about
10 for typical designs.
2.5 Process Efficiency and Reliability - A design for a commercial Pheno-
solvan unit operating on Lurgi gasification wastewater has assumed
99.5% removal of monohydric phenols, 60% removal of polyhydric
ac
,(2)
phenols, and 15% removal of other organics^ '. A commercial facil-
ity at Sasolburg, South Africa reports >96% total phenol removal
2.6 Raw Material Requirements
• Solvent Properties: Solvents which have been used in Phenosolvan
plants include butyl acetate, isopropyl ether, and light aromatic
oil(3). Some properties of these solvents are listed below'3»4).
Solvent
Butyl Acetate
Isopropyl Ether
Aromatic Oil
Phenol
Distribution
Coefficient,
KD*
49
20
^22
Solubility
in Hater
(wt %)
1.0 at 308°K
0.8 at 308°K
'vO.l at 275°K
Boiling Point
°K (°F)
398 (256)
338 (148)
353+ (175+)
*,/ _ wt fraction of substance in solvent phase measurecj at high dilution
^n "" ~*—/- . • T"Z ^..u^4*-anoA in aniiPHMS DnaSG
!/ fv v i I m ^» i* i v»i v • ~* vt~ ~ — — — „ , .—
D " wt fraction of substance in aqueous phase
E-95
-------
For butyl acetate at 300°K (77°F) the following distribution
coefficients for various phenolic compounds have been reported
Compound KD
Phenol 65
3,5-xylenol 540
Pyrocatechol 13
Resorcinol 10
• Solvent Makeup Requirements: Makeup is required to balance
solvent losses in the crude phenol product and to a lesser
extent, in the dephenolized aqueous effluent. The SASOL plant
uses about 15 £ of butyl acetate makeup per 10° fc feed^J.
No data available for other solvents.
• Nitrogen: No data available.
2.7 Utility Requirements
• Low Pressure Steanr2': 75 kg/1000 £ feed (600 lbs/1000 gals)
• Electricity^: 1-1.7 kwh/1000 £ feed (3-6 kwh/1000 gals)
(?\
• Cooling Waterv ': 5.2 A/A feed
3.0 Process Advantages
• Process can recover phenols suitable for sale.
• Process can achieve 99.5$ removal of monohydric phenols and partial
removal of polyhydric phenols.
• Process recovers most of the solvent from the dephenolized waste-
water via nitrogen purging.
4.0 Process Limitations
• The multivessel operation requires relatively large capital investment.
• Only limited removal of non-phenolic organics is obtained.
• Process uses large amounts of cooling water to effect recovery of
phenols and solvent.
5.0 Process Economics
No current data are available on capital and operating costs of commercial
operations.
E-96
-------
6.0 Input Streams
6.1 Gas Liquor Feed (Stream 1) - see Table E-27
6.2 Makeup Solvent (Stream 2) - see Section 2,6
6.3 Nitrogen (Stream 10) - no data available
6.4 Steam (Stream 9) - see Section 2.6
7.0 Intermediate Streams
7.1 Lean Solvent (Stream 3) - No composition data available; see
Section 2.6 for flow rates
7.2 Rich Solvent (Stream 4) - no data available
7.3 Raffinate (Stream 5) - no data available
7.4 Crude Phenol (Stream 7) - no composition data available
7.5 Solvent Rich Vapor (Stream 11) - no data available
7.6 Recycle phenol (Stream 12) - no data available
7.7 Phenol Rich Vapor (Stream 13) - no data available
7.8 Recycle Gas Liquor (Stream 14) - no data available
8.0 Discharge Streams
8.1 Dephenolized Gas Liquor (Stream 6) - No operating data available.
See Table E-27 for properties of a dephenolized effluent following
ammonia and hydrogen sulfide removal.
8.2 Stripped Phenol (Stream 8) - No operating data available. A recent
estimate of the gross composition of organics recovered from gas
liquor generated by Lurgi gasification of western U.S. coals is^ ':
85% monohydric phenols
10% polyhydric phenols
5% other organics
8.3 Filter Backwash Sludge (Stream 15) - no data available.
E-97
-------
TABLE £-27. PROPERTIES OF FEED AND EFFLUENT GAS LIQUOR AT THE
SASOL PHENOSOLVAN PLANTl2)
Parameter/Consti tuent
Total Phenols
Steam Volatile Phenols
COD
Fatty Acids (as CH3COOH)
Total Suspended Solids
Total Dissolved Solids
Suspended Tar and Oil
Total Ammonia
Total Sulfide
Cyanide
Chloride
Fluoride
Sodium
Calcium
Iron
Ortho Phosphate
Conductivity (ymhos/cm)
PH
Phenosolvan
Feed (mg/1)
3250 - 4000
—
—
300
—
—
5000
10800
228
6
—
—
53
—
—
—
—
—
Phenosol van/Stri pped
Effluent* (mg/1)
160
1
1126
560
21
875
<21
215*
12*
1
25
56
—
18
1
2.5
1000 - 1800
8.4
*Dephenolized gas liquor is
The stripping operation is
process.
steam stripped to remove h^S and ammonia.
not considered part of the basic Phenosolvan
E-98
-------
9.0 Data Gaps and Limitations
Data gaps and limitations for the Phenosolvan process relate primarily
to the properties of certain process/waste streams, Limited data are
available for one coal gasification application of the process which
employs butyl acetate as a solvent. No operating data are available
for plants using other solvents.
10.0 Related Programs
Radian Corp., under contract to EPA, is conducting an environmental
sampling and analysis program at a Lurgi gasification facility at Kosovo,
Yugoslavia. This program includes the sampling of gas liquor feed to
and effluent from the Phenosolvan plant at this facility. Data are
expected to be available in 1978.
REFERENCES
1. Beychok, Milton R., Coal Gasification and the Phenosolvan Process, Amer.
Chem. Soc., Div. Fuel Chenu Prepr. 19 (5), 85-93, 1974.
2. Information provided by South African Coal, Oil and Gas Corp, Ltd. to
EPA's Industrial Environmental Research Laboratory (Research Triangle
Park), November 1974.
3. Wurm, H. J., Treatment of Phenolic Wastes, Eng. Bull., Purdue University
Eng. Ext. Serv., 132(11) 1054-73, 1969.
4. Earhart, J. P., et al, Recovery of Organic Pollutants via Solvent
Extraction, Chemical Engineering Progress, May 1977, p. 67.
5. American Lurgi Corp., Dephenolization of Effluents by the Phenosolvan
Process, company brochure.
E-99
-------
ACTIVATED CARBON ADSORPTION PROCESS
1.0 General Information
1.1 Operating Principle - Removal of organic compounds from a wastp^ater
by adsorption on activated carbon. Methods of contact include: (a)
passing the wastewater through a bed of granular carbon and (b) adding
powdered carbon directly to treatment systems. In the case of granu-
lar carbon, and powder carbon separated from treated water, regenera-
tion is usually effected by thermal treatment. Spend powdered carbon
added to biological treatment units usually exits the process as a
component of waste sludge and is not reclaimed.
1.2 Development Status - Activated carbon systems are currently employed
for both municipal and industrial wastewater treatment. In addition
to several existing commercial scale facilities, a number of pilot
scale projects in a wide variety of industries are presently
underway.
(3 8)
1.3 Licensor/Developed ' ' - No specific process developer. A number
of companies supply activated carbon products and related consultant
services. Some of the major vendors include:
Barnebey-Cheney Co., Columbus, Ohio
Calgon Corporation, Catlettsburg, Kentucky
ICI United States, Inc., Marshall, Texas
Husky Industries, Romeo, Florida
Union Carbide Corporation, Carbon Products Division,
Fostoria, Ohio
Westvaco Corporation, Covington, Kentucky
Witco Chemical Co., Petrolia, PA
1.4 Commercial Applications - Activated carbon systems for wastewater
treatment are employed in industries such as coke production, oil
refining, petrochemical production, and pesticide manufacture.
Carbon systems are also used for trace organics, and taste and
odor removal from potable water supplies.
E-100
-------
Refineries which have installed activated carbon process for
wastewater treatment include Atlantic Richfield, Carson, California,
and British Petroleum, Marcus Hook, PA(4). At least one coke plant
has tested activated carbon for treatment of wastewater^.
2.0 Process Information
2.1 Flow Diagram - see Figures E-14 and E-15.
• Granular Activated Carbon systems commonly employ two or more
beds.* The series flow Granular Carbon Adsorption system shown
in Figure E-14 provides for continuous treatment with periodic
removal of one or the other of the adsorbers from service for
backwashing and for carbon removal. Backwashing serves to remove
particulate matter from the carbon which accumulates over time and
increases bed pressure drop.
In Figure E-15 a typical thermal regeneration process is depicted.
Dewatered spent carbon enters the top of a multiple hearth furnace
where it travels downward through progressively hotter zones. The
furnace provides for: (1) drying, (2) thermal desorption,
(3) pyrolysis and carbonization, and (4) gasification. Hot reacti-
vated carbon is quenched and washed to remove fines before return
to the adsorption system. Regeneration offgas may be treated by
venturi scrubbing (as shown in Figure E-15) or by cyclone and
fabric filtration. Incineration may also be required for odor,
carbon monoxide, and hydrocarbon emission control. Wastewaters
resulting from bed backwashing, reactivated carbon quenching and
washing, and venturi scrubbing are usually returned to upstream
treatment systems (e.g., solids removal, activated sludge).
• Powdered Activated Carbon may also be employed as an additive to
biological treatment systems. In such applications a carbon
inventory is maintained by recycle of carbon containing activated
sludge and addition of fresh carbon. The carbon contained in
excess sludge is not ordinarily recovered. In addition to the
feed water, feed carbon, and treated effluent streams, the
powdered carbon system would generate a carbon-containing sludge
stream.
*Fixed beds may be arranged in series or parallel with either upflow or down-
flow design. Pulsed columns with countercurrent flow of carbon and wastewater
have also been used. The two bed series system depicted in Figure E-14 is
perhaps the most common design and is the only one specifically addressed by
this data sheet.
E-101
-------
4.*-
ADSORBER 1
O
ro
ADSORBER 2
LEGEND:
VALVE CLOSED
VALVE OPEN
1. FEED WATER
2. TREATED EFFLUENT
3. BACK WASH FEED
4. BACK WASH EFFLUENT
5. REGENERATED/MAKEUP ACTIVATED CARBON
6. SPENT CARBON
Figure E-14. Two-Vessel Granular Carbon Adsorption System^1^
-------
16
CARBON
DEWATERING
TANK
AFTERBURNER
PARTICULATE
REMOVAL
SYSTEM
17
14
MULTIPLE
HEARTH
FURNACE
O
OJ
LEGEND:
6. SPENT CARBON
7. DEWATERING EFFLUENT
8. STEAM
9. FUEL
10. QUENCH WATER MAKEUP
11. WASHED REACTIVATED CARBON
12. WASH WATER
13. WASH EFFLUENT
14. RAW REGENERATION OFFGAS
15. SCRUBBER FEED WATER
16. CLEANED REGENERATION OFFGAS
17. SCRUBBER EFFLUENT WATER
MO
CARBON
QUENCH
TANK
•THE STREAM NUMBERING SYSTEM IS
A CONTINUATION OF THAT SHOWN IN
FIGURE D-1S.
13
REACTIVATED
CARBON WASH
TANK
11
12
Figure E-15. Multiple Hearth Furnace Carbon Regeneration Systenr '
-------
2.2 Equipment^ ' - Granular Carbon Adsorption employs carbon steel or
concrete vessels and tanks. Corrosion is a big problem but can be
minimized by use of coal tar epoxy paints. Pumps and piping for
slurry transport are required. A refractory lined multiple hearth
furnace is usually required. A venturi scrubber or fabric filter is
usually required for furnace particulate control.
2.3 Feed Stream Requirements
Temperature: No specific requirement, hot wastewater feed may lead
to gasing in bed and decreased adsorption of organics.
Pressure: No specific requirements.
Composition: Inorganic composition not generally important, except
that acidity or alkalinity can influence adsorption
efficiency of certain organics (e.g., phenolics,
carboxylic acids). Suspended solids (inorganic or bio-
floe) tend to clog beds and should be largely removed
upstream. Periodic bed backwashing is usually required.
2.4 Operating Parameters
2.4.1 Granular Carbon Beds
Adsorption^6'7)
Flow Rate - liters/min/m2 (gal/min/ft2): 0.7-3.5 (2-10)
Flow Rate - m/min (ft/min): 0.07-0.4 (0.25-1.34)
Bed Depth - m (ft): 4.5-11.6 (15-38)
Contact Time - min: 15-38 or higher
Contact Time - m3/103l/min (ft3/gal/min): 15-37 (2-5)
or higher
Carbon Capacity - kg COD/kg carbon: 0.2-1.2
Bed Expansion Allowance: 10%-50%
f
Backwashing^ '
Flow Rate - 1/min/m2 (gal/min/ft2): 4.2-7.0 (12-20)
Total Flow Requirement: should not exceed 5% of average
plant flow
E-104
-------
Granular Carbon Regeneration^7'11^
Furnace Temperature: increasing temperature from top to
bottom of furnace: 366°K-1255°K (200°F-1800°F)
Oxygen: maintained at less than 1%
Steam: approximately 1 kg/kg carbon
Residence Time: drying - 15 minutes
pyrolysis - 5 minutes
gasification - 10 minutes
2.4.2 Powdered Carbon Addition to Activated Sludge Treatment
SystemsO.lO)
Steady State Carbon Level in system recovered for maximum
efficiency: 200-2000 mg/1 or higher (depends upon the nature
and strength of wastewater to be treated)
Continuous addition required to maintain needed carbon level:
10-20 mg/1 (depends on the wastewater and sludge washing/
recycle ratio)
2.5 Process Efficiency and Reliability - Activated carbon preferentially
adsorbs high molecular weight and less polar organic compounds.
