ASSESSMENT OF FOREIGN COAL CONVERSION TECHNOLOGIES, VOL III

• Document Type: CREST
• Collection: General CIA Records
• Document Number (FOIA) /ESDN (CREST): CIA-RDP83M00914R001000060018-0
• Release Decision: RIPPUB
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• Document Page Count: 325
• Document Creation Date: December 20, 2016
• Document Release Date: February 20, 2007
• Sequence Number: 18
• Publication Date: July 1, 1982
• Content Type: REPORT

r Re ed iil fl Cl RDP83M00914R00100006Q0-1-&
ASSESSMENT OF FOREIGN COAL
CONVERSION TECHNOLOGIES
Volume III
Appendix B
I
OFFICIAL USE ONLY
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OFFICIAL USE ONLY
INSTITUTE OF GAS TECHNOLOGY
IIT CENTER
CHICAGO, ILLINOIS 60616
ASSESSSMENT OF FOREIGN COAL CONVERSION TECHNOLOGIES
Volume III
Appendix B
by
Martin Novil
Christopher.F. Blazek
Edward J. Daniels
July 1982
I N S T I T U T E O F G A S T E C H N O L O G Y
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1
Coal-synfuels commercialization forecasts have traditionally been based
on the assumption that the technology for implementation does or will exist.
The processes under development worldwide represent an extention of existing
coal synfuels technologies in their potential for greater process efficiency,
tolerance of a variety of coals, or production of a specific product range of
greater interest to the anticipated commercial market. Because of technical,
economic or marketing problems most of the approximately 40 processes under
development outside the United States will not succeed beyond the demonstra-
tion phase. The economic and market related issues have been presented as
constraints in Volume II of this report. In this appendix, the technical and
development status of emerging coal synfuels technologies is presented.
Tables OB1 and OB2 summarize the status of processes analyzed. Technical
details are provided in the text of this volume.
46(3)/overvol3/ER
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Table OB1. COAL GASIFICATION AND COMBINED CYCLE PROCESSES
Largest
Pilot Plant
Size
(metric tons of Largest Proposed
coal/day Pilot Plant Commercial Commercial
Country/Process except as noted) Location Plane Sites
Ruhr-100 High Pressure 135
Lurgi
VEW Partial Gasification 360
(Combined Cycle)
Eergbau-Forschung 10.8
Nuclear Assisted
Rheinbraun Nuclear 25
Lignite Hydrogasification
Steag/Lurgi (Combined-
Cycle)
* Already under construction or completed.
Volklinger-Fur-
stenhausen, FRG
Oberhausen/
Holten
Muscle Shoals, US*
Cool Water Project,UJe
Tennessee Eastman,US
Stockum, FRG 1 Gersteinwerke, FRG
Essen, PEG
Weaseling, FRG
Ruhr region, FRG /
Proposed
Commercial
Plant Size
(metric tons/of
coal/day
except as noted)
Continued in-house research feasibility
studies continue, no commerical plant
announcements. Intended plans have been
cancelled for the FRG.
No government support, commercial plant
under construction by Rheinbraun.
Government support ending soon, no
commercial plane announced.
Government interested in
supporting commercial demon-
stration plant in the FRG.
Government support ending soon,
no commercial plans announced.
State government of Nordrhein, Westfalia,
is expected to fund project through
demonstration phase.
Commercialization prospects are low with
anticipated government support ending.
Research continues, no commercial plans
announced.
Research continues, no commercial plans
announced.
Government interested in supporting
commercial demonstration plant in
the FRG.
Commercial size facility shut down for
economic reasons.
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Table OB1. COAL GASIFICATION AND COMBINED CYCLE PROCESSES, (Cont.)
Country/Process
Great Britain
British Gas
Slagging Lurgi
National Coal Board PFB 5MW
(Pressurized Fluidized
Bed Combustion of Coal)**
Bharat Heavy Electrical 144.0
Japan
Sumitomi Molten Iron 60
Coal Mining Research 40
(Combined Cycle)
Largest Proposed
Pilot Plant Commercial Commercial
Location Plans Sites
Westfield,
Scotland
Abington,
England
Leatherhead, / Grimethorpe,
England England
Trichy, India / Tiruchirapalli, India*
Kashima, Japan /
Kitakyusha, Japan
Yubari, Hokkaido, /
Japan
* Already under construction or completed.
** Noted here because of potential competitaion with combined cycle
coal gasification for electric power generation.
Largest
Pilot Plant
Size
(metric tons of
coal/day
except as noted)
Proposed
Commercial
Plant Size
(metric tons/of
coal/day
except as noted)
British Gas is trying to find a
sponsor for commercialization.
Proposed plant does not apper to be
ready for commercialization.
20NW* Testing of the demonstration size unit
will last throuh 1982.
Experimental research facility in
operation since 1962.
Future plans call for the construction
of a demonstration plant in Japan, and
Australia.
Still in early phases of pilot plant
testing.
Future plans call for construction of a
40 MTPD pilot plant.
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Table OB1. COAL GASIFICATION AND COMBINED CYCLE PROCESSES, (Cont.)
Largest
Pilot Plant
Size
(metric tons of
coal/day
Country/Process except as noted)
Power Engineering 100-200
High Speed Pyrolysis
(Fixed Bed)
Largest Proposed
Pilot Plant Commercial Commercial
Location Plans Sites
Proposed
Commercial
Plant Size
(metric tons of
coal/day
except as noted) Official Status
175 mt/hr Demonstration plant at Karsnoyarsk
(demon- still in the construction phase.
stration Proposed comercial plant at Kansk
plant) Achinsk has capacity of 25 X 106 metric
tons/yr.
M _ r _ - - = i - - = = = - -~ - _.
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Table OB2. COAL LIQUEFACTION PROCESSES
Largest
Pilot Plant
Size
(metric tons of Largest Proposed
coal/day) Pilot Plant Commercial Commercial
Country/Process except as noted) Location Plans Sites
Ruhrkohle/Veba Oil
Catalytic Hydrogenation
Saarbergwerke Catalytic
Hydrogenation
200
6
Volklinge r /
Furatenhausen,
W.G.
11,250
6000
National Coal Board
2.7
Stoke Orchard,
Supercritical Gas
Solvent Extraction
England
National Coal Board
0.7
Stoke Orchard,
Liquid Solvent Extraction
England
CSIRO High Speed
0.5
North Ryde,
Flash Pyrolysis
England
Central Mining Institute 0.12 Tychy-Wyry,
Catalytic Hydrogenation Poland
Proposed
Commercial
Plant Size
(metric tons of
coal/day
except as noted) Official Status
South Africa
Commercial development in doubt
pending decision on FRG support.
Commercial development in doubt pending
decision on FRG support at end of pilot
plant tests in 1983.
Commercial development in doubt pending
decision on FRG support.
Commercial development unlikely in the
near term due to National Coal Board R&D
cutbacks.
Commercial development unlikely in the
near term due to National Caol Board R&D
cutbacks.
Research is still underway but no
commercial plans have been announced.
Sasol (Lurgi/Fischer
Sasolburg, (Sasol II)*
11000
Commercial operation since 1955.
Tropach)
Secunda (Sasol II + III)*
60000
Full commercial operation of both
facilities by end of 1982.
* Under construction or already completed.
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Table OB2. COAL LIQUEFACTION PROCESSES, (Cont.)
Largest
Pilot Plant
Size
(metric tons of Largest Proposed
coal/day) Pilot Plant Commercial Commercial
Country/Process except as noted) Location Plans Sites
South African Explosives
& Chemicals Koppers Totzek
South African Explosives
& Chemicals/Mobil
Sasol Direct Bench scale
Liquefaction
>
CY
).4.
Mitsui Solvent Refined 4.5
Coal
W
Nippon Brown Coal 0.55
"Kominic" Direct
Hydrogenation
Mitsui Eng. & Ship 2.4
Building Direct
Sumitomo Solvent
Extraction
Mitsubishi Heavy Ind. 1
Solvolysis
Inst. of Mineral Fuels 0.1
Catalytic Hydrogenation
Sasolburg
Ohmuta City,
Japan
Ibaraki Prefac-
ture, Japan
* Under construction or already completed.
Proposed
Commercial
Plant Size
(metric tons/of
coal/day
except as noted)
Commercial operation since 1972 for
ammonia production.
Status of this development in doubt.
Methanol production intended.
Still in the early experimental phase
of development.
Recent announcements have indicated that
construction of the demonstration
facility has been postponed indefinitely.
Construction of a 50 MTPD demonstration
plant in Australia should be completed
in 1983.
Pilot plant still in early research
stage.
Pilot plant still in early research
stage.
Pilot plant in construction
at Belkovskaya mine (10 MTPD).
Future plans call for pilot
plant construction at Kansk-Achinsk
(75 MTPD).
M r w -
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Pag e
I Shell-Koppers Coal Gasification Process B-3
I 2. High-Temperature Winkler Coal Gasification Process B-10
3. Saarberg-Otto Coal Gasification Process B-10
4. RCH/RAG Texaco Coal Gasification Process B-23
I 5. Ruhr 100 High-Pressure Lurgi Gasification B-31
6. KGN Fixed-Bed Coal Gasification Process B-38
1 7. VEW Coal Conversion Process B-44
8. Bergbau-Forschung Nuclear-Assisted Coal Gasification Process B-50
I 9. Rheinbraun Nuclear-Assisted Lignite Hydrogasification Process B-57
10. KHD Kloeckner-Humbolt-Wedag Molten-Iron Coal Gasification
Process B-64
1 11. Ruhrkohle/Veba Oil Hydrogeneration Coal Liquefaction Process B-71
12. Saarberg Catalytic Hydrogenation Coal Liquefaction Process B-82
13. Rheinbraun Brown Coal Liquefaction Process B-94
14. Steag/Lurgi Combined Cycle Project B-103
' 15.1 British Gas Slagging Coal Gasification Process B-114
16. British Gas Composite Coal Gasification Process B-126
I 17. National Coal Board Supercritical Gas Solvent Extraction
Coal Liquefaction Process B-128
18. National Coal Board Liquid Solvent Extraction Coal Lique-
faction Process B-138
19. National Coal Board Pressurized Fluidized-Bed Coal
Combustion Process B-149
' 20. Esso Chemically Active Fluidized Bed Coal Gasification
Process B-162
21. CSIRO Flash Pyrolysis Coal Liquefaction Process B-168
' 22. Polish Central Mining Institute Catalytic Hydrogenation
Coal Liquefaction Process B-174
I 23. SASOL (Lurgi/Fischer-Tropsch) Coal Liquefaction Process B-182
24. Modderfontein Coal-to-Methanol Process B-201
I 25. Sasol Direct Coal Liquefaction Process B-211
26. Central Fuels Research Institute (Lurgi) Coal
Gasification Process B-217
I 27. Central Fuels Research Institute (Bergius) Coal
Liquefaction Process B-220
I N S T I T U T E O F G A S T E C H N O L O G Y
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TABLE OF CONTENTS, Cont.
Page
28.
Bharat Heavy Electricals Ltd. Combined Cycle Coal
Gasification Process
B-222
29.
Mitsui Solvent Refined Coal Process
B-228
30.
Nippon Brown Coal "Kominic" Hydrogenation Direct Coal
Liquefaction Process
B-236
31.
Mitsui Engineering and Shipbuilding Direct Coal Liquefaction
Process
B-262
32.
Sumitomo Solvent Extraction Coal Liquefaction Process
B-266
33.
Sumitomo Molten Iron Coal Gasification Process
B-275
34.
Mitsubishi Heavy Industries Solvolysis Coal Liquefaction
Process
B-302
35.
Mitsui M-Gas Coal Gasification Process
B-308
36.
Coal Mining Research Center Combined Cycle Coal
Gasification Process
B-309
37.
Hitachi Coal/Oil Mixture Gasification Process
B-310
38.
Power Engineering High-Speed Pyrolysis Process
B-312
39.
Institute of Fossil Fuels Hydrogenation Coal Liquefaction
Process
B-314
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The following sections survey 39 foreign coal conversion processes.
These processes are numbered consecutively and organized by country of origin:
? Sections 1 through 14 are West German processes.
? Sections 15 through 20 are British processes.
? Section 21 is an Australian process.
? Section 22 is a Polish process.
? Sections 23 through 25 are South African processes.
? Sections 26 through 28 are Indian processes.
? Sections 29 through 37 are Japanese processes.
? Sections 38 and 39 are Soviet processes.
I N S T I T U T E O F G A S T E C H N O L O G Y
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1. Shell-Koppers Coal Gasification Process
Process Description
The coal gasification reactor in the Shell-Koppers process is essentially
an empty pressure vessel equipped with diametrically opposed diffuser guns.
