United States
Environmental Protection
Agency
Industrial Environmental Research
Laboratory
Research Triangle Park NC 2771 1
EPA-600/7-78-186b
September 1978
Environmental Assessment
Data Base for High-Btu
Gasification Technology:
Volume II.
Appendices A, B, and C
nteragency
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:
I 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
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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
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This report has been reviewed by the participating Federal Agencies, and approved
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This document is available to the public through the National Technical Informa-
tion Service, Springfield, Virginia 22161.
-------
EPA-600/7-78-186b
September 1978
Environmental Assessment Data
Base for High-Btu
Gasification Technology:
Volume II. Appendices A, B, and C
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
APPENDIX A - GASIFICATION OPERATION A-l
Dry Ash Lurgi Process A-2
Slagging Gasification Process A-25
Cogas Process A-41
Hygas (Steam Oxygen) Process A-63
C02-Acceptor Process A-88
Synthane Process A-118
Bigas Process A-l37
Battelle-Carbide (Self-Agglomerating Ash) Process .... A-l52
Hydrogasification (Hydrane) Process A-l62
Koppers-Totzek Process A-l 78
Texaco Process A-l96
APPENDIX B - GAS PURIFICATION OPERATION B-l
Acid Gas Removal Module
Physical Solvents
Rectisol Process B-2
Rectisol (Dual Absorption Mode) Process B-14
Selexol Process B-21
Purisol Process B-29
Estasolvan Process ....... B-35
Fluor Solvent Process B-41
Amines
Sulfiban (MEA) Process B-48
MDEA Process B-54
SNPA-DEA Process B-60
ADIP Process B-66
Fluor Econamine (DGA) Process B-73
Alkazid (Alkacid) Process B-80
-------
CONTENTS (Continued)
Mixed Solvents _ oc
. B-ob
Sulfinol Process ................ .
.
Ami sol Process ................
Carbonate Processes
. B-100
Benfield (Hot Carbonate) Process ........
Redox Processes
Giammarco-Vetrocoke (G-V) Process ......... B"
B-l 21
Stretford Process .................
Methanation Guard Module
B-129
Zinc Oxide Adsorption Process .............
Iron Oxide Adsorption Process ............. B-l 36
Metal Oxide Impregnated Carbon Process ........ B-l 45
Activated Carbon Process (Organics Removal from
Gases) ....................... B-151
Molecular Sieves Process ............... B-l 57
APPENDIX C - GAS UPGRADING OPERATION
Shift Conversion Module
Cobalt Molybdate Process ............... C-2
Methanation and Drying Module
Fixed-Bed Methanation Process ............. C-ll
Fluidized-Bed Methanation Process ........... C-25
Liquid Phase Methanation/Shift (LPM/S) Process .... C-31
IV
-------
APPENDIX A
GASIFICATION OPERATION
Dry Ash Lurgi Process
Slagging Gasification Process
Cogas Process
HYGAS (Steam-Oxygen) Process
C02~Acceptor Process
Synthane Process
BIGAS Process
Battelle-Carbide (Self-Agglomerating Ash) Process
Hydrogasification (Hydrane) Process
Koppers-Totzek Process
Texaco Process
A-l
-------
DRY ASH LURGI GASIFICATION PROCESS
1.0 General Information
1.1 Operating Principles - High pressure coal gasification in a gravi-
tating bed by injection of steam plus oxygen with countercurrent
gas/solid flow? ash is maintained below the fusion temperature.
1.2 Development Status - Commercially available since 1940.
1.3 Licensor/Developer - Lurgi Mineral51technik GMbH
American Lurgi Corp.
377 Rt. 17 South
Hasbrouch Heights, New Jersey
1.4 Commercial Applications - See Table A-l.
2.0 Process Information
2.1 Commercial Scale - See Figure A-l for flow sheet.
2.1.1 Gasifier; See Figures A-2 and A-3*.
2.1.1.1 Equipment^1'2)
Construction: vertical, cylindrical steel
pressure vessel
Gasifier dimensions:
- 2.5 to .3.8 m (8.5 to 12.3 ft) in diameter
- 2.1 to 3.0 m (7 to 10 ft) coal bed depth
- 5.8m (19 ft) approximate overall height
of gasifier 3
*
Figure A-2 shows the evolution of Lurai u^n-m^c ,,-n-v,
^fttdiS!:3 * ="« S,T W&TT
A-2
-------
TABLE A-l. LURGI, DRY ASH, COMMERCIAL INSTALLATIONS
(1)
Plant
No.
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
Location
Bohlen,
Central Germany
Bohlen,
Central Germany
Most, CSSR
Zaluzi-Most,
CSSR
Sasolburg,
South Africa
Dors ten,
West Germany
Morwell ,
Australia
Daud Khel ,
Pakistan
Sasolburg,
South Africa
Westfield,
Great Britain
Jealgora, India
Westfield,
Great Britain
Coleshill,
Great Britain
Naju, Korea
Sasolburg,
South Africa
Luenen, GFR
Sasolburg,
South Africa
Year
1940
1943
1944
1949
1954
1955
1956
1957
1958
1960
1961
1962
1963
1963
1966
1970
1973
Type of Coal
Lignite
Lignite
Lignite
Lignite
Sub-Bitum. with 30%
ash and more
Caking Sub-Bitum.
with high chlorine
content
Lignite
High Volatile coal
with high sulfur
content
Sub-Bitum. with 30%
ash and more
Weakly Caking Sub-
Bitum.
Different grades
Weakly Caking Sub-
Bitum.
Caking Sub-Bitum.
with high chlorine
content
Graphitic anthracite
with high ash
content
Sub-Bitum. with 30%
ash and more
Sub-Bitum.
Sub-Bitum. with 30%
ash and more
Gasifier
I.D.
8'6"
8' 6"
8'6"
8'6"
12'1"
8 '9"
8'9"
8'9"
12'1"
8' 9"
N/A
8'9"
8' 9"
10'5"
12'1"
1T4"
12'4"
Capacity
(MMSCFD)
9.0
10.0
7.5
9.0
150.0
55.0
22.0
5.0
19.0
28.0
0.9
9.0
46.0
75.0
75.0
1400 MM
Btu/hr
190.0
No. of
Gasifiers
5
5
3
3
9
6
6
2
1
3
1
1
5
3
3
5
3
A-3
-------
LEGEND:
1.
2.
3.
4,
COAL
0.2
STEAM
FEED LOCK
HOPPER GAS
ASH LOCK
HOPPER GAS
TAR/OIL
SERARATION
FEED LOCK
HOPPER VENT GAS 16. TAR
9. ASH
10. PRODUCT GAS
11. COMBINED LIQUID STREAM
12. SEPARATOR FLASH GAS
13. OIL
14. LIQUOR
15. RECYCLE TAR
RAM GAS
ASH LOCK
HOPPER VENT GAS
17. RECYCLE LIQUOR
Figure A-l. Lurgi Gasifier (Based on Westfield Lurgi Installation)
-------
en
year
first generation
1936 -1954
CAS
coal grade
capacity
MM BTU coal input
hr
lignite
100
second generation
1952-1965
third generation
all coal grades
180-250
GAS * -
non-caking coals
£00-500
from 1969
GAS
all coal grades
450-570
Figure A-2. Stages of Lurgi Gasifier Development
-------
FEED COAL
RECYCLE TAR
DRIVE
GRATE >
DRIVE
\_\'i I1/-/
^^J^^
DISTRIBUTOR
GRATE
SCRUBBING
COOLER
GAS
STEAM*
OXYGEN />
WATER JACKET
Figure A-3. Lurgi Pressure Gasifier
A-6
-------
Bed type and gas flow: gravitating bed;
continuous countercurrent gas flow; lateral
gas outlet near the top of the gasifier.
t Heat transfer and cooling mechanism: direct
gas/solid heat transfer; water jacket provides
gasifier cooling.
Coal feeding: intermittent; pressurized lock-
hopper at the top of the gasifier dumps the
coal onto a rotating, water-cooled coal
distributor.
Gasification media introduction: continuous
injection of steam plus oxygen at the bottom
of the coal bed through a slotted ash
extraction grate.
t Ash removal: rotating, slotted grate at the
bottom of the coal bed; refractory lined,
pressurized lock hopper collects the ash and
dumps it intermittently.
Special features:
- Direct quench gas scrubber and cooler which
knocks out the majority of particulates,
tars, oils, phenols and ammonia, is attached
to the gasifier at the gas outlet.
- Gasifier water jacket supplies approxi-
mately 10 percent of the required gasifi-
cation steam.
- Rotating coal distributor provides uniform
coal bed depth.
- Tar injection nozzle at the top of the
gasifier permits recycle of by-product tar
(separated external to the gasification
module) which also helps to reduce coal
fines carryover in the product gas
(optional features).
- Rotating, water-cooled coal bed agitator
aids the gasification of strongly caking
coals (optional feature).
A-7
-------
(1 2)
2.1.1.2 Operating Parameters '
Gas outlet temperature:
Range - 644°K to 866°K (700°F to 1100°F)
Normal - 727°K (850°F)
Coal bed temperatures:
1255°K to 1644°K (1800°F to 2500°F)
Gasifier pressure:
Range - 2-1 to 3.2 MPa (300 to 465 psia)
Normal - 2.1 MPa (300 psia)
0 Coal residence time in gasifier:
Approximately one hour
(1 2)
2.1.1.3 Raw Material Requirements^ '
t Coal Feedstock
Type: All types; strongly caking coals
require agitator reduced throughput
and increased steam rate.
Size: 3.2 to 38.1 mm (0.125 to 1.5 in):
Coal is usually fed in two size ranges;
coal with up to 10% minus 3.2 mm
(0.125 in) can be accepted.
Rate:* 136 to 544 g/sec-m2
(100 to 400 lb/hr-ft2)
Coal pretreatment - crushing and sizing, dry-
ing to less than 35 percent moisture; partial
oxidation is required for use of strongly
caking coals in gasifiers without agitators.
Steam - 1.11 to 2.59 kg/kg
Oxygen - 0.26 to 0.62 kg/kg coal(3)
Quench water - 3.3 x 10'4 m3/kg coal(2)
Rate varies with gasifier design and coal type.
A-8
-------
2.1.1.4 Utility Requirements^1'
Water
Boiler - 2.42 x 10"3 m3/kg coal
(580 gal/ton coal)
Cooling - ?
Electricity - 25 kwh/metric ton (23 kwh/ton)
2.1.1.5 Process Efficiency
(1}
t Cold gasv ' - 63 to 60 percent
r , [Product gas energy output] 1nn
L~J [Coal energy input]
Overall thermal^ ' - 76 percent
r=i [Total energy output (product gas + HC by-products + steam)] ,g0
J [Total energy input (coal + electric power)]
2.1.1.6 Expected Turndown Ratio"' = 100/25
r=i [Full capacity output]
L J [Minimum sustainable output]
2.1.1.7 Sas Production Rate/Yield^ -
0.37 to 0.68 m3/sec-m2 (4875 to 900 scf/hr-ft2)
0.93 60 1.70 Nm3/kg coal (16 to 30 scf/lb-coal)
2.2 Coal Feed Pretreatment - Coal feed is from pressurized lockhoppers,
no pretreatment required in third-gene ration gasifiers.
2.3 Quench and Dust Removal - Crude gas leaves the top of the gasifier
and flows through a scrubber cooler, where it is washed by recir-
culating quench liquor from the tar-oil separation section. The
gases then pass through a waste heat boiler and a final cooler.
Dust, tars and condensables are collected from these units.
3.0 Process Economics - Due to the advanced state of development of the
Lurgi gasifier, numerous studies related to costs have been com-
pleted^,5,6,7). However, most of these address themselves only to
integrated facilities rather than the gasification module. The one
exception, in which equipment lists are presented and detailed cost
estimates made, is the Bureau of Mines Study^. For a 250 MMSCFD SNG
facility costing a total of $737,538,000 in 1974 dollars, 27.1 percent
A-9
-------
is estimated to be attributable to the gasification section. Lurgi
estimates total plant costs of $440,000,000 also in 1974 dollars. No
gasification section cost estimates are made.
4.0 Process Advantages
0 Present gasifiers can accept caking and non-caking coals.
t Pressurized operation favors the formation of methane in the gasifier
and reduces upgrading costs. The high pressure of the product gas
would also reduce the cost of gas transmission via pipeline. High
pressure may also be advantageous for combined-cycle or synthesis gas
utilization.
Gasifier has been operated commercially for many years.
Small reactor size may be advantageous for small-scale industrial
applications.
5.0 Process Limitations
Caking coals reduce throughput rate and increase steam consumption
which also increases the amount of liquid waste to be treated.
t Maintaining the coal-bed temperature below the ash fusion temperature
limits the maximum process efficiency.
t Process condensate and by-products require additional processing for
environmental acceptability.
Maintaining a low coal bed temperature results in low steam conversion
in the gasifier.
Limited reactor size may necessitate use of multiple units in parallel
tor large installations.
6.0 Input Streams(3'8)
6.1 Coal (Stream No. 1) - See Table A-2.
6.2 Oxygen (Stream No. 2)
Coal No. 12345
Rate: kg/kg 0.26 0.48 0 49 n M
(Includes 6% inerts) °'62
Pressure MPa 3.6(370) 3.5(360) 3.5(360) 3.5(360) -
A-10
-------
TABLE A-2. PROPERTIES OF COAL FEED TO LURGI GASIFICATION (STREAM NO. 1)
Coal No.
Type/Origin
Size: mm (in)
HHV (dry):
Kcal/kg (Btu/lb)
Swel ling No.
Caking Index
Compos i ti on :
Moisture: %
Volatile matter: %
Ash: %
C: %
H: %
n- i
\) . la
C . ol
o . A
No/
A
Trace Elements*(ppm)
Be
Hg
Ca
Sb
Se
Mo
Co
Ni
Pb
As
Cr
1
Montana Rosebud*
Subbituminous A
6.4-31.8
(1/4-1 1/4)
6553
(11,436)
0
0
24.70
29.20
9.73
67.15
4.22
13.02
1.45
1.20
--
--
--
--
--
--
2
Illinois #6*
High Volatile
6.4-31.8
(1/4-1 1/4)
7094
(12,770)
3
15
10.23
34.70
9.10
71.47
4.83
9.02
3.13
1.35
1.6
1.1
< .03
0.1
--
7
4
14
10
1
20
3
Illinois #5*
Bituminous
6.4-31.8
(1/4-1 1/4)
7228
(13,010)
2.2-5
15
11.94
35.21
8.13
72.80
4.95
7.99
3.56
1.39
2.0
0.2
< .03
.2
9
7
4
32
28
2
15
4
Pittsburgh
#8
6.4-31.8
(1/4-1 1/4)
7826
(14,087)
7.5
30
4.58
37.37
7.74
77.71
5.28
4.74
2.64
1.42
--
-_
__
__
__
_ _
--
5
South African1"
Subbituminous
4989
(8,980)
8.0
31.6
52.4
2.6
11.7
0.43
1.2
__
_ _
_ _
~_
__
__
__
"" ~
(continued)
-------
TABLE A-2. Continued
Coal No.
Type /Origin
Trace Elements*(cont)
(ppm)
Cu
B
Zn
V
Mn
F
Cl
1
Montana Rosebud*
Subbituminous A
--
--
400
2
Illinois #6*
High Volatile
12
132
43
29
20
79
600
3
Illinois #5*
Bituminous
10
307
200
21
22
57
800
4
Pittsburgh
#8
--
--
--
--
1000
5
South African
Subbituminous
_ _
_-.
.-
--
--
__
l
ro
* (3}
From trials of American coals at Westfieldv '.
tData from SASOL unit in South Africa^8'.
fData from trials of American coals at Westfield^ '.
-------
6.3 Steam (Stream No. 3)
Coal No. 1 2345
Rate: kg/kg 1.11 1.97 1.84 2.59
Pressure: (psia) (370) (362) (360) (360)
Temperature:
°K (°F)
6.4 Feed Lock Hopper Gas (Stream No. 4) - No data reported.
6.5 Ash Lock Hopper Gas (Stream No. 5) - No data reported.
7.0 Intermediate Streams
7.1 Gaseous Streams
7.1.1 Feed Lock Hopper Vent Gas (Stream No. 6) - No operational
data reported.
7.1.2 Raw Gas (Stream No. 7) - No operational data reported.
7.1.3 Ash Lock Hopper Vent Gas (Stream No. 8) - No operational
data reported.
7.2 Liquid Streams
7.2.1 Combined Liquid Stream (Stream No. 11) - No data reported.
7.2.2 Recycle Liquid (Stream No. 17) - No data reported.
7.2.3 Recycle Tar (Stream No. 15)^
Coal No. 12345
Toluene (wt %) 32.3 8.6 8.0 3.2
Insoluble (dust) 29.2 10.8 11.1 12.2
Ash
Composition (See Tars - Stream No. 16)
8.0 Discharge Streams
8.1 Gaseous
Product Gas (Stream No. 10) - See Table A-3.
Separator Flash Gas (Stream No, 12) - See Table A-4.
A-13
-------
TABLE A-3. PRODUCTION RATE AND COMPOSITION OF LURGI PRODUCT GAS-STREAM
NO. 10\^'°'
Coal No.
Production Rate:
Nm^/kg coal
(C02> N2, and 02
free basis)
Gas Analysis:
H,
02 (includes
N2+Argon)
CO
CH4
CO 2
C H
£_ U
H2S
Total Organic
Sulfur
NH3
HCN
Naphthalene
St. ClairdeVille
Condensable
1
0.98 m3/kg
41.1%
1.2
15.1
11.2
30.4
0.5%
666g/100Nm3
12-40
0.09
0.27g/100Nm3
0.24
389
2
1.36
39.1
1.2
(N2-0.6)
17.3
9.4
31.2
0.7
1510
23
0.18
2.8
0.68
460
, =j
3
1.79
38.8
1.5
(N2-0.7)
17.5
9.2
31.0
0.5
(C2H4-0.3)
1420
30
not
detectable
8.7
1.1
531
4
1.32
39.4
1.6
(N2-0.8)
16.9
9.0
31.5
0.7
(C2Hg-0.1)
1010
15
0.18
0.50
1.2
277
5
1.36
40.05
--
20.20
8.84
28.78
0.54
422
__
__
A-14
-------
TABLE A-4. COMPOSITION OF LURGI SEPARATOR FLASH GAS-STREAM NO. 12 (VOL.
Coal No.
H2S
NH3
co2
CO
H2
02+Argon
N2
CH4
1
Tar Oil
Sep. Sep.
3.8 8.6
6.3 12.0
64.7 59.3
5.9 4.7
2.9 2.3
3.1 2.5
8.0 6.4
5.3 4.2
2
Tar Oil
Sep. Sep.
5.7 5.5
1.0 1.8
84.9 85.5
1.5 0.8
3.5 3.6
0.4 0.6
1.2 1.0
1.8 1.2
3
Tar Oil
Sep. Sep.
6.2 6.8
4.6 2.7
62.9 67.0
4.5 4.2
11.7 13.3
1.3 1.4
5.9 2.3
2.9 2.3
4
Tar Oil
Sep. Sep.
4.4 5.5
2.9 3.5
71.3 73.9
4.7 3.8
12.0 9.6
0.3 0.2
1.0 0.8
3.4 2.7
5
--
--
--
--
--
--
8.2 Liquid Streams
Tars (Stream No. 16) - See Tables A-5 and A-9 .
Oils (Stream No. 13) - See Tables A-6 and A-9,.
Liquors (Stream No. 14) - See Tables A-7 and A- 9.
8.3 Solids Streams
Ash (Stream No. 9) - See Tables A-8 and A-9 .
9.0 Data Gaps and Limitations
Even though the Lurgi gasifier has the most complete data of any gasifier
due to its advanced state of development, the available data are not
comprehensive in that hot all streams (e.g., lockhopper vent gas) are
addressed, and not all potential pollutants and toxicological and ecologi-^
cal properties are identified. An environmental data acquisition effort
which would lead to the generation of the needed data corresponds to the
EPA's phased level approach to multimedia environmental sampling and
analysis(9).
A-15
-------
10.0 Related Programs
Environmental assessments of commercial scale Lurgi SNG facilities have
been prepared by El Paso Natural Gas for its proposed Burnham facility
and by ANG Coal Gasification Company for its proposed North Dakota Coal
Gasification Project. Documents on process and environmental consider-
ations for other projects have also been released. Chief among these is
the Wesco SNG facility. ERDA is presently continuing tests at the
British Coal Boards Lurgi plant at Westfield, Scotland. The present
series of tests involves operating the Lurgi gasifier in the slagging
mode (this is the subject of another gasifier data sheet). The U.S. EPA
has released a report, "Control of Emissions from Lurgi Coal Gasifi-
cation Plants," (EPA 450/2-78-012, March 1978) which is to provide
information to States and regional EPA offices involved in setting
standards for or evaluating impacts from proposed Lurgi gasification
facilities.
A-16
-------
TABLE A-5. PROPERTIES OF LURGI TAR - STREAM NO. 10
Coal No.
Production Rate:
kg/kg coal
Water: wt. %
Tol uene
insoluble wt. %
Density: grams/cc
Phenols: (wet) wt. %
Calorific Value
Gross: Kcal/kg
(Btu/lb)
Ultimate Analysis
(dry, dust- free
basis)
C wt. %
H wt. %
N wt. %
S wt. %
Cl wt. %
Ash wt. %
0 (by difference
wt. %
1
0.02
30.0
22.0
1.025
5.3
8794
(15,830)
83.06
7., 69
0.65
0.28
0.04
0.05
8.23
2
0.03
26.7
4.5
1.145
2
8829
(15,893)
85.48
6.44
1.18
1.70
N.D.
0.03
5.17
3
0.04
10.4
7.1
1.148
4.7
8837
(15,906)
85.85
6.40
1.19
2.39
N.D.
0.01
4.16
4
0.03
11.9
8.5
1.175
1
8956
(16,120)
88.51
5.93
0.87
1.52
N.D.
0.01
3.16
5
0.02*
--
_ _
--
--
__
--
0.3
--
--
--
A-17
-------
TABLE A-6. PROPERTIES OF LURGI OIL - STREAM NO. 13
(3)
Coal No.
Production rate kg/kg
Water: wt. %
Dust: wt. %
Density: grams/cc
Phenols: (dry, dust-
free) wt. %
Calorific Value
Kcal/kg (Btu/lb)
Ultimate Analysis:
C: wt. %
H: wt. %
N: wt. %
S: wt. %
Cl: wt. %
Ash: wt. %
Oxygen: (by
difference) wt. %
1
0.02
22.3
0.4
0.937
19.1
(16,960)
81.34
9.17
0.46
0.50
0.04
0.03
8.46
2
0.003
4.3
0.8
1.015
20.1
(16,482)
84-82
7.77
0.70
2.40
N.D.
0.01
4.30
3
0.007
5.4
0.1
1.011
19.2
(16,578)
8.488
7.65
0.49
2.27
N.D.
0.01
4.70
4
0.01
15.4
0.02
0.991
10.0
(17,134)
87.33
7.61
0.45
1.50
N.D.
0.01
3.10
5
0.004
--
--
--
0.25
A-18
-------
TABLE A-7. PROPERTIES OF LURGI LIQUORS - STREAM NO. 14
(3)
Coal No.
Prod. Rate kg/ kg
Tar: ppm
Analysis on
tar free
basis
Tar free basis
pH
S.G. at 60°F
T.D.S.:
ppm
T.D.S.
after
ignition
ppm
Sulfide
H£S, ppm
Total S;
ppm
Fatty acids:
ppm
Ammonia:
Free: ppm
Fixed ppm
Carbonate:
ppm
1
0.93
350 650
Inlet Inlet
tar oil
sep. sep.
9.6 8.3
1.003 1.025
4030 1765
45 35
130 115
150 265
1250 1670
3990 14015
395 525
4070 19460
2
2.11
1130 2150
Inlet Inlet
tar oil
sep. sep.
9.8 8.5
1.003 1.032
2770 1570
110 35
25 440
180 730
490 280
1700 17650
280 210
1280 6550
3
1.77
2150 2200
Inlet Inlet
tar oil
sep. sep.
9.5 8.3
1.002 1.027
3180 1120
85 25
15 490
160 930
400 260
1520 13970
410 330
680 9210
4
2.60
300 1100
Inlet Inlet
tar oil
sep. sep.
9.3 8.2
1.000 1.026
1550 1240
105 120
65 520
155 720
275 610
1600 14000
320 250
1360 10740
5
1.06
5000
(tar & oil)
_ _
--
0.03%
10,600
150-200
(continued)
-------
TABLE A-7. Continued
Coal No.
Total phenols:
ppm
Cyanide:
ppm
Thiocyanate:
ppm
Cl : ppm
BOD: ppm
COD: ppm
1
4200 4406
2 4
6 15
45 40
9900 13400
22700 20800
2
2200 1900
3 11
65 160
135 75
3800 4700
10100 12000
3
2900 3750
7 14
79 158
290 170
6000 6200
9300 10600
4
1400 2150
1 12
70 185
240 210
4100 5400
650 7500
5
3250-4000
6
--
--
ro
o
-------
TABLE A-8. PROPERTIES OF LURGI ASH - STREAM NO. 9
(3,10)
Coal No.
Production Rate:
kg/ kg
Angle of repose
Bulk Density
Poured:
kg/Nm3 (lb/ft3)
Tapped:
kg/Nm3 (lb/ft3)
Ash Fusion Temp.
Oxidizing:
I.F.: oc
H.P.: oc
F.P.: oc
Reducing:
I.F.: oc
H.P.: oc
F.P.: oc
Partial analysis
Carbon: wt. %
Si 02'- wt- %
A1203: wt. %
FezOs: wt- %
CaO: wt. %
MgO: wt. %
Sulfur (as
S03): wt. %
Cl: wt. %
1
0.097
24°
918 (57.4)
1078 (67.4)
1240
1260
1290
1165
1175
1210
6.5
46.8
17.7
11.2
8.3
3.9
1.7
0.01
2
0.090
330
762 (47.6)
894 (55.9)
1350
1365
1390
1090
1150
1225
3.2
49.6
20.5
17.2
2.1
1.0
1.3
0.01
3
0.087
41°
990 (61.9)
1106 (69.1)
1280
1300
1330
1030
1060
1070
2.0
46.1
18.1
19.7
3.9
0.7
0.6
0.01
4
0.077
43°
(42.1)
(48.9)
1340
1360
1380
1145
1170
1180
7.6
43.6
20.7
15.0
3.0
0.7
0.8
0.01
5
0.313
__
--
52
28
5
7
1.7
0.2
(continued)
A-21
-------
TABLE A-8. Continued
Coal No.
Trace Elementst (ppm)
Be
Hg
Cd
Sb
Se
Mo
Co
Ni
Pb
As
Cr
Cu
B
Zn
V
Mn
F
1
_-
--
--
--
--
--
--
2
14
.04
<0.3
0.2
--
6
.40
456
96
0.1
750
239
622
469
301
200
5
3
20
.016
<0.3
19
--
8
.38
462
200
0.3
592
273
673
1600
181
305
4.6
4
--
--
--
--
--
5*
--
--
--
--
--
--
--
--
--
--
*Trace element balance for SASOL is presented in Table A- 9
tFrom Reference 10.
A-22
-------
TABLE A-9. TRACE ELEMENT BALANCE FOR LURGI AT SASOL*
(35 OF ELEMENT IN COAL)(8)
Element
Be
B
V
Mn
Ni
As
Cd
Sb
Ce
Hg
Pb
Br
F
Cl
Ash
1
36
72
154
154
36
40
40
72
40
180
3.6
54 1
51 1
Liquor
1.6
3.5
0.06
0.36
0.64
90
35
36
0.1
32
3.2
32
42 1
46 1
Tar
0.5
0.8
0.005
0.005
0.05
2.5
0.5
3
0.003
4.9
8.2
0.05
0.08
0.24
Oil
0.01
0.002
<0.001
<0.001
0.01
5.2
1.1
0.5
0.001
0.5
0.02
0.003
0.008
Total
3
40
72
154
155
134
77
80
72
77
191
36
96
97
*Analysis by spark source mass spectrometer (which can give a semi-
quantitative analysis) for El Paso by SASOL.
t% distribution calculated on analyses as done by Sasol previously.
A-23
-------
REFERENCES
1. Handbook of Gasifiers and Gas Treatment Systems, Dravo Corp., ERDA
FE-17772-11, February 1972.
2. The Lurgi Process: The Route to S.N.G. from Coal, presented at the Fourth
Synthetic Pipeline Gas Symposium, Chicago, Illinois, October 1972.
3. Woodall-Duckham, Ltd, Trials of American Coals in a Lurgi Gasifier at
Westfield, Scotland, Final Report, Research and Development Report No. 105,
FE-105; Crawley, Sussex, England, November 1974.
4. Preliminary Economic Analysis of Lurgi Plant Producing 250 Million SCFD
Gas from New Mexico Coal, Report No. ERDA-75-57, Bureau of Mines,
Morgantown, West Virginia, March 1976.
5. Gallagher, J. T., Political and Economic Justification for Immediate
Realization of a Synfuels Industry, Third Annual International Conference
on Coal Gasification and Liquefaction: What Needs To Be Done Now!,
Pittsburgh, Pennsylvania, August 1976.
6. Kasper, S., Lurgi Gasification Process: Prospects for Commercialization,
Symposium on Coal Gasification and Liquefaction, Pittsburgh, Pennsylvania,
August 1974.
7. The Lurgi Pressure Gasification: Applicability, Lurgi Express Information
Brochure No. 01145/6.75, January 1974.
8. Information provided the Fuel Process Branch of EPA's Industrial Environ-
mental Research Laboratory (Research Triangle Park) by South African
Coal, Oil and Gas Corporation, Ltd, November 1974.
9. forsey, J- A., and Johnson, L. D., Environmental Assessment Sampling and
Analysis: Phased Approach and Techniques for Level 1, EPA-600/2-77-115,
10. Sather, N F. et al, Potential Trace Element Emissions from
the
A-24
-------
SLAGGING GASIFICATION PROCESS
1.0 General Information
1.1 Operating Principles - High pressure gasification of coal in a
gravitating bed by injection of steam plus oxygen with counter-
current gas/solid flow. Gasifier operation in the slagging mode
requires lower steam rates producing high thermal efficiency.
Solids are removed from the gasifier as a molten slag.
1.2 Development Status - The slagging gasifier is being developed by
the Department of Energy (DOE) at its Grand Forks Energy Research
Center (GFERC) and by British Gas Corporation (BGC) and Lurgi
at Westfield, Scotland. The latter work is sponsored by a con-
sortium headed by CONOCO. The DOE/GFERC unit is a relatively small
pilot plant (0.4m gasifier diameter); tests performed to date with
this gasifier have been with bituminous char, lignite char and
lignite. The BGC/Lurgi unit is a pilot plant (2.8m diameter)
based on earlier bench-scale work carried out by BGC at Solihull,
England in the 1950's. The slagging Lurgi Process is the basis
for a proposed demonstration plant sponsored by DOE^ '.
1.3 Licensor/Developer - DOE/GFERC: U.S. Department of Energy
Grand Forks Energy Research Center
Grand Forks, North Dakota
BGC/Lurgi: British Gas Corporation
59 Bryanston St.
Marble Arch
London, W-l, England
Lurgi Mineralb'technik GmbH
P.O. Box 119181
Bockhemeyer Landstrasse 42
D-6 Frankfurt (Main), Germany
1.4 Commercial Developments - None.
A-25
-------
2.0 Process Information
2.1 DOE/GFERC Slagging Gasifier - Flow diagram, see Figure A-4 .
2.1.1 Gasifier
(1 2)
Equipment '
Gasifier Construction: vertical, cylindrical steel pressure
vessel with refractory lining.
t Gasifier Dimensions: diameter 0.4 (16.6 in), coal bed depth
1.8 to 4-6m (6 to 15 ft), overall height 11.6m (38 ft).
t Bed type and gas flow: gravitating bed, continuous counter-
current gas/solids flow, lateral gas outlet near top of
gasifier.
t Heat transfer and cooling mechanism: direct gas/solid heat
transfer, water jacket for gasifier cooling.
Coal feeding: intermittent pressurized lockhopper which is
an integral part of the gasifier.
t Gasification media introduction: continuous injection of
steam plus oxygen through tuyeres (injection ports) in the
sides of the bottom of the gasifier.
t Slag removal: a tap hole in the conical bottom of the
gasifier drains the slag into a quench bath from which it
is passed to a slag lockhopper for intermittent removal.
Special features
- Direct gas quench gas scrubbing cooler knocks out
particulates, tars, oils, phenols and ammonia
- Side stream sample line at top of gasifier allows
raw product gas analysis.
Operating Parameters^1'2'
Gas outlet temperature: 358°F to 644°K (185°F to 700°F)
Maximum allowable coal bed temperature: approximately
1644 K (2500 F), depends upon the ash fusion temperature
of the feed coal. Bed temperature depends upon oxyaen
rate, 02/steam ratio, and coal moisture.
A-26
-------
RECYCLED
WATER
I
STEAM
.DRUMj
I
r^
I
?
I WATER
| JACKET
I SYSTEM
I
L_
INS
n
i
i
i 1
OXYGEN
13
FLARE
STACK
( COMPRESSOR h
Ki
-tr
*^~^
^^s r~ ~~~
RAY , .-
HER 1 i
^^f
r^*^ 1
\
GAS METER
ac.
LU
UJ
Q
GAS LIQUOR RECEIVER
(ATM PRESSURE)
LEOENO:
1. COAL FEED
2. OXYGEN
3. (TEAM
4. QUENCH WATER
. CLAGEUmiElt FUEL
«. RAW PRODUCT OAt
r. moDiicroA*
i. (LAO QUENCH WATER OUT
. VENTOM
W. StAO
11. TAB
II. OUINCHOaOLMCONOENMTC
U. "
GAS METER
Figure A-4. DOE/GFERC Slagging Gasifier Pilot Plant
(1,2)
-------
Gasifier pressure: 0.66 to 2.9 MPa (95 to 415 psia).
Coal residence time in gasifier: approximately 15 to
45 minutes.
/I 28^
Raw Material Requirements ' '
Coal feedstock requirements:
- Type: bituminous char, lignite char or lignite
- Size: 6.4 to 19 mm (0.25 to 0.75 in)
- Rate: 262 to 1860 g/sec-m2 (193-1370 lb/hr-ft2)
Coal pretreatment: crushing and sizing
Steam: 0.30 to 0.46 kg/kg coal (MAP)
Oxygen: 0.48 to 0.55 kg/kg coal (MAP)
Utility Requirements
Water: ?
Electricity; ?
Process Efficiency^ '
Cold gas efficiency; 77.3 to 85.4%
[Product gas energy output] ,nn
[Coal energy input] x
Overall thermal efficiency: ?
[Total energy output (product gas + tic hy-nmHurtc + steam)]
LJotal energy input (coal + electric power)] - x 10°
Expected Turndown Ratio: ?
Gas Production Ratqflfleldfl.Z). 0.53 to 2.1 Nm3/ sec -m2
(6566 to 26,060 scf/hr-ftZ); 1.0 to 1.9 Nm3/kg coal (17 to
33 scf/lb coal)
A-28
-------
2.1.2 Coal Feed/Pretreatment(1) - Coal feed from pressurized lock
hoppers, pretreatment as per Raw Material Requirements in
Section 2.2.1.
2.1.3 Quench and Dust Removal^: Cooled gas leaves the gasifier
and flows through a spray cooler where it is washed and cooled
by a recirculating liquor. The gases then pass through an
indirect cooler and demister before being flared.
2.2 BGC/Lurgi Pilot Plant - Flow diagram, see Figure A-5 .
2.2.1 Gasifier
(3 4}
Equipment^ '
Construction: vertical, cylindrical steel pressure vessel
with refractory lining in lower half of gasifier.
Gasifier dimensions:
Diameter - 0.9m (3 ft) at Solihull
2.8m (9.25 ft) at Westfield
Coal bed depth - 3.1m (10 ft)
Bed type and gas flow; same as DOE/GFERC
Heat transfer and cooling mechanism: same as for DOE/GFERC
t Coal feeding: pressurized lockhopper at the top of gasifier
dumps coal intermittently onto a rotating, water-cooled coal
distributor. Coal fines can be injected into the combustion
zone through the tuyeres (steam-oxygen injection ports).
Gasification media introduction: same as for DOE/GFERC
Slag removal: same as for DOE/GFERC
Special features:
- Direct quench gas scrubber knocks out particulates,
tars, oils, phenols, and ammonia at the gas outlet.
- Rotating coal distributor provides uniform coal bed
composition.
- Sampling ports at the side of the gasifier_permit
measurements of temperature and gas composition.
A-29
-------
STEAM
WASTE MEAT
BOILER
CO
o
12
10
1. COAL FEED
2. OXYGEN
3. STEAM
4. QUENCH WATfR
5 SLAG BURNER FUFU
6. RAW GAS PRODUCT
7 PRODUCT GAS
8. SLAG QUENCH WATER
9 SEPARATOR VENT GAS
10 SLAG
11. TAR
17 QUFNCH COOI ER CONDENSATE
1^ I OCKHOPPER VENT GAS
14 I OCK HOPPER PRESSURIZATION
r;AS
1
-------
Operating Parameters^ ' '
Gas outlet temperature: 473°K to 1073°K (390°F to 1470°F)
t Maximum coal bed temperature: greater than 1533°K (2300°F),
depending on the ash fusion temperature of the feed coal.
t Gasifier pressure: 2.07 to 2.76 MPa (300 to 400 psia).
Coal residence time in gasifier: approximately 10 to 15
minutes.
(3 51
Raw Material Requirementsv * '
Coal feedstock:
- Type: Generally all types; only those coals with high
refractory ash content (15% to 20%) are considered
not well suited. Strongly caking coals require the
use of an agitator.
- Size: 13 to 51 mm (0.5 to 2.0 in)
- Rate: 702 to 1958 g/sec-m2 (516 to 1440 lb/hr-ft2)
Coal pretreatment: crushing and sizing, drying to less
than 35% moisture*
Steam: 0.29 to 0.31 kg/kg coal (MAP)
Oxygen'. 0.48 to 0.53 kg/kg coal (MAP)
Quench water: ?
Uti1ity Requirements
Water: ?
Electricity: ?
0 Fuel (for slag burner): ?
(3\
Process Efficiency^ '
Cold gas: 83%
Overall thermal efficiency: ?
*Instead of drying, Injection of tar and powdered coal (fines) into the fuel
bed through tuyeres may be employed.
A-31
-------
Expected Turndown Ratio; ?
Gas Production Rate/Yield(3>5): 2.03 to 2.14 Nm3/kg coal
(34-36 scf/lb coal).
2.2.2 Coal Feed/Pretreatment - Coal feed is from pressurized lock-
hoppers; pretreatment as per Raw Material Requirements in
Section 2.2.1.
2.2.3 Quench and Dust Removal^3'5^ - Raw gas leaves the top of the
gasifier and flows through a scrubber cooler where it is washed
by a recirculating quench liquor. The gases then flow through
a waste heat boiler. Condensates are sent to tar/oil separation
facilities.
(5)
3.0 Process Economics '
A 1.5 x 103 Nm3/day (60 MMscfd) facility producing 8010 kcal/Nm3
(950 Btu/scf) gas is being designed by a consortium* headed by CONOCO.
Funding is to be provided 50% by the consortium and 50% by DOE. Pro-
jected capital costs for the facility are about $190,000,000. This
figure includes development and engineering, construction and 3.5 years
of operation. A credit of $45,000,000 is taken for the operational
period in estimating the required capital cost. Cost of the gas pro-
duced in the demonstration plant is projected to be $19.9/MM Kcal
($4.79/MMBtu's) and would decrease to $13.8/MM kcal ($3.46/MMBtu's)
for a commercial scale facility. All dollar values are given in
1975 dollars.
4.0 Process Advantages
Increased efficiency and throughput over conventional fixed-fed
non-slagging gasifiers.
Reduced steam consumption.
A-32
-------
High pressure operation favors methane production in the gasifier
and reduces subsequent methanation requirements and hence gas
transmission costs.
Coal fines may be utilized through injection into the coal bed.
Smaller reactor size (for same production rate as for fixed-bed
non-slagging gasifier).
5.0 Process Limitations
Caking coals may require pretreatment coals with low ash content
or a high percentage of refractory ash may require addition of
ash fluxing agents (coals of these types have not yet been
tested).
Condensates and by-products will require additional processing.
Slagging gasifiers have been operated only on a pilot plant scale.
t The DOE/GFERC gasifier has experienced trouble with tap hole
erosion in tests to date.
6.0 Input Streams*
6.1 Coal (Stream No. 1) - see Table A-10.
6.2 Oxygen (Stream No. 2)
Gasifier BGC/Lurgi^ GFERC^
Rate: kg/kg (MAP) 0.48 0.29
6.3 Steam (Stream No. 3)
Gasifier BGC/Lurgi GFERC(8)
Rate: kg/ kg (MAP) 0.29 0.16
6.4 Lockhopper Gas (Stream No. 13 DOE/GFERC) - Product Gas
6.5 Slag Burner Fuel (Stream No. 5 BGC/Lurgi) - ?
6.6 Quench Water (Stream No. 4): ?
*Data for the BGC/Lurgi are from tests run at the Soli hull facility; data
from Westfield are unavailable at the present time. Data for GFERC reflect
the particular coal and operating conditions employed; data for other coals
and operating conditions are being generated which may differ from those
presented here.
A-33
-------
TABLE A-10. PROPERTIES OF SLAGGING GASIFIER FEED COALS
Property
Type
Size, mm (in)
2
Rate, g/sec-m
(lb/hr-ft2)
Flux Added
Composition, %
Volatile Matter
Moisture
Ash
Fixed carbon
Ultimate Analysis, %
Hydrogen
Carbon
Nitrogen
Oxygen
Sulfur
Ash
Higher Heating Value,
J/kg (Btu/lb)
Swelling Number
Caking Index
DOE/GFERd7)
Indianhead/Lignite
6.4 to 19
(0.25 to 0.75)
1053
(775)
None
29.4
27.0
8.6
35.0
5.9
45.6
0.7
38.5
0.7
8.6
1.75 x 107 (7620)
0
0
BGC/Lurgi(6)
Donisthorpe/Weakly
Caking Bituminous
35.4 to 38.1
(1.0 to 1.5)
1952
(14.36)
None
N/A*
13.8
5.6
N/A*
N/A*
N/A*
N/A*
N/A*
1.3
5.6
N/A*
N/A*
N/A*
N/A = Data not available
A-34
-------
7.0 Intermediate Streams
7.1 Raw Gas (Stream No. 6); ?
7.2 Quench Cooler Condensate (Stream No. 12, BGC/Lurgi): ?
8.0 Discharge Streams
8.1 Quenched Product Gas (Stream No. 7) - See Table A-ll
8.2 Coal Lockhopper Vent Gas (Stream No. 13) BGC/Lurgi): ?
GFERC - Product Gas
8.3 Separator Vent Gas (Stream No. 9): ?
8.4 Slag Water (Stream No. 8): ?
8.5 Quench Cooler Condensate (Stream No. 12)
BGC/Lurgi: ?
DOE/GFERC: See Table A-12
8.6 Tar (Stream No. 11): See Table A-13
8.7 Slag (Stream No. 10) - See Table A-14
9.0 Data Gaps and Limitations
Limitations in the data available for slagging gasifiers relate
primarily to specific stream compositions. The major limitations for
the gasifier include:
Feed coals - limited data on ash and trace element composition
Raw and cleaned product gas - no data on trace sulfur and nitrogen
compounds. No trace element data.
Water requirements - no data on quench or cooling water requirements.
A-35
-------
TABLE A-ll. SLAGGING GASIFIER PRODUCT GAS COMPOSITION AND PROPERTIES
(STREAM NO. 7)
Consti tuent/ Parameter
__
CO, %
H , %
m 4. f* U °/
CH4 + L2 4'
C2Hg, %t
co2, %
N2 + Ar, %
\Jn 5 %
H2S
cos + cs2
Tar
Participates
(kg/kg coal)
Higher Heating Value
Kcal/Nm3 (Btu/scf)
GFERC<7)*
56.9
29.6
5.3
0.2
7.5
Not Reported
0.1
Not Reported
Not Reported
Not Reported
Not Reported
3364 (339)
BGC/Lurgi(6)
60.85
28.1
7.7
0.55
2.7
Not Determined
Not Determined
Not Reported
Not Reported
Not Reported
2.3 x 10~2
3162 (375)
*Example data for gasification of Indianhead/1 ignite
''"
"Reported as
which may also contain other light olefins.
A-36
-------
TABLE A-12.
PROPERTIES OF RAW GAS QUENCH CONDENSATE FROM
DOE/GFERC GASIFIER (STREAM 12)(7)
Production Rate: kg/kg coal
Consti tuent/Property:
pH
Alkalinity, ppm as CaCOs
Turbidity, JTU
TOC, ppm
TOC, kg/kg coal
NHs, ppm
NH3» kg/kg coal
Suspended tar, ppm
Particulate, ppm in liquor
Liquor particulate, kg/kg coal
Total Dissolved Solids, ppm
Inorganic Dissolved Solids, ppm
Mass Spectrometer Analysis of
Organic Liquor Fraction, %
Phenol
Cresol
Xylenol
Methyl naphthalene
Biphenyl
Dimethyl naphthalene
Fluorene
Carbazole
Dibenzofuran
Phenanthrene
Methylbenzofuran
Methylphenanthrene
Pyrene/Fluoranthene
Methylpyrene
Benzonaphthiofuran
Chrysene
Benzopyrene
0.312
8.9
13,500
43
10,015
0.005
9,605
0.004
1,032
45
15 x 10
2,924
418
-6
56.382
19.616
4.523
0.341
0.188
0.264
0.174
0.091
0.740
3.175
1.009
0.764
1.004
1.776
0.684
0.119
0.711
A-37
-------
TABLE A-13. PROPERTIES OF TAR PRODUCED IN DOE/GFERC
GASIFIER (STREAM II]
Production Rate: kg/kg coal
Moisture content, % as received
Specific gravity at 289°K (60°F)
As received
Dry
Total particulates as received, %
Ultimate analysis, %
C
H
0
N
S
Boiling point distribution, %
Oils to 200°C
200 to 270°C
270 to 300°C
Pitch
Loss
Ash, dry basis %
Heating value, kcal/kg
Mass Spectrometer analysis of tar
sample, dry basis %
Phenol
Indanol
Dibenzofuran
Indanes
Naphthol
Indenes
Pyridines
Qu inclines
Naphthalene
Biphenyl
Fluorene
Phenanthrene
Pyrene
Chrysene
Benzenes
Benzopyrenes
0.25
39.5
0.817
1.017
1.7
83.86
8.72
5.91
0.88
0.49
16.40
30.90
9.30
41.00
2.30
0.37
9320(16776)
13.5
6.8
2.0
3.4
4.9
1.7
2.1
1.2
9.7
2.6
4.1
4.6
2.5
1.6
4.5
1.8
A-38
-------
TABLE A-14. COMPOSITION OF SLAG PRODUCED IN DOE/GFERC
GASIFIER (STREAM NO. 10)(7)
Production Rate, kg/kg coal
Composition, %
Si02
A1203
Fe203
MgO
CaO
Na20
0.67
32.6
13.3
8.1
5.7
21.4
7.7
0.6
-(7)
10.0 Related Programs
Programs for both the DOE/GFERCVM gasifier and the BGC/Lurgi
gasifier^ ' are expected to continue. The DOE/GFERC installation is
expected to perform tests on gasifier operation with different coals.
Also, tests involving the utilization of coal fines, by either agglo-
meration or direct injection into the gasification zone, are to be
performed. The BGC/Lurgi gasifier is involved presently with the
testing of American coals at Westfield. Development work on the
gasifier and demonstration plant is continuing.
Under DOE sponsorship, CONOCO and British Gas Corporation have
conducted tests with American Coals (Pittsburgh No. 8 and Ohio No. 9)
at a Lurgi gasifierin Westfield, Scotland, modified to operate under
slagging conditions^. These tests, which have been aimed primarily
at collecting engineering data for the design of a demonstration plant
in the U.S., have included 48-hour duration runs with (a) Ohio No. 9
premixed with coke; (b) Pittsburgh No. 8 premixed with coke; and
(c) Pittsburgh No. 8 alone. While the runs with Pittsburgh No. 8 have
been very successful, limited success has been obtained with the Ohio
No. 9. Except for one additional "exploratory" run which is planned for
August-September 1978 with Pittsburgh No. 8, the DOE/CONOCO slagging
gasification test program at Westfield is considered complete.
A-39
-------
REFERENCES
1. Ellman, R.C., et al, Current Status of Studies in Slagging Fixed-Bed
Gasification at the Grand Forks Energy Research Center, presented at
the 1977 Lignite Symposium, Grand Forks, North Dakota, May 18-19, 1977.
2. Ellman, R.C., et al, Slagging Fixed-Bed Gasification presented at the
4th Annual International Conference on Coal Gasification, Liquefaction
and Conversion to Electricity, Pittsburgh, Pa., August 2-4, 1977.
3. Savage, P.R., Slagging Gasifier Aims for SNG Market, Chemical
Engineering, September 12, 1977.
4. Schora, Frank C., Jr., Fuel Gasification. Advances in Chemistry Series
69, a symposium sponsored by the Division of Fuel Chemistry at the
152nd Meeting of ACS, New York, N.Y., September 1966.
5. Sudbury, John D., J. R. Bowden, et al, Demonstration of the Slagging
Gasification Process, 8th Synthetic Pipeline Gas Symposium, Chicago,
111., October 18-20, 1976.
6. Lacey, J.A., The Gasification of Coal in a Slagging Pressure Gasifier,
American Chemical Society, Division of Fuel Chemistry, 10 (4),
151-67 (1966).
7. Ellman, R.C. and Johnson, B.C., Slagging Fixed-Bed Gasification at the
Grand Forks Energy Research Center, 8th Synthetic Pipeline Gas
Symposium, October 18-20, 1976.
8. Information provided to TRW by R.C. Ellman and M. Fegley of GFERC,
March, 1978.
9. Information provided to TRW by Mr. Robert Verner of DOE, July 24, 1978.
" Synthetl'C Fuels Q^terly, Vol. 15, No. 1,
A-40
-------
COGAS PROCESS
1.0 General Information
1.1 Operating Principles - Low pressure, fluidized coal pyrolysis,
char gasification and char combustion (to provide heat for char
gasification) steps in series with steam injection in the gasifier
and air injection in the combustor.
1.2 Development Status - Two separate portions of the COGAS process
(coal pyrolysis and char gasification/combustion) have been tested
at two different pilot plants. The coal pyrolysis step (the "COED"
Process), which was developed by FMC Corporation under sponsorship
of the U.S. Office of Coal Research, was tested in a 33-tonne/day
(36-ton/day) pilot plant constructed in 1970 at Princeton, N.J.
Tests at this plant were conducted on lignite, subbituminous and
bituminous coals^ '. The char gasification/combustion portion of
the COGAS process has been tested in a 45-tonne/day (50-ton/day)
steam blown char gasifier and air blown char combustor pilot plant,
constructed in 1975 at laboratories of the British Coal Utilization
Research Association, Ltd. (BCURA) in Leatherhead, England.* The
chars from the COED pilot plant were used at the BCURA facility.
The COED chars have also been successfully gasified in a Koppers-
M 9\
Totzek gasifier in Spalrr ' ;.
*The BCURA pilot plant uses recycled char as the heat carrier (see Section 2.2).
An earlier 2.3-tonne/day (2.5-ton/day) pilot plant using an inert solid heat
carrier was tested for a brief period at Princeton, N.J.; this design,
however, was abandoned in favor of the BCURA design.
A-41
-------
In June 1977 ERDA awarded a contract to Illinois Coal
Gasification Group (ICGG)* to design, construct and operate a
2000 tonne/day (2200 TPD) coal gasification demonstration plant
producing 5.58 Nm3/day (18 x 106 scfd) of SNG and 380 m /day
(2400 bbls/day) of syncrude(3\ The plant is to use the COGAS
process, integrating fluidized bed pyrolysis as developed under
Project COED, and steam gasification of the char from pyrolysis.
The plant is to be built in Perry County, Illinois, and is to
process a blend of Herrin No. 6 and Harrisburg No. 5 coal and
other coals.
1.3 Licensor/Developer - COGAS Development Company
P.O. Box 8
Princeton, N.J. 08540
1.4 Commercial Applications - none.
2.0 Process Information
2.1 COED Pilot Plant
2.1.1 Pyrolysis Units (see Figure A-6 )
Equipment
t Construction: four vertical, cylindrical vessels in series.
t Dimensions: PI - 19.7m (6 ft); P2 - 14.8m (4.5 ft);
P3 - 13.1m (4 ft); P4 - 8.2m (2.5 ft)(6).
Bed type and gas flow: Gas flow to the first pyrolysis unit
(PI) is recycled product gas. Gas from the units P2 and P3
goes to the product recovery unit. All beds are fluidized
and involve countercurrent flow of gas and char. Char is
fed forward from PI through each stage to P4. Char transport
from P2 to P3 is by gravity. Steam and oxygen are fed counter-
current to char in P4U.5).
A-42
-------
VENT
-fc.
CO
LEGEND:
1. COAL FEED
2. STEAM
3. AIR/OXYGEN
4. FLUIDIZING GAS
5. RAW PRODUCT GAS
6. RAW OIL
7. PRODUCT GAS
8. FILTERED OIL
9. CHAR
10. CHAR FINES
11. SCRUB LIQUOR
12. SEPARATED WATER
13. FILTER SOLIDS
FLUIDIZED-BED PYROLYSIS UNITS
Figure A-6 . COED Pilot Plant
(8)
-------
Heat transfer and cooling: Gas/solids transfer in pyrolyzers.
Furnace heats recycle gas prior to feed to Pl(5).
Coal feeding: Coal is fed to PI pneumatically using a portion
of the Ist-stage product gas(6).
Gasification media introduction: Gas enters below fluidized
bed in each pyrolyzer.
Char removal: ?
Operating Parameters
Gas outlet temperature: ?
Coal bed temperature:
°K
Design^ ' Western Kentucky^ '
589 (600)
728 (850)
811 (1000)
1089 (1500)
506 (450)
681 (765)
747 (885)
969 (1285)
Pittsburgh Seam^
506 (450)
664 (735)
728 (850)
922 (1200)
PI
P2
P3
P4
Gasifier pressure
MPa (psia)
Western Kentucky Pittsburgh Seam
PI 0.143 (20.80) 0.145 (21.02)
P2 0.139 (20.11) 0.139 (20.21)
P3 0.146 (21.21) 9.150 (21.81)
P4 0.147 (21.32) 0.157 (22.71)
Coal residence time in gasifier^ :
_ Min. _
Western Kentucky Pittsburgh Sean
P1 95 160
P2 30 55
P3 12 35
P4 10
t1wi) m ^ t0 M agglomerating
A-44
-------
Raw Material Requirements
t Coal feedstock: (See Table A-15)
- Type: Essentially all types; feeds tested included Colorado
Somerset. C, Wyoming Monarch, Illinois No. 6 - both high
volatile B and high volatile C bituminous coals, North
Dakota Lignite, Utah A, Western Kentucky Seam Nos. 9 and
14, and West Virginia coal from Pittsburgh No. 8 Seam(l).
- Size: Smaller than 16 mesh (1.08 mm) fed to 1st stage
pyrolyzer(6).
- Rate (wet): Nominal 1364 kg/hr (36 TPD); Western Kentucky
1159 kg/hr (30.6 TPD); Pittsburgh 636 kg/hr (16.8 tons/day)(2) *
Coal pretreatment: Coal is dried at 463°K (375°F). For
Pittsburgh coal, oxygen pretreatment was used(l).
t Steam(8,9,10)t: 0.15 - 0.69 kg/kg coal feed, dry basis.
See Table A-16.
0 Oxygen(8,9,10); 0.06 - 0.24 kg/kg coal feed, dry basis.
See Table A-16.
Process Efficiency
Cold product efficiency:
Energy in product gas output ,00 _ ?
Total energy in coal input
Overall thermal efficiency^ ':
Total energy in product gas, tar, oil and by-products
Total energy input in coal + electricity
= 84% (Utah coal)
= 79% (Illinois coal)
Oil energy efficiency* '.
Total energy in syncrude 1QO = m utah coal
Total energy in coal
= 23%, Illinois coal
*Difference in feed rates (and residence times) are due to highly agglomerating
.characteristics of Pittsburgh coal.
Used for bed fluidization and heating.
A-45
-------
TABLE A-15. PROPERTIES AND FEED RATES OF COAL FEED (STREAM NO. 1)
Coal Origin
Coal Type
Reference
Moisture *
Ultimate Analysis,
wt. * dry
Ash
C
H
0
S
N
HHV kcal/kg
(Btu/lb)
Proximate Analysis,
wt. % dry
Volatile
matter
Fixed
carbon
Ash
Coal Feed Rate,
kg/hr
(wet)
Colorado
High Volatile
Bituminous
8
6.0-15.0
6.9
73.4
5.4
12.1
0.6
1.6
7,430
13,380
37.7
55.4
6.9
925
1,080
Wyomi ng
Sub
Bituminous
8
14.7-22.6
17.7
57.7
4.5
15.7
0.6
0.8
6,110
11,000
40.3
42.0
17.7
776
1,180
Illinois
High Volatile B
Bituminous
8
11.0
12.4
67.0
4.7
11.0
3.7
1.2
6,980
12,556
38.1
51.9
64.
658
1,383
No. Dakota
Lignite
9
12.2-25.9
7.7
59.8
4.6
26.3
0.6
0.9
5,890
10,610
47.0
44.7
8.3
Utah
High Voltatile B
Bituminous
9
2.7-4.0
6.0
75.0
5.8
10.6
0.6
1.6
7,555
13,600
42.2
50.2
7.6
:
Illinois
High Volatile C
B1 tumi ndus
9
8.3-9.6
10.9
67.5
4.9
11.1
4.3
1.3
6,794
12,230
38.6
50.4
11.0
H. Kentucky
High Volatile B
Bituminous
10
0.88-5.1
9.7
71.8
5.0
8.8
3.2
1.5
7.230
13,020
35.9
54.2
9.7
680
1,270
Pittsburgh
High Volatile B
Bituminous
10
2.3-4.0
8.4
73.8
5.3
2.4
3.9
1.2
7,500
13,500
40.2
51.4
8.4
363
657
3>
*
CTt
-------
TABLE A-16. STEAM AND OXYGEN FEED RATES TO COED PILOT PLANT (STREAM NO. 2)
Coal Origin
Coal Type
Reference
Steam to P4
( kg/ kg coal)
Oxygen to P4
(kg/kg coal)
Colorado
High Volatile B
Bituminous
8
.20*
.27*
.06
.12
Wyoming
Sub Bituminous
8
.19
.20
.13
.16
Illinois
High Volatile B
Bituminous
8
.20
.31
.12
.14
No. Dakota
Lignite
9
.18
.21
.08
.09
Utah
High Volatile B
Bituminous
.9
.15
.18
.13
Illinois
High Volatile C
Bituminous
9
.20
.32
.11
.18
W. Kentucky
High Volatile B
Bituminous
10
.20
.37
.11
.16
Pittsburgh
High Volatile B
Bituminous
10
.39
.69
.14
.24
*Nitrogen used rather than steam.
-------
TABLE A-17. PROPERTIES OF CHAR PRODUCED IN THE COED PILOT PLANT (STREAM NO. 9)
Coal Origin
Coal Type
Reference
Production
Rate
kg/ kg dry
coal feed
Elemental
Analysis,
(wt. % dry)
C
H
0
N
S
Proximate
Analysis,
(wt. % dry)
Volatile
matter
Fixed
carbon
Ash
HHV,
kcal/kg
(Btu/lb
dry)
Colorado
High Volatile B
Bituminous
8*
.48 - .53
--
--
--
Wyomi ng
Sub Bituminous
8*
.32 - .44
--
~
--
~
--
Illinois
High Volatile B
Bituminous
8*
.49 - .58
79.8
0.7
2.1
1.0
2.1
4.2
81.5
14.3
6130
(11030)
No. Dakota
Lignite
9
.56 - .62
78.4
1.4
3.5
0.7
0.8
8.5
76.2
15.3
6720
(12090)
Utah
High Volatile B
Bituminous
9
.53 - .62
80.9
1.3
0.4
1.4
0.5
6.6
77.7
15.7
6767
(12180)
Illinois
High Volatile C
Bituminous
9
.58 - .64
74.0
0.9
0.3
1.1
2.5
3.8
74.9
21.3
6206
(11170)
W. Kentucky
High Volatile B
Bituminous
9
.53 - .67
76.2
1.9
2.1
2.0
3.3
6.5
78.9
14.6
6697
(12055)
Pittsburgh
High Volatile B
Bituminous
10
77.8
2.7
1.7
1.4
3.7
6.4
81.1
12.5
7139
(12850)
*Wet coal feed basis.
-------
Expected turndown ratio; ?
Gas production rate/yield^ >9'10); 0.12 to 0.69 Nm3/kg dry coal
(2 - 12 scf/lb). See Table A-18.
Oil production rates^8'9'10^; 0.04 - 0.20 I/kg (0.005 - 0.025
gal/lb) dry coal. See Table A-19.
2.1.3 Coal Feed/Pretreatment - Coal is ground by hammermills to minus
3 mm (1/8 in) size and fed to a dryer which is heated by recycled
hot product gas^ '. Dried coal exits at 463°K (375°F) for feeding
to Pl(7).
2.1.4 Quench and Dust Removal - Volatile products leaving P2 pass
through three cyclones in series. Fines collected by the first
two cyclones are returned to the second-stage pyrolyzer. Fines
collected by the third-stage pyrolyzer are discarded. The gas
is quenched in a venturi scrubber at 350°K (170°Fr6'.
2.2 Char Gasification/Combustion Pilot Plant^11'12^ - Flow Diagram
(see Figure A-7 )* .
2.2.1 Gasifier (includes the combustor unit)
Equipment
Construction: vertical, cylindrical steel.
Dimensions: ?
Bed type and gas flow:
Gasifier - fluidized, countercurrent, gas/solids flow.
Combustor - entrained, cocurrent gas/solids flow.
Heat transfer and cooling: Water-cooled coils used in com-
bustor. For other units, no data.
Char feeding: ?
t Gasification media introduction: Continuous feeding of steam
and air to bottoms of gasifier and combustor units, respectively.
Ash removal: slagging combustor with ash removal from bottom
of combustor. No quench data.
A-49
-------
16
i
en
O
FUEL
FINES
CHAR
COMBUSTOR
HOT
RECYCLE
t,
N.CHAR />* >.
\
\
i
r 1
GASSIFIER
L J
^
I
.-**
j
1
^
.X
14 H- CHAR FINES
9
LE
2.
3.
9.
14.
15.
16.
17.
2
LEGEND
2. STEAM
3. AIR
9. CHAR FEED
14. RECYCLE CHAR
15. GASIFIERGAS
16. FLUE GAS
17
*The specific stream numbering system conforms to those used in Figures A-6 and A-8.
Figure A-7. Cogas Char Gasification Pilot Plant Unit-Char Heat-Carrier Process
-------
TABLE A-18. PRODUCT GAS PRODUCED FROM THE COED PILOT PLANT (STREAM NO. 7)
Coal Origin
Coal Type
Reference
Production
Rate
Nm3/kg
(scf/lb
dry coal )
Composition,
vol %
N2
co2
CO
H2
CH4
C2H4
C2H6
C3H6
C3H8
V
H2S
HHV
kcal/Nm3
(Btu/scf)
Colorado
High Volatile B
Bituminous
8*
.198-. 246
(3.36-4.17)
--
~
~
--
Wyoming
Sub Bituminous
8*
.342-. 428
(5.80-7.26)
--
--
Illinois
High Volatile B
Bituminous
8*
.296-. 509
(5.03-8.61)
7.8
22.4
15.4
39.6
9.6
0.4
0.2
0.1
0.3
3.2
3020
(340)
No. Dakota
Lignite
9
.468-. 690
(7.96-11.73)
8.42
28.79
8.46
43.75
9.36
0.13
0.22
0.08
0.15
0.15
0.46
2520
(283)
Utah
High Volatile B
Bituminous
9
.313-. 407
(5.33-6.92)
5.77
23.25
18.96
34.38
14.62
0.48
0.90
0.38
0.33
0.52
0.41
3390
(381)
Illinois
High Volatile C
Bituminous
9
.246-. 640
(4.17-10.30)
16.96
22.30
9.83
32.17
13.21
0.74
1.01
0.36
0.30
0.20
2.92
3050
(343)
W. Kentucky
High Volatile
Bituminous
9
.118-. 149
(2.00-2.52)
13.9
4.4
9.2
12.2
0.7
0.8
0.3
0.3
0.3
1.6
2030
(228)
Pittsburgh
High Volatile B
Bituminous
10
.115-. 132
(1.94-2.24)
11.0
6.6
5.3
8.6
0.2
0.9
0.2
0.2
0.3
1.9
1600
(180)
l
tn
*Uet basis
-------
TABLE A-19. PROPERTIES OF RAW OIL PRODUCED IN COED PILOT PLANT (STREAM NO. 6)
Coal Origin
Coal Type
Reference
Trace Elements,
wt %
Iron
Calcium
Sodium
Aluminum
Silica
(as Si02)
Titanium
Colorado
High Volatile B
Bituminous
8
~
--
--
--
Wyoming
Sub
Bituminous
8
-
--
Illinois
High Volatile B
Bituminous
8
--
No. Dakota
Lignite
9
0.033
0.040
0.025
0.013
0.030
0.003
Utah
High Volatile B
Bituminous
9
0.023
0.032
0.047
0.062
0.210
0.0042
Illinois
High Volatile C
Bituminous
9
--
--
W. Kentucky
High Volatile B
Bituminous
9
0.02
0.007
0.001
0.005
0.007
Pittsburgh
High Volatile B
Bituminous
10
--
l
tn
r\>
-------
Operating Parameters
Gas outlet temperature: ?
Char bed temperature; 1200°K (1700°F) in the gasifier; no
data on combustor.
Gasifier pressure: 0.15 - 0.20 MPa (22.0 - 29.4 psia).
Char residence time in gasifier: ?
t Char circulation rate: ?
Raw Material Requirements
Char feedstock*
- Type: obtained from COED pyrolysis pilot plant.
- Size: minus 1.6 mm (1/16 in).
- Rate: 1894 kg/hr (50 tons/day)
Steam: 0.797 kg/kg char (test FT8/2).
t Air: ?