Table E-28 shows the relative adsorbability of several representative
compounds as a function of compound type and molecular weight. In
actual wastewater applications, a wide range of substances would be
encountered and the actual carbon performance would have to be deter-
mined by laboratory and pilot testing.
In commercial refinery applications, from 59%-83% COD removal has
been obtained with granular carbon systems used without prior bio-
logical treatment. A petrochemical pilot plant employing granular
activated carbon treatment of activated sludge effluent has achieved
(4)
50%-68% COD removal, 53%-80% SOC* removal, and 50%-65% BOD removal* '.
Studies of an activated carbon system for treatment of a coke plant
effluent after clarification and filtration reported 80% COD removal,
91% TOC removal and 99%*• phenol removal1 '.
*SOC = Soluble organic carbon
E-105
-------
TABLE E-28. AMENABILITY OF TYPICAL ORGANIC COMPOUNDS TO ACTIVATED
CARBON ADSORPTION*U3)
Compound
Adsorbabilityt
(grams compound/grams carbon)
Ethanol
2-Ethyl Butanol
Acetaldehyde
Benzaldehyde
Di-N-Butylamine
Monoethanolamine
2-Methyl 5-Ethyl Pyridine
Benzene
Hydroquinone
Ethyl Acetate
Butyl Acetate
Isopropyl Ether
Ethylene Glycol
Tetraethylene Glycol
Acetone
Acetophenone
Formic Acid
Valeric Acid
Benzole Acid
0.02
0.170
0.022
0.188
0.174
0.150
0.179
0.080
0.167
0.100
0.193
0.162
0.0136
0.116
0.054
0.194
0.047
0.159
0.183
*Westvaco Nuchar WV-G (12 x 40 mesh, coal based) carbon
t5g carbon added to 1 liter of solution containing 100 mg/1 of
compound
E-1Q6
-------
In pilot plant granular carbon adsorption tests of biologically
treated API separator effluent, 57% BOD removal, 73% COD removal,
and 77% TOC removal were achieved^. Removal of the bulk of Cr,
Cu, Fe, and Zn were also observed. Carbon adsorption does not
ordinarily remove sulfide, ammonia, or cyanide.
A pilot powdered activated carbon/activated sludge system treating
refinery wastewater is reported to achieve 50% suspended solids
reduction, 20%-36% COD reduction, and 51%-76% BOD reduction when
compared to activated sludge treatment alone^10^. Similar results
are reported for powdered carbon tests of several other activated
sludge systems at refineries^.
Available information indicates that both granular and powdered
carbon systems are reasonably reliable. For effective performance,
the systems require routine monitoring of pressure drop, effluent
quality, and carbon activity.
2.6 Raw Material Requirements
Properties of Fresh Activated Carbons - Carbons for wastewater treat-
ment applications are usually made from coals. Some properties of
commercially available granular carbons are shown in Table E-29^ .
M 91
Makeup Requirementsv ' ' - Typical losses during thermal regenera-
tion are 5%-10%, Additional losses result from attrition in the
handling and transport of carbon and from purposeful withdrawal to
minimize ash buildup and to maintain adsorption activity. Exact
makeup requirements will depend heavily upon the nature and strength
of the wastewater treated, since this determines the frequency con-
ditions of regeneration (see Section 2.4.1).
In the case of powdered carbon, dosage depends upon the nature and
strength of the wastewater (see Section 2.4.2).
2.7 Utility Requirements
Electricity Needed for pumping, carbon reactivation, and control
instrumentation. Pumping energy tends to be design specific, but
E-107
-------
TABLE E-29. TYPICAL PROPERTIES OF SEVERAL COMMERCIALLY AVAILABLE
GRANULAR CARBONS*
Parameter
Physical Properties
2
Surface area, m /gm
Apparent density, gm/cc
Density, backwashed and
drained, kg/m3 (Ib/cu ft)
Real density, gm/cc
Particle density, gm/cc
Effective size, mm
Uniformity coefficient
Pore volume, cc/gm
Mean particle diameter, mm
SPECIFICATIONS
Sieve size (U.S. std.
series)
Larger than 'No. 8
(max. %)
Larger than No. 12
(max. %}
Smaller than No. 30
(max. %)
Smaller than No. 40
(max. %}
Iodine No.§
Abrasion No:., minimum
Ash^5^
Moisture as packed
(max. %)
ICI
America
Hydrodarco
3000
600-650
0.43
355 (22)
2.0
1.4-1.5
0.8-0.9
1.7
0.95
1.6
8
t
5
t
650
*
*
*
Calgon
Filtrasorb
300
(8 x 30)
950-1050
0.48
419 (26)
2.1
1.3-1.4
0.8-0.9
1 .9 or less
0.85
1.5-1.7
8
t
5
t
900
70
8
2
Westvaco
Nuchar
WV-L
(8 x 30)
1000
0.48
419 (26)
2.1
1.4
0.85-1.05
1 .8 or less
0.85
1.5-1.7
8
t
5
t
950
70
7.5
2
Witco
517
(12x30)
1050
0.48
484 (30)
2.1
0.92
0.89
1.44
0.60
1.2
t
5
5
t
1000
85
0.5
1
+Not applicable to this size carbon
+NO available data from the manufacturer
5An index of the amount of pore area in the small molecule size range
E-108
-------
would be in the range of 0.04 kwh/1000 £ (0.15 kwh/1000 gals)(12).
Multiple hearth granular carbon regeneration electrical
ranges from 0.02-0.09 kwh/kg (0.01 to 0.04 kwh/lb) JSb
Steam for Regeneration^1 >: About 1 kg/kg carbon
Fuel(11): 3300-4400 kcal/kg (6000-8000 Btu/lb) carbon
3.0 Process Advantages
• Commercially proven in a variety of applications.
• Can remove a wide variety of organic compounds to low levels in water,
including refractory or non-biodegradable substances.
• Adsorption not generally affected by changes in loading, temperature,
or the presence of toxic substances (e.g., Cr, CN~).
• Adsorbed organics are largely destroyed during thermal reactivation of
granular carbon and do not become a sludge disposal problem as in some
of the other organics removal technologies.
c Can be used in conjunction with copper addition to remove cyanide via
catalytic oxidation^5'!4'
t Powdered carbon improves the settleability of solids in activated
sludge systems in addition to enhancing organics removal.
• Powdered carbons and, to a lesser extent, granular carbons can provide
greater removal efficiencies than calculated from simple adsorption
tests due to biological activity promoted on the carbon surfaces.
• Potential for product recovery (e.g., phenols via caustic extraction).
4.0 Process Limitations
• Process is relatively expensive compared to biological oxidation on a
weight COD or BOD removal basis. Carbon systems are usually only
economical for tertiary treatment applications or where the wastewater
is not amenable to biological treatment.
• When thermal reactivation is practiced, potentially valuable organics
are not recovered (e.g., phenols).
• Offgas from carbon regeneration often contains particulate matter,
carbon monoxide, and unburned hydrocarbons which must be removed prior
to atmospheric discharge.
• Trace constituents such as ammonia, cyanide, sulfide, and certain
trace elements are not generally removed by activated carbon.
t Sulfide levels may increase during activated carbon treatment due to
biological activity. This may lead to odor or effluent problems.
E-1Q9
-------
5.0 Process Economics
Capital and operating costs of granular carbon systems depend upon the
specific design and the nature and volume of the wastewater treated. One
estimate of 1976 capital costs are as follows^ ':
Adsorption Equipment
Flow Cost ($)
4 x 105 z/day (1 x 105 gal/day) 180,000
4 x 106 A/day (106 gal/day) 550,000
Regeneration Equipment
Carbon Usage Rate Cost ($)
910 kg/day (2000 Ibs/day) 270,000
8200 kg/day (18000 Ibs/day) 1,000,000
1976 operating costs have been estimated at about $0.68 per 1000 liters
(400 gals) for every 1000 mg/1 of COD removed^12'.
6.0 Input Streams
6.1 Feed Water (Stream 1) - See Section 2.5 and Tables E-30, E-31, and
E-32.
6.2 Regenerated/Makeup Activated Carbon (Stream 5) - See Table E-29 for
typical characteristics of fresh carbon. Regeneration of carbons
tends to cause an increase in the average "pore" size and thus reduce
carbon affinity for small molecules (e.g., phenol). However, lignite-
derived carbons do not undergo as much pore size enlargement as
bituminous-derived carbons (see Section 2.6). Regenerated carbon
loading capacity for organics tends to be lower than fresh carbon
and ash tends to build up since some of the original carbon is burned
during each regeneration.
6.3 Backwash Feed (Stream 3) - Typically treated effluent is used for
backwashing.
E-11Q
-------
TABLE E-30.
"""
Parameter
Total Suspended Solids
Total Dissolved Solids
Total Organic Carbon
Soluble Organic Carbon
Chemical Oxygen Demand
Biochemical Oxygen Demand
Phenols
Cyanide
Ammonia
Thiocyanate
PH
Feed
Wastewater
(mg/A)t
<5
--
1750
1750
6340
--
1950
0.01
4000
700
8.0
— —
Treated
Wastewater
(mg/i)t
<5
—
156
156
1260
--
<0.1
0.01
4000
<700
8.0
Average
Spent Carbon
Loading (%)
30
25
*Wastewater has received clarification/filtration treatment
fExcept pH
6.4 Quench Water Makeup (Stream 10) - Typically treated effluent would
be employed.
6.5 Wash Reactivated Carbon (Stream 11) - See Section 6.2. Washing
removes some of the loose or brittle material.
6.6 Wash Water (Stream 12) - Typically, treated effluent would be
employed.
6.7 Scrubber Feed Water (Stream 13) - No operating data available,1 this
stream would likely be treated effluent.
E-111
-------
TABLE E-31. COMPARISON OF GRANULAR ACTIVATED CARBON ADSORPTION
AND BIOLOGICAL TREATMENT OF REFINERY WASTEWATERS*(4)
Constituent/
Parameter"!"
BOD
COD
TOC
Oil and
Grease
Phenols
Cr
Cu
Fe
Pb
Zn
S=
NH3
CN"
API
Separator
Effluent
97
234
56
29
3.4
2.2
0.5
2.2
0.2
0.7
33
28
0.25
Carbon
Treated
Effluent
48
103
14
10
0.004
0.2
0.03
0.3
0.2
0.08
39
28
0.2
Biologically
Treated
Effluent
7
98
30
10
0.01
0.9
0.1
0.2
0.2
0.4
0.2
27
0.2
Biological/
Carbon
Treated
Effluent
3
26
7
7
0.001
0.02
0.05
0.2
0.2
0.15
0.2
27
0.2
*Pilot scale operation
"'"An units are mg/1
E-112
-------
TABLE E-32. PERFORMANCE OF POWDERED ACTIVATED CARBON ADDITION TO ACTIVATED
SLUDGE SYSTEM*(
Trial 1
Control
Carbon
Added
Trial 2
Control
Carbon
Added
Trial 3
Control
Carbon
Added
Carbon
Dose
(mg/1)
0
24
0
19
0
9
Flow Rate
1/min
(gal/min)
2370 (630)
2370 (630)
2460 (650)
2490 (660)
3180 (840)
3030 (800)
Total Suspended
Solids (mg/1)
Influent Effluent
-f 115
50
164
72
79
42
% TSS
Reduc-
tion
56
00
55
„
49
COD (mg/1)
Influ- Efflu-
ent ent
459 170
457 135
343 266
444 183
367 166
379 112
% COD
Reduc-
tion
20
30
..
36
BOD (mg/1)
Influ- Efflu-
ent ent
152 15
213 13
152 30
227 14
188 12
207 3
% BOD
Reduc-
tion
2
..
52
..
76
I
CO
*Treating API Separator wastewater; steady state aeration system contained 450 mg/1 carbon and steady
state recycle system contained 1000 mg/1 carbon
-- Indicates data not available
-------
7.0 Process Discharge Streams
7.1 Treated Effluent (Stream 2) - See Tables E-30, E-31, and E-32, and
Section 2.5.
7.2 Back Wash Effluent (Stream 3) - No data available. This effluent
would normally be returned to upstream suspended solid removal
operations.
7.3 Spent Carbon (Stream 6) - Limited actual data are available.
Table E-28 shows the capacity of an example carbon for various
compounds. The exact loading and nature of adsorbed organics
depends upon the wastewater being treated (see Section 6.2).
7.4 Dewatering Effluent (Stream 7) - No data available. This stream
would normally be returned to upstream treatment units.
7.5 Washed Reactivated Carbon (Stream 11) - No actual data available.
See Section 6.2.
7.6 Wash Effluent (Stream 13) - No data available. This stream would
normally be returned to upstream treatment units.
7.7 Raw Regeneration Offgas (Stream 14) - No data available. This gas
will contain organics, carbon monoxide, and entrained particulates.
7.8 Cleaned Regeneration Offgas (Stream 16) - No data available.
7.9 Scrubber Effluent Water (Stream 17) - No data available. This
stream would normally be sent to upstream dissolved solids removal
units.
8.0 Data Gaps and Limitations
Data gaps and limitations relate primarily to the properties of various
processes streams associated with activated carbon systems. Carbon
adsorption has never been employed for organics removal from coal gasi-
fication wastewaters and hence no operating data exists. Data from coke
plant and refinery applications are limited and do not necessarily repre-
sent a spectrum of organic substances similar to that likely to be
encountered in coal gasification. Also, for existing carbon adsorption
systems, essentially no information is available for regeneration
E-114
-------
offgases and backwash waters. Finally, little is known about the nature
of organics which remain in the treated effluent from carbon adsorption
systems (e.g., biodegradability, toxicity).