Crushed, ground, and dried coal sized to 90%98%
up to 2,700?C
600 psig
tars, oils
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Table 1, Part 2. TECHNOLOGY FACT SHEET: RCH/RAG Texaco Coal Gasification Process
Z STATUS OF DEVELOPMENT:
OPERATING FACILITIES -
Texaco pilot plant (15 tons/day) in Montebello, Cal. since early 50's. Demon-
stration plant built at Oberhausen-Holten, West Germany in 1978. Tests at this
6 ton/hr facility are still underway. TVA has recently started up a Texaco
gasifier at their Muscle Shoals facility.
MAJOR FUNDING AGENCY Two-thirds funded by Federal German Ministry for Research and Technology.
ANNUAL LEVEL OF FUNDING - Initial plant cost 29 MM DM, Annual operating cost offset by syngas sale to
Ruhrchemie. Government support for commercial develoment seems likely.
TECHNICAL PROBLEMS: From the tests to date, two areas of critical concern are the ash removal systems and
the refractory lining of the gasifier vessel. Molten ash removal has caused a shut-down at the Muscle Shoals
Texaco facility and equipment re-design is required. System re-design was also rqquired at the RCH/RAG pilot
plant in 1978. The most suitable ceramic lining materials tested to date have expected operating lifetimes
of just over 10,000 hours. This relatively short lifetime may call for gasification vessel re-design.
OTHER FACTORS AFFECTING OVERALL FEASIBILITY:
? Testing of alternative components required
? Substitution of water as the suspension agent
? Optimization of heat transfer equipment
? Carbon recycle system needs further demonstration
? Testing of different types of coals required.
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Research on the RCH/RAG version of the Texaco coal gasification process
also includes the following aspects:
? Test of alternative components and systems
? Investigation of fundamental relationships and further system optimization
? Substitution of water as the suspension agent
? Optimization of heat transfer equipment
e Optimization of recycle carbon and noncombustible materials
? Testing of different geographic coals
Relationship to Prior Technologies
The Texaco coal gasification process (TCGP) is an extension of their
synthesis gas generating process (TSGGP) for converting high-sulfur residual
petroleum fuels and tars into synthesis gas. Research on the TSGGP was
started in 1949 and resulted in licensing agreements by 1953. More than 75
plants, in 22 countries, have been constructed since 1955 for the ammonia,
methanol, and oxo-chemical industries.
The feedstock flexibility of the TSGGP suggested to Texaco that this
concept could also be used for converting coal into synthesis gas. The
abundance of coal in the U.S. and its potential role as a leading energy
supply prompted Texaco to initiate process development research on the Texaco
coal gasification process in 1984. In the early 50's Texaco's Montebello
research laboratory, east of Los Angeles, started development work on a 15
ton/day coal gasification pilot plant. This resluted in the construction and
two-year operation of a 100 ton/day demonstration plant in Morgantown, West
Virginia, in 1956. This project was sponsored by the U.S. Government through
the Bureau of Mines. Eastern coal was used in the gasifier to produce
synthesis gas to make ammonia. This air-blown quench-type gasifier operated
at 400 psig and incorporated a refractory lining and water jacket. Due to the
increasing availability of low-cost oil and natural gas, this demonstration
plant research was discontinued.
The Arab oil embargo in 1973 rekindled Texaco's interest in coal-gas-
ification. During this same time period Ruhrchemie AG, a syngas consumer, and
Ruhrkohle AG, a coal producer and processor (RCH/RAG) of West Germany began
their own investigation of coal gasification technologies. At that time,
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their conclusions resulted in the selection of the Texaco process for commer-
cial applications. In 1975, RCH/RAG entered into a licensing agreement with
Texaco Development Corp. to adapt the TCGP to a commercial scale. Ruhrchemie's
and Ruhrhohle's experience in the commercial scale equipment aspects related
to coal gasification, and the experimetnal results from Texaco's Montebello
gasifier were combined to build a 150 ton/day demonstration plant at
Oberhausen-Holten, West Germany.
Operating Facilities
Ruhrkohle and Ruhrchemie have operated a 160 ton/day Texaco coal gasifi-
cation pilot plant since 1977 in Oberhausen, West Germany, More than
50,000 tons/day of coal have been gasified in 10,000 hours of operation.
Texaco, U.S. also operates two 15 ton/day gasifiers at its Montebello
California research facility. Another Texaco gasification facility is under-
going testing at the Tennessee Valley Authority's Muscle Shoals ammonia
facility.
The Texaco coal gasification process has thus far been licensed to a
number of plants around the world. These plants and other organizations
considering Texaco coal gasification technology include following:
? Tennessee Eastman project
? Cool Water Combined Cycle Coal Gasification project
? Tennessee Valley Authority project
? Dow Chemical (Texas) project
? Ruhrkohle and Ruhrchemie project
? Alsands project
? WyCoal Gas project
? SRC-II project
? Mitsubishi Heavy Industries and Central Research Institutes of the
Electric Power Industry (Japan) Combined Cycle project
? Rotterdam Municipal Utility project
? Ube Industries (Japan) Ammonia project
? Nyanes Petroleum (Sweden) project
? Moers-Meerbeck (West Germany) Combined Cycle project
B-28
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Major Funding Agencies
The total demonstration plant costs were 29 million Deutsche Marks (DM),
with approximately two thirds of the funding from the Federal German Ministry
for Research and Technology. This relatively minor project cost is the result
of effective utilization of the Ruhrchemie plant infrastructure and the sale
of the syngas to Ruhrchemie's syngas network, which helps to offset the annual
operating cost expenditures. The companies of Carl Still, Rechlinsghausen,
and Friedrick Uhde, Dortmund FDR, are monitoring the project. Ruhrkohle's
project responsibilities are being carried out by its subsidiary "Gesellschaft
fur Vergasung and Verflussigung von Steinkohle mbH."
In June of 1982 the West German Ministry of Economics requested addi-
tional environmental and economic information from RCH/RAG in order to make a
final decision to assist in the construction of a demonstration facility. The
Ministry can grant up to half of the 220 million Deutsch Mark ($93 million DM)
cost of the project. The Ministry is expected to make a final decision of the
RAG/RCH Texaco demonstration project in September or October of 1982.
Technical Problems
From the test results to date, two areas of critical concern are the ash
removal systems and the refractory lining of the gasification vessel. The
prerent shut-down of the TVA Muscle Shoals Texaco plant due to ash removal
problems is an illustration of those problems. Ash removal problems were also
encountered at the RCH/RAG Texaco pilot plant These problems resulted from
the unexpectedly low density of the slag, which inhibited settling velocities
in the slag bath. The low bulk density ash also created very high transport
volumes. These problems were overcome at the RCH/RAG pilot plant through
system redesign. Another ash removal problem centers on the necessity to use
special valves which must be resistant to the strongly eroding lag solids.
These valves are subject to failure if lockhopper operation is not closely
monitored.
Another area of major concern at the onset of the RCH/RAG pilot plant
operation was reactor lining life. A test program to determine the best
refractory lining material was undertaken at the Holten pilot plant. In
special test zones in the reactor jacket lining, some 50 different ceramic
materials from various manufacturers were tested. Inferior refractory lining
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life times were encountered in the early tests due to early failure of the
weakest materials. The most suitable ceramic lining materials tested to date
have expected lifetimes of more than 16,000 hours. Unfortunately, this is
less than one year of commercial plant operation. Strict temperature control
will be required for commercial operation.
Another problem that was mentioned in the questionnaire responses is that
because the RCH/RAG Texaco is an entrained gasifier, the inventory of coal in
the gasifier is small. Consequently, the stability of the process becomes
very dependent upon the capability to adequately monitor the conditions of the
process. Measurement of the reactor temperature has been a particular problem
of the RCH/RAG pilot plant as the temperature monitoring devices apparently
degrade in their environment of slag and hydrogen.
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WEST GERMAN COAL GASIFICATION PROCESS
5. Ruhr 100 High Pressure Lurgi Coal Gasification Process
Process Description
The Ruhr 100 gasifier is designed to operate between 25 and 100 bars
(360 to 1450 psi) at a coal feed rate of 3 to 7 metric tons per hour. The
135 metric ton/per day pilot plant gasifier has a 1.5 meter inside diameter
and a height of 11.5 meters, excluding the coal lockhoppers. The total
reactor height including the ash and coal lockhoppers is more than 20
meters. Coal feed at the Ruhr 100 pilot plant is stored in bunkers at the
neighboring Furst Leopold mine for transportation to the gasifier by a 350
meter conveyor belt system. The coal arriving at the gasifier plant is then
dedusted and screened to 6.3 mm before being stored in an 85 metric ton coal
bin. Coal from the bin is fed to a weighing system before injection into the
coal lockhoppers. Two alternating mode lockhoppers are used to feed coal into
the gasifier to reduce lockhopper gas losses by 40%.
The pressurized Ruhr 100 reactor vessel consists of three sections with a
cooling water jacket surrounding the middle and bottom sections. The outer
walls of the gasifier vessel are made of a special steel alloy, formerly used
only on a nuclear reactor, to help minimize the wall thickness. The inner
shell of the reactor is comprised of a number of different alloys in order to
collect data for future designs of larger units. Special temperature measure-
ment equipment is also installed in the pilot plant gasifier in order to
monitor the gasification reactions at various points within the gasifer.
As the coal enters the gasifier via the lockhoppers, it descends through
a drying zone, carbonization zone, gasification zone and combustion zone. The
slowly descending coal is evenly distributed within the gasifier by an
externally driven rotating grate. A special feature of the Ruhr 100 gasifier
is the addition of an extra raw gas outlet located below the carbonization
zone. This second outlet makes it possible to influence the temperature and
flow rate of the hot raw gases into the devolitilization zone. The effect of
this control mechanism will be to reduce the amount of fines carryover by a
reduction of gas velocity through the topmost gas outlet. The product gas
exits the "clear gas" outlet below the carbonization zone at a temperature of
approximately 800?C and the carbonization zone outlet is rich in methane and
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hydrocarbon liquids which were formed in the carbonization zone. A technology
fact sheet for this process is presented in Table 1.
Process Goals
The basic objective of the Ruhr 100 project is to increase the operating
pressure of the Lurgi gasifier from its normal operating pressure range of 20
to 28 bars (290 to 400 psi) to a maximum operating pressure of 100 bars
(1450 psi). One positive outcome in this incresed pressure operation will be
an increase in gasifier throughput without increasing the gasifier diameter.
In addition bench scale and theoretical research by G. Baron has indicated
that due to the higher partial pressure of hydrogen of the carbonization pro-
ducts results in a 60% to 80% (by volume) increase in methane formation over
25 bar operation. The increase in methane production within the gasifier
reduces coal and oxygen consumption due to the exothermic reaction of methane
formation and the resulting reduction of the partial combustion of coal within
the gasifier to produce heat. The results of tests made by Sustmann and
Ziescehe also indicate a 75% reduction of tars in the raw gas stream at the
100 (1450 psi) bar operating pressure. This is due in part to the hydrogasi-
fication of some of the tars in the gasifier. In addition to less tar produc-
tion, fewer higher hydrocarbons and phenols are formed as a result of the
100 bar 91450 psi) gasification pressure.
The increase in methane production within the gasifier is of particular
importance to SNG plant design and economics and of obvious concern to the
Ruhrgas AG. The increase in the methane content of the raw gas will reduce
the size of the downstream balance plant, especially the methanation plant, to
further reduce plant costs. An added benefit to increased pressure operation
is the production of SNG above the typical pipeline operating pressure of
70 bars (1015 psi) which will eliminate the need for recompression equipment.
Another special feature of the Ruhr 100 gasifier is the addition of a
second raw gas outlet just above the gasification zone. This second gas out-
let will reduce fines carryover by reducing the gas velocity before the
devolatilization zone. The gas exiting this second outlet will also contain
far fewer condensible products because these condensible products are formed
predominantely in the devolatilization zone and withdrawn from the topmost
(1st) gas outlet. The reduction of fines carryover will enable the Ruhr
100 gasifier to accept a wide range of coal particle sizes including possible
run-of-mine coals.
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I N S T I T U T E O F G A S T E C H N O L O G Y
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Ruhr 100 High-Pressure Lurgi Coal
Table 1 Part 1 TECHNOLOGY FACT SHEET: Gasification Process
NEW TECHNOLOGY (PROCESS DESCRIPTION AND GOALS):
The Ruhr 100 gasifier is a fixed bed, countercurrent
flow reactor consisting of three sections with a cooling water jacket around the middle and bottom sections.
The goals of the project are to increase throughput and methane production by increasing reaction pressures
to 1450 psia. Other goal sinclude operation of the gasifier on run-of-mine coals and "clear" gas removal from
novel side port.