Utility Requirements
Water (26% of the heat in the combustor was lost to water-
cooled coils - test FT8/2)
Electricity: ?
Process Efficiency
Cold gas efficiency:
Product gas energy output m = 53% (t t FT8/2)
Total energy in char feed v ' '
Overall thermal efficiency:
Total energy in product gas, tar, oil and by-products 1QO
Total energy in coal + electricity
= 69% (calculated theoretical)
*Char from COED pilot plant in Princeton, N.J,
A-53
-------
Expected Turndown Ratio; ?
Gasifier Gas Production Rate/Yield; Carbon gasification rate
of 0.203 kg C/hr/kg C in bed (Test FT8/2). 1,66 Nm3/kg
(28.2 scf/lb) of char.
2.2.2 Char Feed/Pretreatment - Char derived from COED pilot plant
runs, no other pretreatment required.
2.2.3 Quench and Dust Removal - Cyclones on both gasifier gas and
flue gas outlets. 7.3% of feed carbon rejected as dust in
gasifier cyclone gas.
2.3 Proposed Demonstration Plant - A process flow diagram for an
integrated COGAS plant is shown in Figure A-8 ^. Additional
updated information about the proposed demonstration plant is
not currently available.
3.0 Process Economics
No updated information is available for the integrated COGAS
process.
4.0 Process Advantages
Does not require use of oxygen.
Substantial quantities of synthetic oil and a qas suitahlP
'
been s"«<^lly demonstrated on a
*
. us1"9
5.0 Process Limitations
of
has notbeenedemtratedeSS (pyrolys1s and gasification/combustion)
A-54
-------
>
in
LEGEND:
1.
2.
3.
4.
5.
6.
7.
8.
9.
10.
11.
12.
13.
14.
15.
16.
17.
COAL FEED
STEAM
AIR/OXYGEN
FLUE GAS
RAW PRODUCT GAS
RAW OIL
PRODUCT GAS
FILTERED OIL
CHAR
CHAR FINES
SCRUB LIQUOR
SEPARATED WATER
FILTER SOLIDS
RECYCLE CHAR
GASIFIER GAS
FLUE GAS
ASH
OIL
DEHYDRATOR
cc
LLI
N
o
cc
DC
HI
N
EC
LU
N
DC
Q.
9 T
15
16
T
14
QC
O
8
17
Figure A-8. COGAS Integrated Demonstration Plant
(9)
-------
The synthetic oil produced requires hydrotreating before use as
substitute for conventional petroleum products.
0 Process is low pressure, therefore the product gas must be
pressurized (at least for methanation and pipelining).
6.0 Input Streams
6.1 Coal Feed Rate (Stream No. 1) - No data for the integrated COGAS
demonstration plant. See Table A-15 for coal feed rates and prop-
erties used in the COED pilot plant.
6.2 Steam (Stream No. 2) - No data for the integrated COGAS demon-
stration plant. Steam rate to gasifier in char gasification pilot
f 1 ?)
plant was 0.79 kg/kg char (see Section 2.3)v '. See Table A-16
for steam rates used in the COED pilot plant.
6.3 Air/Oxygen (Stream No. 3) - No data for the integrated COGAS
demonstration plant. No data for the char gasification pilot
plant. See Table A-16 for oxygen rates used in the COED pilot
plant.*
6.4 Flue Gas (Stream No. 4) - Approximately 0.04 kg flue gas/kg coal
\
(dry) was used in the COED pilot plant to fluidize PI reactor
(Figure A-6 )(8'9'10>.
7.0 Intermediate Streams
7.1 Char Feed (Stream No. 9) - No data for integrated COGAS demonstra-
tion plant. A char feed rate of 30-tonnes/day (34-tons/day) was
used in the char gasification pilot plant^12^. See Table A-17
for properties of char produced in the COED pilot plant.
7.2 Recycle Char (Stream No. 14) - No data for integrated COGAS demon-
stration plant. No data for char gasification pilot plant.
7.3 Raw Oil (Stream No. 6) - See Table A-19 for rates and properties
of raw oil produced in the COED pilot plant.
*ln the integrated COGAS demonstration, air will be fed only to the combustor.
Oyxgen was used in COED pilot plant, but integrated facility ant?cipa?es
using air.
A-56
-------
7.4 Gasifier Gas (Stream No. 15) - No data for the integrated COGAS
demonstration facility. See Table A-20 for gasifier gas produced
in char gasification pilot plant.*
8.0 Discharge Streams
8.1 Raw Pyrolysis Gas (Stream No. 5) - No data for the integrated
COGAS demonstration facility. No data for raw pyrolysis gas
produced in COED pilot plant.
8.2 Flue Gas (Stream No. 16) - No data for the integrated COGAS
demonstration facility. No data for the char gasification pilot
plant.
TABLE A-20. PROPERTIES OF GASIFIER GAS PRODUCED IN THE COGAS CHAR
GASIFICATION PILOT PLANT (STREAM NO. 15)(12)
Production Rate, Nm /kg char
(scf/lb char)
Composition, vol. % (dry)
CO
C00
1.66
(28.2)
53.4
29.4
15.8
1.4
0.0
*This stream will replace the steam and oxygen streams (2&3) used in the
COED pilot plant.
A-57
-------
8.3 Product Gas (Stream No. 7) -See Table A-18 for properties of
product gas produced in the COED pilot plant.
8.4 Filtered Oil (Stream No. 8) - See Table A-21 for trace elements in
filtered oil produced in the COED pilot plant. No other data
available.
8.5 Scrub Liquor (Stream No. 11) - See Table A-22 for properties of
scrub liquor produced in the COED pilot plant.*
8.6 Separated Water (Stream No. 12) - No data for the COED pilot plant.
8.7 Char Fines (Stream No. 10) - No data for the rate of char fines
production. For properties of char fines produced in COED pilot
plant, see Table A-23.
8.8 Filter Solids (Stream No. 12) - No data available.
8.9 Ash/Slag (Stream No. 17) - No data available.
9.0 Data Gaps and Limitations
Data gaps and limitations for the COGAS process relate primarily
to the properties of various process and discharge streams. Since the
coal pyrolysis and char gasification/combustion portions of the process
have been tested at separate pilot plants, no data exist for integrated
COGAS operation. Also, the nature of the downstream quench and gas/oil
separation operations in the proposed COGAS demonstration plant are not
known.
10.0 Related Programs
It is believed that the COGAS Development Company has considerable
data on the COGAS process which are not publicly available at the present.
The recently awarded DOE contract to ICGG for design of a COGAS demon-
stration plant is expected to result in the release of some of these
data. However, release of such data is not expected to be prior to
late 1978 or early 1979.
facilityCh SChOTe 1S n0t knOW" f°r ^ 1ntegrated COGAS demonstration
A-58
-------
TABLE A-21. TRACE ELEMENTS IN FILTERED OIL FROM COED PILOT PLANT (STREAM NO. 8)
Coal Origin
Coal Type
Reference
Production
Rate
(I/kg dry
coal)
Composition
wt. % dry
C
H
0
N
S
Ash
HHV.
kcal/kg
(Btu/lb)
Quinoline
Insolubles,
wt %
Colorado
High Volatile B
Bituminous
8*
.12 - .14
--
--
--
--
--
--
Wyoming
Sub Bituminous
8*
.07 - .09
--
--
--
--
--
Illinois
High Volatile B
Bituminous
8*
.13 - .16
81.4
6.8
7.0
1.2
1.6
2.0
8290
(14920)
No. Dakota
Lignite
9
.04 - .05
77.0
7.0
8.2
0.6
0.6
0.6
--
26.7
Utah
High Volatile B
Bituminous
9
.15 - .20
83.7
8.7
5.6
1.0
0.5
0.5
9110
(16400)
2.9
Illinois
High Volatile C
Bituminous
9
.19 - .20
81.4
7.7
5.9
1.1
1.9
2.0
6.5
W. Kentucky
High Volatile B
Bituminous
9
.12 - .17
82.3
7.5
6.2
1.2
1.6
1.2
8720
(15700)
5.3
Pittsburgh
High Volatile B
Bituminous
10
.14 - .15
83.6
7.2
5.0
1.0
2.4
0.8
8830
(15900)
8.0
in
vo
*Wet coal
-------
TABLE A-22. PROPERTIES OF SCRUB LIQUOR PRODUCED IN THE COED PILOT PLANT (STREAM NO. 11)
Coal Origin
Coal Type
Reference
Production Rate
n/kg coal (dry)
(gal/lb)
Elemental Composition
N, wt %
C, wt %
S, wt %
Dissolved Solids,
wt %
Suspended Solids,
wt %
Phenol Content, wt %
PH
Entrained Oil, wt %
Colorado
High Volatile
B-Bi luminous
8*
0.067-0.133
(0.008-0.016)
--
--
~
--
Wyoming
Sub
Bituminous
8*
0.017-0.080
(-0.002-0.10)
--
--
--
Illinois
High Volatile
B-Bi tumi nous
8*
0.292-0.458
(0.035-0.055)
0.98
0.25
0.77
1.58
0.41
8.8
No. Dakota
Lignite
9
0.031-0.062
(-0.004-0.008)
0.75
2.30
0.52
0.29
1.04
8.9
0.0
Utah
High Volatile
B-Bituminous
9
0.008-0.129
(0.001-0.015)
1.40
2.00
0.02
0.39
0.54
8.9
0.48
Illinois
High Volatile
C Bituminous
9
0.076
(0.009)
0.89
2.00
0.16
0.11
0.68
8.3
0.7
W. Kentucky
High Volatile
B Bituminous
9
0.058-0.210
(0.007-0.025)
0.41
1.10
0.16
0.53
0.17
8.7
0.24
Pittsburgh
High Volatile
B Bituminous
10
0.096-0.016
(0.001-0.002)
0.78
0.86
0.38
1.08
0.20
7.1
0.48
a\
o
*Wet coal basis.
Liquor yields represent total liquor
effluent from plant including fluidizing stream.
-------
TABLE A-23. PROPERTIES OF CHAR FINES PRODUCED IN THE COED PILOT PLANT (STREAM NO. 10)
Coal Origin
i
Coal Type 1
Reference
Moisture %
Ultimate Analysis,
wt %
C
H
N
S
0
Ash
Proximate Analysis,
wt %
Volatile matter
Fixed carbon
Ash
HHV, kcal/kg
(Btu/lb dry char)
Colorado
High Volatile B
Bituminous
8
--
--
--
--
--
--
--
Wyoming
Sub
Bituminous
8
--
--
--
--
--
--
Illinois
High Volatile B
Bituminous
8
--
--
--
--
--
--
--
--
~~
No. Dakota
Lignite
9
2.4
72.2
2.0
0.9
0.5
8.0
16.4
16.7
66.1
17.2
6,150
Utah
High Volatile B
Bituminous
9
1.3
83.6
1.7
1.5
0.8
1.7
10.7
6.4
82.9
10.7
7,150
(11,065) ! (12,870)
Illinois
High Volatile C
Bituminous
9
0.3
68.5
2.1
1.0
3.2
2.8
22.4
11.7
65.9
22.4
6,140
(11,050)
'
W. Kentucky
High Volatile B
Bituminous
9
0.74
77.3
2.6
1.6
2.7
5.6
10.2
13.0
76.8
10.2
6,900
(12,420) [
Pittsburgh
High Volatile B
Bituminous
10
--
en
-------
REFERENCES
1. Merril, R.C., L.J. Scotti, et al, Clean Fuels from Eastern Coals by COED.
Coal Processing Technology, AIChE publication, Vol. 2: 88-93, 1975.
2. Cameron Engineers, Inc. Synthetic Fuels Quarterly Report. Vol. 13, No. 1.
Denver, Colorado, March 1976 pp B-28, B-35.
3. Cameron Engineers, Inc. Synthetic Fuels Quarterly Report. Vol. 14, No. 3.
Denver, Colorado, September 1977. p.B-17.
4. The Dravo Corp. Handbook of Gasifiers and Gas Treatment Systems,
Pittsburgh, Pa., February 1976. pp 39-44.
5. Strom, A.M., and R.T. Eddinger, COED Plant for Coal Conversion. Chemical
Engineering Progress, Vol. 67, No. 3: 75-80, March 1971.
6. Jones, J.F., M.R. Schmid, et al, Char Oil Energy Development, Office of
Coal Research, Washington, D.C., January 1965, 228 pp.
7. Scotti, J.L., L. Ford, et al, The Project COED Pilot Plant, Chemical
Engineering Progress, Vol. 71, No. 4: 119-120, April 1975.
8. L.J. Scotti, B.D. McMunn,et al, Char Oil Energy Development, Research
and Development Report No. 73, Interim Report No. 2, July 1972 -
June 1973, FMC Corp. under contract to the Office of Coal Research
of the Department of the Interior.
9. Scotti, J.L., R.C. Merrill, et al, Char Oil Energy Development, Interim
Report No. 5, July 1973-June 1974, FMC Corp. under contract to ERDA.
10. Jones, J.F., M.J. Brunsvold, et al, Char Oil Energy Development, Vol. I,
Final Report, August 18, 1971-June 30, 1975. FMC Corp. under contract
to ERDA.
11. Bloom, Ralph, Jr., and R. Tracy Eddinger. Status of the GOGAS Process.
In: Sixth AGA Synthetic Pipeline Gas Symposium, Chicago, 111 ,
October 28-30, 1974, 22 pp.
12. Sacks, M.E. and R. Tracy Eddinger, Development of the COGAS Process,
Princeton, N.J., January 1975, 11 pp.
A-62
-------
HYGAS (STEAM-OXYGEN) PROCESS
1.0 General Information'1'4'13^
1.1 Operating Principles - High pressure moderate temperature coal gasi-
fication in four stages of fluidized beds using steam and oxygen.
1.2 Development Status - A 73-tonne/day (80-ton/day) Hygas pilot plant
has been operating in Chicago, Illinois since 1971 under joint
sponsorship by ERDA and the AGA (American Gas Association). The
plant has tested Montana lignite, Illinois bituminous, and Montana
subbituminous coals. Pilot plant operations have demonstrated the
technical feasibility of the Hygas process for the gasification of
lignite and pretreated bituminous coal (based on AGA/ERDA feasibil-
ity criteria). A conceptual design for a demonstration plant is
currently being prepared by Procon under contract to ERDA. This
plant will likely be based on the Hygas process (using the steam-
oxygen version of the process).
1.3 Licensor/Developer - Institute of Gas Technology
3424 South State Street
Chicago, Illinois 60609
1.4 Commercial Applications - None.
2.0 Process Information
2.1 Pilot Plant (Figure A-9 )*(3)
2.1.1 Gasifier (Figure A-10)^
Equipment
t Construction: Vertical, cylindrical steel vessel encompassing
4 gas/solid contacting stages: slurry drying, first-stage
*Acid gss treatment, methanation, wastewater treatment, and certain other
operations at the pilot plant are not shown in the figure.
A-63
-------
>
en
TO OIL-WATER
SEPARATOR
LEGEND
1 DRY COAL
2 AIR
3 OXYGEN
4 STEAM
5 PRETREATER OFFGAS
6 MAKEUP WATER
7 PRETREATER QUENCH WATER
8 PRETREATED COAL
9 PRODUCT GAS CYCLONE SLURRY
10 QUENCHED PRODUCT GAS
11 PRODUCT GAS QUENCH CONDENSATE
12 GASIFIER ASH SLURRY
13 RECYCLE/PRODUCT OIL
14 OIL STRIPPER BOTTOMS
IS OIL STRIPPER VENT GAS
16 SEPARATED WATER
17 SEPARATED SOLIDS/SLUDGES
18 OIL STORAGE VENT GAS
Figure A- 9 . Hygas Pilot Plant Flow Diagram'1'14^
-------
COAL SLURRY
COAL FEED TO FIRST
STAGE
COCURRENT FLOW OF
GAS & SOLIDS
HOT GAS TO FIRST
STAGE
CHAR FEED TO SECOND
STAGE
COUNTERCURRENT FLOW
OF CHAR & GASES
STEAM
OXYGEN
DIAMETER 5'-7" I.D. ASH ^
HEIGHT 132' OVERALL
RAW GAS
> SLURRY DRYING
GAS-SOLIDS
DISENGAGING
FIRST STAGE
HYDROGASIFICATiON
SECOND STAGE
HYDROGASIFICATION
STEAM-OXYGEN
GASIFICATION
Figure A-10. HYGAS Pilot Plant Gasifier with Steam-Oxygen
Gasification(2,14)
A-65
-------
hydrogasification, second-stage hydrogasification and steam-
oxygen gasification.
Dimensions: The outer pressure vessel shell which encloses the
four stages is 1.8m (5.9 ft) in diameter. The overall height,
including skirt, to top gas outlet flange is 40m (131 ft).
Slurry drying top chamber is 0.76m (2.5 ft) in diameter by 4.6ni
(13 ft) high. The first-stage hydrogasifier is a co-current
upflow draft tube O.lm (4 in.) I.D. by about 7.6m (25 ft) long.
The second-stage hydrogasifier is approximately 0.76m (2.5 ft)
I.D. by 9m (30 ft) high. The steam-oxygen gasifier is approxi-
mately 0.60m (2 ft) in diameter by 8m (26 ft) high.
t Bed type and gas flow: Fluidized bed with continuous counter-
current gas flow, horizontal gas outlet after gas proceeds
through slurry drying zone.
t Heat transfer and cooling: Adiabatic reactor with direct gas/
solid heat transfer.
Coal feeding: Coal/light oil mixture charged to gasifier using
slurry pump (see Section 2.1.4).
t Gasification media introduction: Steam and oxygen are injected
into the bottom fluidization zone of the Hygas reactor through
a multiport sparger.
Char removal mechanism: Char exits the bottom of the gasifier
through a full-open ball valve into a quench vessel where it is
first quenched with steam at reactor pressure. The char is then
picked up in a water slurry and discharged through a Willis choke.
Char is also separated from product gas by a cyclone.
Special features: Conditions in the first stage hydrogasifica-
tion zone of the reactor enhance methane formation.
Utility Requirements
Boiler feed water: ?
Cooling water: ?
Quench water blowdown:
Pretreater quench^ - ~8 I/kg (1 gal/lb) 111. bituminous
Product gas quench(1'8) - ~1 l/kg (0.12 gal/lb) lignite
~4 I/kg (0.5 gal/lb) Illinois
bituminous
Electricity: ?
A-66
-------
Process Efficiency
Cold gas efficiency (these data are of less importance than car-
bon conversion efficiency for pilot plant operation) - see
Table A-24:
= (product gas energy output/coal energy input) x 100
lignite: ?
subbituminous (5): 58%
bituminous' ': 67%
Overall thermal efficiency:
_ Total energy output (product gas + HC byproducts + steam)
Total energy input x 1UU
= ?
Raw Material Requirements
Coal feedstock:
Type - Three types of coal have been tested: Montana lignite,
Illinois No. 6, and Montana subbituminous. Technical feasibility
has been demonstrated on all three types.
Size - minus 10 mesh
Rate(3'4'5'14): Run 37 - 1601 kg/hr (3530 Ibs/hr)
Run 54 - 2687 kg/hr (5912 Ibs/hr)
Run 58 - 2490 kg/hr (5478 Ibs/hr)
Run 61 - 1438 kg/hr (3164 Ibs/hr)
Run 63 - 1743 kg/hr (3834 Ibs/hr)
t Pretreatment: Fluid bed pretreater for caking coal at 700 K
(80QOF)
A-67
-------
TABLE A-24. OPERATING PARAMETERS FOR HYGAS PILOT PLANT
Coal bed temperature, °K (°F)
- Slurry drying
- 1st stage hydrogasifier
- 2nd stage hydrogasifier
- Steam-oxygen gasification
Gasifier pressure, MPa (psia)
Carbon conversion (%)
Coal residence time in four
stages:
- Slurry drying
- 1st stage hydrogasificatior
- 2nd stage hydrogasificatior
- Steam-oxygen
hydrogasification
Run 37*^
575 (575)
756 (901)
1009 (1356)
1008 (1499)
7.1 (1040)
90
--
42 min
41 min
Run 54*(4)
610 (580)
772 (854)
982 (1335)
1151 (1640)
6.6 (970)
74
3 min
--
23 min
17 min
Run 58* ^ 5)
--
1118 (1550)
6.2 (911)
67
--
Run 61
596 (616)
868 (1115)
1050 (1424)
1200 (1705)
6.0 (890)
90
--
--
Run 63
608 (634)
894 (1149)
1030 (1393)
1185 (1677)
7.0 (1025)
71
--
--
"
cr>
00
*More than 65 runs have been made at the pilot plant, and a considerable amount of operating data
has been generated. Data for steady periods with lignite (Run 37-1506 hr, 7/5/75 to 0694 hrs,
7/7/75), bituminous coal (Run 54-0000 hrs 7/10/76 to 0000 hrs, 7/11/76; Run 61-1200 hrs 5/9/77 to
1600 hrs 5/9/77; Run 63-1400 hrs 6/24/77 to 1000 hrs 6/25/77), subbituminuous coal (Run 58-0800
hrs 11/15/76 to 0800 hrs 11/16/76) were selected as representative of steady state operations
for the three types of coals.
-------
Steam* and Oxygen:
Run 37(3) Run 54(5) Run 58(5) Run 61(14>
Steam (kg/kg coal) 1.46 1.03 1.25 2.66
Oxygen (kg/kg coal) 0.20 0.14 0.18 0.34
Expected Turndown Ratio
= [Full capacity output/minimum sustainable output] = ?
Gas Production Rate/Yield
f 3} ?
Lignite^': 1.12 NmVkg (18.9 scf/lb) quenched gas (Run 37)
Subbituminous^5': 1.13 Nm3/kg (19.1 scf/lb) quenced gas (Run 54)
Bituminous^4': 1.04 Nm3/kg (17.5 scf/lb) quenched gas (Run 58)
Bituminous^ ': 1.23 Nm3/kg (20.8 scf/lb) quenched gas (Run 61)
Bituminous'14': 1.08 Nm3/kg (18.2 scf/lb) quenched gas (Run 63)
2.1.2 Coal Feed/Pretreatment (Figure A-9 ) - The pretreater is a fluid-
ized-bed type reaction vessel whose major dimensions are 2.5m (8 ft)
in I.D. by 9.9m (30 ft) seam-to-seam. Air is compressed to provide
the fluidizing medium (this air is also used as an in-process oxi-
dant). The offgas from the pretreater enters two internal cyclones
where entrained fines are returned to the bed. In the pilot plant,
the resulting offgas from the pretreater is incinerated.
The pretreater bed operates at about 700°K (810°F). To cool
the pretreated char from 700°K down to 363°K (180°F), a pretreated
char cooler is utilized (closed vessel).
*Steam is also added to the pretreater (along with air) in amounts ranging
from 0.2 - 0.35 Ibs/lb feed coal(14).
A-69
-------
Cooled char is discharged into the char storage hopper and
later sent to the reactor. The vapors from the char cooler
proceed to a quench system.
2.1.3 Quench and Dust Removal (see Figure A- 9 ) - Entrained dust is
removed from the gasifier product gas by dry cyclones and is
discharged through a quench system. The raw gas is quenched
with an externally cooled recycled aqueous condensate. The
resulting oil -water liquor is decanted. Most of the aqueous
phase is cooled and recycled to quench; the net incremental
aqueous condensate is discharged to an oil/water/solids
separation unit. The oil condensate is recycled to the slurry
mix tank.
2.1.4 Coal Feed to Gasifier (see Figure A- 9 ) - Minus 10 mesh coal
feed is mixed with a light aromatic oil (mostly toluene) to
form a slurry which is charged to the high-pressure reaction
system by a slurry pump. The oil in excess of recycle require-
ments is a byproduct of the process. The slurry liquid must
be evaporated before the dried coal enters the actual reaction
zone. This drying is done in a fluidized bed using effluent
gases from the first-stage reactor to provide the heat.
^- '
C O
Based on a preliminary conceptual design of a 7 x 10 Mm /day
3.0 Process Economics
Based on a p
(250 x 10 scfd) commercial Hygas (steam-oxygen) plant, capital
investment for an integrated facility is estimated at $880 x 106 (1976).
The gasification section of such a plant (including coal feeding and
quench) accounts for about 14% of the overall plant cost.
4.0 Process Advantages
The technical feasibility of the process has beejn demonstrated in
pilot plant operations.
Can use essentially any coal, caking or noncaking
Two-thirds of product methane is produced in the gasifier, thus
reducing downstream methanation requirements.
A-70
-------
t Product gas is at pipeline pressure.
Coal feeding system has proved to be a reliable technique for
feeding coal at high pressure.
5.0 Process Limitations
The carbon contained in char collected by the product gas cyclone
system represents a thermal penalty for the process unless the
char can be returned to the gasifier or otherwise utilized.
Hydrocarbon and energy values of pretreater offgas are high and
would require recovery in a commercial operation.
0 Net production of oil has not been demonstrated in the pilot plant.
Toluene solvent is expensive and would affect the economics of
operation if the process is a net consumer of solvent.
6.0 Input Streams (see Figure A-9 )
6.1 Coal (Stream 1) - see Table A-25.
6.2 Steam (Stream 3) - see Section 2.1.1
6.3 Oxygen (Stream 4) - see Section 2.1.1
6.4 Air (Stream 2) - Run 54-0.12 Nm3/kg (1.97 scf 02/lb) coal
- Run 61-0.11 Nm3/kg (1.9 scf 02/lb) coal
- Run 63-0.12 Nm3/kg (2.1 scf 02/lb) coal
6.5 Recycle Oil Product (Stream 13) - see Table A-26. Coal/solvent
f-3\ (A)
ratios are 0.26 for lignitev ' and 0.31 for biturrrinousv '.
6.6 Make-up Water (Stream 6) - see Table A-27.
7.0 Intermediate Streams (see Figure A-9 )
7.1 Pretreated Coal (Stream 8) - Run No. Kg char/Kg coal
54 0.81
61 0.79 - 0.84
63 0.76 - 0.78
7.2 Product Gas Cyclone Slurry (Stream 9) - see Table A-28.
7.3 Gasifier Ash Slurry (Stream 12) - see Table A-29.
7.4 Pretreater Quench Water (Stream 7) - see Table A-30.
A-71
-------
7.5 Product Gas Quench Condensate CStreara 11) - see Table A-31.
7.6 Oil Stripper Bottoms (Stream 14) - see Table A-32.
8.0 Discharge Streams (see Figure A-9)
8.1 Quenched Product Gas (Stream 10) - see Table A-33.
8.2 Pretreater Offgas (Stream 5) - see Table A-34. Sulfur balance
data indicate that about 23-25% of input sulfur is released
from Illinois bituminous coal during pretreatment and exits as
(15)
gaseous sulfur compounds in the pretreater offgasv '.
8.3 Oil Stripper Vent Gas (Stream 15) - see Table A-35.
8.4 Oil Storage Vent Gas (Stream 18) - see Table A-36.
8.5 Separated Water (Stream 16) - no data available.
8.6 Separated Solids/Sludges (Stream 17) - no data available.
A-72
-------
TABLE A-25. HYGAS PILOT PLANT INPUT COAL (STREAM NO. 1 )*
Run No. (Reference)
Coal Type
Feed Rate, kg/kg
(Ibs/hr)
HHV - kcal/kg
(Btu/lb)
Size
Composition, wt % dry
C
H
0
N
S
Ash
Moisture
Volatile matter
Trace elements (ppm)
Fe
Ba
Mn
Ni
Zn
Li
Cr
Cu
Cd
Pb
Hg
Mo
B
Be
F
TV
V
37<3>
Lignite
1600(3530)
minus 8 mesh
61.3
4.2
20.1
0.97
0.88
12.6
15.8
40
«
--
~
~
54<«>
.
Bituminous
2687(5412)
7100(12780)
minus 8 mesh
73.9
5.0
8.3
1.5
2.76
8.5
2.2
35
~
~
--
~
--
--
58(5,8^T =
.
Subbituminous
2490(5478)
6272(11289)
minus 8 mesh
67.5
4.4
16.7
0.90
0.94
9.5
8.8
48
3300
750
150
2.6
22
14
6.2
7.3
.0.06
5.6
0.026
70
0.36
31
0.04
9
61<14>
Bituminous
6833(12300)
minus 8 mesh
59.6
4.93
9.22
1.32
4.32
10.6
5.8
37
~
--
~
-
--
63<14>
Bituminous
6833(12300)
minus 8 mesh
59.6
4.93
9.22
1.32
4.32
10.6
5.8
37
--
--
*Tables A-17 through A-27 contain data for one or more of the following runs at the Hygas Pilot Plant -
Nos. 37, 54, 58, 61, 63. All of the data in these tables represent steady state periods cooperation.
Runs were chosen for data presentation partially on the basis of data availability and partially to
represent the four coal types tested - lignite, subbituminous, medium sulfur bituminous, high suirur
bituminous.
"'Trace element data from Reference 8.
A-73
-------
TABLE A-26. COMPOSITION OF HYGAS RECYCLE/PRODUCT OIL (STREAM NO. 13)
Run No. (Reference)
Coal Type
Oil Composition (wt %}
Aliphatics
Olefins
Benzene
Toluene
Ethyl Benzene
1
C3-Cg Benzene
Xylenes
Indanes, Indenes and
Alkylindenes
Phenol
Cresols
C2 Phenols
Cs Phenols
Naphthalene
Methyl naphtha! enes
C2-C5 Naphthalenes
Biphenyl
Ci-C3 Biphenyls
Acenaphthenes
Fluorenes
Phenanthrene
Anthracene
Pyrene
Acetone
2-Butanone
Furans
Elemental Composition
Carbon
Hydrogen
Oxygen
37(7*)
Lignite
1.08
6.94
85.2
0.17
1.18
0.45
1.41
<.l
0.42
0.47
0.13
0.77
0.40
0.40
0.054
0.032
0.11
0.11
--
--
0.027
--
--
--
91.3^)
8.7
I
Sulfur i
54<4>
Bituminous
1
i
--
--
--
--
--
--
--
--
--
--
--
--
--
--
--
--
--
--
91.3<4>
8.7
56<7t>
Subbituminous
2.39
2.18
7.0
81.0
0.46
1.48
1.36
0.78
<.5
1.79
0.32
0.21
0.06
0.058
0.02
0.076
0.04
0.01
0.02
0.06
0.20
__
__
!
*Single sample July 3, 1975
^Composite for 37 hrs.
A-74
-------
TABLE A-27. CONSTITUENTS CONTRIBUTED BY MAKE-UP WATER*(15)
AT THE HYGAS PILOT PLANT
Constituents
Phenols
NH3
TOC
S=
CN"
SCN"
TDS
TSS
Cl"
Hexane Solubles
Run No.
60
0.004
0.03
0.24
0.00009
0.00003
0.03
3.76
0.36
0.87
0.12
63
0.34
0.31
0.74
0.0004
0.00005
0.010
2.31
0.13
0.35
0.06
*Units are in kg/10 kg of MAP pretreated char; concentration
data not available.
A-75
-------
TABLE A-28. HYGAS PRODUCT GAS CYCLONE SLURRY* (STREAM NO. 9)
Run No. (Reference)
Slurry Flow Rate -
I/kg coal (gal/lb)
Solids Flow Rate -
kg/ kg coal
Slurry Composition -
(mg/1 )
Phenols
NH3
TOC
S=
CN"
serf
IDS
TSS
Cl"
Hexane Solubles
PH
Trace Elements
(mg/1 )
Ca
Mg
Na
Ag
Al
B
Ba
Be
Cd
Co
Cr
Cu
Fe
Mn
Mo
Ni
Pb
Sn
Ti
V
Y
Zn
37<3>
1.4 (0.18)
0.27
~
--
-
~
--
~
-
--
-
~
--
--
--
--
~
-
~
~
--
--
--
-
-
-
--
54(4,7t)
-1 (.12)
0.081
.394
439
509
76
.004
31
432
15,700
«
--
8.0
--
-
--
~
-
--
~
--
--
-
-
-.
-
(*. Iff)
1.1 (0.14)
~
2455
257
1518
34
<.01
198
669
26,000
51
2190
7.1
18,900
12,800
60.600
<25
895
19,000
5,390
<2
28
<10
58
18
851
24
11
<50
<60
<100
131
<200
<10
142
61<15>
-
-. .
~
~
~
--
~
~
--
~
-
~
_..
--
63<15>
~
-
--
-
-
-
~
~
-
--
-
~
~
~
«
..
._
..
__
__
__
__
~
(continued)
A-76
-------
TABLE A-28. Continued.
Run No. (Reference)
Constituent
Production Rate
(kg/MAFs tonne)
Phenol s
NH3
TOC
S=
CN~
SCN"
IDS
TSS
Oil
Solids Composition
(wt*)
C
H
N
S
0
Ash
37<3>
~
--
~
~
~
49.6
3.3
5.6
0.7
0.5
41.5
54(4,7t)
0.43
0.47
0.50
0.06
3.2xlO"6
0.03
0.45
11
0.09
80.5
2.7
4.6
2.6
2.0
8.7
_
58(8t, 12t)
2.7
0.28
1.7
0.04
5xlO"5
0.19
0.72
36
2.3
--
--
61(15)
1.0
0.73
7.8
0.28
5xlO'6
0.16
1.2
43
0.4
~
63<15>
0.85
0.30
8.0
0.25
IxlO"5
0.16
2.7
28
7.7
--
--
*Data represent averages for each run.
Production rate data from 7 and 8.
'Moisture ash free coal basis for Runs 37, 54 and 58; Moisture ash free char basis for Runs 61 and 63.
fTrace element data from Ref. 12.
A-77
-------
TABLE A-29. HYGAS GASIFIER ASH SLURRY*-(STREAM NO. 12)
Run No. (Reference)
Slurry Flow Rate -
i/kg MAP* coal
(gal/lb)
Slurry Composition -
(rag/l)
Phenols
NH3
TOC
S=
CN"
SCN"
TDS
TSS
cr
Hexane Solubles
PH
Constituent
Production Rate
(kg/MAF tonne)
Phenol s
NH3
TOC
S=
CN"
SCN"
TDS
TSS
Oil
Solids Composition
(wt X)
C
H
0
N
S
Ash
37^
5.1 (.64)
0.06
5
243
<.01
<.001
4
815
4700
8
9.2
l.SxlO'4
0.026
1.25
--
~
~
~
~
~
0.002
1.0
6.7
0.08
8xlO"6
0.016
1.6
73
0.09
-
63^5)
--
~
~
~
0.26
0.37
1.4
0.15
2xlO*5
0.02
3.0
59
0.20
i
(continued)
A-78
-------
TABLE A-29. Continued.
Run No. (Reference)
Trace elements
(ppm)
Fe
Ba
Mn
Ni
Zn
Li
Cr
Cu
Cd
Pb
Mo
B
Be
F
Tl
V
Kg
37(l'7t)
--
~
--
--
54^4'7§)
--
~
~
--
"
-
58(6,et)
8,000
2,000
440
10
7.4
45 i
20
17
0.07
15
200
1
54
0.05 ;
26
0.0032
=====
61(15)
--
--
--
_
63'15)
..
..
__
--
--
--
-
-
--
--
--
*Data represent averages for each run
^Moisture ash free coal basis for Runs 37, 54 and 58; moisture ash free char basis for Runs 61 and 63
^Source of production rate data
Itrace element data
A-79
-------
TABLE A-30. HYGAS COAL PRETREATMENT QUENCH WATER* (STREAM NO. 7)
Run No. (Reference)
Flow Rate - l/MAFf kg coal
(gal/MAF Ib)
Composition (mg/i)
Phenols
NH3
TOC
S=
CN"
SCN"
TDS
TSS
Cl"
pH
Oil
Constituent Production
Rate (kg/MAFt tonne)
Phenols
NH3
TOC
S=
CN"
SCN"
TDS
TSS
Oil
»<»
--
--
1.2
0.3
2.1
5xlO"7
l.lxlO"5
1.5
12
2.8
8.1
59<«)
7.86 (0.97)
331
26
1206
1.7
0.003
316
4847
702
262
6.2
239
2.6
0.2
10
0.015
3xlO"5
2.5
38
5.5
1.9
61(15)
--
--
--
1.7
0.2
5.8
2.2xlO"4
5xlO"5
2.0
20
9.6
2.2
63<>5)
--
--
2.0
0.24
7.3
3xlO"4
4xlO"5
2.5
33
2.4
1.8
*Data represent averages for each run
char
A-80
-------
TABLE A-31. HYGAS PRODUCT GAS QUENCH CONDENSATE* (STREAM NO. 11)
Run No. (Reference)
Flow Rate - I/kg MAFf
coal (gal/lb)
Composition (mg/l)
Phenols
NH3
TOC
S=
ClT
SCN"
IDS
TSS
Cl"
Hexane Solubles
PH
Trace elements
(mg/l)
Ca
Mg
Na
Ag
Al
B
Ba
Be
Cd
Co
Cr
Cu
Fe
Mn
Ho
Ni
Pb
Sn
Ti
V
Y
Zn
Production Rate
(kg/MAFt tonne)
Phenols
NH3
TOC
S=
CN"
SCH"
TDS
TSS
Oil
37(l,7l)
1.1 (.14)
2100
3800
4000
125
<0.001
360
1750
32
--
8.3
~
--
-
--
--
--
-.
2.6
4.6
5.0
0.14
1.4xlO~6
0.44
2.2
0.12
0.08
:
54<7S>
--
1230
6550
1010
1080
<0.001
100
1090
36
--
7.8
--
--
--
~
--
--
--
--
--
--
1.1
6.0
0.9
0.7
2.5X10"6
0.09
1.0
0.025
0.013
1
57t(12)
-
--
4200
1600
4800
<13
150
12000
140
<1
<10
<5
<12
15
76
40
<5
<25
<30
<50
5
<100
<5
37
~
-
'
58(65,7,12*)
1.14 (.14)
4390
6048
3324
195
0.01
214
2089
1011
6.5
65
7.4
108000
28200
133000
<25
<200
251
211
<2
<20
<10
<24
<4
2270
206
<10
81
<60
<100
11
<200
<10
63
5.0
6.9
3.8
0.22
IxlO"5
0.25
2.3
1.1
0.08
61(15)
--
..
..
__
..
__
..
__
__
__
__
--
--
--
--
--
--
--
--
--
1.9
7.2
1.7
1.0
2xlO"5
0.2
2.7
1.8
0.021
_
63<15>
-
..
..
..
..
__
__
__
__
__
--
--
--
--
--
--
-
0.8
5.5
2.7
0.75
2xlO"4
0.15
2.2
11
0.11
*Data represent averages for each run
sRun 57 used subbituminous coal
.Source of production rate data .
Wloisture ash free coal basis for Runs 37, 54 and 58; moisture ash free char basis for Runs bi ana
fTrace element for Run 58 is from Ref. 12
A-81
-------
TABLF A-32. HYGAS OIL STRIPPER BOTTOMS* (STREAM NO. 14)
Run No. (Reference)
Flow Rate - l/iwf kg
(gal/lb)
Composition (rng/l )
Phenols
NH3
TOC
S=
CN"
SCN"
IDS
TSS
cr
Hexane Solubles
PH
Production Rate
(kg/MAFt tonne)
Phenols
NH3
TOC
S=
CN"
SCN"
IDS
TSS
Oil
3^,7*)
2.1 (.026)
1400
200
6000
<0.03
<0.003
330
950
41000
--
7000
9.9
3.0
0.42
7.4
__
--
~
--
1.2
0.37
5.7
0.04
2xlO"5
0.28
1.6
57
0.34
63^
--
~
--
1.1
0.75
6.7
0.11
2xlO"5
0.25
2.4
74
1.0
*Data are averages for each run
jSource of production data
(-Moisture ash free coal basis for Runs 37, 54 and 58; moisture ash free char
basis for Runs 61 and 63
A-82
-------
TABLE A-33. HYGAS QUENCHED PRODUCT GAS (STREAM NO. 13)
Run No. (Reference)
Production Rate -
Mm3/ kg coal
(scf/lb)
Composition
H2
co2
C2H6
N2 + Ar
H2S
CH4
CO
COS
cs2
RSH
NH3
HCN
HHV (kcal/Nm3)
Btu/scf
37*^
1.15
(18.9)
37.40
33.03
0.14
8.05
0.23
13.41
7.74
--
--
--
2370 (284)t
-' =
54*(15)
!
i . _
1.13
(19.1)
26.12
27.57
0.4
7.1
0.69
28.14
10.02
--
--
--
3410 (409)i
t==
58<15)
i
!
1.00
(17.0)
29.56
32.93
1.2
4.10
0
21.21
10.91
--
--
L
61(15)
I
1.23
(20.8)
i
! 33.22
27.99
j 0.51
i
7.82
j
1.71
19.16
9.59
--
--
-_
-.
63(15)
- -^
1.07
(18.2)
30.06
30.74
0.58
8.68
1.37
20.49
8.08
--
--
--
--
*Run 37 - 1506 hrs 7/5/75 to 0694 hrs 7/7/75; Run 54 - 0000 hrs 7/10/76 to
0000 hrs 7/11/76; Run 58 - 1400 hrs 11/14/76 to 0800 hrs 11/15/76; Run 61 -
1200 hrs 5/9/77 to 1600 hrs 5/9/77; Run 63 - 1400 hrs 6/24/77 to 1000 hrs
6/25/77.
Calculated from composition
A-83
-------
TABLE A-34. HYGAS PRETREATER OFFGAS (STREAM NO. 5)
(15)
Run No.
Flow Rate, Mm3/ kg'
(scf/lb) coal
Composition (vol %)
H2
co2
C2H6
02/Ar
N2
CH4
CO
H20
34*
1.38
(23.4)
0.20
4.96
0.03
1.98
4.27
0.55
1.48
48.1
6lt
1.19
(20.2)
0.37
3.75
0.05
2.44
43.3
0.19
3.32
46.6
63§
1.25
(21.2)
0.06
4.81
0.04
1.93
44.9
0.16
3.91
44.2
*1800 hrs 7/3/76 to 0030 hrs 7/7/76
f1300 hrs 5/8/77 to 2400 hrs 5/8/77
§
1400 hrs 6/24/77 to 1700 hrs 6/25/77
TABLE A-35. HYGAS OIL STRIPPER VENT GAS* (STREAM NO. 15)
Flow Rate - Nrn /kg coal (scf/lb)
(0.54)
Composition (vol %)
N2
CO
co2
H2
H2S
CH4
C2H6
C3H8
C4H10
cc+
5
20.0
0.7
58.1
10.9
0.1
8.0
0.64
0.23
0.17
A-84
-------
TABLE A-36. HYGAS OIL STORAGE VENT GAS* (STREAM NO. 18)
Run No. (Reference)
Composition (vol %}
N2
CO
co2
H2
H2S
CH4
C2H6
C3H8
C4H10
C5H12
C3H6
C4H8
C5H10
C6+
^^l^r^
5.3
1.4
67.3
9.0
0.6
10.9
1.1
0.44
0.43
0.12
0.14
0.4
0.17
-
*Steady state period 1506 hrs 7/5/75 to 0694 hrs 7/7/75.
A-85
-------
9.0 Data Gaps and Limitations
Data gaps and limitations relate primarily to composition and flow
rates of specific pilot plant streams during representative steady state
periods of operation. Major gaps include the following:
Data on the composition of product or recycle oil during
bituminous coal gasification are not available. Trace organic
sulfur and nitrogen constituents in Hygas oil(s) are not known.
t No data are available for sulfur compounds other than H2$, or for
HCN and NHs in quenched product gases from any of the coals
gasified at the pilot plant.
t For the major liquid discharge (or intermediate) pilot plant
streams, no data are available on trace organic substances.
No composition data are available for solids/sludges (Stream 17)
or water generated by the Edens separator and associated filtration
operations.
Although, as presented above, some data are available on the
characteristics of a number of input and discharge streams of the
subject process, the available data are not comprehensive in that
not all streams are addressed and not all potential pollutants and
toxiciological and ecological properties are identified. An
environmental data acquisition effort which would lead to the
generation of the needed data corresponds to the EPA's phased level
approach to multimedia environmental sampling and analysis(13).
10. Related Programs
As part of the environmental assessment of high-Btu coal gasifica-
tion processes, DOE has contracted with the Institute of Gas Technology
for sampling and analysis at the Hygas plant. Data will be generated for
pilot plant operations at least through the calendar year 1978. Reports
prepared by IGT and by Carnegie-Mellon University are expected to
contain additional data on intermediate/discharge streams.
REFERENCES
1. M. J. Massey, R. W. Dunlap, et al, Characterization of Effluents from the
Hygas and ^-Acceptor Pilot Plant, an interim report for the period
July-September 1976 prepared by Carnegie-Mellon University, ERDA FE-2496-1,
November 1976.
A-86
-------
2. Handbook of Gasifiers and Gas Treatment Systems, a final reoort hv
Dravo Corp. to ERDA, ERDA Document No. FE-1772-11, February 1976
3. Pipeline Gas from Coal-Hydrogasification (IGT Hydrogasification Pror«O
Interim Report No. 2, July 1974-June 1975, ERIV i Document Nc f. FE-1221-
1 .7 / D -
4. Bernard S Lee, Current Development of the Hygas Program, presentation
to the Eighth Synthetic Pipeline Gas Symposium, October 18-20, 1976.
5. Bernard S. Lee, Hygas Process Achieves 92% Coal Conversion, Oil and Gas
Journal, August 1, 1977.
6. Anastasia, L.J., Environmental Assessment of the Hygas Process, Monthly
Report for the Period Dec. 1 to Dec 31, 1976, ERDA Document No
FE-2433-7, February 1977.
7. Anastasia, L.J., Environmental Assessment of the Hygas Process, Quarterly
Progress Report No. 2, Oct. 1 to Dec. 31, 1976, ERDA Document No.
FE-2433-8, May 1977.
8. Anastasia, L.J., Environmental Assessment of the Hygas Process, Monthly
Report for the Period April 1 to April 30, 1977, ERDA Document No.
FE-2433-13, August 1977.
9. Anastasia, L.J., Environmental Assessment of the Hygas Process, Monthly
Report for the Period March 1 to March 31, 1977, ERDA Document No.
FE-2433-11.
10. Anastasia, L.J., Environmental Assessment of the Hygas Process, Monthly
Report for the Period Jan. 1 to March 31, 1977, ERDA Document No.
FE-2433-12, August 1977.
11'. Detman, R.F., Preliminary Economic Comparison of Six Processes for
Pipeline Gas from Coal, 8th Synthetic Pipeline Gas Symposium, Chicago,
111., October 18-20, 1976.
12. Massey, J.M., et al, Environmental Assessment in the ERDA Coal Gasifi-
cation Development Program Progress Report for the Period July 1976 to
December 1976, Carnegie-Mellon University, ERDA Document No. FE-2496-6,
March 1977.
13. Dorsey, J.A., and Johnson, L.D., Environmental Assessment Sampling
and Analysis: Phased Approach and Techniques for Level 1, EPA-600/2-
77-115, June 1977.
14. Pipeline Gas from Coal-Hydrogenation (IGT Hydrogasification Process)
No. FE-2434-23, April 1978.
15. Anastasia, L.J., Environmental Assessment of the Hygas Process, Quarterly
Progress Report No. 5 for the period July 1 to September 30, is//,
DOE No. FE-2433-20, March 1978.
A-87
-------
C02 - ACCEPTOR PROCESS
1.0 General Information
1.1 Operating Principles - Moderate pressure gasification of coals
in a bed fluidized by steam. Hot calcined limestone or dolomite
Cthe "acceptor") is injected into the top of the gasifier to
provide heat for gasification and to absorb carbon dioxide
(CaO + C02 = CaCO, + A). Carbonated acceptor and gasifier char
are transferred from the gasifier to a regenerator vessel where
residual carbon is burned in a bed fluidized by air and the
acceptor is regenerated (CaC03 + A = CaO + C02).
1.2 Development Status^6'10^ - The C02 - Acceptor process has been
under development by Consolidation Coal Company (CONSOL, now
CONOCO) since the early 1960's. Under joint sponsorship by ERDA
and the American Gas Association (AGA), a 36 tonne/day (40 ton/
day) pilot plant was constructed in Rapid City, South Dakota in
1971. Since 1972, more than 42 runs have been made under the
direction of CONOCO, using both North Dakota lignites and a
Montana subbituminous coal. Two types of acceptors (Ohio dolo-
mite and South Dakota limestone) have also been tested. During
1975, methanation of product gas was successfully demonstrated.
Pilot plant operations demonstrated the technical feasibility of
the C02-Acceptor process for gasification of lignite (based on
AGA/ERDA technical feasibility criteria).
The pilot plant was closed down in the fall of 1977. (This
facility has been modified for testing the Westinghouse gasifi-
cation process.) CONOCO has prepared conceptual designs for a
commercial plant based on the C02 - Acceptor process, although
no commercial plant is currently planned.
A-J
-------
1.3 Licensor/Developer - CONOCO Coal Development Company
Research Division
Library, Pennsylvania 15129
1.4 Commercial Applications - None.
2.0 Process Information
2.1 Pilot Plant Csee Figure A-ll , Flow Diagram)
2.1.1 Gasifier Csee Figure A-12)
Equipment'- '
Construction: vertical, cylindrical pressure vessel flanged at both
ends. The pressure bearing wall is low carbon alloy steel with a
stainless steel liner. Vessel is water jacketed and contains an
internal cyclone at product gas exit.
t Dimensions v J: 100cm (40 in) I.D. width
166cm (77 in) O.D. width
21mm (70 ft) total height
13mm (42 ft) bed height
Bed type and gas flow: Gasifier coal bed is fluidized by steam
and recycle gas (during startup) to prevent acceptor agglomeration.
Product gas flows counter-currently to acceptor solids and exits
the gasifier at the top through an internal cyclone.
Heat transfer and cooling: Hot calcined acceptor provides heat
(sensible as well as chemical heat via exothermic reaction with
C02 in the gasifier) for the gasification. Heat transfer is via
direct gas-solids contact. Water jacket provides gasifier cooling.
Coal feeding: Dried and preheated coal is transferred by gravity
to the gasifier via a lockhopper system pressurized with recycle
gas. (In the commercial plant C02 or flue gas would most likely be
used.)
Gasification media introduction: Steam for gasification is intro-
duced at the lower portion of the fluidized bed near the point of
coal injection. Steam is also injected at the bottom of the gasifier
to strip char from the carbonated acceptor. Calcined acceptor is
injected at the top of the gasifier and showers through the
fluidized coal bed. Carbonated acceptor collects in a narrow throat
at the bottom of the gasifier.
A-89
-------
V
^
_, CYCLONE
QUENCH
TOWER
*
1
so2
SCRUBBER
1. Feed Coal 14.
1. Raw Product Gas 15.
3. Feed Acceptor 16.
4. Quenched Product Gas (to flare) 17.
5. Gasifier Char Slowdown 18.
6. Recarbonated Acceptor 19.
7. Recycle Gas 20.
3. Steam 21.
9. Air 22.
10. Carbon Dioxide 23.
11. Reject Acceptor (may be withdrawn 24.
from either the gasifier or regenerator) 25.
12. Calcined Acceptor 26.
13. Regenerator flue gas (to atmosphere or
to recycle)
Product Sas Quench Water Slowdown
Ash from Flue Gas Cyclone
Pond Solids
Flue Gas Quench Water
SO? Scrubber Slowdown
Asn Slurry Tank Offgas
Ash Slurry Tank Effluent
Venturi Scrubber Slowdown
Drier/preheater Flue Gas
Depressurization Offgas (to flare)
S02 Scrubber Offgas (to flare)
Pond Effluent (to city sewer)
Product Gas Quench Solids
Figure A-11. C02- Acceptor Pilot Plant Flow Diagram
(4)
A-90
-------
PRODUCT GAS
HOTACCEPTOR
3"
CHAR
REFRACTORY
STEAM
en REJECT
*«. ACCEPTOn
J
STEAM
^CARBONATED ACCEPTOR
WATER OR STEAM M
(3)
figure A-12. C02-Acceptor Gastfier
A-91
-------
Char and acceptor removal: Residual char Is withdrawn from the
gasifier near the top of the fluidized bed and is fed to the
regenerator as fuel. A small fraction of the char also leaves the
gasifier as overhead loss through the internal cyclone. Carbonated
acceptor collects and is withdrawn at the bottom of the gasifierUJ.
Special features:
Flow and pressure control - Solids flow through the three standlegs
(transfer lines) between the gasifier and the regenerator (see
Figure A-19) is controlled by butterfly valves. Seals between the
two vessels are maintained by careful recycle gas pressure control.
Operating Parameters
Gas outlet temperature: 1090°K (1500°F)
Bed temperature^: 1090°K (1500°F)
Pressure^: 1 MPa (10 atm)
Coal residence time in gasifier (calculated from data in
Reference 1): 250 rain
Gas residence time in gasifier^ ': -30 sec in fluidized bed
Acceptor circulation rate^ ': 3.6 kg/kg coal (based on Velva Lignite
from N. Dakota and Tymochetee dolomite from Ohio)
Raw Material Requirements
Coal feedstock:
Type -non-caking coals (lignite, subbituminous)
SizeU _ +100 - 8 mesh (.147-2.3 mm)
RateU): .004 kg carbon/kg carbon in bed/min, pilot plant can
gasify about 1136 kg (2500 Ibs) coal/hour
Coal pretreatment: sized coal is preheated to 500°F for drying to
less than 5% moisture. Natural gas is used as fuel for preheating
at the pilot plant.
Steam requirements to gasifiert1'2); l.l kg/kg dry coal (lignite)
Recycle gas to gasifier^'2): .4-.5 Nm3/kg (7-3 scf/lb coal)
(required to maintain steam partial pressure less than 1.3 MPa
[13 atm]).
A-92
-------
Utility Requirements
Water: ?
Electricity: ?
Boiler feed water: ?
Process Efficiency
Cold gas efficiency (2) = (product gas energy output/ coal energy
input x 1.00): 77% (dry lignite @ 6290 kcal/kg or 11.350 Btu/lb)
Overall thermal efficiency: ? (Pilot plant used natural gas to
dry and preheat lignite. Sensible heat was not recovered from
either gasifier product gas or regenerator flue gas as would be the
case in a commercial facility.)
Expected Turndown Ratio
Little actual data are available from pilot operations. Fluidized-
bed operation and stringent pressure control requirements limit the
turndown which can be accomplished without system failure.
Gas Production
1.35 Nm3/kg (23 scf/lb) based on lignite or 3135 Nm3/hr/m3
(10840 scf/hr/ft2} bed area
2.1.2 Regenerator
Equipment^ '
Construction: Vertical, cylindrical pressure vessel with refractory
lining and a ring type air distributor. An external cyclone is
provided for the regenerator.
Dimensions: ?
Bed type and gas flow: Acceptor bed is fluidized by air and recycled
flue gas. Char introduced into the lower section of the regenerator
is burned and the resulting ash becomes entrained in the regenerator
flue gas.
t
Heat transfer and cooling: Heat of char combustion TS utilized to
reverse the acceptor reaction and raise the temperature or tne
acceptor for recycle to the gasifier. Heat transfer is via direct
gas-solids contact.
A-93
-------
t Ash removal: Ash leaves the regenerator with, flue gas and is
collected by an external cyclone and wet S02 scrubbing system.
Acceptor removal: Regenerated acceptor is withdrawn near the top
of the fluidized bed for return to the gasifier. Reject acceptor
is withdrawn from the bottom of the regenerator.*
Operating Parameters^ '
Bed temperature: 1280°K (1860°F)
Pressure: 1 MPa (10 atm)
Acceptor residence time: ?
Raw Material Requirements
Air^1'2^: 2.35 kg air/kg lignite feed to gasifier
Acceptor make-up^- ': 0.24 kg/kg coal to gasifier (based on
Tymochetee dolomite from Ohio). Start-up requires burned limestone
or dolomite to prevent agglomeration. After start-up, make-up
limestone is added directly to the regenerator.
Flue Gas Production Rate
3.7 Nm3/kg lignite (62 scf/lb). (In the pilot plant the flue gas
is vented to the atmosphere after passing through a cyclone and S02
scrubbing system.)
Special Features - Partial Oxidation of Sulfur in the Regenerator
Calcium sulfide associated with carbonated acceptor and sulfur
contained in char are partially oxidized to S02 (and other trace gaseous
sulfur species - S03, COS, CS2, S2) in the regenerator. S02/S03 then
partially react with CaO to form calcium sulfate and calcium sulfite.
*In early pilot plant runs, the reject acceptor was withdrawn from the
gasifier. Withdrawal of the reject from the regenerator bottom is preferred
fnr nnorat-irmal vaacnnrlRl ?*>
j ***« i v> >*iviiviiMinrui \j \ 1*11
for operational reasons(6).
A-94
-------
In order to minimize the formation of deposits resulting from low
melting CaS/CaSOs/CaSQ4 mixtures, air to the regenerator is kept some-
what below stoichiometric. This results in conversion of CaSO /CaSO
to CaS by the reactions:
4CO H- CaS04 B 4C02 + CaS
SCO + CaS03 = 3C02 + CaS
The net result is that sulfur leaves the regenerator mainly as CaS with
reject acceptor or with cyclone ash. The remainder of the sulfur leaves
as gaseous species (mainly S02) in the flue gas.
2.1.3 Ash Slurry Tanlr ': The major purpose of the ash slurry opera-
tion is to minimize the potential for oxidation and/or leaching
of reduced sulfur after disposal of spent acceptor/ash. Ash
from the flue gas cyclone and reject dolomite from the gasifier
(or regenerator) are slurried with blowdown waters from product
and flue gas quenching and from the SCL scrubber. The resulting
slurry is transferred to an ash tank where COp is added and hLS
is released:
CaS/CaO + C02 + H20 = CaC03 + H2S
At the pilot plant, the H2S stream is sent to a flare; the
ash/carbonated lime sludge, to a holding pond.
2.1.4 Coal Feed/Pretreatment^1'2': Coal is crushed and screened to
+100 - 8 mesh and dried at about 533°K (500°F) in a hot gas
sweep mill to less than 5% moisture. Coal is fed to the gasi-
fier via lockhoppers pressurized with inert gas. (In commercial
applications, recycle product gas or flue gas from the regen-
erator will probably be used.)
A-95
-------
2.1.5 Quench and Dust Removal;
Product gas^: Gasifier has an internal cyclone and a
venturi scrubber/quench tower system. Make-up water require-
ments for the product gas scrubbing system at the pilot plant
are about 8 I/kg coal (1 gal/lb coal).
Regenerator flue gas: Flue gas cleaning system consists of
a cyclone, a quench tower, and $03 scrubber at the pilot
plant.
2.2 Conceptual Commercial Scale Design - The U.S. Bureau of Mines has
prepared two conceptual designs for 7 million Mm /day (250 million
SCF/day) commercial plants using Montana sub-bituminous coal and
North Dakota lignite^. The gasification and regeneration systems
in these designs are essentially scaled-up versions of the pilot
plant design. C.F. Braun (under ERDA contract) and CONSOL have
also prepared conceptual designs of commercial C02 - Acceptor
plants, but these designs are not publicly available at present.
3.0 Process Economics
Based on the U.S. Bureau of Mines conceptual designs of 7 million
Nm3/day (250 million SCF/day) C02 - Acceptor plants^4), the following
capital costs are estimated for the gasification/regeneration systems:
Capital Cost Percent of
Feed Coal (July 1975 dollars) Overall Plant Cost
Sub-bituminous Coal $149,000,000 28
Lignite $118,000,000 26
C.F. Braun and Company has also prepared economic estimates for a
7 million Nm3/day (250 million scf/day) conceptual commercial scale
C02 - Acceptor plantW. The gasification/regeneration system,
including power recovery and raw gas quench, are estimated to cost
$257,000,000 in 1976 dollars, or 29% of total plant investment.
4.0 Process Advantages
Production of medium heating value gas without the use of oxygen.
t Raw product gas contains essentially no hydrocarbons other than
methane. Tar/oil separation ^teps are not necessary in raw gas
processing.
A-96
-------
. Product gas has a H2/CO ratio greater than three. All of the en
shift! 2 "" methanated without P>ior water g"as
Process results in nearly complete coal carbon utilization Ash
removed from the regenerator contains less than one percent of the
carbon in the feed coal.
Carbon dioxide and hydrogen sulfide are partially removed from
product gas in the gasifier by reaction with the acceptor This
minimizes the requirement for removal of acid gases from the product
gas.
Compared to other processes which circulate solids for heat transfer,
solids circulation rate is low since the acceptor provides both
sensible and chemical heat for the endothermic coal gasification
reactions.
5.0 Process Limitations
Process is limited to noncaking coals (e.g., lignite and sub-
bituminous coals.)
0 Process is mechanically complicated, requiring careful pressure
balance control to maintain simultaneous circulation of several
solids streams.
Acceptor make-up requirements amount to about 0.25 kg/kg of lignite.
Unless reject acceptor can be reconstituted, the make-up require-
ment represents a large operational cost. (Natural deposits of
suitable acceptors are generally not located near lignite deposits.)
t Regenerator offgas contains both sensible and chemical energy which
is lost in pilot plant operation. In a commercial facility, this
gas might be incinerated and expanded through a turbine to capture
this energy. If not used, this energy represents a significant
thermal penalty.
t Raw product gas contains fine char particulates which are not
collected by the internal gasifier cyclone. This material amounts
to about 4% of the input carbon and would represent a thermal
penalty for a commercial plant unless the char is salvaged from the
quench system and returned to the regenerator.
6.0 Input Streams (see Figure A-ll)
6.1 Coal (Stream 1) - see Table A-37
6.2 Acceptor (Stream 3) - see Table A-38
6.3 Stream (Stream 8) - see Section 2.1.1
A-97
-------
6.4 Air (Stream 9) - see Section 2.1.2
6.5 C02 (Stream 10} - no data available (intermittent operation)
7.0 Intermediate/Discharge Streams
7.1 Gaseous Stream (see Figure A-ll)
7.1.1 Product gas (Stream 4) - see Table A-39
7.1.2 Regenerator flue gas (Stream 13} - see Table A-40
7.1.3 Ash slurry tank offgas (Stream 19) - no data available
7.1.4 Drier/preheater flue gas (Stream 22) - no data available
7.1.5 Organic waste tank depressurization offgas (Stream 23) -
no data available
7.1.6 S02 scrubber offgas (Stream 24) - no data available
7.2 Liquid Streams (see Figure A-19)
7.2.1 Preheater venturi scrubber water (Stream 21) - see Table A-41
7.2.2 Product gas quench water blowdown (Stream 14) - see Table A-42
7.2.3 Flue gas quench water blowdown (Stream 17) - see Table A-43 and
A-37
7.2.4 S02 scrubber water blowdown (Stream 18) - see Table A-45
7.2.5 Ash slurry tank effluent (Stream 20) - see Table A-46
7.2.6 Holding pond effluent (Stream 25) - see Table A-47
7.3 Solids Streams (see Figure A-19)
7.3.1 Gasifier char blowdown (Stream 5) - see Table A-48
7.3.2 Product gas quench solids (Stream 26) - see Table A-49
7.3.3 Reject acceptor (Stream 11) - see Table A-50
7.3.4 Ash from flue gas cyclone (Stream 15) - see Table A-51
7.3.5 Pond solids (Stream 16) - see Table A-52
A-98
-------
TABLE A-37. ^-ACCEPTOR PILOT PLANT INPUT COALS (STREAM NO.
Run No. (Reference)
Coal Type1"
Dry HHV- kcal/kg
(Btu/lb)
Size (mesh)
Feed Rate kg/hr
Obs/hr)
Composition %
Moisture (as
received)
Carbon
Hydrogen
Oxygen
Su.1 fur
Nitrogen
Ash
Trace Elements (ppm)
Sb
As
Be
Cd
Cr
Pb
Hg
Ni
Se
Te
V
....-
21(D
Velva
Lignite
5975
(11190)
+100 - 8
1136
(2500)
40
66.2
4.5
21.6
0.7
0.8
6.2
.
26(2)
Velva
Lignite
6066
(11360)
+100 - 8
1134
(2450)
40
66.2
4.6
20.7
0.54
1.0
6.8
1
27c(5>
Velva
Lignite
--
--
--
--
--
5
<10
0.11
<0.4
1.4
=====
28b(5>
Velva
Lignite
_
__
__
..
2
40
1
30
10
10
<0.4
<0.1
<50
33b(5>
VeTv^
Lignite
..
__
__
__
..
-
39(8)
Glenharold
Lignite
-
__
__
The runs listed in this table are those which had steady state periods and/or for which specific stream
composition data are available.
Velva lignite is a low sodium coal; Glenharold lignite is a high sodium coal.
A-99
-------
TABLE A-38. C02-ACCEPTOR PILOT PLANT ACCEPTOR (STREAM NO. 3)
Run No. (Reference)
Makeup Feed Rate,
kg/kg coal
Composition (%)
MgO
CaO
co2
Inert
Trace Elements
(ppm)
Sb
As
Be
Cd
Cr
Pb
Hg
Ni
Se
Te
V
21d)
0.24
19.8
29.5
44.8
5.8
--
--
--
--
--
26<2>
0.24
22.4
27.0
45.7
4.9
--
--
--
--
--
--
27C<5>
_
3
<10
2
15
25
--
16
< 0.4
< 0.1
<50
28b(5>
_ _
--
--
6
--
--
34
0.08
0.4
1.4
33b<5>'
__
--
1.6
1.6
<5
4.0
10
<0.5
<0.05
30
0.3
0.2
55
A-TOO
-------
TABLE A-39. C02-ACCEPTOR PILOT PLANT PRODUCT GAS (STREAM NO. 4)
Run No. (Reference)
Production Rate,
Nm3/kg coal (SCF/lb)
Temperature After
Quench (°C)
HHV kcal/Nm3 (Btu/SCF)
Composition (Vol. %, dry)
CO
H2
CH4
co2
N2 + Ar
°2
H2S
COS
cs2
RSH
c2+
NH3
HCN
Parti cul ate Matter
(mg/Nm3)
H£0
Z1u>
1.34 (22.7)
--
3160 (375)f
15.45
55.98
14.14
10.88
3.0
--
0.132
--
--
0.01
0.427*
--
--
34.0*
26<2>
1.29 (21.8)
3200 (380)1"
15.47
58.8
13.75
9.08
2.91
--
0.12
__
0.01
0.69*
--
31*
33b<7>
--
--
--
--
0.04 -0.09
0.0015-0.004
--
--
~ "
-
"Composition before quench, other data represent quenched gas.
Calculated from composition data.