9.0 Related Programs
No programs are known to be underway or planned which are specifically
aimed at the environmental assessment of carbon adsorption in coal gasifi-
cation applications. However, as part of the ongoing work with the
Synthane PEDU at the Pittsburgh Energy Research Center, the treatability
(15}
of Synthane wastewaters is being investigated^ '. One aspect of this
work involves bench scale adsorption tests of biologically treated efflu-
ent using Synthane char (physically and chemically similar to commercial
activated carbons).
E-115
-------
REFERENCES
1. U.S. EPA, Process Design Manual for Carbon Adsorption, EPA 625/1-71-002a,
October 1973.
2. Kerr, R. S., Pilot Plant Activated Carbon Treatment of Petroleum Refinery
Wastewater, Open Forum on Management of Petroleum Refinery Wastewaters,
January 26-29, 1976, Tulsa, Oklahoma.
3. Environmental Science and Technology, Environmental Control Issue, Vol. 10,
No. 11, October 1977, p. 49.
4. Ford, D. L., Current State of the Art of Activated Carbon Treatment, Open
Forum on Management of Petroleum Refinery Wastewaters, sponsored by EPA,
January 26-29, 1976, Tulsa, Oklahoma.
5. Van Stone, G. R., Treatment of Coke Plant Waste Effluent, Iron and Steel
Engineer, April 1972, pp 63-66.
6. American Petroleum Institute, Manual on Disposal of Refinery Wastes -
Volume on Liquid Wastes, Chapter 10, Washington, D.C., 1973.
7. Rizzo, J. L. and Shepherd, A. R., Treating Industrial Wastewater with
Activated Carbon, Chemical Engineering, January 3, 1977, pp 95-100.
8. Activated Carbon Heads for Sell-out Year, Chemical and Engineering News,
July 22, 1974.
9. DeJohn, P. B. and Adams, A. D., Activated Carbon Improves Wastewater
Treatment, Hydrocarbon Processing, October 1975, pp 104-111.
10. Rizzo, J. A., Case History: Use of Powdered Activated Carbon in an
Activated Sludge System, Open Forum on Management of Petroleum Refinery
Wastewaters, January 26-29, 1976, Tulsa, Oklahoma.
11. Loven, A. W., Perspectives on Carbon Regeneration, Chemical Engineering
Progress, Vol. 69, No. 11, November 1973, pp 56-62.
12. Water Purification Associates, Water Conservation and Pollution Control
in Coal Conversion Processes, ongoing work under EPA Contract No.
68-03-2207.
13. Giusti, D. M., et al, Activated Carbon Adsorption of Petrochemicals,
Journal of Water Pollution Control Federation, Vol. 46, No. 5, May 1974.
14. Huff, J. E. and Bigger, J. M., Cyanide Removal from Petroleum Refinery
Wastewater Using Powdered Activated Carbon, Illinois Institute for Envi-
ronmental Quality, Document No. 77/08, June 1977.
15. Johnson, G. E., et al, Treatability Studies of Condensed Water from
Synthane Coal Gasification, PERC/RI-77/13, 1978.
E-116
-------
Sludge Treatment Module
Gravity Thickening
Centrifugation
Vacuum Filtration
Drying Beds
Emulsion Breaking
E-117
-------
GRAVITY THICKENING PROCESS
1.0 General Information
1.1 Operating Principle - Removal of excess water from sludges to
reduce their volume and to increase solids concentration, using
gravity settling.
1.2 Development Status - Commercially available. Numerous units are in
operation throughout the world for municipal and industrial sludge
thickening.
1.3 Licensor/Developer - Gravity thickener systems and equipment are
offered by many suppliers. Listings of the suppliers are presented
in technical and trade journals (e.g.. Reference 1).
1.4 Commercial Applications - Gravity thickening is in widespread use
in municipal and industrial waste treatment plants. At the SASOL
gasification plant in South Africa, gravity thickening is used to
(2)
concentrate sludge resulting from wastewater treatment .
2.0 Process Information
2.1 Flow Diagram (see Figure E-16) - The thickening is carried out
(usually in a circular tank) on a batch or continuous basis. In
circular tank designs, the influence sludge is distributed at the
center of the tank, the clarified liquid is collected at the sur-
face near the periphery and the concentrated sludge is withdrawn
at the bottom. The tank is usually equipped with a gently rotating
agitator with a sludge scraping mechanism to increase thickening
efficiency and to divert the settled sludge to the sludge hopper
at the bottom for removal.
E-118
-------
DRIVE UNIT
,WEIR
EFFLUENT
LINE
TO PUMP
EFFLUENT CHANNEL
KAKE ARM
SLUDGE
DISCHARGE
LINE
INFLUENT.
LINE
BLADES AND SQUEEGE
CENTER SCRAPERS
INSIDE TANK DIAMETER
Figure E-16. Schematic of Gravity Thickener and Section of Tank
(3)
-------
2.2 Equipment - Thickening tank and associated mechanical devices (mixer,
drive unit, sludge influent and withdrawal structures, pumps, etc.);
the surface area (hence the diameter of the tank is dictated by the
design surface area loading, see Section 2.4); tank depth is gener-
ally in the 3.0-7.5 m (10-25 ft) range^3'.
2.3 Feed Stream Requirements - Good settleability and relatively high
solids content are primary requirements for effective thickening.
The solids in the sludge would be sufficiently compressible and
(3)
porous to permit escape of waterv
2.4 Operating Parameters
• Surface area loading - Determines tank surface area. Varies with
the waste and solid concentrations and underflow concentrations
desired. Ranges from 118-1301 kg/m2 (day) (5-55 Ib/ft? (day))
have been reported^).
2.5 Process Efficiency and Reliability - The levels of sludge concentra-
tion achieved depends on the characteristics of the raw sludge and
the thickener design. For waste activated sludges with a solids load-
ing of 142-237 kg/m2 (day) (6-10 lb/ft2 (day)), waste underflow con-
centrations of 5%-8% solids are typically achieved^ .
2.6 Raw Materials Requirements - When sludge requires preconditioning to
improve settleability, chemical coagulants (i.e., ferric chloride,
aluminum chlorhydrate) may be required.
2.7 Utility Requirements
• Electricity (for control drive mechanism, pumps, raking mechanism,
etc.) - Requirements are design-specific.
3.0 Process Advantages^3'6'
• Widely used commercial process for which extensive operating experience
is available.
• Little maintenance required.
• Little or no raw materials required except for preconditioning
chemicals.
E-120
-------
4.0 Process Limitations
(3,4)
• Not all sludges can be thickened efficiently by gravity thickening
In certain cases, preconditioning may be required (see Section 2.6).
• Laboratory/bench-scale tests may be required to define sludge
thickening characteristics and to generate basis for thickener
design.
• For highly biodegradable sludges, long solids retention time may lead
to the production of odor and floating sludge.
5.0 Process Economics
For many sludges, thickening is considered to be the most economical way
of effecting major sludge volume reduction'3^. The capital cost of sludge
thickeners has been estimated as^ ':
Capital Cost ($) = (18.8 + 9.1/exp [SAT/13,300]) • (SAT)
where SAT = surface area of thickener (ft ).
6.0 Input Streams
6.1 Influent Sludge (Stream 1) - Typical influent streams include pri-
mary and secondary sludges and chemical sludges (e.g., alum, line);
solids concentrations of these sludges vary from less than I0/ to as
much as
7.0 Discharge Streams
7.1 Clarified Effluent (Stream 2) - Consists of wastewater containing
some suspended solids.
7.2 Thickener Underflow (Stream 3) - Consists of the thickened sludge;
solids concentration depends on thickener loading and influent sludge
characteristics.
8.0 Data Gaps and Limitations
No data available on the thickeners used in the SASOL plant for handling
sludges originating from coal gasification and associated operations.
9.0 Related Programs
Not known.
E-121
-------
REFERENCES
1. Environmental Control Issue, Control Equipment, Environmental Science and
Technology, October 1977.
2. Information provided by South African Coal, Oil,and Gas Corp., Ltd. to EPA's
Industrial Environmental Research Laboratory (Research Triangle Park),
November 1974-
3. Weber, W. M., Jr., Physiochemical Processes for Water Quality Control,
Wiley-Interscience Publishers, Inc., New York, p. 547-558.
4. Azad. H. S., Industrial Wastewater Management Handbook, McGraw-Hill Book
Co., New York, 1976, pp 3-30 to 3-32.
5. Newton, D., Thickening by Gravity and Mechanical Means, Sludge Concentra-
tion, Filtration and Incineration, University of Michigan School of Public
Health, Continued Education Series, 113, 1964, p. 4.
6. Reid, G. W., and L , E. Streebin, Evaluation of Waste Waters from Petroleum
and Coal Processing, Oklahoma University, Norman, Oklahoma, PB-214-610,
December 1972, 218 pp.
7. Smith, R., Cost of Conventional and Advanced Treatment of Wastewater,
Journal of Water Pollution Control Federation, Vol. 40, No. 9, p. 1546-
1574, September 1968.
E-122
-------
CENTRIFUGATION
1.0 General Information
1.1 Operating Principle - Physical liquids-solids separation by means
of sedimentation and centrifugal force.
1.2 Developmental Status - Commercially available. Numerous units are
in operation throughout the world in industrial applications and
for municipal and industrial waste treatment, line!uding petroleum
refinery sludges.
1.3 Licensor/Developer - Many centrifuge treatment systems and equip-
ment are offered by numerous suppliers. A complete listing of these
systems and their applications is available in the literature
(e.g., Ref. 1).
1.4 Commercial Applications - Method is in widespread use in municipal
(2}
wastewater treatment . Has also been used in treatment of
refinery wastes, including oily sludges such as storage tank and
7o\
gravity separator bottoms . Sometimes used to dewater sludges
following treatment by coagulation-flocculation or emulsion-breaking
techniques. No known applications to coal gasification wastes.
2.0 Process Information
2.1 Flow Diagram - See Figure E-17
• Process Description - Influent wastewater (Stream 1) is fed
through a stationary feed pipe into the centrifuge from which
it is thrown out through feed parts into the conveyor hub.
The solids (Stream 3) are settled out against the outer 'bowl1
wall by centrifugal force, and are continuously conveyed by a
screw moving at a speed slightly different than the bowl to the
end of the centrifuge and discharged. A pool volume is main-
tained in the equipment. Liquid effluent (Stream 2) discharges
out of adjustable effluent ports or weirs after passing the
length of the pool under centrifugal force.
E-123
-------
LIQUID POOL
CONVEYOR / SCROLL
HELIX ROTATING CONVEYOR
BOWL
Legend:
1. Influent Wastewater/Sludge
2. Liquids Discharge (centrate)
3. Dewatered Solids
Figure E-17. Schematic of Continuous Solid Bowl Centrifuge
(4)
E-124
-------
2.2 Equipment
t Centrifuge equipment - solid bowl, basket, nozzle, or disk types.
• Pumps
2.3 Feed Stream Requirements^4' - The dewaterability of sludges by cen-
trifugation depends on factors such as the concentration, size, shape,
and surface characteristics of the sludge particles, the extent of
aggregation, the structural characteristics of the particles, and the
viscosity, ionic strength, and pH of the suspending water. Perfor-
mance parameters which reflect the combined influence of these vari-
ables and which are calculated from measurable variables for
determination of optimum operating conditions for a given centrifuga-
tion system are the specific resistance and the coefficient of com-
pressibility of the waste. Pre-treatment of the sludge by coagulation-
flocculation, emulsion breaking and thickening techniques may
facilitate centrifugation operations.
2.4 Operating Parameters^ ' ' - See Table E-33 for listing of operating
parameters and their effect on percent solids recovery and cake
solids concentration.
2.5 Process Efficiency and Reliability^4' - Efficiency depends on the
type and design of the system used, and on the nature of the sludge
treated. Tables E-34 and E-35 present the results of centrifugation
of various industrial and municipal sludges. Centrifugation
processes have been widely used and proven highly reliable for
treatment of a range of sludges.
2.6 Raw Materials Requirements
• Sludge conditioning chemicals (e.g., chemical flocculants) - May
be required to enhance removal of fine, difficult-to-remove
solids. See Table E-35.
2.7 Utility Requirements
• Electricity - used for driving pumps and central screw feed
mechanism. Requirements vary with the specific design and
removal efficiency desired.
E-125
-------
TABLE E-33. EFFECT OF AN INCREASE IN VARIOUS CENTRIFUGATION
VARIABLES ON SOLIDS CAPTURE AND DEWATERING(4)
Variable Parameters
Effect of Increase in Variable On
Solids Recovery
Cake Solids Concentration
Machine Parameters
Bowl Speed
Pool Depth
Scrolling Speed
Process Parameters
Feed Rate
Feed Concentration
Temperature
Increase
Increase
Decrease
Decrease
Decrease
Increase
Increase
Decrease
Decrease
Increase
Increase
Increase
3.0 Process Advantages
(4,7)
t
•
•
•
Simple to operate; units are compact and require little space.
Totally enclosed to minimize odor dispersion.
Minimal to nil raw materials requirements.
Suitable for treatment of a wide variety of sludges with differing
physical and chemical properties.
Minimal supervision requirements.
Operation can be adjusted to permit concentration of relatively vola-
tile material in sludge in the centrate and concentration of nonvola-
tile solids in the dewatered solids, thus permitting some selectivity
in waste segregation.
E-126
-------
TABLE E-34. RESULTS OF CENTRIFUGATION OF SLUDGES^
Type of Sludge
Raw Primary
Digested Primary
Activated
Raw Primary and
Activated
Digested Raw and
Activated
Pulp and Paper
Wastes^
Box Board
Hard Board
White Water
Barker
Kraft
Specialty
Paper
Softening
Sludge
-.
Cake
Concentration
(X Solids)
— .