RELATIONSHIP TO PRIOR TECHNOLOGY (INCLUDING STATE OF DEVELOPMENT OF PRIOR TECHNOLOGY): The "Ruhr .100" project
is an extension of the continuous research effort to improve the Lurgi coal gasification process. This
research has led to the development of the'Mark IV gasifier, the British Gas Slagging Lurgi gasifier and
the Ruhr 100 gasifier. The Lurgi?gasifier has been commercially available for more than 40 years. The Ruhr
100 gasifier is currently in the pilot plant stage of development with a 135 metric ton/day plant operating in the FRG.
CHARACTERISTICS OF THE TECHNOLOGY:
PRIMARY OUTPUT (DESIGN CASE) ........................
TYPE OF PROCESS .....................................
FEEDSTOCK REQUIRFIENTS..............................
OVERALL THERMAL EFFICIENCY (INCLUDING BY-PRODUCTS)..
CARBON CONVERSION EFFICIENCY. .......................
OPERATING TEMPERATURE ...............................
OPERATING PRESSURE ..................:...............
BY-PRODUCTS .........................................
*
Represents efficiency for high Btu gas production
**
Cold gas efficiency for syngas
NEW TECHNOLOGY
340 Btu/SCF
Screened to - 6.3 mm
**
79.7%
98.9%
800?C
1450 psia
tar, oils, ash
PRIOR TECHNOLOGY
309 Btu/SCF
Below coal's ash fusion
temperature
370 psia
Ash, steam, tars, oils,
phenols
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Table 1 Part 2 TECHNOLOGY FACT SHEET: Ruhr 100 High-Pressure Lurgi Coal
Gasification Process
STATUS OF DEVELOPMENT:
OPERl1TING FACILITIES -
The Ruhrgas/Ruhrchemie Lurgi coke-oven gas facility at Dorsten, FRG, was
selected for the 135 metric ton/day Ruhr 100 pilot plant site: this decision
was based on the availability of offsite facilities and experienced operators that
were still available trom. the ormer urg opera ons a c ose in
MAJOR FUNDING AGENCY 75% funding by German Ministry of Research and Technology until the end of 1982.
ANNUAL LEVEL OF FUNDING - Four year funding program at 150 million DM level until the end of 1982.
TECHNICAL PROBLEMS: The areas of critical concern include lock hopper operation at high pressures, fouling
due to fines carryover, utilization of caking coals, and plugging of downstream processes. Of these problems
lock hopper operation at high pressures for coal feeding and ash removal and operation which involves large
amounts of gas, high speed operation, erosion of valves, control system reliability and sealing are areas of
OTHER FACTORS AFFECTING OVERALL FEASIBILITY:
Ruhrkohle is an owner/parent company of Ruhrgas. Incentive
for development is to provide SNG for Ruhrgas while providing a market for Ruhrkohle's coal.
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Relationship to Prior Technology
The "Ruhr 100" project is an extension of a continuous research effort to
improve the Lurgi coal gasification process. This research, which started
with the theoretical work of Drawe and Danulat, produced the first generation
Lurgi gasifier for the gasification of lignite under pressure in the early
19301s.
Due to the loss of availability of lignites from Central Germany after
World War II, Lurgi and Ruhrgas continued research to adapt the Lurgi process
to the gasification of hard coals such as the mildly caking subbituminous
variety available in the Ruhr district of West Germany. This renewed research
led to the construction of a pilot plant in 1950 which was equipped with a
1 meter diameter gasifier at the Ruhrchemie site in Oberhausen-Holten, West
Germany. The pilot plant tests resulted in two different second generation
Lurgi designs.
The Dorsten site acted as a proving ground for technical improvements to
the Lurgi process. These improvements were in the area of corrosion control,
grate water cooling systems, raw gas treatment, and coal and ash lockhopper
design. Although the second generation designs represented a marked improve-
ment to gasifier throughput, further increases in gasifier throughput were
still desirable. Based on the operating experience gained in the Dorsten
facility, a third generation design with a larger internal diameter was
introduced in 1969.
In the wake of the Arab oil embargo of 1973, a sense of urgency developed
to further improve the Lurgi gasification design. This research effort has
centered on increasing gasifier capacity, especially for SNG plants, and
increasing the range of gasifiable coal type and coal size distributions,
including run of mine coals. This has led to the development of the Mark V
gasifier, the British Gas slagging Lurgi gasifier, and the Ruhr 100 gasifier,
all based in part on the basic Lurgi design philosophy. Work on the Ruhr 100
project started in 1973 when Ruhrgas AG, Ruhrkohle AG, and Steag AG formed a
joint venture to explore a high pressure Lurgi option for hard coal gasifi-
cation to increase gasifier throughput and methane conversion rates.
I N S T I T U T E O F G A S T E C H N O L O G Y
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Operating Facilities
The Ruhrgas/Ruhrchemie Lurgi coke-oven gas facility at Dorsten, West
Germany was selected as the Ruhr 100 pilot plant site. From the first test in
September of 1979 to September of 1981 a total of 15 tests were conducted.
Over 2360 hours of operation were logged which consumed about 16,000 tons of
coal and produced nearly 9.6 million m3 of gas. The tests lasted from 3 to
20 days. During the most recent test conducted at 96 atms, 60 to 100% tar
recycle was achieved over a ten day period. Previous tests proved gasifier
operation with run-of-mine coal having an ash content of up to 50%.
After the pilot plant stage of development which is scheduled to be
partially funded by the German Government through 1982. Plans call for the
possible construction of a 3-million tonne/year commerical facility to produce
1.5 billion m3/yr of SNG.
Major Funding Agencies
The total engineering design and cost of the pilot plant in 1977 was
about DM 60 x 106. An additional D14 90 x 106 was spent to cover four to five
years of operating costs. The West German Ministry of Research and Technology
has funded 75% of this cost to date. Funding is expected to last through
1982. It does not appear that the Ministry will fund this project after this
period.
Technical Problems
Although pilot plant testing is still underway, areas of critical concern
have surfaced. These areas include lockhopper operation at high pressures,
fouling due to fines carryover, utilization of caking coals, and plugging of
downstream processes. Of these problems, lockhopper operation at high pres-
sures for coal feeding and ash removal and operation with caking coals are the
most important. Due to rapid pressurization and depressurization which
involves large amounts of gas, high speed operation, erosion of valves, con-
trol system reliability and sealing are areas of great concern. The utiliza-
tion of caking coals will involve better control and operation of the devola-
tilization section in order to avoid coal caking in the lower sections of the
gasifier.
The Ruhr 100 high pressure Lurgi gasifier is, in general, a more complex
gasifier with regard to control systems because of the high pressure
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However, the Ruhr 100 operates at temperatures below the coal's ash fusion
temperature and some of the problems expected with the British gas slagging
Lurgi type gasifier are avoided. The most critical problems avoided are the
reactor lining corrosion/erosion problem and the ash removal (slag tap) design
problem.
During testing in the fall of 1981, the coal feed preparation was changed
from 6 mm to 35 mm feedstock, to a feed containing 40 to 45% by weight below
6 mm material. Plugging of the clear gas outlet occurred and all gas was
withdrawn from the carbonization gas outlet.
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WEST GERMAN COAL GASIFICATION PROCESS
6. KGN Fixed Bed Coal Gasification
Process Description
The Kohlegas Nordrhein GmbH (KGN) fixed bed gasifier consists of a con-
ventional revolving grate with a carbonizaton shaft as shown in Figure 1. The
gasifier has been designed to operate in both a cyclic and continuous mode.
However, synthesis gas (medium Btu-gas) can he only produced in the continuous
mode of operation. In the continuous mode the gasifier can be operated between
atmospheric and 30 bars of pressure depending on end use application.
Due to the nature of the counter flow fixed bed design this process can
only operate on non caking or slightly caking coals. In addition, coal feed
size requirements are also critical. To overcome coal feed size requirements
and the ever increasing production of fines in modern mining techniques, the
KGN process utilized anthracite briquettes of approximately 40 grams in weight
produced from fines. These briquettes are introduced into the gasifier via a
coal-lockhopper.
The coal which enters the gasifier passed through a rotating distributor
grate at the top into the carbonization zone where carbonization gas is pro-
duced. This crude gas containing tars and higher hydrocarbon gases is drawn
into the recycle carbonization shaft by steam injection. By drawing the tar
and soot laden gases into the gasification zone through the carbonization
shaft, all tars and higher hydrocarbons are thermally cracked. In addition,
this technique also guarantees that only coke reaches the gasification zone as
it is formed in its downward movement in the gasifier.
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The oxygen and part of the steam required for the gasification reaction
are injected into the bottom of the gasifier under the revolving ash grates.
These grates separate the ash which is then collected in an ash lockhopper
underneath the gasification vessel.
The superheated steam (580?C) required for the process is partially pro-
duced by heat exchange with the exiting gas and from heat collected by the
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gasifier cooling jacket.
are injected at 35 bars.
The oxygen and the steam required for gasification
As the tar-free synthesis gas exits the gasifier at
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800?C and 30 bars, most of the soot in the gas is removed in hot cyclones.
The gas is then cooled to 280?C before entering a final soot removed stage
using venturi scrubbers.
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1
Drying zone
Distillation
zone
Gasification
zone
I
Driving device
for rotating grate
Ash lock hopper
Figure 1. KGN FIXED-BED GASIFIER
I N S T I T U T E O F G A S T E C H N O L O G Y
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Part I TECHNOLOGY FACT SHEET: KGN Fixed Bed Coal Gasification
NEW TECHNOLOGY (PROCESS DESCRIPTION AND GOALS):
The KNG gasifier consists of a conventional revolving
grate fixed bed gasifier which incorporates a carbonization shaft (Figure 1). Coal, fed through lockhoppers,
descends slowly through the integrated pyrolysis/char bed to the gasification zone. Raw synthesis gas con-
taining tar is recycled to the combustion zone for thermal cracking. Project goals include the production
of a tar-free gas which would simplify gas conditioning. Other goals include the use of fines with a high
ash content, the development of a briquetting stage, and the development of a suitable gas cleaning system.
RELATIONSHIP TO PRIOR TECHNOLOGY (INCLUDING STATE OF DEVELOPMENT OF PRIOR TECHNOLOGY):
Based on a conventional counter flow-fixed-bed principle-.--.-
PRIMARY PRIMARY OUTPUT (DESIGN CASE) ........................
TYPE OF PROCESS .....................................
FEEDSTOCK REQUIREMENTS ..............................
OVERALL THERMAL EFFICIENCY (INCLUDING BY-PRODUCTS)..
CARBON CONVERSION EFFICIENCY ........................
OPERATING T1WERATURE ...............................
OPERATING PRESSURE .................. ................
BY-PRODUCTS .:....................... ................
NEW TECHNOLOGY
234 Btu/SCF
Pressurized fixed bed
Non-caking coals
80%
850?C
30 bars
Steam, ash, sulfur
PRIOR TECHNOLOGY
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 ~ No 1W M M t vec Re 20Wd2/2f lA-FW3M"4R"000 18-OM M = M
aLa ub ur- ur.vtLvrmNT: Pilot Plant
u1U-INU FACILITLE5 - Iwo LO11/11L p11UL plant capacie or operating in the cyclic or continuous
1
C
mode.Pilot Plant is located near Huchelhoven in the colliery of Sophia-
ANNUAL LEVEL OF FUNDING - Total cost approximately 19 million DM.
TECHNICAL PROBLEMS: Early operating problems with the grate drives and lockhoppers were reported but
too have since been fixed.
OTHER FACTORS AFFECTING OVERALL FEASIBILITY: Briquetting requirements may prove uneconomical unless
z inexpensive source of coal fines is available.
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After further cooling the 75?C the gas can be cleaned of CO2 and sulfur
compounds in a Stretford unit. The final gas processing step involves gas
drying in a Glycol unit.
The objective of the Kohlegas Nordrhein project is to develop a commer-
cial size fixed bed gasifier which is easy to operate and which can produce
tar-free gases. In addition to the production of tar free gas the sensible
heat of the crude gas will also be available.
Relationship to Prior Technology -
The KGN process is based on the commercially proven fixed bed counter
flow gasifer principle.
Operating Facilities
Construction of a 2.0 tonne/hour pilot plant was completed in February
1979. The reactor, which has an internal diameter of 2.1 meters is capable of
operating at pressures of up to 7 bars. Over 200 tons of briquette coal have
been processed in the facility with a turndown ratios of nearly 20%. The
longest run of 1000 hours was achieved without problems. Only two operators
per shift are required for this fully automatic pilot plant design. Briquettes
for the process are produced on site by the colliery in which the pilot plant
is located. This facility, located near Huchelhoven, West Germany, is
operated by Gewerkschaft Sophia-Jacoba. In actuality Kohlegas Nordhein GmbH
is a joint venture company of Gewerkschaft Sophia-Jacoba and Projektierung
Chemische Verfahrenstechnik GmbH. This site was selected due to the availa-
bility of nearly 1.8 million tons/year of coal fines from the colliery opera-
t ion.