A-101
-------
TABLE A-40. C02-ACCEPTOR REGENERATOR
FLUE GAS (STREAM NO. 13)
Run No. (Reference)
0
Flow Rate, Mm /kg coal
(SCF/lb)
Composition*
CO
co2
H2
N2 + Ar
H20
SO 2
s2
H2S
COS
NH3
Parti culates (mg/Nm )
210)
3.7 (62)
1.98 Vol.%
28.32 Vol.%
0.04 Vol.%
68.66 Vol.%
0.98 Vol.%
121 ppmv
3 ppmv
28 ppmv
46 ppmv
--
26<2>
__
2.2 Vol.%
29.35 Vol.%
0.06 Vol.%
68.40 Vol.%
1.3 Vol.%
92 ppmv
--
39 ppmv
46 ppmv
33b<5)
__
--
55-320 ppmv
95-150 ppmv
--
After S02 scrubber
A-102
-------
TABLE A-41. C02-ACCEPTOR PREHEATER VENTURI SCRUBBER WATER (STREAM NO. 21)
Run No. (Reference)
Flow Rate, I/kg coal (gal/lb)
Composition*
COD
TDS
TSS
NH3
S=
so3=
so4=
Cyani des
Phenols
Total Alkalinity1"
Total Hardness
N03~
Total P04= .
cr
Ca++
M ++
Mg
Na+
K+
Fe(+2 & +3)
Mn++
PH
27c<5>
8 (1)
30
590
--
1.8
0.23
0.68
310
<.02
<.004
48
329
0.1
7.9
8.3
83
29
14
14
0.27
0.13
7.0
28b<5>
8 (1)
10
588
5110
3.01
294
0.4
300
<.02
0.027
350
.03
4.5
7
83
36
18
6
.03
.03
6.8
33b(5>
8 (1)
272
950
2684
162
1.96
4.1
38
<.02
<.001
714
290
0.53
2.9
2
55
37
85
94
0.2
<.05
8.0
(continued)
A-103
-------
TABLE A-41. Continued
Run No. (Reference)
Trace El ements i n Fi 1 tered Water
Sb
As
Be
Cd
Cr
Pb
Hg
Ni
Se
Te
V
Trace Elements in Scrubber
Water*
Sb
As
Be
Cd
Cr
Pb
Hg
Ni
Se
Te
27c<5>
0.4
0.009
<0.005
<0.005
0.1
<0.01
--
<0.1
<0.005
<0.002
<2
1
8
1.8
10
<10
0.4
15
1
0.2
28b<5)
0.5
<0.01
<0.01
<0.01
0.1
<0.05
0.1
<0.005
<0.005
<2
0.8
3
<10
<1
10
<10
0.3
10
0.8
<-l
33b(5>
0.001
:0.015
<0.02
<0.02
0.1
<0.05
0.001
<0.01
0.003
0.001
<1
2
4
<10
<2
10
2
13
5
2
<-l
t
mg/1 except pH
as CaCO,
A-104
-------
TABLE A-42. C0?-ACCEPTOR PRODUCT GAS QUENCH WATER SLOWDOWN (STREAM NO. 14)
Run No. (Reference)
Flow Rate, I/ kg coal
(gal/lb)
Composition*
COD
TDS
TSS1"
NH3
S=
S03=
504=
Cyanides
Phenols
Total Alkalinity^
Total Hardnesst
N03-
Total P04 =
ci-
Ca++
Mg++
Na+
K+
Fe(+2 & +3)
Mn++
SCN-
PH
27c(5>
8 (1)
120
1210
4825
1180
0.01
33.3
58
0.02
0.004
5520
1998
0.03
11.4
56
30
478
18
17
0.18
0.13
8.2
28b^
8 (1)
100
608
1290
1505
1.19
10.7
56
0.02
0.05
186
400
0.07
3.1
70
19
93
15
18
0.01
0.03
--
8.7
33b<5)
8 (1)
300
426
2000
1250
84
12.7
335
0.001
5540
240
3.8
42
39
35
15
9
0.03
0.05
7.7
39 (8)
1330
60
350
77
1.6
0.003
11
31
__
--
--
--
-
5.6
If
mg/1 except pH
see Table A-42 for composition of solids
CaCOa
A-105
-------
TABLE A-43. C09-ACCEPTOR F:LUE GAS QUENCH WATER SLOWDOWN (STREAM NO. 17)
Run No. (Reference)
Flow Rate, I/kg coal (gal/lb)
Composition*
COD
TDS
TSS
NH3
S=
S03=
S04=
Cyanides
Phenols
Total Alkalinity1'
Total Hardness^
N03-
Total P04 =
CT
Ca++
Mg++
Na+
K+
Fe(+2 & +3)
Mn++
PH
27c^
6 (.75)
30
1300
150
37.9
<0.01
5.4
585
<0.02
< 0.004
154
592
0.05
12.1
48
141
58
95
90
0.12
0.25
6.6
28b<5>
6 (.75)
5
912
119
33
0.04
6.65
282
<0.02
<0.004
__
453
0.01
4.2
63
140
26
95
6
0.01
0.05
8.0
33b<5)
6 (.75)
70
1098
1630
292
3.24
43.8
685
0.02
--
1502
636
6.2
1.2
1.5
134
73
57
0.7
0.03
<0.05
7.2
mg/1 except pH
'"as CaC00
A-106
-------
TABLE A-44. SOLIDS IN C02-ACCEPTOR FLUE GAS QUENCH WATER
(STREAM NO. 17)
SLOWDOWN
Run No. (Reference)
Flow Rate, kg/ kg coal
Composition (wt %}
H
C
N
0
S
Ash
CaS-CaO
co2
Trace Elements (ppm)
Sb
As
Be
Cd
Cr
Pb
Hg
Ni
Se
Te
V
27c<5>
0.0009*
--
--
0.4
--
140
50
0.4
28b(5>
0.0007*
--
--
--
--
5.3
176
10
70
160
500
--
50
5.3
1,4
102
33b(5>
0.00978*
--
--
--
2
23
10
2
70
19
0.06
65
2
0.7
26
Calculated from TSS and flow rate in Table A-36.
A-107
-------
TABLE A-45. C02-ACCEPTOR S02 SCRUBBER WATER SLOWDOWN (STREAM NO. 18)
Run No. (Reference)
Flow Rate, I/kg coal (gal/lb)
Composition*
COD
TDS
TSS
NH3
S=
so,=
J
so4=
Cyanides
Phenols
Total Alkalinity1"
Total Hardness
N03
Total P04=
Cl"
Ca**
Mg++
Na+
K+
Fe(+2 & +3)
Mn++
PH
27c^
0.4 (.05)
120
5678
26
10
10.8
57.9
395
0.02
0.004
284
555
0.6
9.3
95
47
107
45
45
0.16
0.30
7.4
28b(5>
0.4 (.05)
77292
--
11.5
2.17
99.2
172
--
--
34800
3400
0.67
11.3
48
642
437
35
32800
0.13
0.05
7.9
33b(5>
0.4 (.05)
55
6704
49
95.8
3.96
125
246
0.02
0.004
__
248
0.02
0.8
3.0
190
61
23
50
190
0.1
7.9
T
mg/1 except pH
as CaCOo
A-108
-------
TABLE A-46. (^-ACCEPTOR ASH SLURRY TANK EFFLUENT (STREAM NO. 20)
Run No. (Reference)
Flow Rate
Composition*
COD
TDS
TSS
NH3
S=
S03=
so4=
Cyanides
Phenols
Total Alkalinity1"
Total Hardness
N03-
Total P04=
cr
Ca++
Mg++
Na+
K+
Fe(+2 & +3)
Mn++
~_ PH
27c(5)
Intermittent
Operation
290
6244
67980
18
230
102.5
195
<.02
<.004
4630
6440
0.0
3.7
235
1323
763
425
495
1.3
<.05
12.2
28b(5>
Intermittent
Operation
135
2772
4478
200
18.4
8.56
60
<.02
<.004
--
2228
0.3
0.8
400
922
0
205
115
0.0
0.02
12.3
mg/1 except pH
as CaCOo
A-109
-------
TABLE A-47.
- ACCEPTOR HOLDING POND WATER AND OUTFLOW (STREAM NO. 25)
Run No. (Reference)
Pond Outflow*
I/kg coal (gal/lb)
Parameters
TSS (mg/1)
BOD (mg/1 as 02)
Pond Water*
Composition!
COD
TDS
TSS
NH3
S=
so3=
so4=
Cyanides
Phenols
Total Alkalinity1^
Total Hardness
N03-
Total P04=
ci-
Ca++
Mg++
Na+
K+
Fe(+2 & +3)
Mn++
PH
27c(5)
24-35(3-4)
20-90
2.5-3.5
15
1038
257
242
<0.01
<0.01
315
<0.02
<0.004
900
1036
7.0
12.5
17
83
201
95
100
0.10
0.13
7.8
28b(5)
2-30(3-4)
5-200
8-33
15
2988
1120
277
0.48
4.13
238
<0.02
0.017
182
129
0.03
3.3
208
32
12
130
36
0.01
<0.02
8.9
33b(5)
8-21(1-3)
70-375
5-24
35
1002
296
<0.03
<0.01
380
<0.02
< 0.001
1030
259
4.6
2.5
60
26
89
108
0.09
<0.05
7.9
f
The exact location of these samples is unknown.
as
T
mg/1 except pH
A-110
-------
TABLE A-48. COg-ACCEPTOR GASIFIER CHAR SLOWDOWN (STREAM NO. 5)
*- , , =
Run No. (Reference)
Flow Rate, kg/ kg coal
Composition (wt %)
H
C
N
0
S
Ash
CaS-CaO
C09
c
Trace Elements (ppm)
Sb
As
Be
Cd
Cr
Pb
Hg
Ni
Se
Te
V
26<"
0.03
0.83
73.29
0.36
0.00
0.00
18.70
0.07
6.75
--
--
--
--
--
--
--
27c<5>
__
--
--
--
_ _
--
21
--
--
25
0.05
0.4
1.4
__ ,
2*M
--
^ _
--
--
__
--
3
<10
1
40
20
28
0.4
<-l
<50
33bW
M_
--
--
--
0.2
19
5
12
30
13
0.62
20
4
0.1
7
A-lll
-------
TABLE A-49. SOLIDS IN C02-ACCEPTOR PRODUCT GAS QUENCH WATER (STREAM NO. 26}
Run No. (Reference)
Flow Rate kg/kg coal
Composition (wt %)
H
C
N
0
S
Ash
CaS-CaO
co2
Trace Elements
(ppm)
Sb
As
Be
Cd
Cr
Pb
Hg
Ni
Se
Te
V
210)
0.024
0.39
68.0
0.24
0.86
0.66
29.85
--
--
--
--
--
--
--
--
--
26<2>
0.0457
0.78
59.1
0.22
0.00
0.00
33.05
0.49
6.36
--
--
--
__
--
__
--
27c<5>
0.039*
--
--
--
--
0.4
20
--
4.0
20
40
0.09
45
0.4
<0.1
--
28b<5>
0.010*
--
--
--
--
--
--
--
--
--
__
--
33b<5>
0.016*
..
--
--
--
4
19
<5
12
30
14
0.62
20
4
<0.1
7
Calculated from TSS and flow in Table A-35.
A-112
-------
TABLE A-50. COg-ACCEPTOR REJECT ACCEPTOR (STREAM NO. 11)*
Run No. (Reference)
Flow Rate, kg/kg coal
Composition (wt %}
MgO
CaO
Inert
co2
CaS-CaO
Trace Elements (ppm)
Sb
As
Be
Cd
Cr
Pb
Hg
Ni
Se
Te
V
21d)
0.15
M
17
16
0.14
--
--
--
--
--
26^)
0.13
36.02
43.42
7.93
12.61
0.03
--
--
27c(5)
v
0.2
18
1.8
40
14
0.1
55
0.2
<0.1
A
Withdrawn from gasifier in these runs,
A-H3
-------
TABLE A-51. ASH FROM C02-ACCEPTOR FLUE GAS CYCLONE (STREAM NO. 15)
Run No. (Reference)
Flow Rate kg/kg coal
Composition (wt %)
H
C
N
S
Ash (Oxides)
CaS-CaO
Trace Elements
(ppm)
Sb
As
Be
Cd
Cr
Pb
Hg
Ni
Se
Te
V
21")
0.058
7.55
--
5.14
88.3
--
--
--
--
--
26<2>
0.058
0.00
7.38
0.00
0.00
90.79
1.83
--
--
--
--
27C<5'
--
--
--
--
--
--
56
0.04
--
0.6
1.4
--
- -
28b'5'
--
--
6
<10
3
50
33
--
59
3
<0.1
<50
33b<5>
» H
--
--
--
0.1
12
<5
5.3
74
<0.5
<0.05
69
3.1
<0.2
71
A-1H
-------
TABLE A-52. COg-ACCEPTOR PROCESS WATER HOLDING POND SOLIDS (STREAM NO. 16)
Run No. (Reference)
Trace Elements
Sb
As
Be
Cd
Cr
Pb
Hg
N1
Se
Te
V
Pilot Plant Run 28b^
ppm
5
<10
2
60
70
--
<0.4
<0.1
<50
--
8.0 Data Gaps and Limitations - Data gaps and limitations for the C02 -
Acceptor process relate primarily to the properties of specific inter-
mediate and discharge streams.
Raw product gas: Levels of H2S, COS, NH3> and HCN are not known.
t Raw regenerator flue gas: Levels of sulfur and nitrogen species are
not known.
Ash tank offgas: Composition of this stream is unknown. The commer-
cial counterpart of this stream would be processed for sulfur
recovery. Composition data (especially H2S and CO?) from the pilot
plant batch operation could provide some useful information for
downstream system design.
Quench and scrubber blowdowns: Constituents not reported for these
streams include total organic carbon, thiocyanate, and fixed cyanides,
Flow data for these streams is either lacking or semiquantitative.
A-115
-------
Reject acceptor and ash: The forms and quantities of sulfur in these
waste solids are not completely known.
t Pond effluent: Available data for pond water and pond effluent are
not entirely consistent (Table A-47).
0 At present, data relating to recent pilot plant runs with sub-
bituminous coal(s) are unavailable.
Although, as presented above, some data are available on the characteris-
tics of a number of input and discharge streams, the available data are
comprehensive in that not all streams are addressed and not all potential
pollutants and toxicological and ecological properties are identified.
The type of data needed correspond to that which are obtained using the
EPA's phased level approach to multimedia environmental sampling and
analysis'^'.
9.0 Related Programs
0 Radian Corporation, under contract to CONSOL, prepared a test plan
for sampling at the pilot plant and conducted sampling/testing during
run No. 39 (September 27 through October 2, 1976). The test results
are not currently available to the public.
The Environmental Studies Institute of the Carnegie-Mellon University
(CMU) has been the environmental program coordinator for the C02-
Acceptor sampling and analysis efforts which have been conducted.
Data for pilot plant runs with Glenharold lignite are expected to be
published by CMU in the near future.
A-116
-------
REFERENCES
1. Fink, C., G. Curran, et al, CO^-Acceptor Process Pilot Plant - 1974, pre-
sented at the 6th Synthetic Pipeline Gas Symposium, Chicago, Illinois,
October 28, 1974.
2. Fink, C., 6. Curran, et al, COg-Acceptor Process, Symposium Proceedings:
Environmental Aspects of Fuel Conversion Technology, II. EPA-600/2-76-149,
June 1976.
3. Dravo Corp., Handbook of Gasifiers and Gas Treatment Systems, ERDA FE-1772-
11, February 1976.
4. Preliminary Economic Analysis of C02-Acceptor Process Producing 250 Million
Standard Cubic Feet Per Day of High-BTU Gas from Two Fuels: Sub-Bituminous
Coal and Lignite, U.S. Bureau of Mines, Morgantown, W. Va., March 1976.
5. Massey, M. J., R. W. Dunlap, et al, Characterization of Effluents from the
Hygas and CC>2-Acceptor Pilot Plants - Interim Report July-September 1976,
Carnegie-Mellon University under ERDA Contract E(49-18)-2496, November 1976.
6. McCoy, D.C., ^-Acceptor Process Pilot Plant, 1976, 8th Synthetic
Pipeline Gas Symposium, Chicago, Illinois, October 18-20, 1976.
7. Dunlop, R. W., and Massey, M. J., Gas-Phase Environmental Data from
Second Generation Coal Gasification Processes, Interim Report No. 1,
ERDA Document No. FE-2496-2, February 1977.
8. Massey, M. L., et al, Environmental Assessment in the ERDA Coal
Gasification Development Program - Progress Report for the Period
July 1976-December 1976, Environmental Studies Institute of Carnegie-
Mellon University, Pittsburgh, Pennsylvania, March 1977.
9. Detman, R., Preliminary Economic Comparison of Six Processes for
Pipeline Gas from Coal, 8th Synthetic Pipeline Gas Symposium,
Chicago, Illinois, October 18-20, 1976.
10. 1975-1976 Fossil Energy Program Report, Vol. II - Coal Gasification,
ERDA Document No. 76-10.
11. Dorsey, J.A., and Johnson, L.D., Environmental Assessment Sampling
and Analysis: Phased Approach and Techniques for Level 1, EPA-
600/2-77-115, June 1977.
A-117
-------
SYNTHANE PROCESS
n 2)
1.0 General Information^' '
1.1 Operating Principles - High pressure coal gasification in a
fluidized bed by injection of steam and oxygen with counter-current
gas/solids flow.
1.2 Development Status^12) - A 65 tonnes/day (72 tons/day), 6.8 MPa
(1000 psia) pilot plant has been constructed at Bruceton,
Pennsylvania by Lummus Co. for ERDA. The plant operated from
July to December 1976 in the free fall injection mode.* Several
problems with this mode of operation led to a modification of the
gasifier to allow deep bed injection of feed coal. The plant has
operated in the latter mode since February 1977, with fewer
operating problems. Tests with agglomerating coals are planned
and downstream systems are to be brought on line during the latter
part of 1978.
f
i
1.3 Licensor/Developer - U.S. Department of Energy (DOE), Pittsburgh
Energy Research Center, 4800 Forbes Avenue, Pittsburgh, Pa.
15213.
1.4 Commercial Applications - None.
2.0 Process Information
2.1 Bench Scale/Process Development Unit (see Figure A-13^4'5h
2.1.1 Gasifier
Equipment
t Construction: vertical, cylindrical stainless steel
pipe
e Dimensions: a 183 cm (6 ft) high, 10 cm (4 in.) I D
pipe inside a 25 cm (10 in.) I.D. pipe
*
uizeddlPvl! the Ejection of coal above the
fair far Dluaalna of ±/a?ld h?atl"9 and dev°^tilization during "free-
tan ar plugging of internal cyclone and overloading quench systems led
'6 heis9;echn?qu MM*?"- C°al 1s now 1nj2t2HSo^h^f?u1d
, this technique has minimized tar generation and associated problems.
A-118
-------
PRETREATER
COAL
FEED
HOPPER
CONDENSER
WATER
STEAM
GENERATOR
-»
V
CONDENSER
FILTER
GASIFIER
ASH EXTRACTOR
n
0
10
11
LEGEND:
1. COAL FEED
2. NITROGEN FOR PRETREATER
3. OXYGEN FOR PRETREATER
4. OXYGEN FOR GASIFICATION
5. STEAM FOR GASIFICATION
6. RAW PRODUCT GAS
7. FILTER DUST
8. ASH
9. PRODUCT GAS
10. CONDENSATE
11. TARS
ASH HOPPER
Figure A-13. Bench Scale/Process Development Unit Synthane Gasified3'
-------
Bed type and gas flow: fluidized bed with continuous
counter-current gas flow, horizontal gas outlet
Heat transfer and cooling: direct gas/solid heat
transfer: electric heaters and refractories are con-
tained in the annul us between the two pipe sections
Coal feeding: feeding is by overflow from the fluid
bed pretreater. The pretreater is fed from a feed
hopper by injection of oxygen and nitrogen.
Gasification media introduction: continuous flow of
steam plus oxygen at the bottom of the bed
Ash removal: an ash "extractor" is used to remove
and transfer ash to a hopper
Operating Parameters
Gas outlet temperature: ?
t Coal bed temperature: Max 1258°K (1800°F)
Gasifier presssure: 4.0 MPa (600 psia)
Coal residence time in gasifier: 18 sec*
.Raw Material Requirements
Coal feedstock:
Type - all types, caking coals partially oxidized
in attached pretreater
Size - minus 20 mesh
Rate - 44 kg/hr (20 Ibs/hr)
Coal pretreatment: attached fluid bed pretreater
t Steam: 1.54 - 1.68 kg/kg coal
Oxygen: 0.30 - 0.37 kg/kg coal
Other materials: None
*
Calculated based on data in Reference 5.
A-120
-------
Utility Requirements
Water: ?
Electricity: ?
Process Efficiency
Cold gas efficiency: ?
Overall thermal efficiency: ?
Carbon conversion: 71% - 72%
Expected Turndown Ratio - ?
Gas Production Rate/Yield - 0.69 - 0.73 Nm3/kg (11.7 - 12.3
12.3 scf/lb)
2.1.2 Coal Feed/Pretreatment - Dry pulverized (-20 mesh) feed
coal is transported from a feed hopper by injection of
oxygen and nitrogen and carried to the pretreater. The
pretreater operates at gasifier pressure and 636°K (690QF)
to partially oxidize the feed and destroys its caking
properties. The feed coal overflows the pretreater to the
gasifier.
2.1.3 Quench and Dust Removal - Dust removal is by a filter with
condensibles removed in a two-stage condenser.
2.2 Pilot Plant (Figure A-14 )^]'2'7^
2.2.1 Gasifier (Figures A-14 andA-15)
Equipment
Construction: vertical, cylindrical steel vessel
t Dimensions: 31 m (101 ft. 9 in.) high, 0.15 m (5 ft)
I.D.
e Bed type and gas flow: fluidized bed with continuous
counter-current gas flow, horizontal gas outlet from
internal cyclone near top of vessel
t Gasification media introduction: continuous flowing of
steam plus oxygen at the bottom through multiple
orifices in a cone shaped distributor
A-121
-------
ro
ro
SCRUBBER SURGE \STfAM
LEGEND:
1. FEED COAL
2. OXYGEN TO PRETREATER
3. OXYGEN TO GASIFIER
4. STEAM TO PRETREATER
5. STEAM TO GASIFIER
6. RAW PRODUCT GAS
7. TARS
8. CHAR
9. QUENCHED PRODUCT GAS
10. CONDEIMSATE SLOWDOWN
11. LOCKHOPPER PRESSURIZATION GAS
12. DECANTER WASTEWATER
13. MAKEUP WATER
14. LOCKHOPPER VENT GAS
15. PRETREATER OFFGAS
Figure A-14. Synthane Pilot Plant Flow Diagram^11>12)
-------
t Afagi--,0*S TO VENTUBI
CHAR TO LOCK HOPPER
TRICKLE VALVE
CH.'.ft BOTTOM OUTLET
(2}
Figure A-15, synthane Pilot Plant Gasified '
A-123
-------
t Char removal mechanism: char is discharged into char
lock hoppers and transported in a low pressure steam
stream
Special features:
- Internal cyclone removes particles larger than 50 ym
- High pressure steam produced in char cooler section
or gasifier
Operating Parameters^ ' '
Run 1-T Run 1-DB
Gas outlet temp. 593°K (608°F) 840°K-880°K
(10500F-1126°F)
Coal bed temp. 1000°K (1440°F) 950°K-1090°K
(1280°F-1500°F)
Gasifier press. 4.2 MPa (615 psia) 4.2 MPa (615 psia)
Raw Material Requirements
t Coal feedstock:
Type - all types, caking coals partially oxidized by
attached pretreater
Size - minus 20 mesh
Rate - Run 1-T Run 1-DB
2.3 tonnes/hr 1.8-2.6 tonnes/hr
(2.5 tons/hr) (2.0-2-8 tons/hr)
Run 1-T Run 1-DB
Steam (kg/kg coal): 1.12 0.99-1.75
Oxygen (kg/kg coal): 2.32 3.25 - 3.85
Utility Requirements
Water: Boiler - ?
Quench - ?
Cooling - ?
Electricity - ?
A-124
-------
Process Efficiency
Cold gas efficiency: ?
t Overall thermal efficiency: ?
Carbon conversion efficiency: Run 1-T 57; Run 1-DB 43-62
Expected Turndown Ratio: ?
Gas Production Rate/Yield Run 1-T Run 1-DB
1.18 Nm3/kg 1.36-1.48 Nm3/kg
(20 scf/lb) (23-25 scf/lb)
2.2.2 Coal Feed/Pretreatment (see Figure A-14) - Dry pulverized
(-20 mesh) feed coal is transported to the pressurized
feed hopper through a system of cyclones, bin and lock-
hoppers. The coal is then entrained by high pressure
steam and a small amount of oxygen and fed to the pre-
treater. The pretreater is a separate fluid bed (main-
tained at 6.8 MPa and 973°K) in which the caking properties
of the feed coal are nullified. The coal from the pre-
treater overflows into the gasifier. Noncaking coals by-
pass the pretreater.
2.2.3 Quench and Dust Removal (see Figure A-14) - After removal
of large (greater than 50 pm) particles in the gasifier's
internal cyclone, product gas exits the gasifier and
passes through a water spray venturi scrubber. The gas
condensates and particulate matter such as carbon fines
enter a scrubber surge tank where the gas and liquid
phases are separated. The gas stream then passes to a
scrubbing tower for further cleanup. The scrubbing tower
contains both water and oil wash sections. The conden-
sates and carbon fines from the venturi scrubber and
scrubbing tower collected in the scrubber surge tank are
pressurized and sent to a decanter.
A-125
-------
2.2.4 Miscellaneous Operations - Systems for the recovery of
heat and recirculation of wash water (and oil) are pro-
vided. Various compressors, pumps and heat exchangers
are employed.
3.0 Process Economics^3'9^ - Capital investment requirements and costs for
the produced SNG have been estimated by the Bureau of Mines and
C.F. Braun^ (for ERDA-AGA). The Braun data is considered the most
reliable and will be the basis for the discussion. Total capital require-
ments for 250 MMSCFD facility are estimated to be 1.5 billion dollars.
Of the total, 11.3% is accountable to coal feed, gasification and quench.
The cost of the produced gas is estimated to be $4.69/MM Btu average
by the utility financing method.
4.0 Process Advantages
9 Caking coals can be used
A high H2/CO ratio is obtained, minimizing or eliminating shift
requirements
A high percentage of methane is produced in the gasifier, reducing
methanation requirements
The process flow system and equipment are relatively simple
Gas is produced at high pressure
Deep bed injection mode of operation minimizes tar formation, thus
simplifying downstream processing systems
5.0 Process Limitations
Limited experience with high pressure vessels-, process has not
operated at pipeline pressure to date
Tar condensation during start-up may cause plugging of the internal
cyclone drip leg in the pilot plant gasifier
Process has not been demonstrated with caking coals to date
6.0 Input Streams
6.1 Coal (Stream No. 1) - (see Table A-53 )
A-126
-------
TABLE A-53- FEED COAL (STREAM NO. 1) COMPOSITIONS FOR SYNTHANE BENCH SCALE
UNIT AND PILOT PLANT
Coal Type/Origin
Proximate and Ultimate
Analysis (%)
Moisture
Volatile Matter
C
H
N
S
0
Ash
Trace Elements (ppm)
Hg
U
As
Zn
Mn
Cr
V
F
B
Be
Bench
Scale,
Unit(5>
Bituminous/
Illinois No. 6
3.7
38.1
65.0
4.9
1.2
3.6
11.6
13.7
0.1
1.4
0.87
25
160
170
100
490
86
1.5
- -
Pilot Plant^
Run 1-T
Subbituminous/
Rosebud
5.2
33.9
62.8
4.6
0.9
0.5
19.6
11.6
--
--
--
Run l-DB-B/4
Subbituminous/
Rosebud
6.9
34.7
66.4
4.4
1.0
0.9
15.9
11.4
__
__
--
....
A-127
-------
6.2 Oxygen
6.3
Stream No. 2
for pret
Rate, kg/kg
(5)
(02 for pretreatment)
Bench Scale
Included in Stream
No. 5
Pilot Plant
01,12)
Pretreater not used
Stream No. 3
(02 for gasification)
Rate, kg/kg 0.37
Steam
Stream No. 4
(steam for pretreatment) Bench Scale
3.25 - 3.85
(5)
Pilot Plant
(11,12)
Rate, kg/kg Not used in test
Pressure, MPa (psia) (1.5 kg/kg N2 used
Temperature in place of steam)
Stream No. 5
(steam for gasification)
Rate, kg/kg 1.68
Pressure, MPa (psia) 4.0 (600)
Pretreater not used
(C02 used to trans-
port coal to the
gasifier)
0.99 - 1.75
4.0 (600)
6.4 Makeup Water (Stream 13) - no data available
6.5 Lockhopper Pressurization Gas (Stream 11) - C02 used at pilot plant
7.0 Intermediate Streams
7.1 Raw Product Gas (Stream 6) - no data avaiTable
8.0 Discharge Streams
8.1 Quenched Product Gas (Stream 9) - see Table A-54
8.2 Tars (Stream 7) - see Tables A~55 and A-56 for tar elemental and
organics composition data from bench scale unit. No data avaiTable
for pilot plant. Essentially no tars are produced in the process
when deep bed coal injection is employed.
A-128
-------
TABLE A-54. SYNTHANE QUENCHED PRODUCT GAS (STREAM NO. 9)
Production Rate,
Nm3/kg (scf/lb)
Composition
(Vol %)
co2
CO
H2
CH4
C2H6
H2S
N2
COS
RSH
HCN
Bench Scale^
0.76 (13)
37.4
12.0
35.1
12.8
1.29
1.43
--
0.014
0.008
8.8 x 10"9
Pilot Plant(15>
Run 1-T
20.2
51.5*
6
31
10
0.9
0.3
0.3
--
Run 1-DB
23-25
50-64*
2.8-9.6
23-28
7-11
0.2-1.3
0.1-0.8
0.0-0.3
--
--
--
Includes about 0.25 Mm3 of C0£ per kg feed coal (4.3 scf/lb) used as trans-
port, petrocarb, and purge gas. This accounts for about one-third to one-
half of the C02 found in the product gas.
A-129
-------
TABLE A-55. ELEMENTAL COMPOSITION OF TAR (STREAM NO. 7) PRODUCED IN
IN THE SYNTHANE GASIFIER(5)
Tar Production Rate,
kg/kg coal
C, %
H, %
0, %
N, %
S, %
Ash, ppm
U, ppm
As, ppm
Zn, ppm
Mn, ppm
Cr, ppm
V, ppm
P, ppm
F, ppm
B,x ppm
Be, ppm
0.03
83.0
6.4
6.4
1.2
2.7
1.2
0.01
0.71
0.48
2.2
7.1
0.21
14
0.97
12
0.03
A-130
-------
TABLE A-56. COMPOSITION OF BENZENE SOLUBLE TARS (STREAM NO. 7) PRODUCED
IN THE SYNTHANE GASIFICATION PROCESS")
Compound/Class
Mono Aroma tics
Benzene
Phenols
Di Aroma tics
Naphthalenes
Indans/Indenes
Naphthols & Indanols
Tri Aromatics
Phenyl naphthalenes
Acenaphenes
Fluorenes
Anthracenes/
Phenanthrenes
Acenaphthols
Phenanthrols
Tetracyclic Aromatics
Peri condensed
(benzanthracenes,
chrysene)
Catacondensed
(pyrene, benz-
phenanthrenes)
Pentacyclic Aromatics
Heterocyclics
Dibenzofurans
Dibenzothiophenes
Benznapththothiophenes
N-Heterocyclics
Type/Origin of Coal
litunrinous
Illinois)
Lignite
(N. Dakota)
Subbituminous
(Montana)
Bituminous
(Pennsylvania)
Volume %
2.1
2.8
11.6
10.5
0.9
9.8
13.5
9.6
13.8
2.7
7.2
4.0
trace
6.3
5.2
10.8
4.1
13.7
19.0
5.0
11.4
3.5
12.0
7.2
10.5
2.5
«
3.5
1.4
not
detected
5.2
1.0
3.8
3.9
5.5
15.3
7.5
11.1
6.4
11.1
9.7
9.0
4.9
0.9
4.9
3/1
.0
not
detected
5.6
1.5
5.3
_
1.9
3.0
16.5
8.2
2.7
7.6
If* rt
5.8
10.7
14.8
2.0
7.6
n -\
H . I
trace
4.7
2 A
.4
8.8
A-131
-------
8.3 Char (Stream 8) - see Table A-57
8.4 Condensate Slowdown (Stream 10) - see Tables A-58 and A-59
8.5 Decanter Wastewater (Stream 12) - no data available
8.6 Lockhopper Vent Gas - no data available
8.7 Pretreater Offgas - no data available; has not operated to date
9.0 Data Gaps and Limitations - The limitations in the data presented fall
into two categories. First, most of the available data are for the PDU
gasifier (which has been extensively tested); the pilot plant may yield
data somewhat different than those for the PDU. Second, although, as
presented above, some data are available on the characteristics of a
number of input and discharge streams, the available data are not com-
prehensive in that not all streams are addressed and not all potential
pollutants and toxicological and ecological properties are identified.
An environmental data acquisition effort which would lead to the gener-
ation of the needed data corresponds to the EPA's phased level approach
to multimedia environmental sampling and analysis^ .
10.0 Related Programs^6' - PERC and Carnegie-Mellon University are presently
performing programs to determine process and effluent stream character-
istics. An ambient sampling program is also being undertaken by PERC
to determine baseline conditions as well as impacts caused by the pilot
plant during operation.
A-132
-------
TABLE A-57.
SYNTHANE FILTER DUST (STREAM NO. 11) AND CHAR
(STREAM NO. 8) FLOW RATES AND COMPOSITIONS
Filter Dust Flow
Rate,* kg/kg coal
C, %
H, %
0, %
N, %
S, %
Ash
Hg, ppm
U, ppm
As, ppm
Zn, ppm
Mn, ppm
Cr, ppm
V, ppm
P, ppm
F, ppm
B, ppm
Be, ppm
Char/ Ash Flow Rate,
kg/ kg coal
C, %
H, %
0, %
v 5 I
-------
TABLE A-58. CONDENSATE (STREAM NO. 10} FLOW RATE AND COMPOSITION
FOR THE SYNTHANE BENCH SCALE UNIT*(4)
Constituent/
Parameter^"
pH
Suspended
solids
Phenol
COD
Thiocyanate
Cyanide
Armenia
TOC
Condensate
Rate, kg/kg
coal
Elemental
Analysis
C, %
H, %
0, %
s, %
Hg, ppm
As, ppm
Zn, ppm
Mn, ppm
Cr, ppm
P, ppm
F, ppm
B, ppm
Illinois
No. 6
Coal
8.6
600
2600
14000
152
0.6
8100
6800
1.44
1.6
10.9
87.0
0.5
0.027
0.001
0.13
0.20
0.043
0.04
39
43
Wyomi ng
Sub-
Bituminous
8.7
140
6000
43000
23
0.23
9500
--
--
--
--
--
--
North
Dakota
Lignite
9.2
64
6600
38000
22
0.1
7200
--
--
--
--
--
--
--
--
--
Western
Kentucky
Coal
8.9
55
3700
19000
200
0.5
10000
--
--
--
--
--
--
--
--
--
--
Pittsburgh
Seam Coal
9.3
23
1700
19000
188
0.6
11000
4980
--
--
--
--
--
--
--
__
Montana
Rosebud
Coal
9.2
68
3000
22000
31
0.07
9500
9090
--
--
--
--
--
--
__
__
*No data reported on the pilot plant
"'"All data are mg/£ except pH
A-134
-------
TABLE A-59. CONDENSATE SLOWDOWN (STREAM NO. 10) COMPOSITION* AND FLOW
RATES AT SYNTHANE PILOT PLANT
Flow Rate (kg/kg coal)
Consti tuent/Parameter
Total Suspended Solids
Phenols
Chemical Oxygen Demand
Ammonia
Sulfide
Pilot Plant Run No. (Ref.
1-T(11)
0.6 -
10,000 -
220 -
4600 -
79 -
0 -
1.2
80,000
3160
22,500
3018
210
)
1-DB(12>
0.6 -
0 -
0 -
70 -
45 -
4 -
0.8
20,000
5
8000
4400
276
mg/1
A-135
-------
REFERENCES
1. Massey, M.J., R.W. Dunlap, et al, Environmental Assessment tn the ERDA
Coal Gasification Development Program, prepared for ERDA, Carnegie-
Mellon University, Pittsburgh, Penn., March 1977, ERDA FE-2496-6.
2. Handbook of Gasifiers and Gas Treatment Systems, prepared for ERDA,
Dravo Corporation, Pittsburgh, Penn., February 1976, ERDA FE-1772-11.
3. Preliminary Economic Analysis of Synthane Plant Producing 250 Million
SCFD High-Btu Gas from Two Coal Seams: Wyodak and Pittsburgh,
prepared for ERDA, Bureau of Mines, March 1976, ERD-76-59.
4. Forney, A.J., et al, Analysis of Tars, Chars, Gases, and Water Found
in Effluents from the Synthane Process, Symposium Proceedings:
Environmental Aspects of Fuel Conversion Technology, May 1974.
5. Forney, A.J., et al, Trace Elements and Major Component Balances
Around the Synthane PDU Gasifier, Symposium Proceedings: Environ-
mental Aspects of Fuel Conversion Technology II, December 1975.
6. Scott, R.L. and Milvihill, J.W., Ambient Air Quality Assessment for
the Synthane Coal Gasification Pilot Plant, Proceedings of the Fourth
National Conference on Energy and the Environment, October 1976.
7. Haynes, W.P., et al, Synthane Process Update, Mid-'77 presented at
the 4th Annual International Conference on Coal Gasification,
Liquefaction and Conversion to Electricity, University of Pittsburgh,
August 2-4, 1977.
8. Kalfadelis, C.D. and Magee, E.M., Evaluation of Pollution Control in
Fossil Fuel Conversion Processes - Gasification: Section 1: Synthane
Process, Esso Research and Engineering, Jun 1974.
9. Detman, R., Preliminary Economic Comparison of Six Processes for
Pipeline Gas from Coal, presentation to the 8th Synthetic Pipeline
Gas Symposium, October 18-20, 1976.
10. Dorsey, J.A., and Johnson, L.D., Environmental Assessment Sampling
and Analysis: Phased Approach and Techniques for Level I, EPA-600/
2-77-115, June 1977.
11. CE. Lummus Co., Synthane Pilot Plant, Bruceton, Pa., Run Report
No. 1 for operating period July-December 1976.
12. C.E Lummus Co., Synthane Pilot Plant, Bruceton, Pa., Run Report
No. 1-DB for operating period February-August 1977.
A-136
-------
BIGAS PROCESS
1.0 General Information
1.1 Operating Principles - High pressure, two-stage gasification by
high-velocity entrainment of coal in a steam and synthesis-gas
mixture in the upper stage (Stage 2) of the gasifier (hydrogasifi-
cation) and injection of oxygen plus steam and char in the lower
stage (Stage 1). Char gasification occurs under slagging
conditions.
fl 2)
1.2 Development Statusv ' ' - The Bigas gasifier has been under develop-
ment since 1965 by Bituminous Coal Research, Inc. (BCR). From
1969 to 1971, a 45-kg/hr (100-lb/hr) Stage 2 PEDU was operated. The
operation demonstrated that high yields of methane righ gas could
be obtained with both low and high rank coals. Under ERDA and
AGA (American Gas Association) sponsorship, a 110-tonne/day
(120-ton/day) integrated pilot plant was constructed (starting in
1972) in Homer City, Pennsylvania. The pilot plant construction
was completed in June 1976 but start-up problems delayed the first
gasification tests until late 1976. Steady state operations with
char gasification have not been achieved to date. Test during
1977 have been aimed primarily at maintaining controlled slag
flow from the bottom of the Stage 1 section of the gasifier. The
successful demonstration of the Bigas process at the pilot plant
awaits the resolution of several mechanical and operational/
monitoring problems including maintenance of continuous slag flow,
measurement of solids and slurry flow rates and measurement of
gasifier internal temperatures.
1.3 Licensor/Developer - Bituminous Coal Research, Inc.
350 Hochberg Road
Monroeville, PA 15146
1-4 Commercial Applications - none.
A-137
-------
2.0 Process Information
2.1 Process Engineering Development Unit (PEDU - see Figure A-16)
2.1.1 Gasifier
(3)
Equipment
Construction: Vertical, cylindrical steel pressure
vessel, single stage refractory lines, for Stage 2.
Stage 1 gasification is simulated by using a burner
which partially oxidizes an aromatic solvent.
Dimensions: 20 cm (8 in.) inside diameter downflow
reactor. Volume 0.056 m3 (1.96 cu ft).
Bed type and gas flow: Coal particles entrained in
continuous concurrent gas flow; vertical gas outlet at
bottom of reactor. Continuous injection of steam plus
coal and CO-rich gas which is generated in the Stage 1
burner.
Heat transfer and cooling: The reactor uses direct gas-
solid heat transfer. Helical water coils in the outer
refractory wall provide gasifier cooling.
Coal feeding: An 8-hour batch pressurized hopper supply
is used with a metering feeder.. The large amount of
transport nitrogen used to pressurize the coal lockhopper
affects product gas composition. Using a superheated
steam jet, coal is injected vertically downwards at about
15 m/sec (50 ft/sec).
Gasification media introduction: Saturated steam,
oxygen and burner fuel form 1900°K (3000°F) synthesis gas
in the PEDU Stage 1 burner. More steam arrives as the
coal transport media.
Char removal: Char in product gas is quenched in a water
stream under the gasifier and sluiced to a separator.
A stream of char in water is let down through a valve
to a flash tank, and this flows to a settling tank.
Operating Parameters' '
Gas outlet temperature: Test program (50 to 60 tests in
1969-71) covered the range 1020°K-1450°K (1375°F-2160°F).
For low-rank coals 110°K (1500°F) was typical; 1200°K
(1700°F) for Pittsburgh seam coal.
A-138
-------
Oxygen
CO
vo
Ash-Free
Fuel
Steam
J
Coal Weigh
Tank Feeder
Stand By
Gas Burner
Recycle
Compressor
Dryer
t
w
i
C
Flor*
Product Gas
Scrubber
Settling
Tank
Sample
Filter
Dissolved
"^ Gas
->» Solids
Liquid
Figure A-16, Flow Diagram of 45-kg/hr (100 lb/hr| Stage 2 Bigas Process ana
Equipment Development Unit (PEDU)(3)
-------
Coal particle temperature: Injection at below 590°K
(600°F) to avoid caking. After injection, the coal
rapidly equilibrates with 1920°K (3000°F) burner gas.
Gasifier pressure: Test program covered range 15 to
9.65 MPa (235-1435 psia). 6.8 MPa (1000 psia) was
typical.
0 Coal residence time in Stage 2: Test program covered
a range of 3 to 22 seconds.
(3 4)
Raw Material Requirementsv ' '
Coal feedstock
Type: essentially any type, caking or noncaking.
Size: 70 percent passing 200 mesh.
Feed rates: covered range of 20-49 kg/hr (43 to
100 Ib/hr).
Coal pretreatment: none.
Steam-to-coal ratio: ranged from 0.9 to 2.78 kg/kg.
Oxygen-to coal ratio: ranged from 0.695 to 1.91 kg/kg.
Others:
Burner fuel (creosote or benzene) - 3.6-45 kg/hr
(8-100 Ib/hr).
Startup fuel (natural gas) - 0.3-3.7 kg/hr (0.2 to
8.2 Ib/hr); primary and secondary burner air 3-164
kg/hr (7 to 360 Ib/hr).
Coal hopper pressurization - 0-11 kg/hr (0-25 Ib/hr).
Utility Requirements^
Boiler feedwater: 6-159 kg/hr (13-350 Ib/hr)
Quench water: 68-250 kg/hr (150-550 Ib/hr)
Scrubber water: 2700-5680 kg/hr (6,000 to 12,500 Ib/hr)
Electricity: ?
Cooling water: ?
A-140
-------
(3)
Process Efficiency^ '
. Co,d gas efficiency -
= 38-85% depending on coal rank and operating conditions;
generally lignites show higher efficiencies than
bituminious coals
. Overa,, «,«.! efficiency - energy output
Turndown Ratio In Dcwnf^ Reactor'^
output
127% of nominal capacity down
to 14% of capacity, based on product gas weight.
Gas Production Rate/Yield: 3-5 Nm3/kg (51-85 scf/lb).
2.1.2 Coal Feed/Pretreatment - Coal is ball-milled to size? no
pretreatment is needed.
2.1.3 Quench and Char Removal - Product gas and entrained char
are quenched in an ambient temperature water stream under
the gasifier and sluiced to a separator. A stream of char
in water is let down through a valve to a flash tank, and
then flows to a settling tank.
2.1.4 Miscellaneous Operations - Product gas is scrubbed with
water, filtered and vented through a letdown valve, and
burned in a flare.
2.2 Pilot Plant
2.2.1 Gasifier
Equipment (see Figures A-17 and A-18 r '
Construction: Vertical, cylindrical steel pressure
vessel, 2-stage configuration, refractory lined in top
two stages.
A-141
-------
STAGE 2
PROCESS GAS
OUTLET
COOLING WATER
OUTLET
GAS SAMPLE
SUPPORT LUGS
REFRACTORY
TWO COAL
INJECTION NOZZLES
THREE CHAR
BURNERS
STAGE 1
COOLING WATER INLET
SLAG TAP BURNER
AND VIEW PORT
SLAG QUENCH ZONE
TWO SLAG
OUTLET NOZZLES
Figure A-17 Bigas Pilot Plant Gasifier^
A-142
-------
co
STEAM
1
s
/
SLURRY
TANK
\P
"
LEGEND-
1. FEED COAL
2. STEAM
3. OXYGEN
4. SLAG
5. RAW PRODUCT GAS
6. STEAM
7. SLAG LOCK HOPPER VENT GAS
8. MAKEUP SLURRY WATER
9. SLURRY TANK VENT GAS
10. OVERHEAD PRODUCT GAS
11. QUENCH SLOWDOWN
12. MAKEUP CONVEYING GAS
13. GASIFIER QUENCH WATER
14. QUENCH WATER DEPRESSURIZATION
VENT GAS
15. MAKEUP QUENCH WATER
16. HEATER STACK GASES
15
L',
VENT
GAS^
WASHER
11
Figure A-18, Bigas Pilot Plant Flow Diagram
-------
t Dimensions: Overall !6.4 m (53.75 ft) high; 0.9 m
(3 ft) inside diameter of refractory sections, 1.5 m
(5 ft) inside diameter of pressure shell. 4.0 m (13 ft)
height of bottom (slag quench) zone; 1.8 m (6 ft) height
of the center (char combustor) zone (Stage 1); 4.3 m
(14 ft) height of the top (coal hydrogasification) zone
(Stage 2). Wall thickness 9 cm (3.56 in.) for all .3
MPa (1165 psia) design pressure.
t Gasification media introduction: continuous injection
of steam, oxygen, and char in the lower stage of the
gasifier. Continuous injection of steam plus coal in
the upper stage.
Coal/char transport and gas flow: Coal/char particles
entrained in continuous concurrent gas flow; vertical
gas outlet at top of reactor.
Ash-removal mechanism: Slag particles fall into a
quench tank in the bottom of the reactor. The slag in
water is removed and let down thorugh a lockhopper.
Special features: The reactor uses direct gas/solid
heat transfer. Vertical water tubes in the walls of
the upper two stages provide gasifier cooling. Coal
is fed by slurry injection with steam through nozzles
in the upper stage of the gasifier. The slurry-feeding
system elmi nates any moisture-content restrictions for
coal feed. The external char cyclone removes entrained
char particles from the raw gas and recycles the char
to the gasifier lower stage to maximize carbon con-
sumption. The 2-stage design maximizes methane produc-
tion in the gasifier.
Operating Parameters
Gas outlet temperature: 1200°K (1700°F)
Maximum char/slag temperature: 1755°K (2700°F) (Stage 1)
, Pressurec 8 MPa (1175 psia) nominal; up to
10 MPa (1500 psia)
Coal residence time in Stage 1: 2 sec.
Stage 2: 6 sec.
Mixing temperature (Stage 2): 1470°K (2200°F)
t Space rate: 14670 kg/hr/m2 (300 Ib coal /hr/ft2)
A-144
-------
Raw Material and Utility Requirements
Coal feedstock
Type - all types
Size - 7Q% less than 75 jim (0.003 in. or 200
Rate - up to 110 tonnes/day (120 tons/day): initial
operations at 76 tonnes/day (84 tons/day) U)
Coal pretreatment: Oxi dative pretreatment of caking
coals is not necessary.
Steam*: 0.72 kg/kg
Oxygen*: 0.56 kg/kg coal
Flue gas*: 0.22 kg/kg coal'2'
Flux"1": Limestone addition in the range of 15% - 30% of
coal ash improves slag flowv^)
Utility Requirements
Water: boiler feed - ?
cooling water - ?
Electricity: ?
Process Efficiency*
Cold gas efficiency: ?
Overall thermal efficiency: ?
Expected Turndown Ratio*:
Gas Production Rate/Yield^: ?
Test 6-1: Char gasification simulated by partial oxidation of natural gas
injection with steam and oxygen in char burners (Stage 1). Steam and oxygem
consumption do not represent steady state coal and char gasification.
"[Tests G-3 and G-4: Flux addition with char gasification
T
Steady state operations with char gasification have not been achieved at the
Pilot plant to date.
A-145
-------
2.2.2 Coal Feed/Pretreatment (see Figure A-26) - Water slurry
from coal preparation (70% through 200 mesh) is compressed
to slightly greater than gasifier pressure with a triplex
plunger pump, heated to about 518°K (475°F) with a stream
preheater, and sprayed into a hot recycle gas stream to
vaporize slurry water plus moisture in the as-received
coal. The vaporized water plus recycle gas are separated
from the dried coal in a cyclone. The recycle gas is
water-washed, compressed and heated for return to the dry-
Ing step. Water is then returned to coal preparation for
reuse.
/I o}
2.2.3 Quench and Dust Removal (see Figure A-18)v ' ' - Gases
leaving Stage 2 with entrained residual char (carbon and
ash) are quenched by water spray and flow to an external
cyclone separator. Char is separated and drained into a
char hopper.
Raw gas and uncoilected fine char from the outlet of the
cyclone separator enter the raw-gas scrubber and pass
upward through a curtain of downward flowing water. The
water scrubs the char dust, cools the gas, and condenses
the moisture. Water is circulated by reflux pumps, and
heat is rejected through an air cooler. A sHp stream of
the wash water is let down to atmospheric pressure through
a valve, to vent-gas washer. Dissolved gases are
released in the washer, and vented to a thermal oxidizer,
and the slurry is drained to a waste pond.
(8)
3.0 Process Economics^ - Investment capital costs for a commercial scale
Q
7 million Nm°/day (250 million scfd), Bigas plant have been estimated at
$1.03 million (1976 dollars). The gasification, power recovery and
raw-gas quench sections are estimated to account for 11.6 percent of
the total installed capital cost.
A-146
-------
4.0 Process Advantages
Gasifier can accept all types of coal directly without oxidative
pretreatment.
No by-product tars, oil, or char which require additional processing
are produced.
Gas is produced at pipeline pressure.
e Cyclone char-recycle system should permit 99+ percent carbon conversion.
t Fine coal particles are not rejected as feed since the gasifier is
designed to use pulverized fuel.
The size and number of gasifier vessels in a commercial plant are
minimized since entrainment design maximized throughput (lb/hr-ft2)
compared to fluidized- or fixed-bed design.
A high percentage of methane is produced directly from coal in the
gasifier, reducing shift and methanation duty.
5.0 Process Limitations
Steady state hydrogasification and char gasification at the pilot
plant have not been demonstrated.
Ability to control slag flow has yet to be demonstrated.
t Durability of materials for slagging,high-temperature service in
presence of reductants is not known.
Pilot plant has limited ability to measure internal temperatures and
various stream flow rates.
The high-throughput, high-temperature process vessel has relatively
high surface area/volume ratio and heat loss potential, a characteris-
tic of entrainment gasifiers.
6.0 Input Streams
6.1 Coal (Stream 1) - see Table A-60.
6.2 Slurry Water (Stream 8) - no data available.
6.3 Steam (Streams 2 and 6) - see Sections 2.1.1 and 2.2.1.
6.4 Oxygen (Stream 3) - see Sections 2.1.1 and 2.2.1.
6.5 Makeup/Recycle Gas (Stream 12) - no data available.
6.6 Quench Water (Streams 13 and 15) - no data available.
A-147
-------
TABLE A-60. PROPERTIES OF COALS WHICH HAVE BEEN GASIFIED IN THE BIGAS PEDU
AND PILOT PLANT (STREAM NO 1)(2>3)
Coal Type
Rate - kg/hr
(Ibs/hr)
HHV - kcal/kg
(Btu/lb)
Volatile
Matter
(Wt %)
Moisture
(wt %)
Ash (wt 85)
Carbon
(wt %)
Hydrogen
(wt %)
Oxygen
(wt %)
Nitrogen
(wt %)
Sulfur
(wt %)
Subbituminous*
(Montana
Rosebud)
3180
(7000)
6360
(11400)
34.7
6.9
10.6
66.4
5.03
15.9
1.5
0.8
Bituminous*
(Pittsburgh
No. 8)
23-35
(50-77)
7720
(13890)
37.5
1.4
6.2
78.1
5.4
6.0
1.4
1.5
Lignite*
(Mercer Co, N.D.)
28-49
(62-108)
4430
(7980)
33.7
25.4
8.2
48.2
3.2
13.3
0.6
1.1
Subbituminous*
(Lincoln Co, Wyo)
47
(104)
5690
(10250)
34.3
17.9
3.5
59.2
4.3
13.4
1.1
0.7
Only Rosebud
pilot plant;
coal has been employed for initial testing at the Homer City
the other coals have been tested in the PEDU.
7.0 Intermediate Streams
7.1 Raw Product Gas (Stream 5) - no data available.
8.0 Discharge Streams
8.1 Quenched Product Gas (Stream 10) - see Table A-61
8.2 Slag (Stream 4) - no data available.
A-148
-------
8.3 Slurry Tank Vent Gas (Stream 9) - no data available.
8.4 Vent Gas Scrubber Slowdown (Stream 11) - no data available.
8.5 Vent Gas (Stream 14) - no data available.
8.6 Slag Lockhopper Vent Gas (Stream 7) - no data available.
8.7 Heater Flue Gas (Stream 16) - no data available; natural gas used
as fuel at the pilot plant.
9.0 Data Gaps and Limitation - The Bigas process has not achieved steady
state operation at the pilot plant, and available PEDU data do not re-
flect char gasification. Stream flow rates and composition data will
have to be obtained through sampling and analysis during steady state
periods of operation (when such conditions occur).
10.0 Related Programs - Penn Environmental Consultants (PEC), an
service contractor, is to conduct an environmental sampling and analysis
program for the pilot plant. Effluent water, gas and solids samples are
to be obtained^6'9).
A-149
-------
TABLE A-61. PROPERTIES OF PRODUCT GAS PRODUCED IN THE BIGAS PEDU
GASIFIER (STREAM NO. 10) ^
Gas Production
Rate-Nm3/kg
(scf/lb)
HHV-kcal/Nm3
(Btu/scf)
Composition
(wt 55)
CO
H2
CH4
co2
N2 + Ar
H2S
NH3
Subbituminous
(Montana
Rosebud)
--
-
Bituminous
(Pittsburgh
No. 8)
(50-85)
--
18.6
31.5
8.1
21.1
20.6*
Lignite
(Mercer Co, N.D.)
(41-71)
15.0
38.3
5.0
22.6
19.0*
Subbi tumi nous
(Lincoln Co, Wyo)
(54)
18.5
35.8
7.5
22.6
15.6*
*Nitrogen used to pressurize coal lockhopper.
A-150
-------
REFERENCES
1. Dravo Corp., Handbook of Gasifiers and Gas Treatment Systems. ERDA Docu-
ment No. FE-1772-11, February 1976.
2. Walker, K. E., Status of the Bi-Gas Pilot Plant Program, Ninth Synthetic
Pipeline Gas Symposium, October 31 - November 2, 1977, Chicago, Illinois.
3. BCR, Inc., Gas Generator R&D - Phase 2. Process and Equipment Develop-
ment, Office of Coal Research R&D Report 20 - Final, August 1971.
4. Young, R. K., Current Status of Bi-Gas Process, presented to 3rd Inter-
national Conference on Coal Gasification and Liquefaction: What Needs
to be Done Now, University of Pittsburgh, August 1976.
5. Bituminous Coal Research, Quarterly Reports to ERDA, FE-1207-21, September
1976; ERDA FE-1207-25, January 1977; and ERDA FE-1207-29, April 1977.
6. Miles, J. M., Status of the Bi-Gas Program - Part 1, Pilot-Plant
Activities, presentation to the 8th Synthetic Pipeline Gas Symposium,
October 18-20, 1976.
7. Phillips Petroleum Co., FY1976 Annual Report to ERDA on Bigas Homer City
Operation, ERDA FE-1207-P21, August 1976.
8. Detman, R., Preliminary Economic Comparison of 6 Processes for Pipeline
Gas from Coal, presentation to the 8th Synthetic Pipeline Gas Symposium,
October 18-20, 1976.
9. Massey, M. J., et al, Environmental Assessment in the ERDA Coal Gasifi-
cation Development Program, Progress Report July 1976-December 1976,
ERDA Document No. E(49-18)-2496, March 1977.
A-151
-------
BATTELE-CARBIDE PROCESS
(Self-Agglomerating Ash)
1.0 General Information
1.1 Operating Principles - Moderate pressure gasification of coal in
a steam-fluidized gasifier. Heat for gasification is provided by
a circulating body of agglomerated ash heated in a separate
burner vessel by combusting char or coal with air.
1.2 Development Status^1'2'5' - Early development work was done by the
Union Carbide Corporation. A 23-tonne/day (25 TPD) process
development unit has been constructed at West Jefferson, Ohio
under the sponsorship of DOE (ERDA) and the AGA (American Gas
Association). Tests of the integrated process (burner and
gasifier) have indicated where modifications must be made in
order to achieve performance goals.
1.3 Licensor/Developer - Battelle Memorial Institute
505 King Avenue
Columbus, Ohio 43201
1.4 Commercial Applications - None
2.0 Process Information
2.1 Process Development Unit (See Figure A-19)^3^
2.1.1 Gasifier (Figure A-20)^3^
Equipment
t Construction: vertical, cylindrical steel. The
gasifier has three sections with I.D.s progressively
increasing from bottom to top.
A-152
-------
cn
CA)
LEGEND:
1. COAL FEED
2. STEAM
3. AIR
4. HOT ASH TO GASIFIER
5. CHAR TO BURNER
6. COOL ASH TO BURNER
7. RAWGS
7. RAW GASIFIER GAS
8. BURNER OFF-GAS
9. ASH SLOWDOWN
10. PRODUCT GAS
11. SCRUB WATER SLOWDOWN
12. CHAR FINES
13. FLUE GAS
14. ASH FINES
SCRUBBER
10
Figure A-19. Battalia-Carbide Gasification Process
(3)
-------
DISTRIBUTOR
DETAILS
GAS TO
CYCLONE
CHAR TO
BURNER
STEAM
INJECTION
RING
DISTRIBUTOR
CLEANOUT
ASH TO BURNER
Figure A-20. Battelle-Carbide
A-154
-------
Dimensions: lower section vessel I.D. 45.7 cm (18 in)
lower section shell I.D. 86.4 cm (34 in)
middle section vessel I.D. 91.4 cm
(36 in)
upper section vessel I.D. 121.9 cm (48 in)
upper section shell I.D. 162.6 cm (64 in)
length 9.7 m (31.8 ft)
Bed type and gas flow: fluidization is provided by
continuous counter-current steam flow. Gas outlet on
top of gasifier.
Heat transfer and cooling: direct gas/solid heat
transfer. Gasifier vessel lined with castable
refractory.
Coal feeding: from pretreater to lockhopper to
gasifier pneumatically.
Gasification media introduction: superheated steam
injected 1) through a distributor plate located at the
bottom of the gasifier and 2) at a higher level in the
gasifier through an externally mounted, ring-type
distributor.
Ash removal: ash agglomerates are removed through a
special opening in the steam distributor plate. Most
of the ash is returned to the burner via an air lift;
the remainder is removed via a lockhopper system for
disposal.
(3 51
Operating Parametersv ' '
Gas outlet temperature: 1144°K - 1255°K (1600°F -
1800°F)
Coal bed temperature: 1088°K - 1255°K (1500°F -
1800°F)
Gasifier pressure: 0.79 MPa (115 psia)
Coal residence time: approximately 30 minutes
A-155
-------
Raw Material Requirements'- '
Coal feedstock
Type: Eastern bituminous coal
Size: +100-8 mesh
Rate: 557 kg/hr (1225 Ibs/hr)
Coal pretreatment: pulverized to +100-8 mesh.
Pretreatment expected to be unnecessary for caking
coals but this has not been demonstrated to date since
only noncaking coals have been tested at the PDU.
Steam: 773 kg/hr (1700 Ibs/hr)
t Air or oxygen: ?
Utility Requirements'' '
Water (Non optimum)
Boiler: 31 1/min (8 gal/min)
Quench: 310 1/min (80 gal/min)
t Cooling: 78 1/min (20 gal/min)
Electricity:
Process Efficiency:
t Cold Gas Efficiency^5)
Energy in product gas output
Total energy in coal input
Thermal Efficiency
X 100 = 66%
Total energy in product gas, tar,
oil and bv-products v inn - ?
Total energy in coal and * iuu - .
electricity
Turndown Ratio:
Full capacity output
Minimum sustafnable output
= ?
A-156
-------
Gas Production
(2)
1.18 Mm /kg dry coal medium Btu-gas
(20 scf/lb dry coal) (estimated)
2.1.2 Burner (Figure A-21)^
Construction: vertical, cylindrical, in two sections
Dimensions: lower vessel I.D. 61.0 cm (24 in)
lower shell I.D. 106.7 cm (42 in)
upper shell I.D. 91.4 cm (36 in)
upper shell I.D. 137.2 cm (54 in)
length 7.14 m (23.4 ft)
Bed type: fluidized bed, co-current coal and air flow
t Heat transfer and cooling: direct solid/gas heat
transfer. Vessel lined with castable refractory.
Coal/char feeding: coal is transferred to the lock-
hopper from the pretreater pneumatically. Feed coal
is then mixed with recycle char from the gasifier and
injected into the burner with air.
Air feeding: air is fed with coal and recycle char
and through a distributor plate at the bottom of the
burner.
Ash removal: ash particles overflow out of the
burner and are transferred to the gasifier by a
steam lift.
(3 5)
Operating Parameters^ '
Gas outlet temperature: ^1366°K (2000°F)
t Bed temperature: 1366°K - 1422°K (20008F - 2100°F)
Burner pressure: 791 MPa (115 psia)
t Coal residence time in burner: estimated at less
than 2 seconds.
A-157
-------
T
TT(
FLUE GAS
TO CYCLONE
o
.T
CO
V
y^
36
^B
V4
\
1
'l.D.^
'ID.
DISTRIBUTOR-DETAILS
JVSH AGGU
TO GASIFIER
ASH AGGLOMERATES & AIR -J ASH WITHDRAWAL
i INTERMITTENT
Figure A-21. Battelle-Carbide Burner
A-158
(3)
-------
Raw Material Requirements^ '
Coal feedstock
Type: Originally designed for eastern bituminous;
only western subbituminous tested to date.
Size: Minus 0.147 mm (minus 100 mesh)
Rate: 444 kg/hr (977 Ib /hr)
Coal pretreatment - dried and pulverized to minus
100 mesh. No other pretreatment necessary.
t
Air: approximately 70 Mm /min (2000 scf/min)
Utility requirements: ?
2.1.3 Coal Feed/Pretreatment^ ' - The sized coal is pneumatically
conveyed with inert gas to two identical lockhopper trains
maintained at a pressure higher than the system pressure.
Coal is transferred to the gasifier from one train, and
coal from the other train is mixed with recycle char
from the gasifier and injected into the burner with air.
f<3\
2.1.4 Quench and Dust Removalv ' - Both gasifier gas and flue
gas from the combustor pass through cyclones and venturi
scrubbers. Slowdown from the venturi scrubbers is sent
to a settler.
3.0 Process Economics
No data currently available.
4.0 Process Advantages
No need for an oxygen plant
Hot flue gases can be expanded through a turbine
Self-agglomerating coals (such as Eastern bituminous) can be
converted to synthesis gas in this process.
t Gas produced is under pressure.
A-159
-------
5.0 Process Limitations
Limited experience with high pressure operations.
Limited experience with recirculating agglomerated ash.
(5)
6.0 Input Streamsv '
6.1 Coal (Stream 1, Figure A-19) - Eastern bituminous (2.5%) sulfur
557 kg/hr (1225 Ibs/hr)
6.2 Steam (Stream 2, Figure A-19) - 773 kg/hr (1700 Ibs/hr)
6.3 Air (Stream 3, Figure A-19) - 70 Nm3/min (2000 scf/min)
7.0 Intermediate Streams^ '
7.1 Ash to Gasifier (Stream 4, Figure A-19) - Approximately
18 tonnes/hr)
7.2 Char to Burner (Stream 5, Figure A-19) - No data available.
7.3 Recycle Ash to Burner (Stream 6, Figure A-19) - Approximately
18 tonnes/hr (20 tons/hr)
7.4 Raw Gasifier Gas (Stream 7, Figure A-19) - Approximately
1230 kg/hr (2700 Ibs/hr) design rate.
8.0 Discharge Streams'4'*
8.1 Ash (Stream 9, Figure A-19) - 1.8-2.3 tonne/day (2-4 TPD) of
coal ash.
8.2 Product Gas (Stream 10, Figure A-19) - 37,520 Nm3/day (1.4 X
10 scfd) of (dry) composition:
Nitrogen 0.6%
Hydrogen 59.0%
Carbon Monoxide 36.5%
Carbon Dioxide 3.1%
H2S 0.1%
*Stream>rates and compositions are projected based on an operating rate equal
to design capacity - 23 tonne/day (25 TPD).