28-35
25-35
6-10*
18-24
18-24
22-33
26-28
21-30
32-40
36-43
15
53-57
===============
_^ojJd^_R^c^ve£v_[%]__
Without With
Chemicals Chemicals
™ " —
85-90 >95
80-90 >95
50-80 >95
50-70 >95
86-94
85-95
78-94
90-93
78-89
90
79-93
Cost of
Chemicals,
$/tonne
($/ton) of
Dry Solids
3.3-8.8 (3-8)
3.3-8.8 (3-8)
8.8-22 (8-20)t
6.6-22 (6-20)
11-22 (10-20)
-
-
-
-
-
-
—
*Without chemicals.
''"Cost of chemical conditioning to improve upon the 6%-1Q% cake.
*For pump and paper sludges polymers could be used to increase capture
to 95%-99% at a cost of $3 to $8 per ton.
E-127
-------
TABLE E-35. CENTRIFUGE PERFORMANCE AND OPERATING COSTS
FOR MUNICIPAL SLUDGE TREATMENT^)*
Plant Flow
MLD (MGD)
Process
Number of
Units
Machine Size,
cm (in.)
Performance,
Percent Solids
Feed Solids
Cake Solids
Recovery
Chemical Used
Dosage,
$/tonne ($/ton)
Cost, $/kg
Operating Cost,
$/tonne ($/ton)
Maintenance
Operating Labor
Amortization'*'
Plant A
68 (18)
Primary Plus
Trickling
Filter with
Anaerobic
Digestion
One
61 x 96
(24 x 38)
4-6
18-24
95-97
3.3-6.6 (3-6)
2.2 (2)
|
2.79 (2.53)
2.98 (2.71)
1.43 (1.30)
Plant B
19 (5)
Primary
Treatment
with
Anaerobic
Digestion
One
61 x 96
(24 x 38)
7.5-8.5
30-35
65-75
1.93 (1.75)
'
1.04 (0.94)
14.55 (13.20)
Plant C
10.2 (2.7)
Primary
Treatment
Plus Acti-
vated with
Anaerobic
Digestion
One
61 x 152
(24 x 60)
4.5-5
20-25
90-95
8.8 (8)
1.76 (1.60)
2.90 (2.63)
7.91 (7.17)
14.1 (12.80)
Plant D
30.2 (8.0)
Primary Plus
Trickling
Filter
Two
61 x 96
(24 x 38)
8
30
65-75
-
-
1.92 (1.74)
1.10 (1.00)
3.56 (3.23)
*A11 costs based on 1973 dollars.
'''Amortization based on 6% interest cost and amortized 25 years
7.823% of the capital cost as yearly cost.
(Continued)
equals
E-128
-------
TABLE E-35. Continued
j Plant A
Operating Cost,
I/tonne ($/ton)
(Continued)
Power
Chemicals
Total Cost
Ultimate
Disposal
Years of
Service
Operating
Schedule
Tonnes (Tons)
Dry Sludge
Solids
Dewatered
0.77 (0.70)
10.2 (9.30)
18.17 (16.54)
Landfill
4-5
24 hr/day
54.9 (60.5)
|
I
Plant B
0.43 (0.39)
None
17.95 (16.28)
Fertilizer/
Compost
7
9 hr/week
7.8 (8.6)
:
i
Plant C
0.54 (0.49)
14.1 (12.80)
39.56 (30.74)
Landfill/
Fertilizer
5
21 hr/week
14.3 (15.8)
Plant D
0.39 (0.35)
None
6.97 (6.32)
Incineration/
Landfill
13.2
85 hr/week
31.8 (35.0)
*A11 costs based on 1973 dollars.
"^Amortization based on 6% interest cost and amortized 25 years equals
7.823% of the capital cost as yearly cost.
E-129
-------
4.0 Process Limitations
(7)
• Scrolling of solids up the beach of the centrifuge must be carefully
regulated by controlling the operating speeds, or high shearing forces
caused by fluid drag from escaping liquid and by agitation of the
scroll may carry solids back into the liquid pool.
t Relatively high maintenance requirements.
t Dewatered sludge is generated which requires disposal by incineration,
landfill, or other method.
5.0 Process Economics - No capital cost data available; see Tables E-34 and
E-35 for operating and chemical cost data.
6.0 Input Streams
6.1 Influent Wastewater/Sludge (Stream 1) - Sludge characteristics vary
depending on the source. Will contain suspended and dissolved solids,
oils, emulsions, heavy metals, etc.
7.0 Intermediate Streams
7.1 Liquids Discharge (Centrate) (Stream 2) - Will vary, depending upon
composition of Stream 1. Will contain unreacted, excess coagulation
chemicals.
8.0 Discharge Streams
8.1 Dewatered Solids (Stream 3) - See Tables E-34 and E-35.
9.0 Data Gaps and Limitations
Centrifugation has not been tested on sludges generated in coal gasifica-
tion operations to determine optimum operating conditions.
10.0 Related Programs
None known.
E-130
-------
REFERENCES
1. Smith, J. C., Centrifugation Equipment Applications, Ind. Enq. Chem. 53 (6),
439 (1961). —
2. White, W. F., Fifteen Years of Experience Dewatering Municipal Wastes with
Continuous Centrifuges, AIChE Symposium Series, No/129, Volume 69, 1973,
p. 211-216.
3. Cavanaugh, E. C., J. D. Colley, et al., Environmental Problem Definition
for Petroleum Refineries, SNG plants and LNG Plants, Radian Corporation,
Austin, Texas, EPA-600/2-75-068, PB-252-245, 1975, p. 318.
4. Weber, Jr., W. J., Physiochemical Processes for Water Quality Control,
Wiley-Intersciences, New York, 1972, p. 572-575.
5. Eckenfelder, Jr., W. W., Industrial Water Pollution Control, McGraw-Hill
Book Company, New York, 1966, p. 250.
6. Albertson, 0. E., and E. J. Guidi, Jr., Advances in the Centrifugal
Dewatering of Sludges, Water Sew. Works, 114, R.N., R-113, (1967).
7. Azad, H. S., Industrial Wastewater Management Handbook, McGraw-Hill
Book Company, New York, 1976, p. 3-35.
E-131
-------
VACUUM FILTRATION
1.0 General Information
1.1 Operating Principle - Use of an applied vacuum to dewater a slurry
or sludge by means of a rotary filter drum containing porous medium
which retains the solid but allows the liquid to pass. Media used
include cloth made of natural or synthetic filters, steel mesh, and
tightly wound coil springs. Filter drum may be precoated with
diatomaceous earth to facilitate breaking of emulsions, removal of
suspended solids and traces of oil.
1.2 Development Status - Commercially available. Numerous units in
operation throughout the world for municipal and industrial waste
treatment, such as petroleum refinery sludges. Vacuum filtration is
the most commonly used mechanical sludge dewatering method in the
u.s.W.
1.3 Licensor/Developer - Many vacuum filtration treatment systems and
equipment are offered by numerous suppliers; a complete listing of
these systems and their applications is available in the literature^ '•
1.4 Commercial Applications - Numerous applications to municipal and
industrial wastewaters. Commonly used in treatment of boiler treat-
ment and blowdown and chemical or biological treatment sludges at
fo\
petroleum refineriesv '. Often used following treatment of sludges
by coagulation-flocculation or emulsion-breaking techniques. No
known applications to coal gasification wastes.
2.0 Process Information^ '
2.1 Flow Diagram - See Figure E-18.
t Process Description - Influent sludge (Stream 1) is fed to a sludge.
tank containing a rotating drum. As the drum passes through the
sludge, solids are retained on the drum surface under an applied
E-132
-------
WATER SPRAY
CO
VACUUM CONTROL
REGULATORS
CONTINUOUS ROTARY
FILTER DRUM
LEGEND:
1. SLUDGE FEED FROM MIX TANK
2. DEWATERED SLUDGE
3. SLUDGE OVERFLOW TO SLUDGE WELL
4. FILTRATE
MUFFLER
VACUUM
PUMPS
Figure E-18. Schematic of Vacuum Filtration Process
-------
vacuum; a cake of solids is built up, and filtrate (Stream 4)
is removed by filtration through the deposited solids and the
filter medium., As the drum emerges from the sludge tank, the
deposited cake is further dried by liquid transfer to air drawn
through the cake by the applied vacuum. At the end of the cycle,
a knife edge scrapes the filter cake from the rotary filter drum
to a conveyor for removal (Stream 2). Overflow sludge (Stream 3)
is sent to a sludge well for recycle. The rotary filter drum is
usually washed with water sprays at the end of a cycle before it
is re-immersed in the sludge tank.
2.2 Equipment
• Filtration device (e.g., rotary drum, scroll-discharge, tlltlng-
pan, disk, and batch leaf. Variations in the rotary drum
include multicompartment, single compartment, belt, precoat,
Corrco, hopper dewater, and top feed units.)
• Sludge tank.
• Water spray apparatus.
• Pumps (filtrate, vacuum, water-wash).
• Filtrate receivers.
t Vacuum control regulators.
• Miscellaneous equipment (pipes, mufflers, etc.).
2.3 Feed Stream Requirements^ '
t The dewaterability of sludges by vacuum filtration depends on fac-
tors such as the concentration, size, shape, and surface character-
istics of the sludge particles, the extent of aggregation, the
structural characteristics of the particles, and the viscosity,
ionic strength and pH of the suspending water. Performance param-
eters which reflect the combined influence of these variables and
which are calculated from measurable variables for determination
of optimum operating conditions for a given vacuum filtration sys-
tem are the specific resistance (5) and the coefficient of com-
pressibility (s) of the waste.
2.4 Operating Parameters^5'6^
Operating parameters and design consideration include: sludge feed
concentration, sludge viscosity, filtrate viscosity, operating
vacuum, type and porosity of filter media, degree of sludge thicken-
ing preceding filtration, thickening chemical, drum submergence time,
and drum speed. See Table E-36.
E-134
-------
TABLE E-36. TYPICAL SEWAGE SLUDGE FILTRATION
CHARACTERISTICS AND RATES(7)
Sludge Type
Primary Sludge
Raw
Digested
Digested - Elutriated
Primary - Trickling
Filter
Raw
Digested
Digested - Elutriated
Primary - Activated
Sludge
Raw
Digested
Digested - Elutriated
Activated Sludge -
Concentrated
Feed
Solids,
%
8
8
8
7
8
8
5
6
6
3
Filtration Rate,
Dry kg/hr-m3
(Dry lbs/hr-ft2)
48.9 (10.0)
39.1 (8.0)
31.8 (6.5)
43.9 (9.0)
34.2 (7.0)
31.8 (6 5)
21.9 (4.5)
21.9 (4.5)
21.9 (4.5)
9.8 (2.0)
Average
Cake
Moisture
66
70
71
68
71
72
79
76
78
84
Chemicals
FeCl3
1.5
3.0
2.5
1.5
3.0
2.5
4.0
4.0
5.0
-
CaO
7.0
8.5
4.0
8.0
8.5
4.0
4.0
9.0
5.0
0
E-135
-------
2.5 Process Efficiency and Reliability - See Table £-37.
Vacuum filtration techniques are widely used and proven highly
reliable for dewatering a range of sludge wastes.
2.6 Raw Materials Requirements
• Sludge Conditioning Chemicals - Chemicals such as FeCl^ and lime
reduce the specific resistance of sludge and increase their fil-
tration rate. See Tables E-36 and E-37 and data sheet on
coagulation-flocculation.
t Diatomaceous Earth - Filtration media for precoat vacuum
filtration.
t Water - For water spray used to wash filtration apparatus.
2.7 Utility Requirements
t Electricity - Used for driving pumps and for central drive unit
on filtration apparatus. Requirements vary with the specific
design and removal efficiency desired.
TABLE E-37. TYPICAL VACUUM FILTRATION RESULTS
(5)
Chemical
Sludge Type
Raw
Digested
Raw + Trick-
ling Filter
1 M 1 L. NCI icru
Solids,
Weight
Percent
6-10
6-10
5-7
Raw + | 4-6
Activated Sludge^
i
Activated Sludge! 2-4
KCLju i rei
wt.
FeCl3
1-2
1-4
2-4
2-4
8-10
lien ti ,
%
CaO
5-7
6-10
8-12
8-12
Filter Yield,
kg/hr-m2
(Ib/hr-ft2)
24.4-34.2 (5-7)
29.3-39.1 (6-8)
29.3-39.1 (6-8)
14.7-24.4 (3-5)
2.4-9.8 (0.5-2)
Cake
Moisture,
Wt. %
65-70
70-75
75-80
75-80
80-85
E-136
-------
3.0 Process Advantages^ ' '
• Widely used commercial process for which extensive operating experience
is available.
• Minimal raw materials requirements.
• Suitable for treatment of a wide variety of sludges with differing
physical and chemical properties.
4.0 Process Limitation^5'8)
• Dewatered sludge is generated which requires disposal by incineration,
landfill, or other methods. For precoat vacuum filtration, spent
diatomaceous earth is generated which also requires ultimate disposal.
• Evaluation and operating conditions of vacuum filters on specific
sludges require determination by laboratory "leaf" tests.
t Continuous pilot-scale tests may be required when design information
for large vacuum filtration installation is needed.
t Certain vacuum filtration systems, especially precoat vacuum filtra-
tion systems, have large capital investment and high operating costs.
5.0 Process Economics - See Tables E-38 and E-39.
6.0 Input Streams
6.1 Sludge Feed From Mix Tank (Stream 1) - Sludge characteristics vary
depending on the source. May contain dissolved and suspended
solids, oils, emulsions, heavy metals, etc.
7.0 Intermediate Streams
None.
8.0 Discharge Streams
8.1 Dewatered Sludge (Stream 2) - Moisture content of dewatered sludge
is typically between 60%-80% (see Tables E-36 and E-37). Other
characteristics will vary depending on those of Stream 1.
8.2 Sludge Overflow to Sludge Well (Stream 3) - Same as Stream 1.
8.3 Filtrate (Stream 4) - Will vary, depending upon composition of
Stream 1. Will contain unreacted, excess coagulation chemicals.