Major Funding Agencies
The project is sponsored by the Ministry of Economics of the state of
Nordrhein Westfalia, West Germany. Project cost is estimated to be approxi-
mately 19 million DM. Operation of the pilot plant phase of development is
expected to continue through 1982. The State Ministry of Economics is
expected to continue funding of this process through the demonstration phase.
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During early pilot plant operation mechanical difficulties were encoun-
tered with the rotating grate drive units and the coal-lockhoppers. Operating
problems with the carbonization recycle tube also occurred. However, all of
these problems were overcome and the pilot plant has since sustained smooth
operation.
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WEST GERMAN COAL GASIFICATION PROCESS
7. VEW Coal Conversion Process
Process Description
The Verinigte Elektrizitatswerke Westfalen (VEW) coal conversion process
has been developed for combined cycle power generation applications. However,
it can also be applied to synthesis gas and reducer gas production when oper-
ated with oxygen.
In this process, shown in Figure 1, coal is partially gasified in the
entrained flow gasifier. Prior to-gastfication the coal must be pulverized to
a partical size of 90 um or less. In the pilot plant operation, the coal is
pneumatically transported to a heat exchanger where it is heated to a tempera-
ture of 350 to 400?C. Next the coal undergoes pre-treatment with steam for 2
to 3 seconds in a reaction tube to reduce its caking tendencies. Coal pre-
treatment may not be necessary since pilot plant operation seems to indicate
that the coal's caking properties do not affect gasifier operation.
The pulverized coal is injected through a gasification burner at the top
of the reactor with preheated air or oxygen and 600?C steam (see Figure 2).
Since the gasifier operates at atmospheric pressure, the use of a lockhopper
feed system is not required. The coal entering the burner undergoes rapid
devolatilization and partial combustion to form a tar free gas and coke. Due
to the rapid reaction rates the coke which is formed has a large surface
area. This large surface area promotes desulfurization. The resulting crude
gas and coke exit the bottom of the gasifier at 1200 to 1300?C and the coke is
separated out for use as a boiler fuel. Waste heat is recovered from the hot
crude gas at a level of 100 to 150?C. Heat is also recovered from the hot
exiting coke stream. The gas can then undergo further treating depending upon
application. At a coal conversion rate of 50% to 60%, approximately 80% of
the gasifier steam requirements can be produced from waste heat recovery.
Coke formed in this process has been tested as a boiler fuel in separate
tests. It's combustion characteristics show very good ignition and burning
qualities comparable to hard coal firing.
A technology fact sheet for this process is shown in Table 1.
t
 M M M M . f veccg Re 2 2/2iJIA-F 3MM14R"00(("18-(bo l? m m m
Crushing
Gasification
F-
Heat
Heat Recovery
Steam Turbine Flue Gas Heat
Electric Process GasTurbine
Pbwer
m
0
x
z Ash Electric Power
0
r
0
0
Figure 1. VEW-COAL CONVERSION PROCESS
Dust Sulphur
Purification 7
Gas
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Coal, Air
Steam. Air (600?C) (02)
Air, Gas
Burner for Starting
t
Steam forCooling
Steam for Cooling
t To Combustion Chamber
To CharBunker
Figure 2. VEW-COAL CONVERSION PROCESS REACTOR
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I N S T I T U T E O F G A S T E C H N O L O G Y
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Table 1
Part 2 TECHNOLOGY FACT SHEET: VEW Coal Conversion Process
NEW TECHNOLOGY (PROCESS DESCRIPTION AND GOALS):
The VEW process has been developed primarily for combined
cycle applications. This partial gasification, entrained flow gasifier is claimed to significantly reduce the
sulfur content of the by-product coke thereby making it more attractive for direct combustion applications.
RELATIONSHIP TO PRIOR TECHNOLOGY (INCLUDING STATE OF DEVELOPMENT OF PRIOR TECHNOLOGY):
CHARACTERISTICS OF THE TECHNOLOGY:
CHARACTERISTICS NEW TECHNOLOGY
PRIOR TECHNOLOGY
PRIMARY OUTPUT (DESIGN CASE)....... sXngas........... @30 to 60 coal conversion-44.4 to 48.5
CO, 1.1 to 8.3 Vol. % CH4
TYPE OF PROCESS .....................................
partial gasification - entrained bed
FEEDSTOCK REQUIREMENTS .............................. all types of coal
OVERALL THERMAL EFFICIENCY (INCLUDING BY-PRODUCTS).. 41% coal to electricity
CARBON CONVERSION EFFICIENCY........................ designed operation 30 to 60%*
OPERATING TEMPERATURE. .............................. 1300>?C
OPERATING PRESSURE .................................. 1 atm
BY-PRODUCTS .............................. .........., coke, sulfur
*
low carbon conversion efficiency due to partial gasification
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Vol.% H29 approx. 44.6 Vol.%
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Table 1 TECHNOLOGY FACT SHEET:
OPEMTING FACILITIES -
A one tonne/hr pilot plant has been in operation at Stockum, West Germany
since 1977. A 15 tonne/hr demonstration plant began operation in 1981.
Plans call for the construction of a 1.8 million tonne/yr commercial
MAJOR FUNDING AGENCY The-Federal Ministry for Research and Technology, West Germany.
The 1 tonne/hr pilot plant cost 9 million DM to construct and the
ANNUAL LEVEL OF FUNDING -
15 tonne/hr demonstration plant cost 37 million DM to construct.
Coal caking tendencies and heat recovery from dust laden eases were originally spen.ac lrnhlPm ara~c
td However, these areas posed no technical problems during pilot plant operation.
is
00
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M OTHER FACTORS AFFECTING OVERALL FEASIBILITY:
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Process Goals
The main objective of the VEW development work is to apply this process
to combined cycle power generation. In addition, the partial gasification
principle is being tested for the production of low sulfur coke. The favor-
able sulfur scavenging effect in the reactor during gasification produces a 70
to 80% sulfur removal rate based on a 50 to 60% coal conversion efficiency.
Relationship to Prior Technology
The VEW gasification process is based on the entrained flow principle of
coal gasification. However, the VEW process is not an extension of a commerc-
ially available gasification process (i.e., Koppers-Totzek).
Operating Facilties
The VEW gasification process is being tested in a 9 million DM 1 tonne/hr
pilot plant in Stockum, West Germany. The pilot plant gasifier which has been
in operation since 1977 consists of a refractory lined vertical shaft of about
12 meters in- length with a diameter of 0.5 meters. The plant does not have
gas purification facilities. Operation of this facility with air and
different coals thus far has proved satifactory. VEW is all being developed
by the company of Stein Miller, Gummersbach, which is responsible for the
gasification plant. The company of Still, Recklingshausen, is responsible for
the gas cleaning system.
In 1981 construction was completed on a 15 tonne/hour demonstration plant
which cost 37 million DM. The plant also began operation in 1981. Future
plans call for the construction of a 1.8 million tonne/yearly VEW combined
cycle plant in 1983. This plant will be located at Gersteinwerke, Lippe and
Emstand. Operation is scheduled for 1985.
Major Funding Agencies
The Federal Ministry of Research and Technology in West Germany is the
major funding source of the VEW process. It does not appear that government
support of this process will continue.
Technical Problems
None reported to date.
I N S T I T U T E O F G A S T E C H N O L O G Y
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8. Bergbau-Forschung Nuclear Assisted Coal Gasification Process
Process Description
Bergbau Forschung GmbH (BF) is developing a nuclear heat assisted coal
gasification [Prototypanlage Nukleare Prozesswarme (PNP)] process for the
production of synthesis gas or SNG. Based on the study of coal kinetics it
was determined that the heat produced from a high-temperature gas-cooled
nuclear reactor (HTGR) is sufficiently high (950?C) to gasify coal. With the
gasifier heat requirements being met by nuclear energy, a substantial
reduction in oxygen requirements and coal could be accomplished.
The heat produced in the HTGR is transferred by a helium circuit at a
temperature of 950?C. For safety reasons, the heat is transferred to a
secondary helium loop which flows through the gas generator and the power
plant. The gas generator which is shown in Figure 1, is a fluidized bed
reactor with a submerged helium heat exchanger. The hot helium flows through
this immersed heat exchanger to supply the heat of reaction. All gas loops
and the gasifier reactor are operated at a pressure of 40 bars.
The gasifier itself is a horizontally mounted cylindrical pressure
vessel. Feed coal below 0.5 mm is fed into the top of the reactor through
lockhoppers. The bed is designed with a perforated trough where steam enters
to provide fluidization. The hot (900?C) helium heat exchanger tubes drop
from above into the fluidized bed. The coal as it is introduced into the
gasifier moves longitudinally through the bed as it is gasified. The ash
which accumulates at the other end as well as entrained solids which are
removed overhead in a cyclone are removed through another set of lockhoppers.
In the pilot plant noncaking coals are fed to the free board while caking
coals are fed directly into the fluidized bed. The difficulty of feeding into
the bed is probably a reason for using this technique only when absolutely
necessary. In large scale commercial operation a pretreatment step may be
required for caking coals. A schematic of the pilot plant process is shown in
Figure 2. A summary of this process is presented in the technology fact sheet
of Table 1.
Conventional coal gasification processes consume 30 to 40% of the input
coal for process energy needs. The remaining 60 to 70% is then available for
B-50
I N S T I T U T E O F G A S T E C H N O L O G Y
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 M ~ M M M vec Re6W21 2/2"1A-FW3M"14R"00{"18-(Jo M M M M
A C1
1 cooling water
Dampf
steam
Figure 1. GAS GENERATOR
Helium
DDS
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5 x 5470
Rohgas/raw gas
Kuhlwasser
Koks
char
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SCHWELKOKS
CHAR
GENERATION
u
HELIUM ERHITZUN
HEATING
STEAM
KONDENSAT
CONDENSATE
DAMPFERZEUGUNG
ROHGAS 0 RAW GAS
ZYKLON GASKUHLUNG GASWASCHE
CYCLONE GAS COOLING GAS CLEANING
ASCHE/ASH
H2O
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Bergbau-Forschung Nuclear Assisted
Table 1 Part 2 TECHNOLOGY FACT SHEET: Coal Gasification Process
NEW TECHNOLOGY (PROCESS DESCRIPTION AND GOALS): The Bergbau-Forschung process is being developed to use the energy
from a high temperature gas cooled nuclear reactor to supply process heat to a steam coal gasification process.
The process utilizes a fluidized bed gasifier which has an internal heat exchange that circulates.900?C helium from
the HTGR. The use of an outside heat source will lower oxygen demand and coal use as well as reduce gaseous pollutant
RELATIONSHIP TO PRIOR TECHNOLOGY (INCLUDING STATE OF DEVELOPMENT OF PRIOR TECHNOLOGY):
No relation to priortclinology.
PRIMARY OUTPUT (DESIGN CASE) ...... syngas.(Yol;%) .....
TYPE OF PROCESS .....................................
FEEDSTOCK REQUIREMENTS ..............................
OVERALL THERMAL EFFICIENCY (INCLUDING BY-PRODUCTS)..
CARBON CONVERSION EFFICIENCY ........................
OPERATING TEMPERATURE ...............................
OPERATING PRESSURE ..................................
BY-PRODUCTS... * ....... 9 .............................
NEW TECHNOLOGY
H2=53%, CO=11%, C02=26%, CH4=10%
Nuclear assisted fluidized bed
Coal less than 0.5 mm
95%
900?C
40 bars
PRIOR TECHNOLOGY
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Table 1 Part 1 TECHNOLOGY FACT SHEET: Bergbau-Forschung Nuclear Assisted
Coal Gasification Process
OPERPtTING FACILITIES -
450 kg/hr pilot plant at Bergbau Forschung Facility in Essen, West Germany.
MAJOR FUNDING AGENCY West German Federal Ministryof Research and Technology/The state of Nordhrein-
Westfalia.
ANNUAL LEVEL OF FUNDING - 3.0 x 106 DM/yr
Ability to utilize caking coals without pretreatment in question. Long residence
time in reactor will require large reactor size.
m OTHER FACTORS AFFECTING OVERALL FEASIBILITY: Coupling of HTGR heat source with fluidized bed coal
gasification process still unproven. Also, productio n of low cost process heat with an HTGR relative to coal
Z still need analysis.
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
1
conversion into gas. To overcome the use of feed coal to supply process
energy, the goal of this development work is to use process heat from an HTGR
as a source of process energy. This offers the following process advantage:
? Better use of coal supplies
? Higher process efficiency as well as nuclear cycle efficiency
? Decrease of gaseous pollutants per unit of gaseous energy produced
? Lower production cost if coal prices are regionally high and nuclear enrgy
is inexpensive.
Relationship to Prior Technology
This technology represents a novel approach to coal gasification and does
not have any relationship to prior technology.