A-160
-------
8.3 Scrub Water (Stream 11, Figure A-19) - 94.6 1/min (25 GPM)
8.4 Char Fines (Stream 12, Figure A-19) - No data currently available
8.5 Flue Gas (Stream 13, Figure A-19) - 120,600 Nm3/day
(4.5 X 10 scfd) of (dry) composition
Nitrogen 81.0%
Oxygen 4.2%
Carbon Dioxide 14.7%
Sulfur Dioxide 2700 ppm
8.6 Ash Fines (Stream 14, Figure A-19) - No data
9.0 Data Gaps and Limitations
Although the PDU has been built, tests have not been completed. It is
expected that two years are needed (January 1980) before an evaluation
of the self-agglomerating ash process can be completed.
10.0 Related Programs
Scientific Design, Inc., N.Y. is presently completing an economic
63 6
evaluation of a commercial scale 7 x 10 Nm /day (250 x 10 scfd)
self-agglomerating ash plant. Results should be available during 1978.
REFERENCES
1. Adams and Corder, Agglomerating Burner Gasification Process; Design,
Installation, and Operation of a 25 Ton-A-Day Process Development Unit.
Quarterly Report - July-September 1976, November 1976, ERDA.
2. Information provided to TRW by H. Feldman of Battelle Columbus
Laboratory, Nov 15, 1977.
3. Dravo Corp., Battelle/Carbide in: Handbook of Gasifiers and Gas
Treatment Systems, pp. 27-31, Pittsburgh, Penn., 1976.
4. Letter, Carl Lyons (Battelle, Columbus) to L. Jablonsky (ERDA), Response
to Questionnaire on Environmental Safety and Health.
5- Information provided to TRW by R. D. Litt of Battelle, March, 1978.
A-161
-------
HYDROGASIFICATION (HYDRANE) PROCESS
1.0 General Information
1.1 Operating Principles - Direct reaction of coal with hydrogen
(hydrogasification) to produce methane.
1.2 Development Status - The hydrogasification process has been tested
in a special two-stage bench-scale reactor (the "Hydrane" gasifier)
at Pittsburgh Energy Research Center (PERC), Bruceton, Pa. The
two stages of the bench-scale Hydrane gasifier (free-fall dilute
phase, FDP; and fluid/moving bed phase, FB/MB) were constructed as
separate sections and were tested first separately and then in
combination as a semi-integrated unit. A feasibility study
providing the preliminary design for a 9.1 tonne (10 ton) per day
PDU and a 27.2 tonne (30 ton) per day hydrogasification process
using the Hydrane reactor design has been prepared for DOE by
Dravo Corporation (Pittsburgh, Pa.). Based on studies conducted
for and by DOE, in 1975 DOE concluded the Hydrane process was not
feasible for commercialization^ '.
Under DOE sponsorship, the Rocketdyne Division of Rockwell
International Corporation (Canoga Park, Ca.) is currently testing
a 0.23 tonne (0.25 ton) per hour, short residence time high
throughput hydrogasification reactor which uses a proprietary
"rocket injection" technology design for feeding pulverized coal
into the reactor (see Section 2.2). This design appears to be
superior to the Hydrane reactor design.
1.3 Licensor/Developer - The U.S. Department of Energy
20 Massachusetts Avenue
Washington, D.C. 20545
1.4 Commercial Applications - None
A-162
-------
2.0 Process Informations
2.1 Hydrane Bench-Scale Reactor*^ '
2.1.1 Gasifier (Free-Fall Dilute Phase Stage1")
Equipment
0 Construction: metal pipe; no additional information
Dimensions: 8.3 cm (3.26 in) I.D. x 1.5 m (5 ft)
L enclosed in 25.4 cm (10 in) I.D. pressure pipe.
Bed type and gas flow: free-fall, concurrent
downward gas-solid flow.
Heat transfer and cooling: referred to as heated pipe
reactor; no further description.
t Coal feeding: coal fed from a coal hopper by a
rotary vane feeder and a nozzle.
Gasification media introduction: continuous feeding
of equimolar mixture of methane and hydrogen into
top of reactor.
Ash/char removal: char collected in an air-cooled
receiver located below the reactor.
Operating Parameters
Gas outlet temperature: ?
Coal bed temperature: 998°K - 1173°K (1337°F -
1652°F); temperature measured in reactor wall
Gasifier pressure: 3.5-20.8 MPa (515-3015 psia)
Coal residence time in gasifier: 1.9-3.8 minutes,
calculated from published data at 1173°K (1652°F) and
pressures of 6.9 and 13.8 MPa (1000 and 2000 psia),
respectively.
*Even though commercial hydrogasification facilities may not use the Hydrane
reactor design, the Hydrane "process" is included here as most of the
available hydrogasification test data have been obtained using the Hydrane
bench-scale reactor. The PERC bench-scale unit is also still used on an
as-needed basis to generate support data for hydrogasification research and
Development programs. .
Used in tests involving only the first stage of two-stage reactor system.
A-163
-------
Raw Materials Requirements
Coal feedstock
- Type: Pittsburgh seam hvAb and Illinois #6 hvCb
bituminous coals and North Dakota lignite
- Size: Pulverized 0.15 - 0.25 mm (50-100 mesh)
- Rate: 2.26 to 8.16 kg/hr (5 to 18 Ib/hr)
Coal pretreatment: none
Simulated feed gas (Hg + CH4): 0.724 - 0.825 Nm3/kg
coal (12.3 - 14.0 scf/lb)*
Process Efficiency
t Cold gas efficiency:
= (product gas energy output/coal energy input) x 100
= ?
t Overall thermal efficiency:
_ total energy output
total energy input
Expected Turndown Ratio (range tested): 2.26 to 8.16 kg/hr
(5 to 18 Ib/hr) coal feed rate.
Coal agglomeration beyond 8.61 kg/hr (18 Ib/hr) with caking
coals.
Gas Production Rate/Yield: 0.624 - 0.995 Nm3(dry)/kg coal
fed (10.6 - 16.9 scf/kg).
* In the 2-stage Hydrane reactor, gas fed to the dilute phase comes from the
2nd stage. In this case, simulated feed gas containing an equimolar mixture
of methane and hydrogen, without carbon monoxide, was fed.
A-164
-------
2.1.2 Gasifier (Fixed/Moving Bed Stage)1"
Equipment
t Construction: ?
Dimensions: 0.79 cm (5/16 in) I.D. x 81.3 cm
(32 in) L
Bed type and gas flow: inapplicable
t Heat transfer and cooling: gas/solids contact
Gasification media introduction: inapplicable
Operating Parameters
Gas outlet temperature: ?
0 Coal bed temperature: (reactor temperature) 1073°Kand
1173°K (14726F and 1652°F) measured by thermocouples
embedded in reactor wall
Gasifier pressure: (reactor pressure) 7.0 MPa
(1015 psia)
t Coal residence time in gasifier: coal conversion
efficiencies were measured after 2, 10, 12 and
20 min
Raw Material Requirements
Char feedstock
- Type: char from bench scale dilute phase reactor
described in Section 2.1.1
- Size: 0.09 - 0.17 cm (0.035 - 0.067 in) particle'
size
- Rate: single batch charge of 8 g (0.02 Ib)
Char pretreatment: none
H~: 0.03 Nm3/hr (1.3 scf/hr)
These data correspond to separate testing of the 2nd stage of the gasifier.
The tests were primarily aimed at evaluating the effectiveness of carbon con-
version in the second stage. An 8-gram (0.018 Ib) charge of char (from the
stage) was reacted with hydrogen under the conditions described.
A-165
-------
Process Efficiency
Coal gas efficiency: not given directly; 54% carbon
conversion at 1173°K (1652°F), 35% at 1073'K (1472°F)
No information available on gas production rate or yield
2.1.3 Gasifier (Integrated Laboratory Two-Stage Reactor) (see
Figure A-22)
Equipment
Construction: vertical, cylindrical two-section pipe
with FDP section above MB section. (See Sections 2.1.1
and 2.1.2)
t Dimensions: overall length = 1.83 m (6 ft). See
Sections 2.1.1 and 2.1.2 for dimensions of each
section.
Bed type and gas flow:
- 1st stage: free-fall dilute phase (FDP); gas and
coal flow concurrent downward.
- 2nd stage: moving bed (MB): char from FDP and
fresh hydrogen feed flow countercurrently - gas
upwards, solids downwards.
Heat transfer and cooling:
- 1st stage: 0.92 m (3 ft) of this section is heated
externally.
- 2nd stage: gas/solids contact.
Char feeding: from a coal hopper to a rotary vane
feeder through a nozzle to top of FDP section. From
1st to 2nd stage, char feed is gravimetric through a
5.1 cm (2.0 in) necked down char transfer tube.
Gasification media introduction:
- 1st stage: continuous feeding of gas into the top
of the gasifier along with fresh coal. Simulated
gas was used instead of actual gas from the
2nd stage.
- 2nd stage: continuous feed of hydrogen into the
bottom of the unit.
A-166
-------
COAL
FEED
HOPPER
DILUTE
PHASE
REACTOR
MOVING
BED
SAMPLE
0.64 CM (0.25 IN) ID
FEED NOZZLE
8.3 CM (3.26 IN) ID
X1.8M (6 FT)
MOVING
BAD
REACTOR
DILUTE PHASE SAMPLE
(PRODUCT GAS)
5.1 CM (2 IN) CHAR
TRANSFER TUBE
DISENGAGING ZONE
8.3 CM (3.26 IN) ID
X 3.05 M (10 FT)
CHAR
EJECTOR
THROAT
CHAR
RECEIVER
1. COAL INPUT
2. CHAR OUTPUT
3. HYDROGEN TO GASIFIER INPUT
4. SIMULATED FDP FEED GAS INPUT
5. MB INTERMEDIATE GAS (OUTPUT FROM MB)
6. FDP INTERMEDIATE GAS (OUTPUT FROM FDP)
Figure A-22. Integrated Two-Stage Hydrane Reactor-Bench Scale Unit
(2)
A-167
-------
Ash/char removal: gravity fall into converging char
ejector throat and char receiver.
Special features: although gasifier is integrated,
simulated H2/CH4 feed gas is fed to FDP while product
gas from 2nd stage exits through side vent. However,
char from FDP serves as feed to 2nd stage.
Operating Parameters*
Gas outlet temperatures: ?
Coal bed temperatures:
- 1st stage: 1073°K-1173°K (1473°F-1652°F)
- 2nd stage: 973°K-1173°K (1292°F-1652°F)
Gasifier pressure: 7.0 MPa (1015 psia), both stages.
Coal residence time in gasifier:
- 1st stage: data not given specifically, but should
be about 1.9 min at 7.0 MPa (1015 psia) as was
indicated in Section 2.1 for bench scale dilute
phase reactor.
- 2nd stage: 10.4 min.
Raw Material Requirements
Coal feedstock
- Type: Illinois #6
- Size: pulverized coal of unspecified size was fed
to 1st stage.
- Rate: 1st stage - 4.7 kg (10.3 Ib) per hr (dry
basis)
2nd stage - 2.27-3.02 kg (6.01-6.68 Ib) per hr char
(dry basis)
*Numbers below and other experimental data for 2nd stage refer to MB tests.
FB tests were attempted but were plagued by problems and failed.
A-168
-------
Coal pretreatment: ?
CH4/H2 mixture: as feed gas to 1st stage, 46/54 per-
cent mixture fed at a rate of 4.46-4.86 Nm^/hr
(166-182 scf/hr).
H2= as feed to 2nd stage, 99% pure, 3.78-4.06 Nm3/hr
(141-152 scf/hr)
Process Efficiency
cold gas efficiency - P"gff ^^put*""* x 100
overall thermal 'efficiency - ** W fff x 100
= ?
Although thermal efficiency is not given, total carbon
conversion is greater than 50% for both stages
combined.
Expected Turndown Ratio: ?
Gas Production Rate/Yield
Nm3/kg coal (SCF/lb coal)
CH4 0.4465 - 0.4594* (7.58 - 7.80)
C2H6 0.0077 - 0.0100 (0.13 - 0.17)
CO 0.0406 - 0.0471 (0.69 - 0.80)
C02 0.0053 - 0.-065 (0.09 - 0.11)
H2Sf 0.0018 - 0.1141 (0.03 - 0.07)
*0nly 0.92 m (3 ft) of the normal 1.52 m (5 ft) length available for the 3
moving bed section was used. The shorter residence time yielded 0.4t>y Mm
methane/kg dry coal (7.80 scf/lb) instead of 0.585 NnvVkg (9.93 scf/lbj
expected.
fAbout 50% of the converted sulfur appeared in the gas product after water
scrubbing.
A-169
-------
2.2 Short Residence Time High Throughput Reactor^ '
2.2.1 Gasifier (the 0.23 tonne/hr Rocketdyne unit)
Equi pment
Construction: metal pipe; no additional information.
Dimensions: 7.6 cm (3 in) I.D. x 4.5 m (15 ft) high
Bed type and gas flow: entrained bed (flow).
Heat transfer and cooling: gas solid heat transfer.
e Coal feeding: the proprietary coal injection system
is based on rocket nozzle technology; pulverized coal
is pressure-fed to the reactor in the form of a dense
coal stream (99.5% coal, 0.5% hydrogen carrier gas).
Gasification media introduction: hot hydrogen gas is
fed to the reactor; feeding mechanism not known.
Ash/char removal: ?
Operating Parameters
Gas outlet temperature: ?
Coal bed temperature: reactor temperature 1200°K
(1700°F)
Gasifier pressure: ?
Coal residence time in gasifier: 0.5 to 1 second.
As was indicated in Section 1.2, under a 1-1/2 year contract with DOE,
Rocketdyne is currently testing a 0.23-tonne (0.25-ton) per hr short resi-
dence time high throughput hydrogasification reactor. Under a subcontract
to Rocketdyne, Cities Service Company of New Jersey is operating a much
smaller scale "research" unit (0.9-0.18 kg/hr or 2-4 Ib/hr) conducting
parametric studies and "screening" tests prior to scale-up verification in
the larger Rocketdyne unit. Very little information is available on the
design features of the two units and on the test results obtained to date.
The Rocketdyne study does not address char gasification which will be
employed in commercial facilities to produce hydrogen for use in the
gasification.
A-170
-------
Raw Material Requirements
Coal feedstock
- Type: ?
- Size: ?
- Rate: unit design capacity 0.23 tonne (0.25 ton)
per hr; the range of coal feed rates tested is
not known.
Coal pretreatment: none.
Hydrogen feed rate: ?
Process Efficiency
Cold gas efficiency:
= (product gas energy output/coal energy input) x 100
= 7
Overall thermal efficiency:
_ total energy output x -|00
total energy input
= ?
Expected turndown ratio: operating range not known.
Gas production rate/yield: not known; a 40% to 50%
coal conversion has been achieved in the gasifier.
3.0 Process Economics
No data are available on the economics of the hydrogasification process
using the Rocketdyne Reactor design.
A detailed economic analysis including sensitivity analysis has been
published for a 6.25 x 109 kcal/day (250 x 109 Btu/day) two-stage
Hydrane plant(2). Total capital is $310,057,000 (Dec. 1974 dollars)
based on Pittsburgh hvAb coal. The gasifier accounts for 14.5% of total
capital or 17.3% of bare equipment cost. Dust removal equipment was
not listed separately.
A-171
-------
4.0 Process Advantages
Methane production directly from coal is high, reducing the
requirement of catalytic methanation.*
No pretreatment, thus minimizing coal losses and reducing process
costs.
Potentially very high thermal efficiencies and carbon utilization
rates.
Residual char from hydrogen plant (0.137 kg/kg dry coal) can be
used as fuel for power plantU). This char is also low in sulfur.
Product gas C02 concentration low (1%), therefore recovered H2S
can be fed directly to Claus plant.
Caking coals can be used.
t Gas is produced at or near pipeline pressure.
5.0 Process Limitations
The Hydrane process has been tested only in a laboratory bench scale
unit; the Rocketdyne hydrogasifier is also in the bench scale
development stage.
Only pulverized coal can be fed to the hydrogasifier.
6.0 Input Streams
6.1 Coal/Char - see Tables A-62 and A-64
6.2 Hydrogen - see Table A-64.
6.3 Simulated FDP Feed Gas - see Tables A-62 and A-64.
6.4 Intermediate Char from FDP - see Tables A-63 and A-65.
7.0 Discharge Streams
7.1 Gaseous -
7.2 Liquids -
7.3 Solids -
see Tables A-63 and A-65
*Jot^erpSnH Proce?s« about 95% "ethane equivalent (CH4 + C2H6) of the
total required for pipeline gas is produced in the gasifier.
A-172
-------
TABLE A-62. OPERATING DATA AND FEED RATES FOR THE INPUT STREAMS FOR
BENCH SCALE DILUTE PHASE REACTOR TESTSU)
Coal Type/Origin*
Test No.
Coal Feed Rate
kg/hr
(Ib/hr)
Pressure
MPa
(psia)
Temp., Gasifiert
Max., °K(°F)
Avg., °K(°F)
Simulated Feed Gas
RateJ
Nrrr/kg coal
(scf/lb coal)
Pittsburgh
156
5.4
(12)
7.0
(1015)
1173(1650)
1100(1520)
0.724
(12.3)
166
5.4
(12)
8.4
(1215)
1173
1100
0.724
(12.3)
Seam hvAb
160
5.4
(12)
10.4
(1515)
1173
1100
0.772
(13.1)
157
5.4
(12)
13.9
(2015)
1173
1100
0.730
(12.4)
Illinois
#6 hvCb
164
5.4
(12)
8.4
(1215)
1173
1100
0.825
(14.0)
North
Dakota
Lignite
184
5.4
(12)
7.0
(1014)
1173
1100
0.724
(12.3)
*No coal composition data have been reported.
^Thermocouples were placed in the wall of the gasifier,
^Simulated gas is 50/50 CH4/H2-
A-173
-------
TABLE A-63 DISCHARGE STREAM RATES AND COMPOSITIONS FOR HYDRANE BENCH SCALE
DILUTE PHASE REACTOR TESTSUi
Test No.*
Product Gas Rate
Nm3/kg coal
(as CO+H2+CH4+C2H6)
(SCF/lb coal)
Carbon Converted, %
Gas Analysis, %
(dry basis)
H2
°2
CO
CH4
C02
C2H6
H2S
N2
HHV
kcal/Nm3
(Btu/SCF)
Condensate Ratet
Kg/kg coal
Char/Ash Rate (Total)
kg/kg dry coal
Oil Rate
kg/kg coal
156
0.789
(13.4)
25.0
22.4
3.2
71.4
0.7
0.5
0.4
1.4
6889
(817)
0.70
166
0.783
(13.3)
25.6
22.7
2.3
72.2
0.5
0.1
0.2
2.3
6872
(815)
-0.01 to
0.70
160
0.795
(13.5)
24.2
19.7
1.4
75.2
0.8
0.3
0.1
2.4
7040
(835)
n n°
u . uo
0.71
-0.01 - 0
157
0,819
(13.9)
30.0
18.1
0.5
79.0
0.4
0.1
0.2
1.7
7277
(863)
n
U
0.66
164
0.995
(16.9)
27.8
21.9
2.4
72.8
0.7
0.0
0.3
1.9
6897
(818)
.05-0.09
0.68
184
0.624
(10.6)
32.1
27.9
6.3
57.5
5.9
0.1
0.1
2.1
5860
(695)
--
--
*See Table A-62 for coal type used.
'The main contaminants in the water were phenols, cresols, xylenols,
naphthalene, anthracene and indole.
A-174
-------
TABLE A-64. OPERATING DATA AND INPUT STREAM RATES AND COMPOSITIONS FOR
HYDRANE BENCH SCALE TWO-STAGE REACTOR TESTS (FIGURE A-22)?2)
Test No.
Reactor State
Coal or Char Feed Rate
kg/hr (dry)
(Ib/hr)
Pressure
MPa
(psia)
Temp. , Gasifier
°K
(°F)
Residence Time, min.
Run Time, min.
Hydrogen Rate to MB
Nm3/hr
(SCF/hr)
Simulated Feed Gas Rate
to FDP
Nm3/hr
(SCF/hr)
Vol % H2
CH4
He
46
FDP* MBt
1.33 0.84
(10.51) (6.68)
7.0 7.0
(1014) (1014)
1123 957
(1562) (1263)
0
187
3.78
(141.4)
4.4 3.8
(164.4) (141.4)
56.2 99.4
37.2
1.05 0.50
5.45
48
FDP MB
1.30 0.64
(10.26) (5.08)
7.0 7.0
(1014) (1014)
1123 1073
(1562) (1472)
10.4
193
4.07 ,
(152.0)
4.9 4.1
(181.7) (152.0)
52.0 99.0
42.1
1.10 1.00
4.7
49
FDP MB
1.30 0.63
(10.32) (5.01)
7.0 7.0
(1014) (1014)
1123 988
(1562) (1319)
10.4
187
4.03
(150.7)
4.5 4.1
(166.2) (150.7)
50.9 98.6
42.8
1.5 1.30
4.7
*FDP: free-fall dilute phase reactor (0.9 m or 3 ft heated length)
' MB: moving-bed reactor
A-175
-------
TABLE A-65. DISCHARGE STREAM RATES AND COMPOSITIONS FOR HYDRANE BENCH SCALE
TWO-STAGE REACTOR TESTS (FIGURE A-22)(2)
Test No.
Reactor Stage
Product Gas Rate
(CO+H2+CH4+C4H6)
Nm3/kg dry coal
(SCF/lb dry coal)
Gas Yields (dry coal
basis)
CH4, Nm3/kg
(SCF/lb)
C2H6, Nm3/kg x TO"3
(SCF/lb)
CO, Nm3/kg
(SCF/lb)
C02, Nm3/kg x 10'3
(SCF/lb)
H2S, Nm3/kg x 10'3*
(SCF/lb)
Oil Yield, kg/kg dry coal
Char/Ash Rate
kg/kg coal
Solids Conversion, wt. %
MAF Coal
C
H
S
N
0
46
FDP MB
0.209
(3.55)
0.177
(3.01)
6.48
(0.11)
0.0253
(0.43)
5.89
(0.10)
2.36
(0.04)
0.048
--
43.1
33.0
75.4
66.7
59.4
91.0
48
FDP MB
0.510
(8.66)
0.459
(7.80)
10.01
(0.17)
0.0406
(0.69)
5.30
(0.09)
1.77
(0.03)
0.041
--
60.2
50.7
96.4
74.8
89.7
99.6
49
FDP
0.501
(8.51)
0.446
(7.58)
7.66
(0.13)
0.0471
(0.80)
6.48
(0.11)
4.12
(0.07)
0.026
--
--
_-
--
MB
--
--
--
--
--
--
--
--
--
--
--
--
--
60.4
53.6
93.5
76.3
86.4
90.0
*About 50% of the converted sulfur appears in the gas product after water
scrubbing.
A-176
-------
8.0 Data Gaps and Limitations
The limitations in the data presented are primarily due to the fact that
the hydrogasification process (Hydrane and Rocketdyne Reactor Designs)
is only in the bench scale development stage with effort primarily
aimed at evaluating the feasibility of hydrogasification and testing the
proposed reactor designs. Even the bench scale Hydrane reactor for
which some data are available has only been tested as a semi-integrated
unit (i.e., a simulated gas and not the gas produced in the second stage
was fed to the first stage).
9.0 Related Programs
The Rocketdyne current 1-1/2 year contract with DOE to build and test
a 0.23 tonne/hr (0.25 ton/hr) short residence time high throughput
hydrogasifier (see Sections 1.2 and 2.2) will expire by the end of
FY 1978. Pending favorable test results, the effort may be followed by
design, construction and testing of a PDU, perhaps 9 to 18 tonne
(10 to 20 ton) per day capacity. The bench scale Hydrane gasifier at
PERC's Bruceton, Pa facility is used periodically on an "as needed" basis
for parametric hydrogasification studies.
REFERENCES
1. Information provided to TRW by Mr. Louis Jablansky, Department of Energy
Oct. 30, 1977.
2. Gray, J. A., and P. M. Yavorsky, The Hydrane Process. In: Clean Fuels
from Coal II, 6th Synthetic Pipeline Gas Symposium, U.S. ERDA, PERC,
Pittsburgh, Pa., 1974. pp 159-175.
3. Information provided to TRW by Mr. Joe Friedman, Rocketdyne Division,
Rockwell International Corporation, November 14, 1977.
A-177
-------
KOPPERS-TOTZEK PROCESS
1.0 General Information
1.1 Operating Principles - High temperature gasification of coal at
atmospheric pressure with co-current flow of coal, oxygen and
steam.
1.2 Development Status - Commercially available since 1952.
1.3 Licensor/Developer - Krupp Koppers, GmbH
Essen, W. Germany
In U.S. - Koppers Company, Inc.
Koppers Building
Pittsburgh, Pa. 15219
1.4 Commercial Applications - 54 gasification units are currently
in operation, 47 using coal as feedstock (see Table A-66).
Existing coal gasifiers are used entirely to make synthesis gas
for the production of ammonia.
2.0 Process Information
2.1 Commercial Units - see Figure A-23, Flow Diagram
2.1.1 Gasifier (see Figures A-23 and A-24)
Equipment
Gasifier construction: horizontal ellipsoidal, double
walled steel vessel with refractory lining. There are
two gasifier designs. The two-headed gasifier
(Figure A-32) has heads shaped as truncated cones
mounted on either end of the ellipsoid. The four-
headed gasifier (Figure A-.33) resembles two inter-
secting ellipsoids with heads at the ends of the
ellipsoids oriented 90° apart(8).
Gasifier dimensions: (see Figure A-24)
t Bed type and gas flow: entrained bed; continuous
co-current gas/solids flow; vertical gas outlet at
the top of the gasifier in the center of the
ellipsoid.
A-178
-------
TABLE A-66. GASIFICATION PLANTS USING THE K-T PROCESS
(6)
Location
Carbonnages de France, Paris,
Mazingarbe Works (P.d.C.)
France
Typpi Oy, Oulu
Finland
Nihoh Suiso Kogyo Kaisha, Ltd.,
Tokyo
Japan
Empresa Nacional "Calvo Sotelo"
de Combustibles Liquidos y
Lubricantes, S.A., Madrid,
Nitrogen Works in Puentes de
Garcia Rodriguez, Coruna
Spain
Typpi Oy, Oulu
Finland
S.A. Union Chimique Beige,
Brussels, Zandvoorde Works
Belgium
Amoniaco Portugues S.A.R.L.,
Lisbon, Estarreja Plant
Portugal
The Government of the Kingdom
of Greece,
The Ministry of Coordination,
Athens,
Nitrogenous Fertilizer Plant,
Ptolemais,
Greece
Fuel
Coal Dust,
Coke-Oven-Gas,
Tail Gas
Coal Dust,
Oil , Peat
Coal Dust
Lignite Dust
Coal Dust,
Oil, Peat
Bunder-C-Oil
Plant convertible for
Coal Dust Gasification
Heavy Gasoline,
Plant extendable to
Lignite-and
Anthracite Dust
Gasification
Lignite Dust,
Bunker-C-Oil
Number of
Gasifier
Units
1
3
3
3
2
2
2
4
Capaci ty
CO + H2
in 24 Hours
75,000-
150,000 Nm3
2,790,000-
5,580,000 SCF
140,000 Nm3
5,210,000 SCF
210,000 Nm3
7,820,000 SCF
242,000 Nm3
9,000,000 SCF
140,000 Nm3
5,210,000 SCF
176,000 Nm3
6,550,000 SCF
169,000 Nm3
6,300,000 SCF
629,000 Nm3
23,450,000 SCF
Us'e of
Synthesis
Gas
Methano,l
and
Ammonia
Synthesis
Ammonia
Synthesis
Ammonia
Synthesis
Ammonia
Synthesis
Ammonia
Synthesis
Ammonia
Synthesis
Ammon i a
Synthesis
Ammonia
Synthesis
Year
of
Order
1949
1950
1954
1954
1955
1955
1956
1959
Continued
-------
TABLE A-66. Continued.
CO
o
Location
Empresa Nacional "Calvo Sotelo"
de Combustibles Liquidos y
Lubricantes, S.A., Madrid,
Nitrogen Works in Puentes de
Gracia Rodriguez, Coruna,
Spain
The General Organization for
Executing the Five Year
Industrial Plan, Cairo,
Nitrogen Works of
Societe el Nasr d1 Engrais
et d' Industries Chimiques,
Attaka, Suez
United Arabian Republique
Chemical Fertilizer Company Ltd.,
Thailand,
Synthetic Fertilizer Plant
at Mao Moh, Lampang
Thailand
Azot Sanayii T.A.S., Ankara
Kutahya Works
Turkey
Chemieanlagen Export-Import
G.m.b.H., Berlin fur VEB Germania,
Chemieanlagen and Apparatebau,
Karl-Marx-Stadt
VEB Zietz Works
Kobe Steel Ltd., Kobe, Japan
for Industrial Development Corp.,
Zambia, at Kafue near Lusaka
Zambia, Africa
Fuel
Lignite Dust
or Naphtha
Refinery Off-Gass,
L.P.G. and Light
Naphtha
Lignite Dust
Lignite Dust
Vacuum residue
and/or
fuel oil
Coal Dust
Number of
Gasi fier
Units
1
3
1
4
2
1
Capacity
CO + H2
in 24 Hours
175,000 Nm3
6,500,000 SCF
778,000 Nm3
28,950,000 SCF
217,000 Nm3
8,070,000 SCF
775,000 Nm3
28,850,000 SCF
360,000 Nm3
13,400,000 SCF
214,320 Nm3
7,980,000 SCF
Use of
Synthesis
Gas
Ammonia
Synthesis
Ammonia
Synthesis
Ammonia
Synthesis
Ammonia
Synthesis
Raw gas to
produce
hydrogen
for hydro-
generation
Ammonia
Synthesis
Year
of
Order
1961
1963
1966
1966
1967
Continued
-------
TABLE A-66. Continued
Location
Nitrogenous Fertilizers
Industry S.A. , Athens,
Nitrogenous Fertilizers PTant
Ptolemais,
Greece
The Fertilizer Corporation
of India Ltd, New Delhi ,
Ramagundam Plant, India
The Fertilizer Corporation
of India Ltd., New Delhi
Talcher Plant, India
Nitrogenous Fertilizers
Industry S.A., Athens
Nitrogenous Fertilizers Plant
Ptolemais, Greece
The Fertilizer Corporation
of India Ltd., New Delhi,
Korba Plant, India
AE & Cl Ltd., Johannesburg,
Modderfontein Plant,
South Africa
Indeco Chemicals Ltd.,
Lusaka, Kafue Works,
Zambia
Indeco Chemicals Ltd.,
Lusaka, Kafue Works
Zambia
Fuel
Lignite Dust
Coal Dust
Coal Dust
Lignite Dust
Coal Dust
Coal Dust
Coal Dust
Coal Dust
Number of
Gasifier
Units
1
4
(1 of them
as standby)
4
(1 of them
as standby)
1
4
(1 of them
as standby)
6
1
2
Capacity
CO + \\2
in 24 Hours
165,000 Nm3
6,150,000 SCF
2,000,000 Nm3
74,450,000 SCF
2,000,000 Nm3
74,450,000 SCF
242,000 Nm3
9,009,000 SCF
2,000,000 Nm3
74,450,000 SCF
2,150,000 Nm3
80,025,000 SCF
220,800 Nm3
8,220,000 SCF
441,660 Nm2
16,440,000 SCF
Use of
Synthesis
Gas
Ammonia
Synthesis
Ammonia
Synthesis
Ammonia
Synthesis
Ammonia
Synthesis
Ammonia
Synthesis
Ammonia
Synthesis
Ammonia
Synthesis
Ammonia
Synthesis
Year
of
Order
1969
1969
1970
1970
1972
1972
1974
1975
00
-------
CO
ro
LEGEND:
1. COAL
2. STEAM
3. OXYGEN
4. SLAG FROM QUENCH TANK
5. COOLED PRODUCT GAS
6. WASHER COOLER SLOWDOWN
7. QUENCH SLOWDOWN
8. MIST ELIMINATOR SLOWDOWN
9. QUENCHED PRODUCT GAS
10. COMBINED EFFLUENT TO CLARIFIER
11 CLARIFIER SLUDGE
12. CLARIFIER EFFLUENT (TO DISCHARGE
OR RECYCLE)
Figure A-23. Koppers-Totzek Coal Gasification Process
(8)
-------
LOW
PRESSURE
STEAM
T
3.4 M (8')
oo
co
COAL -s
STEAM -I
OXYGEN-
BOILER
FEED
WATER
BOILER
FEED
WATER
POSITION OF
HEADS FOR
FOUR HEADED
GASIFIER
BURNER
COOLING
WATER
INTERNAL VOLUME
2-HEADED GASIFIER - 28 m3 (1000 FT3)
4-HEADED GASIFIER - 59 m3 (2100 FT3)
7.6M (25')
Figure A-24. Koppers-Totzek Gastfier
(8)
-------
WASTE HEAT
BOILER
SYSTEM
COAL, STEAM
AND OXYGEN
ASH TO
DISPOSAL
ASH
QUENCH
TANK
GAS TO
COOLING AND
CLEANING
SYSTEM
Figure A-25.
Koppers-Totzek Gasifier with Ash Extractor
and Waste Heat BoilerO)
A-184
-------
Heat transfer and cooling mechanism: Direct gas/solid
heat transfer; the gasifier is water jacketed to
provide gasifier cooling and generate low pressure
steam.
Coal feeding mechanism: continuous screw conveyor
feeds the pulverized coal to mixing nozzles at the ends
of the gasifier heads; the coal is entrained in a pre-
mixed stream of steam and oxygen and the mixture is
injected into the gasifier through sets of two
adjacent nozzles. Injection speeds are higher than
speed of flame propagation to prevent flashback.
Gasification media introduction: continuous injection
of steam plus oxygen, with entrained coal feed.
fo\
Ash removal mechanisnr ;: approximately 50% of the
ash flows down the gasifier walls as molten slag and
drains into a slag quench tank where circulating water
causes it to shatter into a granular form; a conveyor
lifts the slag granules out of the quench tank (see
Figure A-24). The remainder of the ash leaves the
gasifier as fine particles entrained in the exit gas.
The particles are solidified at the gasifier exit by
water sprays. After treating the gas for heat
recovery, particulate matter is removed by a washer
cooler and disintegrator scrubber. The slag is sub-
sequently separated from the scrubber water as a
sludge by a clarifier.
Special features(2>4>5):
- water sprays at gasifier exit and in the washer
cooler system solidify entrained ash particles for
collection by the scrubbing system
- screw feeding system provides for continuous coal
feeding
- slag produced in the quench tank is granular,
allowing for belt conveyor transport
- opposing burners provide for:
high turbulence and mixing
continuous ignition should one burner become
temporarily blocked
A-185
-------
directing the flue into center of gasifier, thus
minimizing hot spots in refractory lining
particles which pass through one flame region
unreacted are gasified in the opposing flame
Operating Parameters
Gas outlet temperature: 1750°K-1780°K(2700°F-27506F)(1)
Coal bed temperature: 3280-3410°K (3300-3500°F)(1)
Gasifier pressure: 0.1 MPa (1 atmr
(4)
t Coal residence time in gasifier: a few seconds
Raw Material Requirements
Coal:
Type - essentially all types with ash contents up to
40%(2)
Size - 70% less than 200 mesh (0.074 mm)
Rate - two headed gasifiers handle up to 360 tonnes/
day (400 tons/day); four headed gasifiers handle up
to 770 tonnes/day (850 tons/day)
Pretreatment - pulverizing and drying to about 2%
moisture for bituminous coals and 8% moisture for
lignites(3,5). por coals with high ash fusion temper-
ature, fluxing agents such as lime, silica, or soda
ash are added to lower ash fusion temperature below
gasifier operating temperature.
Typical Steam and Oxygen Requirements:
Kg steam/ Kg 02/
Coal Type Kg coal Kg coal Ref. #
Montana lignite
Illinois bituminous
Eastern bituminous
Wyoming subbituminous
111 . high volatile
0.14
0.41
0.41
0.14
0.27
0.73
0.86
0.85
0.65
0.70
5
5
5
6
6
bituminous
A-186
-------
kg steam/ kg 02/
Coal Type _ kg coal kg coal Ref. #
Eastern high volatile 0.29 0.82 6
bituminous
South African 0.30 0.79 7
bituminous
(0\
By-Products^ ;, based on Illinois Bituminous Coal Feed
Jacket steam
Waste heat
Pressure
MPa (psig)
0.37 (55)
6.1 (900)
1
Temp °C(°F) 1
141 (287)
480 (900)
-------
Overall thermal efficiency:
[Total energy output (product gas + by-products + steam)] x 1QQ
' [Total energy input "(coal + electric power)]
68%, based on Eastern bituminous coal (1447 kcal/kg or
= 12,640 Btu/lb), quenched and cooled product gas, and
reference temperature 300°K (80°F)U)
Expected Turndown Ratio/
_ [Full capacity output]
[Minimum suitable output]
100/60 for two-headed gasifier,
100/30 for four-headed gasifier
Gas Production Rate^ '
Dry Dry
Coal Type Nm3/kg scf/lb
Montana subbituminous 1.53 (25.9)
Illinois bituminous 1.75 (29.7)
Eastern U.S. bituminous 2.02 (34.2)
2.1.2 Coal Feed/Pretreatment - Coal is dried to 2% - 8%
moisture, depending on rank, and crushed to about 70%
passing 200 mesh. Coal is conveyed with nitrogen to
gasifier service bins which supply the screw feeding
system. Screw feeders continuously discharge coal into
a mixing head where it is entrained in oxygen and low
pressure stream and delivered through transfer pipes to
the burner head of the gasifier.
2.1.3 Quench and Dust Removal^ ' - Product gas is sprayed with
water at the exit of the gasifier to solidify molten
entrained particulates and prevent their adherence to
waste heat boiler tubes. Radiant surface boiler followed
by a fire-tube boiler cool the gas to about 1600°F. Bulk
particulates are then removed by water sprays in a venturi
scrubber/washer cooler. Finer particulates are removed
in a Theisen disintegrator and a mist eliminator.
A-188
-------
fn\
3.0 Process Economicsv ;
Basis:
15 four-headed gasifiers with capacity of 8820 tonnes/day (9700 tons/
/day) producing 1.26 x 10' Nm3/day (3.7 x TflSscfd) of gas at 1 2 MPa
170 psig). Heating value of gas is 2810 kcal/Nm3 (300 Btu/scf).
Includes coal preparation and gas cleaning facilities as depicted
in Figure 1.
Capital - 454 million dollars (1976)
Annual Operating Costs - 95 million dollars/year
4.0 Process Advantages
0 Gasifier can accept all types of coal.
The absence of tars, oils, naphthas and phenols in the raw gas and
quench waters simplifies by-product recovery and pollution control
technology requirements.
Gasifiers can be started in 30 minutes, can be shut down instantly,
and restarted in 10 minutes(3).
Gasifier uses pulverized coal; no unusable fines are generated
during crushing.
Gasifiers have been operated commercially for many years and have
shown high reliability and low maintenance requirements.
5.0 Process Limitations
High temperature of exit gases and slag requires heat recovery in
order to maintain satisfactory thermal efficiency.
Low operating pressure is a disadvantage for transmission of the
product gas or utilization in combined-cycle applications.
Relatively high particulate loadings after quench requires further
processing for many applications.
Low H2/CO ratio in product gas requires extensive shift and C02
removal for methanation or for use in ammonia and methanol
synthesis.
A-189
-------
6.0 Input Streams (see Figure A-23)
6.1 Coal - Stream 1 (Table A-67}
6.2 Low Pressure Steam - Stream 2 (see section 2.1.2 for quantities)
6.3 Oxygen - Stream 3 (see section 2.1.2 for quantities)
7.0 Intermediate Streams (see Figure A-23)
7.1 Cooled Product Gas - Stream 5 (Table A-68)
7.2 Slag Quench Tank Blowdown (Stream 7) (Table A-69)
7.3 Washer Cooler Blowdown (Stream 6) (Table A-69)
7.4 Mist Eliminator Blowdown (Stream 8) (Table A-69)
7.5 Combined Flow to Clarifier (Stream 10) (Table A-69)
8.0 Discharge Streams (see Figure A-23)
8.1 Quenched Product Gas (Stream 9)
composition - expected to be similar to that reported in
Table A-68. Washing operation removes unknown
amounts of NH3, HCN, H2S, COS and S02=
particulates - 9.5 mg/Nm3^; 63 mg/Nm3^
8.2 Clarifier Effluent (Stream 12) (Table A-69)
8.3 Slag (Stream 4) - Composition should be similar.to that of coal
ash. Limited data available from actual operations.
8.4 Clarifier Sludge (Stream 11) - The solids contained in this
stream are a combination of slag particulates from the slag
quench tank and ash particulates from the gas quench/washing
systems. Very limited data are available on the composition of
Clarifier sludge. Composition of the solids should reflect coal
ash composition (Table A-67) and degree of carbon conversion in
the gasifier (Table A-68). Metallic elements in clarifier solids
based on Turkish lignite feed are listed below^11^.
A-190
-------
TABLE A-67. PROPERTIES OF SOME COALS WHICH HAVE BEEN USED IN KOPPERS-TOTZEK GASIFIERS - STREAM 1
Coal Type
Coal Origin
(Reference)
Dry HHV
(kcal/kg) (Btu/lb)
Dry LHV
(kcal/kg) (Btu/lb)
Size
Coal Composition (%l
C
H
N
S
0
Ash
Moisture
Totals
Ash Composition (%)
Si02
A1203
CaO
MgO
Fe203
S03
Totals
Lignite
Turkey
(2)
--
70%<200 mesh
39.9
3.27
1.36
0.95
19.2
32.3
7
104
48.14
13.71
6.73
6.23
16,29
8.18
99.28
Lignite
Montana
(3)
1151 (10050)
__
58.12
4.3
1.1
1.5
14.2
12.7
8.0
100
--
--
--
~
Subbi turn nous
Montana
(6)
1154 (9983)
~
56.76
4.24
1.01
0.67
13.18
22.14
2.0
100
--
--
--
--
--
--
Bituminous
Illinois
(6)
1304 (11390)
--
61.94
4.36
0.97
4.88
6.73
19.12
2.0
100
--
--
--
--
Bituminous
Illinois
(8)
1294 (11310)
62.98
4.23
1.22
4.23
7.90
13.63
6.0
100
41.7
19.8
6.8
1.0
21.2
90.5
Bituminous
"3J
1447 (12640)
69.88
4.90
1.37
1.08
7.05
13.72
2.0
100
--
--
--
Bi tumi nous
South Africa
(7)
1234 (10780)
70% < 200 mesh
68.2
4.3
1.7
1.6
9.9
14.5
1.0
101.2
--
--
--
--
--
-------
TABLE A-68. PROPERTIES OF KOPPERS-TOTZEK RAW PRODUCT GAS - STREAM 5
ro
Coal Type
(Reference)
Dry Composition (%)
CO
H2
CH4
C02
N2+Ar
H2S
COS
CS2
R-5H
S02
NH3
HCN
NOX
Totals
Moisture
HHV dry-(kca1/Nm3)
(Btu/scf)
LHV dry-(kca1/Nm3)
fltu/scf)
Dry Gas Production
Nm3/kg(scf/lb)
Parti culates (wet)
grams/Nm3(grains/ s cf )
Particulate Composition (%)
Si02
A1203
CaO
MgO
Fe203
Carbon
Lignite
(2)
58.4
26.1
--
12.5
2.2
0.5
--
-
--
--
--
--
--
100
--
1.35(22.8)
--
--
--
-
~~
Lignite
(3)
56.87
31.3
--
10.0
1.2
0.6
0.5
--
-
--
--
--
100
2705(289)
--
1.62(27.4)
--
--
--
--
--
--
Subbituminous
(6)
58.68
32.86
--
7.04
1.12
0.28
0.02
--
--
100
2762(295)
1.52(25.9)
--
--
Bituminous
(6)
55.38
34.62
--
7.04
1.01
1.83
0.12
-
--
--
--
100
2716(290)
--
1.75(29.7)
--
--
--
--
Bituminous
(8)
57.35
32.74
--
7.05
1.16
1.59
0.114
-
--
-r
"
--
--
100
9.48
2575(275)
1.80(30.4)
52(22)
30.5
14.48
4.97
0.73
15.51
33.8
Bituminous
(3)
52.8
35.5
0.11
10.1
0.87
0.32
0.025
-
--
0.0031
0.24
0.0407
0.0010
100
29 .'2
2678(286)
--
1.91(32.4)
27(11)
--
--
--
--
--
Bituminous
(7)
56.0
30.0
0.1
11.7
0.15
0.52
0.074
--
-
0.0002
0.0090
0.004
0.0006
99
5.58
--
1.57(26.5)
97(40)
--
--
17
--
-------
TABLE A-69.
KOPPERS-TOTZEK LIQUID PROCESS AND DISCHARGE STREAMS
Stream Number
(Reference)
Coal Type
Coal Origin
Stream Parameter f
TSS
TDS
COD
Alkalinityt
Total Hardnesst
Conducti vity(umho)
pH
Stream Composition**
Ca^
4-4-
Mg++
Na+
K+
Zn++
F6++
Cu^
NH|
NOj
NO-3
Total PO^
ci-
so4=
CN-
Si02
S =
As, Br, Cr, F
Stream 10
(10) (3)
Lignite
Turkey
3072
706
16
--
__
1800
8.8
96
10
18
8
0.02
0.2
0.01
137
0.24
25
0.8
57
255
1.4
20
Not
Detected
__
Bituminous
S. Africa
--
2769
--
--
681
--
8.9
177
55
408
--
0.02
0.2
<0.01
15
6.2
488*
--
284
342
<0.01
69
Not
Detected
Not
Detected
-^ _ __ _
.
Stream 8
(3)
Lignite
Turkey
278
606
18
--
__
970
7.5
60
60
18
7
0.03
0.26
0.06
25
5.3
34
1.7
53
147
7.0
31
Not
Detected
~ =
Stream 6
(3)
_
Lignite
Turkey
5084
940
128
2000
7.5
55
114
18
10
0.02
2.0
0.01
184
4.5
3.7
1.2
96
155
12.5
15
Not
Detected
__
=
Stream 7
(3)
_
Lignite
Turkey
4612
812
18
1800
8.8
71
95
18
9
0.03
0.22
0.01
157
0.13
3.3
0.81
85
216
0.52
16
Not
Detected
_
=====
Stream 12
(31 ln\
\*> j
Lignite
Turkey
50
724
63
2400
8.9
127
80
18
8
0.02
.64
0.06
122
4.4
23
2.7
46
109
14
43
Not
Detected
__
w
' .
Bituminous
S Afrira
"* niri La
.
..
..
15
__
-_.
--
i
-.
Not
Detected
__
*High NOj partially reflects NO^ contained in raw make-up water.
fAs CaCOs
Fmg/1 except pH and conductivity
A-193
-------
Percent of Dry
Element Clarlfler Solids
Fe 6.8 - 8,4
Ni 0.22 - 44
Cu 0 - 0.05
Mn 0.028 - 0.069
9.0 Data Gaps and Limitations
Limitations of the data for the K-T process relate primarily to the
specific properties of input, intermediate, and waste streams. These
limitations include the following:
Feed coals - limited data on ash and trace element composition
of coals which have been gasified in K-T gasifiers.
t Raw and cleaned product gas - limited data on trace sulfur and
nitrogen compounds (C$2, R-SH, S02, NHs, HCN, NOX). No trace
element data for cleaned gas.
t Clarifier effluent and sludge - some data is available for these
streams from the gasification of Turkish lignite. Parameters/
constituents such as TOC, phenols, oil and grease, SCN", and various
trace elements are not included. No data for these streams from
gasification of American coals are available.
Quench tank slag - no data are available on carbonaceous material
or trace elements contained in gasifier slag. The Teachability of
organics and trace elements from such slags is also essentially
unknown.
10.0 Related Programs
Although no K-T gasifiers are currently operating in the U.S., DOE has
recently awarded a contract to Air Products and Chemicals, Inc. of
Allentown, Pa. for design, construction and operation of a Koppers-
Totzek facility to produce hydrogen from coal for industrial use. The
demonstration facility, possibly to be located at Cedar Bayou, Texas,
would use Texas lignite as feed.
No programs specifically aimed at environmental assessment of K-T
operations are known to be under way at present.
A-194
-------
REFERENCES
1. Handbook of Gasifiers and Gas Treatment Systems, ERDA document
No. FE-1772-11, Dravo Corp., February 1976.
2. Wintrell, R., The K-T Process: Koppers Commercially Proven Coal and
Multi-fuel Gasifier for Systematic Gas Production in the Chemical and
Fertilizer Industries, 78th National AIChE Meeting, Salt Lake City Utah
August 1974. '
3. Farnsworth, J. F., et al, Clean Environment with K-T Process, presented
at the EPA Symposium on Environmental Aspects of Fuel Conversion
Technology, St. Louis, Missouri, May 13-16, 1974.
4. Gas Processing Handbook, Hydrocarbon Processing, Vol. 54, No. 3, April
1975.
5. Farnsworth, J. F., Application of the K-T Coal Gasification Process in
the Steel Industry, 104th Annual AIME Meeting, February 16-20, 1975.
6. Information provided by Koppers Company, 1977.
7. Sharpe, R. A., Gasify Coal for Syn Gas, Hydrocarbon Processing, Vol. 55,
No. 11, November 1976.
8. Farnsworth, J. R., et al, K-T: Koppers Commercially Proven Coal and
Multi-Fuel Gasifier, Association of Iron and Steel Engineers Annual
Convention, Philadelphia, Pa., April 22-24, 1974.
9. Mitsak, D. M., et al, Koppers-Totzek - Economics and Inflation, 3rd
International Conference on Coal Gasification and Liquefaction,
Pittsburgh, Pa., August 3-5, 1976.
10. Mitsak, D. M. and Kamody, J. F., Koppers-Totzek: Take a Long Hard Look,
2nd Symposium on Coal Gasification, Liquefaction, and Utilization,
Pittsburgh, Pa., August 5-7, 1975.
11. Information provided by South African Coal, Oil and Gas Corp. Ltd., to
EPA's Industrial Environmental Research Laboratory, Research Triangle
Park, No. Carolina, November 1974.
12. Caution Marks Progress in Coal-Conversion Plan, Chemical Engineering,
October 10, 1977, p. 77.
A-195
-------
TEXACO PROCESS
1.0 General Information^
1.1 Operating Principles - High pressure, high temperature gasification
of coal entrained in oxygen and steam, with co-current gas/solids
flow.
1.2 Development Status - Since 1953, the Texaco process has been in com-
mercial use for the production of synthesis gas from petroleum feed-
stocks and is currently used in approximately 70 plants in over 20
countries^'2). The application of the process to coal is currently
at the pilot plant stage. However, there are plans to convert an
existing European oil gasification plant to the Texaco coal gasifica-
tion process; scheduled start-up of the plant is late 1977^ '. ERDA
has also recently awarded a contract to W. R. Grace for conceptual
design of an 1,800 tonne per day (2,000 TPD) synthesis gas demonstra-
tion plant for the production of 1,088 tonne per day (1,200 TPD)
ic ip oi\
ammonia from high sulfur agglomerating coaP * ' . Minnkota
Power Cooperative, Inc. of Grand Forks, No. Dakota and Northern
States Power Company, Minneapolis, Minn., have also recently circu-
lated a proposal to collectively undertake a feasibility study of a
lignite-fueled methanol plant to be located in western North
Dakota^13'18'. The Louisiana Municipal Power Company (LAMPCO) has
recently proposed construction of a facility at Baldwin, La. to pro-
duce 1.25 kcal/SCM (150 Btu/scf) gas from bituminous coal and residual
oil for power generation^4). (See Table A-63.) Recently, the
Tennessee Valley Authority has awarded a contract to design and con-
struct a Texaco gasifier to produce synthesis gas for ammonia pro-
duction^ '. Southern California Edison Company has announced that
it will test the coal gasification/combined cycle process using a
f
Texaco gasifier at a utility station near Barstow, California11
A-196
-------
TABLE A-70. DEVELOPMENT STATUS OF TEXACO COAL GASIFICATION PROCESS
Facility
Operator
Location
Capaci ty
Status/Miscellaneous
3>
.j
UD
(1)
Pilot Plant
Pilot Plant
(1,3,5,11,17)
Planned Commercial1
Planned /,- 19 9
Commercial(6>12>2
(demonstration)
Planned Demonstration
Planned Commercial'
Planned Commercial
Planned Comnercial
Texaco
Texaco (?)
(Olin-Mathieson)
W. R. Grace &
Co. (ERDA-sponsored)
Montebello Research Laboratory
Montebello, California
Morgantown, W. Va.
Germany (?)
Probably western Kentucky
Tennessee Valley
Authority
Minnkota Power
Cooperative, Inc.
Louisiana Municipal
Muscle Shoals, Alabama
Western No. Dakota
Baldwin, Louisiana
Southern California
Barstow, California
13.6 tonne per day
(15 TPD); Single Train
90.7 tonne per day
(100 TPD)
144 tonne per day
(159 TPD)
Plant would utilize
1,800 tonne per day
(2,000 TPD) of high
sulfur agglomerating
coal for production
of 1,088 tonne per day
(1,200 TPD) ammonia
153 tonnes/day
(168 tons/day)
Plant would utilize
22,700 tonne per day
(25,000 TPD) coal for
production of 2.4 million
liters (7.5 million gals)
of methanol per day.
Plant would produce
1.25 Kcal/SCM (140
Btu/SCF) gas from
bituminous coal and
residual oil for
power generation.
Plant would produce
fuel gas for a 90 MW
gas turbine/30 MW stream
turbine electric generator
In operation since 19 ?-
present.
Operational from
1956-58
Scheduled for start-up
in late 1977
Phase 1, conceptual
design, was awarded
by ERDA in August
1977. Phase II,
construction and
operation, is expected
to be completed in
1981. Project cost
estimated at S320 million.
Contract for enaineers
and construction awarded
June, 1978.
Proposal for feasibility
study issued mid-1977
Economic and engineering
studies are completed.
Project cost estimated
$62 mil lion.
Preliminary engineering
work is underway
*At the present time, two European companies are converting an existing gasification plant to the Texaco coal gasification
process at this site. (2)
-------
1 3 Licensor/Developer - Texaco Development Corporation
135 East 42nd Street
New York, N. Y. 10017
1.4 Commercial Applications - The Texaco process has been in commercial
use for the production of synthesis gas from petroleum feed since
1953. There is no present commercial application to coal. Pro-
posed commercial-scale developments have been discussed in Section
1.2 above.
2.0 Process Information
2.1 Pilot Plant (See Figure A-26, Flow Diagram)
Ground coal is slurried with water (or oil) and the slurry is
pumped to the gasifier. Steam (as optional moderator) and
oxygen (or air) are injected into the gasifier to effect
partial oxidation of the coal, with relative steam, oxygen and
slurry rates chosen to control gasification temperature.
Gasifier products exiting the outlet at the bottom of the
gasifier are cooled to solidify slag, and the product gas
proceeds to an external water cooler (knockout pot) where
ungasified solids are removed.* Water and entrained solids
from the gasifier and the knockout pot are continuously
removed to a soot-water clarifier. Slag is periodically
removed from the bottom of the gasifier via a lockhopper
system and is separated from water by screens and a clarifier
for disposal; the aqueous phase proceeds to a slag fines
clarifier. The settled solids from the slag fines clarifier
and the soot water clarifier are recycled to the process or
wasted. Clarified waters from the two clarifiers are
recycled to the gasifier or slurry tank or wasted (after
depressurization).
2.1.1 Gasifier (see Figure A-27)
« Construction: vertical, cylindical pressure vessel with
carbon steel shell. The top section where gasification
occurs, is refractory lined. The lower section (slag
quench chamber) which contains a reservoir of water for
quenching of gas, is unlined steel(1).
t Dimensions: 1 5m (5 ft) outside shell diameter and 6m
(20 ft) height(2).
rpofarj h. o^9 n f °^d f the Texaco Montebello pilot plant would be
t?ol 2? y 9 coole:/teat recovery for applications such as power genera-
economi ^ 9 Pr6SSUre Steam is necessa^ for overall process
A-198
-------
_ _ jjim_ _ _ _
_SJEAM.
OXYGEN
t
AL
JRRY
NK
^
^^
\
f
t
;
i
/ _
'. G/
/
/
t
t
t
i
t
1
GASIFIER
-j SCREENS |
DARSE SLAG
CLARIKIER
KNOCK
OUT
PLOT
I SOOT L.
I WATER I
CLARIFIEHJ
FLASH
DRUM
10
11
INDEX TO STREAMS
1. COAL
2. COAL SOOT-SLAG FINES
(TO RECYCLE OR DISCARD)
3. OIL (OPTIONAL)
4. MAKEUP WATER
5. RECYCLE WATER (ALSO
MAY BE DISCHARGED)
6. COARSE SLAG
7. SOOT
8. QUENCHED PRODUCT GAS
B. DEPRESSURIZATION GAS
10. SLAG FINES WATER (TO
RECYCLE OR DISCHARGE)
11. SLAG FINES
Figure A-26. Process Flow Diagram for Texaco Coal Gasification Pilot Plant at Montebello, California
(25)
-------
COAL SLURRY FEED
OXYGEN
COOLING
WATER
IN
WATER IN
BURNER
(IGNITION MECHANISM)
COOLING
WATER
OUT
WATER QUENCH
SECTION
SYNTHESIS
GAS
GENERATOR
REFRACTORY
LINING
GAS
-«»SOOT WATER OUT
SLAG OUT
Figure A-27. Texaco Gasifier^
A-200
-------
Bed type and gas flow: entrained bed; continuous
co-current downward gas/solid flow; lateral gas outlet
chamber"?!^ °f the Unit' at the toP of the s1^ quench
Heat transfer and cooling mechanism: direct gas/solids
heat transfer. Water jacket at the top of the gasifier
provides cooling for the burner nozzles(S).
Coal feeding mechanism: continuous injection of coal and
steam (supplied by water in the slurry feed) tangentially
or axially near the top of the gasifier through a water-
cooled burner nozzleW*.
t Gasification media introduction: continuous feeding of
preheated oxygen through a separate water-cooled burner
nozzle tangentially or axially near the top of the
gasifier(4).
0 Ash removal mechanism: molten ash flows through an opening
at the bottom of the gasifier burner section into the slag
quench chamber. The quenched slag is discharged from the
bottom of the quench chamber through lockhoppers'lS). The
water used for quenching is sent to a clarifier for removal
of suspended solids.
t Special features: gas quenching and cooling, as well as
slag removal, are accomplished simultaneously in the slag
quench chamber. The coal /water feeding mechanism elimi-
nates any moisture content restrictions for coal feed.
Operating Parameters
Gas outlet temperature: 478°K to 533°K (400°F to 500°F)(1)
Internal gasifier ("reaction zone") temperature: 1370°K to
1640°K (2000*F-2500-F)(9>2°)
Gasifier pressures: 2.4 - 8.2 MPa (350 to 1200 psig)(1'20)
(4)
Coal residence time in gasifier: a few seconds
(23)
t Coal slurry solids loading: 48X - 66%
*When petroleum or coal liquefaction residues are used as feedstock, the
feedstock is pumped to the gasifier as a liquid; steam is fed into tne
gasifier separately rather than as a liquid/steam mixture.
A-201
-------
Raw Materials Requirements
Coal feed stock requirements
Type - all types of coal (also, hydrocarbon-containing
residuum, such as H-coal liquefaction
res1dues)U,2,lO)
Size - 70 percent less than 0.074 mm (0.003 in)^ ' '
Rate - ~410 kg/sec-m2 (300 lb/hr-ft2) (Calculated from
data in 1)
Steam requirements: 0.1 to 0.6 kg/kg coal (ordinarily
supplied by water in the coal slurry feedUJ. (0.24 to
0.43 kg/kg coal for Illinois #6 H-coal liquefaction
residues; 0.25 to 0.32 kg/ kg coal for Wyodak H-coal
liquefaction residues)^1").*
Oxygen requirements: 0.6 to 0.9 kg/ kg coal '.
(0.98 to 0.10 kg/kg coal for Illinois #6 H-coal liquefac-
tion residues; 1.0 to 1.11 kg/kg coal for Wyodak H-coal
liquefaction residues)(19).
Utility Requirements
Boiler feed water: ?
Cooling water: ?
Electricity: ?
Process Efficiency
Cold gas efficiency:
= (product gas energy output/coal energy input) x 100
= 66 -
= 83 - 84% with H-coal liquefaction residues (Illinois
#6 bituminous and Wyodak coals) C2, 19)
Overall thermal efficiency:
[Total energy output (product gas + HC byproducts + steam)] inn
[Total energy input (coal + electric power)] x
= ?
*When liquefaction residues are gasified, the feed is heated and pumped
directly to the gasifier (a small quantity of light oil may be employed to
clean lines). Steam is injected directly into the top of the gasifier.
A-202
-------
Gas Production Rate/Yield
t 1.5 - 2.1 Nra3/kg (26-36
(Approximately 2.2 - 25 Nm3/kg (35-40 SCF/lb) for H-coal
liquefaction residues)(.2,19)
2.1.2 Coal Feed/Pretreatment(2'4'9) - A thickener is used to
prepare a water slurry of coal containing 40% to 70% coal
by weight. The slurry is then pumped through a heater where
the mixture is heated to 823°K (1000°F) at a pressure of
1.5 MPa (225 psia). The steam to coal ratio is controlled
by reducing excess steam through the use of a cyclone ahead
of the gasifier.*
2.1.3 Quench and Dust Removal^ '5'9^ - Molten slag is discharged
into quench water in the lower half of the gasifier unit
(slag quench chamber). The solidified slag is removed at
the bottom of the gasifier through a lockhopper system.
"Soot water," which contains dispersed soot and other
suspended and dissolved matter, is drawn off near the
bottom of the quench chamber and sent to a clarifier (see
Figure A-34.)
2.2 Conceptual Commercial-Scale Design^ - A typical commercial
Texaco gasifier 2.7m (9 ft) O.D. and 5m (15 ft) high is projected
to gasify 1,700 tonnes (1,900 tons/day) of coal to produce about
3MM SCM/day (100MM SCF/day) of medium-Btu gas at 4.5 MPa
(650 psig).
3.0 Process Economics
A recent estimate of the capital cost of a Texaco gasifier producing
fuel gas for combined cycle power generation has been made by EPRI
The plant investment for coal handling (9,090 tonnes/day or
*WherTpetroleum or coal liquefaction residues are used as feedstock, the
feedstock is pumped to the gasifier as a liquid. Steam is tea to me
gasifier separately rather than as a liquid/steam mixture.
A-203
-------
10,000 tons/day), oxidant feed, gasification, ash handling, and gas
cooling sections of the conceptual facility is estimated at about
$230,000,000 (mid-1976 dollars).
4.0 Process Advantages^
All types of coals, chars, and many other organic materials can be
gasified.
Gasifier can be operated with either oxygen or air.
t Tars, oils, naphthas and phenols are present in the raw gas only in
trace amounts, reducing downstream gas treatment requirements.
t Use of the water slurry feeding mechanism eliminates the need for
coal drying and any restriction on coal moisture content.
When the coal is slurried with water, grinding and pulverizing
operations may be carried out in a wet mi)l» thus avoiding emissions
and hazards associated with dry coal dust(4).
The use of pulverized coal does not require rejection of coal fines
from the feed, as is the case with some other processes.
Gas quench and slag quench are conducted simultaneously in the
bottom of the gasifier vessel.
Essentially all coal carbon is gasified in the process
5.0 Process Limitations
(1)
High temperature of exit gases and slag slurry requires heat
recovery for maintenance of satisfactory thermal efficiency.
t High carryover of slag particles in the raw product gas may lead
to operating problems in the waste heat boiler.
t CO to \\2 ratio in product gas is about 1. Extensive shift is
necessary prior to methanation or for ammonia or methanol synthesis.
6.0 Input Streams (see Figure A-26)
6.1 Coal (Stream No. 1) - See Table A-71.
6.2 Make-up Water for Slurry and Steam (Stream No 4) - See
Section 2.1.1
6.3 Oil (Stream 3) - Optional for use as purge when heavy liquefaction
or petroleum residues are gasified.
A-204
-------
TABLE A-71. PROPERTIES OF INPUT COAL (AND COAL RESIDUE) TO TEXACO GASIFIER
Property
Size
Volatile matter, %
Moisture, %
Composition (dry), %
C
H
N
S
Ash
0
Cl
Ash Composition, %
Si02
A12°3
Fe203
Ti02
P2°5
CaO
MgO
Na20
K20
B2°3
so3
HHV-kcal/Kg(Btu/lb)
Illinois #6
Bituminous
Coal (1,16)
70% <0. 074mm
(0.003 in.)
38.1
3.7
65.0
4.9
1.2
3.6
13.7
11.6
--
--
--
--
--
--
--
--
--
--
7,305
(13,150)
Eastern
Coal
(23)
__
--
72.7
5.03
1.4
3.0
8.7
9.1
--
--
--
--
--
--
--
7,027
(12,650)
Western
Coal
(23)
--
--
74.6
5.31
1.0
0.46
7.2
11.5
--
--
--
--
--
--
--
--
--
7,297
(13,134)
Illinois #6 H-Coal
Liquefaction
Residues (19, 2)*, t
Run 1-2 Run I-5c
--
--
73.1 71.2
5.8 5.4
0.73 0.76
1.37 1.74
16.8 18.6
1.7 2.0
0.5 0.3
Average
46.9
19.3
18.9
0.91
0.15
4.33
1.16
1.29
1.98
0.15
3.67
7,746 7.453
(13,943) (13,416)
Wyodak H-Coal
Liquefaction
Residues (19, 2)*, t
Run W-6 Run W-7
--
78.3 79.7
5.8 5.6
0.9 0.9
0.06 0.01
10.4 9.0
4.6 4.8
.00 .00
Average
31.4
15.8
5.83
0.86
1.63
23.83
5.79
2.26
0.27
0.13
7.38
8,042 8,108
(14,476) (14,594)
*Includes aromatic purge solvent.
tA total of 17 runs were performed using Illinois #6 H-coal liquefaction residues, and a total of 8
runs were performed using Wyodak H-coal liquefaction residues. Runs 1-2, I-5c, W-6andw-/ represent
the extremes of slag and soot production obtained under different input rates (gasification
temperatures). Oxygen input rates were as follows: 0.56 SCM/kg (9.1 SCF/lb) for Run 1-2,
0.61 SCM/kg (9.9 SCF/lb) for Run I-5c; 0.62 SCM/kg (10.1 SCF/lb) for Run W-6; and 0.68 SCM/kg
(11.1 SCF/lb) for Run W-7.