E-137
-------
TABLE E-38. VACUUM FILTRATION COSTS OF PRIMARY ACTIVATED DIGESTED
SLUDGE(7)* IN 1973 DOLLARS
Parameter
Chemicals
Direct Labor
Supervision and
Maintenance Labor
Power
Supplies
Total Operating Cost
Amortization and Interest
Grand Total
Cost, $/Tonne
($/Ton) Dry Solids
6.37 (5.79)
2.57 (2.34)
2.57 (2.34)
1.2 (1.09)
0.25 (0.23)
12.97 (11.79)
2.00 (1.82)
14.97 (13.61)
Percent of Total
Operating Cost
49
20
20
9
2
100
*For municipal treatment system handling 12,400 tonnes (13,700 tons)
solids/year, using 4.8 m (16 ft) diameter rotary vacuum filter.
TABLE E-39. COSTS FOR SLUDGE THICKENING AND VACUUM FILTRATION OF
PETROLEUM REFINERY SLUDGES (1967 DOLLARS)(3)
Older
Technology
Typical
Technology
Newer
Technology
Small
Capital
Costs
120,500
59,000
35,000
Refinery*
Annual
O&M Costs
22,000
11,500
9,500
Medium Refinery"!"
Capital Annual
Costs ' O&M Costs
150,000 50,000
82,500 20,500
62,500 12,000
Large Refinery^
Capital Annual
Costs O&M Costs
265,000 58,750
108,500 22,500
82,500 20,500
"Up to 4.2 x 106 liters (35,000 bbl) per day capacity.
f4.2 x 106 to 1.2 x 107 liters (35,000-100,000 bbl) per day capacity.
^Greater than 1.2 x 107 liters (100,000 bbl) per day capacity.
E-138
-------
9.0 Data Gaps and Limitations
Vacuum filtration has not been tested on sludges generated in coal
gasification operations to determine the most suitable operating
conditions.
10.0 Related Programs
None known.
REFERENCES
1. Weber, Jr., W. J., Physiochemical Processes for Water Quality Control,
Wiley-Interscience, New York, 1972, p. 563-571.
2. Environmental Control Issue; Control Equipment, Environmental Science and
Technology, October 1977.
3. The Cost of Clean Water, U.S. Department of the Interior, Federal Water
Pollution Control Administration, Washington, D.C., 1967, p. 39.
4. Eckenfelder, Jr., W. W., Industrial Water Pollution Control, McGraw-Hill
Book Company, New York, 1966, p. 236-256.
5. Azod, H. S., Industrial Wastewater Management Handbook, McGraw-Hill Book
Company, New York, 1976, p. 3-33.
6. Powers, P. W., How to Dispose of Toxic Substances and Industrial Wastes,
Noyes Data Corporation, Park Ridge, N.J., 1976, p. 22.
7. Sherwood, R. J., and D. A. Dablstrom, Economic Costs of Dewatering
Sewage Sludges by Continuous Vacuum Filtration, in AIChE Symposium-Water,
1972, Volume 69, 1973, p. 192-203.
8. Cavanaugh, E. C., J. D. Colley, et al, Environmental Problem Definition for
Petroleum Refineries, SNG Plants and LNG Plants, EPA-600/2-75-068, NTIS
No. PB-252-245, November 1975, pp. 317-318.
E-139
-------
DRYING BEDS
1.0 General Information
1.1 Operating Principle - Dewatering of a sludge by application to beds
consisting of a top layer of sand and a bottom layer of gravel under-
lain by drainage laterals leading to sumps. Initial water loss is
due primarily to filtration of the water through the sludge and
percolation into the sand; after several days, water loss is due
mainly to evaporation. •
1.2 Development Status - In use on a commercial scale. Drying beds are
used for dewatering principal and industrial sludges^ '.
1.3 Licensor/Developer - Sludge drying bed treatment systems and con-
struction materials are offered by numerous design firms and sup-
pliers. Sources are available in the literature.
1.4 Commercial Applications - Method is in widespread use in municipal
wastewater treatment systems. In use at SASOL Lurgi-type coal
(2)
gasification facility in Sasolburg, S. Africav . Also in use in
numerous industrial facilities, including petroleum refineries.
2.0 Process Information
2.1 Flow Diagram - See Figure E-19
• Process Description - Sludge to be dried (Stream 1) is applied to
the surface layer of sand in the drying bed. The bed is sur-
rounded by low walls to retain the sludge and to segregate the
beds from neighboring beds. Water from the sludge percolates
through the sand and gravel layers of the bed and drains through
open-jointed tiles to underground laterals, then is conveyed to
sumps for removal (Stream 2). Dried sludge (Stream 3) is
periodically removed, usually by manual methods, and the bed is
returned to service.
E-140
-------
rn
Open-Jointed
Drainage Tiles
Sludge Layer
Sand Layer
Gravel Layer
•1 CL
Legend:
1. Applied Sludge
2. Filtrate (to Sump)
3. Dried Sludge to Disposal
Figure E-19. Schematic of Sludge Drying Bed
-------
2.2 Equipment'
• Drying bed consisting of 10-22.5 cm (4-9 in.) of sand over
20-45 cm (8-18 in.) of graded gravel over open-jointed tiles
for drainage.
• Retainer walls to enclose drying bed.
• Lateral drainage system, sumps, etc.
• Bed cover material (e.g., glass, plastic, etc.)
(3}
2.3 Feed Stream Requirements^ '
Organic sludges should be pretreated (e.g., by digestion, congula-
tion flocculation, etc.) prior to application to enhance draina-
bility and to prevent the formation of undesirable odors. The
dewaterability of the sludge is a function of its concentration,
size, surface characteristics, extent of aggregation and other struc-
tural characteristics, as well as the quantity, viscosity, ionic
strength and pH of the suspending water.
(A\
2.4 Operating Parameters^ '
Principal operating parameters are sludge loading (wt/unit area and
depth of application), and length of stay. See Table E-40.
(5\
2.5 Process Efficiency and Reliability '
Efficiency depends on the type and design of the system used, on the
nature of the sludge treated, and the duration of its residence time
in the bed. Typically, the applied sludge allowed to dry 10-15 days
to achieve approximately 60% moisture content.
2.6 Raw Materials Requirements - None specific to the process. However,
sludge conditioning chemicals may be required to enhance sludge
dewaterability (e.g., alum, ferric chloride, etc.).
2.7 Utility Requirements - None (except for pumping)
E-142
-------
TABLE E-40. SLUDGE DRYING BED DESIGN
Sludge
Sludge Loading
kg dry solid/m2-yr
(Ib dry solids/ft2-yr)
Primary
Primary + Trickling Filter
Primary + Activated Sludge
97.8-146.6 (20-30)
97.8-146.6 (20-30)
48.9-73.3 (10-15)
3.0 Process Advantages' '4'
• Minimal raw materials requirements.
t Simple to operate.
9 Suitable for treatment of a wide variety of sludges with differing
physical and chemical properties.
9 Costs are usually low.
(1 41
4.0 Process Limitations*'5'
t Significant amounts of labor are required to lift and remove dried
sludge from the beds.
• Large land area required.
• Efficiency of drying is dependent upon climatic conditions.
« Dewatered sludge is generated which requires disposal by incineration,
landfill, or other methods.
e Can cause an odor problem.
5.0 Process Economics
No actual data available. Costs depend on land value; sludge volume;
equipment and labor for dry sludge removal.
6.0 Input Streams
6.1 Applied Sludge (Stream 1) - Sludge characteristics will vary, depend-
ing on the source. May contain dissolved and suspended solids, oils,
emulsions, heavy metals, etc.
E-143
-------
7.0 Intermediate Streams
None.
8.0 Discharge Streams
8.1 Filtrate (to sump) (Stream 2) - Will vary, depending on composition
of Stream 1. May contain excess sludge conditioning chemicals.
8.2 Dried Sludge to Disposal (Stream 3) - Moisture content of sludge
dried 10-15 days is approximately 60 percent. Other characteristics
will vary, depending on composition of Stream 1.
9.0 Data Gaps and Limitations^3'6'
In the past, bed requirements have been based only on empirical relation-
ships or experience factors. Investigators have recently attempted to
derive design criteria from laboratory experiments and pilot operations.
Although several selected variables on sludge drying have been studied in
the laboratory and in pilot-scale operations, further studies are needed
to develop engineering criteria for the design of full-scale systems.
REFERENCES
1. Weber, Jr., W. J., Physiochemical Processes for Water Quality Control,
Wiley-Interscience, New York, 1972, p. 575-6.
2. Information provided by South African Coal, Oil and Gas Corp., Ltd., to
EPA's Industrial Environmental Research Laboratory (Research Triangle
Park), November 1974.
3. Powers, P. W., How to Dispose of Toxic Substances and Industrial Wastes,
Noyes Data Corporation, Park Ridge, N.J., 1976, p. 20.
4. Azad, H. S., Industrial Wastewater Management Handbook, McGraw-Hill Book
Company, New York, 1976, p. 3-36.
5. Cavanaugh, E. C., J. D. Colley, et al, Environmental Problem Definition for
Petroleum Refineries, SNG Plants and LNG Plants, Radian Corporation, Austin
Texas, EPA-600/2-75-068, NTIS No. PB-252-245, 1975, p. 317.
6. Carnes, B. A., Masters' Thesis, Department of Engineering, University of
Texas, 1966.
E-144
-------
EMULSION BREAKING
1.0 General Information
1.1 Operating Principle - Coalescence and separation of the oil and water
phases in a wastewater emulsion by physical methods (e.g., heating,
distillation, centrifuging, precoat vacuum filtration and electro-
lytic methods) and by chemical methods.*
1.2 Developmental Status - Commercially available.
1.3 Licensor/Developer - Equipment and chemicals for emulsion breaking
processes are offered by numerous suppliers. Some licensed or patented
versions of physical emulsion breaking processes are the Oliver pre-
coat vacuum filter and the Cottrell electrical precipitator. Japan's
Mitsubishi Petrochemical Company has developed and is currently
to]
operating an electrolytic coagulation system using iron anodesv '.
A listing of other manufacturers is presented in technical and trade
journals (e.g., Ref. 1). Chemical agents for emulsion breaking are
available through chemical supply firms.
1.4 Commercial Applications - Many applications to petroleum refinery
effluents, including recovered oil from API separators;and other oily
emulsions. No known applications to coal gasification.
2.0 Process Information
2.1 Flow Diagram - Figure E-20 depicts emulsion breaking by chemical
treatment combined with precoat vacuum filtration and heating.
Influent oil emulsion (Stream 1) is heated in a heat exchanger and
discharged into a settling tank maintained at 338°K-350°K (150°F-
170°F).
*See draft data sheet on centrifugration and vacuum filtration for additional
data in these processes.
E-145
-------
HEAT
I
tr>
LEGEND:
1. INFLUENT OIL EMULSION
2. STEAM
3. SEPARATED OIL
4. ACID
5. RESIDUAL EMULSION AND SLUDGE
6 OIL THEATER EMULSION SLUDGE
7.8. TREATED EMULSION AND SLUDGE
9. CHEMICALS
10. WASTEWATER TREATMENT SEDIMENT
11. RESIDUAL EMULSION AND WASTEWATER
SEDIMENT MIXTURE
12. FILTER CAKE
13. DRV OIL
14. WATER (INCLUDING PRECOAT VACUUM
FILTRATION FILTRATE!
IB. ASH
Figure E-20.
Treatment of Recovered Oil by Chemical Derail sification
and Precoat Vacuum Filtration
-------
After about 24 hours, the separated oil layer (Stream 3) is pumped
to an oil treater, where the oil is heated to 355°K-365°K (180°F-
200°F) and acid (Stream 4) is added to assist in emulsion breaking.
After about 48 hours, the oil is skimmed off (Stream 13) and pumped
to the refinery for reprocessing. The residual emulsion and sludge
(Stream 5) are pumped to a bottom sediment and water (BS&W) treater,
where they are heated by steam coils, then pumped (Stream 7) to an
emulsion treater tank for chemical treatment; separated water
(Stream 14) is removed. After two weeks, the oil separated in the
emulsion treater (Stream 8) and the separated water are removed.
Unbroken emulsions and sludge are pumped to a mixing tank where
other plant sludges (Stream 10) are added. The mixture (Stream 11)
is then fed to continuous rotary vacuum precoat filters. The
filtrate is combined with the separated water and sent to plant oil/
water separators and the filter cake (Stream 12) is fed to an
incinerator.
In electrolytic coagulation processes, the influent emulsion is
passed through a tank containing two electrodes. A high potential
pulsating electrical current is applied, which causes the water
globules in the emulsion to coalesce. When the masses attain a
certain weight, they settle by gravity and are withdrawn.
The distillation process of emulsion breaking involves the use of
heat to weaken the interfacial films of emulsions and permit
coalescence and separation of the oil and water phases. Waste
emulsions enter the distillation column, where water and light
ends of the oil are vaporized, then are condensed and withdrawn
as liquid. Residual oil remains in the bottom of the apparatus
and is removed for recycle or disposal.
2.2 Equipment - Depending on the process used, equipment may include:
rotary vacuum precoat filters, centrifuge apparatus, heat exchangers,
pumps, distillation column or tower, electrical precipitator
apparatus (e.g., metal electrodes) or chemical treater unit.
E-147
-------
2.3 Feed Stream Requirements
• Loading - varies with the specific emulsion treatment system,
removal efficiencies desired, and specific design.
• pH - proper pH facilitates breaking of certain emulsions. Adjust-
ment of pH to optimum value may be accomplished by addition of
caustic or acid.
2.4 Operating Parameters^ ' - Flow rates, temperatures, and retention
times vary with the specific process used and the waste treated.
Retention times of two weeks or longer in chemical demulsification
units are typical for API separator emulsions in petroleum refineries.