Operating Facilities
Bergbau-Forshung has been developing this nuclear assisted coal gasifi-
cation process since 1969. This research is being conducted at BF's Essen
facilities. Besides coal gasification research, BF is also involved in coal
liquefaction, production of active carbon, flue gas treatment, water
purification, sewage treatment, gas separation by molecular sieves, and coke
and electrode carbon production. The support research laboratories for coal
gasification were predominatly built in the last six years. These facilities
include a high pressure thermogravimetric analyzer, a high pressure wire
screen reactor, a curve-point pyrolysis reactor, a high-pressure hot-stage
microscope, and an externally heat steam-char gasification PDU.
The development of the nuclear assisted coal gasification unit has
proceeded in a step wise manner since 1969. The reaction kinetics of steam
and hydrogasification of different coals and chars using particle sizes
smaller than 2 mm has been investigated since 1969 in a fixed-bed differential
sweep gas reactor at temperatures up to 1000?C, total pressures up to 70 bar
(70 x 105 N/m2), and various steam-hydrogen mixtures. A small-scale pilot
plant has been operated since 1973 at 40 bar (40 X 105 N/m2) using an inter-
nally heated fluidized bed. This unit processes up to 5 kg/h and gives result
concerning reaction kinetics, gas composition, reaction heat, fluidized bed
density, and heat transfer under quasi-realistic conditions.
In the mid-70's BF started construction of a 450 kg/hr pilot plant which
started operation in 1979. The pilot plant gasifer is 0.9 m in inside
I N S T I T U T E O F G A S T E C H N O L O G Y
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
diameter and approximately 4 meters in height. The overall pilot plant
structure is about 60 ft tall by 40 ft square. Since the beginning of pilot
plant tests in 1979, they have logged over 7525 hours of coal gasification.
The longest run lasted for 40 days in 1979. A typical raw gas is reported to
contain CH4=14%, C0=14%, H2=51%, and C02=21%. BF claims 95% carbon conversion
but with very long residence times (in excess of two hours). The steam
distributor is made up of an incolloy 800 pipe with holes pointing downwards.
Major Funding Agencies
The nuclear assisted steam coal gasification process is being jointly
developed by the following companies:
? Bergbau-Forschung GmbH, Essen, in cooperation with:
? Gesellschaft fur Hochtemperaturreaktor-Technik GmbH
? Hochtemperature-Reaktorbau Gmbh
? Kernforschungsanlage Julich GmbH
? Rheinische Braunkohlenwerke AG
The major funding agency for this project is the West German Federal
Ministry of Research and Technology as well as the State of Nordrhein-
Westfalia. The pilot plant was designed and built at a cost of DM 13.5 X 106
and the current annual operating cost is DM 9 X 106. The project manger is
Dr. van Heek who also directs what is called the pyrolysis laboratories
(including devolatilization and char gasification work) and the materials
testing laboratory. The annual costs for these two support research
activities are about DM 2.0 x 106 and DM 4.0 X 106, respectively. This
research group has a total of 30 to 35 people out of the DM 15 X 106 annual
budget. The group spends nearly DM 4.68 X 106 per year towards salaries and
the rest towards equipment, materials, and supplies and utilities.
Technical Problems
The fluidized bed gasifier used in this process is very sensitive to the
caking properties of the coal feedstock. In addition, the required long resi-
dence time in this reactor will require a large reactor size to attain the
desired coal throughput. Finally, it has yet to be proven whether or not two
different processes such as the HTGR and the fluidized bed coal gasification
process can operate in harmony given their inherently different operating
conditions.
1
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I N S T I T U T E O F G A S T E C H N O L O G Y
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WEST GERMAN COAL GASIFICATION
9. Rheinbraun Nuclear Assisted Lignite Hydrogasification Process
Process Description
The Rheinbraun process is being developed as part of the West German
Prototype Plant Nuclear Process Heat Project (PNP). In this process, a
fluidized bed coal gasifier is used to hydrogasify coal. Unlike conventional
hydrogasification techniques that react a part of the coal input to produce
the required hydrogen, the Rheinbraun process reforms part of the end-product
methane to produce hydrogen. The heat for the endothermic reforming reaction
is produced by a high temperature gas-cooled nuclear reactor (HTGR). A flow
diagram of this process is shown in Figure 1.
Figure 1. RHEINBRAUN HYDROGASIFICATION PROCESS
I
In the pilot plant arrangement shown in Figure 2, low sulfur coal is
crushed and fed into the fluidized bed gasifier via a lockhopper. Prior to
entering the gasifier the coal powder is also dried and pretreated to reduce
caking properties. The hydrogasifier reactor operates at a pressure of 65 to
100 bars and a temperature of 1000?C with a coal throughput capacity of up to
400 kg/hr of raw brown coal. The gasifier consists of a steel pressure vessel
with an alumina lining.
I N S T I T U T E O F G A S T E C H N O L O G Y
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
1
 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
Because of the long residence time required for complete gasification,
only a fraction of the coal is gasified. This fraction of gasified coal is in
the range of 50 to 70% of the coal input. The resulting char is removed from
the gasifier via lockhoppers and can be used as a boiler fuel to raise steam.
Hydrogen for the hydrogasification reaction along with steam and oxygen are
injected into the bottom of the gasifier. These reaction products also pro-
vide the fluidization in the bed. A synthesis gas with a methane content of
30 to 50% can be produced in this gasifier.
In a commercial operation, the hot raw gas exiting the gasifier is cooled
and separated into its prime constituents. The carbon dioxide and hydrogen
sulfide are scrubbed out and processed while the hydrogen is recycled back
into the gasifier. Part of the captured methane exits the plant as product
while the remainder is sent to a catalytic steam reformer. The heat required
for the endothermic steam reforming reaction is supplied by an HTGR via heat
exchange with a 900?C loop of circulating helium. This step produces the
additional hydrogen required by the hydrogasification reaction. The resulting
synthesis gas is then mixed with the carbon monoxide from the gasifier and
reacted in a water gas shift reactor to produce additional hydrogen. This
hydrogen is then sent to the gasifier and the carbon dioxide is scrubbed and
released to the atmosphere.
The above described process can be modified by the addition of a High-
Temperature Winkler gasifier. This gasifier would eliminate the use of
nuclear heat by consuming the unreacted char from the hydrogasifier to produce
the additional hydrogen required by the process. A technology fact sheet
describing the HTGR coupled process is shown in Table 1.
The goal of this project is to produce SNG with an integrated HTGR and
hydrogasification process. This offers the advantage of reducing gaseous
pollutants and conserving coal reserves. This project will demonstrate the
use of a novel gasifier, methane steam reformers, and integrated HTGR.
Relationship to Prior Technology
Although the concepts of each major component are not new, the designs
and application of these concepts have not been demonstrated by any previous
process.
 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
Operating Facilities
Successful operation of this concept has been demonstrated at the 0.4
metric ton/hr process development unit operated by Rheinbraun since 1975 at
their Wesseling facilities. The nuclear heat steam reformer concept has been
tested since 1972 in a single tube design by Kernforschungsanlage (KFA) in
Julich. Construction is underway of a one-tenth scale, 25 ton/day, pilot
scale gasification plant and a 30 tube nuclear assisted methane steam reformer
reactor.
Major Funding
The nuclear assisted lignite hydrogasification process is being developed
by Union Rheinische Braunhoklenwerke AG (Rheinbraun) in association with the
Gesellschaft fur Hochtemperatureaktor - Technik mbH, Hocktemperature -
Reactorbau GmbH, Kernforschungsanlage Julich GmbH, and Bergbau-Forschung
GmbH. Funding for this project is being sponsored by the Federal Ministry of
Research and Technology (FMRT) as well as the State of Nordrhein-Westfalia.
Total cost of the project is 150 million DM through the semi-commercial stage
of development. Seventy-five percent of this cost is sponsored by the FMRT.
Technical Problems
Blockages occurred in the upper gasifier section when agglomerates formed
during operation on caking coals. This problem can be overcome by the
selective use of coals or the pretreatment of caking coals.
Capital Costs
The capital cost for a nuclear assisted lignite hydrogasification fac-
ility consuming 17.4 X 106 metric tons/yr of brown coal is shown in Table 2.
This facility would cost 3,615 million DM in 1976 DM. This cost includes a
3000 MW HTGR to supply process heat.
I N S T I T U T E O F G A S T E C H N O L O G Y
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.Table 2. CAPITAL COST OF THE RHEINBRAUN INTEGRATED HTGR-HYDROGASIFICATION
PROCESS (1976 Basis)
HTGR Thermal Output
3000 MW
Utilization Factor
7500 hr/yr
Coal Input
17.4 X 106 tonne/yr
Gas Output
2.96 X 109 m3/yr
Char Output
2.13 X 106 tonne/yr
Electricity Production
877 X 106 kWhr/yr
Project Cost
3,615 million DM
Coal Price
7 DM/G cal
I N S T I T U T E O F G A S T E C H N O L O G Y
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WEST GERMAN COAL GASIFICATION PROCESS
10. KHD Kloeckner-Humbolt-Wedag Molten Iron Coal Gasification Process
Process Description
The Kloeckner process is based on the use of a molten iron bath for
gasification. The gasifier is a refractory lined vessel equipped with gas
cooled tuyeres. Crushed coal sized up to 3 mm is injected into the bath
through the bottom tuyeres which are similar to the Q-BOP technology proven by
the steel-making industry. The coal which has been dried to a moisture con-
tent of 1.5% is also mixed with lime of the same partical size before being
pneumatically injected with nitrogen into the molten bath. Additional gasi-
fying agents such as oxygen, air steam, or CO2 can also be injected simul-
taneously into the bath through the special tuyeres. A diagram of the tuyeres
design is shown in Figure 1.
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Figure 1. TRIPLE-FLOW TUYERES DESIGN
To protect these tuyeres which are made of a high grade steel, from the
excess heat, a cooling gas is blown through the outermost annular gap. This
gas can be propane, methane, carbon dioxide, steam or purified recycled syn-
thesis gas. A simplified flow diagram of the Kloechner process is presented
in Figure 2.
As the coal enters the 1350 to 1400?C iron bath it undergoes rapid
devolitization and cracking of the hydrocarbons. The carbon in the coal is
dissolved in the bath and the coal ash rises to the bath surface. Any sulfur
I N S T I T U T E O F G A S T E C H N O L O G Y
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in the coal reacts with the iron to form FeS. This FeS reacts further with
the injected lime to form CaS. The calcium sulfide and ash are removed from
the gasifier as liquid slag. The dissolved carbon in the iron bath reacts
with the injected gasification agents such as oxygen and steam to produce a
synthesis gas consisting of CO and H2. The high process temperature produces
a gas which is relatively free of carbon dioxide, volatiles, and sulfur com-
pounds. When the gasifier is operated under continuous conditions the carbon
content of the iron bath is approximately 3.5%. The carbon conversion effici-
ency of the process is 98% with a raw gas production of 2100 m3 per ton of
coal. Table 1 shows the operating characteristics of the Kloechner pilot
plant tests. The technology fact sheet for this process is shown in Table 2.
Figure 2. FLOW DIAGRAM OF THE KLOEKHNER MOLTEN IRON
GASIFICATION PROCESS
I
I N S T I T U T E O F G A S T E C H N O L O G Y
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Table 1. OPERATING CHARACTERISTCS OF THE KLOECKNER PROCESS
Coal 250 - 400 kg/hr ? tonne Fe
Oxygen 0.58 m3/kg of coal
Propane 0.1 m3/m3 02
Transporting Gas (N2) 0.1 m3/kg of solids (coal & lime)
Coal Composition: Fixed carbon 67.50%; sulfur 1.0%; vola-
tile matter 22.0%; ash 8.0%; moisture
1.5%, net heating value 7,5000 kcal/kg
Raw Gas Composition CO 65 +; H2 25-30%; C02 > 0.3%
S < 20 ppm; CH4 > 1%.
The objective of the Kloeckner-process is to produce a synthesis which is
relatively gas free of unwanted components for use in the chemical industry or
as a combined cycle fuel. This is accomplished by the use of a molten iron
gasifier equipped with specially designed tuyeres similar to those used in the
Q-BOP technology of the steel making industry. This process also offers the
stated advantage of being economical due to the high gasifier throughput and
the dramatic reduction of gas cleanup facilities.
Relationship to Prior Technology
The concept of molten iron gasification was first patented and tested by
the Applied Technology Corporation (USA) in the early seventies. However, the
Applied Technology Corp. concept introduced the coal and oxygen feedstock
through lances from the top into the molten iron bed. Cooling, material
problems, and mechanical stability of these lances made this approach techni-
cally unfeasible. In 1978 Kloeckner acquired the patent rights and modified
the process to use the bottom tuyeres feed arrangement.