A-205
-------
7.0 Intermediate/Discharge Streams (see Figure A-26)
7.1 Quenched Product Gas (Stream 8) - See Table A-72
7.2 Recycle Water (Stream 5) - No data available
7.3 Coal Soot-Slag Fires (Stream 2) - See Table A-73 for soot and
slag production rates and carbon contents associated with the
gasification of H-coal liquefaction residues. No data are
available for soot and slag from coal gasification.
7.4 Coarse Slag (Stream 6) - No data available for coal gasification;
see Table A-73 for properties of coarse slag from gasification
of H-coal liquefaction residues.
7.5 Soot (Stream 7) - No data available for coal gasification; see
Table 4 for properties of soot from gasification of H-coal
liquefaction residues.
7.6 Slag Fines Water (Stream 10) - No data available
7.7 Slag Fines (Stream 11) - No data available for coal gasification;
see Table A-73 for properties of fine slag from gasification of
H-coal liquefaction residues.
7.8 Depressurization Offgas (Stream 9) - No data available
8.0 Data Gaps and Limitations
Limited data are available to provide an accurate and complete
description of the Texaco process. Data gaps currently exist in
the characterization of most of the gaseous, liquid and solid streams
generated in Texaco pilot operations using coal feed (see Section 7).
Also, limited data are available on the characteristics of the waste
streams from gasification of liquefaction residues.
9.0 Related Programs
The Electric Power Research Institute and Texaco, Inc. are sponsoring
a program for determination of the performance of the Texaco gasifier
for fuel production for combined cycle electric power generation. A
major objective of the program is the characterization of certain emis-
sions generated by the process by means of environmental sampling and
analysis^25).
A-206
-------
TABLE A-72. TEXACO PRODUCT GAS PROPERTIES AND PRODUCTION RATES FOR COAL (AND COAL RESIDUES) (STREAM NO. 12)
Dry Composition
(Vol %)
CO
H2
CH4
co2
N2 + Ar
°2
H2S
COS
HHV-kcal/SCM
(Btu/SCF)
Production Rate
SCM/kg (SCF/lb)
Illinois #6
Bi tuminous
Coal
(1)
37.6
39.0
0.5
20.8
0.6
--
1.5
--
2250
(253)
1.5
(26)
Eastern
Coal
(23)
44.6
36.2
0.4
20.6
0.4
--
0.8
0.05
2283
(271)
1.95
(.33)
Western
Coal
(23)
50.7
35.8
0.09
13.1
0.2
--
0.02
0.01
2488
(295)
2.13
(36)
Illinois #6 H-Coal
Liquefaction
Residues (2, 19)*
Run 1-2
53.1
41.0
0.5
5.2
0.07
--
0.20
0.01
2740f
(308)
2.1
(35)
Run I-5c
51.4
39.9
0.06
8.2
0.04
--
0.40
0.01
26 10t
(294)
2.2
(37)
Wyodak H-Coal
Liquefaction
Residues (2,19)*
Run W-6
39.2
54.9
0.2
0.5
0.18
--
0.00
0.00
2710t
(305)
2.4
(40)
Run W-7
38.0
54.2
0.02
7.6
0.11
--
0.00
0.00
2740t
(308)
2.4
(40)
ro
o
*A total of 17 runs were performed using Illinois #6 H-coal liquefaction residues, and a total of
8 runs were performed using Wyodak H-coal liquefaction residues. Runs 1-2, l-5c, W-6 and W-7 represent
the extremes of slag and soot production obtained under different oxygen input rates (gasification
temperatures). Oxygen input rates were as follows: 0.56 SCM/kg (9.1 SCF/lb) for Run 1-2;
0.61 SCM/kg (9.9 SCF/lb) for Run l-5c; 0.62 SCM/kg (10.1 SCF/lb) for Run W-6; and 0.68 SCM/kq
(11.1 SCF/lb) for Run W-7.
^Calculated from composition.
-------
TABLE A-73. PRODUCTION RATES AND CARBON CONTENTS OF SLAG AND SOOT FOR TEXACO GASIFICATION OF H-COAL
LIQUEFACTION RESIDUES
3=>
ro
o
03
Gasification Residue
Coarse Slag (Stream 6)
Carbon (wt %)
Production Rate
(dry kg/kg feed)
Fine Slag (Stream 11)
Carbon (wt %)
Production Rate
(dry kg/kg feed)
Soot (Stream 7)
Carbon (wt %)
Production Rate
(dry kg/kg feed)
Feed*
HI- #6 , }
H-Coal Residue1 iy;
Run 1-2
12.6
0.023
31.3
0.10
31.3
0.145
Run I-5c
<0.50
0.055
2.50
0.12
16.76
0.029
Wyodak
H-Coal Residue^)
Run W-6
0.5
0.005
8.4
0.044
17.7
0.072
Run W-7
0.45
0.07
7.4
0.005
12.7
0.017
*A total of 17 runs were performed using Illinois #6 H-coal liquefaction
residues, and a total of 8 runs were performed using Wyodak H-coal liquefac-
tion residues. Runs 1-2, I-5c, W-6 and W-7 represent the extremes of slag
and soot production obtained under different oxygen input rates- (gasification
temperatures). Oxygen input rates were as follows: 0.56 SCM/kg (9 1 SCF/lb)
for Run 1-2; 0.61 SCM/kg (9.9 SCF/lb) for Run I-5c; 0.62 SCM/kg (10.1 SCF/lb)
for Run W-6; and 0.68 SCM/kg (11.1 SCF/lb) for Run W-7).
-------
REFERENCES
1. Dravo Corporation, Handbook of Gasifiers and Gas Treatment Systems
ERDA FE-1772-11, Washington, D. C., February 1976, p. 25-26
2. Texaco Reports Results of Gasification Tests on H-Coal Residue
Synthetic Fuels Quarterly Report, Cameron Engineers, Inc., Denver
Colorado, 14 (2): p. 4-34 to 4-40, June 1977.
3. Hall, E. H., et al, Fuels Technology: A State-of-the-Art Review, NTIS No
PB-242 535, U.S. Environmental Protection Agency, Hashington, D. C ,
April 1975, p. 5-72 through 5-73; p. 5-93 through 5-95.
4. Eastman, D., Preliminary Report on Coal Gasification, presented at
Annual Meeting of the American Institute of Mining and Metallurgical
Engineers, New York, February 1952, 7 p.
5. Conn, A. L., Sulfur Developments: Low Btu Gas for Power Plants,
Chemical Eng. Prog., 69 (12): 56-61, 1973.
6. An Ammonia-from-Coal Demonstration Plant Proposal, Chemical Eng., 84
(19): 87, 1977.
7. Hahn, 0. J., Present Status of Low-Btu Gasification Technology,
Institute for Mining and Minerals Research, University of Kentucky,
Lexington, Kentucky, January, 1976.
8. Hoy, H. R. and D. M. Wilkins, Total Gasification of Coal, Brit, Coal
Utility Research Association Monthly Bulletin, 22: 57-110, 1958.
9. Katz, D. L. et al, Evaluation of Coal Conversion Processes to Provide
Clean Fuels, Final Report, EPRI 206-0-0, PB-234 202 and PB-234 203,
University of Michigan, College of Engineering, Ann Arbor, Michigan,
1974, p. 198-200.
10. Glazer, F. et al, Emissions for Processes Producing Clean Fuels,
EPA-450/2-75-028, U.S. Environmental Protection Agency, Washington, D.C.,
March 1974. p. X-l to x-20.
11. Ferrell, J. and G. Poe, Impact of Clean Fuels Combustion on Primary
Particulate Emissions from Stationary Sources, PB-253 452, Accur^x
Corporation, Aerotherm Division, Mountain View, California, March 1976,
p. 3-2 to 3-9.
12. Government Concentrates, Chemical and Engineering News, 55(36);16, 1977.
13. Feasibility Study Proposed for North Dakota Methanol Plant, Synthetic
Fuels Quarterly Report, Cameron Engineers, Inc., Denver, Colorado,
14 (2): p. 4-16 to 4-17,. June 1977.
A-209
-------
14. Status of Synfuels Projects/Coal-Louisiana Municipal Power Commission.
Synthetic Fuels Quarterly Report, Cameron Engineers, Inc., Denver,
Colorado, 14 (2): p. B-ll, June 1977.
15. Child, E. T. and C. P. Marion, Recent Developments in the Texaco
Synthesis Gas Generation Process, for presentation at the Fertilizer
Association of India, National Seminar, New Delhi, India, December
1973.
16. Forney, A. J., et al., Trace Elements and Major Component Balances
Around the Synthane PDU Gasifier, Symposium Proceedings: Environmental
Aspects of Fuel Conversion Technology-II, December 1975.
17. As Coal Bids for a Chemical Comeback, Chemical Week, 78: 76-80,
June 30, 1956.
18. Osur, J. D., Consultant, and E. C. Glass, Northern States Power Company,
and A. L. Freeman, Minnkota Power Cooperative, Report on a Proposed
Study of a Lignite-Fueled Methanol Plant, March 25, 1977, 65 p.
19. Robin, A. M., The Production of Synthesis Gas from H-Coal Liquefaction
Residues, Texaco Inc., Montebello Research Laboratory, for presentation
at 83rd National Meeting of the American Institute of Chemical
Engineers, Houston, Texas, March 20-24, 1977, Texaco Document No. 2085,
32 p.
20. Evaluation of Coal Gasification Technology, Part II, Low and Inter-
mediate Btu Fuel Gases, Panel on Evaluation of Coal Gasification
Technology, Office of Coal Research, U.S. Department of Interior,
Washington, D.C., 1974, p. 35.
21. ERDA Awards Contract to W. R. Grace and Co. for Design of Medium Btu
Demo Plant, Information from ERDA, Weekly Announcements, Washington,
D. C., Article No. 77-150, 3, (35): p. 2, September 2, 1977.
22. Chemical Engineering, June 19, 1978.
23. Crouch, W. G. and G. Klapatch, Solids Gasification for Gas Turbine Fuel
100 and 200 Btu Gas, llth Intersociety Energy Conversion Conference,
Lake Tahoe, September 16, 1976.
24. Economic Studies of Coal Gasification Combined Cycle Systems for
Electric Power Generation, Electric Power Research Institute, EPRI
AF-62, Project 239, January 1978.
25. Request for Proposal Entrained Gasification-Combined Cycle Power
Systems - Environmental Baseline Studies, Texaco, Inc., November 8, 1977.
26. "Chementators," Chemical Engineering, Vol. 85, No. 7, March 27, 1978,
p. 68.
A-210
-------
APPENDIX B
GAS PURIFICATION OPERATION
Acid Gas Removal Module
Physical Solvents
Rectisol (Single Absorption Mode)
Rectisol (Dual Absorption Mode)
Selexol
Purisol
Estasolvan
Fluor Solvent
Amines
Sulfiban (MEA)
MDEA
SNPA-DEA
ADIP
Fluor Econamine (DGA)
Alkazid (Alkacid)
Mixed Solvents
Sulfinol
Ami sol
Carbonate Processes
Benfield (Hot Carbonate)
Redox Processes
Gianmarco-Vetrocoke (G-V)
Stretford
Methanation Guard
Zinc Oxide Adsorption
Iron Oxide Adsorption
Metal Oxide Impregnated Carbon
Activated Carbon
Molecular Sieves
B-l
-------
RECTISOL PROCESS
(Single Absorption Mode)
1.0 General Information
1.1 Operating Principles - Physical absorption of the sour components
(H2S, COp, COS, mercaptans, etc.) of a gas stream using methanol as
the sorbent. Selective regeneration can provide a rich sulfur con-
taining gas stream and a relatively pure C02 stream.
1.2 Development Status - Commercially available.
1.3 Licensor/Developer - Lurgi Mineralb'ltechnik GmbH
American Lurgi Corporation
377 Rt. 17 South
Hasbrouck Heights, N.J.
1.4 Commercial Applications
Purification of low/medium Btu gas produced from coal gasifica-
tion. Gasification plants using the process include Sasolburg,
South Africa; Westfield, Scotland; and Pristina, Yugoslavia.
Carbon dioxide removal and drying of coal-derived ammonia synthe-
sis gas. One of the facilities using this process is located in
Kutahya, Turkey.
t Carbon dioxide removal from low-temperature fractionation feed
gas. The locations of facilities using the process are not known.
Carbon dioxide and water removal from a feed gas to LNG plants.
Plant location(s) are unknown.
2.0 Process Information^1'2'3'6^
2.1 Flow Diagram (see Figure B-l, B-2 and B-3) - The Rectisol Process
can be used in a variety of "modes" to achieve different treatment
objectives. Only three operation modes which have been used or pro-
posed appear most pertinent to coal conversion and are discussed
here. The pertinent features of these operation modes are summa-
rized in Table B-l.
B-2
-------
DO
I
OJ
m
cc
CC
o
a:
LU
LU
CC
13
LEGEND:
RAW GAS
WATER
MAKE-UP MeOH
PREWASH FLASH GAS
REGENERATOR FLASH GAS
PRODUCT GAS
,. EXPANSION GAS
8. COMBINED FLASH GAS
9. REGENERATOR OFF-GAS
10. STILL BOTTOMS
11. NAPHTHA
128.13. STRIPPING GAS
C/J
X
o
CC
o
<
CC
O
X
-^ 11
Figure B-l. Rectisol Type A' (Removal of C02 from Gas Mixtures Containing Little or
No
H2S)
-------
CO
-p.
f
1
5
O
O
T
PREWAS
*
1
L
<
<
-
L
i
C.
\
0
<
J-
£
/)
S.
>
LI
C
1_
*-
12
L*
I
co2
ABSORBER
ULFUR
SORBER
CO 03
4
1
J
'
c
i
1
5
4;
j
CM
D
^
1
(
k
b
1
LL
LL
LO
I
1
1
cc
O
LU
LU
0
LU
CC
O
1
1
_J
C
I
3
a
^
^
)
N
j
3
NAPHTHA
SEPARATOR
LEGEND:
RAW GAS
WATER
MAKE-UP MeOH
PREWASH FLASH GAS
SULFUR FLASH GAS
PRODUCT GAS
7. LEAN H2S COMBINED GAS
8. CO2 FLASH GAS
9. REGENERATOR GAS
10. STILL BOTTOMS
11. NAPHTHA
12, 13, & 14. STRIPPING GAS
Figure B-2. Rectisol Type B (Removal of C02 and H2S with Separate Recovery)
-------
11-*
CO
01
LEGEND:
1.
2.
3.
4.
5.
6.
7.
8.
9.
10.
11.
12.
13.
14.
12
RAW GAS
WATER
MeOH
NH3 COOLANT
COOLANT
PRODUCT GAS
INTERMEDIATE GAS
REGENERATOR OFF-GAS
5TH STAGE FLASH GAS
6TH STAGE FLASH GAS
1ST, 2ND, 3RD, 4TH STAGE FLASH GAS
AROMATICS
STILL BOTTOMS
CONDENSATE
10
Figure B-3. Rectisol Type
^ '
(Removal of CO and
H2S with Separate Recovery)
-------
TABLE B-l. PROCESS DESCRIPTIONS FOR RECTISOL TYPES A, B, AND C OPERATING MODES
Type
Process Application/
Treatment Objective
Process Description
(Fig. B-l)
CO
B
(Fig. B-2)
(Ftg. B-3)
Removal of C02 from
gas mixture con-
taining little or
no sulfur.
Simultaneous removal
of C02 and sulfur
compounds with sep-
arate recovery.
Same as Type B
A methanol stream rich in CO? and HgS is used in the prewash column to remove water,
naphtha, ammonia and residual heavy hydrocarbons from the raw gas. The exit
solvent enters the prewash flash column where a flash stream lean in I^S and
rich in C02 is produced (Stream 4). The liquid bottoms from the flash vessel
are routed to a separator where water (Stream 2) is added so that the naphtha
and heavy hydrocarbons can be separated. In the main absorber raw gas contacts
a pure methanol stream from the hot regenerator. A slipstream of saturated
methanol is sent to the prewash column. The remaining methanol is sent to a
flash regenerator where the absorbed gases are removed. Methanol from flash is
sent to the hot regenerator where the final traces of C02 and h^S are removed.
Water is removed from the prewash methanol in the methanol/water still with off
gases going to the hot regenerator. Stripping gases (usually nitrogen) may be
used.
Except for the use of a two-stage absorber and two separate flash columns, Type B
Rectisol is very similar to Type A. The raw gas (after leaving the prewash
absorber) is first contacted with a C02~saturated methanol stream. This first
stage absorber removes H2S. In the second stage a pure methanol stream removes
C02- The fflethanol for first stage comes from the second stage absorber. The two
methanol streams are flashed separately to create a stream rich in H^S (No. 5)
and a nearly pure C02 stream (No. 8). Regeneration is the same as in the Type A.
The primary difference in Type C as compared to Type B is in the regeneration pro-
cess. The first stage acts like the prewas'h in Type B with second and third stages
similar to first and second in Type B. A multistage flash unit is used to desorb
gases from first and second stage absorption. First stage methanol is first com-
bined with heavy hydrocarbons and water removed from the raw gas and sent to the
separator. The separator works in the same manner as the separators in Types A
and B. The multistage flash reduces the regeneration requirements. The third
stage methanol is handled in a conventional hot regenerator to provide a pure
methanol for final absorption. A split stream regeneration section is also shown
in Figure B-3. Similar gas cooling sections are used in Types A and B but are not
shown on the figures.
-------
2.2 Equipment - Conventional absorbers, stripping columns, distillation
columns, heat exchangers, separators and regenerators.
t Construction - vessels may be fabricated from carbon steel,
dimensions dependent on application.
2.3 Feed Stream Requirements* - Gas should be cooled to reduce solvent
losses; high pressures (close to 2.0 MPa or 300 psia) are usual.
Gas temperatures between 253°K-213°K (-5°F to -75°F) are usual,
depending on conditions^).
2.4 Operating Parameters
Absorption: 0.3-7.1 MPa (45 to 1066 psia)
approximately 303aK (80°F)
t Regeneration: see discharge streams, Section 8.0
2.5 .Process Efficiency and Reliability
C02 better than 97%^
H2S better than 99. 9% ^
Reliability is considered high with a simple solvent and
construction.
2.6 Raw Material Requirements
Solvent - CH3OH; purity - ?
Solvent losses can be estimated using equilibrium constants;
however, considerable errors could be involved. No informa-
tion available on solvent losses based on actual operating data.
2.7 Utility Requirements - ?
2.8 Miscellaneous - ?
3.0 Process Advantages
Lower energy consumption than conventional amine solvent acid gas
removal processes (2).
Can be adapted for removal of all impurities in one pass or for
selective removal (2).
ThiI? conditions are for optimum performance; other input conditions can be
handled with increased solvent losses and reduced efficiency.
B-7
-------
(2)
Production of a product gas with very low water content' '.
(3)
Noncorrosive nature of the solventv '.
(3)
Unlimited solubility of methanol in waterv '.
Chemical stability and low freezing point^ '.
Good for high pressure applications.
4.0 Process Disadvantages
(2)
t Complex flow schemev .
(2)
Solvent carryover losses may be higlv .
Not suited for operation at pressures below 1.1 MPa (165 psia)^ '.
5.0 Process Economics - ?
6.0 Input Streams
6.1 Gaseous
Stream No. 1 - Raw Gas: see Table B-2.
(1]
Stream Nos. 12, 13, and 14 - Types A and B Stripping Gas*v ;:
When used, the stripping gas is nitrogen from an oxygen plant.
Rate: 231,300 - 693,500 Nm3/hr (153,400 - 430,000 SCFM)
Temperature: ?
Pressure: 0.1 - 0.5 MPa (20-80 psia)
6.2 Liquid
0 Stream No. 2 - Water to Separator: quantity ?
Stream No. 3 - Methanol Makeup: quantity ?
7.0 Intermediate Streams
7.1 Gaseous
Stream Nos. 4 and 5 - Types A and B Flash Gases: ?
Stream No. 7 - Type C Intermediate Gas: ?
*This corresponds to the range from Type A and B facilities reported in
Table B-2 from Reference 3.
B-8
-------
TABLE B-2.
Constituents/
Parameters
H *
H2
CO
CH4
co2
N2 + Ar
H2S
COS
cs2
RSH
Thiophene
c2+
MeOH
Temp: °K (°F)
Pressure:
MPa (psia)
Rate: Nm3/hr
(SCFM)
i(5)
Type A
40.05
20.20
8.84
28.78
1.59
4220 mg/Nm3
10 ppm
--
20 ppm
--
0.54
303 (86)
2.5 (380)
381,000
(236,000)
RECTISOL GASEOUS INPUT STREAMS
Stream Number Reference
1<3> i(3) jO)
Type A Type B Type B
__
58-4 62.31 61.59
0.3
0.2
21.9
19.2
--
--
--
--
--
--
2.4 (356)
153,100
(94,300)
3.25
0.17
33.25
0.53
0.49
10 ppm
--
--
--
3.2 (480)
142,340
(88,250)
2.60
0.33
34.55
0.41
0.52
--
--
-_
--
--
7.1 (1066)
137,000
(84,940)
~" =
-
i<4>
Type c
63.74
4.13
0.13
31.62
0.12
0.26
63 ppm
--
__
--
303 (86)
0.3 (45)
80,000
(49,600)
*A11 values, unless otherwise noted, are in volume percent.
B-9
-------
8.0 Discharge Streams
8.1 Gaseous
Stream No. 6 - Product Gas: see Table B-3.
Stream Nos. 7, 8 and 9 - Types A and B Off-Gases: see Table B-4.
Stream Nos. 8, 9, 10 and 11 - Type C Off-Gases: see Table B-4.
8.2 Liquid
Stream No. 10 - Types A and B Still Bottoms: ?
Stream No. 11 - Types A and B Hydrocarbons and Stream No. 12 -
Type C Hydrocarbons: ?
(4^
Stream No. 13 - Type C Still Bottomsv ':
Rate: 16 m3/hr
pH: 9.7
Phenol: 18 mg/1
Cyanide (as CN): 10.4 mg/1 (includes thiocyanate)
Ammonia (as N): 42 mg/1
Sulfides (as S): Trace
Oxygen Absorbed: 286
COD: 1,606 mg/1
Conductivity: 1,111 ymhos/cm
t Stream No. 14 - Type C: ?
9.0 Data Gaps and Limitations
The major limitation in the data is that not all input and dis-
charge streams are characterized and the characterizations are not com-
prehensive in that all potential pollutants and toxicological and ecolog-
ical properties are not identified. An example is the total lack of data
on MeOH carryover.
10.0 Related Programs - ?
B-10
-------
TABLE B-3. RECTISOL PRODUCT GAS STREAMS
Constituents/
Parameters
H2
CO
CH4
co2
N2 + Ar
H2S
COS
cs2
RSH
Thiophene
c2+
MeOH
Temp: °K (°F)
Pressure:
MPa (psia)
Rate: Nm3/hr
(SCFM)
6(5)
Type A
57.30
28.40
11.38
0.93
1.77
0.05 mg/Nm3
total sulfur
--
288(59)
2.3(345)
263,000
(163,000)
' ' .
6<3)
Type A
74.8
0.38
0.25
60 ppm
24.57
--
2.2(327)
118,500
(73,500)
"
Stream No.
- i ,
6<3)
Type B
94.08
4.86
0.24
10 ppm
0.82
--
--
--
--
--
3.0(450)
94,040
(34,300)
~'
.
6(3)
Type B
94.92
3.94
0.47
50 ppm
0.67
1 ppm
__
--
6.9(1037)
88,530
(54,890)
~
~7w
Type C
i -
93.58
6.06
0.19
--
0.17
_ _
--
--
295(72)
2.9(440)
54,500
(33,800)
B-11
-------
TABLE B-4. RECTISOL OFF-GAS STREAMS
Constituents/
Parameters
H2
CO
CH4
co2
N2 + Ar
H2S
COS
C2+
HeOH
cs2
RSH
Thiophene
Temp: °K (°F)
Pressure:
MPa (psia)
Rate: Nm3/hr
(SCFM)
Stream Number Reference
Type A(3) Type B(3) Type B(3) Type B(4) Type C(4)
897 8 9 789 789 11 9 10 8
0.4
0.014
0.017
73.95
25.62*
--
..
-.
..
0.1(15) --
45,090
(27,956)
0.15 0.79
0.04 0.22
0.05
76.81 98.91 64.6
23.0* 0.05 0.1
2 ppm 2 ppm 35.2
0.1
--
0.1(15) 0.24(36) 0.24(36]
41,480 14,130 1980
(25,845) (8,760) (1230)
0.76 --
0.11
0.06 --
90.85 68.31
8.22* -- . 1.92
5 ppm 29.77
--
--
0.1(16) -- 0.2(28)
50,280 2390
(31,170) (1480)
0.33
0.14
0.00
80.19 -- 68.46
19.34* -- --
<5 ppm 30.78
8 ppm -- 0.76
--
..
295(72) - 322(121)
0.1(15) -- 0.5(73)
30,800 673
(19,100) (417)
21.4 2.6 0.14
18.2 4.8 0.0
11.4 7.2 0.9
46.7 83.4 97.2
1.5 0.8 0.03
3176 ppm 4941 ppm 8824 ppm
0.003
0.7 1.1 0.7
--
0.0002
0.028
0.0002
273(32) 273(32) 268(23)
1.3(195) 0.46(70) 0.1(15)
4500 15,000 98,000
(2852) (9,300) (60,760)
CO
I
ro
*Includes N2 stripper gas.
-------
REFERENCES
1 Sinor, O.E., Evaluation of Background Data Relating to New Source Perform-
ance Standards for Lurgi Gasification, EPA-600/7-77-057, June 1977.
2. Kohl, A. and Riesenfeld, F., Gas Purification, Gulf Publishing Co.,
Houston, Texas, 1974.
3. Scholz, W.H., Rectisol: A Low-Temperature Scrubbing Process for Gas Puri-
fication, Advances in Cyrogenic Engineering, Vol. 15, 1969.
4. Draft: Standards Support and Environmental Impact Statement Volume 1:
Proposed Standards of Performance for Lurgi Coal Gasification Plants,
November 1976.
5 Information provided by South African Coal, Oil & Gas Corporation,
Limited, to the Fuel Process Branch of EPA's Industrial Environmental
Research Laboratory (Research Triangle Park), November 1974.
6. Maddox, R.N., Gas and Liquid Sweetening, Campbell Petroleum Series, 1974.
B-13
-------
RECTISOL PROCESS (DUAL ABSORPTION MODE)
1.0 General Information
1.1 Operating Principles - Physical absorption of acid gases (C02, H2S,
COS, CS2, etc.) using methanol. When operated in the dual absorp-
tion mode, C02 saturated methanol is used in the first absorption
step to remove H2S and other sulfur compounds. In the second absorp-
tion step, pure methanol is used for the absorption of C02-
1.2 Development Status - Commercially Available.
1.3 Licensor/Developer - Lurgi Mineraloltechnik GmbH
American Lurgi Corporation
377 Rt. 17 South
Hasbrouck Heights, New Jersey
1.4 Commercial Applications - A Rectisol of this type is installed at
Modderfontein, South Africa for purification of synthetic gas from
coal for manufacture of ammonia.
2.0 Process Information
2.1 Flow Diagram - see Figure B-4 .
Process Description - C02 and H2S are absorbed in separate col-
umns with CO shift occurring between operations. In essence two
separate Rectisol units, each with its own stripper column (but
with common still and regenerator) are employed. C02 saturated
methanol is used to absorb H2S in the first absorber. Pure
methanol from the regenerator is used in the COp absorber.
2.2 Equipment - Conventional absorbers, stripping columns, distillation
columns, heat exchangers and knockout drums.
t Construction - vessels may be fabricated from carbon steel;
dimensions dependent on application.
2.3 Feed Stream Requirements - ?
B-14
-------
en
01
LEGEND:
1. RAW GAS
2. N, STRIPPER GAS
3. N2 STRIPPER GAS
4. MnOH MAKE-UP
5. H2S SCRUBBER GAS
6. INPUT TO CO2 REMOVAL
7. PRODUCT GAS
8. LEAN H2S FROM NO. 1 STRIPPER
9. LEAN H2S FROM NO. 2 STRIPPER
10. COMBINED LEAN HjS
11. CONCENTRATED H2S
12. PUHECOj
13. CONDENSATE
14. COj SATURATED METHANOL
15. PURE METHANOL
16. LEAN METHANOL
o
N t
GC
<
Q.
LU
|
I | «
Oi
V5 »-
1 41
1 1
1-
cc
LU
z
LU
0
CC
l_
-c|
>J
1
1
«
^
IN
o
CC
LU
CL.
0.
DC
I
W3
|
1 m
1
u
-e-
Figure B-4. Rectisol - Dual Absorption Flow Diagram (as installed at Modderfontein, South Africa)
-------
2.4 Operating Parameters'1'2'3'
Absorption - H2S: 297°K (75°F) 3.0 MPa (440 psia)
C02: 213°K (-75°F) 4.9 MPa (720 psia)
t Regeneration - ?
2.5 Process Efficiency and Reliability - Removal of acid gases to a few
micrograms per cubic meter. Reliability is high due to relatively
simple operation.
2.6 Raw Material Requirements
Solvent - Methanol
2.7 Utility Requirements - Utility requirements are high due to large
refrigeration requirements. Exact amounts are unknown.
2.8 Miscellaneous - ?
3.0 Process Advantages
A single solvent (methanol) is used for absorption of both C02 and f^S.
Noncorrosive environments.
H2S streams rich enough to be processed in a Claus unit can be obtained.
Good selectivity between acid and product gases.
t Unlimited solubility of solvent in water.
Solvent is chemically stable and has a low freezing point.
4.0 Process Limitations
Solvent retains heavy hydrocarbons.
Solvent losses during regeneration may be high.
High utility requirements.
5.0 Process Economics - ?
6.0 Input Streams
All stream data based on the Modderfontein plant.
B-16
-------
6.1 Gaseous
6.1.1 Stream No. 1
Composition, wt % Ref. 1 Ref. 2
co2
CO
H2
N2
Ar
CH4
H2S
COS
MeOH
3
Volume Nm /(scfm)
11.6
55.02
31.2
1.0
0.5
0.1
0.5
0.8
0
91,700
(53,370)
13.37
54.45
30.00
0.95
0.54
0.10
0.59
(includes COS)
0
Pressure, MPa (psia)
Temperature, °K (°F)
6.1.2 Stream Nos. 2 and 3 - Nitrogen from air separation plant, rate
unknown.
6.2 Liquid
Stream No. 4 - Methanol makeup, rate unknown.
7.0 Intermediate Streams
7.1 Gaseous
7.1.1 Stream No. 5
Composition, wt % Ref. 1 ReJUL
CO, 12.00 11-27
CO
'2 54-60 56.02
B-17
-------
Composition, wt %
Ar
CH4
H2S
COS
MeOH
o
Volume, Nm /hr (scfm)
Pressure, MPa (psia)
Temperature, °K ( F)
7.1.2 Stream No. 6
Composition, wt %
co2
CO
H
Ar
CH4
H2S
COS
MeOH
3
Volume, Nm /hr (scfm)
Pressure, MPa (psia)
Temperature, °K (°F)
7.1.3 Stream No. 8 - ?
7.1.4 Stream No. 9 - ?
8.0 Discharge Streams
8.1 Gaseous
Ref. 1
Ref. 2
31.80
1.00
0.50
0.10
31.06
0.98
0.57
0.10
93,300
(58,370)
3.0(440)
298(75)
Ref. 1
41.30
3.00
54.64
0.70
0.30
0.06
*
--
-
Ref. 2
41.29
3.00
54.63
0.64
0.37
0.07
140,000
(87,590)
5.0(735)
308(95)
B-18
-------
8.1.1 Stream No. 7
Composition, wt % Ref. i
co2
CO
4.60
93.50
1.20
0.60
0.10
5.02
93.14
1.12
0.61
0.11
Ar
CH4
H2S
COS
MeOH
Volume, Nm3/hr (scfm) 80,000
(50,110)
Pressure, MPa (psia) 4.9(720)
Temperature, °K (°F) 213(-75)
8.1.2 Stream No. 10 - ?
8.1.3 Stream No. 12 - Mostly C02> trace constituents unknown.
8.1.4 Stream No. 11
Composition, wt % Ref. 1 Ref. 2
C02 75.00
CO
Ar
CH4
H2S 22.00
COS 3.00
MeOH
Volume, Nm3/hr (scfm) 21,000
(13,140)
Pressure, MPa (psia) 0.1 (15)
Temperature, °K (°F) 313(105)
B-19
-------
8.2 Liquid
Stream No. 13 - ?
9.0 Data Gaps and Limitations
Limitations in the data for the selective absorption Rectisol
relate primarily to the stream compositions. These limitations include
the following:
Input gas streams - little data on minor component concentrations.
No data on N2 stripper gas rates.
Makeup methanol - no data on amount of makeup methanol required.
Intermediate and product gas streams - limited data on minor
components.
t Discharge gas streams - very limited data on compositions of offgas
streams from the strippers and regenerator.
Condensate stream - no data on compositions and rates of regenerator
condensate stream.
t Operating parameters - utility requirements, regeneration parameters,
etc. are not reported.
10.0 Related Programs
No known programs are presently being undertaken to assess the
discharges from this process.
REFERENCES
1. Staege, Hermann, Ammonia Production on the Basis of Coal Gasification,
Chemical Industry Developments, 1973.
2. Schellberg, Wolfgang, Coal Based Ammonia Plants, ICI Operating Symposium
1974, Paper 21.
3. Goeke, E.K., Status of Coal Gasification Technology, FAI Symposium on Coal
as Feedstock for Fertilizer Production, New Delhi, India, 1974.
B-20
-------
SELEXOL PROCESS
1.0 General Information
1.1 Operating Principle^1' - The physical absorption of the sour compo-
nents (H2S, C02, COS, mercaptans, etc.) of a gas stream using the
Selexol (dimethyl ether of a polyethylene glycol ) solvent.
1.2 Development Status - Commercially available.
1.3 Licensor/Developer - Allied Chemical Corporation
Gas Purification Department
P.O. Box 1013R
Morristown, New Jersey 07960
1.4 Commercial Applications
Eleven commercial plants have been put into operation. Applications
include: The purification of sour natural gas; synthesis gas from
(2)
the gasification of coal, oil and light hydrocarbons v '.
Applicability to coal gasification^3'5': The Bi-Gas pilot plant of
Bituminous Coal Research, Inc. at Homer City, Pennsylvania incorpo-
rates the Selexol process to remove the H2S and C02 components of the
gas from the CO-shift.
2.0 Process Information
2.1 Flow Diagram^ - see Figure B-5 for one Selexol design for treatment
of coal or oil gasification product cases.* Sour feed gas (Stream 1)
containing H2S and C02 enters the H2S absorber, where H2S is absorbed
selectively to maximize its content in the off-gas to a sulfur recov-
ery plant (Stream 2). Solvent is regenerated by reboiled steam. The
essentially sulfur-free product then enters the C0£ absorber where
for nonselective acid gas removal are different than the one shown in
Figure B-5.
B-21
-------
CO
I
ro
r\»
INDEX TO STREAMS
1. FEED GAS
2. CONCENTRATED HjS STREAM
3. PURIFIED GAS
4. CONCENTRATED COj STREAM
5. STRIPPING GAS (AIR OR NITROGEN)
Figure B-5. Selexol Solvent Process (with three flashing stages)
-------
C02 is removed so that the final product gas (Stream 3) is suitable
for feed to a methanator. The solvent is regenerated by pressure
letdown and inert gas stripping (air or nitrogen). Hydraulic letdown
turbines extract energy from the high pressure solvent streams during
letdown. Recycle loops return coabsorbed product to the absorbers,
enhancing selectivity.
2.2 Equipment
Conventional absorbers, flash vessels, and stripping columns.
Absorbers and strippers employ both packing and trays.
2.3 Feed Stream Requirements
Pressure: 2.0-10.7 MPa (300-1500 psia) typical
Temperature: Usually air cooled or exchanged
2.4 Operating Parameters
2.4.1 Absorption Step
Pressure: 2.0-10.7 MPa (300-1500 psia) typical
Temperature: 270°K-310eK (20°F-100GF) typical
2.4.2 H2S Regeneration Step
Pressure: 2.0-10.7 MPa (300-1500 psia)
Temperature: Up to 450°K (350°F)
2.4.3 (XL Regeneration Step
Pressure: 0.14 MPa (20 psia) typical
Temperature: 255°|<-450°K (0°F-350°F)
2.5 Process Efficiency and Reliability
Process can reduce H2S, COS and mercaptans concentrations to less
than 1 ppm each. The C09 level can be reduced to any desired level
(5}
by adjusting the solvent mix^ '.
In over thirteen years of commercial application, the process has
been reported to be dependable, flexible, and relatively maintenance
free(2,5)_
B-23
-------
2.6 Raw Material Requirements^ '
Solvent makeup: 8g/1000 Nm3 (0.5 Ib/MMSCF} of gas.
2.7 Utility Requirements^5' - Typical requirements when treating gas con-
taining 0.5% H2S and 35% C02 at ,3. 5 MPa (500 psig) with <0.01 ppm
H2S and 11% C02 in purified gas are as follows:
Steam: 50,800 kg/106 Nm3 (3000 Ib/MMSCF)
Cooling Water: 4.8 x 106 £/106 Nm3 (3500 gallons/MMSCF)
Electricity: 33,500 kwh/106 Nm3
Economically attractive for bulk removal of CO,^ ',
2.8 Miscellaneous
Solvent discharged from stack or from drips and spills will readily
remove paint; therefore, good housekeeping is required.
Possible problems due to build-up of inert materials; e.g., glycol,
compressor oil, and heavy hydrocarbons in solvent.
3.0 Process Advantages
t Economic removal of H2S of <1 ppnr .
Economically attractive for bulk
Selective removal of COS to <1 ppm.
NHg can be reduced to low ppm levels^ '.
( 2 &}
Feed gas can be varied over a broad range in existing equipment^ ' .
Regeneration can be accomplished by flashing, inert gas stripping
and/or heat treatment (2, 7).
t Solvent is noncorrosive, nonfoaming, nontoxic, and biodegradable.
Low heat of absorption, low specific heat and low vapor pressure mini-
mize solvent losses(2).
t Process is highly selective to sulfur compounds and yields a high sul-
fur feed gas to sulfur recovery(2). (See Table B-5 for relative solu-
bility of selected chemicals in the Selexol solvent.)
t
Process dehydrates gas during H2S/C02 removal
(8)
Process reduces heavy hydrocarbon content of feed gas to meet hydrocar-
bon dewpoint(S).
B-24
-------
Solvent is not degraded by impurities in the feed gas; thus, no sol-
vent reclaimer is required(S). 3 ' u b01
TABLE B-5. RELATIVE ORDER OF SOLUBILITIES OF GASES
IN SELEXOL SOLVENTO)
1. HCN (most)
o HO
3. C4H4S
4. S02
5. C6H6
6. Cy
7. CS2
8. CH3SH
9. Cg
10. H2S
11.
12.
13.
14.
15.
16.
17.
18.
19.
20.
4.0 Process Limitations
0 Hydrocarbons are absorbed proportional to their
(See Table B-5.)
NH3
C5
COS
C4
C3
co2
c2
cl
CO
H2 (least)
partial pressures*^ ' '
0 Not designed to treat gas at low pressures and/or with low acid gas
concentrations I5).
Solvent is expensive ($8.60/gallon, 1976)
(3)
0 Selectivity of Selexol Process for H^S over CO? is somewhat decreased
if COS is also to be removed along with H2S(8).
5.0 Process Economics* '
The estimated cost of treating 2,830,000 Nm3 (100 MMSCF) of gas con-
taining 43% C02 and no H2S at a temperature of 291°K to 322°K (65 F to
120°F) and pressure of 6.22 MPa to 7.58 MPa (900 to 1000 psig) is
$3.66/103 Nm3 (9.8 cents/103 SCF).
B-25
-------
6.0 Input Stream^3'
6.1 Feed Gas (Stream 1) - See Tables B-6 and B-7.
6.2 Stripping Gas (Stream 5) - Air or nitrogen, no quantitative data
available.
TABLE B-6. OPERATING CONDITIONS AND MOLAR BALANCE FOR THE SELEXOL PROCESS
FOR REMOVAL OF C02 FROM NATURAL GAS(4)
Operating Conditions
Volumetric Flow
Rate -Nm3/d (MMSCFD)
Pressure, MPa (psia)
Temperature, °K (°F)
Components
co2
N2
CH4
C2H6
C3H8
H2S
Total
C02 Mole %
CH4 Mole %
hLS ppm
Stream 1
Feed Gas
2.83 x 106 (100)
6.8 (100)
302 (85)
moles/day
116,095.0
1,556.9
144,337.3
1,447.8
369.4
15.8
263,852.2
44
54.7
60.0
Stream 3
Purified Gas
1.58 x 106 (56)
6.63 (975)
297 -(75)
moles/day
4,094.1
1,556.0
140,895.3
1,183.0
107.5
0.8
147,836.7
2.8
95.3
5.4
Stream 4
Acid Gas
1.25 x 106 (44)
0.11 (16)
297 (75)
moles/day
112,000.9
0.9
3,442.0
294.8
261.9
15.0
116,015.5
96.5
3.0
129.3
7.0 Discharge Streams^ '
7.1 Purified Gas Stream (Stream 3)
7.2 Concentrated H2S Stream (Stream 2)
7.3 Concentrated C02 Stream (4)
B-26
See Tables B-6 and B-7
-------
TABLE B-7. STREAM COMPOSITIONS FOR SELEXOL PROCESS DESIGNED FOR SFIFnn/r
ACID GAS REMOVAL FROM COAL GASIFICATION PRODUCT GAS(8)
Component*
co2
H2S
CH4
H2
CO
Stream 1
Feed Gas
31
7000 PPMV
8
46
15
Stream 3
Purified Gas
0.5
<1 PPMV
11
67
22
L=
Stream 2
Concentrated HpS
68
32
--
--
. .
Stream 4
Concentrated C0?
98
5 PPMV
--
2
*Volume % unless specified otherwise.
8.0 Data Gaps and Limitations
Data gaps exist in the following areas:
Applicability to coal conversion processes:
- Process reliability and efficiency
- Feed stream" requirements with regard to the concentrations of
various contaminants (e.g., COS, HCN) temperature and pressure.
- The effect that various contaminants (NH3, HCN, CS2, trace metals,
etc) have on the process, and the ultimate fate of such contami-
nants in the system.
Operating parameters and feed stream requirements in refinery and
natural gas processing:
- Applicable temperature, pressure and-concentration ranges for
feed gas streams and appropriate operating conditions for various
steps in the process (absorption step, regeneration step, etc.).
- The effect that various contaminants (NH3, carbonaceous matter,
trace metals, etc.) have on the process, and the ultimate fate
of such contaminants in the system.
9-0 Related Programs
The joint ERDA/AGA funded Bi-Gas Pilot Plant, operated by Phillips
Petroleum, at Homer City, Pennsylvania has a Selexol system incorporated
B-27
-------
in its design. The Selexol system removes the H^S and C(L components
of the gas from the CO-shift.
The original contract has expired, and at this time Phillips is
awaiting further funding from ERDA prior to continued operation of the
Homer City facility.
REFERENCES
1. Riesenfeld, F.C., and Kohl, A.L., Gas Purification, Second Edition, Houston,
Texas, Gulf Publishing Co., 1974.
2. Gas Processing Handbook, Hydrocarbon Processing, Vol. 54, No. 4, April 1975.
3. Valentine, J.P., New Solvent Process Purifies Crude, Coal Acid Bases, The
Oil and Gas Journal, 18 November 1974.
4. Raney, D.R., Remove Carbon Dioxide with Selexol, Hydrocarbon Processing,
April 1976.
5. Handbook of Gasifiers and Gas Treatment Systems, Dravo Corp. for ERDA,
FE-1772-11, February 1976.
6. Hegwer, A.M., and Harris, R.A., Selexol Solves High H?S/C02 Problem,
Hydrocarbon Processing, April 1970.
7. Maddox, R.N., Gas and Liquid Sweetening, Campbell Petroleum Series, 1974.
8. Information supplied to TRW by J. P. Valentine of Allied Chemical Corpora-
tion, June 28, 1978.
B-28
-------
PURISOL PROCESS
1.0 General Information
1.1 Operating Principles - The physical absorption of the sour components
(e.g., C02 and H2S) of a gas stream using N-methyl-2-pyrrolidone
(NMP)(5,6).
1.2 Development Status - Commercially available.
1.3 Licensor/Developer - American Lurgi Corporation
377 Route 17
Hasbrouck Heights, New Jersey 07604
M 9\
1.4 Commercial Applications^ ' '
Four plants are in operation:
- Two in high pressure hydrogen manufacturing
- Two in natural gas treating
2.0 Process Information^ '
2.1 Flow Diagram (see Figure B-6) - Feed gas, Stream 1, enters the
absorber where it is dehydrated with a slipstream of rich NMP, and
then scrubbed with regenerated NMP. Lean NMP enters the top of the
absorber. The feed gas passes counter-flow to the lean NMP solution.
Entrained NMP is recovered by a water wash before the purified gas,
Stream 2, exits the top of the absorber. Rich solvent is flashed at
high pressure in the lower section of the absorber. The gases
evolved in this flashing step are separated, recompressed, cooled
and recycled to the feed gas stream.
The rich solvent -exits the bottom of the absorber; it is cooled and
piped to a stripping column. It is regenerated in this column by
two-stage flashing at atmospheric pressure. Acid gases are evolved
during this step. The gases evolved are separated and piped to sul-
fur recovery. The lean solvent is pumped to the absorber.
B-29
-------
CO
o
FEED GAS
2. PURIFIED GAS
3. ACID GAS
4. HIGH PRESSURE FLASHING
5. RECYCLED HIGH PRESSURE
FLASH GAS
6. RICH SOLVENT
7. LEAN SOLVENT
8. ACID GAS STREAM
9. SEMI-STRIPPED SOLVENT
Figure B-6. Purisol Process
-------
The NMP/water mixture from the absorber and stripper are combined
and sent to the solvent dryer. Acid gases evolved during the drying
step are combined with the acid gas stream from the stripper to form
Stream 3, and piped to sulfur recovery. Dehydrated NMP is returned
to the stripper.
2.2 Equipment - Conventional absorbers, stripping columns and flash
vessels.
2.3 Feed Stream Requirements
t Temperature: ?
t Pressure: ?
Others: ?
2.4 Operating Parameters
2.4.1 Absorption Step
Temperature: ?
Pressure: ?
Others: ?
2.4.2 Stripping Step
Temperature: ?
Pressure: ?
t Others: ?
2.4.3 Solvent Drying Step
Temperature: ?
Pressure: ?
Others: ?
2.5 Process Efficiency and Reliability^
Process can reduce H2S concentration to 4 ppm and C02 concentra-
tions to 2-3 vol %.
2.6 Raw Material Requirements^ '
. Solvent makeup 35 g/1000 Nm3 (2.1 Ibs/MMscf) acid gas treated.
B-31
-------
2.7 Utilities Requirements^ ' - Typical requirements for a plant operating
with feed conditions of 2,830,000 Mm /day (100 MMscfd) at a pressure
of 7.4 MPa (1070 psig) and temperature of 317°K (110°F) are:
Electric power: 2100 kW
Steam: 4.06 MPa (45 psig) 1.7 tonne/hr (1.87 ton/hr)
Cooling water: 297°K (75°F) 300 m3/hr (10,600 ft3/hr)
Condensate: 1.3 tonne/hr (1.43 ton/hr)
Solvent makeup: 3 kg/hr (6.6 Ib/hr)
2.8 Miscellaneous - ?
3.0 Process Advantages^3'5'6)
Solvent is noncorrosive and nonfearning
t Low vapor pressure minimizes solvent losses
Solvent is readily available
t Solvent preferentially absorbs sulfur compound and CO,,.
4.0 Process Limitations
,(4)
At pressures of 2.8 MPa (400 psig) and lower, process is uneconomical
iole<
5.0 Process Economics^ '
Absorbs high molecular weight hydrocarbons^ ',
Typical requirements per 1000 Nm (MMscf) for a feed gas containing
6 vol. % H2S and 15 vol.% C02 at 7.38 MPa (1070 psig) with 2 ppm H2S and
13.6 vol. % C02 in purified gas are:
Steam: 50 kg/1000 Nm3 (3125 Ib/MMscf)
Cooling water: 1780 £/1000 Nm3 (13,300 gal/MMscf)
t Electric power: 9.33 kWh/1000 Nm3 (264 kwh/MMscf)
Solvent loss: 35 g/1000 Nm3 (2.1 Ib/MMscf)
6.0 Input Stream
6.1 Inlet Gas Stream (Stream 1) - see Table B-8.
B-32
-------
TABLE B-8. PURISOL GAS STREAM DATA*
Component
H2
co2
CO
cl
No + Ar
H2S
Temperature
Pressure
Flow Rate
Stream 1
Sour Gas
64.53%
15% 33,15%
1,50%
0.44%
0.38%
6%
7.38 MPa 317°K
1.38 MPa
2.83 x 106 Nm3
"'"
Stream 2
Purified Gas
Ref. (1) Ref. (2)
96.44%
0.10%
2.24%
0.59%
0.63%
Stream 3
Acid Gas
Ref. (1) Ref. (2)
13.6%
2 ppm
*The information contained in this table is for two different applications.
Reference 1 is natural gas and reference 2 is for syngas. No complete
description of all streams; e.g., sour gas, purified gas, and acid gas, for
each type of application was given.
7.0 Discharge Stream
7.1 Purified Gas Stream (Stream 2) - Table B-8.
7.2 Acid Gas Stream (Stream 3)
7.3 Solvent Slowdown
8.0 Data Gaps and Limitations
Several gaps exist, they are as follows:
No data on maximum allowable concentrations of various contami-
nants; e.g., CS2, COS, mercaptans.
. No data on removal efficiency for various contaminants at various
concentrations.
B-33
-------
No information on reliability and maintenance characteristics of
process and facility.
No information which would indicate applicability to coal gasifi-
cation process.
Characterization of gaseous and liquid streams for refinery/
natural gas applications (temperature, pressure, composition,
etc.).
The effect that various contaminants (NH3, carbonaceous matter,
trace metals, etc.) have on the process, and the ultimate fate of
such contaminants in the system.
9.0 Related Programs
No information available.
REFERENCES
1. Handbook of Gasifiers and Gas Treatment Systems, Dravo Corp. for ERDA,
FE-1772-11, February 1976. p 120-121.
2. Gas Processing Handbook, Hydrocarbon Processing, Vol. 54, No. 4, April 1975.
3. Hochagesand, G., Rectisol and Purisol, Industrial and Engineering Chemistry,
Vol. 62, No. 7, July 1970.
4. Beavon, O.K. and T.R. Roszkowski, Purisol Removes CO? from Hydrogen,
Ammonia Syngas, The Oil and Gas Journal, 14 April 1969.
5. Kohl, A.L.. and F.C. Riesenfeld, Gas Purification, Gulf Publishing Company,
6. Maddox, R.N., Gas and Liquid Sweetening, Campbell Petroleum Series, 1974.
B-34
-------
ESTASOLVAN PROCESS
1.0 General Information
1.1 Operating Principle(1) - The physical absorption of the sour compo-
nents of a gas stream (H?S, CO,,, C02, COS, RSH, etc.) using an aqueous
solution of tri-n-butylphosphate (TBP) as the solvent.
1.2 Development Status^' - Commercially available.
1.3 Licensor/Developer - Institut Francais du Petrole
1 et 4, av. de Bois-Preau
92-RUEIL-MALMAISON
(.Hauts-de-Seine) France
1.4 Commercial Application'1'3'*
Natural gas desulfurization
Natural gas desulfurization and liquid hydrocarbon recovery.
2.0 Process Information
2.1 Flow Diagranr ' - see Figure B-7.
t Sour feed gas, Stream 1, enters the bottom and lean solvent enters
the top of the absorber. The sour gas passes counter-flow to the
solvent. The purified product gas, Stream 2, exits the top, and
the rich solvent exits the bottom of the absorber. The rich sol-
vent is piped to the flashing step. Acid gas exits the top of the
flashing vessel and the semi-lean solvent exits the bottom of the
vessel. The solvent is then piped to a regenerator where the
remaining acid components are boiled off. These components exit
the top of the regenerator and combine with the acid gas stream
from the flash vessel; the combined stream is then piped to sulfur
*As of May 1970 this process was untried in commercial applications. However,
extensive pilot plant operation had been conducted in both France and West
Germany. No information is available describing this process in a commercial
operation.
B-35
-------
CO
1
CO
tr
LU
00
tr
O
V)
00
QC
I
QC
Ul
EC
LEGEND:
1. SOUR FEED GAS
2. PURIFIED PRODUCT GAS
3. ACID GAS TO SULFUR PLANT
4. RICH SOLVENT
5. LEAN SOLVENT
6. SEMI-LEAN SOLVENT
7. ACID GAS
8. SOLVENT SLOWDOWN
Figure B-7. Flow Diagram Estasolvan Process
-------
e"1ts the ~»"«"«t»p and is pumped
2.2 Equipment
t Conventional absorbers, flash vessels and stripping columns.
2.3 Feed Stream Requirements
Pressure: 1.0 MPa - 20 MPa (10-200 atm)
t Temperature: 253°K - 333°K (-40°F - 140°F)
Others: ?
2.4 Operating Parameters
2.4.1 Absorption Step
Temperature: 283°K to 303°K (50°F to 86°F)
Pressure: 2.0 to 10.0 MPa (20 to 100 atm)
Others: ?
2.4.2 Regeneration Step
Temperature: 253°K to 333°K (-4°F to 140°F)
Pressure: 1.0 to 20.0 MPa (10 to 200 atm)
Others: ?
2.4.3 Stripping Step
Temperature: 373°K to 423°K (212°F to 302°F)
Pressure: 1 MPa to 0.5 MPa (1 to 4.8 atm)
2.5 Process Efficiency and Reliability^ '
t
Pilot plant and design data indicate that process can reduce H2S
component of purified gas to 3 ppm. See Table B-9.
No information is available on reliability of a commercial plant.
(3\
2.6 Raw Material Requirements^ ;
Inert gas: 1200 Nm3/hr (42,360 scf/hr)
B-37
-------
TABLE B-9. STREAM DATA FOR TYPICALLY DESIGNED ESTASOLVAN PROCESS
(1)
Composition
H2S
COS- (mg/Nm3)
RSH- (mg/Nm3)
co2
N2
CH4
c2+
Percent by Volume
Feed Gas Purified Gas Acid Gas
Stream 1 Stream 2 Stream 3
10.0
500
1500
7
7.5
75.5
Traces
3 ppm
6
50
6.4
8.0
85.6
--
85.75
0.65
11.40
2.20
--
Gas Flow Rate - 4 x 106 Nm3/d (115 MMscfd)
Pressure - 6.9 MPa (1000 psig)
i o\* 3
2.7 Utility Requirementsv ' - For a plant processing 75,000 Nm /hr
(2.65 MMSCF/hr) of sour gas at a pressure of 7.0 MPa (995 psia) and
temperature of 303°K (86°F)
Steam at 0.34 MPa (35 psig) 1996 Kg/hr (4400 Ib/hr)
Cooling water 75,700 liters (20,000 gph)
Electric power 1200 kwh
2.8 Miscellaneous - No information available.
3.0 Process Advantages'1'3^
Solvent has low specific gravity
Solvent has low viscosity even at low operating temperatures
Solvent has low vapor pressure
Solvent has low foaming tendency
*This information is taken from pilot plant operation. No information avail-
able for actual operating commercial facility.
B-38
-------
t Solvent has no toxicity
Solvent is noncorrosive.
^ '
4.0 Process Limitations
(lo
<40
Gas with high C02 content may require a separate C02 removal process
Under normal operating conditions solvent absorbs hydrocarbons As
process is adjusted for high C02 removal by increasing the TBp'concen-
tration in the absorption solvent the amount of Ci, c?, and Co
removed will be increased. 6
The gas stream compositions which are most efficiently treated by the
Estasolvan process are given in Table B-10.
5.0 Process Economics^ '
Studies indicated that this process will be 10 to 15 percent less
expensive than chemical absorption, provided the partial pressure of
H2S in the feed gas is in excess of 0.27 MPa (40 psia).
6.0 Input Streams^ ' - see Figure B-7.
Sour feed gas (Stream 1) - see Table B-9.
7.0 Discharge Streanr ' - see Figure B-7.
0 Purified gas stream (Stream 2) - see Table B-9.
Acid gas stream (Stream 3).
TABLE B-10. GAS STREAM MOST EFFICIENTLY TREATED BY THE
ESTASOLVAN PROCESS(2)
Constituents
Methane
Hydrogen Sulfide
Ethane and higher hydrocarbons
Other gases such as steam, nitrogen, C02, mercaptans,
CO, COS and hydrogen
Vol. %
50-90
1-25
1-25
0-40
^responds to gas at operating pressure of 6.9 MPa (1000 psig) with H2S con-
tent of 4 to 5 percent.
B-39
-------
8.0 Data Gaps and Limitations
Data supplied above is from pilot plant operation and design informa-
tion; no information was obtained for a commercial plant. Due to this
fact, reliability and commercial process plant efficiency are not
reported.
Data gaps exist in the following areas:
- Applicability of the process to coal conversion processes; e.g.,
efficiency, reliability, and feed stream requirements. Character-
ization of gaseous and liquid streams (e.g., purified gas, feed gas,
acid gas, and sludge) for refinery/natural gas commercial'
applications.
- Definition of the maximum allowable concentrations of various con-
taminants in the feed gas; e.g., COS, CS~, NFL, and mercaptans.
- The effect that various contaminants (NH3, carbonaceous matter, trace
metals, etc.) have on the process, and the ultimate fate of such
contaminants in the system.
9.0 Related Programs
No data available.
REFERENCES
1. Franckowiak, S. and Nitschke, E., Estasolvan: New Gas Treating Process,
Hydrocarbon Processing, May 1970.
2. Strecher, P.G., Hydrogen Sulfide Removal Processes, Noyes Data Corporation,
3. New Process Has Wide Scope, Oil and Gas Journal, 20 May 1968.
B-40
-------
FLUOR SOLVENT PROCESS
1.0 General Information
1.1 Operating Principles(1) - The physical absorption of the sour compo-
nents (C02, H2S, etc.) from a gas stream using propylene carbonate
as the sorbent.*
1.2 Development Status - Commercially available.
1.3 Licensor/Developer - The Fluor Engineers and Constructors
3333 Michel son Drive
Irvine, California 92730
1.4 Commercial Applications - Process is presently used in 9 plants^6'.
t Five in natural gas application
One in hydrogen production
Three in ammonia production
Table B-ll gives a list of seven plants using the Fluor process with
their respective owners and locations.
2.0 Process Information
2.1 Flow Diagram (see Figure B-8)'3' - The sour feed gas, Stream 1, and
hydrocarbon containing flashed gas, Stream 6, are combined and
injected at the bottom of the absorber. The lean solvent enters at
the top of the absorber. Purified product gas, Stream 2, and rich
solvent, Stream 4, exits the top and bottom of the absorber, respec-
tively. The solvent is let down through hydraulic turbines. The
acid gases containing some evolved hydrocarbons, Stream 6, are
*The process is primarily used for the removal of C02 and O^-^S from high-
Pressure natural gas or synthesis gas. However, by proper selection of the
operating conditions, it can also be used for selective removal of H2S from
gases that contain both C02 and H2S.
B-41
-------
TABLE B-ll. SITES WHERE FLUOR SOLVENT IS USED IN GAS TREATMENT
(2)
Location
Lost Hills,
California
Grand County,
Utah
Trinidad
Terrell County,
Texas
Pecos County,
Texas
Bedrijven,
Belgium
Pascagoula,
Mississippi
Woodward,
Oklahoma
Pecos County,
Texas
Plant Owner
Standard Oil Co.
of California
Sinclair Oil and
Gas
W.R. Grace
El Paso Natural
Gas
El Paso Natural
Gas
Union Chemique
Chemische
Standard Oil Co.
W.R. Grace
Intratex Gas
Company
Application
Natural gas treating
Natural gas CO,,
removal
Synthesis gas (MH.J
C02 removal
Natural gas C02
removal
Natural gas treating
Synthesis gas sulfur
removal
Hydrogen plant
Synthesis Gas (MH3)
C0~ removal
Natural gas
Through-Put
230,000 Nm3/d
(10 MMSCFD)
566,000 Nm3/d
(20 MMSCFD)
1909 Tonnes
MH3/d
(1200 tons/day)
Not reported
2.5 x 106
Nm3/d
(88 MMSCFD)
Not reported
Not reported
1090 Tonnes
MH3/d
(1200 tons/day)
2.8 x 106
nm-Vd
(100 MMSCFD)
recycled to the absorber and the rich solvent is sent to the strip-
ping column for regeneration.* The regenerated solvent, Stream 5,
is pumped back to the absorber, and the acid gas evolved in the
stripper, Stream 3, is sent to the sulfur recover, flared or vented,
depending on acid gas composition.
2.2 Equipment - Conventional absorbers, flash vessels and stripping
columns.
*te1inii?L?pdtl!nrfafhi1S?t1hn/nd,0perat1ng conditions, the stripping step may
reuse S°lvent d1rectly returned to the absorber for
B-42
-------
00
-f=»
oo
fr"
LEGEND:
1. SOUR GAS FEED
2. PURIFIED PRODUCT GAS
3. ACID GAS
4. RICH SOLVENT
5. LEAN SOLVENT
6. RECYCLED FLASHED GAS
CONTAINING HYDROCARBONS
7. INERT GAS (E.G.. N2)
Figure B-8. Fluor Solvent Processes
-------
2.3 Feed Stream Requirements^
Temperature: Not critical
Pressure: Acid gas partial pressure > 0.5 MPa (75 psia)
2.4 Operating Parameters^ '
t Absorption Step:
Temperature: Ambient or below
Pressure: Acid gas partial pressure > 0.5 MPa (75 psia)
Solvent loading: ?
Stripping Column:
Temperature: Typically 269°K (25°F)
Pressure: Atmospheric
(1 4)
2.5 Process Efficiency and Reliabilityv ' '
In treating natural gas, the process can produce a product of
pipeline specifications.
Over 15 years of commercial application.
2.6 Raw Material Requirements
Chemicals: (Solvent makeup, etc.)
Inert gas (e.g., N2):
2.7 Utility Requirements
Water: Little or none
Electricity: Normally low, depends upon acid gas content
Steam: ?
2.8 Miscellaneous - Available information indicates no unusual mainte-
nance problems, or potential hazardous conditions created by process.
3.0 Process Advantages
Feed gas is dehydrated as acid gas is removed.
Low vapor pressure at operating temperature which minimizes solvent
losses(4).
Solvent has a low viscosity which minimizes pumping costs'3'.
Solvent has good thermal and chemical stability'3'.
B-44
-------
all°WS the use °f "rt°" ««1 throughout
Solvent is readily available'2^.
Solvent reclamation is not required.
4.0 Process Limitations
t Solvent absorbs heavy hydrocarbons^'.
Refrigeration may be required because solvent carrying capacity is
increased at lower temperatures (see Table B-12)(3). However, partial
or complete energy requirements may be provided by flashing of the
acid gasv°;.
t Inert gas stripping, or vacuum flashing may be required for solvent
regeneration(3).
5.0 Process Economics
The cost of treating a synthesis gas containing 32 vol. % OL is
approximately $1.70/tonne ($1.55/ton) of carbon dioxide removed* in 1969
dollars^.
TABLE B-12. SOLUBILITY OF C02 IN PROPYLENE CARBONATE
(3)
Mm3 C02/m3 soln at 300°K
(scf C02/cf soln at 80°F)
Mm3 C02/m3 soln at 278° K
(scf C02/cf soln at 40°F)
C02 Pressure MPa (psig)
0.51 (60)
16 (17)
26 (26.5)
1.5 (200)
50 (52.5)
89 (90.0)
2.8 (400)
102 (107)
311 (327)
4.2 (600)
212 (223)
00
licensor cautions that these data are not to be used for design purposes
Tills cost includes royalty fees, capital investment, utilities cost, labor,
and various fixed charges.
B-45
-------
6.0 Input Streams
Sour feed gas stream, Stream 1, see Table B-13.
Inert gas stream, Stream 7. Only required to achieve very low acid
gas in the treated gas.
7.0 Discharge Stream
Purified product gas (Stream 2) - See Table B-13.
Acid gas stream (Stream 3) - No data available.
TABLE B-13. EXAMPLE STREAM DATA FOR FLUOR PROCESS*
(4)
Plant Capacity
Feed Gas Pressure
Feed Gas Composition:
co2
N2
CH
C2H6
Mercaptans
H20
Treated Gas Specifications:
CO,,
H20
2150 Nm3/min (110 MMSCFD) feed gas
5.5 MPa (800 psig)
53.2% by volume
0.7% by volume
45.7% by volume
0.4% by volume
nil
nil
294°K (70°F) dewpoint
2% maximum (without stripper)
273°K (32°F) maximum dewpoint
*These requirements are for plants designed to primarily reduce C09 content
of purified gas. *
B-46
-------
8.0 Data Gaps and Limitations
Applicability to coal conversion processes.
Applicable feed gas stream requirements
- Temperature and pressure ranges
- Concentration of various contaminants
0 The effect that contaminants such as NH3, carbonaceous matter trace
minerals, etc., have on the process and the ultimate fate of such con-
taminants in the system.
t Operating parameters of various stages of the process.
t Characterization and volume of various discharge streams (e.g., blow-
down, acid gas).
Process Economics
- Typical utilities cost or usage
- Inert gas requirements
- Solvent makeup requirement
- Other operating costs
9.0 Related Programs
No information available.
REFERENCES
1. Gas Processing Handbook, Hydrocarbon Processing, April 1975.
2. Buckingham, P.A., Fluor Solvent Process Plants: How They Are Working,
Hydrocarbon Processing and Petroleum Refiner, Vol. 43, No. 4, April 1964.
3. Maddox, R.N., Gas and Liquid Sweetening, Campbell Petroleum Series, 1974.
4. Kohn, A.L., and Buckingham, P.A., Fluor Solvent C02 Removal Process,
Petroleum Refiner, May 1960.
5. Cook, T.P., and Tennyson, R.N., Improved Economics in Synthesis Gas Plant,
Chemical Engineering Progress, Vol. 65, No. 11, November 1969.
6. Information provided to TRW by R. L. Schendel of Fluor Engineers and Con-
tractors, March 7, 1978.