Temperatures of 338°K-365°K (15°F-200°F) are used in heating emul-
sions both with and without simultaneous chemical treatment. See
Table E-41 for operating parameters for precoat vacuum filtration.
2.5 Process Efficiency and Reliability - Efficiency depends upon the type
and design of the method used, and on the nature of the emulsion.
Emulsions consist of mixtures of water and oil phases and a third
phase known as the stabilizing interfacial film which binds the
oil and water phases together and must be removed or destroyed for
effective emulsion breaking. Emulsions may be ionic, non-ionic,
colloidal (hydropholic or hydrophilic), or may consist of solid
particles which are surface active. It is essential that the
chemical or physical method used for emulsion breaking suit the
characteristics of the specific emulsion being treated; laboratory-
scale testing of an emulsion is sometimes required in order to iden-
tify the appropriate method. Emulsion breaking processes have been
widely used and proven highly reliable for treatment of a range of
industrial emulsions.
2.6 Raw Materials Requirements
• Emulsion breaking chemicals - include acids (sulfuric acid),
caustics (sodium hydroxide, lime), salts (iron sulfate, calcium
chloride, sodium silicate, sodium sulfate, alum), and commercial
organic treatment chemicals.
• Diatomaceous earth - filtration media for precoat vacuum
filtration.
E-148
-------
TABLE E-41. OPERATING DATA FOR PRECOAT VACUUM FILTRATION OF API
SEPARATOR EMULSIONS*4* An
Parameters
Waste Characteristics:
Specific Gravity
Solids, % by Weight
Water, % by Weight
Oil, % by Weight
Kg/MM l(lb/MM Gal)
Sludge/Flow
Volumetric Filtration Rate,
l/m2 . hr (gal/ft2 - hr)
Solids Removed by Filtration,
kg/m2 hr (Ib/ft2 . hr)
Sludge Cake Characteristics:
Percent Water
Percent Oil
Fuel Value, kcal/lb (Btu/lb)
Value
1.0-1.08
2.1-79
2.4-85
0.29-20
110-1285 (64-748)
24.4-58.9 (0.6-1.45)
1.14-1.58 (0.233-0.323)
21
24
5800 (10,500)
2.7 Utility Requirements
• Steam - used for facilitating coalescence and separation of
emulsion phases. Quantity used depends on waste being treated
and loading.
• Electricity - used for emulsion breaking by application of strong
electric fields in which the waste is passed between two elec-
trodes and subjected to a high-potential pulsating current. Also
used for driving centrifuges, pumps, compressors, etc. Require-
ments vary with the specific process design and efficiency
desired.
E-149
-------
3.0 Process Advantages^3'4)
• Effective for separation and recovery of oil from emulsions produced
in API separators or other pollution control processes.
• Suitable for treatment of a wide variety of waste emulsions with
differing physical and chemical properties.
t Little or no raw materials requirements for some methods (e.g.,
distillation, centrifugation, heating).
(•? Q.)
4.0 Process Limitationsv ' '
• Emulsion breaking technology is only quasi-scientific and trial-and-
error experimentation and lab-scale testing are required to determine
the specific process and the proper operating conditions to be
implemented.
• Unresolved emulsions and sludge are generated requiring disposal,
often by incineration with landfill of residual ash.
• Certain processes (e.g., distillation, precoat vacuum filtration)
generate residuals which require subsequent disposal.
5.0 Process Economics
No data available. Centrifugation and precoat vacuum filtration have
greater operating and capital investment costs compared to other emulsion
breaking processes.
6.0 Input Streams
6.1 Influent Oil Emulsion (Stream 1) - Characteristics will vary depend-
ing on the waste source. See Table E-41 for petroleum refinery API
separator emulsion characteristics.
6.2 Steam (Stream 2) - See Section 2.7.
6.3 Acid (Stream 4) - Used to assist in emulsion breaking in conjunction
with heating. See Section 2.6.
6.4 Chemicals (Stream 9) - See Section 2.6.
6.5 Other Plant Sludges (Stream 10) - Characteristics will vary depending
on the waste source.
E-150
-------
7.0 Intermediate Streams
7.1 Separated Oil (Stream 3) - No data available on characteristics; will
be similar to dry oil (Stream 7).
7.2 Residual Emulsion and Sludge (Stream 5) - No actual data available
on characteristics. Will contain oil, water, dissolved and suspended
matter.
7.3 Oil-Treater Emulsion Sludge (Stream 6) - Consists of oil sludge,
residual dissolved and suspended matter, and residual acid from the
oil treater operations. No composition data available.
7.4 Treated Emulsion and Sludge (Streams 7 and 8) - No composition data
available. Similar to dewatered Streams 5 and 6, plus residual treat-
ment chemicals (Stream 8).
7.5 Residual Emulsion and Wastewater Sediment Mixture (Stream 11) -
Combination of Streams 5 and 9.
8.0 Discharge Streams
8.1 Separated Oil (Stream 13) - Consists of separated oil from treater
units. Composition will vary, depending on waste source. May con-
tain residual acid and chemicals from treatment operations.
8.2 Filter Cake (Stream 12) - Consists of suspended solid material,
diatomaceous earth, precoat filter chemicals, and occluded emulsion
treatment chemicals. Cake can usually be burned with the production
of heat in excess of that required to sustain combustion. See
Table E-41.
8.3 Water (Including Precoat Vacuum Filtration Filtrate) (Stream 14) -
Consists of wastewater separated from the influent emulsion in the
BS&W and emulsion treater units, and filtrate from precoat vacuum
filtration. Constituents will vary, depending on the waste and
treatment chemicals used.
8.4 Ash (Stream 15) - Produced by incineration of filter cake (Stream 12).
Suitable for land disposal.
E-151
-------
9.0 Data Gaps and Limitations
Emulsion breaking processes have not been tested on emulsions produced
in coal conversion operations.
10.0 Related Programs
None known.
REFERENCES
1. Environmental Control Issue, Control Equipment, ES&T, October 1977.
2. Chemical and Engineering News, January 23, 1978.
3. Manual on Disposal of Refinery Wastes, Chapter 8, Treatment of Reserved
Oil Emulsions, American Petroleum Institute, Washington, D.C., First
Edition, 1969, p. 8-1 to 8-13.
4. Weston, R. F., Separation of Oil Refinery Waste Water, Ind. Eng. Chem,
42, 607-12, April (1950).
E-152
-------
APPENDIX F
SOLID WASTE MANAGEMENT
Incineration
Land Disposal
Chemical Fixation/Encapsulation
F-l
-------
INCINERATION
1.0 General Information
1.1 Operating Principle - Controlled combustion of waste to destroy
organics and decrease waste volume. Incineration results in the
formation of carbon dioxide, water, ash and other inorganic compounds.
1.2 Development Status - Commercially available; numerous units in
worldwide operation for disposal of municipal and industrial solid
wastes and sludges.
1.3 Licensor/Developer - Numerous incinerator systems and equipment are
available through various suppliers. These include: 1) the Dorr-
Oliver fluidized bed incinerator; 2) the Bartlett-Snow rotary kiln
incinerator; and 3) multiple hearth incinerator systems. Complete
listings of these systems and their suppliers are presented in the
literature^.
(2 3)
1.4 Commercial Applications ' ' - Applications include: a) refinery
wastes, such as spent caustic solutions, API separator bottoms,
DAF float, waste biosludge and slop oil emulsion solids; b) municipal
sewage sludges; c) industrial waste activated sludges; d) neutral
sulfite semi-chemical paper mill waste liquors; and e) pharma-
ceutical wastes. No known application to coal gasification wastes.
2.0 Process Information^ '
2.1 Flow Diagram (See Figure F-l) - Most incineration systems consist of
four basic components: a) a waste storage facility; b) a burner and
oxidation chamber, where the influent waste (Stream 1) is combusted
in the presence of air or oxygen (Stream 2) and secondary pollutants
(CO, C0?, NO , SO and halogen-containing compounds) are formed;
£ A A
F-2
-------
Legend:
1. Influent Waste
2. Air
3. Flue Gas
4. Residuals
Figure 1. Portable Rotary Kiln Incineration Unit
F-3
-------
c) an effluent purification system, when warranted; d) a vent or
stack, for discharge of combustion gases (Stream 3); and e) an ash
removal mechanism (Stream 4).
2.2 Equipment
• Incinerator - Varies with the type of waste incinerated. Three
types of incinerators (i.e., fluidized bed, rotary kiln and
multiple hearth) are commonly used in solids and sludge combus-
tion; multiple chamber and retort incinerators are also used for
the incineration of solids. Are constructed of refractory mate-
rials suited to the desired operating temperature of the
incinerator. See Table F-l for design parameters.
• Oxidant Source Equipment - For supply of air. Required equip-
ment may include blowers, pumps, plenums, etc.
• Bed material, such as sand (for fluidized bed combustors).
t Auxiliary Burners (oil or gas-fired)
2.3 Feed Stream Requirements
• Combustibility - Waste must contain sufficient carbonaceous matter
to be combustible. Depending on the water content, combustion of
biosludges and other wastes produced during gasification of coal
may require supplemental fuel.
• Calorific Value - Adequate heat balance (e.g., the difference
between heat evolved from combustion and heat absorbed due to
vaporization, radiation, etc.).
0 Moisture Content - Excessive moisture must be removed from slur-
ried wastes to minimize the amount of auxiliary fuel required;
moisture content less than 60% typically required.
• Corrosiveness - The corrosiveness of the waste must be accommodated.
by the materials of construction of the incinerator chamber.
• Sulfur, Halogen, and Inorganic Ash Content - Wastes containing
these constituents from combustion products (e.g., SOX, halogen
acids, and inorganic oxides) which may require removal by pollu-
tion control equipment, such as wet scrubbers, electrostatic
precipitators and fabric filters.
2.4 Operating Parameters - See Table F-l - Feed rates depend on feed
characteristics and the furnace design. Feed rates for multiple
hearth incinerators vary from 35-60 kg/hr/m2 (7-12 Ib/hr/ft2)^.
F-4
-------
TABLE F-l. TYPICAL DESIGN AND OPERATING PARAMETERS FOR FOUR TYPES OF INCINERATOR UNITS
(4)
Parameter
Rotary Kiln*
Incinerator Type
Multiple Hearth*
Fluidized Bed*
Multiple Chamber+
Dimensions
Residence
Time
Temperature
Air
Requirement
Capacity
Length/Diameter
Kiln Ratio
of 1:5
Seconds to Hours
1144-1922°K
(1600-35000°F)
37.3-746 kg/hr
(100-2000 Ib/hr)
(7.9 - 279 m2) 85-3000 ft2
Hearth Area Overall Height:
4.75 - 16.6 m (15 ft 7 in. -
54 ft 7 in.)
Seconds to Hours
Combustion Zone: 1033-
1255°K (1400-1800°F)
Upper Hearth: 589-801°K
(600-1000°F)
Cooling Hearth: 513-589°K
(400-600°F)
(22.9-39.4 kg/m -hr)
7-12 Ib/ft2-hr
Bed Diameter
<15.3 m (<50 ft)
Seconds to Hours
760-871°C
(1400-1600°F)
1.5-2.1 m/sec
(5-7 ft/sec)
Length/Width Ratio
of Retort: 2:1
Seconds to Hours
811°K (1000°F)
300% excess air
280-373 kg/hr
(750-1,000 Ib/hr)
*Suitable for sludge and solids incineration.
•^Suitable for solids incineration.
-------
2.5 Process Efficiency and Reliability - Efficiency depends upon the type
wastes incinerated, and temperature and residence time in the combus-
tion chamber. Incineration of waste solids and sludges is widely
performed and has been proven reliable. Occasional operational diffi-
culties may arise due to the type of waste incinerated; e.g., the
incineration of sludge containing chlorides and alkali elements at
low temperatures may lead to plugging in the exhaust gas ducts by
(2)
ash depositsv '.
2.6 Raw Material Requirements
• Fuel - For incineration of sludges having inadequate heat value,
and for start-up; for a waste sludge with a moisture content of
95 percent and a dry heating value of 5500 kcal/kg (10,000 Btu/lb),
3.1 Nm3/tonne (100 scf/ton) of natural gas is required!4).
• Air - Required to support combustion. See Table F-l.
• Water - May be required for pollution control equipment (e.g.,
scrubbers); requirements depend on specific system used and
emission restrictions.
2.7 Utility Requirements
• Electricity and Water - May be required for pollution control
equipment; requirements vary with system used.
3.0 Process Advantages^ '
• Suitable for disposal of many types of wastes, including organic and
partially inorganic sludges and solids, including biosludges.
• Reduces sludges and solids to inert, sterile gases and residuals;
also eliminates waste odors.
t Widely used method for which extensive commercial-scale operating
experience on sludges and solid is available.
• Minimal raw materials requirements (except for auxiliary fuel, when
required).
4.0 Process Limitations
• Generates residual solids (ash) requiring disposal.
• Can generate air pollutants such as particulates, sulfur dioxide,
nitrogen oxides, and metals such as mercury, as well as hydrocarbons
and carbon monoxide.
F-6
-------
• To minimize air pollutant discharges, expensive pollution control
equipment may be required.
• Process has high capital and operating costs.
5.0 Process Economics
Capital investment costs will vary depending on the type of incinerator,
the type and quantity of waste being incinerated, and the nature of pollu-
tion control equipment. Operating costs are a function of the amount of
secondary fuel required, the replacement of refractory linings, and labor.
6.0 Input Streams (See Figure F-l)
6.1 Influent Waste (Stream No. 1) - Will vary depending on the source.
Coal gasification wastes such as chars, tars/oily sludges and bio-
sludges are candidate wastes.
6.2 Air (Stream No. 2) - See Section 2.6.
7.0 Discharge Streams
7.1 Flue Gas (Stream No. 3) - Consists primarily of CO,,, water, and air.
Also contains particulates which vary in quantity and size depend-
ing on the type of waste incinerated, operating procedures, and
completeness of combustion. May also contain hydrocarbons and CO
due to incomplete combustion. Sulfur dioxide, nitrogen oxides, and
metals such as Hg may also be components of the flue gas.