Operating Facilities
A 2.4 tonne/hr process development unit with an iron bath capacity of
6 tonnes has been tested at the Maxhutte Research facilities in Bavaria (West
Germany). A larger 10 tonne/hr facility is scheduled to begin operation at
the end of 1982 in the Ruhr region. Testing in this 20,000 m3/hr capacity
1
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
gasifier will end in 1984. This facility is capable of operating at 10 bars
and at a temperature of 1350 to 1400?C.
Major Funding Agencies
The initial process development unit testing was funded by the Maxhutte
Research Institute. The 15 million U.S. $ funding for the 10 tonne/hr pilot
plant is being sponsored by the State of Nordrhein Westfalia. Future plans
call for the construction of a molten iron coal gasification commercial size
plant in 1985 in Bremen, West Germany. This plant will have a 200 tonne
molten iron reactor with a rated capacity of 200,000 m3/hr of synthesis gas.
The Federal Ministry of Research and Technology will fund 40% of the 750
million D.M. construction cost.
Technical Problems
Much of this or similar technology has been commercially proven in the
steel industry and as such no major technical problems have been reported.
Capital Costs
The capital costs for a direct reduction, iron ore facility incorporating
a Kloechner molten iron gasification process has been published by KHD Humbolt
Wedag AG. Figure 3 presents the layout of the iron ore reduction/gasification
system. The operating and cost specifications for this plant are shown in
Table 3. The gasification plant investment cost of 60 million D.M. was
assumed to be on a 1980 basis. However, no basis was presented in the
literature. Based on varying coal costs and the information presented in
Table 2, Kloeckner has calculated the cost of producing reduction gas. These
costs are presented in Figure 4, and vary from a gas cost of approximately
4/m3 at a coal cost of $25/tonne to a gas cost of approximately 8.74~/m3 for
a coal cost of $125/tonne.
I N S T I T U T E O F G A S T E C H N O L O G Y
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
Table 3. OPERATING AND COST DATA FOR A KLOECKNER MOLTEN IRON
GASIFICATION PLANT CONNECTED TO A IRON ORE REDUCTION FACILITY
Plant Capacity 450000 t/a Fe
Reducing Gas Demand
1500 m3 N/t
by coal gasification
-750 m3 N/t
by gas recycling
1
Operating Time 8000 h/a
Gasification Plant
Investment 60 mil. DM
Capital Change 14 %/yr
gas price
cts/rrr3
reduction gas from natural gas x'
, 5,00 /MM Btu
I x 4,40 S /MM Btu
3,80 S /MM Btu
reduction gas from coal
25 50 75 100 125 coal price
Figure 3. REDUCTION GAS COSTS USING THE KLOECKNER PROCESS
I N S T I T U T E O F G A S T E C H N O L O G Y
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1
1
WEST GERMAN COAL LIQUEFACTION PROCESS
11. Ruhrkohle/Veba Oil Catalytic Hydrogenation Coal Liquefaction Process
Process Description
This catalytic hydrogenation process is a modification of the Bergius-
Pier (IG Farben process) hydrogenation technology which was a commercial
reality in Germany prior to 1945. Based on this working experience which was
documented by BASF, Ruhrkohle (RAG) and Veba Oil in cooperation with KRONIG
made the following modifications to the IG process:
? A mixture of heavy oil and middle distillate oil from the process is used
as a solvent
? Separation of solids and asphaltenes is accomplished by distillation
? The residual material containing solids and asphaltenes is used to produce
hydrogen.
These changes are expected to improve the IG process as follows:
? Reduce the process pressure from 700 bars to 300 bars
? Raise the specific coal throughput by 50%
? Improve heat recovery
? Raise thermal efficiency by 25%
? Reduce process capital costs.
Based on a renewed interest in the IG process by RAG and Veba Oil in
1974, a process development unit was designed in 1975 and completed in 1976 by
Bergbau-Forschung in its Essen Laboratory. Test runs are reported to last a
minimum of 4 days with an estimated total of over 21,000 hours operating
experience gained since startup.
The process development unit has two 11 liter reactors, designed for
400 atm and 500?C with a rate coal throughput of 250 kg/day. About 600 m3/day
of hydrogen is supplied from a huge bank of (about 200) 140 psig cylinders.
The recycle gas flow rate is about 1000 m3/day. A wide variety of coals are
tested in the pilot plant with carbon conversions ranging up to 95% and hydro-
gen consumption estimated at about 6% of wt. by d.a.f. coal.
The PDU involves mixing ground coal (less than 0.2 mm) with recycled oil,
and a disposable catalyst (red mud) of 2% to 3% by weight of coal, resulting
I N S T I T U T E O F G A S T E C H N O L O G Y
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in a 45% solids content slurry. The slurry is pressurized in stages (using
Moyno pumps to 6-8 atm and high pressure piston pumps to 300 atm), with high
pressure hydrogen, preheated to 400 to 450?C, and transferred to the hydro-
genation reactor. The isothermal reactor contains an agitator oscillating
axially. The resulting products are a mixture of oils of a wide boiling
range, gases, water and residual char. These are separated in a series of
vacuum flash distillation columns to produce ash and sulfur-free distillate
(50% by weight of d.a.f. coal), byproduct gases (20%), sour water (10%), and
residual char. Part of the distillate oil and hydrogen, separated by
scrubbing the product gases, are recycled to the front end. BF observed that
some H2 in the recyled hydrogen aids liquefaction, possibly by activating the
catalyst by sulfurization. However, use of pyrite in catalyst resulted in
scale deposition in the slurry preheater.
A typical catalyst composition is given at 30% Fe203, 40% A1203. 2-6%
Ti02, and the rest silica and other inerts.
In November 1977, RAG and Veba Oil decided to jointly sponsor the con-
struction of a 200 tonne/day pilot plant in Bottrop close to RAG's "Prosper"
coking plant which supplies some of the support facilities. The flow diagram
of this 200 tonne/day pilot plant, which is similar to the PDU operated by BF,
is shown in Figure 1. Construction of the pilot plant started in 1979. The
plant output will be processed in a oil refining pilot plant at Veba Oil's
Scholven facility nearby. The fresh hydrogen required by the process will be
obtained from a nearby hydrogen pipeline operated by Chemische Werke Huls AG.
In the pilot plant operation coal is delivered by train or truck and
stored in the raw coal bins. From here the coal is sent to the grinding/drying
mill where the coal is reduced to a grain size of less than 1 mm and a residue
moisture of 0.5% by weight. Any airborne coal dust is recovered by a cyclone
and electrostatic filter. All dried and sized coal is sent to the dry coal
storage bin at a temperature of 120?F. The catalyst (red mud) is also stored
in bins before it undergoes drying and sizing.
From the dry coal bins the coal is sent to the mixing section where it is
admixed with the solvent, which is a 40 to 60 weight percent ratio of medium
oil (boiling temperature range 200 - 325?C) and heavy oil (boiling temperature
greater than 325?C) which is recyled from the process. At this point the coal
is wet ground with the catalyst to less than 0.2 mm size. This slurry, which
1
 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
1
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I N 5 T I T U T E O F G A S T E C H N O L O G Y
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
contains 40 weight percent solids, is pumped to a pressure of 300 bars and
mixed with recycle gas, make up hydrogen, and recycled reacted slurry. This
mixture is then heated to 380?C by heat exchange and then to the 425?C reactor
inlet temperature by a gas fired heater.
The hydrogenation reactions are performed in a series of three reactors.
The slurry is introduced in the topmost reactor and flows downward through the
remaining two reactors. Due to the exothermic nature of these reactions the
reactors are maintained at 475?C through the injection of cooled recycle gas
at various levels of the reactors.
Upon leaving the last reactor at 450?C, the slurry gas mixture enters a
separator where hot gases and vapors are drawn off the top and liquids and
solids are removed from the bottom. The gaseous and vapor section is cooled
by exchange with the incoming feed and then fed to the distillation section
where naphtha, middle oil, and heavy oil are extracted. The resulting gas
phase is oil washed to remove gaseous hydrocarbons, H2S, C02, CO, and N2. The
remaining process gas is partially recycled to the feed/product heat exchanger
and the rest is sent to the flash evaporation unit or sold as product to the
cokery.
The liquid and solids separator bottoms are flashed to separate the gas
and vapor before it is sent to the vacuum distillation unit. The slurry
contains about 50 weight percent solids after distillation. If this weight
percent ratio cannot be maintained the slurry is heated and compressed to 50
to 100 bars with hydrogen and further processed in a high pressure flash unit.
The residue is sent to a granulation units where it is shaped into tablets for
further processing outside the plant. The excess oil which is not recycled is
then sent to a refinery for further processing. The feed and projected pro-
duct slate for this pilot plant are shown in Table 1. The technology fact
sheet for this process is presented in Table 2. The refined product specif i-
cations after coal-oil processing are shown in Table 3.
Process Goals
The objective of this process is to improve the efficiency, operating
characteristics, and product slate of the old IG Farben Process. New equip-
ment which will be tested in the pilot plant includes the slurry feed piston
pumps, feed/product preheaters, reactors, bottoms separation flash units, and
process oil upgrading.
I N S T I T U T E O F G A S T E C H N O L O G Y
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
Table 1. FEED AND PROJECTED PRODUCTS OF BOTTROP PILOT PLANT
Coal., m.a.f .
Feed
t/d
200
Hydrogen H2, makeup
m3/d (ft3/d)
220,000
(7.8x106)
Process water
t/d
41.2
Catalyst (Fe203)
t/d
4.0
Power
kWh/d
108,000
Pro j ected Products
/d
t
61.3
Naphtha (stab.) I.B.P. 200?C (390?F) t/d
24.4
Middle oil 200-325?C (390-620?F) t/d
Heavy oil >325?C (620?F) t/d
Residue t/d
I N S T I T U T E O F G A S T E C H N O L O G Y
Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
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I N S T I T U T E O F G A S T E C H N O L O G Y
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1
Specif ications
Gravity at 15?C g/cn3
Sp. g. (59?F)
Aromat. C, wt Z
Naphth. C, wt %
Paraff. C, wt Z
Carbon, wt 2
Hydrogen, wt Z
Sulfur, wt 2
Nitrogen, wt 2
Oxygen, wt 2
Base value, mg NH3/1
g/cc)
CFPP, ?C (OF)
Flash point, ?C (?F)
Viscosity, mm2/S
Centipoise
Heating value, kJ/kg
Btu/lb
Fuel value, kJ/kg
Btu/lb
Phenols, wt 2
Bases, wt 2
Benzene, wt %
Toluene, wt Z
Ethyl benzene, wt Z
Xyl ene , wt 2
ROZ (clear)
Boiling range
Initial boiling point,
?C(?F)
10 vol 2, ?C(?F)
50 vol 2, ?C(?F)
End point, ?C(?F)
Light oil Middle oil
Crude Refined Crude Refined
0.865
0.865
46
10
43
0.827 0.990 0.993
0.827 0.990 0.993
31.6 64 50.0
27 28 31.3
41.4 28 18.7
85.25 87.80 87.40 88.5
11.15 12.25 9.10 10.75
146 ppm 2 ppm 0.60 0.23
0.24 2 ppm 0.60 0.23
3.5 0.1 3.0 0.4
1100 3 8800 2900
1.1x10-3 3x10-6 8.8x10-3 2.99x10-3
-26 (-15) -47 (-53)
93 67
3.1 2.1(50?C)(12?F)
3.1 2.1
41,000 43,000 38,500 40,900
17,625 8,485 16,550 17,585
43,000 46,000 40,500 43;250
1,455 19,775 17,410 18,595
3.9
5.3
2.0
4.4
76 (169) 43 (109) Z12 (414) 169 (336)
102 (216) 81 (178) 225 (437) 205 (401)
158 (316) 149(300) 253 (487) 244 (471)
206 (403) 211(412) 324 (615) 303 (579)
r
s
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1
I
Relationship to Prior Technology
The Ruhrkohle/Veba Oil Process is based on the Bergius Pier Process which
was further developed into the IG Farben process in the 1930's. The first IG
Farben Process plants started in 1936 and 1939 with an output capacity of
680,000 tons per year of gasoline and middle distillates. Production capacity
using this process reached 4 million tons/year before complete shutdown in
1945.
Operating Facilities
In 1974 RAG and its subsidiary STEAG AG started evaluating coal liquefac-
tion technologies with the goal of selecting the best process for further de-
velopment. In the same year Veba-Chemie AG began a research program which
involved the hydrogenation of coal in a continuously operating bench-scale
autoclave unit. A larger process development unith with a capacity of
250 kg/day was designed in 1975 and completed in 1976 by Bergbau Forschung it
its Essen facility. In 1977 Ruhrkohle and Veba Oil formed a joint venture to
construct a 200 tonne/day pilot plant in Bottrop. The design and construction
of this plant, which was completed in 1981, was based on the operating data
collection by BF.