B-47
-------
SULFIBAN (MEA) PROCESS*
1.0 General Information
1.1 Operating Principle - Absorption of H^S and organic sulfur compounds
from a gas stream using an aqueous solution of monoethanolamine (MEA)
as the sorbent.
1.2 Development Status - Commercially available.
1.3 Licensor/Developer - Applied Technology Corporation
4242 S.W. Freeway
Houston, Texas 77027
1.4 Commercial Applications - The Sulfiban process has been used to desul-
furize refinery gas, coke oven gas, synthesis gas, natural gas and
hydrogen. Approximately 100 units have been constructed treating a
combined gas volume total of 4.7 x 107 Nm3/d (1.73 x 109 SCFD) with
individual plant capacities ranging from 5300 Mm /d to 5.26 x 10
Nm3/d (0.2 to 200 MMSCFD)^.
2.0 Process Information
2.1 Flow Diagram (see Figure B-9)
Sour feed gas (Stream 1) enters the bottom of the absorber and the
lean amine solution* (Stream 5) enters the top. The purified prod-
uct gas (Stream 3) exits the top of the absorber and the rich
amine solution (Stream 4) which exits the bottom is pumped to a
stripping column where it is heated and the chemical reaction is
The Su fiban process is a patented variation of the basic ("standard") mono-
ethanolannne (MEA) system whicli incorporates the use of a variety of propri-
Lc7
-------
to
1
1
ABSORBER
V
^y
%
L
5
4
STRI
V_^
3PER
_J
i
(~
i
6 _
,
)_ f
\
\ ^ _
EGEND:
RECLAIMER
i
1. FEED GAS STREAM
2. ACID GAS TO TREATMENT
3. PURIFIED PRODUCT GAS
4. RICH AMINE SOLUTION
5. LEAN AMINE SOLUTION TO ABSORBER
6. LEAN AMINE SOLUTION TORE-BOILER
7. RECLAIMER FEED
8. STEAM
9. DEGREDATION PRODUCTS
Figure B-9. Sulfiban Process
-------
reversed.* Acid gases evolved (Stream 2) exit the top of the
stripper and are piped to acid gas treatment (for recovery of
sulfur, sulfuric acid, or pure H?S) and the lean amine (Stream
5) is returned to the absorber.
2.2 Equipment - Absorber column, stripping column reboiler, and reclaimer.
2.3 Feed Stream Requirements^
Temperature: 289°K to 311°K (60°F to 110°F)
t Pressure: Atmospheric
2.4 Operating Parameters^ '
2.4.1 Absorption Step:
0 Temperature: 311°K (100°F)
0 Pressure: 0.1 to 6.9 MPa (0 to 1015 psia)
0 MEA concentration: 12% to 20% by weight in water
0 MEA circulation rate: Dependent on gas flow and composition
2.4.2 Stripping Step:
0 Temperature: 366°K to 394°K (200°F to 250°F)
0 Pressure: 0.15 to 0.17 MPa (22 to 25 psia)
2.4.3 Redistillation Step:
0 Temperature: 394°K to 422°K (250°F to 300°F)
0 Pressure: 0.15 to 0.17 MPa (7 to 10 psia)
2.4.4 Reclaimer - High temperature is employed to minimize degradation
and maximize sorbent recovery.
*The primary reactions of MEA with H2S and C02 are as follows^3': H2S absorp-
tion solvent regeneration (RNH2 = monoethanolamine) reactions:
0 2 RNH2 + H2S = (RNH3)2S and
0 (RNH3)2S + H2S - 2 RNH3HS
C02 absorption solvent regeneration reactions:
0 2 RNH2 + H20 + C02 = (RNH3)2 C03 and
0 (RNH3)2C03 = H20 + C02 + 2 RNH3HC03 or
0 2 RNH2 + C02 = RNHCOONH3R
B-50
-------
2.5 Process Efficiency and Reliability - For an inlet H2S loading of
12.5 grams/Mm3 (500 gr/100 scf) to the Sulfiban unit, the overall
removal efficiency for the combined Claus-Sulfiban treatment system
can range from 87% to 98%(4'5). The basic "MEA System" and the
Sulfiban system have been in use for approximately 50 years and
25 years, respectively; they have proven to be reliable gas puri-
fication systems.
2.6 Raw Material Requirement^ '
Amine: 136 to 408 kg/day (300 to 900 Ibs/day)*
2.7 Utility Requirements
Cooling water: 67 to 200 £/sec (1060 to 3180 gpm)*
/
Power: 1300 to 4148 kwh/day*
Steam: 4900 to 14,800 kg/hr (10,913 to 32,746 lbs/hr)*
2.8 Miscellaneous - No information available which indicates special
maintenance problems or unusual hazardous condition created by the
process.
3.0 Process Advantages
Not pressure sensitive^ '
(?)
Low solvent costv '
(5)
Can remove COS and CS2 in addition to \\£
High carrying capacity for acid gases^
Does not remove hydrocarbons from feed gas^
(5)
t Compatible with HCN in feed gasv '.
4.0 Process Limitations
Feed gas temperature must be no greater than 316 K (110 F)
HCN in the feed gas causes some degradation of solvent.
Reclaimer sludge requires proper handling and disposal.
«eTupon 12.5 grams H2S per Nm3 (500 gr H2S/100 scf) at.^Jf*
efficient Sulfiban system; the plant sizes are 0.52 to 1.6 mill
(20 MMSCFD and 60 MMSCFD), respectively.
B-51
-------
5.0 Process Economics
(4)
In late 1974 it was estimated that a Sulfiban plant handling 1.6 mil-
lion Nm3/d (60 MMscfd) would cost between $1.00 to $3.00 per 1000 Mm3
(3.9U/MSCF to 7.60£/Mscf) for 90% and 98% removal efficiencies, respec-
tively (1975 dollars).*
6.0 Input Streams - see Table B-14.
0 Feed gas stream (Stream 2)
Steam (Stream 8)
t Solvent makeup: 255 kg/Mm3 (15 lbs/106 scf)
/
7.0 Discharge Streams - see Table B-14.
Purified product gas (Stream 3)
Acid gas (Stream 2)
t Sludge blowdown (Stream 7)
TABLE B-14. STREAM INFORMATION FOR A SULFIBAN PROCESS TREATING COKE OVEN GAS
WITH A PLANT THROUGH-PUT OF 1.6 x 10& Nm3/DAY (60 MMSCF) AND
SULFIBAN EFFICIENCY OF 98%(4)
Stream
Number
Stream Name
Stream Composition
3
7
8
Feed Gas
Acid Gas
Product Gas
Sludge/Degradation Products
Steam
12.5 gm H2S/Nm3 (500 gr H2S/100 scf)
1.5 gm HCN/Nm3 (60 gr HCN/100 scf)
50% H2S
4% HCN
46% C02
250 mgm H2S/Nm3 (10 gr/100 scf)
140 I/day (37 gal/day)
327,000 kg/day of 1.12 MPa steam
(720,000 #/day of 150 psig steam)
indication as to whether this cost includes royalty fees.
B-52
-------
8.0 Data Gaps and Limitations
Limited data are available on the maximum allowable concentrating «f
various contaminants in the feed gas (e.g., COS, CS2' Erace melals
carbonaceous matter NH3, COS, CS2, HCN) have on thl process anJ the
ultimate fate of such contaminants in the system.
* "e °" the """tWP'rtl" °f «cl.1«r sludge/
9.0 Related Programs
No data available.
REFERENCES
1. Dravo Corporation, Handbook of Gasifiers and Gas Treatment Systems, ERDA
FE-1772-11, Washington, D.C., February 1976.
2. Riesenfeld, F.C. and Kohl, A.C., Gas Purification, Second Edition, Gulf
Publishing Co., Houston, Texas, 1974.
3. Maddox, R.N., Gas and Liquid Sweetening, Campbell Petroleum Series, 1974.
4. Massey, M.J. and Dunlap, R.W., Economics and Alternatives for Sulfur
Removal from Coke Oven Gas, Journal of the Air Pollution Control Associ-
ation, Vol. 25, No. 10, October 1975.
5. Information provided to TRW by W. M. Peters of Applied Technology
Corporation, February 27, 1978.
B-53
-------
MDEA PROCESS
1.0 General Information
1.1 Operating Principle - Chemical absorption of acid gases (H2S, organic
sulfides and C02) from industrial gases using methyldiethanolamine
(MDEA) as the sorbent. The process is highly selective toward H2$*
(in the presence of C02)> and can also be modified to remove substan-
tial quantities of COS'*'.
1.2 Development Status - Commercially available.
1.3 Licensor/Developer - Dow Chemical Company
Freeport, Texas 77541
1.4 Commercial Applications - Three or four U.S. refineries have con-
verted their arotne systems (MEA/DEA) to MDEA and one commercial
plant is under construction^ . Location of facilities and opera-
ting parameters not reported.
2.0 Process Information
2.1 Flow Diagram (see Figure B-10) - The raw gas (Stream 1) enters the
bottom of the absorber while a lean solution of MDEA (Stream 3) is
delivered to the top. The lean solution absorbs most of the H2S and
some C02 as it passes counter-current to the gas. The treated gas
(Stream 2) exits at the top of the absorber. The rich solution
(Stream 4) flows to the stripper for regeneration. The acid gases
(Stream 9) go to sulfur recovery. A sidestream of the lean solution
(Stream 10) is sent to a distillation unit for purification. The
remainder of the lean solution is returned to the absorber.
*The degree of selectivity is affected by the partial pressure of C02 in the
gas stream, solvent loading, and the temperature in the absorption unit (see
Section 4.0).
B-54
-------
12
11
CONTACTOR
11
STRIPPER
f CONDENSATE *\
I ACCUMULATOR J
13
10
LEGEND:
1. RAW GAS
2. TREATED GAS
3. LEAN SOLUTION
4. RICH SOLUTION
S. MDEA Makeup
6. AMINE SLUDGE
7. STEAM IN
8. STEAM CONDENSATE
9. ACID GASES
10. LEAN SOLUTION SIDESTREAM
11. COOLING WATER IN
12. COOLING WATER OUT
13. REFLUX WATER
-^Jk
REDISTILLATION
UNIT
Figure B-10. MDEA Acid Gas Removal Process
B-55
(2)
-------
2.2 Equipment - Conventional contactor, stripping column, and stripper
reboiler .
2.3 Feed Stream Requirements^
Temperature: 288°K to 323°K (60°F to 120°F)
t Pressure: Atmospheric or greater
2.4 Operating Parameters
2.4.1 Absorber
Temperature: 300°K to 325°K (80°F to 125° f^1'
Pressure: 0.1 to 0.7 MPa (0 to 1000 psigr1'
0 Solvent loading: Usually less than 0.03 Nm acid gas/£ (3.2 scf
acid gas/gal)(2). Some commercial applications
have involved'loading of 30% MDEA to 0.7 Nm3 of
acid gas/£ (7.9 scf/gal)(4).
2.4.2 Stripper^
Temperature: 388°K to 393°K (240°F to 250°F)
0 Pressure: 0.14 to 0.17 MPa (7 to 10 psig)
2.4.3 Redistillation Unit^4'
0 Temperature: 393°K to 423°K (250°F to 350°F)
Pressure: 0.14 to 0.17 MPa (up to 10 psig)
2.5 Process Efficincy - hLS in treated gas can be reduced to less than
4 ppm. The exact amount of H9S, C09 and other acid gases removed
( 1 ^}
depends upon actual operating conditions*v .
2.6 Raw Material Requirements
3 I \ ]
MDEA solvent makeup: 9 g/1000 Nm of sour gas (0.5 lb/mSCF)v '.
Foam inhibitors: Dow Corning DB-313 Exxon-Corexit 350.
Utility Requirements^ ' - Basis
(105°F) and 0.41 MPa-(60 psia).
2.7 Utility Requirements'1^ - Basis: 28,3000 Nm3 (MM scf) gas at 314°K
*In applications where selective removal of H2S from a sour gas stream is a
major objective the following can be obtained: Streams of 92% to 95% pure
can be produced while producing sweet gas with from 5 to 30 ppm H£S and retain
ing 60% to 70% of the C02. Best results are achieved when the absorber is
operating at high solvent loading and short contact times(2)
B-56
-------
2.7.1 Case 1: Gas feed composition of 0.6 volume % \\£ and 10 volume %
C02 with 50 ppmv H£S and 3.3 volume % C02 in treated gas.
Steam: 4900 kg (10,700 lb)
Cooling water: ?
Electricity: 15 kWh
2.7.2 Case 2: Gas feed composition same as above, but 965 ppmv H S and
7.3 volume % C02 in treated gas.
Steam: 2270 kg (5000 lb)
Cooling water: ?
Electricity: 8 kWh
2.8 Miscellaneous - No information available which indicates special
maintenance problems or unusual hazardous conditions created by the
process.
3.0 Process Advantages
(1 3)
Removes most sulfur compounds without solvent degradation^ ' '.
Can be operated over a wide range of pressure^ ' .
(2)
More selective than most amines in the preferential removal of \\^ ' .
Lower energy requirements than most chemical type processors ^ .
In the stripping operation reflux rates as low as 0.7 kg H20/kg acid
gas (0.7 Ib H20/lb acid gas) can be used without noticeably increasing
the H?S in the sweet gas(2).
4.0 Process Limitations^ '
Process selectivity for H2$ is maintained in contactor operating from
0.45 to 1.7 MPa (50 to 250 psi) and having 2% to 24% C02>
Selectivity decreases markedly after a critical solvent loading level
is surpassed. Pilot plant studies indicate the critical loading to be
approximately 0.03 Nm3 acid gas/£ (3.2 scf acid gas/gal).
Selectivity is sensitive to temperature of lean amine entering the
absorber. Pilot plant studies indicate absorber temperature should
be less than 316°K (HOOF).
B-57
-------
Potential corrosion problems exist at high MDEA concentrations (>4535)
and high loadings (>0.5 moles acid gas/mole MDEA).
t Some potential foaming problems can occur requiring the use of foam
inhibitors.
5.0 Process Economics
No data available.
6.0 Input Streams (see Figure B-10)
6.1 Raw Gas (Stream 2) - see Table B-15.
6.2 MDEA Makeup (Stream 6)(4): - 8.5 kg MDEA/Nm3 (0.51 Ibs MDEA/106 scf)
7.0 Discharge Streams
7.1 Treated Gas (Stream 2) - see Table B-15.
7.2 Acid Gas (Stream 9) - The acid gas produced during regeneration is
composed primarily of CO,,, H2S, organic sulfur and some hydrocarbons.
This stream will require further processing in a sulfur recovery
unit, such as a Claus or Stretford.
TABLE B-15. OPERATING DATA AND STREAM CHARACTERISTICS FOR THE APPLICATION
OF A MDEA PROCESS IN A REFINERY SETTING(2)
Absorber Size
Absorber Parameters
Feed Stream
Composition - Stream 1
Product Gas Acid
Components - Stream 2
0
0
1.83 m (6 ft) in diameter
20 type A Koch Flexi trays at 0.61 m
(24 in. ) spacing
Solvent temperature: 316°K (110°F)
Pressure: 0.63 MPa (78 psig)
MDEA concentration: 15 wt %
Solvent loading: approximately
0.03 Nm3 acid gas/£ (3.0 scf acid
gas/gal)
- 8.5%
- 1.4%
H2
CH4
C2H6
- 38%
- 21.5%
- 12%
C02 - approximately 64% of original C02
H2S - 7 ppm
B-58
-------
7.3 Amine Sludges (Stream 6) - Stream composed mainly of solvent with
traces of degradation products and other components scrubbed from
gas stream. Quantitative rate and actual composition of this stream
not known.
8.0 Data Gaps and Limitations
Process applicability to coal conversion process gas purification sys-
tems not established.
No data on composition of by-product acid gas.
t No data available on maximum allowable concentrations of various con-
taminants in the feed stream (COS, CS2, NH3, HCN, mercaptans, etc.).
t The effects that various contaminants (NH3, carbonaceous matter, trace
metals, etc.) have on the process, and the ultimate fate of such con-
taminants in the system.
Characterization of gaseous and liquid streams for natural gas and/or
refinery applications (temperature, pressure, composition, etc.).
9.0 Related Programs
Pilot plant tests have been conducted with the MDEA process for
sweetening simulated gases which are typical of those produced by coal
(4)
gasification from a Koppers-Totzek gasifierv .
REFERENCES
1. Dravo Corporation, Handbook of Gasifiers and Gas Treatment Systems, Final
Report. Task Assignment No. 4 Report FE-1772-11, ERDA Contract No.
E(49-18)-1772, Pittsburgh, PA, Chemical Plants Division, February 1976.
2. Vidaurri, F.C. and Kahre, L.C., Recover H2S Selectively from Sour Gas
Streams, Hydrocarbon Processing, November 1977.
3. Frazier, H.D. and Kohl, A.L., Selective Absorption of Hydrogen Sulfide from
Gas Streams, Industrial and Engineering Chemistry, November 1950.
4. Information provided to TRN by R. L. Pearce of Dow Chemical Company,
March 3, 1978.
B-59
-------
SNPA - DBA PROCESS
1.0 General Information
1.1 Operating Principle - The chemical absorption of acid impurities
(H2S, C02, COS, CS2, etc.) of a gas stream using diethanolamine (DEA)
as the sorbent.
1.2 Development Status - Commercially available.
1.3 Licensor/Developer
Developer - Societe Nationale Elf d'Aquifaine (Production)
U.S. Licensor - The Ralph M. Parsons Co.
100 W. Walnut Street
Pasadena, California 91124
1.4 Commercial Applications - The SNPA -DEA process has been used exclu-
sively in sour natural gas treatment to remove H9S and C09; the total
3
world capacity of the process is approximately 125 million Mm
(4.2 billion scf) per day^ '. There is no present commercial applica-
tion to acid gases generated in coal gasification.
2.0 Process Information
2.1 Flow Diagram (see Figure B-ll) - Raw gas (Stream 1) is contacted
counter-currently with lean DEA solution in a contactor where the
acid gases are removed. Purified product gas (Stream 2) exits the
top of the contactor. Rich solvent (Stream 4) exits the bottom of
the contactor and is flashed to remove dissolved and entrained hydro-
carbons. From the flash tank, the rich DEA is heated and then
charged to the stripper. In the stripper the acid gases are stripped
from the DEA solution, then cooled and the acid gas is evolved
(Stream 8). Lean solution exits the bottom of the stripper and is
B-60
-------
I FILTER AID
CO
I
RAW GAS
THE ATEO GAS
LEAN SOLUTION
RICH SOLUTION
RFF LUX WATER
STEAM IN
STEAM CONDEMSATE
ACID GASES
FUFL GAS
COOLING WATER IN
COOLING WATER OUT
FILTER RESIDUE
Figure B-ll. SNPA-DEA Acid Gas Removal Process
-------
cooled before returning to the contactor (Stream 3). Solution stor-
age and conditioning* are provided on the lean DEA stream.
2.2 Equipment - Conventional contactor, flash vessel, stripping column
and stripper reboiler.
2.3 Feed Stream Requirements*1 '
Temperature: 300°K to 330°K (77°F to 130°F)
Pressure:* 2 to 15 MPa (300 to 2250 psia)
2.4 Operating Parameters^ '
2.4.1 Absorption Step
Temperature: 523°K to 343°K (480°F to 160°F)
t Pressure: 2 to 15 MPa (300 to 2250 psia)
Solvent Loading: ?
2.4.2 Regeneration Step
Temperature: 383°K to 393°K (230°F to 250°F)
Pressure: 0.14 to 0.2 MPa (21 to 30 psia)
2.4.3 Flash Vessel
Temperature: 523°K to 533°K (160°F to 176°F)
Pressure: 0.3 to 0.5 MPa (44 to 74 psia)
2.5 Process Efficiency and Reliability - Treated gas will meet the con-
ventional pipeline specification of 0.25 grain H0S per 100 scf maxi-
(-}} t-
mum and C02 of 2 volume % or lessv '.
2.6 Raw Material Requirements
DEA Makeup: Approximately 30 to 140 kg/1000 Nm3 (2 to 9 lb/
MMscf)(l).
Conditioning consists of solvent filtration for the removal of trapped impur-
ities (liquid hydrocarbon, pipe scale, glycol, corrosion products and well
head additives)(1). Activated carbon may also be used in conjunction with
filtration to improve solvent rejuvenation.
Partial pressure of the acid gas in the feed stream must be at least
0.4 MPa (59 psia)(2).
B-62
-------
Filter aids (activated carbon, etc.): ?
Blanketing Gas^ ' - Nitrogen or oxygen free gas.
Foam Inhibiter: Typical silicone fluid.
2.7 Utility Requirements^ '
Electricity: Depends on design (e.g., availability of high pres-
sure steam for driving pumps).
t Water: Depends on operating conditions (e.g., feed tempera-
ture and moisture content)
t Steam: 0.08 to 0.18 kg/£ (0.7 to 1.5 Ibs/gal) of circulated
solution for reboiling.
3.0 Process Advantages
Greatest advantage is in treating raw gases having high acid gas con-
centrations, at high pressures.
0 Utility requirements for a DEA unit generally run substantially below
those required for MEA units.
COS is removed without DEA degradation.
4.0 Process Limitations^ '
Process is not effective at low pressures.
Conditioning of lean solvent is required.
t Gas blanketing of the pure DEA solution during storage may be necessary
to prevent oxygen degradation of the solution.
Royalty fees are required.
Potential foaming problems exist which requires the use of foam
inhibitors and an anti-foam-injection facility.
Potential corrosion problems exist in the rich solution piping down-
stream of large pressure reduction stations. Expected life of pipe in
this area has been reported to be approximately 3 to 4 years UK
Corrosion problems may occur in the lean solution heat exchanger which^
will require the use of stainless steel rather than carbon steel tubes .
5.0 Process Economics
No data available.
B-63
-------
6.0 Input Streams (see Figure B-ll)
6.1 Raw Gas (Stream 1) - see Table B-16.
6.2 Filter Aids (activated carbon, etc.) - Dosages/quantities not known.
6.3 Foam Inhibitor - (Not shown in Figure B-ll.)
7.0 Discharge Streams (see Figure B-ll)
7.1 Treated Gas (Stream 2) - see Table B-16.
7.2 Acid Gas (Stream 8) - Acid gases removed from raw gas are at adequate
pressures and proper temperature to serve as a direct feed for a
Claus-type sulfur recovery unit. No data available on the composi-
(2)
tion of this streairr '.
7.3 Fuel Gas (Stream 9) - Usually low grade fuel, no data available on
the composition and generation rate for this stream.
7.4 Filter Residue - Stream usually contains various small particles
(e.g., iron sulfide, liquid hydrocarbons, reaction products and
TABLE B-16. OPERATING CONDITIONS FOR A SNPA-DEA PROCESS
Feed Gas Flow Rate
Feed-Gas Analysis
H2S %
co2 %
COS ppm
CS2 ppm
CH4
Outlet Gas Analysis
H2S gr/100 scf
C02 gr/100 scf
COS gr/100 scf
11.5 NnT/sec
15.0
10.0
300
600
Balance
0.28
1.6
0
B-64
-------
degradation products). Jsk> data available on the composition and
generation rate for this stream.*
8.0 Data Gaps and Limitations
Data gaps exist in the following areas:
Process applicability to coal conversion process gas purification
systems not studied/established.
No data available on maximum allowable concentrations of various
contaminants in the feed stream (COS, CS?, NH,, HCN, mercaptans,
etc.). * J
No data available on process economics.
No data available on the composition, amount and disposition of
the residues collected by filters.
9.0 Related Programs
No data available.
REFERENCES
1. Daily, L.W., Status of SNPA-DEA, The Oil and Gas Journal, May 4, 1970.
2. Goar, B.G., Today's Gas-Treating Processes-1, The Oil and Gas Journal,
July 12, 1971.
3. Gas Processing Handbook SNPA-DEA, Hydrocarbon Processing, 54(4):95, 1975.
4. Wendt, C.J. Jr., and L.W. Dailey, Gas Treating: The SNPA Process, Hydro-
carbon Processing, 46(10):155-7, 1967.
5. Riesenfeld, F.E. and A.L. Kohl, Gas Purification, Second Edition, Gulf
Publishing Company, Houston, Texas, 1974, page 30-1, 78.
6. Information provided to TRW by C. L. Black of Ralph M. Parsons Company,
June 20, 1978.
^HidUc-tion of 1 PPm by weight of materials continuously entering the contactor
of a typical 100 MMscfd plant will represent approximately 1 ton or
contaminants(1).
B-65
-------
ADIP PROCESS
1.0 General Information
1.1 Operating Principle - The chemical absorption of acidic gases (H2S,
C02, COS, etc.) from a gas stream using a relatively concentrated
(20 to 35%) aqueous solution of diisopropanolamine (DIPA) as the
absorbent.
1.2 Development Status - Commercially available.
1.3 Licensor/Developer - Shell Development Company
One Shell Plaza
P.O. Box 2463
Houston, Texas 77001
(2 3)
1.4 Commercial Applications^' '
Removal of H^S from industrial gases.
Commercial applications include natural gas, synthesis gas and
refinery fuel gas.
More than 130 units for hLS removal are in operation or under
construction.
Selective adsorption of H2S to upgrade sulfur plant feed streams.
2.0 Process Information^2'3^
2.1 Flow Diagram (see Figure B-12) - The raw gas (Stream 1) enters the
bottm of the contactor, and the regenerated DIPA enters the top.
Raw gas passes counter-flow to the DIPA and the purified gas (Stream
2) exits the top of the contactor. The rich solution (Stream 4)
from the contactor bottom flows to the stripper. Stripping occurs
with the steam generated in the reboiler. Steam is condensed and
separated from the acid gases and refluxed to the regenerator, while
the acid gases (Stream 7) go to sulfur recovery. The regenerated
solvent is returned to the contactor.
B-66
-------
CONDENSER
STRIPPER
CONTACTOR
DO
CTi
-J
CONDENSATE
ACCUMULATOR
10
LEGEND
1. RAW GAS
2. TREATED GAS I
3. LEAN SOLUTION
4. RICH SOLUTION
5. STEAM IN
6. STEAM CONDENSATE
7. ACID GASES TO SULFUR RECOVERY
8. COOLING WATER IN
9. COOLING WATER OUT
10. REFLUX WATER
11. SOLVENT MAKE-UP
STRIPPER
REBOILER
Figure B-12. ADIP Acid Gas Removal Process
(3)
-------
2.2 Equipment - Contactor, stripping column, and stripper reboiler.
Because of low solvent corrosivity, carbon steel can be used
throughout the plant.
(4)
2.3 Feed Stream Requirements
Temperature*: 283°K to 322°K (50°F to 120°F)
Pressure"1": 0 to 5.2 MPa (0 to 915 psia)
Other Impurities: ? Water washing or treatment of the feed gas to
remove impurities such as HCN may be necessary to
minimize DIPA degradation. Presence of a signifi-
cant concentration of 02 in the feed gas results
in degradation of DIPA, requiring use of a
reclaimer to recover the degraded DIPA.I
(2 4}
2.4 Operating Parametersx ' '
2.4.1 Contactor
0 Temperature: 310°K to 322°K (100°F to 120°F) (see also
Table B-17).
Pressure: 1 to 5.2 MPa (15 to 915 psia) (see also
Table B-17).
Solvent Loading: Up to 0.6 mole H2S/mole amine in absorber.
2.4.2 Stripper
Temperature: 394°K to 408°K (250°F to 247°F)
Pressure: Usually near atmospheric
*Like all amine solvent processes, the ADIP process is not suitable for high
temperature operation because of significant reduction in absorption effi-
ciency and increased solvent losses. (Note: amine solvents are regenerated
thermally.)
fln general, the absorption efficiency with chemical absorption systems such as
ADIP is independent of pressure.
^No existing ADIP system has required a reclaimer.
B-68
-------
TABLE B-17. OPERATING DATA FOR ADIP PLANTS
(4)
Feedstock
Feed Gas Flow MM Mm3/ day (MMSCFD)
H2S Vol. %
C0£ Vol. %
Molecular Wt
Absorption
Pressure MPa (psia)
Absorption
Temperature °K (°F)
Treated Gas
C02 Vol. %
H2S ppra (vol)
Low Pressure
Steam Consumption
MKg/hr (M Ib/hr)
Cracked Gas
Gas From a
Catalytic Cracker
0-.27 (10)
10.4
2.5
28
1.9 (285)
308 (95)
0.2
<10
4.2 (9.2)
Residual Gas From
Hydrodesulfurization
2.7 (100)
15.6
25
0.98 (70)
314 (105)
100
260 (134)
Gas From ,
Hydrogen
Purification
2.8 (104)
3.0
91.2
42
0.1 (15)
283 (50)
98.1
500
47 (104)
Gas From FCCU*
5.9 (20)
2.4
1.9
24
+0.3 (195)
311 (100)
1.0
160
3.3 (7.3)
Natural Gas
0.02 (3.. 7)
7.1
21.3
28
2.7 (40)
316 (10)
17.8
130
2.6 (5.7)
U3
I
10
Fluidized Catalytic Cracking Unit
-------
2.5 Process Efficiency and Reliability - The following product specifi-
cations are readily attained in actual applications (see also
Table B-17 for data on three actual applications):
H2S content in fuel gas: less than 100 ppmv
H2S content in natural gas: less than 5 ppmv
Other acid gases as well as COS are removed depending on their con-
centrations and the operating conditions.* Selective amounts and
rates of absorption of H?S and C02 can be achieved by proper selec-
tion of DIPA concentration, flow rates, etc.). Commercial plants
have demonstrated high stream factors.
No reliability data available on the operation of commercial plants.
2.6 Raw Material Requirements - DIPA solvent makeup (including mechanical
losses): 2-5 x 10 kg/kg of H^S removed.
o
2.7 Utility Requirements - Typical requirements per 28,300 Mm (1 MMSCF)
of a gas containing 10 vol.* H2S and 2-5 vol. % C02 at 1.96 MPa
(285 psia) with 2 ppm H2$ and 0.2 vol.% C02 in purified gas are:
Steam: 997 kg (2200 Ib)
Electric power: 85 kwh
Cooling water: not applicable
3.0 Process Advantages
DIPA solvent is noncorrosive
Solvent is not degraded by COS
Low steam consumption compared to other processes (e.g., MEA and DEA)
0 Process can be operated over a wide range of pressures^
The DIPA solvent in the ADIP process is not degraded by COS and the process
can be used for regenerative COS removal. In gas treatment application, how-
ever, a very long residence time would be required for COS removal. In proc-
/^i^»f*^\**i-'ii^%t*«N^.4.L.j»f*..TX1-! T___ /.i .. _ \
- - ..«. w,_ i ^.^ u i i \_vj I V I \^\J*J I dllV/ VU I . All f-M '
euses *s the Sulfino1 Process (the subject of a separate data sheet)
where DIPA is dissolved in Sulfolane (instead of water as in ADIP), a more
effective contact between DIPA and COS is possible because of th
solubility of COS in Sulfolane.
Higher pressure operation has favorable effect on reaction rate.
B-70
-------
4.o Process Limitations
Presence of HCN and significant concentration of 0? in the feed aas
may require feed gas pretreatment to minimize solvent degradation.
$.0 Process Economics
No overall cost data available. DIPA cost is $1.0/kg ($0.46/lb) in
1977 dollars.
6.0 Input Streams (see Figure B-12)
6.1 Raw Gas (Stream 1) - (see Table B-17)
6.2 Makeup Solvent (Stream 12) - Quantity not known.
7.0 Discharge Streams
7.1 Treated Gas (Stream 2) - (see Table B-17)
7.2 Acid Gas (Stream 7) - Acid gas stream produced during solvent regen-
eration is composed primarily of H2S, C02- No actual data available.
7.3 Solvent Slowdown (Stream 11) - Containing sludges, degraded solvent,
trapped impurities, etc. Rate and composition not known.
8.0 Data Gaps and Limitations
Data gaps exist in the following areas:
No data available on maximum allowable concentrations of various
contaminants in the feed stream (NH3, HCN, mercaptans, etc.).
t No information on process costs including solvent makeup
requirements.
No data on solvent losses, composition of by-product acid gas and
"post treatment" requirements to minimize solvent losses.
9.0 Related Programs
The ADIP process is featured in the design of the Wesco Lurgi SNG
facility for removal of hydrocarbons from the acid gas stream from the
Rectisol unit^5'.
B-71
-------
REFERENCES
1. Riesenfeld, F.C. and Kohl, A.L., Gas Purification, 2nd Edition, Gulf Pub-
lishing Company, Houston, Texas, 1974-
2. Dravo Corporation, Handbook of Gasifiers and Gas Treatment Systems, ERDA
FE-1772-11, Washington, D.C., February 1976, pp 101-103.
3. Gas Processing Handbook, ADIP, Hydrocarbon Processing, 54 (4) : 84, 1975.
4. Information provided to TRW by J. M. Duncan of Shell Development Company,
December 8, 1977.
5. Control of Emissions from Lurgi Coal Gasification Plants, EPA Office of
Air Quality Planning and Standards, Research Triangle Park, North Carolina.
EPA-450/2-78-012, OAQPS 1.2-093, March 1978.
B-72
-------
FLUOR ECONAMINE (DGA) PROCESS
1.0 General Information
1.1 Operating Principles - Chemical absorption of acid impurities (C0?,
H2S, organic sulfur) from gas streams using a relatively concentrated
aqueous solution of diglycolamine (DGA) as the sorbent.
1.2 Development Status - Commercially available.
1.3 Licensor/Developer - Fluor Engineers and Constructors, Inc.
3333 Michaelson Drive
Irvine, California 92730
1.4 Commercial Applications (see Table B-18) - The Fluor Econamine Process
has been in commercial use for removing h^S and C0? from natural, syn-
thesis or refinery gas streams since 1965(1). There is no present com-
mercial application to gases produced from coal gasification.
2.0 Process Information
2.1 Flow Diagram (see Figure B-13) - Raw gas (Stream 1) is purified in a
contactor vessel where acidic impurities are absorbed by the DGA.
Purified product gas (Stream 2) exits at the top. Rich solution
(Stream 4) leaves the bottom of the contactor and is flashed to
remove absorbed hydrocarbons. The rich solution is heated by inter-
change with hot lean solution before being introduced into the stripper
for solution regeneration. Acid gases and water vapor pass overhead
to the condenser. Condensed water is refluxed to the stripper and the
acid gases (Stream 8) are sent to flare or sulfur recovery. Regener-
ated solution leaves the bottom of the stripper and is then cooled
before being introduced in the absorber top. A side stream (Stream
9) of lean solution is taken to the reclaimer for purification.
B-73
-------
TABLE B-18. COMMERCIAL PLANTS USING FLUOR'S ECONAMINE PROCESS
Company
El Paso Natural Gas Co.
El Paso Natural Gas Co.
El Paso Natural Gas Co.
Getty Oil Company
El Paso Natural Gas Co.
El Paso Natural Gas Co.
Texaco Inc.
Getty Oil Co.
Mountain Fuel Supply
Ark-La Gas Company
Texaco Inc.
Kansas-Neb Natural Gas
Marathon Oil Company
Hunt Oil Company
Texaco Inc.
Natural Gas Pipeline of
America
Montana Dakota Utilities
Rocky Mountain Natural Gas
Corporacion Venezuela de
Petroleos
El Paso Natural Gas
Claiborne Gasoline Co.
ARAMCO
Canadian Superior Oil Co.
ARAMCO
Confidential Client
Nat. Iranian Oil Co.
ARAMCO
Nat. Iranian Oil Co.
ARAMCO*
ARAMCO*
Confidential Client*
Plant i
Monument
Wasson 'A1
Waha 'A1
New Hope
Waha 'B'
Wasson 'B'
Knight 'A1
Teas
Church Buttes
Gilmer
Knight 'B1
Pawnee Rock
Yates
Nordheim
Henry
Haley
Riverton
Slick Rock
Ful lerton
Antioch
Plant 415
Harmattan Stg 7
Plant 464
Tehran Ref.
Berri - Pit 474
Esfahan Ref.
Shedgum
Uthmahiyah
Location
Monument, New Mexico
Denver City, Texas
Coyanosa, Texas
Scroggins, Texas
Coyanosa, Texas
Denver City, Texas
Pecos, Texas
Streatman, Texas
Church Buttes, Wyoming
Gilmer, Texas
Pecos, Texas
Pawnee Rock, Kansas
Iraan, Texas
Nordheim, Texas
Erath, Louisiana
Kermit, Texas
Riverton, Wyoming
Slick Rock, Colorado
Lake Maracaibo, Venezuela
Fullerton, Texas
Lisbon, Louisiana
Ras Tanura, Saudi Arabia
Udhaliyah, Saudi Arabia
Victoria, Australia
Tehran, Iran
Jubail , Saudi Arabia
Esfahan, Iran
Saudi Arabia
Saudi Arabia
Middle East
Comments
Conversion of MEA-DEG
Conversion of MEA-DEG
New
Conversion of MEA
New
New
New
New
New
Conversion of MEA
New
Conversion of MEA
New
New
New
New
New
New
New - Ethane Product
Conversion of MEA-Dehydration
New
New - 2 Plants
New
New - 2 Plants
New - COS Removal
New
New - 3 Plants
New
New - 4 Plants - Improved DGA
New - 3 Plants - Improved DGA
New - Improved DGA
-Inlet
Capacity
HHSCFD
180
105
285
44.5/5.5
285
45
30
12
3
15
30
20
10
10
23.5
30
20
25
14.2
35
10
70 ea.
150
38 ea.
40
5.5/6.3
220
6.1/8.0
547 ea.
547 ea.
141
Inlet
Pressure
psig
625
165
875
1,100/400
875
1,100
165
800
800
1,100
1,100
190
125
850
445
1,100
960
800
400
550
1,100
15
930
230
1,050
320/57
125
330/63
160
16C
220
H9S MOL.
i %
0.75
0.32
0.04
8.48/16.6
0.04
0.6
0.32
31.7
6.4
0.31
0.6
0.2
5.2
0.21
nil
0.48
4.95
0.2
nil
1.25
0.8
12.28
0.01
1.7
Trace
8.7/26.7
9.02
5.5/24.9
2.2
2.2
1.88
CO, MOL.
'%
2.75
5.9
3.9
4.01/3.4
3.9
6.5
5.9
3.8
23.4
2.49
6.5
1.25
6.3
7.38
7.6
1.5
0.35
9.0
11.41
3.80
2.0
2.09
4.45
10.2
0.64
0/0
7.72
0/0
9.6
9.6
4.34
CO
I
*Under Construction
-------
COOLER
LEGEND:
t. RAW GAS
2. TREATED GAS
3. LEAN SOLUTION
4. RICH SOLUTION
5. FUEL GAS
6. STEAM IN
7. STEAM CONDENSATE
B. ACID GASES (TO FLARE
OR SULFUR RECOVERYI
9. LEAN GAS TO RECLAIMER
10. REFLUX WATER
Figure B-13. Fluor Econamine Acid Gas Removal Process
(3)
B-75
-------
2.2 Equipment - Conventional contactor, stripper column, flash tank,
stripper reboiler and a reclaimer.
2.3 Feed Stream Requirements*
Temperature1": Usually less than 344°K (160°F)
Pressure^: No limit
Loading: Depends on acid gas partial pressure and temperature
of rich amine.
(3 5)
2.4 Operating Parametersv ' '
2.4.1 Contactor
Temperature1": 305°K to 378°K (90°F to 220°F)
Pressure^: No limit
q
Solution loading: 0.035 to 0.14 Nm of acid gas per liter (5 to
2.0 scf/gal)
2.4.2 Stripper^
Temperature: 377°K to 465°K (220°F to 280°F)
Pressure: Near atmospheric
2.4.3 Solvent Reclaimer
Temperature: 433°K to 465°K (320°F to 380°F), kettle
temperature
Pressure: Same as stripper
Reclamation rate: 1 to 2% of circulation solution
2.5 Process Efficiency and Reliability - Can reduce H0S and C00 in gas
Ml
to less than 4 ppmv and 0.01 volume %, respectively^ '. Satisfactory
*Fluor Econamine (DGA) is considered suitable for all applications where
aqueous monoethanolamine (MEA) process can be used. The process shows signi-
ficant savings over MEA in applications where the inlet gas contains more
than 1.5 to 2% acid gasO).
Like all amine solvent processes, the Fluor Econamine process is not suitable
for high temperature operation because of reduction in absorption efficiency
and increased solvent losses. (Note: amine solvents are regenerated
thermally.)
In general, in chemical absorption systems, the absorption efficiency is
independent of absorption pressure.
B-76
-------
long-term operating experience at commercial installations confirm
the acceptability of the process from the standpoint of system cor-
rosion, efficiency and maintenance requirements.
2.6 Raw Material Requirements^ - DGA solvent makeup: Approximately 35
to 100 kg/Mm3 (2 to 6 Ibs/MMSCF) depending on gas throughput and
temperature.
2.7 Utility Requirements^ - Based on feed stream rate of 2.7 MM Nm3/
day (100 MMSCFD) at 305°K (90°F) and 5.9 MPa (850 psig) and a feed
stream composition of 90% CH4, 5% H2S and 5% Ok:
Steam: 32.2 (35 ton) per hour at 0.48 MPa (70 psia)
Electricity: ?
Cooling water: ?
3.0 Process Advantages
Low solvent volatility losses relative to other amine systems.
(21
Substantial removal of mercaptans and organic disulfidesv '.
(2\
Low absorption of heavy hydrocarbons^ '.
The relatively low vapor pressure of DGA permits its use in relatively
high concentrations, typically 40% to 70% (compared to 15% to 20% for
MEA). This results in higher acid gas pick-up capacity, lower solution
circulation rate, lower stripper reboiler steam consumption and signifi-
cant reductions in capital investments and operating costs(2>3>4»5).
Savings of as much as 15% to 25% in both capital and operating costs
have been claimed for Fluor Econamine process, compared to conventional
MEA.
DGA is suitable for process streams containing COS and CS2 since (unlike
MEA) reaction products are thermally regenerated.
4.0 Process Limitations^
Relatively high cost of DGA.
Somewhat high corrosiveness toward carbon steel, although probably less
than MEA.
Higher solution losses (compared to MEA) due to higher concentration in
circulating solution.
Requirement for vacuum distillation for solvent purification.
B-77
-------
5.0 Process Economics
No actual cost data available for the process. Savings of as much
as 15% to 25% in both capital investment and operating costs have been
(2)
claimed for the process when compared to the conventional MEAX ;.
6.0 Input Streams (see Figure B-13)
6.1 Raw Gas (Stream 1) - Typical case from Reference 4. 3.2 MM Mm /day
(121.3 MMSCFD, 5.9 MPa (865 psia)) natural gas containing 2% to 5%
(41
total acid gas^ '.
6.2 Solution Makeup: Quantity unknown.
7.0 Discharge Streams (see Figure B-13)
7.1 Treated Gas (Stream 2)^
C02, 0.01 mole %
H2S, 0.006 g/Nm3 (0.25 grains/100 scf)
7.2 Acid Gases (Stream 8) - No data available on actual rates and com-
position. For a plant handling 2.68 MM Nm3/day (100 MMSCFD of nat-
ural gas, the acid gas volumetric rate would be close to 0.268 MM Nm /
day (10 MMSCFD). The gas contains primarily CO,, H?S, and probably
£- £-
traces of organic sulfur compounds, hydrocarbons and solvent.
7.3 Lean Solution Side Stream Sent to Reclaimer (Stream 9) - Volumetric
rate 1% to 2% of circulating solution; no data available on
composition.
7.4 Fuel Gas Generated in Flash Tank (Stream 5) - No data available on
quantity and composition. (The gas contains mainly volatile
hydrocarbons.)
7.5 Solvent Slowdown (from reclaimer and/or stripper) - Intermittent
only.
8.0 Data Gaps and Limitations
Data gaps exist in the following areas;
Process applicability to coal conversion process gas purification
systems not studied/established.
B-78
-------
Limited data available on maximum allowable concentrations of
various contaminants in the feed stream (COS, CS0, NH,, HCN
mercaptans, etc.). ^ 3
No information on process costs including solvent makeup
requirements.
No data on solvent losses, composition of by-product acid gas
and "post treatment" requirements (if any) to minimize solvent
losses.
No data on rates and composition of blowdown solvent and flash
gas.
9.0 Related Programs
No data available.
REFERENCES
1. Holder, H.L., Diglycolamine - A Promising New Acid-Gas Remover, The Oil and
Gas Journal, 64 (May 2): 83-86, 1966.
2. Fluor Corporation, Gas Treating - 95/J340/6/2, Los Angeles, California,
5 pages.
3. Dingman, J.C., and T.F. Moore, Compare DGA and MEA Sweetening Methods,
Hydrocarbon Processing, 47 (7): 138-140, 1968.
4. Riesenfeld, F.C. and A.L. Kohl, Gas Purification, Second Edition, Gulf
Publishing Company, Houston, Texas, 1974, pp 31-2.
5. Information supplied to TRW by R. L. Schendel of Fluor Engineers and Con-
structors, Inc., March 7, 1978.
B-79
-------
ALKAZID (ALKACID) PROCESS
1.0 General Information
1.1 Operating Principle^ ' ' - The absorption of the sour components of a
gas stream (H2S, C02, CS2) using an aqueous solution of alkali salts
of weak-nonvolatile amino acids. Three different absorption solutions
are used:
t Alkazid "M" for the removal of H2S and COg from gas streams con-
taining one or both components. The absorption solvent is an
aqueous solution of potassium monoethyl amino propionate.
Alkazid "DIK" for the selective removal of h^S from gases that
also contain C02- The absorption solvent is an aqueous solution
of the potassium salts of dimethyl amino acetic acid.
Alkazid "S" used in cases where feed gas contains appreciable quan-
tities of impurities; e.g., HCN, ammonia, CS2, The absorption
solvent contains sodium phenol ate.
1.2 Development Status - Commercially available.
1.3 Licensor/Developer - BASF AG
6700 Ludwigshafen
West Germany
1.4 Commercial Applications^ ' '
Process has been used commercially since the 1930's; 60 to 80 plants
in operation throughout the world, of which 20 were built by Davy
Powergas GMBH.
No commercial application known in the United States.
Process has been used to treat natural gas, synthesis gas, refin-
ery gas, flue gas and coal gas (commercially applied to Lurgi,
Koppers-Totzek and Winkler gasification plants).
B-80
-------
(3 *>}
2.o Process Information^ * '
2.1 Flow Diagram - see Figure B-14.
The lean solvent and the sour feed gas enter the top and bottom
of the absorber, respectively. The purified product gas exits
the top of the absorber. The rich solvent is pumped from the
absorber, heated by interchange with hot lean solvent and enters
the top of the stripper, within which the absorption reactions are
reversed with heat and acid gases are evolved. Steam is condensed
and separated from the acid gas and the acid gas is piped to sul-
fur recovery. The hot lean solvent is pumped from the regenerator
to a heat exchanger and then to the absorber.
(2)
2.2 Equipment - Conventional absorbers, stripping columns and conden-
sate accumulators.
2.3 Feed Stream/Requirements^ '
Temperature: 293°K to 313°K (68°F to 104°F)
Pressure: 0.11 to 7.0 MPa (16 to 1000 psia)
2.4 Operating Parameters^ '
2.4.1 Absorption Step
Temperature: 293°K to 313°K (68°F to 104°F)
Pressure: 0.11 to 7.0 MPa (16 to 1000 psia)
2.4.2 Stripping Step
Temperature: 378°K to 393°K (220°F to 250°F)
Pressure: 0.11 to 0.16 (16 to 23 psia)
2.5 Process Efficiency and Reliability^ - Purity of product gas depends
upon factors such as pressure, temperature, ratio between H2S and C02
in the feed gas. In high pressure applications a product gas purity
of less than 1 g/100 Mm3 (0.5 gr/100 scf) can be attained and in cer-
tain low pressure application as much as 46 g/100 Nm (20 ft/100 scf)
in the purified gas can be realized.
2.6 Raw Material Requirements - Solvent makeup requirements not known.
o
2.7 Utility Requirements^ - Typical requirements per 1000 Nm - (MMscf) of
feed gas with 0.7 vol % H2S and 30 vol % C02 at 7.58 MPa (1100 psig)
B-81
-------
fa
03
ro
LEGEND:
1. SOUR FEED GAS
2. PURIFIED PRODUCT GAS
3. ACID GAS
4. RICH SOLVENT SOLUTION
5. LEAN SOLVENT SOLUTION
6. SEMI-LEAN SOLVENT RECYCLE
7. STEAM-ACID GAS MIXTURE
8. CONDENSATE
Figure B-14. Alkazid Process
-------
and 298 K (77 F) with 5 ppm H2S in purified gas are as
follows:
Steam: 248 kg/1000 Mm3 (15,500 Ib/MMscf)
Cooling water: 1.25 x 104 */1000 Mm3 (93,500 gal/MMscf)
Electric Power: 8.13 kwh/1000 Mm3 (230 kWh/MMscf)
(2)
3.0 Process Advantagesv '
Solvent has a low vapor pressure
Solvent has high H2S carrying capacity
COS, CS2, and NH3 and mercaptans in feed gas do not affect performance
of Alkazid solvent.
4.0 Process Limitations
Solvent is degraded by certain contaminants (e.g., HCN, 02)'3 .
t Aluminum and special alloys are normally used for the hot-solution
pumps and lines, the reactivator, and the reboiler
the
(3).
Possible foaming problems, particularly during start-up (as is common
with any amine system).(2)
5.0 Process Economics
(7)
The capital investment for a 3.84 x 106 Nm3/day (134 MMscfd) Alkazid
plant is estimated at about $4 x 106 (1977 dollars).
6.0 Input Streams (Stream 1)
See Table B-19.
6.1 Feed Gas (Stream 1) - See Table B-19.
7.0 Discharge Steams
7.1 Product Gas Stream (Stream 2) - see Table B-19.
7.2 Acid Gas Stream (Stream 3) - see Table B-19.
8.0 Data Gaps and Limitations
Data gaps exist in the following areas:
Definition of the maximum allowable concentrations of various con-
taminants in the feed gas; e.g., COS, CS2, NH3, and mercaptans.
B-83
-------
TABLE B-19. TYPICAL PERFORMANCE DATA FOR ALKAZID
(7)
Gas quantity MM Nm3/day (MMSCFD)
Pressure MPa (psia)
Temperature °K (°F)
Analysis:
C02 Vol. %
CO Vol. %
H2 Vol. %
CH4 Vol. %
N2 + Ar Vol . %
H20 ppm
Streams
1 2 3
Feed Gas Product Gas Acid Gas
3.84 (143)
2.1 (308)
298 (77)
21
36
39
2
2
7835
3.74 (139)
2.05 (301)
298 (77)
19
36.9
40.1
2.0
2.0
100
0.128 (4.77)
4.0 (588)
377 (221)
74.7
--
2.0
--
--
23.3
OO
-------
The effect that various contaminants (NH3, carbonaceous matter,
trace metals, etc.) have on the process, and the ultimate fate
of such contaminants in the system.
9.0 Related Programs
No current information available.
REFERENCES
1. Reed, R.M. and Updegraft, N.C., Removal of Hydrogen Sulfide from Industrial
Gases; Industrial and Engineering Chemistry, Vol. 42, No. 11, 1950.
2. Wainwright, H.W., Egleson, G.C., et al, Selective Absorption of Hydrogen
Sulfide from Synthesis Gas; Industrial and Engineering Chemistry, Vol.
45, No. 6, 1953.
3. Gas Processing Handbook, Hydrocarbon Processing, April 1975.
4. Riesenfeld, F.C., and Kohl, A.L., Gas Purification, 2nd Edition, Gulf Pub-
lishing Company, Houston, Texas, 1974.
5. Handbook of Gasifiers and Gas Treatment Systems, Dravo Corp. for ERDA, FE-
1772-11, February 1976.
6. Maddox, R.N., Gas and Liquid Sweetening, Campbell Petroleum Series, 1974.
1. Information provided to TRW by L. H. Greives of Davy Powergas, Inc.,
June 16, 1978.
B-85
-------
SULFINOL PROCESS
1.0 General Information
1.1 Operating Principle - A combination of physical and chemical absorp-
tion of sour components of gas streams (e.g., h^S, CO^, COS, mercap-
tans) using the Sulfinol solvent (a mixture of Sulfolane and DIPA).
The physical absorption done by Sulfolane (cyclotetramethylene sul-
fone) and the chemical absorption done by DIPA (diisopropanolamine).
1.2 Development Status - Commercially available.
1.3 Licensor/Developer - Shell Development Company
One Shell Plaza
P.O. Box 2463
Houston, Texas 77001
1.4 Commercial Applications
1.4.1 110 plants in operation or under construction. Worldwide applica-
(9)
tions are as follows :
Approximately 70% are in natural gas sweetening.
The remaining are in the purification of refinery gases, synthesis
gases, LNG feedstock, and hydrogen(l).
1.4.2 Application to Coal Gasification
(2)
The Azot Sanayii, Koppers-Totzek, Coal Gasification plant located
at Kutahya, Turkey has a Sulfinol unit on line. At this facility
gas is taken from a gas holding tank, compressed to 2.72 MPa
(400 psi) then piped to the Sulfinol unit for hLS removal. No
further data are available.
2.0 Process Information
2.1 Flow Diagram - see Figure B-15 - The feed gas, Stream 1, enters the
bottom of the absorber, and the stripped solvent enters the top.
Sour gas passes counter-flow to the solvent. Purified product gas,
Stream 2, exits the top of the absorber and rich solvent exits the
B-86
-------
ACID GAS
PURIFIED GAS
CO
I
00
SOLVENT 1
RECLAIMER ,
LEGEND:
1. FEED GAS
2. SULFINOL MAKE-UP
(LOCATION NOW KNOWN}
3. PURIFIED GAS
4. ACID GAS
5. SLUDGE
STRIPPER BOTTOMS
PUMP
Figure B-15. Sulfinol Process
-------
bottom. The rich solvent enters the flashing stage, and the acid gas
evolved, Stream 3, exits the top of the stripper. Lean solvent exits
the bottom of the stripper, and is pumped to the absorber. A small
slip-stream of lean solvent is pipe to a reclaimer unit for removal
of degradation products.
2.2 Equipment - Absorbers, flash vessel, stripping columns, pumps, heat
exchangers.
(9)
2.3 Feed Stream/Requirementsv '
Temperature: 289°K to 330°K (60°F to 140°F)
Pressure: 0.15 to 9 MPa (22 to 1330 psia)
(9}
2.4 Operating Parameters^ '
Absorption step
Temperature: 289°K to 330°K (60°F to 140°F)
Pressure: 0.15 to 9 MPa (22 to 1330 psia)
Stripper column
Temperature: 380°K to 422°K (225°F to 300°F)
Pressure: 0.11 to 0.24 MPa (16 to 35 psia)
Solution Loading: 0.03 to 0.124 Nm3/l (4 to 17 scf/gal)
2.5 Process Efficiency and Reliability
Process is capable of producing pipeline specification gas.
Solvent has good affinity for sour components at low-to-medium
partial pressures and has high affinity for sour components
at high partial pressures. See Table B-20 for solubilities
versus partial pressure.
Sulfinol units presently on line display an ease of operation with
minimum problems due to system upsets(5;.
In natural gas and. synthesis gas applications with feed gas as
follows(8):
H2S: 0 to 53 MOL %
C02: l.l to 46 MOL %
B-88
-------
TABLE B-20. SOLUBILITY OF HYDROGEN SULFIDE IN SULFINOL SOLVENT
(1)
==
Partial Pressure H?S
MPa (psia) L
0.14 (20)
0.54 (80)
0.68 (100)
1.02 (100)
1.36 (200)
Equilibrium Solvent
Nm3 HzS/lOO 1 of
(SCF/gal)
Loading
Solvent
3.7 (5)
7.5 (10)
9.0 (12)
11.2 (15)
13.5 (18)
purified gas attained:
hLS:
-------
Viton material should be used for resilient seats in butterfly
valves in acid-gas service.
t Solvent will remove paint unless immediately removed; good house-
keeping of spills and drips is necessary.
t Solvent is relatively expensive ($3.50 to $5.00/gal in 1977).
(3 5 9)
3.0 Process Advantagesv ' ' '
Process can produce pipeline specification gas.
The presence of heavy liquid hydrocarbons, including crude oil in the
contactor does not cause foaming.
0 Solvent is noncorrosive which allows carbon-steel materials to be used
throughout the system except high parts which contact rich solvent.
t The heat capacity of the solvent varies with composition, but is about
0.38 cal/gram °K at 339°K (0.678 Btu/lb OF at 15QOF).
t As the partial pressure of H2S in the feed gas increases, the flow rate
of the solvent on a volume-to-volume basis decreases.
The preceding two items will cause a relative lowering of utility costs
compared to other alcoholamine processes.
Solvent does not expand when frozen.
In certain cases, H2S can be selectively removed in the presence of CO,,.
4.0 Process Limitations
Greatest appllability when:
1. H2S/C02 ratio in feed gas is 1:1 or greater
2. Acid gas partial pressure is 0.68 MPa (100 psia) or greater
Solvent is expensive
Some hydrocarbons are absorbed in the solvent and may appear in the con-
centrated acid gas*
*In subsequent sulfur recovery using the Claus process, discoloration of pro-
duct sulfur can be experienced if the feed gas has more than 0.3 mole % of
aromatics or if the acid gas contains more than 2.0 mole % hydrocarbons or
if the C5+/H2S mole ratio is greater than 0.005. To avoid such problems, a
carbon adsorption unit may be used ahead of the sulfur plant to remove
organics^3'.
B-90
-------
5.0 Process Economics
Sulfinol costs - $2.90 per 3.78 a (gallon)
Royalty must be paid to licensor
3
Typical requirements per 1000 Mm (MMscf) of gas at 2.8 MPa (397 psig)
and 2950K (720F) containing 0.46% H?S and 4.9 mole % C02, with 3 ppm
H2S and 0.05 mole % C02 in purified gas are estimated:
Steam: 16 kg/1000 Nm3 (10,000 Ib/MMSCF)
Cooling water: 1
Electric power: 2.12 kwh/1000 Nm3
Solvent: 1
6.1 Input Streams
6.1 Feed Gas (Stream 1) - see Table B-21.
6.2 Sulfinol Makeup (Stream 2)
t Typically less than 600g/1000 Nm3 (<35 Ib/MMscf)
7.0 Discharge Streams
7.1 Purified Gas Stream (Stream 3) - see Table B-22.
7.2 Acid Gas Stream (Stream 4) - see Table B-23.
7.3 Sludge* (Stream 5)
8.0 Data Gaps and Limitations
Data gaps exist in the following areas:
Process information from the Azot Sanayii coal gasification plant
at Kutahya, Turkey.
Characterization of gaseous and liquid feed streams for refinery/
natural gas applications.
Characterization of off-gas and liquid waste streams for refinery/
natural gas applications; e.g., purified gas, acid gas and sludge.
*Reclaimer bottoms generated vary from installation to installation. In coal
conversion plants, they are generally mixed with waste ash.
B-91
-------
TABLE B-21. SOUR GAS FEED STREAM TO SULFINOL UNIT AT PERSON
GAS PLANT, KARNES, TEXAS
Flow rate: 90.56 x 104 Nm3/day (32 MMSCFD)
Pressure: 69 MPa (1000 psig)
Component
Cl
c2
C3
C4
C5
C6
C7
C8
c9+
Benzene
Toluene
Xyl ene
Total Hydrocarbon
N2
co2
H2S
COS
RSH
Mole Percent
81.57
5.82
1.85
1.03
0.45
0.15
0.06
0.043
0.004
0.013
0.010
Trace
91.00
0.50
6.90
1.60
(7 ppm)
(19 ppm)
B-92
-------
TABLE B-22. OPERATING DATA FOR SULFINOL UNIT AT PERSON GAS PLANT
Item
Feed gas composition, vol. %
co2
Feed gas rate, Nm3 (MMSCFD)
Solvent flow, 1/m (gpm)
Solvent loading
vol. acid gas/vol. solvent
3
Residue gas, g/Nm
(gr H2S/100 scf)
Observed Case I
1.6
6.9
90.56 x 104 (32)
1192 (315)
45
0.01 (0.6)
Observed Case II
1.6
6.9
90.56 x 104 (32)
1268 (335)
42.5
<0.002 (<0.1)
TABLE B-23. COMPOSITION OF THE ACID GAS STREAM FROM SULFINOL UNIT
AT PERSON GAS PLANT, KARNES, TEXAS
Component
C1-C4
C5+
Aromatics
£ Hydrocarbons
H2S
co2
Acid Gas Composition,
Vol.% Observed
1.20
0.50
0.20'
1.90
18.0
80.1
B-93
-------
t Definition of the maximum allowable concentration of various con-
taminants in the feed gas; e.g., CS2, NH3, mercaptans, etc,
0 The effect that various contaminants (NF^, carbonaceous matter,
trace metals, etc.) have on the process, and the ultimate fate'
of such contaminants in the system.
9.0 Related Programs
No data available.
REFERENCES
1. Reisenfeld, F.C. and Kohl, A.L., Gas Purification, 2nd Edition, Gulf Pub-
lishing Company, Houston, Texas, 1974.
2. Dunn, C.L., E.R. Freitas, et al, First Plant Data from Sulfinol Process,
Hydrocarbon Processing, 44 (4), p. 137-140, April 1965.
3. Stecher, P.G., Hydrogen Sulfide Removal Process, Noyes Data Corporation,
1972.
4. Goar, B.G., Sulfinol Process has Several Key Advantages, The Oil and Gas
Journal, p. 117-120, 30 June 1969.
5. Gas Processing Handbook, Hydrocarbon Processing, p. 96, April 1975.
6. Dunn, C.L., E.R. Freitas, et al, Shell Reveals Commercial Data on Sulfinol
Process, The Oil and Gas Journal, p. 89-92, 29 March 1965.
7. Dravo Corporation, Handbook of Gasifiers and Gas Treatment Systems, ERDA
FE-1772-11, Washington, D.C., February 1976, p. 139-141.
8. Information supplied to TRW by J. M. Duncan of Shell Development Company,
December 8, 1977.
B-94
-------
AMISOL PROCESS^
1.0 General Information
1.1 Operating Principle - A combination of physical and chemical absorp-
tion of the sour components (H2$, C02, HCN, etc.) of a gas stream
using an aqueous solution of methanol and diethanolamine or
monoethanolamine.
1.2 Development Status - Commercially available.
1.3 Licensor/Developer - American Lurgi Corporation
377 Route 17
Hasbrouck, New Jersey 07604
1.4 Commercial Applications - One system on stream in an ammonia and
methanol complex treating raw gas produced by partial oxidation of
fuel oil.
2.0 Process Information
2.1 Flow Diagram (see Figure B-16) - Sour feed gas, Stream 1, enters the
bottom and lean solvent and water enter the top of the absorber.
The sour gas passes counter-flow to the solvent and the purified
product gas, Stream 2, exits the top of the absorber. Rich solvent
exits the bottom of the absorber and is piped to the regenerating
unit. Flashing occurs in the regenerating unit and the acid gas
evolved, Stream 3, exits the top of the regenerator and is piped to
sulfur recovery. Lean solvent exits the bottom of the regenerator.
A portion of the lean solvent is piped to the absorber and a portion
is piped through a heat exchanger back to the regenerator.
A methanol/water solution is piped from the absorber to a methanol
distillation process where methanol and water are separated and
recycled back to the system.
B-95
-------
10
LEGEND
1. SOUR FEED GAS
2. PURIFIED PRODUCT GAS
3. ACID GAS
4. RICH SOLVENT
5. METHANOL AND WATER MIXTURE
6. LEAN SOLVENT TO ABSORBER
7. LEAN SOLVENT TO REGENERATOR
8. METHANOL
9. METHANOL AND WATER TO DISTILLATION
10. METHANOL ANDWATER FROM DISTILLATION
11. WATER/METHANOL TO SYSTEM
12. WATER/METHANOL RECYCLE
13. WATER
14. STEAM
Figure B-16. Flow Diagram - Ami sol Process
-------
2.2 Equipment - Conventional absorbers, stripping columns, and flash
vessels.
2.3 Feed Stream Requirements
Temperature: ?
Pressure: ?
Others: ?
2.4 Operating Parameters
Temperature: ?
Pressure: ?
Others: ?
2.5 Process Efficiency and Reliability - Process can reduce hLS content
of gas to less than 0.1 ppm, and CO,, content to less than 5 ppm.
2.6 Raw Material Requirements - Amine makeup is less than 160 kg/
1000 Nm3 (10 Ib/Mscf) of sour gas treated.
2.7 Utility Requirements: ?
2.8 Miscellaneous: ?
3.0 Process Advantages
Due to the solvent's noncorrosive nature, equipment used in the
Amisol process can be constructed of carbon steel.
4.0 Process Limitations
Due to low boiling point of methanol, methanol vapors may contami-
nate purified gas stream.
5.0 Process Economics
6-0 Input Streams
6.1 Sour Gas Stream (Stream 1) - Table B-24.
7-0 Discharge Streams
7.1 Purified Gas Stream (Stream 2) - Table B-24.
7.2 Acid Gas Stream (Stream 3) - Table B-24.
B-97
-------
TABLE B-24. AMISOL INPUT AND DISCHARGE STREAM COMPOSITION*
Composition
co2
H2S
COS
CO
H2
H2
CH4
Pressure
Temperature
Streams
1
Feed Gas
6.6 Vol %
0.38 Vol %
152 ppm
44.9 Vol %
47.6 Vol %
0.2 Vol %
0.3 Vol %
2.98 MPa
293°K
2 3
Purified Gas Acid Gas
10 ppm
0.3 ppm
0.1 ppm
48.2 Vol %
5 .3 Vol %
90.7 Vol %
4.4 Vol %
0.15 Vol %
2.3 Vol %
2.4 Vol %
*Based on tests with a gas produced from residual oil by pressure gasi-
fication with oxygen and steam.
See Figure B-16 for stream locations in process.
8.0 Data Gaps and Limitations
Applicability to coal conversion processes
Applicable feed gas stream requirements
Temperature and pressure ranges
Concentration of various contaminants
t The effect that contaminants such as NH3, carbonaceous matter,
trace minerals, etc., have on the process, and the ultimate fate
of such contaminants in the system.
The reliability and efficiency of the process.
B-98
-------
Process economics
Typical utility cost or usage
Solvent makeup required
Other operating cost
Miscellaneous process unique maintenance requirements,
9.0 Related Programs
REFERENCES
1. Bratzler, K. and Doerges A., Amisol Process Purifies Gases, Hydrocarbon
Processing, April 1974.