7.2 Residuals (Stream No. 4) - Consist of inorganic, noncombustible
materials including ash and other materials present in the influent
waste (e.g., iron/steel, glass ceramics in municipal wastes).
Require ultimate disposal in landfills or by other suitable methods.
8.0 Data Gaps and Limitations
• Limited data are available on the fate of high molecular weight
organics in tars/oils in incinerators. Materials not destroyed by
incineration may remain as vapors in the flue gas or may be associ-
ated with the ash.
• The combustibility of chars and tars/oily sludges from coal gasifica-
tion has not been studied/established.
9.0 Related Programs
None known.
F-7
-------
REFERENCES
1. Pauletta, C., Incineration, Pollution Engineering, March/April 1970.
2. Becker, K. P. and C. J. Wall, Waste Treatment Advances: Fluid Bed
Incineration of Wastes, Chemical Engineering Progress, p. 61-68, October
1976.
3. Rosenberg, D. G., R. J. Lofy, H. Cruse, E. Weisberg and B. Butler,
Assessment of Hazardous Waste Practices in the Petroleum Refining
Industry, Jacobs Engineering Company, Pasadena, California, NTIS No.
PB-259-097, June 1976, 367 p.
4. Ottinger, R. S., et al., Recommended Methods of Reduction,Neutralization,
Recovery or Disposal of Hazardous Waste, Volume 3, TRW Systems, Inc.,
EPA Contract No. 68-03-0089, February 1973, p. 99-303.
5. Burns, D. E. and G. I. Shell, Physical-Chemical Treatment (PCT) of Waste
Water, Envirotech Corporation, April 1970, 23 pages.
6. Air Pollution Aspects of Sludge Incineration, EPA Technology Transfer
Seminar Publication, EPA-625/4-75-009, June 1975, 16 p.
7. Baum, B., C. H. Parker and DeBell and Richardson, Inc., Solid Waste
Disposal, Volume 1, Incineration and Landfill, Ann Arbor Science
Publishers, Inc., Ann Arbor, Michigan, 1974, p. 56.
F-8
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LAND DISPOSAL PROCESS
(Landfilling, Land Burial, and Application to Soils)
Four major methods of land disposal of coal gasification waste solids and
sludges are: (a) return of the wastes to deep mines; (b) return to sur-
face mines; (c) conventional landfill ing techniques; and (d) soil applica-
tion. Key features of these methods are presented below.
A. Return to Deep Mines^1'2^
Involves return of solid wastes/sludges directly to an underground mine for
ultimate disposal. The applicability of the method would be a function of
the haul distance to the mine, as well as the physical and hydrogeological
conditions of the mine.
Applicability: All types of solid wastes from coal gasification (including
ash, tars, and oily sludges, and biosludges) could be disposed of in deep
mines. The more hazardous wastes, such as heavy metal catalysts, should
be containerized prior to deposition in the mine to minimize environmental
contamination.
Development Status: The procedure of waste return to deep mines is an
untried method which as not yet been field tested. However, procedures
for this type of disposal are currently being developed for the oil shale
industry for the Bureau of Mines .
Operating Considerations/Disadvantages:
• Time Delay - Wastes could not be returned to the mine until sufficient
space became available so that the disposal operation would not inter-
fere with the mining operation.
• Technology Modification - Return of the wastes to an underground mine
would require extensive changes in mine operation procedures which were
originally designed to remove rather than insert large quantities of
material. Underground backfilling would require design and use of
permanent haul ways and access routes for trucks and/or conveyors.
F-9
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• Potential for groundwater contamination - Returning coal gasification
wastes underground carries the potential for groundwater contamination
if water which percolated into the backfilled areas or water contained
in the waste sludge were to leach the wastes and exit the mine area.
Contaminants which could be mobilized would depend on the nature of the
waste, but could include soluble salts and organics and suspended
solids. These contaminants could eventually reach surface waters. The
most suitable mines for disposal would be those whose geological and
hydrological characters would minimize the potential for groundwater
contamination.
• Compaction - Compaction would: (a) maximize the volume of the waste
that would be returned to the mines; (b) minimize leaching of soluble
inorganics/organics and suspended solids. Solids may have to be com-
pacted to reduce volume prior to deposition in mines. When in place,
compacted wastes may help control surface subsidence of the mine.
Compaction could be accomplished either above ground or within the mine.
The operation of heavy mechanical compaction equipment below ground
would require special operating techniques and considerable void space
to accommodate equipment maneuvering. Special safety precautions would
also be needed to minimize the hazards to personnel associated with
equipment moving and working in close quarters.
• Haulage Requirements - Waste can be returned to the mine by means of
short-haul trucks or pumped into the mine as a slurry. Drainage pipes
or pumps would be needed to collect and control the excess water used
in the slurrying operation. The slurry method would have a greater
potential for groundwater contamination due to possible leaching of the
waste and migration of the slurry water.
t Transportation Costs - Costs would be very high, if the gasification
plant is located at a significant distance from the mine.
Advantages:
• Suitable for many types of wastes, if adequate provisions for environ-
mental protection are employed.
• Method is flexible; wide variations in waste loads are readily
accommodated.
• Does not require the availability of large tracts of surface land, as
in landfill ing.
• Waste is sheltered from many natural forces of erosion, such as wind
and precipitation.
• Method requires no revegetation or other surface stabilization procedures.
• Can be beneficial from the standpoint of reducing surface subsidence.
F-K)
-------
v Costs;
Costs associated with return of wastes to deep mines would be functions of:
"" riJcn^2nnf°h TSte ret"rned. *? the mine; the more waste that could be
di'sposal. W 9 * 6SS that W0uld requ1re costly surfa'e
nd -ab0r; the use of haul tnjcks> Pneumatic conveyors and
wi dispsT1^ W°Uld be Slgnif1cant operating costs associated
-- Mine modification; the return of processed solids and other wastes to
the mine may require extensive modification of the mine work area,
ventilation systems, etc.
B. Return to Surface Mines^
As with deep mines, surface mines are likely to be available near coal
gasification facilities and may be suitable and economical for use in
waste disposal, depending on the distance of the mine from the site and
on the environmental suitability of the site.
Applicability: Same as for return to deep mines - see Section A.
Development Status: Although coal gasification wastes have not been
disposed of using this method, the method has been used for the disposal
of municipal refuse, sewage sludges, and power plant fly ash. Reclamation
of surface coal mines using earthen fill has been suspended at several
sites.
Operating Consideration/Disadvantages:
• Time Delay - Wastes cannot be returned to the mine until sufficient
room is available so that the disposal operation does not interfere
with mining operations.
t Erosion Control - Erosion control measures would be required to mini-
mize exposure of the wastes to wind, rain, snow, and other natural
forces. Once filled, the waste disposal area of the mine could be
levelled by bulldozers and covered with topsoil and/or other suitable
materials and vegetated for further erosion control and to improve
the appearance of the site. Interim stabilization techniques may be
implemented, such as application of straw or mulch to cover the
deposited waste.
F-ll
-------
• Compaction - Compaction of the waste would increase mine capacity to
receive waste and would further reduce the erosion potential due to
wind and water action. Compaction could be accomplished more easily
at a surface rather than a deep mine, since special below ground opera-
ting techniques and safety precautions would not be required.
• Haulage Requirement - Same as for return to deep mines.
Advantages:
t No additional acreage required for disposal.
• Suitable for disposal of many types of wastes, if adequate provisions
for environmental protection are employed.
• Method is flexible; wide variations in waste loads are readily
accommodated.
• Less complicated and hazardous than return to deep mines.
• Technology for reclamation of surface mines using non-waste materials
is known and has been utilized on a commercial scale.
Costs:
Are site specific and depend on the quantity and type of waste handled,
and on the number and type of erosion control measures implemented. For
example, costs associated with the establishment of vegetation include
cost of surface preparation, topsoil, mulching, seed and seeding, ferti-
lization, irrigation and maintenance. Costs for vegetating disposal
sites for the commercial-scale oil shale operation proposed for the Colony
development operation have been estimated at $0.50 per square meter
($2,000 per acre), including topsoil^ . Certain costs associated with
disposal in underground mines, such as modifications to ventilation sys-
tems, would not be applicable to surface mines.
Conventional Landfill ing
Method involves disposal of solid waste/sludges on land with provisions
for minimizing environmental contamination. Landfill operations range
from open dumping of debris to controlled disposal in "secure" or "sani-
tary" landfills. Open dumps, in which wastes are piled on the surface
of the terrain, are prohibited in most states and are to be totally
phased out under the provisions of the recently enacted Resource Conserva-
tion and Recovery Act (RCRA). In sanitary landfills, the wastes are
usually compacted to confine them to the smallest practical area, and
F-12
-------
then are covered with a layer of soil at regular intervals (usually at
(3}
the end of a day's operations)v .
Applications: Landfills have been widely used for the disposal of munici-
pal refuse and a range of industrial wastes. Landfill ing is currently
the most prevalent method of disposal of petroleum refinery solid wastes
(e.g., solids and sludges from pollution control processes, spent catalyst,
tars, fly ash, and miscellaneous plant refuser • Although there are no
known applications to each gasification facilities, landfill ing would be
suitable for the disposal of ash, other inorganic solids/sludges, chars,
sludges from biological treatment, and possibly unrecyclable spent cata-
lysts and related materials. Highly hazardous waste may be "chemically"
fixed ("passified") or encapsulated prior to placement in landfills.
Development Status: Commercially available.
Methods of Operation: The principal methods used in landfill ing are
classified as: (a) area; (b) trench, and (c) depression. In the area
method, wastes are spread on the surface of the land in long, narrow strips
that vary in depth from 0.40 - 0.65 m (16-30 in). Each layer is compacted
until the thickness of the compacted wastes reaches 1.8 - 3.1 m (6-10 ft).
A 0.15 - 0.30 m (6-12 in) layer of soil is then placed over the waste.
In this trench method, wastes are placed in trenches varying from 30.5 -
122 m (100-400 ft) in length, 0.9 - 1.8 m (3-5 ft) in depth, and 5 - 8 m
(15-25 ft) in width. The waste is compacted and added until the desired
height is reached, then is covered with soil. The depression method is
similar to the trench method, except that natural or artifical depressions
are used to contain the waste^ .
Design Factor: Factors that must be considered in designing and evaluating
landfill sites include: (a) available land area; (b) soil conditions
(which affect pH and sorptive capacity) and topography; (c) geologic con-
ditions (rock type, geologic structure, and weathering characteristics);
and (d) hydrology (permeability, depth to water table, direction and rate
of groundwater flow; (e) climatological conditions; and (f) potential
ultimate uses for the completed site. Provisions must be made in landfill
F-13
-------
design for diversion and control of surface waters, for leachate collection,
for gas venting, for inclusion of impermeable liners, and for monitoring
wells. Cover materials or liners may be required to suppress air emissions.
Incompatible wastes may require segregation prior to compaction.
Economics: Investment costs are usually low. Operating costs depend upon
the method of operation, the cost of labor and equipment (e.g., motorized
machinery, tools, facilities, fences, drainage pipes, cover material, etc.),
and the efficiency of the operation. Cost for various liners are shown in
Table F-2.
TABLE F-2. LINER COSTS
Liner Type
Cost per Acre (1978)
Clay
Asphalt
Rubber
Hypalon
Polyvinyl Chloride (PVC)
$ 1,185
$ 6,000 - $12,000
$11,000 - $22,000
$11,000 - $22,000
$ 4,840 - $ 9,680
Advantages^1'7)
• Usually the most economical method of solid waste disposal; initial
investment is usually low compared to other methods.
• Suitable for many types of wastes, if adequate provisions for environ-
mental protection are employed.
• Method is flexible; wide variations in waste loads are readily
accommodated.
• Land may be reclaimed for use as parking lots, playgrounds, golf
courses, etc.
Disadvantages^5'7)
• Leachate generated from the waste during compaction or filling activi-
ties, or due to rainwater/snowmelt seepage, may contaminate groundwater
unless adequate leachate containment methods are employed.
F-14
-------
• Suitable land may be unavailable within economic hauling distance of
a coal gasification facility.
• Requires daily and periodic maintenance to prevent environmental
contamination.
• Landfills located in or near residential areas can evoke public
opposition.
• If improperly vented, landfills may generate explosive or hazardous
concentrations of methane and other gases, which may interfere with
the use of the landfill or create a nuisance.
D. Soil Application^8'9'10^
Disposal of solid waste by mixing into topsoil. Organic material in the
waste undergoes degradation through microbial action, and inorganic compo-
nents of the waste are slowly released into the soil, thereby increase in
its nutrient content. Soil application may incorporate production of crops
(e.g., alfalfa) which would offset commercial disposal costs.
Applicability: Method is applicable to the disposal of waste biosludge,
oily sludges and ash. Alkaline ash is particularly useful in soils con-
taining pyritic sulfur such as that near coal mines, which slowly decompose
to acidic products.
Operating Conditions: Waste is distributed on the land in one of three
ways: the "spreading" method, the "flooding" method, and the injection
method. In the spreading method, waste is spread over the land directly
from the tank trucks, or pumped or gravity fed through pipelines to the
agricultural land or land to be reclaimed. In the flooding method, a
plot of land is flooded with the waste and allowed to remain idle until
most of the water is evaporated. Once applied to the land and dried,
rototillers or plows can be used to homogenize the waste into the soil to
depths or up to 51 cm (20 in.).
Some design consideration and process variables involved are: waste com-
position, including toxics concentration, soil composition, nutrient con-
tent and moisture, proximity to surface waters and distance to groundwater
table, nutritional value of the waste, land availability, transportation
costs, effects on vegetation, and atmospheric and climatic conditions.