Major Funding Agencies
Early work by both Ruhrkohle and Veba was sponsored by the Ministry of
Ecnomics of the State of Nordrhein-Westfalia. The Begrbau Forschung process
development unit was completed for a total cost of DM 3 X 106. The operating
costs for this PDU have been around DM 3 X 106 per year. This research was
also partially funded by the State of Nordrhein-Westfalia. The construction
investment cost for the 200 tonne/day Bottrop pilot plant amounts to nearly
200 million DM. The operating cost for this plant is expected to cost another
200 million DM during the three-year demonstration testing period. The State
of Nordrhein-Westfalia sponsored 90% of the construction phase costs and will
sponsor approximately 80% of the operating phase costs. A decision to proceed
with a larger demonstration size plant will be made by the West German
Ministry of Science and Technology by the end of 1982.
Technical Problems
Although no technical problems have been reported, a number of new
approaches being tested in the pilot plant still need to be proven. These new
 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
technologies are the slurry feed pumps, preheaters, reactors, vacuum distilla-
tion units, and upgrading coal-oil equipment.
The slurry pumps will be driven by SCR controlled AC motors and operated
in parallel. For test purposes. the pumps will be operated in the piston ver-
tical and horizontal position. These pumps will be equipped with remote valve
boxes with inlet and outlet valves of different design and materials to test
for performance in regard to corrosion and erosion. The external valve box
will also provide for abrasion free action, although this will result in oil
loss which will be made up by clean oil on each stroke.
The preheater section will consists of two different units that are each
rated to handle the full plant capacity. The first heater is a shell and tube
design which is similar to the ones used in the IG Farben process until 1945.
At the expense of size and cost this heater provides for gradual feed heating
to prevent coking. The second preheater utilizes a radient furnace design
which is frequently used in oil refineries. The higher heat transfer charac-
teristics of this design increase the possibility of coking but decreases the
unit size and cost.
The next new area to be tested at the Bottrop plant is reactor construc-
tion. The old IG Farben process coal hydrogenation reactors were made from
forgings whereas the new reactors will be of a new multilayer design. Insula-
tion will be accomplished using a layer of refractory brick which will be
protected from abrasion by a stainless steel liner. A new type of reactor is
in the design phase which will utilize the multilayer concept but the required
insulation will become an integral part of the multilayer wall.
The fourth area of testing will focus on slurry bottoms separation using
a combination of flash evaporation and vacuum distillation. These methods
will be employed to maximize volatile matter separation from the hot bottom
slurry which contains 65 to 75% oil.
The last area of investigation will determine the best upgrading approach
for the coal derived oil. The raw oil from the coal liquefaction process
varies from mineral oil in many respects. These include the distribution of
boiling range fractions, sulfur, nitrogen and oxygen contents, aromatic and
density properties. A comparison of the coal-oil properties to "Arabian
light" as shown in Table 4.
1
i
1
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Table 4. COMPARISON OF SOME CHARACTERISTIC PROPERTIES OF OIL FROM COAL AND
ARABIAN LIGHT
spec. gravity at 15?C
API gravity
carbon
hydrogen
sulfur
nitrogen
oxygen
H/C ratio
gasoline, ibp-200 ?C
middle distillate, 200-325?C
vacuum gasoil, 325-500?C
vacuum residue, 500 ?C
oil from coal
Arabian light
g/ml
0.950
0.856
17.5
34
% wt
86.6
85.5
% wt
9.05
12.6
% wt
0.1
1.7
% wt
0.75
0.2
% wt
3.50
--
1.26
1.77
% wt
22
23
% wt
70
23
% wt
8
28
% wt
-
25
Capital Costs
Veba Oil has published the cost of a liquefaction plant capable of both
liquefying coal and for processing heavy refinery oil residues into light oil
products such as gasoline. The plant would have a coal capacity of 3.7 million
tonnes per year with a rated output of 1.95 million tonnes per year of liquid
product. The overall efficiency would probably be nearly 55%. The liquid
product would consist of 250,000 tonnes per year of LPG, 850,000 tonnes per
year of gasoline, and 850,000 tonnes per year of heating oil. According to a
Veba study, this plant would require a capital investment of 6 billion DM on a
1981 basis. Design and construction of this facility would take eight years
which with interest charge, would bring the final plant costs to 7.8 billion
DM.
22(3)/process/ER
I N S T I T U T E
 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
WEST GERMAN COAL LIQUEFACTION PROCESS
12. Saarberg Catalytic Hydrogenation Coal Liquefaction Process
Process Description
The Saarberg coal liquefaction process is a modification of the com-
mercially proven IG Farben technology. The main differences are a lower
pressure, reduced hydrogen consumption, improved heat transfer techniques and
more efficient products separation equipment. A process flow diagram of the 6
tonne/day pilot plant operated by Saarbergwerke AG in Volkingen-Furstenhausen,
West Germany is shown in Figure 1.
Figure 1. FLOW DIAGRAM OF THE 6 TONNE/HR SAARBERG PILOT PLANT
In this pilot plant the coal is first crushed and dried to a particle
size of less than 3 mm before it is mixed into a slurry. Only high quality
coal with an ash content of less than 15% can be used in this process. The
dried coal powder is then fed into a ball mill where it is wet ground to a
maximum particle size of 0.1 mm with recycled solvent, and catalyst. The
catalyst consists of a mixture of iron based sulfates (FeSO4), red mud (Fe203),
and-sodium sulfide (Na2). The catalyst cost was quoted as approximately 54/lb
B-82
I N S T I T U T E O F G A S T E C H N O L O G Y
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in 1981 dollars. The catalyst mixture composition and quantity depend on the
hydrogen pressure used in the process as well as the hydrogenation reactivity
of the coal feedstock. With some coals that have a sufficient autocatalytic
effect, the additional amount of catalyst that must be added may be negli-
gible. This may be true of many non-German coals that have a high pyritic
sulfur content. This feature makes the Saarberg coal liquefaction process
flexible since the optimal hydrogenation conditions can be adjusted for by
changes in the catalyst composition and process pressure. The slurry paste
contains approximately 60 weight percent coal, FeSO4 (1.21% of maf coal), red
mud (2%), sodium sulfide (1%) and the remainder is chiefly distillate oil
which is recycled from the vaccuum distillate section.
This thick slurry paste is then sent to a high pressure injection pump
before it is mixed with hydrogen and heated by a unique preheating section.
The preheat section is designed to heat the paste by injecting hot condensed
oil, which is obtained from the hydrocarbon vapors leaving the top of the hot
separator feed section. The hot diluent oil, which was collected in an inter-
mediate catchpot before injection into the preheat section, helps to heat the
slurry and reduced the coal content to 50 weight percent. The paste is
further diluted and heated by a second hot oil injection preheat section. By
this means much of the exothermic heat generated in the hydrogenation section
can be captured in the preheat injection oil and recycled in the process. The
slurry which leaves the preheat section has a temperature of 400?C and a coal
weight of 38%.
The initial boiling point of the diluent intermediate entering the pre-
heat section is above 200?C. The remaining oil vapors which enter the preheat
section and are condensed in heat exchange with the slurry has a boiling point
of about 400?C. This internal hot oil recycle is about the same weight as the
incoming fresh coal and helps to eliminate the need for atmospheric distilla-
tion of the oil fraction. One benefit of this method is that the oil recycle
fraction is kept at a high temperature and pressure. Although this method
reduces heat exchanger requirements, a bundle heat exchanger is still neces-
sary for slurry preheat. A diagram of this techngiue is shown in Figure 2.
Further attempts to simplify the process led to another idea which eli-
minates the heat exchangers needed for preheating the slurry. In this method
the coal slurry is mixed directly with the overhead vapors from the hot sep-
arator. For this purpose the coal slurry is passed to a mixing zone which is
B-83
I N S T I T U T E O F G A S T E C H N O L O G Y
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B-84
I N S T I T U T E O F G A S T E C H N O L O G Y
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also entered by the hot separator vapors. The vapors are cooled by the coal
slurry, the heavy oils are condensed, and the slurry is preheated and diluted
by the condensed oil. This also permits the use of coarser coal. Since the
coal slurry is being transported only in tubes, there is little danger of
sedimentation.
Coal particle diameters up to 1 mm (0.04 in.) are possibly applicable.
The effluent of the mixing zone is separated into the hot coal slurry which is
fed into the first reactor, and into the overheat vapors containing the pro-
ducts and the excess hydrogen. Application of this intermediate catchpot
results in another advantage. The coal while being heated normally splits off
C02, H20, CH4, etc. These compounds enter the reactor and increase the total
pressure by their partial pressure. Now these compounds are eliminated from
the coal slurry prior to entering the reactor. Also the physically adsorbed
water is stripped off and withdrawn from the intermediate separator overhead.
This preheating system has the advantage that there is a reduction of the
pressure resulting from a higher concentration of the hydrogen in the reactor
and a lower pressure drop of the slurry compared with the use of heat
exchangers. This new method is pictured in Figure 3. Saarberg believes that
this arrangement can help make coal liquefaction simpler and more economic.
The final slurry preheat output is sent at a temperature of 430?C to the
reaction section which consists of four reactors in series which operate at a
pressure of 300 bars. Oil or quench gas in injected into these reactors to
maintain an operating temperature range of 470 to 475?C. The reactor product
is then sent to a separator which produces a bottom containing oil, ash, cat-
alyst and unreacted coal. These bottoms are then reduced in pressure to
50 bars and fed to a vacuum distillation unit. An attempt will be made to
recover energy from this pressure letdown stage via a piston engine device.
The residual material from the vacuum distillation section which contains 50%
solids and 50% bitumen with a melting point of 80?C can be gasified to gener-
ate process hydrogen. Gases which exit the top of the separator are used to
supply part of the feed preheat energy via a concentric tube hat exchanger.
The pilot plant is designed for two different operations to form a dis-
tillate synthetic crude oil. In the first technique coal is processed only by
high pressure hydrogenation. The expected results are shown in Table 1 as
compared to that of the IG Farber process. In the second mode of operation
the feed will also undergo carbonization after hydrogenation. The milder
B-85
I N S T I T U T E O F G A S T E C H N O L O G Y
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hydrogenation conditions of this second mode of operation seems more attrac-
tive from the standpoint of catalyst and hydrogen cosumption. The byproduct
char can then be gasified to produce hydrogen.
In the bench scale PDU the cut point of the atmospheric distillation was
increased from 325?C in the old IG Farben process to 400?C.
Classical
I.G. Process
Saarberg
Modified
(Bench Scale)
Conditions
Pressure, bar
700
300
Temperature, ?C
490
475
Catalyst
iron
iron
CHSV(X)
0.62
0.65
Product Distribution
wt. % of M.A.F. coal)
C1 - C4 hydrocarbons
20.0
15.0
C5 - 200?C distillate
12.1
14.6
200 ?C - 325?C distillate
31.4
30.2
325?C +
17.4
9.4
H209H2, NH3, CO, C02
13.9
11.7
Carbonization:
Coke and Gas
12.5
Vacuum distillation:
Residual oil
--
Unreacted coal;
(3)
5.6
H2 reacted
(7.3)
(5.5)
Total
107.3
105.5
(X)coal hourly space velocity
I N S T I T U T E O F G A S T E C H N O L O G Y
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
This increase reduces the cracking which occurs in the sump phase which
is only slightly selective compared to a fixed-bed catalyst. This lowers the
reaction temperature 10?C which results in the suppression of undesirable
light hydrocarbons which have a higher hydrogen content. These changes have
led to an improvement in liquefaction operating conditions. Hard coal has
been liquefied at a pressure of 285 bars and a hydrogen consumption rate of
5.5 weight percent of the m.a.f. coal. Oil yields have increased to more than
56%. A technology fact sheet for this process is presented in Table 2.
The Saarberg process is a modification of the commercially proven IG
Farben (Bergus-Piers) process. The major project goals of the Saarberg pro-
cess are to reduce the operating pressure to a maximum of 30 bars compared to
the previous 600 bars, reduced hydrogen consumption, and improved heat
economy. This will be accomplished by employing novel slurry preheat con-
cepts, the use of vacuum distillation, novel energy recovery equipment and a
new recycle solvent appraoch.
Simultaneously with the experimental program, Saarberg has followed the
upgrading of the sump phase oil to marketable products through cooperation
with BASF. The main objective is the production of gasoline, particularly of
high-octane blending components. For this purpose the coal oil is first
refined to eliminate nitrogen, oxygen and sulphur. In a hydrocracker the
refined middle distillate is cracked to naphtha. In a subsequent power
former, both naphtha from the hydrocracker and sump-phase naphtha are trans-
formed into high octane gasoline. The gasoline which was produced from the
bench scale oils in the laboratories of BASF had a research octane number
(clear) of about 104.