B-99
-------
BENFIELD HOT CARBONATE PROCESS
1.0 General Information
1.1 Operating Principles - Chemical absorption of carbon dioxide and
hydrogen sulfide at an elevated pressure by a hot potassium carbonate
solution containing activators such as diethanolamine as (absorption
catalyst).* Absorbed gases are stripped from the carbonate solution
by steam at low pressure.
1.2 Development Status - Commercially available since early 1960's.
1.3 Licensor/Developer - The Benfield Corporation
615 Washington Road
Pittsburgh, Pennsylvania 15228
(i 2}
1.4 Commercial Applications^ ' ' - Over 400 Benfield systems are operat-
ing worldwide. Several units are operating on synthesis gases pro-
duced by partial oxidation of heavy petroleum fractions, and one unit
has operated successfully at a Lurgi coal gasification plant in
Westfield, Scotland.
2.0 Process Information
2.1 Flow Diagrams (see Figure B-17 and B-18f)
*Another hot carbonate process, Catacarb (for which Fickmeyer & Associates of
Prairie Village, Kansas are the licensors), employs amine borates as absorption
catalysts. The process flow diagrams and performance capabilities of Catacarb
units are essentially the same as those of Benfield units. The Benfield proc-
ess is covered by this data sheet since it is much more widely used than Cata-
carb and for which more extensive data are available.
Several variations of the Benfield process have been used commercially, de-
pending on the degree and selectivity of C02/H2S removal required. These
variations are primarily in the solution circulation patterns and degree of
regenerated solution coding. Two of the more commonly used designs of the
Benfield process are covered in this data sheet.
B-100
-------
ABSORBER
LEGEND:
1. FEED GAS
2. STEAM
3. PURIFIED GAS
4. LEAN CARBONATE SOLUTION
5. RICH CARBONATE SOLUTION
6. ACID GAS
7. MAKE-UP CARBONATE SOLUTION
(LOCATION NOT KNOWN)
Figure B-17. Benfield Split Stream Process
B-101
-------
ABSORBER
LEGEND:
1. FEED GAS
2. STEAM
3. PURIFIED GAS
4. LEAN CARBONATE SOLUTION
5. RICH CARBONATE SOLUTION
6. ACID GAS
7. MAKE-UP CARBONATE SOLUTION
(LOCATION NOT KNOWN)
Figure B-18. Benfield Hipure Process
B-102
-------
Benfield Split Flow Process (for bulk C02 and/or H2S removal)^ -
Gas feed enters the bottom of the absorber and flows counter-
current to lean carbonate solution. The bulk of the carbonate
solution is introduced to an intermediate point of the absorber
at about 383°K (230°F); the remainder is introduced at the top of
the absorber at 339°K to 386°K (150°F to 200°F). Purified gas
leaves the top of the absorber while rich solution is pumped to
a regenerator. Rich solution is stripped with steam generated
in a reboiler. Acid gases leave the regenerator through a reflux
drum.
Hipure Process (for obtaining low residual C02 and fyS levels)
Gas feed enters the bottom of the absorber and flows counter-
current to solutions which enter at the top and at the middle of
the absorber. The solution entering the top is cooler than the
one entering the middle. The two solutions, which may differ in
composition, are handled in separate circuits (i.e., separately
transported to the regenerator and are stripped of acid gases
with steam). The separate solution circuits allow temperatures,
solution compositions, and flow rates to be employed which take
advantage of kinetic and equilibria differences between C02 and
H2S absorption to effect high H2S or C02 removal.
2.2 Equipment - Packed or trayed tower design for absorber and regener-
ator. Carbon steel is used for tower and packing (provided that
corrosion inhibitors are added to the circulating Benfield solution).
Stainless steel is recommended for reboiler tubes, control valves
fo\
and solution pumpsv ' .
2.3 Feed Stream Requirements
Tempera ture(s): Feed stream temperature may vary from ambient to
450°K (350°F). Maximum carbonate solution temperature is usually
less than about 400°K (280°F).
Pressure"': Commercial installations have operated with feed
gas pressures ranging from 0.7 to 7.6 MPa (100 to 1100 psia).
Feed Gas Composition: Depending on the mode of operation, feed
gases with 0)2' partial pressures below 34 to 100 KPa (5 to 15
psia) are not economically handled by the Benfield process(°).
Hydrogen sulfide containing gas streams cannot be economically
treated unless some carbon dioxide is also present. This is
because carbon dioxide would be lost from the stream during
regeneration, allowing build-up of nonregenerable potassium
sulfide (
B-103
-------
2.4 Operating Parameters
Temperature(s): Absorption and regeneration usually occur iso-
thermally; practical operating temperatures range from 383°K to
4000K (2300F to 2800F).
Pressure' ': Absorber can operate above 0.7 MPa (100 psia).
Economics of operation below 0.7 MPa (100 psia) are not favor-
able. Regenerator operates at slightly above atmospheric pressure.
Loading: Stoichiometric maximum amounts of carbon dioxide that
can react with the Benfield solution range from 0.04 Nm3/ji 3
(5.5 scf/gal) for 20% weight potassium carbonate to 0.095 Mm /a
(13 scf/gal) for 40% potassium carbonate. The amounts of solution
circulated have been about 32 £/Nm3 (0.23 gal/scf) of CO? removed
in the split stream mode and 34 Jl/Nm3 (0.24 gal/scf) of C02
removed in the Hi pure mode for a feed gas containing 34%
C02(2,10).
2.5 Process Efficiency and Reliability - The split flow mode of operation
is capable of 95% plus removal of both C02 and H^9'10'. Hipure
operation is capable of reducing COp levels to less than 10 ppmv and
H2S levels to about 1 ppmv^2'3\ From 75% to 95% removal of both COS
and CS2 and essentially complete removal of small amounts of mercap-
tans have been reported for commercial operations^ . About 85%
removal of thiophene has been reported in a commercial Benfield
unit^1 '. HCN and SO^ are essentially completely removed from feed
gases by Benfield units, although some buildup of hydrolysis/
oxidation products (formate, S0,~, SO,") in the carbonate solutions
(10) 3 4
can occurv '.
The Benfield process can be designed to effect a high degree of H2S
removal while restricting C02 removal to 10% to 40%' '. hLS reacts
approximately 3.6 times as fast as C02 with the Benfield solution,
and gas-solution contact times can be adjusted to take advantage of
these absorption rate differences. The relative absorption capacities
and absorption rates for several feed gas constituents are listed in
Table B-25(8'10).
2.6 Raw Material Requirements - Makeup Benfield solution contains 20% to
(8^
40% I^COg in water with activator and corrosion inhibitor^ . Makeup
requirements depend upon contaminant buildup (primarily formate and
B-104
-------
sulfate), and leak prevention. Commercial plants typically use one
solution inventory in a 3- to 5-year period.
TABLE B-25. CAPACITY AND ABSORPTION RATES OF CONSTITUENTS
IN BENFIELD SOLUTION
Component
co2
H2S
COS
cs2
CH3SH
NH3
Approximate Capacity
Nm3/I03 1 (scf/gal)*
0.027 (3.7)
0.038 (5.2)
(hydrolyzes to H2S)
(hydrolyzes to H2S)
0.08 (0.11)
0.27 (0.37)f
Relative
Capacity
1
1.4
0.03
0.10t
Relative
Absorption Rate*
1
3.6
0.36
0.10
1.2
3.5t
*35% KzCOs solution, 383°K (230°F; capacity at equilibrium partial pressure
of 13.6 KPa (2 psia). Rate measured at solution loadings equivalent to
equilibrium partial pressure of 13.6 KPa (2 psia).
Solution capacity for ammonia according to recent data^ ' is reported to
be much greater than given in this table, although exact numbers are not
publicly available.
2.7 Utility Requirements*
2.7.1 Split Flow Mode^ ':
Basis: feed gas at 393°K (250°F) and 4.3 MPa (630 psia) and
containing 34% C02; outlet gas 1000 ppmv C02
Steam at 0.44 MPa -,
(65 psia): 1.3 kg/NrrT (0.092 Ibs/scf) of C02
Electric power: 0.036 kwh /Nm3 (0.001 kwh /scf) of C02
Cooling water: ?
Utility requirements are primarily dependent on feed and output acid gas
concentrations.
B-105
-------
2.7.2 Hipure Mode
Basis: feed gas at 393°K (250°F) and 4.3 MPa (630 psia) and
containing 5.4% C02, 1.5% ^S; purified gas contain-
ing 100 ppmv C02, 2 ppmv H2S
Steam at 0.44 -
MPa (65 psia): 3.6 kg/Nnr (0.22 Ibs/scf) total acid gas
Electric power: 0.057 kwh /Nm (0.0020 kWh /scf) of total acid gas
Cooling water: 4.1 1/Nm (0.03 gal/scf) of total acid gas
Basis: feed gas at 393°K (250°F) and 4.3 MPa (630 psi) and
containing 44% COg, 9.4% H2S; purified gas contain-
ing <1500 ppmv C02, 2 ppmv H2S.
Steam at 0.44 o
MPa (65 psia): 1.9 kg/Mm (0.13 Ibs/scf) of total acid gases removed
Electric power: 53 kwh /Nm3 (0.0019 kwh /scf) of total acid gases
removed
Cooling water: quantity not available; a cooling load of 371 kcal/
Nm3 (44 Btu/scf) is reported.
3.0 Process Advantages
Commercially available and proven; over 400 Benfield units are in com-
mercial operation.
Benfield hot carbonate solution does not degrade significantly in the
presence of COS, CS,,, mercaptans, HCN, NH3, or particulate matter.
t Benfield systems (Hipure and split flow) can be designed so as to
selectively remove bulk H2S with minimal C02 removal by taking advan-
tage of different absorption rates of COg and h^S. In this mode of
operation, and depending on the feed composition, the bulk H2S removal
can generate a feed suitable for sulfur recovery in a Claus plant.
Hipure mode of operation can achieve low levels of C09 and H?S in out-
let gas. * L
Being a hot carbonate system, regeneration energy required in the
Benfield process is generally lower than that required in the amine
based systems.
As with other carbonate systems, the Benfield process does not absorb
hydrocarbons or other organics to the extent that physical solvents do.
Depending on the mode of operation, the nearly isothermal regeneration
and absorption eliminates or minimizes extensive heat exchange
equipment.
B-106
-------
4.0 Process Limitations
t Benfield solution can cause corrosion problems; however, addition of
corrosion inhibitors to the Benfield solution allows mostly carbon
steel construction.
High removal efficiency requires moderate to high pressure operation.
t Hot carbonate systems cannot be operated in a regenerate manner for
H2S removal if C02 is not present in the gas stream (a minimum of about
1000 ppmv of C02 should be present in the feed gas).
Energy costs for selective H2S removal are strongly dependent on
desired outlet concentration and could be prohibitive if a very high
removal efficiency is desired.
Depending on the H2S and C02 levels in the feed gas and the mode of
operation, the regenerated acid gas stream may not contain a high
enough concentration of H^S to allow sulfur recovery in a Claus plant.
5.0 Process Economics
(31
Capital costs reported in a 1974 publicationv ' for typical Benfield
systems (either split flow or Hipure) are estimated at $66-88/daily kg-
mole C00 removal capacity ($30-40/daily Ib mole). Operating costs
(?}
reported in a 1974 publicationv ' for these systems are approximately
$30/104 Nm3 of C02 removed ($1/103 scf).
The capital cost for a Hipure system designed for selective H2$
removal (feed -7000 ppmv H2$, outlet gas -200 ppmv) was estimated at about
$200/103 Nm3 of daily feed capacity ($6/103 scf)^. Operating cost for
selective H?S removal is highly dependent on energy costs since outlet gas
purity and system pressure strongly influence energy consumption in the
process. No operating cost data have been reported for the selective H2S
removal Hipure system.
6.0 Input Streams
6.1 Feed Gas (Stream 1) - see Tables B-26 and B-27.
6.2 Steam (Stream 2) - see Section 2.7.
6.3 Makeup Carbonate Solution (Stream 7) - The makeup aqueous solution
contains 20% to 40% potassium carbonate and an undisclosed quantity
of activator(8). The solution usually also contains corrosion
B-107
-------
TABLE B-26. LEVELS OF C02 AND H2S IN COMMERCIAL BENFIELD PROCESS FEED, PRODUCT, AND ACID GAS
STREAMS (STREAMS 1, 3 AND 6)
Applications (Ref.)
Natural Gas^ '
Natural Gas^" '
Natural Gas*6'
LNG<2>
* (2)
H? Purification
Partial Oxidation^1'
Partial Oxidation^
Coal Gasification^1'
Feed Flow Rate
106 Nm3/day
(106 scf/day)
4.1 (150)
4.4 (165)
0.8 (31)
13 (500)
1.6 (58)
0.17 (6.4)
1.4 (52)
0.64 (24)
Feed Gas
co2 (%)
7
12.3
44.4
4.9
30.6
26
5.2
28
(Stream 1)
H2S (%)
16
0.8
9.4
4.7
--
1.6
0.93
0.6
Purified Gas
CO, (%) h
c.
1.0
0.5
<0.1
0.005
0.1
0.011
0.95
0.8
(Stream 3)
\2S (ppmv)
500
20
1.2
<1
--
<1
45
70
Acid Gas*
co2 (%)
27
94
83
52
--
94
85
98
(Stream 6)
HQ 1 °/\
/} O I to i
C.
72
6
17
48
--
6
15
2
oo
i
o
00
Calculated levels based on feed and purified gas concentrations.
-------
TABLE B-27. LEVELS OF TRACE SULFUR COMPOUNDS IN BENFIELD FEED
AND PRODUCT GASES (STREAMS 1 AND 3)
Mode of Operation
Feed Flow Rate
106 Mm3/ day
(106 set/day)
Composition (ppmv)
COS1"
cs2
CH3SH
Thiophene
Application (Ref.)
Pilot Plant^10)
Single
0.53
Feed
1300
331
75
Stage*
(20)
Product
11
135
34
Pilot Plant^10)
Hi pure
0.53 (20)
Feed Product
1300 <1
331 50
75 6
Natural Gas^
Hi pure
0.84 (31)
Feed Product
820 27
--
Coal Gasification' '
Split Flow
0.64 (14)
Feed Product
-120* -30*
2.3 0.5
4 <0.4
§
ro
t
*Single stage operation is similar to the split flow system shown in Figure B-17, but with lean
solution injection at the top of the absorber only.
From 73% to 99% COS removal has been reported for commercial Benfield systems
±
'Actual concentration data not available; typical COS levels in Lurgi product gas are about 2% of
H2S levels(H). About 75% removal of COS has been reported for the Benfield system operating on
Lurgi gas at Westfield, Scotland(6).
§About 85% thiophene removal reported.
-------
inhibitors (such as metavanadate). Quantitative data on makeup
requirements are not publicly available.
7.0 Intermediate Streams
7.1 Lean Carbonate Solution (Stream 4) - No actual operating data
available.
7.2 Rich Carbonate Solution (Stream 5) - No actual operating data avail-
able. See Section 2.5 for the capacity of a Benfield solution for
various components.
8.0 Discharge Streams
8.1 Purified Gas (Stream 3) - see Tables B-26 and B-27.
8.2 Acid Gas (Stream 6) - No actual operating data available. See
Table B-26 for calculated C02 and \\^S levels in this stream for
several applications.
8.3 Slowdown Carbonate (Stream 8) - See Table B-28 for the composition
of a Benfield solution after two years of operation at a Lurgi coal
gasification facility.
9.0 Data Gaps and Limitations
Data gaps for the Benfield process relate primarily to the degree of
removal of various trace constituents from gases likely to be encountered
in coal gasification. Actual operating data for the Benfield unit which
handled Lurgi gas at Westfield, Scotland are not available at present.
10.0 Related Programs
Under DOE sponsorship, test of a slagging Lurgi Gasifier are cur-
rently under way at Westfield, Scotland. If the product gas is treated
in the existing Benfield unit, data may be generated which could help
in the evaluation of hot carbonate process performance when applied to
coal gasification.
The Synthane coal gasification pilot plant at Bruceton, Pa. has an
installed Benfield unit for CCL and hLS removal from product gas after
ship conversion. As far as is known, this unit has not been operating
to date.
B-110
-------
TABLE B-28. COMPOSITION OF BENFIELD SOLUTION FOR THE UNIT OPERATING ON LURGI
COAL GASIFICATION PRODUCT GAS AT WESTFIELD, SCOTLAND*(10)
Constituent
Concentration
K2C03 (wt %)
KHC03 (wt %)
Formate1" (wt %) (as HC02K)
Sulfide (ppm) (as KSH)
S203= (ppm)
SCN" (ppm)
S03 (ppm)
Suspended Solids (ppm)
22.0
9.8
2.1
1500
350
Trace
610
15 to 800
*After two years of operation.
fFormate may be formed by hydrolysis of either CO and HCN in
the alkaline carbonate solution according to the following
reactions:
CO + OH"
HCO,
HCN + H20 + OH
HC02" + NH3
B-111
-------
REFERENCES
1. McCrea, D.H., The Benfield Activated Hot Potassium Carbonate Process:
Commercial Experience Applicable to Fuel Conversion Technology, Symposium
Proceedings: Environmental Aspects of Fuel Conversion Technology, II.
(December 1975, Hollywood, Florida) EPA-600/2-76-149, June 1976, p. 217-223.
2. Benson, H.E. and Parrish, R.W., Hipure Process Removes CO?/H?S, Hydrocarbon
Processing, April 1974, p. 81-22
3. McCrea, D.H. and Field, J.H., The Purification of Coal Derived Gases:
Applicability and Economics of Benfield Processes, 78th National.AIChE
Meeting, Salt Lake City, Utah, August 18-20, 1974, Paper No. 29b.
4. Ruziska, P.A., Packings for Hot Carbonate Systems, Chemical Engineering
Progress, Vol. 69, No. 2, February 1973, p. 67-70.
5. Dravo Corp., Handbook of Gasifiers and Gas Treatment Systems, ERDA Document
No. FE-1772-11, February 1976.
6. Parrish, R.W. and Field, J.H., The Benfield Process in Coal Gasification,
24th Annual Gas Conditioning Conference, University of Oklahoma (Norman),
March 14-15, 1974.
7. Kohl, A. and Riesenfeld, F., Gas Purification, Gulf Publishing Co.,
Houston, Texas, 1974.
8. Maddox, R.N., Gas and Liquid Sweetening, Campbell Petroleum Series, Norman,
Oklahoma, 1974.
9. Benson, H.E., J. H. Field, et al, Improve Process for C02 Absorption
Uses Hot Carbonate Solutions, Chemical Engineering Progress, Vol. 52, No.
10, 1956.
10. Parrish, R.W. and Nelson, H.B., Synthesis Gas Purification Including
Removal of Trace Contaminants by the Benfield Process, ACS Symposium,
Los Angeles, California, March 31-April 5, 1974.
11. Woodall-Duckham, Ltd, Trials of American Coals in a Lurgi Gasifier at
Westfield, Scotland, Final Report, R&D Rpt. No. 105, FE-105, Crawley,
Sussex, England, November 1974.
12. Information provided to TRW by D. H. McCrea of the Benfield Corporation,
January 9, 1978.
B-112
-------
GIAMMARCO-VETROCOKE (G-V) PROCESS
1.0 General Information
1.1 Operating Principles^1' - Hydrogen sulfide and carbon dioxide are
removed from feed gas in two stages using alkaline solutions con-
taining arsenate/arsenite. Hydrogen sulfide is first absorbed
and oxidized indirectly by the arsenate solution and the resulting
sulfur is recovered by flotation. The H2S absorbing solution is
regenerated by air. C02 is then absorbed from the H2S free gas
by a second alkaline arsenite solution. Rich C02 solution is
regenerated at reduced pressure using hot air and/or steam.
1.2 Development Status - Commercially available.
1.3 Licensor/Developer^ ' - Vetrocoke, Sp.A.
Marghera, Italy
10}
1.4 Commercial Appl i cations v ' - Several hundred plants have operated
worldwide for treating natural gas, hydrogen and synthesis gas.
One plant is known to treat natural gas in the United States.
2.0 Process Information
2.1 Flow Diagram (see Figure B-19 for the design of a G-V plant for
removal of both C02 and H2S from feed gas) - Although other versions
of the process have been used commercially, this design illustrates
the basic units both for H2S and for C02 removal.
2.1.1 H2S Removal - Feed gas enters the H2S absorber where a solu-
tion containing potassium arsenite/arsenate absorbs H2S and
reacts, according to the reaction:
KH2As03 + 3H2S -> KH2AsS3 + 3H20
The trithioarsenite then reacts with arsenate in a digester:
KH2AsS3 + 3KH2As04 -
B-113
-------
C02
IVBSORBER
I SOLUTION A
1 REDUCER 1
13
CO
I
CO2 FLASH TANK
LEAF
FILTER
HjS
IVBSORBER
oc
«*l
8 3 I
« 2
OC
17
15
J1
HEAT
RECOVERY
TOWER
16
T
12
= 2
0 0
V
11
t
FILTER
20
10
OC
UJ
s
IB
19
1. FEED GAS
2. HjS FREE GAS
3. COj FREE GAS
4. LEAN M2S REHKOVAL SOLUTION
5. RICH H^ REMOVAL SOLUTION
S. LEAN COj REMOVAL SOLUTION
7. RICH CO2 REMOVAL SOLUTION
8. PRODUCT COj
9. AIR
10. FILTERED Hf REMOVAL SOLUTION
11. SULFUR
12. H2S SOLUTION REGENERATION OFF GAS
13. FILTER SOLIDS
14. CO2 SOLUTION REGENERATOR OFF GAS
16. HOT WATER
1C. COOL WATER
17. HOT, MOIST AIR
18. LOWOOWN SOLUTION
19. MA«ur SOLUTION
20. FILTER WASH WATER
Figure B-19. Giammarco-Vetrocoke for h^S and CC^ Removal
(3)
-------
The monothioarsenate/arsenite solution is pumped to an acidi-
fication drum where C02 is added to lower the solution
alkalinity to a point where the auto oxidation/reduction
reaction below can occur:
3KH2As03S * 3KH2As03 + 3S
The solution is filtered to separate elemental sulfur and
is sent to a regenerator where air is added to oxidize part
of the arsenite according to:
6KH2As03 + 302 * 6KH2As04
2.1.2 C02 Removal - Gas from the H2S absorber enters a C02 absorber
where it encounters a K2C03/As203 solution. C02 is absorbed
according to the following reaction:
6CO~ + 2K,As07 + 3H90 -» 6KHCO,, + As00, and
£ O j L. o £ o
C02 + K2C03 + H20 -» 2KHC03
The C02 rich solution is flashed to release part of the
absorbed CCU, and is pumped to a regenerator column to be
stripped of the remaining C02 by air (or stream) at slightly
above atomospheric pressure. Flashed C0? is either vented
or used to acidify digested rich HoS absorber solution.
2.2 Equipment^2' - Packed towers for absorption and C02 regeneration;
open towers for sulfur flotation. Carbon steel is commonly employed
without undue corrosion problems. Sulfur is recovered by rotary
type filters.
2.3 Feed Stream Requirements^ - For absorption of H2S feed, pressures
as low as 0.1 MPa (15 psia) and temperatures as low as 410°K (100°F)
can be handled. For C02 absorption, higher feed pressures are
necessary (C02 removal efficiency decreases with decreasing C02
partial pressure). The upper temperature limit is about 523°K
(300°F), the approximate atmospheric boiling point of the G-V
solution (regeneration is conducted at near atmospheric pressure).
B-115
-------
Feed composition (e.g., trace sulfur and nitrogen species, particu-
lates, organics) is reported not to affect 6-V performance.
(2 3)
2.4 Operating Parametersv ' '
Absorbers:
Temperature: 410°K - 523°K (100°F - 300°F)
Pressure: 0.1 - 6.7 MPa (15 - 1000 psia)
3
H2S absorption solution loading: 0.0068 Mm H,,S/1
C02 absorption solution loading: .030 Mm3 C02/l
(.94 scf H2S/gal)
.030 Nm3 C02/l
(4.09 scf C02/gal)
Regenerators:
Temperature: 410°K - 523°K (100°F - 300°F)
Pressure: slightly above atmospheric
3
Air for hLS solution regeneration: 6.2 - 8.3 Nm /kg H2$
(106 - 141 scf/lb H2S)
Air for C02 solution regeneration: ?
2.5 Process Efficiency and Reliability^ ' - Process is capable of
reducing H2S levels to <1 ppmv and C02 to less than 1000 ppmv.
COS and C$2 and mercaptans are reported to be partially removed,
although no actual operating data are available.
At one large natural gas treating plant, corrosion was reported
not a problem, and the G-V process was reliable and met design
specifications. No data for other facilities are currently
available.
(2}
2.6 Raw Material Requirementsv ' - Makeup solution requirements -
quantities not known. H2S absorption solution contains from 0.5 to
15% K2C03; arsenic (as As203) concentration not known. C02 absorp-
tion solution contains 20% - 40% K2C03; arsenic concentration
not known.
B-116
-------
2.7 Utility Requirements
(2)
C02 Circuit H2S Circuit
Steam 0.56 kg/Nm3 SO,
">
removed
(.033 Ibs/scf)
Air ? 6.2-8.3 Nm3/kg H2S
removed
(106 - 141 scf/lb H2S)
Electricity 23 ktoh/Nm3 CO-
removed
(0.672 kwh/103 scf)
2.8 Chemical Hazards - Alkaline arsenite/arsenate solutions are highly
toxic and present both an occupational health and an environmental
hazard. Care must be exercised in handling G-V solutions and in
the prevention and cleanup of spills and leaks.
3.0 Process Advantages
Can achieve lower levels of H-S and CO- than conventional or amine
activated hot carbonate systems.
Has relatively low utility requirements.
Produces elemental sulfur rather than a gaseous H2S steam.
Produces a C02 offgas with a very low H2S concentration.
Unlike conventional carbonate systems, G-V can successfully treat
feed gases containing H2S without significant amounts of C02(4).
4.0 Process Limitations
The arsenite/arsenate solution(s) require special handling and
precautions.
Discharge streams such as the solution blowdown and filter wash water
will contain arsenic compounds.
High C02 removal requires moderate to high pressure operation.
t Product sulfur may require dearsenation prior to sale.
t H2S removal is practical only if the feed gas contains less than
15% H2S and sulfur to be removed is under 14 tonnes (15 tons) per
B-117
-------
5.0 Process Economics
No current data are available.
6.0 Input Streams
6.1 Feed Gas (Stream 1) - see Table B-29.
6.2 Air (Stream 9) - see Section 2.6.
6.3 Makeup Solution (Stream 19) - No data available.
7.0 Intermediate Streams
7.1 Lean hUS Removal Solution (Stream 4) - No data available.
7.2 Rich H2S Removal Solution (Stream 5) - No data available.
7.3 Lean C02 Removal Solution (Stream 6) - No data available.
7-4 Rich C02 Removal Solution (Stream 7) - No data available.
7.5 Filtered H2S Removal Solution (Stream 10) - No data available.
8.0 Discharge Streams
8.1 H2S Free Gas (Stream 2) - See Table B-29.
8.2 C02 Free Gas (Stream 3) - See Table B-29.
8.3 Product C02 (Stream 8) - 99+% C02 with less than 1 ppmv H2S can be
obtained^/.
(2)
8.4 Sulfur (Stream 11) - Washed sulfur contains about 0.3% arsenic^ ' -
No other data available.
8.5 H2S regeneration offgas - No data available.
8.6 Filter Solids - No data available; may consist primarily of ele-
mental sulfur.
8.7 C02 Solution Regeneration Offgas (Stream 14) - No data available;
will consist primarily of air and C02.
8.8 Slowdown Solution (Stream 18) - No data available.
8.9 Filter Wash Water (Stream 20) - No data available.
B-118
-------
TABLE B-29. PERFORMANCE OF(£ GjAMMARCO-VETROCOKE PLANT OPERATING
Stream
Flow Rate
Nm3/day
(scfd)
Temperature ,
°K (°F)
Pressure,
MPa (psia)
Composition
co2
H2S
Feed Gas
5xl06 (ISOxlO6)
311 (100)
6.9 (1015)
28%
0.2%
H2S Free Gas
5xl06 (ISOxlO6)
^28%
40 ppmv
C02 Free Gas
3.8xl06 (133xl06)
323 U22)
<2%*
<4 ppmv
*Levels of 0.05%C02 can be obtained.
9.0 Data Gaps and Limitations
Data gas pertains primarily to the composition and flow rates of various
intermediate and discharge streams. In particular, little data are avail-
able on product sulfur, filter solids, blowdown solution, and sulfur
wash water. Also, no data are available regarding trace constituents
in feed and product streams. Certain features of the process (e.g., the
solution reducer shown in Figure B-19) are not well defined by available
information.
10.0 Related Programs
No programs are known to be under way or planned which are aimed at an
environmental assessment of the G-V process.
B-119
-------
REFERENCES
1. Maddox, R. N., Gas and Liquid Sweetening, Campbell Petroleum Series,
Norman, Oklahoma, 1974.
2. Kohl, A., and Riesenfeld, F., Gas Purification, Gulf Publishing Company,
Houston, Texas, 1974.
3. Sweet-Gas Process Makes U.S. Debut, Chemical Engineering, September 19,
1960, p. 166-69.
4. Parrish, R. W. and Field, J. H., The Benfield Process in Coal Gasification,
24th Annual Gas Conditioning Conference, University of Oklahoma (Norman),
March 14-15, 1974.
B-120
-------
STRETFORD PROCESS
1.0 General Information
1.1 Operating Principles - Sulfur recovery based upon the liquid phase
oxidation of H2S to elemental sulfur in an alkaline solution of
metavanadate and anthraquinone disulfonic acid (ADA) salts.
1.2 Developmental Status - Commercially available.
1.3 Licensor/Developer - Originally developed by the British Gas
Corporation. Licensors include:*
Woodall-Duckham (USA) Limited
Division of Babcock Contractors, Inc.
921 Penn Avenue
Pittsburgh, Pennsylvania 15222
Peabody Engineered Systems
39 Maple Tree Avenue
Stamford, CT 06906
Wilputte Corporation
152 Floral Avenue
Murray Hill, N.J. 07974
Black, Sivalls and Bryson, Inc.
B.S. & B. Process Systems Division
4242 S.W. Freeway
Houston, TX 77027
1.4 Commercial Applications^ - 50 Stretford units are currently in
333
operation, with capacities ranging from 2.7x10 to 864x10 Mm D
(0.1 to 32 MMscfd). The range of installed Stretford units includes
purification of coke oven and producer gases as well as H2S removal
from natural gas. It has also been applied to clean Claus plant tail
gas in petroleum refineries.
*Each licensor incorporates its own process refinements. These refinements
are generally aimed at reducing the quantity of wastes generated and reagent
used in the sulfur recovery section. The following data sheets represent
the process by Woodhall-Duckham.
B-121
-------
2.0 Process Information
2.1 Flow Diagram - See Figure B-20.
Process Description - A raw gas stream is contacted counter-
currently with an aqueous solution of ADA (anthraquinone disul-
fonic acid), vanadium, anhydrous citric acid and sodium carbonate.
H2S in the gas is oxidized to elemental sulfur by the vanadic
salt, while the salt is reduced to the vanadous form. The
reactions involved are: H2S + Na2C03 J NaHS + NaHCOo in the
absorber and 4 NaVOs + 2NaHS + 2H20 j Na^Og + 2S + 4NaOH and
Na2V40g + 2NaOH + H20 + 2ADA £ 4NaV03 + 2ADA (reduced) in the
holding tank. The reduced liquor flows to the oxidizers* where
the vanadium is restored to the vanadic form by a redox reaction
with the ADA. Air is blown through the oxidizers to reoxidize
the ADA and separate the sulfur by froth flotation. This reac-
tion is: 2ADA (reduced) + 02 $ 2ADA + H20. The sulfur float
is sent to a centrifuge and separator where the product sulfur
(99.5% purity) is obtained. Side reactions involving HCN and
other (than H2S) sulfur and nitrogen compounds require that a
portion of the solution be blown down to prevent buildup of
these contaminants.
2.2 Equipment^ ' - Conventional absorbers, oxidation tanks and elemental
sulfur recovery equipment. The sulfur equipment employed varies
by licensor. Although not indicated in Figure B-20, it is often
necessary to incorporate a heater and/or evaporator in the Stretford
circuit to control water inventory.
2.3 Feed Stream Requirements
Pressure: unaffected by pressure
Temperature: ambient to 322°K
Loading: usually for H,,S loading (concentration) up to 15%
Other: process modifications, such as prewashing, may be required
if HCN, COS or CS2 are present^) since COS and CS2 are
not removed in the process and HCN acts in an irreversible
manner with the Stretford solvent (prewashing is used for
high HCN concentrations and increased blowdown for low,
below 50 ppm, HCN concentrations)^).
*When the Stretford process is operated on certain high pressure streams, a
flash drum may be incorporated between the absorber and the primary oxidizer
such that hydrocarbons are recovered at a low pressure fuel gas stream rather
than released to the atmosphere.(5)
B-122
-------
1. RAW GAS
ca
CA>
CENTRIFUGE
2.
3.
4.
5.
6.
7.
8.
9.
10
AIR
WATER MAKE-UP
CHEMICAL MAKE-UP
WATER
PRODUCT GAS
SOLVENT SLOWDOWN
SEPARATOR EFFLUENT
SULFUR
OXIDIZER VENT GAS
Figure B-20. Stretford Process Flow Sheet (.1)
-------
2.4 Operating Parameters
Absorption - Temperature: ambient to 322°K (120°F)
Pressure: to 7.0 MPa (1000 psia)
Loading: to 15% by volume
Regeneration - Temperature: ambient
Pressure: atmospheric
2.5 Process Efficiency and Reliability - H2S concentrations can be
reduced to less than 1 ppm; COS and CS9 are not removed to a signif-
/l \ c
icant degreev '. Reliability is high, with all stages free of
corrosion tendencies.
2.6 Raw Material Requirements^ '
t Make-up Chemicals - Basis: gas containing 1% H2S and 0.235 HCN
without separate HCN removal in kg/106 Mm3 (Ib/MMscf), with
reductive incineration for spent solution for its regeneration.*
ADA: 84 (5)
Sodium Vanadate: 1.0 (0.06)
Citric Acid: 168 (10)
2.7 Utility Requirements^ ' - Same basis as Section 2.6.
Steam @ 0.4 MPa (65 psia): 12,400 kg/106 Mm3 (730 lbs/106scf)
Electricity: 700 kwh/106 Nm3(19 kwh/106 scf)
Water: 0.3 x 1061/106 Nm3 (2160 gal/106 scf)
Fuel: 36 kcal/106 Nm3 (4 x 106 Btu/106 scf)
(A)
3.0 Process Advantagesv '
Can reduce the H2S concentration to less than 1 ppmv.
t Capable of relatively high turndown ratios.
Relatively low make-up chemical requirements.
*Fuel is used for reductive incineration of fixed salts in Stretford solution
blowdown in the specific design consideratedO).
B-124
-------
Relatively low maintenance requirements.
t Not pressure sensitive
4.0 Process Limitation^ '
Except for removal of minor quantities of methyl mercaptans, sulfur
compounds other than H2S (e.g., COS and C$2) are not removed.
High C02 concentrations in the feed gas can cause pH reduction and
reduced efficiency. Absorbers are enlarged to accommodate high C02
concentrations.
Undesirable side reactions (such as 2NaHS + 202-*Na2S203 + H20) can
occur causing thiosulfate formation and increasing solvent blowdown
requirements. This occurs where the system is operating beyond design
limits and 02 is contacted in the absorber or holding tank.
HCN in feed gas reacts with absorption solution to form thiocyanate.
Thiocyanate is stable in solution and must be purged to avoid reduc-
tion in absorption efficient.
Not usually economical for gas streams containing greater than 15%
H2S.
5.0 Process Economics
Capital costs are reported for a 0.45 x 10 Nm3/day (15 MMscfd) natural
gas sweetener operating at 0.3 MPa (45 psia) and better than 99.8% effi-
ciency as $1 million in April 1975 on a West Coast basiV . Capital
costs in 1964 dollars of a Stretford unit operating on a coal gas stream
of 0.33 x 106 Nm3/day (11 MMscfd) with better than 99.9% efficiency are
reported-as $330,OOo'3'. Operating costs will be dependent on the credit
(3)
taken for salable sulfur^ .
6.0 Input Streams- The following data are from a design for a Stretford
unit treating lean acid gas from the Rectisol unit at El Paso Natural
Gas Burnham coal gasification facility.
B-125
-------
6.1 Gaseous
Raw gas (Stream No. 1)
Vol % Vol %
C02 96.0 CH4 0.53
H2S 0.74 C2H4 0.22
COS 77ppmv C2Hg 0.30
CS2 2ppmv H2 0.43
HCN SOppmv H20 1.6
CO 0.17
Air (Stream No. 2): ?
6.2 Liquid
t Water make-up (Stream No. 3): ?
Chemical make-up (Stream No. 4): ?
t Water to centrifuge (Stream No. 5): ?
7.0 Intermediate Streams - None
8.0 Discharge Streams^ ' - Based on treatment of the feed gas reported in
Section 6.1.
8.1 Gaseous
Product Gas (Stream No. 6)
Vol % Vol %
0.52
0.22
0.29
0.42
4.32
Oxidizer vent gas (Stream No. 10): ?
B-126
co2
H2S
COS
cs2
HCN
CO
99.0
8 ppmv
75 ppmv
2 ppmv
0
0.16
CH4
C2H4
C2H6
H2
H20
-------
8.2 Liquid
Solvent blowdown (Stream No. 7)
Production Rate: 24.7 kg/106 Nm3(1.46 lbs/106 scf)
Composition
H20 80.0
Na2S203 10.8
NaSCN 4.4
NaV03 0.7
ADA 1.1
NaHC03+Na2C03 3.0
Separator effluent (Stream No. 8): ?
8.3 Solids
Sulfur (Stream No. 9): Nominally 99.5% sulfur with small
amounts of "contaminants" such as vanadium salts, sodium
thiocyanate.
9.0 Data Gaps and Limitations
Several limitations exist in Stretford operating data:
Lack of stream characterizations for most effluent streams, including
trace and minor constituents.
Actual operating data is limited from the many varied installations
where the Stretford process has been employed.
Lack of updated cost information for Stretford designs suitable for
treating high C02, low H2S gases.
10.0 Related Programs - The Synthane pilot plant at Bruceton, Pennsylvania,
incorporates a Stretford unit for H2S removal from concentrated acid
gases generated by a Benfield unit(6). Also, a commercial scale
Stretford unit is being installed at the SASOL S.A. Lurgi coal gasifi-
cation facility^. The operation of these units could generate useful
data on the applicability of the Stretford process to coal gasification.
B-127
-------
REFERENCES
1. Handbook of Gasifiers and Gas Treatment Systems, Dravo Corp., ERDA FE-1772-
11, February 1976.
2. Gas Processing Handbook, Hydrocarbon Processing, April 1975.
3. Ellwood, P., Meta-Vanadates Scrub Manufactured Gas, Chemical Engineering,
July 20, 1964.
4. Catalytic, Inc., The Stretford Process, unpublished work performed for
the EPA, Contract No. 68-02-2167.
5. Information supplied to TRW by A. J. Grant of Woodall-Duckham, December 5,
1977.
6. Haynes, W. P., Synthane Process Update, Mid 1977, 4th International Confer-
ence on Coal Gasification, Liquefaction and Conversion, August 2, 1977.
7. Atkins, T. W., Problems Associated with Controlling Sulfur Emissions from
High-Btu Coal Gasification Plants, C. F. Braun and Company report to ERDA,
under Contract No. E(49-18)-2240, December 1976.
B-128
-------
ZINC OXIDE ADSORPTION PROCESS
1.0 General Information
1.1 Operating Principles - Removal of hydrogen sulfide from a gas by
reaction with zinc oxide (ZnO) to form zinc sulfide. Zinc oxide
beds are used only for removing the last traces of sulfur after some
other treatment has removed all but a few parts per million. When
hydrogen is present in the gas, ZnO catalyzes the reduction of COS,
CS2 and organic sulfur to H2S which is adsorbed.
1.2 Development Status - Zinc oxide guard beds have been used as the
final sulfur cleanup to protect reforming catalysts since the mid-
19301 s. Their performance is proven in hundreds of commercial
plants.
1.3 Licensor/Developer^ ' - The technology of using hot zinc oxide
cleanup beds for hydrogen sulfide removal is widely known and
readily available from catalyst manufacturers, engineering design
and construction contractors, and private consultants. No dominant
patent exists.
Some suppliers of catalyst-grade zinc oxide are:
Catalysts and Chemical Inc., Louisville, Kentucky
Girdler Chemicals, Louisville, Kentucky
Harshaw Catalysts, Cleveland, Ohio
Katalco Corp., Oak Brook, Illinois
New Jersey Zinc Co., Bethlehem, Pennsylvania
1.4 Commercial Applications
Guard beds to protect reforming catalysts.
§ Guard beds to protect methanation catalysts.
Guard beds to protect other Raney nickel catalysts.
B-129
-------
2.0 Process Information
2.1 Flow Diagram (see Figure B-21) - The gas flow piping is designed
so that either vessel may be first in series or may be isolated for
catalyst dumping and re-loading. When the sour feed gas to the
zinc oxide guard beds is sufficiently pure, it may be possible to
operate only a single bed with removal and replacement of zinc
oxide during routine plant shutdown.
2.2 Equipment - One or more pressure vessels containing catalyst
granules.
2.3 Feed Stream Requirements
Composition: Total sulfur in all forms should be as low as can be
achieved with the upstream acid gas treatment step.
The zinc oxide beds are an expensive means of sulfur
removal; they should only be used to remove the final
traces which are beyond the capability of other
processes.
Water vapor content should be well below saturation.
Zinc oxide beds have been wrecked by an accidental
spill of liquid water from the upstream scrubber.
(1 2}
Temperature:v ' ' Both the reaction rates and sulfur loading
capacity improve with rising temperature (see
Figure B-22). Cold beds would have to be greatly
enlarged because of reduced sulfur loading capac-
ity. Furthermore, cold beds of zinc oxide would
have a negligibly slow rate of destruction of
carbonyl sulfide and thiophene. Therefore, in
practice most zinc oxide guard beds are operated
hot, 589°K to 722°K (600°F to 840°F).
Pressure: Pressure is not critical.
2.4 Operating Parameters
Temperature: See Section 2.3.
Pressure: As required for other steps of coal gasification process.
Catalyst Loading: The maximum sulfur loading capacity of zinc oxide,
as shown in Figure B-22. The maximum recommended
loading is only three percent when the desired
exit gas specification is 0.02 ppm H2s(2).
B-130
-------
CO
I
OJ
[X| OPEN VALVE
CLOSED VALVE
Legend:
1. Feed Gas Stream
2. Purified Gas Stream
3. Spent ZnO '
Figure B-21. Zinc Oxide Adsorption Process
12)
-------
O
UJ
25
X
20
15
X
O
u
N
Z
O
O
Z
3
O
10
ID
CO
273
373
473
573
673
773
TEMPERATURE, °K
Figure 8-22.
Sulfur Loading Capacity of Zinc Oxide as a
Function of Temperature^)
B-132
-------
Space Velocity: 2000 to 20,000 reciprocal hours^. Pressure
drop considerations and the rate of destruction
of the more refractory contaminants present in
the feed gas, particularly carbonyl sulfide, carbon
disulfide, and the thiophenes, are likely to deter-
mine space velocity for most coal-derived gases(5>.
2.5 Process Efficiency and Reliability - ZnO is a very effective adsor-
bent for H2S. The equilibrium concentration of hLS over ZnO can
be as low as 0.0005 ppm at 573°K (5730F)^5^. In practice, process
efficiency is determined by system design and operating conditions.
Zinc oxide absorption of sulfur is a thoroughly proven commercial
process.
2.6 Raw Material Requirements - In a two-bed system, the first bed
should be dumped just before breakthrough. This would typically
be at about 90 percent of maximum loading. Then the beds should
be reversed in sequence. In a single-bed system, the entire inven-
tory of zinc oxide is usually replaced annually.
2.7 Utility Requirements - Zinc oxide guard beds use no utilities.
3.0 Process Advantages
Proven process for sulfur removal.
Highest purity product gas of any sulfur guard process available.
Low capital cost.
Low operating cost when sour feed gas contains only residual traces
of sulfur.
Process is not pressure sensitive.
4.0 Process Limitations
Sulfur-rich gases cannot be economically treated with zinc oxide.
Hot operation is preferred.
Arsenic, halogens, and ammonia are not removed by the process.
Loading capacity is temperature dependent; see Figure B-22 for
detailed data.
. Liquid water will severely damage the process, with possible complete
degradation of zinc oxide bed.
B-133
-------
5.0 Process Economics
*
Costs depend on specific application. The capital investment and
operating cost of zinc oxide treatment are likely to be small items in
a coal gasification plant budget.
6.0 Input Streams
See Table B-30.
7.0 Discharge Streams
7.1 Product Gas - See Table B-30.
7.2 Spent Adsorbent - ZnO is typically discarded when sulfur content
reaches 15% - 20%. Quantity generated depends on specific appli-
cation. In smaller applications, spent ZnO is disposed of; for
larger applications it may be reclaimable (e.g., at zinc smelters.)
8.0 Data Gaps and Limitations
0 Process applicability to coal conversion process gas purification
system not studied/established.
Definition of maximum allowable concentrations of various contaminants
in the feed gas (e.g., trace metals, HCN, carbonaceous matter) has not
been determined.
9.0 Related Programs
No data available.
REFERENCES
1. Kohl, A. L. and Riesenfeld, F. C. Gas Purification, 2nd ed., Gulf
Publishing Co., Houston, Texas, 1974.
2. Dravo Corp., Handbook of Gasifiers and Gas Treatment Systems, ERDA
No. FE-1772-11, Pittsburgh, PA, 1976.
3. Institute of Gas Technology, Pipeline Gas from Coal-Hydrogenation
(IGT Hydrogasification Process) Project 9000 Quarterly Report No. 1,
July-September 1976, ERDA No. FE-2434-4, Chicago, Illinois, 1976.
4. Lee, B. S., Status of the HYGAS Program, 7th Synthetic Pipeline Gas
Symposium, Chicago, Illinois, October 27, 1976.
5. Katalco Corporation, Catalyst Handbook, Oak Brook, Illinois, 1970.
B-134
-------
TABLE B-30. TYPICAL PERFORMANCE DATA FOR THE SULFUR GUARn
AT THE HYGAS PILOT PLANT*(3,4) GUARD
Temperature, °K
Pressure, 106 Pa
Major Components, % (dry)
Hydrogen
Carbon Monoxide
Carbon Dioxide
Methane
Ethane
Total Sulfur, as H2S
Sulfur Compounds, ppmv
H2S
COS
cs2
CH3SH
CH3SCH3 and CH3CH2SH
Oxygen Compounds, ppmv
Methanol
Ketones
Acids
Aldehydes
Feed Gas
i^
___,
617
7.64
50.2
31.11
t
__t
__t
0.6ppm
0.53
0.02
0.00
0.04
0.03
0
0
0
0
T '
Product Gas
617
7.56
50
31
--t
--t
t
__t
0.003
0.045
0.00
0.002
0.000
280
7
15
15
*Include a caustic wash and ZnO guard bed.
fNot available.
B-135
-------
IRON OXIDE ADSORPTION PROCESS
1.0 General Information
1.1 Operating Principle - Removal of H^S from a gas stream by
adsorption on a fixed or fluidized bed of iron oxide ^Fe^O, + 6H,,S
-* 2Fe2S3 + 6hL). The bed is regenerated by treatment with air to
oxidize the chemisorbed sulfide to either elemental sulfur (low
temperature) or sulfur dioxide (high temperature). The chemisorp-
tion reaction is not pressure sensitive.
1.2 Development Status^ '''' - Use of the fixed bed for low tempera-
ture applications dates back to 1849 and the process is widely
used currently. Appleby-Frodingham Steel Co. at Appleby, England
operates a 673°K (752°F), 63.750 Nm3/D (2.5 x 106 SCFD) fluidized
bed coke oven gas treatment plant (plant operation started in
April 1956). Two other gas-works plants (at Nottingham and Exetar)
were also built. As of 1969, none of these plants were in opera-
tion. The fluidized-bed high temperature process is commerically
offered by Woodal1-Duckham but no U.S. plant has been constructed.
The Morgantown Energy Research Center (MERC/DOE) has been working
on the development of a high temperature process for application
to coal-derived synthesis gas since January 1974.
(3 5^
1.3 Licensors/Developers/Suppliersv '
Woodal1 Duckham (USA) Ltd.
200 Manor Oak One
1910 Cockran Road
Pittsburgh, Pa. 15220
Connelly-GPM, Inc.
200 South Second St.
Elizabeth, New Jersey 07206
Portable Treaters Co.
Box 3669
Odessa, Texas 79760
B-136
-------
1.4 Commercial Applications - Low-temperature fixed bed - natural-gas
sweetening. Fluidized bed - foreign, coke-oven gas and town gas;
3 plants have operated prior to 1969.
2.0 Process Information
(5)
2.1 Flow Diagramv ' - See Figure B-23 (for fixed bed design)
2.2 Equipment
Fixed Bed: Two adsorber vessels per train; blowers.
Fluidized Bed: Adsorber vessel, regenerator, conveyor, seal
leg, blowers, tar arrester, feed hopper, heat
exchangers.
2.3 Feed Stream Requirements^ '
Best suited for low inlet H^S (160 to 1200 ppmv) and small volumes;
for large treatment volumes and sulfur contents, another conven-
tional desulfurization process should precede this process.
t Feed stream should be low in dust and tar content to prevent bed
fouling. (Fluidized bed less sensitive to dust in the feed.)
Pressure: No specific requirements.
2.4 Operating Parameters
2.4.1 Adsorption Step^4'7^
Fixed Bed: See Table B-31
Temperature generally between 289°K - 316°K (60°F to 110°F)
in natural gas applications.
Pressure variable: (up to 7 MPa (1000 psig))
Space velocity: 7 to 35 hr
Fluidized
Temperature: 673°K (750°F) for one application to coke
oven gas
Pressure: No limitations (one coke oven application
uses near atmospheric pressure)
Space velocity: ?
B-137
-------
ro
i
oo
oo
LEGEND:
1. FEED GAS
2. STEAM
3. WATER INTERMITTENT)
4. AIR
5. PRODUCT GAS
6. REGENERATION OFFGAS
7. CONDENSATE
8. SPENT SORBENT
M CLOSED VALVE
[XI OPEN VALVE
-M-
cc
UJ
CO
cc
§
8
r
i
DC
LJJ
O
UJ
oc
1!
Figure B-23. Fixed Bed Iron Oxide Adsorption
(5)
-------
TABLE B-31. TYPICAL OPERATING CONDITIONS FOR IRON OXIDE GAS TREATMENT SYSTEMS
(4)
Parameter
Gas Volume, 106 Nm3/day
(106 SCFD)
H2S Content of Feed,
ppmv
Pressure, psia (MPa)
Temperature, °K (°F)
IN
OUT
Space Velocity, hr
Type of System
Conventional
Boxes
0.35 (6)
16,000
Atmos.
289 (60)
294 (70)
7.15
Deep
Boxes
0.25 (4.3)
11,800
Atmos.
296 (73)
307 (93)
6.66
High
Pressure
0.89 (15)
200
23 (340)
--
--
37.4
Tower
Design
1.4 (24)
8000 - 15,000
Atmos.
302 (85)
311 (100)
5.38
Continuous
Process
0.12 (2.0)
16,000
--
--
--
9.4
I
_J
OJ
-------
2.4.2 Regeneration Step
Fixed Bed - Batch Operation
Switch valves and depressurize to 1 atm
Circulate air through the bed; ambient temperature (below
220°K or 120°F)
Monitor offgas for On content
Fluidized Bed
Continuous withdrawal of hot solids from adsorber and
conveying at controlled rate to regenerator (roaster)
vessel
Temperature in the regenerator: ?
Pressure: atmospheric
2.4.3 Sulfur Recovery
In low temperature applications, sulfur accumulated in the
bed as the result of regeneration may be recovered by
steam, hot gas or solvent treatment. The spent iron oxide
containing sulfur can also be discarded directly.
In high temperature applications, the roasting of the bed
during regeneration releases SOp which may be directly
discharged to the atmosphere, converted to elemental sulfur
or sulfuric acid.
2.5 Process Efficiency and Reliability^4'5^
2.5.1 Natural Gas Sweetening Application (low temperature, fixed bed)
0.3 to 4.0 ppmv H2S in output (96% removal); depends on
space velocity
20 ppmv (0.15 grain/100 scf) CHgHS (89% removal)
Effective HCN removal
C02 and organic sulfur not removed
t Reliable and effective, small scale applications; control
or moisture and surface "alkalinity" required for best
efficiency (NaoC03 solution and water or steam may be
injected into the bed for control)
B-140
-------
2.5.2 Coke Oven Gas Treatment (high temperature-fluidized bed)
No data available.
2.6 Raw Material Requirements
t Unmixed oxide type (fixed bed): Prepared from iron ore, contains
75% ferric oxide, 10% water of hydration, 15% inert impurities;
or prepared from red mud (bauxite purification residue), up to
50% ferric oxide; or natural bog ores, hydrated ferric oxide
plus fibrous and peaty material, 45% FLO.
Mixed oxide type (fixed bed): Prepared from wood shavings or
granulated slag used to support pulverized iron oxide; this
material is commonly referred to as iron sponge.
Maker-up iron oxide: Quantity depends on specific application
and mode of operation (fixed vs. fluidized bed and throwaway vs.
bed regeneration).
^2^3 conditioner solution: Quantity not known.
Air for bed regeneration: ?
Hot gas or solvent for sulfur recovery: ?
2.7 Utility Requirements
Steam for sulfur recovery: ?
Steam/water for bed conditioning: ?
Electricity: ? (power for blower is small)
2.8 Miscellaneous - Operational Safety-Low Temperature Fixed Bed -
Eventual bed replacement of fixed beds is required in the iron sponge
process. Vessels must be designed to minimize difficulties in
replacement. Change-out of the beds is hazardous. Exposure to air
when dumping a bed can lead to an exotherm (finely divided oxidizable
materials). Spontaneous combustion can result. Care must be used in
opening the tower and admitting air. The entire bed should be wetted
before beginning the change-out operation.
3.0 Process Advantages
Low temperature fixed bed iron sponge system is a well known widely
applied technology.
Regenerable low cost sorbent. w
B-141
-------
Good removal efficiency (competitive with wet purification processes)
in low space velocity, small volume, low l^S content applications.
Byproduct sulfur can be produced (may be uneconomic on small scale).
HCN can be removed.
Low temperature absorption/regeneration has relatively low utility
requirements.
4.0 Process Limitations
Space requirements: Relatively large for fixed beds (due to low
space velocity); economic disadvantage compared to continuous wet
processes.
Requirement for the disposal of spent bed and the hazards of handling
spent iron oxide (Section 2.8). For trace sulfur removal not as
effective as ZnO guard.
t Best suited for sweetening small volumes of gas with low h^S contents
(low temperature batch process).
C02 not removed.
Cold weather gas-hydrate formation (iron sponge process).
Process effectiveness has not been demonstrated for very high tempera-
ture applications (such as for low Btu fuel gas cleanup).
Process efficiency is affected by the moisture content of the feed
gas (unlike ZnO guard).
In cold weather, moisture can condense on bed, and in some cases,
reduce H,>S removal efficiency.
5.0 Process Economics
Duckworth and Geddes in 1965 reported comparative costs of iron sponge
and MEA treatment for natural gas sweetening applications as follows^ .
Inlet H2S; 112 ppmv (7 grains/100 cu ft)
Inlet mercaptan: 20 ppmv (1.3 grains/100 cu ft)
MEA process capital investment - $270,000
Iron sponge capital investment - $110,000
MEA process direct operating expense - $28,000 (no depreciation)
Iron sponge direct operating expense - $23,000 (no depreciation)
B-142
-------
6.0 Input Streams
6,1 Feed Gas - See Table B-31 for some typical applications.
6.2 Air - ?
6.3 Steam - ?
6.4 Make-up Iron Oxide - Quantity varies with application.
6.5 N^CO-j Conditioning Solution - ?
7.0 Process/Discharge Streams
7.1 Gaseous
Product Gas: Natural gas 0.3 to 4.0 ppm H0S, 2 ppmv or less
CH3HS(5). 2
Product Gas: Coke oven gas from fluidized bed hot process, down
to 270 ppmv and 300 ppmv total sulfur*2'.
Offgas from hot fixed bed regeneration: Contains S02-
7.2 Solid
t Spent bed containing iron oxide, sulfur, and inert carriers from
fixed bed process. Rate of production not known.
Fines from attrition of iron ore, fluidized process, 1 kg/300 Nm
(1 Ib per 5000 scf of gas processed)\2).
7.3 Liquid
Generally no liquid stream, except when a solvent (such as
ammonium sulfide or carbon disulfide) is used for sulfur
recovery.
8.0 Data Gaps and Limitations
Limited data on input and discharge streams characteristics for high
temperature applications.
No data on quantities and characteristics of spent adsorbent.
Applicability of process to gases containing a relatively high concen-
tration of hydrogen is not known.
B-143
-------
9.0 Related Programs
Morgantown Energy Research Center (MERC/DOE) has an on-going pro-
gram for the testing of the iron oxide adsorbent process for applica-
tion to hot low Btu gas desulfurization. The work to date has been
with the development of adsorbent support and bench-scale testing with
coal-derived gas.
REFERENCES
1. Morgantown Energy Research Center, Quarterly Report, April-June 1977,
p. 12.
2. Reeve, L., Desulphurization of Coke Oven Gas at Appleby-Frodingham,
J. Inst. Fuel, July 1958, p. 319-324.
3. Grant, A. J., Applications of the Woodall Duckham 2-Stage Coal Gasifica-
tion presented to 3rd International Conference on Coal Conversion. What
Needs to be Done Now, Pittsburgh, PA, August 3, 1976.
4. Kohl, A. and Riesenfeld, F., Gas Purification, Gulf Publishing Co.,
Houston, Texas, 1974.
5. Dravo Corp., Handbook of Gasifiers and Gas Treatment Systems, Report
FE-1772-11, February 1976, pp. 154-56.
6. Duckworth, G. C. and Geddes, J. H., Oil and Gas Journal 63, September
13, 1965, p. 94.
7. Maddox, R. N., Gas and Liquid Sweetening, Campbell & Co., 1974, p. 182.
B-144
-------
METAL OXIDE IMPREGNATED CARBON PROCESS
1.0 General Information
(1 2}
1.1 Operating Principles''^' - Adsorption of trace quantities of sulfur
species (mainly H2S) from a gas using activated carbon impregnated
with metal oxides* (e.g., copper, zinc, chromium). During subse-
quent regeneration with steam or air, the metallic sulfide is
restored to its original oxide form and elemental sulfur is
produced.
M + H2S -> MS + H20 adsorption
MS + 1/2 02 -» MO + S regeneration
Periodic removal of sulfur from bed is accomplished using steam or
hot inert gas, or using solvents (e.g., carbon di sulfide or ammonium
sulfide).
1.2 Development Status^ ' - Commercially available; used for industrial
gas desulfurization. Sixty plants in the U.S. Original industrial
development by I.G. Farben Industries in 1920-1929, based on chemical
warfare applications in 1915-1918.
1.3 Suppliers - Several companies, including:
Calgon Corp., Box 1346, Pittsburgh, PA 15230
Girdler Chemical Inc., Box 337, Louisville, KY 40201
Barnebey-Cheney Inc., Box 2526, Columbus, OH 43216
(non impregnated) activated carbon has been used for HzS removal. In
this application air is added to the feed gas and the carbon catalyzes the
oxidation of H2S to elemental sulfur which becomes trapped on the carbon
surface. Tests have indicated that carbon has removed greater than 20% of
its weight of H2S. The deposited sulfur can be removed by solvents or heat
transfer.
B-145
-------
1.4 Commercial Applications - Natural gas sweetening and synthesis gas
desulfurizatio.n.
2.0 Process Information
2.1 Flow Diagram (see Figure B-24) - The two-tower arrangement shown in
Figure B-24 allows for operation of one tower in the adsorption
mode while the adsorbent in the second tower is being regenerated.
(See Section 1.1 for process description),
2.2 Equipment - Adsorbers, air blower, coolers, steam supply.
2.3 Feed Stream Requirements
Any pressure
0 Temperature: ? Volatilization of elemental sulfur (produced
during regeneration) may determine upper temperature limitation.
0 Best applicability to gas with £30 ppm H?S.
Low content of higher hydrocarbons, ammonia, and tar required.
2.4 Operating Parameters
2.4.1 Adsorption Step
Any pressure
Temperature: ?
Inlet I^S 30 ppm or lower
Downward vertical gas flow
Space velocity: 350 to 400 hr'1
2.4.2 Regeneration
t At or near 1 atm.
Steam at 447°K to 533°K (400°F to 500°F). Bed preheat to
450°K (350°F) followed by 4 - 6 hours steaming.
Air addition to oxidize sulfides.
Upward vertical steam/air flow (opposite to previous
flow of feed gas).
B-146
-------
DCh-
LjjO.
T
I
L
LEGEND:
1. FEED GAS T
2. STEAM/AIR *
3. PRODUCT GAS
4. REGENERATION VENT GAS
5. STEAM OR SOLVENT FOR REGENERATION
6. PRODUCT SULFUR/CONDENSATE/SOLVENT
7, SPENT ADSORBENT
8. REGENERATION CONDENSATE
CLOSED VALVE
OPEN VALVE
Figure B-24. Metal Oxide Impregnated Carbon Adsorption
(1)
B-147
-------
2.4.3 Sulfur Removal
t Sulfur removal from bed necessary when sulfur build-up
reaches 13 to 25 weight %.
Sulfur removal methods include purging with steam or hot
jas (570°K) or extraction with carbon disulfide or aqueous
gas (b/o
(NH4)2S.
Sulfur impregnated carbon may also be directly discarded
without sulfur recovery.
2.5 Process Efficiency^ ' - Typically 99 percent or better removal before
breakthrough; natural gas of 30 ppm l-^S can be purified to 0.2 ppm
or less.
2.6 Raw Material and Utility Requirements
Metal oxide impregnated activated carbon: quantity required
depends on sulfur concentration in feed gas and nature of metal
oxide/carbon used.
Steam for regeneration: 2.8 to 5.9 kg/m (50 to 100 Ibs per
cubic foot) of carbon.
Electric power for blowers: ? (relatively small)
Cooling water: ? (relatively small)
Air for regeneration: ?
3.0 Process Advantages
High removal efficiency.
Recovery of pure sulfur by-product is possible.
Can be designed to operate at any convenient feed stream pressure.
4.0 Process Limitations
Treatment/disposal of H2S-laden regenerant stream.
Applicable only to low H^S content feed streams; otherwise too
frequent regeneration is required.
Reagent deactivated by contaminants (e.g., ammonia, tars and polymers).
Disposal of the vent gas (depressurizing preceding regeneration) stream
and product sulfur is required.
B-148
-------
Auxiliary facility/equipment for sulfur/solvent recovery necessary.
Regeneration steam may contain mercaptans, COS and other sulfur com-
pounds. Treatment of this stream may be necessary before venting.
5.0 Process Economics
As with most adsorption processes the cost of activated carbon process is
a function of the volume of gas treated, the temperature, and the chemical
composition. No data available on the cost of this process for acid gas
treatment.
6.0 Input Streams
6.1 Feed Gas (Stream 1) - No operating data available.
6.2 Steam/Air (Stream 2) - See Section 2.6.
7.0 Process/Discharge Streams
7.1 Product Gas (Stream 3) - No actual data available; total sulfur level
of less than 0.2 ppm can be obtained.
7.2 Regeneration Vent Gas (Stream 4) - No data available.
7.3 Sulfur/Condensate (Stream 6) - No data available.
7.4 Regeneration Condensate or Solvent (Stream 8) - No data available.
(Condensate formed in bed during steam regeneration operation.
Condensate may contain organics.)
7.5 Spent Adsorbent (Stream 7) - No data available.
8.0 Data Gaps and Limitations
Actual operating data for process (specifically on characteristics of
discharge streams not known) not available.
9.0 Related Programs
No serious proposal to actually use carbon in SNG manufacture was found.
Other desulfurization processes appear to be preferable. Data presented
reflect experience in other industries, possibly not relevant to the
SNG situation.
B-149
-------
REFERENCES
1. Dravo Corporation, Handbook of Gasifiers and Gas Treatment Systems, ERDA
Report FE 1772-11, February 1976, pp. 151-153.
2. Lovett, W. D. and Cunniff, F. T., Air Pollution Control by Activated
Carbon, Chemical Engineering Progress, Vol. 70, No. 5, May 1974.
B-150
-------
ORGANICS REMOVAL FROM GASES USING ACTIVATED CARBON
1.0 General Information
1.1 Operating Principles - Removal of hydrocarbons and other organics
(particularly the odor-producing compounds) by adsorption on a solid
bed of activated carbon ("adsorbent"). The molecules of impurities
("adsorbate") adhere to (adsorb on) the surface of the adsorbent
by cohesion and/or chemical reaction. The extent and nature of
adsorption depend on the properties of the particular carbon,
adsorbate and the medium from which adsorption takes place. In
practice, adsorption systems are designed to provide adequate con-
tact between adsorbate and adsorbent and under conditions (e.g.,
temperature) most favorable for adsorption. The spent carbon is
generally regenerated by physical (e.g., application of heat) and/or
chemical treatment, with the adsorbed material often recovered in
the form of a concentrated stream. In practice, adsorption systems
are generally used for the removal of residual pollutants in a gas
stream after bulk of such pollutants are removed by more conventional
types of gas treatment.
1.2 Development Status - Commercially available.
1.3 Licensor/Developer - Many companies offer carbon adsorption systems
for gas treatment. 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^2'3^ - The removal of organics from a gas
stream to eliminate odor, control air pollution, or recover valuable
products (e.g., benzene). Potential applications of adsorption sys-
tems in a coal gasification plant may be in connection with the
control of emissions from lockhoppers, Claus plant and regeneration
or process catalysts.
B-151
-------
03
I
on
ro
LEGEND:
1. FEED GAS 4. REGENERATION VENT GAS
2. STEAM 5 SPENT ADSORBENT
3. PRODUCT GAS
(X3 OPEN VALVE
CLOSED VALVE
Figure B-25. Activated Carbon Adsorption
-------
2.0 Process Information
2.1 Flow Diagram (see Figure B-25) - The two-tower arrangement shown in
Figure B-25 allows for operation of one tower in the adsorption
mode while the adsorbent in the second tower is being regenerated.