The actual depth of application is determined by experience. The rate of
degradation and disappearance of the waste depends upon the thickness of
F-15
-------
of the waste deposit, the frequency of tilling and the amount of fertilizer
used.
Development Status: Soil application of digested sludges from small
municipal wastewater treatment plants is common in the U.S. and Europe,
particularly in arid and semi-arid regions. The method is also being used
by a number of petroleum refineries. For example, in the Bakersfield, CA.
area in drilling wastes, miscellaneous oily sludges and acid sludges from
petroleum refineries are being applied to land at waste disposal "farms"
(California Class II-l disposal sites).HO)
Advantages:
t Nutrients present in the waste tend to improve soil texture, water
retention, and overall ability to support vegetation.
• Minimal or no formation of undesirable odors or leachates.
• Minimal disturbance of the land.
• Method is flexible; wide variations in waste loadings are readily
accommodated.
Disadvantages:
• Wastes containing high concentrations of toxic compounds or having an
unfavorably high or low pH cannot be successfully treated.
• Method is dependent upon availability of land and proximity to
waste generation site.
• Aerobic conditions are usually maintained only within the top
10 - 15 cm (4-6 in.) of the soil; hence, periodic plowing of the soil
and rotation of the waste-receiving plots may be required to enhance
oxygen transfer between the ambient atmosphere and the wastes. Waste
accumulations and odor problems may occur under anaerobic conditions.
• Inadequate design of land application sites may result in run-off
of material into receiving waters or contamination of groundwaters.
Cost: Depend on the soil application process utilized, on the quantity
of waste handled, haulage distance to the disposal site, and costs of
periodic plowing to enhance oxygen transfer capabilities of the soil.
Costs also include disposal fees charged by private operators; at the
Bakersfield waste disposal farms, general rates charged ranged from
0.10 to 3<£/liter (15<£ to 35<£/bbl ).(10)
F-16
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REFERENCES
1. Bureau of Mines, U.S. Patent No. 456,509.
?.. Management of Solid Waste Residuals from Oil Shale Recovery Processes,
TRW Systems, Inc., EPA Contract No. 68-01-1881, May 1977, 194 p.
3. G. Tchobanoglous, H. Thiesen, et al, Solid Wastes, McGraw-Hill Book Bo.,
New York, 1977, p. 316.
4. D. G. Rosenberg, R. J. Lofy, et al, Assessment of Hazardous Waste Prac-
tices in the Petroleum Refining Industry, Jacobs Engineering Co., Pasadena,
Ca., NTIS No. PB-259-097, June 1976, p. 117.
5. K. E. Bush, Refinery Waste Treatment, Chemical Engineering, April 12,
1976, p. 113.
6. T. Field, Jr., and A. W. Lindsey, Landfill Disposal of Hazardous Wastes:
A Review of Literature and Known Approaches, U.S. Environmental Protec-
tion Agency, Washington, D.C., EPA-530/SW-165, September 1975.
7. R. S. Ottinger, et al., Recommended Methods of Reduction, Neutralization,
Recovery or Disposal of Hazardous Waste, Volume 3, EPA Contract No.
68-03-0089, 1973.
8. P. W. Powers, How to Dispose of Toxic Substances and Industrial Wastes,
Noyes Data Corporation, Park Ridge, N.J., 1976, p. 134.
9. E. C. Cavanaugh, J. D. Colley, et al, Environmental Problem Definition for
Petroleum Refineries, Synthetic Natural Gas Plants, and Liquified Natural
Gas Plants, Radian Corporation, Austin, Texas, EPA-600/2/75-068, NTIS No.
PB-252-245, November 1975, p. 327.
10. M. Ghassemi, Trip Report - Off-site Industrial Waste Disposal, Environ-
mental Protection Corporation and Associates, Inc., Bakersfield, California,
TRW Systems, Inc., November 1974, 17 p.
11. W. D. Striffer, I. F. Wymore, et al, Surface Rehabilitation of Land Dis-
turbances Resulting from Oil Shale Development, Final Report, Colorado
State University, Fort Collins, Colorado, 1974, 300 p.
F-17
-------
CHEMICAL FIXATION AND ENCAPSULATION
1.0 General Information
1.1 Operating Principle - Chemical fixation (also known as cementation,
waste passification or waste immobilization) employs fixation
chemicals which are mixed with the waste for the purpose of
solidifying the wastes prior to encapsulation and/or disposal.
Encapsulation is a process in which the fixed or untreated wastes
are containerized or coated with inert materials in preparation for
ultimate disposal.
1.2 Development Status - Only a few processes are commercially avail-
able and have been used both domestically and abroad (e.g., pri-
marily Europe and Japan); most processes are in early developmental
stages.
1.3 Licensor/Developer - Several chemical fixation and encapsulation
processes are available through commercial suppliers, such as
Chemfix, Inc. (Pittsburgh, Pa.) and Crossford Pollution Services,
Ltd. (Sole, England). A complete listing of available processes
is available in the literature*1 '.
1.4 Commercial Applications - Chemical fixation and encapsulation
processes have been applied to wastes from numerous industries,
including chemical, petrochemical, and metal finishing industries
(Most applications to date have been abroad; however, usage is gain
ing interest in the U.S.) No known application to coal gasifica-
tion wastes.
2.0 Process Information
2.1 Flow Diagram - See Figure F-2 for the Chemfix Process
• Process Description - The specific operations and equipment
employed in chemical fixation vary from process to process and
F-18
-------
Figure F-2. Schematic of Chem-Fix Process
(2)
-------
in many cases are proprietary. In the schematic of the
Chemfix Process shown in Figure F-2, the waste is pumped through
a reaction tank located in a mobile van. Proper amount of the
fixation chemical (a soluble silicate formulation containing_
setting agents) is mixed with the waste. After proper reaction
time, the mixture is discharged to the final disposal area.
Plastic and metal drums and concrete, asphalt and resins have
been used for containerization and encapsulation of untreated
wastes or chemically fixed wastes.
2.2 Equipment - Vary with the process; equipment may include mixing
chamber, pumps, metering devices, mechanical stirring devices, and
chemical storage tanks.
(3\
2.3 Feed Stream Requirementsv ' - Wastes may require stabilization
prior to fixation/encapsulation for two purposes: 1) to make the
waste more compatible with the solidification step, and 2) to con-
vert the wastes into a chemical form that is more resistant to
leaching in the ultimate disposal site. The most common stabiliza-
tion process is pH adjustment; most cementitious fixation processes
require a pH between 9 and 11.
2.4 Operating Parameters - Vary with the specific process used and
waste being treated. Major parameters include waste: chemical
ratio, retention and drying times, and temperature.
2.5 Process Efficiency and Reliability - The effectiveness of a fixa-
tion process depends upon the type of process used, and on the
nature of the waste being treated. The most important criteria of
effectiveness are mechanical strength and resistance to chemical
attack (e.g., by leachate in a landfill environment) and biodegrada-
tion. Standard laboratory leaching tests have been devised to
evaluate the effectiveness of fixation/encapsulation methods' .
Table F-3 presents typical leaching study results for some
refinery wastes stabilized by the Chemfix process.
F-20
-------
TABLE F-3. LABORATORY LEACHING RESULTS OF CHEM-FIXED
REFINERY WASTES(2)*
Constituent
Total
Chromium (Cr)
Iron (Fe)
Zinc (Zn)
Nickel (Ni)
Copper (Cu)
Manganese (Mn)
Cone, in the
Raw Sludge
ppm
43.5
1310
88.0
8.9
0.62
i
!
0-62
<0.10
<0.25
<0.25 j
j
<0.25
<0.25 |
i
<0.25 ;
Cyanide (Cn) - <0.10
i i
Cm. of Leachate Water"*"
62-125
<0.10
<0.10
<0.10
<0.10
<0.10
<0.10
<0.10
125-188
<0.10
<0.10
<0.10
i
<0.10
<0.10 i
1
t
<0.10
<0.10
188-250
<0.10
<0.10
<0.10
<0.10
<0.10
<0.10
<0.10
Concentration of the constituents in ppm in the leachate water after
application of the specified amount of distilled water.
tEach 62 cm (25 in.) of leachate water represents approximatly 80 ml of
distilled water.
2.6 Raw Materials Requirements
• For chemical fixation: Portland cements; pozzolanic cements;
lime-based mortars; asphalt; polybutadiene; silicate; ion-
exchange resins; epoxies, and various proprietary formulations
(e.g., Chemfix Process).
t For encapsulation: concrete, metal or steel containers, or
self-setting resins.
• pH adjustment chemicals (e.g., sulfuric acid, sodium hydroxide,
etc.).
2.7 Utility Requirements
Electricity: For driving pumps, mechanical stirring apparatus,
and other equipment as required. Requirements dependent upon the
specific process used and volume of waste handled.
F-21
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3.0 Process Advantages^2'3)
• Highly hazardous wastes can be disposed of in a landfill after chemical
fixation/encapsulation.
• Chemicals in the solidified/encapsulated wastes are not accessible
to biodegradation or leaching; minimizes leachate formation from
landfills.
• In some processes, wastes with high water content can be processed
without water discharge from the process.
• Some process applicable over wide ranges of waste composition.
(2 3}
4.0 Process Disadvantages^ ' '
• Relatively high cost.
• Applications generally limited to small volume, high toxicity
wastes.
• Durability and long-term performance of most processes under influence
of environmental conditions (e.g., weather, microorganisms, light) are
not known.
5.0 Process Economics^ ' '
Costs of chemical fixation and encapsulation processes are generally high.
An engineering estimate for the chemical fixation of flue gas desulfuriza-
tion sludge (including final disposal) is $8 to $13/tonne ($7.2 to
$11.8/ton).
6.0 Input Streams
• Influent waste - may include heavy metals, and complex mixtures of
organic and inorganic materials.
t Chemical fixation materials - see Section 2.6.
7.0 Discharge Streams
• Solidified, encapsulated waste.
8.0 Data Gaps and Limitations
Essentially nothing is known about the applicability of fixation/
encapsulation of wastes from coal gasification.
9.0 Related Programs
None known.
F-22
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REFERENCES
1. Powers, P. W., How to Dispose of Toxic Substances and Industrial Wastes,
Noyes Data Corporation, Park Ridge, N.J., 1976, p. 14-18.
2. Conner, J. R., Disposal of Liquid Wastes by Chemical Fixation, Waste Age,
September 1974, p. 26-45.
3. Pojasek, R. B., Stabilization, Solidification of Hazardous Wastes, Envi-
ronmental Science and Technology, Vol. 12 (No. 4), April 1978, p. 382-388.
4. Subramanian, R. V., and R. Mahalingam, Immobilization of Hazardous
Residuals by Encapsulation, Washington State University, Pullman,
Washington, PB-262-648, 46 p'.
5. Fling, R. B., et al., Disposal of Flue Gas Cleaning Wastes: EPA Shawee
Field Evaluation - Initial Report, The Aerospace Corporation, El Segundo,
California, EPA-600/2-76-070, PB-251-876, March 1976.
6. Rossoff, J., and R. C. Rossi, Flue Gas Cleaning Waste Disposal - EPA
Sharonee Field Evaluation, presented at Sixth EPA Symposium on Flue Gas
Desulfurization, New Orleans, March 8-11, 1976.
F-23
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TECHNICAL REPORT DATA
(Please read Inunctions on the reverse before completing!
. REPORT NO.
EPA-600/7-78-186c
3. RECIPIENT'S ACCESSION NO.
4. TITLE AND SUBTITLE Environmental Assessment Data Base
for High-Btu Gasification Technology: Volume in.
Appendices D, E, and F
5. REPORT DATE
September 1978__
6. PERFORMING ORGANIZATION CODE
7. AUTHOR(S)
8. PERFORMING ORGANIZATION REP
M. Ghassemi, K.Crawford, and S. Quinlivan
9. PERFORMING ORGANIZATION NAME AND ADDRESS
TRW Environmental Engineering Division
One Space Park
Redondo Beach, California 90278
10. PROGRAM ELEMENT NO.
EHE623A
11. CONTRACT/GRANT NO.
68-02-2635
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 ANJD PERIOD COVERED
Final; 6/77 - 8/78
14. SPONSORING AGENCY CODE
EPA/600/13
15.SUPPLEMENTARYNOTESJERL-RTP project officer \s William J. Rhodes, Mail Drop 61,
919/541-2851.
is. ABSTRACT^ report jg part of a comprehensive EPA program for the environmental
assessment (EA) of high-Btu gasification technology. It summarizes and analyzes the
existing data base for the EA of technology and identifies limitations of available data.
Results of the data base analysis Indicate that there currently are insufficient data for
comprehensive EA. The data are limited since: (1) there are no integrated plants, (2)
some of the pilot plant data are not applicable to commercial operations, (3) available
pilot plant data are generally not very comprehensive in that not all streams and
constitutents/parameters of environmental interest are addressed, (4) there is a lack
of experience with control processes/equipment in high-Btu gasification service, and
(5) toxicological and ecological implications of constituents in high-Btu gasification
waste streams are not established. A number of programs are currently under way or
planned which should generate some of the needed data. The report consists of three
volumes: Volume I summarizes and analyzes the database; Volume II contains data
sheets on gasification, gas purification, and gas upgrading; and Volume III contains
data sheets on air and water pollution control and on solid waste management.
17.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b.lDENTIFIERS/OPEN ENDED TERMS
c. COSATI Field/Group
Pollution
Coal
Coal Gasification
Assessments
Pollution Control
Stationary Sources
Environmental Assess-
ment
High-Btu Gasification
13B
2 ID
13H
14B
18. DISTRIBUTION STATEMENT
Unlimited
19. SECURITY CLASS (ThisReport)
Unclassified
21. NO. OF PAGES
340
20. SECURITY CLASS (This page I
Unclassified
22. PRICE
EPA Form 2220-1 (9-73)
F-24
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