(See Section 1.1 for process description.)
2.2 Equipment - Conventional adsorption vessel, adsorbent, and heating
system for adsorbent regeneration.
2.3 Feed Stream Requirements
Temperature: Varies with the specific application and the
carbon used.
Pressure: Varies with the specific design. Generally higher
pressures favor adsorption.
Gas Composition: Dependent upon the adsorbent used; however, high
concentrations of the contaminants improve
adsorption efficiency.
2.4 Operating Parameters
2.4.1 Adsorption Step
t Temperature: Generally <366°K (<200°F) for most appli-
cations(2)
Pressure: 0.1 to 0.6 MPa (atmospheric to 90 psigr -
Loading (space velocity/contact time): Depends on the
type of carbon used and the specific design; in one
foundry application, a space velocity of 150 min-1 was
used(4).
2.4.2 Regeneration Step
Temperature: Generally in the 366°K to 811°K (200°F to
1000°F) range; low molecular weight adsorbed organics are
generally driven off at lower temperatures, whereas high
temperature regeneration is primarily aimed at in situ
destruction (cracking) of higher molecular weight organics.
Vacuum can also be applied to the system to reduce the
necessary regeneration temperature.
B-153
-------
2.5 Process Efficiency and Reliability - Depending on the type of
carbon, carbon loading and the material to be adsorbed, removal
efficiencies of up to 99% can be attained. In general, the adsorp-
tion efficiency is higher for larger molecular weight and unsaturated
and aromatic compounds. Carbon adsorption has been in use for a
number of years and has proven successful in removing a wide variety
of organic contaminants from gas streams.
2.6 Raw Material Requirements
Carbon makeup/replacement: Depends upon the material adsorbed
(adsorbate), regeneration method and conditions, operating
temperature, pre-filtration and treatment of gas stream.
For an activated carbon system used to recover an organic solvent,
typical carbon makeup requirements are reported as follows:
0.2 to 0.4 kg of carbon per tonne of solvent recovered
(0.5 to 1.0 lb/ton)(2).
2.7 Utility Requirements
Vary with the nature of application; typical values for a carbon
adsorption system recovering an organic solvent are as follows(^):
- Electricity: 4.5 to 6.75 kwh/100 kg of solvent recovered
(0.10 to 0.15 kwh/lb recovered)
- Cooling water: 60 to 85 I/kg recovered (70 to 10 gal/lb
recovered)
- Steam: 1 kg/kg of solvent recovered (1 Ib/lb)
3.0 Process Advantages
Activated carbon is highly efficient in the removal of trace quantities
of organics, particularly for some of the most objectionable odiferous
compounds.
Purified gas from an adsorption unit is generally suited for discharge
directly to atmosphere.
t A product stream suitable for by-product recovery or use as fuel can
be produced.
4.0 Process Limitations
Adsorbent requires regeneration, normally by heat.
Some adsorption capacity is lost during each regeneration cycle.
B-154
-------
0 For most applications involving use of activated carbon, the inlet qas
temperature must be kept below 366°K (200°F),
t Rapid pressure changes can physically disturb the catalyst bed and the
flow regime through the bed, with the resultant reduction in adsorption
efficiency.
Participate matter can clog the sorbent surface, resulting in increased
pressure drop and deterioration of carbon activity.
0 In some applications, pretreatment of the gas stream may be necessary
to extend the bed life and reduce regeneration requirements.
0 Carbons used for gas treatment are generally more expensive than those
used in water pollution control.
0 Depending on the nature of the gas stream handled, spent carbon regener-
ation and ultimate disposal may present special hazards due to the
hazardous nature of the adsorbed organics.
5.0 Process Economics
0 Cost of system varies as a function of gas flow rate and composition,
regeneration method and removal efficiency.
6.0 Input Streams
0 Gas composition is dependent upon the source of the waste gas.
0 Make-up carbon: Carbon for gas treatment applications are generally
manufactured from coconut shells, fruit pits and couhene and babassu
nut shells; the make-up quantity depends on nature, effectiveness and
frequency of regeneration.
7.0 Discharge Streams
0 Purified gas stream (Stream 2): Composition depends on application.
0 Regeneration off-gas (Stream 3): Can contain a high concentration of
the contaminants which were adsorbed.
0 Spent carbon: Depending on the material adsorbed or the nature of
regeneration, may contain potentially hazardous organics and trace
elements. The quantity and characteristics of spent carbon vary
with the application. Spent carbon may be incinerated (or gasified
in a gasification plant) to recover fuel value, or discarded as solid
waste.
B-155
-------
8.0 Data Gaps and Limitations
Extensive performance data are available for the various types of
adsorption systems and media as applied to a variety of industrial gas
cleaning operations. Evaluation of expected performance of adsorption
systems in application to coal gasification plant waste gas streams
requires data on detailed characterization of the gas to be treated.
Such data include gas temperature, chemical characteristics including
a trace element survey, and are generally either unavailable or are
incomplete.
9.0 Related Programs^
A carbon adsorption system is featured at the Synthane pilot plant at
Bruceton, Pennsylvania, for the removal of heavy organics from product
gas after acid gas treatment in a Benfield unit (Benfield process does
not remove organics). When operational, the Synthane integrated operation
should provide useful information about carbon performance in gasification
service.
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. LeDuc, M. F., Adsorption Equipment, Air Pollution Engineering Manual,
2nd Edition (AP-40), U.S.E.P.A., May 1973.
4. Lovett, W. D. and Cunniff, F. T., Air Pollution Control by Activated
Carbon, Chemical Engineering Progress, Vol. 70, No. 5, May 1974, p. 43.
5. Anon, A History of FGD Systems Since 1850, Journal of the Air Pollution
Control Association, Vol. 27, No. 10, October 1977.
6. Haynes, W. P., et. al., Synthane Process Update, Mid-1977, 4th International
Conference on Coal Gasification, Liquefaction and Conversion, Pittsburgh,
Pennsylvania, August 2-4, 1977.
B-156
-------
MOLECULAR SIEVES PROCESS
1.0 General Information
1.1 Operating Principles - Physical adsorption of polar and small
molecules (H20, H2S, mercaptans) from a gas on a fixed bed of porous
synthetic Na/K/Ca aluminosilicate (zeolite) granules. H20 is more
strongly adsorbed than any other component. Molecular sieves con-
sist of a geometric network of cavities connected by pores. Pores
are of molecular dimensions, 0.25 to 12.7 A (1 x 10"9 to 5 x 10"8
inch). Regeneration is performed using a hot gas.
M ?}
1.2 Development Statusv ' ' - Natural zeolites have been used in gas
dehydration since the 18th century; synthetic zeolites (cracking
catalysts) have been used since 1940. Linde molecular sieves have
been commercially offered since 1954. Extensively used worldwide
for dehydration, H^S removal, etc. (over 150 units for desulfurization),
1.3 Commercial Applications
Mercaptan removal from natural gas
Cryogenic C^Hg and helium extraction from natural gas
Air separation plants (recovery of argon)
LNG production
t Annealing-oven inert gas purification
Ethylene purification
2.0 Process Information
2.1 Flow Diagram (see Figure B-26) - The two unit arrangement shown in
Figure B-26 allows for operation of one unit in the adsorption mode
while the molecular sieves in the second unit are regenerated (see
Section 1.1 for process description).
B-157
-------
cn
co
ADSORPTION
l-CO I M
M "El
* I ^^ u~J
^ ' M i M
M-
-M-
-M-
-M-
LEGEND:
1. FEED GAS
2. PRODUCT GAS
3. REGENERATION GAS
4. CONDENSATE
5. REGENERATION OFFwGAS
6. SPENT SIEVES
HEATER
REGENERATION
(HEATING-COOLING)
COOLER
CLOSED VALVE
OPEN VALVE
cr
O
cc.
<
O.
Figure B-26. Molecular Sieve Process for Sulfur Removal from Fuel Gases
(1)
-------
2.2 Equipment - Adsorbers, heater, cooler, knockout drum, sulfur
recovery train. Number of beds in system depends on inlet H?S
concentration.
2.3 Feed Stream Requirements^ '
Any pressure or C0? content (C02-to-H2S mole ratio of up to 1050
can be tolerated).
Temperature: 289°K to 322°K (60°F to 120°F).
Best applicability to dried gases, since large beds would be
required for H2S removal from wet gases.
Total sulfur content up to 10,000 ppm.
t Feed should be very low in organics (e.g., glycol carryover from
a prior dehydration step).
2.4 Operating Parameters^4'6'8)
2.4.1 Adsorption Step
Any pressure
Temperature: 289°K to 322°K (60°F to 120°F) (each 5.6°C
[10°F] increase reduces bed capacity by 20 percent) -
see Table B-32.
Bed density: typically 673 kg/m3 (42 Ib/cu ft).
t Space velocity: ?
2.4.2 Regeneration Step
Uses a hot gas (commonly the product gas): gas tempera-
ture 561 °K to 489°K (550°F to 600°F).
Can run at adsorption pressure but works best at low
pressure (0.1 MPa [1 atm].
Flow in direction reverse to that of adsorption.
(4 7)
2.5 Process Efficiency^ ' '
H2S 4 ppm or less (down to 0.3 ppm) in product.
t CH3SH removed more efficiently than H2S; 95 percent CH3SH removal
is typical.
B-159
-------
TABLE B-32. SIEVE PERFORMANCE ON H2$ IN BINARY SYSTEM (H2$ AND DILUENT)
Temperature
H2S Partial
Pressure
Amount Absorbed,
kg H9S/kg Sieve
298°K (77°F)
348°K (167°F)
423°K (302°F)
(1 mm Hg) 0.13 KPa
(10) 13.2 KPa
(100) 132 KPa
(I) 0.13 KPa
(10) 13.2 KPa
(100) 132 KPa
(1) 0.13 KPa
(10) 13/2 KPa
(100) 132 KPa
298°K (77°F)
(1) 0.13 KPa
(200) 264 KPa
4.4*
10.6
15.7
2.4
5.6
10.6
0.4
2.2
5.2
*At incipient
breakthrough
(UCC type 5A pellets)
5.0
17.0
(UCC type 13X pellets)
UCC = Union Carbide Corp.
Not outstanding on C02 (non-polar); but C0? applications in
natural gas are common.
C2H5SH not removed effectively due to pore size limitation.
H2S removal efficiency depends on inlet feed stream HoO and H?S
levels.
2.6 Raw Material and Utility Requirements
Heat for regeneration: ?
Cooling water: ?
B-160
-------
' appl&s'f "^ °"Ce 1n 3 t0 5 ^ars Ww-fc on the specific
3.0 Process Advantages
Relatively low operating cost.
Selectivity of H2S and CH3$H over CCL,
Relatively safe and maintenance free.
0 Long sorbent life.
Typically smaller size units than carbon beds,
Works well over a wide range of pressure, H2S or CCL level.
Specific pore sizes marketed to suit specific applications.
4.0 Process Limitations
Narrow operating temperature range (optimum temperature 289°K to 322°K
(60°F to 120°F).
0 Sieves are expensive.
0 Strong affinity for H20 may make sieves uneconomical for sulfur removal
when the gas stream contains large amounts of water.
0 The contaminanted regenerant gas may require treatment prior
to disposal.
0 When CO? is present, sieves can act as catalyst for COS formation
(C02 + HzS + COS + HgO). The COS is only weakly adsorbed and quickly
appears in the product. The equilibrium which is normally unfavorable
to COS is upset by the strong retention of product H20 by sieves.
0 Disposal of vent gas, and complex valving, when regeneration pressure
not equal to adsorption pressure. (Pressure swing type operation.)
0 Presence of large quantities of C02 has a detrimental effect on the
sweetening process.
0 Complete removal of water (production of bone dry gas) may not be
desirable in certain applications.
5.0 Process Economics
0 In applications to natural gas,sweetening heater and flared regenera-
tion gas (combined) typically consume 1 to 2 percent of daily
throughput.
B-161
-------
t In a one-sieve application to natural gas, a capital cost of $1 million
was reported for a 2,5 x 106 Nm3/D (100 x 10° scfd) plant. The natural
gas treated contained 6% C02 and 3.5 to 4,7 g/100 Nmd (15 - 20 grains/
100 scf) sulfur. Approximately 54 tonnes (60 tons) of sieves were
used^6'. This cost compares to $1,5 million for an equivalent amine
treating plant(^).
6.0 Input Streams
No data available for applications to H2S removal from gases; the charac-
teristic of the input gas (Stream 1) to a 5.1 x 106 Nrn3/D (2 x 108 scfd)
(9)
natural gas mercaptan removal plant are as followsv ':
Feed pressure: 5.2 MPa (750 psig)
Feed total sulfur concentration: 470 mg/100 Nm (2 grains/100 scf)
(32 ppm H2S equivalent)
HLO, 1.2 kg/100 Nm3 (7 lb/1000 scf) or less
Trace of glycol, glycol degradation products and oil
Zero H2S and C02
Butane, 300 ppm volume
t Balance is ChL, No. ethane, propane
7.0 Process/Discharge Streams
7.1 Product Gas (Stream 2): No data available for applications to H2S
removal from gases (for the mercaptan removal system described in
Section 6.0, 97% total sulfur removal has been reported).
7.2 Purge Gas/Condensate (Streams 4 and 5): ?
7.3 Spent Sieves (Stream 6): ?
8.0 Data Gaps and Limitations
Very limited data available for molecular sieves application to H2S
removal, specifically on the characteristics of discharge streams.
9.0 Related Programs
Sieves have not been seriously proposed for use in SNG plants. In one
instance, sieves were proposed for treatment of C09 vent gas from Benfield
(2}
C02 removal in a Bi-Gas plantv '. The alternative Rectisol process was
deemed preferable to the Benfield and sieve system.
B-162
-------
REFERENCES
1. Dravo Corporation, Report FE-1772-11, Handbook of Gasifiers and Gas
Treatment Systems, pp. 157-59, February 1976.
2. Kohl and Riesenfeld. Gas Purification, 2nd Edition, Gulf Publishing,
1974, pp. 563-575.
3. Jahnig, C. E., Evaluation of Pollution Control in Fossil Fuel Conversion
Processes. Gasification; Section 5,Bi-Gas Process, EPA Report 650/2-75-
009g, NTIS No. PB 293-694, p. 11.
4. Chi, C. W. and Lee, H., Natural Gas Purification by 5A Molecular Sieves
and its Design Method, AIChE Symposium Series No, 134, Vol. 69,
pp. 95-101, 1973.
5. Maddox, R. W., Gas and Liquid Sweetening, Campbell and Company,
pp. 188-200, 1974.
6. Thomas, T. L. and Clark, E. L., Proceedings of the 46th Annual NGPA
Meeting, no date.
7. Harris, T. B., Natural Gas Treating with Molecular Sieves, Pipeline and
Gas Journal, June-August 1972.
8. Lee, N. N. Y. and Collins, J. J., Ammonia Plant Feed Desulfurization with
Molecular Sieves, Presentation to Tripartite AIChE Meeting, Montreal,
Canada, September 25, 1968.
9. Conviser, S. A., Oil and Gas Journal 63, pp. 130-135, December 6, 1965.
B-163
-------
APPENDIX C
GAS UPGRADING OPERATION
Shift Conversion Module
Cobalt Molybdate Process
Methanation and Drying Module
Fixed-Bed Methanation
Fluidized-Bed Methanation
Liquid Phase Methanation/Shift
C-l
-------
COBALT MOLYBDATE PROCESS
1.0 General Information
1.1 Operating Principles - The water-gas shift reaction is the con-
version of steam and carbon monoxide to make hydrogen and carbon
dioxide (CO + H20 = H2 + C02). It is a reversible reaction,
easily promoted by a variety of catalysts. Since it is mildly
exothermic, the equilibrium tends toward more hydrogen at lower
temperatures. The equilibrium is independent of pressure.
1.2 Development Status - Shift conversion for hydrogen production has
been practiced on large scale commercially since about 1930. The
classical high-temperature catalyst, still the most popular, is
iron oxide. In sour gas service iron catalysts suffer some severe
handicaps. Therefore, in 1970 Badische Anil in and Soda Fabrik
announced a cobalt molybdate catalyst specifically designed for
shift conversion of gases contaminated by large percentages of
hydrogen sulfide. Since then several other catalyst manufacturers
have offered cobalt molybdate catalysts for sour gas shift.*
Combined shift and methanation processes are also under develop-
ment, using nickel based catalysts (see data sheets for fixed-
and fluidized-bed methanation).
1.3 Licensor/Developer - The fact that the sulfide derivative of
cobalt molybdate is an efficient shift conversion catalyst has been
known for so long that it is improbable that any dominant patent
now exists. Catalyst manufacturers have performed the development
Although catalysts other than Co-Mo based may be used in shift conversion,
the appropriate temperature range applicability and resistance to sulfur
poisoning make Co-Mo catalysts most attractive for SNG production. This
data sheet will thus be restricted to such catalysts.
C-2
-------
work and are, in general, reluctant to reveal details of operation
except under secrecy agreement with a bonafide catalyst purchaser.
A partial list of catalyst manufacturers who offer cobalt-moly
sulfide catalysts for sour gas shift purposes is as follows:
BASF Wyandotte, Parsippany, New Jersey
Catalysts and Chemicals Inc., Louisville, Kentucky
Hal dor Topsoe, Houston, Texas
Katalco Corp., Oak Brook, Illinois
1.4 Commercial Applications - At least two sour gas shift units (Co-Mo
type) are operating today on commercial scale, but the identity of
one of them is confidential. The coal gasification demonstration
at Westfield, Scotland, between May and September 1974, used a sour
gas shift process provided by LurgiO). The catalyst of that
demonstration was cobalt-moly(2).
2.0 Process Information
2.1 Flow Diagram - See Figure C-l.
2.2 Equipment - The basic equipment consists of an adiabatic fixed-bed
high temperature reactor, feed stream heaters, product stream
coolers, and a condensate separation vessel. Some of the heat
exchangers are recuperative. The reaction vessel for a one-third
demonstration size plant (2.27 x 106 m3 per day) was a horizontal
cylinder with hemispherical heads, 3.20 meters i.d. by 14.63 meters
t.t. It contained two beds of catalyst; the smaller fore-bed could
be by-passed when clogged^3'.
2.3 Feed Stream Requirements
Composition(4): The ratio of steam to dry gas must be at least
0.7. The sulfur content of the dry gas feed must be at least
10 ppm. There is no maximum limit to sulfur content. During
startup the sulfur content of the feed gas should be far higher
than the minimum for normal operation. H2S may be deliberately
injected during startup to convert the cobalt molybdate to the
active sulfided form.
C-3
-------
START-UP HEATER
SHIFT REACTOR
1. Sour Feed Gas
2. Shifted Gas
3. Condensate
CONDENSATE
K.O. POT
FEED-EFFLUENT
EXCHANGER
4. H.P. Steam
5. Foul Water
6. Spent Catalyst
Figure C-l. Sour Gas Shift Conversion at the Hygas Demonstration Plant(3)
-------
Temperature(4): The feed gas to the reactor should be between
533°K (500°F) and 575°K (578°F) during normal operation with fresh
catalyst of high activity.
Pressure: The pressure of shift conversion can be at any con-
venient level determined by other steps in the coal gasification
process. However, at 533°K (500°F) the feed gas should not exceed
11.39 MPa (1790 psia) because the water dew point may be exceeded.
2.4 Operating Parameters^)
Temperature: Normal operation is between 533°K (500°F) and
728°K (851°F). As catalysts age, the operator should gradually
raise the temperature to compensate for loss of activity. Start-
up is rather elaborate, to avoid the two hazards of carbonyl
formation and runaway methanation. The catalyst vendor should
dictate the startup program.
Pressure: The operating pressure can be set for the convenience
of other steps in the coal gasification process.
Space Velocity: ?
2.5 Process Efficiency and Reliability - At least three commercial-
scale coal gasification plants have successfully operated with
cobalt-moly sour gas shift catalysts (see Section 1.2 above).
The catalyst manufacturers who have independently developed this
catalyst are all reliable companies with adequate experience in
gas processing.
2.6 Raw Material Requirements - Catalyst cobalt molybdate (see
Section 1.3). Useful life varies with catalyst, feed composition,
and operating conditions.
2.7 Utilities - High-pressure steam is the only utility consumed in
significant quantity by the shift reaction. The process designer
can trade off capital investment versus steam consumption to
achieve an optimum economic balance. Maximum heat conservation
for the process shown in Figure C-l has lead to a design steam
C-5
-------
demand of 0.419 kilograms per standard cubic meter of raw gas fed
to the shift converter^).
3.0 Process Advantages
The sour gas shift can salvage some of the enthalpy of the hot raw
gas from the gasifier. The raw gas need only be partially quenched
to a degree where tar and dust can be excluded.
The sour gas shift uses less steam than the classical high-temperature
shift.
Part of the steam requirement of the sour gas shift can be furnished
by sparging feed gas through a pool of foul water. This serves as a
disposal outlet for the foul water generated elsewhere.
Cost savings are realized when acid gas removal is located entirely
downstream of the sour gas shift. If the conventional shift were
used, the acid gas removal system would have to be split into two
scrubbers. Absorption of the hydrogen sulfide would have to be up-
stream of the conventional shift; that portion of the carbon dioxide
generated by shift conversion would have to be scrubbed out
downstream.
The cobalt-moly catalyst is active for destructive hydrogenation of
carbonyl sulfide when relatively clean gas is processed.
The sour gas shift catalyst can be regenerated easily. The con-
ventional iron-oxide shift catalyst cannot be regenerated.
4.0 Process Limitations
Reducing and sulfiding the catalyst is a delicate operation at initial
startup. Subsequent cold startups are routine.
Spent catalyst may present handling and disposal problems.
5.0 Process Economics
The capital cost estimate for the shift conversion section of the
HYGAS demonstration plant was $31,000,000(4). The demonstration
plant size was 7.08 million cubic meters per day. The whole demon-
stration plant, on the same basis, was estimated at $681,000,000.
Therefore, the shift conversion unit represents about four and one-
half percent of plant investment.
Steam consumption depends on process design and degree of heat
recuperation pursued.
Catalyst service life is not yet predictable. There is inadequate
industrial experience with this catalyst in coal gasification
applications.
C-6
-------
6.0 Input Streams (Figure C-l)
Feed Gas (Stream 1) - See Tables C-l and C-2*.
High Pressure Steam (Stream 4) - No data available.
Foul Water (Stream 5) - No data available.
7.0 Discharge Streams (Figure C-l)
Shifted Gas (Stream 2) - See Tables C-l and C-2*.
Condensate (Stream 3) - No data available.
Spent Catalyst (Stream 6) - See Sections 1.3 and 2.6.
8.0 Data Gaps and Limitations
The full details of sour gas shift conversion using proprietary cobalt
molybdate catalysts are known only to the catalyst manufacturers and
those catalyst purchasers who have signed secrecy agreements. No data
are currently available regarding the composition of condensate
generated by shift conversion units.
Although bench scale experiments with simulated gas mixtures indicate
that shift catalyst may be active for the hydrolysis of COS and C$2,
reported experience with "dirty" coal gases has not shown that these
sulfur compounds are converted to ^2$ Data to suPPort these
findings for "dirty" gases are not currently publicly available.
9.0 Related Programs
The operation of the shift conversion units at the HYGAS pilot plant is
expected to provide additional data on the performance of the process
in SNG applications.
'Data found in Tables C-l and C-2 are from ench p ot n t or s oper
ing times and small gas volumes. The catalyst used was BASF K8 1H^. mese
data are presented since they indicate catalyst act vity for both
reaction and for hydrogenation of carbonyl sulfide and propylene.
C-7
-------
TABLE C-l. BENCH SCALE SHIFT CONVERSION REACTOR FEED AND PRODUCT GAS
COMPOSITION, 627°K(5)
Feed Gas Rate, g-mol/hr (Ib-mol/hr)
Steam-to-Gas Ratio
Gas Analysis:
Feed Gas Composition, mol %
Carbon Monoxide
Carbon Dioxide
Hydrogen
Methane
Ethane
Propane
Ethylene
Propylene
Carbonyl Sulfide
Hydrogen Sulfide
Sulfur Dioxide
Nitrogen
Water
Total
Reactor Temperature, °K (°F)
Reactor Pressure, MPa (psig)
Steam Temperature, °K (°F)
Product Gas Rate, g-mol/hr
(Ib-mol/hr)
Product Gas Composition, mol %
Carbon Monoxide
Carbon Dioxide
Hydrogen
Methane
Ethane
Propane
Ethylene
Propylene
Acetylene
Carbonyl Sulfide
Hydrogen Sulfide
Sulfur Dioxide
Nitrogen
Water
Total
Space Velocity, hr-1
COS Converted, mg-mol/hr (Ib-mol/hr)
% COS Conversion
Propylene Hydrogenated, mg-mol/hr
(Ib-mol/hr)
% Conversion
5.2 (1
Dry
8.7
6.9
31.2
49.1
1.3
0.04
0.01
0.18
0.15
1.00
0.08
1.3
0.00
99.96
627
6.9
656
5.7 (1
Dry
0. 7
14.3
37.3
43.5
1.2
0.13
0.01
0.04
0.01
0.00
1.00
0.04
1.80
0.00
100.03
2174
7.7 (1.7
7.3 (1.6
.1470 x 10"2)
0.78
-
Wet
OiD
4.12
18.63
29.32
0.78
0.02
0.01
o.n
0.09
0.60
0.06
0.78
40.28
99.99
0 (670°)
(1025)
0 (720°)
.26 x 10"^)
Wet
0.41
8.38
21.86
25.49
0.70
0.08
0.00
0.02
0.00
0.00
0.59
0.02
1.05
41.40
100.00
, 3860
x 10"5)
100 ,
x TO'5)
75.6
*
Bench pilot unit operation with propylene added to feed gas to
investigate hydrogenation of unsaturated hydrocarbons during shift
reaction.
C-8
-------
TABLE C-2. BENCH SCALE SHIFT CONVERSION REACTOR FEED AND PRODUCT GAS
COMPOSITION, 438°K(6)* UHi
Feed Gas Rate, g-mol/hr (Ib-mol/hr)
Steam- to- Gas Ratio
Gas Analysis:
Feed Gas Composition, mol %
Carbon Monoxide
Carbon Dioxide
Hydrogen
Methane
Ethane
Propylene
Butene
Hydrogen Sulfide
Carbonyl Sulfide
Nitrogen
Water
Total
Reactor Temperature, °K (°F)
Reactor Pressure, MPa (psig)
Steam Temperature, °K (°F)
Product Gas Rate, g-mol/hr
(Ib-mol/hr)
Product Gas Composition, mol %
Carbon Monoxide
Carbon Dioxide
Hydrogen
Methane
Ethane
Propylene
Butene
Hydrogen Sulfide
Carbonyl Sulfide
Nitrogen
Water
, -1
Space Velocity, hr
COS Converted, mg-mol/yr (Ib-mol/hr)
% COS Conversion
% Propylene Conversion
4.7
Dry.
21.4
17.8
36.8
20.3
0.37
o.io
0.13
2.0
0.20
0
0
100.00
4
Dry
20.0
17.3
38.5
19.6
0.35
0.11
0.14
2.40
0.20
0
:
(1.04 x 10"2)
0.70
Wet
12.6
10.5
21.7
12.0
0.22
0.06
0.08
1.18
0.12
0
41.0
SOO.OO
438 (330)
1.36 (200)
473 (390) ?
.54 (1.05 x 10"^)
Wet
11.0
9.5
21.1
10.7
0.19
0.06
0.08
1 .32
o.n
45 2
2000
t_ijuu
0
n
U
^SO-minute operation of bench pilot unit
C-9
-------
REFERENCES
1. Hebden, D., and Brooks, C. 1., Westfield--The Development of Processes
for the Production of SNG from Coal, Communication 988 at the 113th Annual
General Meeting of the Institution of Gas Engineers, Edinburgh, 1976.
2. Sudbury, J. E., 0. R. Bowden, et al, A Demonstration of the Slagging
Gasifier Process, Proceedings of Eighth Syntehtic Pipeline Gas Sympos-
ium, A.G.A. et al., Chicago, Illinois, 1976, pp. 483-496.
3. Institute of Gas Technoloay, HYGAS: 1964 to 1972 Pipeline Gas From Coal-
Hydrogenation (IGT Hydrogasification Process) Part VIII: Commercial Plant
Design, ERDA Number FE-381-T9-P4, Chicago, Illinois, 1975.
4. Detman, R. Factored Estimates for Western Coal Commercial Concepts, ERDA
Number FE-2240-5, C. F. Braun & Co., Alhambra, California, 1976.
5. Environmental Assessment of the HYGAS Process, Monthly Report (March 1 to
March 31, 1977) prepared by Institute of Gas Technology, for ERDA, ERDA
Number FE-2433-11.
6. Environmental Assessment of the HYGAS Process, Monthly Report (Sept. 1-30,
1977) prepared by Institute of Gas Technology, for ERDA, ERDA Number
FE-2433-19.
7. Private communication with Don Fleming of the Institute of Gas Technology,
May 31, 1978.
C-10
-------
FIXED-BED METHANATION PROCESS
1.0 General Information
1.1 Operating Principle - Carbon oxides and hydrogen are reacted to
produce methane and water over a fixed bed or surface of Raney
nickel catalyst. The methanation reactions are:
CO + 3H2 = CH4 + H20
C02 + 4H2 = CH4 + 2H20
1.2 Development Status - The reaction itself has a long history of
commercial utility. It is used, for example, in the purification
of ammonia feed gas by extinction of residual carbon monoxide. As
applied to high Btu gas production from coal, fixed-bed metha-
nation generally requires a recycle of product methane to control
heat generation and to dehydrate the reaction atmosphere. Several
different arrangements of accomplishing this recycle have been
tested. One version was demonstrated at Westfield, Scotland, on
commercial scale between May and September, 1974^.3) ^ Another
version was demonstrated on pilot plant scale in Chicago between
April 1973 and November 1976 as a part of IGT's HYGAS demon-
stration program^6'?*8*9'10'11). No facilities currently produce
high Btu gas via methanation on a commercial scale. Combined
shift and methanation has also been tested at the pilot plant
1.3 Licensor/Developer - The technology of fixed-bed methanation is
widely known and readily available from catalyst manufacturers,
engineering design and construction contractors, and private
consultants. No dominant patent exists.
C-ll
-------
1.4 Commercial Applications - Commercial applications have been for the
purification of hydrogen for synthesis of ammonia and other chemi-
cals, and as a final step in manufacture of synthetic natural gas.
2.0 Process Information
2.1 Flow Diagram - See Figure C-2 for a generalized fixed-bed metha-
nation process. Depending on the feed gas composition and the
method of extracting heat from the system, feed gas may either be
fed to the first-stage methanator (1) undiluted, (2) diluted with
recycle gas or (3) diluted with steam. A second-stage methanator
is commonly employed for final methanation followed by cooling/
heat recovery and condensation of moisture.
2.2 Equipment - The equipment consists of steel pressure vessels to
hold the catalyst beds, heat exchangers, the recycle gas com-
pressor, and a condensate flash drum. Equipment sizes are listed
on Table C-3 for three alternate designs of the methanation step
of the HYGAS process.
2.3 Feed Stream Requirements
(13)
Compositionv ': The ratio of hydrogen to carbon oxides should be
slightly greater than stoichiometric* (e.g., 3 moles hydrogen per
mole carbon monoxide and 4 moles hydrogen per mole carbon dioxide).
Residual excess hydrogen at completion of the reaction should be
in the range of 2% to 10%.
The feed stream to an adiabatic reactor may be no richer than
about 4 percent carbon monoxide. This means the feed gas from a
coal conversion process must be diluted by recycle of product gas
and flows carefully controlled. Moisture can be tolerated in feed
up to the saturation level at about 328°K.
The feed gas must be essentially free of sulfur in all forms.
Catalysts differ in their tolerance for sulfur, but a reasonable
specification might be 0.1 ppm measured as H2S.
*
Except where combined shift and methanation are to be effected.
C-12
-------
co
cc
0
<
li. CO
£. "J ^
83*
w r*- in
I COOLING
J AND/OR
HEAT
RECOVERY
KNOCKOUT
DRUM
LEGEND:
1. FEED GAS
2. STEAM (NOT NECESSARY
IN ALL DESIGNS)
3. RECYCLE GAS
4. METHANATED PRODUCT GAS
5. CONDENSATE
6. SPENT CATALYST (PERIODIC
REPLACEMENT)
Figure C-2. Schematic Flow Diagram for Typical Fixed-Bed Methanatiorr ' '
-------
TABLE C-3. FIXED-BED METHANATION - TYPICAL EQUIPMENT SIZES (HYGAS VERSION)
General Scale
Status
Reference
Volume of Product Gas,
cubic meters per day
Reactor Dimensions,
dia./ht. (meters)
Number 1
Number 2
Number 3
Number 4
Recycle Compressor
displacement, cu m/sec
kilowatts (steam drive)
Waste Heat Steam Generator(s)
kg/hr
Feed-Product Exchanger
sq m
Total Surface, All Other
Coolers, sq m
Condensate Flash Drum
dia. /length (meters)
Pilot Plant
Built & Running
(4)
9.7 x 103
0.61/2.95
0.61/4.47
None
None
0.026
30
None
One-third of
Demonstration
Plant Site
Conceptual
(5)
2.26 x 106
3.35/1.22
3.35/2.44
3.35/4.88
3.35/9.14
1.50
(1200)
73400
44.2
2.13/4.57
One-fourth of
Demonstration
Plant Site
Conceptual
(12)
1.94 x 106
3.96
4.11
4.27
None
(785)
198200
685
280
1.83/3.96
C-14
-------
The non-reagent portion of the feed gas should be almost entirely
methane. Nitrogen or other inert diluent should be as low as
feasible.
Temperature/ ': Ideal operating temperature is about 670°K
(750°F). Feed temperatures of 600°K (600°F) are desirable to
allow for generated heat to be used in making superheated steam.
Nickel carbide (and carbonyl) can form from catalyst and carbon
monoxide at temperatures below about 623°K (665°F).
Carbon monoxide can disproportionate to carbon dioxide and
elemental carbon at low temperatures, leading to catalyst
deactivation.
Pressure^ ' ': Since the methanation reactions are favored at
high pressures, the pressure level of fixed-bed methanation is
preferred as high as possible. Most coal gasification process
designs have contemplated pressures in the range of 3.2 MPa
(480 psia) to 8.3 MPa (1200 psia).
2.4 Operating Parameters
Temperature^ ': Typical feed temperatures are about 600°K (600°F).
Exit temperatures are kept below about 760°K (900°F) to inhibit
carbon formation on catalyst.
Pressure: 3.2 MPa to 8.3 MPa (480 to 1220 psia).
(14)
Space Velocity: Fixed-bed HYGAS pilot plant methanator^
320 - 4000 hrs'1; tube wall C02 Acceptor pilot plant methanator^ '-
38,100 - 39,500 hrs"1.
2.5 Process Efficiency and Reliability - Hundreds of fixed-bed units
have operated commercially, mainly for applications such as puri-
fication of ammonia synthesis gas. In such applications essen-
tially complete methanation of carbon monoxide can be obtained.
Commercial SNG production via methanation has yet to have any
widespread application. Available data indicate that fixed-bed
reactors can produce gas containing greater than 90%
C-15
-------
methane and less than 1000 ppmv carbon monoxide. Methanation of
C02 does not occur until essentially all CO has reacted04).
The reliability of the process is influenced by a number of
factors. Sulfur, chlorine and metals can poison methanation
catalysts. Careful control of bed temperatures is necessary to
inhibit carbon formation, catalyst deactivation and nickel carbonyl
formati on.
2.6 Raw Material Requirements^ '' - Catalyst life is normally
scheduled for 2 years. Longer service periods, up to 10 years,
have been achieved by careful operation. Catalyst properties and
composition vary with the source and are usually proprietary
(nickel content varies from 25 to 70 percent).
2.7 Utilities
Steam: There is a net surplus of high-pressure steam generated by
the exotherm of the methanation reaction. The amount of steam
available for export varies widely with process design and equip-
ment selection. Steam may also be used as a diluent to feed gas.
Such steam is partially converted to hydrogen via the water gas
shift reaction, with subsequent methanation.
Cooling Water: No data available; quantities depend on specific
design.
Electric Power: The recycle gas compressor is usually driven by a
steam turbine. There are no other significant power consumers in
this process, except possibly the fans of finned-tube air coolers.
No actual operating data are available.
2.8 Distribution/Formation of Trace Constituents During Methanation -
Unsaturated hydrocarbons and alcohols are completely hydrogenated
over nickel catalyst^). Nitric oxides and ammonia are com-
pletely converted to elemental nitrogen, while HCN is only
partially destroyed. Sulfur compounds are converted to H2S and
subsequently react with and deactivate the catalyst. Chlorine in
feed also deactivates nickel catalysts.
C-16
-------
Nickel carbonyl (NKCOty can form in methanators by reaction of
CO with the catalysttU). This reaction 1s favored at ^ ^^
tures and is not especially important in the steady state operating
range of 600°K-760°K (620°F-900°F). Transient operations, however,
can lead to carbonyl formation.
Iron carbonyl can form by reaction of CO at low temperatures with
iron in carbon steel piping. If carried into the methanation
reactor, iron and carbon can be deposited onto the catalyst, caus-
ing deactivation. Stainless steel piping upstream of methanators
can eliminate this problem^5).
3.0 Process Advantages
3.1 Fixed-Bed Methanation
Fixed-bed methanation can produce pipeline-quality high-Btu gas.
Fixed-bed methanation is a commercially proven process.
There is very little pressure drop from feed gas to product gas
through a properly designed fixed-bed methanation system.
There is little shift conversion in fixed-bed methanation;
carbon dioxide is not generated in large amounts. On the
contrary, any carbon dioxide in the feed gas will tend to be
destroyed by hydrogenation to methane if sufficient hydrogen
is present in the feed.
3.2 Fixed-Bed Methanation/Shi ft
Elimination of a recycle system and large mass and heat flow
rates
Elimination of a separate shift conversion section
t Operation at conditions removed from regions of carbon formation
Production of more steam at higher pressures with less surface
areas
Reduction in catalyst sensitivity to sulfur
t C02 removal from a reduced volume of gas
C-17
-------
4.0 Process Limitations
Essentially no sulfur can be tolerated in the feed.
Cold startup can be complex and requires skilled operators. Hazards
such as Ni(CO)4 production require special precautions.
The reagent gases are thermodynamically unstable outside certain
boundary conditions and will decompose to carbon. Methane cracks
when too hoc or with insufficient residual hydrogen. Carbon
monoxide disproportionates when too cold. The methanation reactors
must be operated within well defined limits of temperature and
reagents' partial pressures. Fortunately, these limits are known
precisely.
Methanation/shift has not been demonstrated on a commercial scale.
5.0 Process Economics
The 1974 capital cost estimate for the methanation section of the steam-
oxygen version of a HY6AS demonstration plant was $29,000,00o(12).
Demonstration plant size is 7.08 million cubic meters per day. The whole
demonstration plant, on the same basis, was estimated at $681,700,000.
Therefore methanation constitutes 4.25% of coal gasification capital
cost. (Presumably this figure includes a small amount for the ante-
cendent zinc oxide guard chamber.)
Operating costs for methanation are not known. Generally, the value of
steam produced is expected to exceed the value of other utilities
consumed.
6.0 Input Streams
6.1 Feed Gas (Stream 1) (see Tables C-4, C-5, C-6, C-7) - The data are
for gases produced by the HYGAS process, the C02-Acceptor process,
a test gas containing essentially no methane, and a simulated
Lurgi product gas, respectively. The former two gases have been
treated for C02 removal and have a ratio of H2 to CO of about 4.
The third contains -about 25 percent C02, and has a ratio of H2 to
CO of about 1:4. The fourth gas has been treated for C02 removal
but has not been shifted prior to methanation.
C-18
-------
RECYCLE
Methanation
Recycle Ratio
Space Velocity (hr'1)
Catalyst Used
Stream
Flow Rate
Nm3/hr (scf/hr)
Composition (dry)
H,
CO
CH4
C2H6
co2
N2
H2S
COS
RSH
1st Stage
.
^ ^___ _____ _
3.3
4200
Harshaw pelleted nickel
Feed Gas
1660 (28100)
vol %
51.6
12.7
23.6
1.4
0
10.4
ppmv
.003
.045
.002
~
2nd Stage
1.0
2800
on Kieselguhr
Product Gas
1020 (17260)
vol %
15.5
0
67.4
0
0
17.1
ppmv
Condensate produced o-i n?7)
kg/hr (Ibs/hr)
*Run #37, 7/1/75 (1300 hrs) to 7/2/75 (0600 hrs).
C-19
-------
TABLE C-5. PERFORMANCE DATA FOR THE C02-ACCEPTOR PILOT PLANT PACKED TUBE
METHANATION REACTOR*06)
Recycle Ratio
Space Velocity (hr )
Catalyst Used
Stream
Flow Rate
Nm3/hr (scf/hr)
Temperature, °K (°F)
Composition (dry)
H2
CO
CH4
co2
N2
Moisture in Gas,
kg/hr (Ibs/hr)
1.24
38,100
N/A
Feed Gas
1415 (23900)
373 (206)
62.8
15.5
13.8
4-2
3.7
100
53 (177)
'
39,500
N/A
Product Gas
549 (9300)
723 (840)
2.5
0
88.6
0
8.9
100
162 (356)
Downstream coolant used to remove reactor heat.
C-20
-------
USING STEAM*0)
Gas Composition (dry)
CO
CH4
co2
N2
Moisture (as percent
of dry gas volume)
FOR METHANATION/S
,
Feed Gas
. ,
42.91
31.14
0.08
24.66
1.21
100
67.3
HI FT OF SYNTHESIS GAS
Product Gas
5.83
0.34
29.13
62.70
2.00
100
118
3 adiabatic stages employed; steam used to dilute feed gas,
high nickel (60% NiO) co-precipitated formula catalyst.
Table C-7. PERFORMANCE OF THE RM PROCESS FOR COMBINED SHIFT/METHANATION*
(19)
Temperature, °K
Inlet
Outlet
Composition
(mole %)
H2
CO
C00
2
CH4
N9+Ar
c.
Moisture (as
percent of
dry gas volume
at inlet)
Feed
Gas
--
--
58.9
24.3
0.1
14.7
1.6
_
.,
Methanation Reactor Product
755 (900)
1060 (1469)
48.7
13.6
6.4
23.6
--
24
_.
755 (900)
1000 (1342)
35.1
7.5
7.5
28.5
""
24
~
588 (600)
800 (128)
18.1
1.8
7.6
34.2
33
..
588 (600)
750 (890)
6.8
0.2
5.9
37.4
IT T
51
=====:
i i v»/"if
Four fixed-bed reactors in
in series operating at different temperatures.
C-21
-------
6.2 Steam (Stream 2) - see Tables C-6 and C-7 for quantities used in
small scale tests.
7.0 Process/Discharge Streams
7.1 Recycle Feed (Stream 3) - See Tables C-4 and C-5 for recycle ratios
which have been employed. Composition of recycle is same as
product gas (Stream 4).
7.2 Methanated Product Gas (Stream 4)(see Tables C-4, C-5, C-6, C-7) -
Note that CH4 to H2 ratios in product gases in Tables C-4 and C-6
are similar. More complete methanation is represented by the data
in Table C-5.
7.3 Condensate (Stream 5) - Quantities produced (or potential) are
indicated in Tables C-4, C-5, C-6 and C-7. Condensate will contain
dissolved gases which can be released upon depressurization,
although quantities are expected to be small. Only traces of
hydrocarbons (other than methane), sulfur and nitrogen compounds,
and suspended material (e.g., catalyst fines) are expected. No
actual operating data are available.
7.4 Spent Catalyst (Stream 6) - See Table C-8 for properties of spent
catalyst used in the HYGAS pilot plant methanator. Spent catalyst
may be disposed of (1) directly as landfill, (2) returned to
catalyst vendor for reclamation of nickel value, (3) used as sulfur
guard bed adsorbent. In the latter application nickel is almost
as active as zinc for trace sulfur removal, and a spent methanation
catalyst still has a considerable capacity for sulfur.
8.0 Data Gaps and Limitations
Data gaps and limitations relate primarily to the properties and compo-
sition of spent methanation catalyst(s) and process condensate(s). The
compositions of product gases during transient or unsteady state
operation are also not well documented (e.g., presence of NI(CO)»).
9.0 Related Programs
No programs are known to be underway, or planned for specifically obtain-
ing environmental data on methanation operations. On-going pilot scale
C-22
-------
TABLE C-8. SPENT HARSHAW Ni-OlQ4-T-l/4 CATAI V.T
PLANT METHANATOR)U8J ' CALYST ANALYSIS (USED
i
IN HYGAS PILOT
Sulfur, %
Carbon, %
Nickel, %
Surface Area,
m2/g
Total Pore
Volume, cc/g
Ni Crystallite
Size, A
methanation is being conducted at DOE-sponsored pilot coal gasification
plants (HYGAS, Synthane, BIGAS). Data generated as part of this work
could fill some of the data gaps.
REFERENCES
1. Woodward, Colin, Catalyst Available for High Temperature Methanation,
Hydrocarbon Processing, January 1977, p. 136.
2. Landers, James E., Review of Methanation Demonstration at Westfield,
Scotland, Proceedings of Sixth Synthetic Pipeline Gas Symposium, A.G.A.,
et al, Chicago, Illinois, 1974, pp.297-304.
3. Hebden, D., and Brooks, C. T.t WestfieldThe Development of Processes
for the Production of SNG from Coal, Communication 988 at the 113th Annual
General Meeting of the Institution of Gas Engineers, Edinburgh, 1976.
4. Institute of Gas Technology, HYGAS: 1964 to 1973 Pipeline Gas from Coal-
Hydrogenation (IGT Hydrogasification Process) Part III: Pilot Plant
Development, ERDA Number FE-381-T9-P2, Chicago, Illinois, 19/b.
5. Institute of Gas Technology, HYGAS: 1964 to 1973 Pipeline Gas from Coal-
Hydrogenation (IGT Hydrogasifi cation Process Part VIII: Commercial Plant
Design, ERDA Number FE-381-T9-P4, Chicago, Illinois, 19/b.
C-23
-------
6. Lee, B. A., Status of HYGAS ProcessOperating Results, Proceedings of
Fifth Synthetic Pipeline Gas Symposium, A.G.A. et al, Chicago, Illinois,
1973, pp.5-17.
7. Lee, B. A., Status of HYGAS Program, Proceedings of Seventh Synthetic
Pipeline Gas Symposium, A.G.A. et al, Chicago, Illinois, 1974, pp.313-355.
8. Lee, B. A., Current Development of the HYGAS Program, Proceedings of
Eighth Synthetic Pipeline Gas Symposium, A.G.A. et al, Chicago, Illinois,
1976, pp.13-32.
9. Institute of Gas Technology, Pipeline Gas from Coal-Hydrogenation (IGT
Hydrogasification Process) Project 8907 Final Report, August 1972-June 1976,
ERDA Number FE-1221-145, Chicago, Illinois, 1976.
10. Institute of Gas Technology, Pipeline Gas from Coal-Hydrogenation (IGT
Hydrogasification Process) Project 9000 Quarterly Report No. 1, July-
September 1976. ERDA Number FE-2434-4, Chicago, Illinois, 1976.
11. Institute of Gas Technology, Pipeline Gas from Coal-Hydrogenation (IGT
Hydrogasification Process) Project 9000 Quarterly Report No. 2, October-
December 1976. ERDA Number FE-2434-8, Chicago, Illinois, 1976.
12. Detman, R., Factored Estimates for Western Coal Commercial Concepts,
ERDA Number FE-2240-5, C. F. Braun & Co., Alhambra, California, 1976.
13. Seglin, L., Methanation of Synthesis Gas, American Chemical Society,
Advances in Chemistry Series 146, Washington, D. C., 1976.
14. Cameron Engineers, Synthetic Fuels Quarterly, Vol. 13, No. 1, March 1976,
pp.4-7 to 4-13.
15. Mueller, F. W., Methanation of Coal Gas for SNG, Hydrocarbon Processing,
April 1974.
16. McCoy, D. C., The C02 Acceptor Process Pilot Plant 1976, Eighth Synthetic
Pipeline Gas Symposium, Chicago, Illinois, October 18-20, 1976, p. 33.
17. Allen, D. W., and Yen, W. H., Methanator Design and Operation, Chemical
Engineering Progress, Vol. 69, No. 1, January 1973.
18. Leppin, D., Operating Experience with the IGT Cold-Gas Recycle Methanation
Process in the HYGAS Pilot Plant, Ninth Synthetic Pipeline Gas Symposium,
Chicago, Illinois, October 31-November 2, 1977.
19. Chow, T. K. et al, The RM Process, A Methanation System, Ninth Synthetic
Pipeline Gas Symposium, Chicago, Illinois, October 31-November 2, 1977.
C-24
-------
in
FLUIDIZED-BED METHANATION PROCESS
1.0 General Information
1.1 Operating Principle - Methanation of carbon oxides and hydrogen
a bed of nickel-based catalyst particles fluidized by feed gas.
1.2 Development Status^1'2' - Fluidized-bed methanation has been tested
at the PEDU level under sponsorship of DOE and Bituminous Coal
Research (BCR). Early tests were aimed at determining catalyst
suitability and attrition and heat transfer characteristics. Dur-
ing 1976 and 1977, PEDU methanation operations at the BIGAS Homer
City, Pa, pilot plant were conducted to evaluate the most promising
catalyst under a wide variety of conditions. As of November 1977,
over 30 days of operating experience with the catalyst have been
logged. This experience is expected to be useful in further
operation and scale-up of the pilot-plant methanator.
Thyssengas GmbH of Duisburg, West Germany has recently tested a
pilot fluidized-bed reactor for combined shift and methanation^5).
Initial results indicate that 75 to 100 percent of equilibrium
methane and carbon dioxide formation can be obtained in the reactor
with acceptable catalyst attrition losses.
1 3 Licensor/Developer - Bituminous Coal Research, Inc
350 Hochberg Road
Monroeville, PA 15146
1.4 Commercial Applications - Thyssengas GmbH
Duisburg, West Germany
2.0 Process Information
2.1 Flow Diagram - Figure C-3 is a diagram of the pilot plant methana-
tor system of the BIGAS process in Monroeville, Pennsylvania.
C-25
-------
Product Gas
Blewback Nitrogen
Catalyit Filters
Coolant
Finned Cooling Tubes
Intermediate Feed Gas_j
Coolant
Disengaging Zone
Reaction Zone
Gas Distribution Zone
I
Figure C-3. Pilot Plant Fluidized-Bed Methanator
C-26
-------
2.2 Equipment - Figure C-3 shows the 0.1524 meter pilot methanator
Note the elaborate internal cooling coils. Equipment outside the
reactor is conventional.
2.3 Feed Stream Requirements
Composition(3): The ratio of hydrogen to carbon oxides should be
slightly greater than stoichiometric. The stoichiometry requires
three moles hydrogen per mole carbon monoxide and four moles
hydrogen per mole carbon dioxide. Residual excess hydrogen at
completion of the reaction should be in the range of 2 to 10 per-
cent. Moisture can be tolerated up to the saturation level at
328°K (130°F).
The feed gas must be quite free of sulfur in all forms. Catalysts
differ in their tolerance for sulfur, but a reasonable specifica-
tion might be 0.1 ppm measured as H^S.
The non-reagent portion of the feed gas should be almost entirely
methane. Nitrogen or other inert diluent should be as low as
feasible.
Temperature^ ^: Ideal operating temperature is about 670°K(750°).
Feed temperatures of 600°K (620°F) are desirable to allow for
generated heat to be used in making superheated steam. Nickel
carbide (and carbonyl) can form from catalyst and carbon monoxide
at temperatures below about 623°K (655°F).
Carbon monoxide can disproportionate to carbon dioxide and elemen-
tal carbon at low temperatures, leading to catalyst deactivation.
Pressure^: Since the methanation reactions are favored at high
pressures, the pressure level of fixed-bed methanation is pre-
ferred as high as possible. Most coal gasification process designs
have contemplated pressures in the range of 3.2 MPa (480 psia) to
8.3 MPa (1200 psia).
C-27
-------
(1 2)
2.4 Operating Parametersv ' '
Temperature: Typical feed temperatures are about 600°K (620°F).
Exit temperatures are kept below about 760°K (900°F) to inhibit
carbon formation on catalyst via methane decomposition.
Pressure: 3.2 MPa to 8.3 MPa (680 to 1200 psia)
Space Velocity: PEDU runs have been in the range of 1500 to 3000
reciprocal hours. However, it may be necessary to go to much
longer residence times or to multiple-staged beds to achieve pipe-
line quality product gas.
2.5 Process Efficiency and Reliability - Fluidized-bed methanation is
still in an early stage of development. Little information is
available regarding efficiency and reliability.
2.6 Raw Material Requirements^ ' - The only raw materials inherently
required for methanation are the feed gas and the catalyst make-
up. Harshaw catalyst (nickel, copper, molybdenum on alumina
support) has been used at the Bigas pilot plant. Catalyst losses
in the runs conducted so far appear to be in the normal range for
fluidized-bed operations.
2.7 Utilities - It is too early in the development of fluidized-bed
methanation to predict utilities requirements. There should be
a substantial credit for surplus high-pressure steam generated in
waste heat boilers.
3.0 Process Advantages^ '
No recycle of product gas is required for a fluidized-bed methanator.
Heat transfer from a fluid bed is excellent; there is little danger
of an uncontrolled exotherm.
t Cold startup should present minimal problems. The internal cooling
coils in the fluidized-bed methanator can be used as a startup heater.
4.0 Process Limitations"'2'
In common with alternate modes of methanation, the fluidized bed
suffers from catalyst sensitivity to sulfur and from the potential
risk of reagent instability and decomposition to carbon.
C-28
-------
Fluid!zed bed reactors suffer from back-mixing and short-
circuiting^}. _ In the case of methanation reactors, back-mixinq
causes a relatively high concentration of water vapor to be present
ab initio at the injection point of the CO-rich feed gas This
causes the water gas shift reaction to take place. Indeed BCR's
PEDU reactor turned out to be a much better shift reactor than a
methanator.
Having the shift reaction occur in a methanator need not be a
serious liability if the resulting C02 can later be methanated.
However, it is a well-known phenomenon that C02 hydrogenation is
inhibited by the presence of CO (at levels greater than about
200 ppmv)(3).
t The short-circuiting characteristic of fluidized beds makes it vir-
tually impossible to get the carbon monoxide concentration of dry
product gas down to the pipeline specification (0.1 percent). At
least it cannot be done in a single stage; perhaps multiple stages
of reaction could overcome this difficulty.
5.0 Process Economics
No dependable cost figures are available for fluidized-bed methanation
because no credible large-scale plant design has yet been published.
6.0 Input Streams
Feed Gas - See Table C-9.
7.0 Discharge Streams
t Product Gas - See Table C-9.
Scent Catalyst - No composition data available. Attrition of catalyst
appears to be a minor problem after the first few hours of methanation
operation" '.
8.0 Data Gaps and Limitations
Fluidized-bed methanation needs considerable development. Attention
should be focused on overcoming the inherent disadvantages of fluidi-
zation operations, i.e., back-mixing and short-circuiting.
9.0 Related Programs
No programs are currently underway to assess environmental problems
associated with fluidized-bed methanation.
C-29
-------
TABLE C-9. TYPICAL PERFORMANCE DATA FOR THE FLUIDIZED-BED METHANATION PLANT
OPERATION AT THE BIGAS FACILITY, HOMER CITY, PA*(4)
Temperature, °K (°F)
Flow Rate, kg-moles/hr
Composition (vol %)
H2
CO
co2
CH4
C2H6
Water Condensed, kg-moles/hr
Feed Gas
673° (752°)
2.434
59.4
19.4
.082
20.2
0.4
--
Product Gas
700° (800°)
1.416
30.8
1.0
7.2
59.8
1.2
0.254
Period 17, Run 22 (Oct. 7, 1976),
REFERENCES
1.
2.
3.
4.
5.
Streeter, R. C., Recent Developments in Fluidized Bed Methanation Research,
Ninth Synthetic Pipeline Gas Symposium, October 3-November 2, 1977.
Streeter, R. C., D. A. Anderson, et al, Status of the BIGAS Program, Part
II - Evaluation of Fluidized-Bed Methanation Catalysts, Proceedings of
Eighth Synthetic Pipeline Gas Symposium, A.G.A. et al, Chicago, Illinois,
1976, pp. 95-127.
Seglin, L., Methanation of Synthesis Gas, American Chemical Society,
Advances in Chemistry Series 146, Washington, D. C., 1975.
Bituminous Coal Research, Inc. Gas Generator Research and Development:
Bi-Gas Process, Quarterly Report October-December 1976, ERDA Number
FE-1207-25, Monroeville, Pennsylvania, 1977.
Flockenhaus, C., One Stage Shift-Conversion and Partial Methanation Process
for Upgrading Synthesis Gas to Pipeline Quality, Ninth Synthetic Pipeline
Gas Symposium, Chicago, Illinois, October 31-November 2, 1977.
C-30
-------
LIQUID PHASE METHANATION/SHIFT (LPM/S) PROCESS
1.0 General Information
1.1 Operating Principles - Methanation is accomplished in a fluidized
bed with feed gas bubbling up through a non-volatile fluid
(aliphatic or naphthenic mineral oil). The reaction is promoted
by nickel catalyst suspended in the oil. Filtered circulating
oil carries the reaction heat to an external cooler, generating
high-pressure steam.
1.2 Development Status^1'2'3' - Under contract to DOE, Chem Systems
Inc. has operated a bench scale and a PDU unit. A pilot plant
was built by Davy Powergas to operate at 1000 psi, handling
118 x 103 Nm3/day (2 x 106 scfd) feed gas. This plant was skid-
mounted for tests as an operating high-Btu gasification pilot
plant (HYGAS) starting in October 1976. Tests with two differ-
ent feed gas streams and two catalysts have been performed to
date.
1.3 Developer - Chem Systems Inc.
New York, N. Y. (201-575-8820)
(ERDA [DOE] Fossil Energy Contract 2036)
1.4 Commercial Applications - None.
2.0 Process Information
2.1 Flow Diagram (see Figure C-4) - Feed gas bubbles through a sus-
pension of nickel methanation catalyst in a high-temperature boil-
ing oil. The oil absorbs the heat of the methanation reaction and
is cooled by heat exchange to regenerate steam. Catalyst fines are
filtered from the oil before recycle to the main reactor. Mois-
ture and volatile oil are condensed from the methanation gas in
a separator prior to final methanation in a small fixed-bed
adiabatic reactor. Condensate is either recycled or discharged.
C-31
-------
J U
LIQUID
PHASE
METHANATOR
O
CO
ro
10
FINES FILTER
SEPARATOR
1
0-1
POLISHING
REACTOR
I. FEED GAS
2. LPM RAW REACTOR PRODUCT
3. MAKEUP OIL
4. FILTERED RECYCLE OIL
5. OILY WATER CONDENSATE
6. SEPARATED RECYCLE OIL
7. POLISHING REACTOR FEED
8. PRODUCT GAS
9. REJECT CTA CATALYST
10. REPLACEMENT CATALYST
(PERIODIC)
Figure C-4. Flow Diagram for Liquid Phase Methanation Pilot
-------
2.2 qu,pment - AH vessels are designed for Mgh pressure operat,on
(main reactor, separator, fiHer syste., polishing reactor, heat
exchangers). No data are available regarding aterials of
construction.
2.3 Feed Stream Characteristics^
(0'5 Ppm ₯ °r les*) to inhibit
. A ZnO g6ard bed is used
upstream.
Pressure: 2 MPa (300 psia) or more pipeline pressure is
suitable.
Temperature: 573°K-653°K (572°F-716°F)
H2/CO ratio between 1:1 and 3:1
Some H20 addition if H?/CO ratio is less than 3, to promote
water-gas shift reaction.
2.4 Operating Parameters^ ' '
t Pressure: 2 MPa - 6.8 MPa (SOOpsia - 1000 psia)
Space velocity: 500 to 10,000 hr'1
Catalyst size: .08 - .42 cm (.03 - .19 inches), previously
hydrogen-activated in dry state
t Temperature: 573°K-653°K (572°F-716°F); upper temperature
limited by thermal stability of liquid. Paraffin or aromatic
oils have been tested. Paraffins and/or naphthenics of low
pour point are preferred and vapor pressure 68 Pa (10 psia)
at 588°K (600°F).
2.5 Process Efficiency'1' - CO conversion in the main-reactor effluent
is about 96 percent, depending on H2/CO ratio in effluent. A
fixed-bed adiabatic polishing reactor follows the main reactor,
similar to commercial naphtha-reforming SNG plants. Space
velocity in the polishing reactor is 9500 hr" , which is a high
rate for fixed-bed techniques. No detectable CO exists in the
final product gas.
C-33
-------
2.6 Raw Material
Suspension catalyst: Two catalysts have been pilot plant
tested: Calsicat NI-230-S catalyst spheres (requires hydro-
gen activation) and INCO catalyst #087H Cno activation
required).
Oil: Small rate of replacement,' depends on operating temper-
ature and volatility of the oil. Pilot plant has used
FREEZENE-100 oil.
Polishing catalyst: Laporte-Davison CRG-A catalyst pellets
2.7 Utilities
Steam: Small internal requirement, process generates excess
steam.
Electricity: ?
Cooling water: ?
3.0 Process Advantages
No recycle compressor necessary.
Potential for increased by-product steam production.
Enhanced catalyst life.
Applicable to wide range of H^/CO in feed.
Capital cost saving (shift reaction combined).
Upstream desulfurization systems need not remove C02- Resulting
product stream can be high in ^S, aiding sulfur plant operation.
4.0 Process Limitations
Catalyst is sulfur-intolerant; feed-gas pretreatment required.
t 300 psi minimum pressure.
Maximum 650°K (705°F) temperature limitation to avoid degradation
of liquid.
t Operation of a LPM/S system with heavier hydrocarbons in feed gas
has not been tried.
Process not demonstrated on a commercial scale. Test results to
date indicate that gas distribution in the reactor is poor, lead-
ing to poor catalyst efficiency and catalyst losses from the reactor.
C-34
-------
5.0 Process Economics
(1)
Chen Systems has estimated cost savings for comparable systems of LPM/S
over conventional shifting plus Lurgi methanation. About 9 cents/106
Btu are calculated to be saved (product-gas basis, H2/CO initially 2.0;
utility financing method). Capital investment and power consumption
savings contributed as did increased steam output to the overall differ-
ence. Overall plant efficiency rises 3 percent with the LMP/S process
and about $30 million is saved in constructing a 7 x 106 Nm3/day (250
x 106 SCFD) plant.
6.0 Input Streams
6.1 Feed Gas (Stream 1) - Gases with H2/CO ratios of 1:3 have been
methanated at the bench or PDU scale. Pilot plant methanator
feeds have had H2/CO ratios of 4:6 and about 3. No other data
available.
6.2 Makeup Oil Containing Catalyst (Stream 3) - Aliphatic or naphthenic
oils are employed (C-|5-C2i). No actual operating data available.
6.3 Replacement Catalyst (Stream 10) - See Section 2.6.
7.0 Intermediate Streams
7.1 IMP Raw Reactor Product (Stream 2) - No data available.
7.2 Filtered Recycle Oil (Stream 4) - No data available.
7.3 Oily Water Condensate (Stream 5) - No data available.
7.4 Separated Recycle Oil (Stream 6) - No data available.
7.5 Polishing Reactor Feed (Stream 7) - PDU scale feed contains 2-
3 percent CO, 6-14 percent H2, 22+ percent C02, 55+ percent CH4,
and up to 15 percent moistureO). No data available from pilot
operations.
8.0 Discharge Streams
8.1 Product Gas (Stream 8) - No data available.
8.2 Reject Catalyst (Stream 9) - No data available.
C-35
-------
9.0 Data Gaps and Limitations
Available bench scale, PEDU, and pilot plant operating data very limited.
No data are available about the properties and composition of oily
condensate or spent catalyst.
10.0 Related Programs
Ongoing process and waste stream monitoring programs at the HYGAS pilot
plant are expected to generate data about the LPM/S system during 1978.
REFERENCES
1. Frank, M. E., et al, LPM/S PDU Results and Pilot Plant Status. In
Proceedings of the Eighth Pipeline Gas Symposium, American Gas Association,
Chicago, Illinois, October 1976, pp.161-182.
2. Frank, M. E., and Mednick, R. L., Liquid Phase Methanation Pilot Plant
Results, Ninth Synthetic Pipeline Gas Symposium, Chicago, Illinois,
October 30-November 2, 1977.
3. Best Way to Methanate Gas from Coal Sought, Chemical and Engineering News,
Vol. 56, No. 3, January 16, 1978, p. 30.
C-36
-------
EPA-600/7-78-186b
3. RECIPIENT'S ACCESSION NO.
Appendices A, B, and C
5. REPORT DATE
September 1978
6. PERFORMING ORGANIZATION COCfc
J7. AUTHOR(S)
M.Ghassemi, K.Crawford, and S.Quinlivan
8. PERFORMING ORGANIZATION REPORT NO.
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
.GEN'CY NAME AND ADORES
EPA, Office of Research and Development
Industrial Environmental Research Laboratory
Research Triangle Park, NC 27711
13. TYPE OF REPORT ASLD PERIOD COVERED
Final; 6/77 - 8/78
14. SPONSORING AGENCY CODE
EPA/600/13
o1o/c,ooc
91y/o41-28ol.
Pr°3ect officer * WiHiam J. Rhodes, Mail Drop 61,"
16. ABSTRACT
TThe report is 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 data base; Volume H contains data
sheets on gasification, gas purification, and gas upgrading; and Volume m contains
data sheets on air and water pollution control and on solid waste management.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b.IDENTIFIERS/OPEN END
Pollution
Coal
Coal Gasification
Assessments
Pollution Control
Stationary Sources
Environmental Assess-
ment
High-Btu Gasification
13B
21D
13H
14B
18. DISTRIBUTION STATEMENT
Unlimited
^^^^^^^^^
EPA Form 2220-1 (9-73)
19. SECURITY CLASS (This Report)
Unclass if led
21. NO. OF PAGES
20. SECURITY CLASS (This page)
Unclassified
22. PRICE
C-37
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