Relationship to Prior Technology
The Saarberg process is a modification of the IG Farben process. In this
process, crushed and dried coal was mixed with a combination of red mud, FeSO4
and Na2S and recycle oil to form a paste. This was done in rotating mills
with steel balls as grinding elements in the presence of oil to prevent oxida-
tion. The homogenized paste was then fed to high pressure pumps for injection
into the slurry feed heat exchangers. The discharge pressure in the injection
pumps was 700 bars (10,500 psi). In the heat exchangers, slurry and hydrogen
1
 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
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flowed up outside the tubes with overhead vapors from the hot separator
flowing countercurrently inside the tubes. Next, the coal paste and hydrogen
were preheated to about 430?C (806?F) in a gas-fired heater. The gas/slurry
mixture was then passed through four vertical reactors in which coal liquefac-
tion took place at about 480?C (896?F). The exothermic reaction was con-
trolled by the addition of cold recycle hydrogen. The effluent from the last
reactor was separated into an asphaltene-free overhead product and a bottom
product containing the unconverted coal, ash, catalyst, heavy distillate and
residue oil. The gases and oil vapors form the top of the separator passed
through the tubes of the slurry feed heat exchangers and were finally cooled.
While the remaining gas was scrubbed to recycle purified hydrogen back into
the process, the liquid was withdrawn from the cold catchpot and passed to an
atmospheric distillation tower. All of the 325?C+ (617?F+) distillate was
recycled as pasting oil. The main product was middle distillate oil which
normally was converted to naphtha by hydrocracking. The hydrocracked naphtha
and refined sump-phase naphtha were then reformed to high octane gasoline by
the IG DHD (German acronym for pressure hydrogen dehydrogenation) process.
In the old system, processing of the hot separator bottoms consisted of
two operations, centrifugation and carbonization. Both asphaltene-containing
centrifuged filtrate and carbonizer oil were used as pasting media. The
application of more severe hydrocracking conditions meant that the ashpaltenes
produced had to be recycled to the reactors since they could not be used else-
where. In one German plant asphalt was taken out of the sump phase and used
as a binder for weakly caking coals.
This process was first developed in Germany by Friedrich Bergius. In
1910 he received a Nobel Prize for his discovery of the hydrogenation
process. The process was further developed and commercialized by the IG
Farben chemical conglomerate. The first commercial size plant started
operation in 1927 using brown coal as the feedstock at Leuna. By 1943, 12
plants were in operation with a combined capacity of 4 million metric tons of
production.
Operating Facilities
In 1974 Saarbergwerke AG started the develompent of coal liquefaction
technologies. This organization is a conglomerate owned by the Federal and
State Governments (Federal Republic of Germany 74%, State Government of the
I N S T I T U T E O F G A S T E C H N O L O G Y
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
Saar 26%) with interests in mining, industrial and commercial operations.
These include the mining of coking and steam coals, operation of coal prepara-
tion facilities, coking plants, electric power plants, oil and tar refineries,
the manufacture of tools and machines, plastic and plastic products, packaging
and building materials, construction and operation of central heating plants,
sales, distribution, engineering, consulting, insurance, mortgages, and
travel-agency activities.
In 1975 a continuous 10 kg/hr bench scale unit was constructed by
Saarberg for initial testing of improvements to the IG Farben Process. This
work led to the construction of a 6 tonne/day pilot which began operation in
July of 1981. This pilot plant located at Volklingen-Furstenhausen is a 50-50
joint venture of Saarbergwerke AG and Gelsenberg AG, a subsidiary of Deutsche
BP. The new joint venture firm, Gesellschaft fur Kohleverflussigung (GFK),
will operate the pilot plant for a total of three years from commission to
gather data for a full scale commercial plant. The Saarberg liquefaction
plant is located adjacent to the Saarberg/Otto coal gasification pilot plant
and near the Furstenhausen coke works, The Saarland oil refinery, the Fenne
power plant, and the Saarbergwerke AG central laboratories.
Future plans call for the construction of a large demonstration scale
plant that will convert 2.3 million tonne/year of coal into one million tonne/
year of synthetic gasoline and other liquids. This would represent nearly 5%
of, West Germany's current gasoline consumption. A final decision in the con-
struction of this plant will be made in 1983 when pilot plant testing will be
nearly complete. This plant will take four or five years to construct.
Major Funding Agencies
The 6 tonne/day pilot plant at Volklingen-Furstenhausen was constructed
at a cost of approximately $21 million. The West German Ministry for Research
and Technology sponsored 75% of this cost and the State of Saarland contri-
buted an additional DM 1.5 million.
Technical Problems
No technical problems have been reported.
1
1
I N S T I T U T E O F G A S T E C H N O L O G Y
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
Capital Costs
The Saarberg liquefaction process and the Ruhrkohle/Veba Oil liquefaction
process are both modifications of the commercially proven IG Farben Process.
Both processes will operate at similar temperatures and pressures, and both
will have similar efficiencies. Based on equipment modification of both these
processes it would be logical to conclude that they would have similar capital
cost requirements.
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I N S T I T U T E O F G A S T E C H N O L O G Y
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
WEST GERMAN COAL LIQUEFACTION PROCESS
13. Rheinbraun Brown Coal Liquefaction Process
Process Description
The Rheinische Braukohlen Werke AG (Rheinbraun) direct coal liquefaction
process (also known as the HVB process - Hydrierende Verflussigung von
Braunkohle) is based on the IG Farben (Bergius-Piers) process.
In the HVB process, liquefaction of coal is carried out in two stages.
In the first stage (called sump-phase hydrogenation), dry brown coal is cata-
lytically converted into coal oil. In the second stage (gas-phase hydrogena-
tion), the coal oil is converted into motor fuels by conventional oil refining
techniques such as hydrocracking and reforming, or into feed materials for the
chemical industry. Development work has been centered on improvements to
sump-phase hydrogenation stage.
In the sump-phase hydrogenation (see Fig. 1), coal with a grain size of

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Proceedings of the Royal Society of London. The enhanced solvent properties
of polar gases in the supercritical region was demonstrated by Hagenback in
Leipzig in 1901. Work by Frank (1956) investigated the solvent properties of
supercritical water and the investigation of Eisenbeiss (1964) and Weale
(1967) demonstrated the solubility of solids in compressed nonpolar and polar
gases. The solvent properties of supercritical fluids was applied to the
deasphalting of petroleum fractions even before 1950. Work at the National
Coal Board's Coal Research Establishment at Stokes Orchards in the mid-1960's
focused on liquids recovery from coal when heated above 750?F. In 1971 Paul
and Wise in London described the principles of gas extraction as a means of
liquefying coal. This technique has undergone further research by the
National Coal Board in various size test units since the early '70's.
Operating Facilities
All work on the supercritical gas extraction process is being carried out
at the National Coal Board's coal research Establishment facilities, Stoke
Orchard, near Cheltenham. This research has progressed from a benchscale unit
to the construction and operation of a 5 kg/hr continuous feed process
development unit. This unit started operation in 1977. Design of a 25
tonne/day pilot plant have been completed which was to be located at the Point
of Ayr in North Wales. The National Coal Board has recently made the decision
not to construct this pilot plant. It is not certain as to whether or not
research on this process will continue.
Major Funding Agencies
The National Coal Board has sponsored all funding for this process.
However, the NCB has recently announced that it will shelve this process as
well as their Liquid Extraction Process.
The major technical problems facing this technology is scale-up of the
current 0.1 tonne/day PDU to the proposed 25 tonne/day pilot plant.
Capital Costs
Capital costs for a commercial scale facility have not been published.
However, the capital costs for the 25 tonne/day pilot plant have been
estimated at # 14.8 m in 1978.
I N S T I T U T E O F G A S T E C H N O L O G Y
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
BRITISH COAL LIQUEFACTION PROCESS
18. National Coal Board Liquid Solvent Extraction Coal Liquefaction Process
Process Description
The National Coal Board's Liquid Solvent Extraction (LSE) process
utilizes anthracene and coal derived solvents to liquefy coal at a temperature
of 420 to 450?C and at 20 bars pressure. The process actually consists of two
stages, the liquefaction stage, and the hydrotreating stage. In the lique-
faction stage the solvent also acts as a mild hydrogen donor. A process flow
diagram for the liquefaction and hydrotreating sections is presented in
Figures 1 and 2.
In this process run-of-mine coal is fed to the coal preparation section
where it is cleaned and dried. From here it is sent to a pulverizer where it
is reduced to less than 0.2 mm in size. All coals, except for anthricite, can
be used as feed regardless of caking properties. However, extract yields can
vary widely depending on the coal feed. A cleaned fraction of coal containing
approximately 6% ash is then fed to the slurry preparation section. The
remaining diameter coal is sent to a gasifier for hydrogen generation or to
the boiler for steam and power generation.
The pulverized coal entering the slurry preparation section is mixed with
a high boiling solvent, recycled from the hydrocracking unit, in the ratio of
3.5 ton of solvent to one ton of coal. This slurry is then pumped to a pres-
sure of less than 20 bars and sent to the preheat section. This section
consists of a heat exchanger where the slurry is heated by the liquefaction
digester effluent, and a fired heater. The slurry enters the digester at
400?C where it is reacted for 30 to 90 minutes. Under optimum conditions as
much as 95% of the coal's non-mineral matter can be dissolved. During
reaction the solvent acts as a mild hydrogen donor to increase the extract
hydrogen content as much as 2 percent. The effluent from the digester is then
sent to the feed preheat exchanger for cooling and then into a gas/liquid
separation vessel.
In this flash vessel the pressure of the slurry is reduced and the
lighter products are separated. The overhead stream is then cooled to
condense out the C5-250?C fraction and water. The gaseous product can then be
recycled or sold while the liquid bottoms from the condenser are separated to
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420' C
EXTRACTION
PRESSURE
LET-MM
NYDROCRACKER FEED
(EXTRACT
r
r
INKS M
Figure 1. LIQUID SOLVENT EXTRACTION PROCESS SECTION
SOLUTION
LOW PREASUREISEEARATORN
TO FRACTIONATION
bm-
TO FRACTIONATION
_.T
OAS
Figure 2. HYDROGENATION UNIT FOR LIQUID SOLVENT EXTRACTION PROCESS
B-139
O F G A S T E C H N O L O G Y
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
remove the water. The liquid bottoms from the flash vessel, which contain
ash, are sent to a holding tank where they are fed batch-wise into a commer-
cially available filter. The dried filtrate is then sent to the gasifier for
hydrogen generation and the filtered coal extract is sent to storage.
From storage the coal extract is mixed with recycled oil to adjust its
concentration before it is pressurized and sent to the second stage hydro-
treating section shown in Figure 2. This mixture entering the hydrogenation
section is heated to 750 to 840?F and pressurized to greater than 2900 psi.
Recycled impure hydrogen is added to the high pressure liquid, resulting in
two-phase flow through the heat exchange and furnace heating zones. The
hydrogen is scrubbed to remove some of the sulfur, nitrogen and chlorine com-
pounds prior to injection. Make-up hydrogen for the proposed 25 ton/day pilot
plant will be obtained by steam reforming natural gas. However, in a commer-
cial size facility much of the required hydrogen will be generated by a gasi-
fier from recycled filtrate, pitch and dirty coal.
The hot extract and hydrogen gases are passed through a guard reactor to
remove hydrocracker catalyst poisons. Two or more of these units will be
operated in parallel to allow for continuous operation when one unit is regen-
erated.
The primary hydrocracking reactors are of conventional trickle bed
design. A number of catalyst beds are used to limit temperature rise to a
maximum of 45?F in each bed. Intercooling between beds with recycled hydrogen
is also used to limit the temperature rise. Facilities are also included to
regenerate the hydrocracker catalyst.
Effluents from the hydrocracker are cooled by heat exchange with extract
feed and are sent to a high pressure separator. Hydrogen rich gases con-
taining some methane and other gases are recycled to the hydrocracker units
after repressurization and scrubbing. Liquid bottoms are sent to the frac-
tionation system. This system contains a vacuum column as well as atmospheric
pressure crude fractionation units. The principle cuts include gases,
C5(345?F) liquids, 345 to 480?F boiling range liquids and 480?F plus liquids.
The primary product cuts are the two lower boiling range liquids. The 480?F +
cut is used as recycle solvent. Alternatively, 700?F + cut can be produced
and fed to a thermal cracker for production of a premium cut. The material
balance for this process is shown in Figure 3. A process technology fact
sheet summarizing the process is shown in Table 1.
B-140
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T U T E O F G A S T E C H N O L O G Y
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 pproved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
The overall process efficiency and product slate varies with coal feed
type and end product requirements. Tables 2 and 3 present the product mix
when using a bituminous coal or lignitic coal. Table 2 also presents process
efficiencies for different plant configurations when pitch is used as a gasi-
fier feed, SNG byproduct is recycled for process gas, and when both pitch and
SNG are recycled. Overall process efficiencies can vary from 68% in the base
case using bituminous coal to 60.2% when operating on a lignitic coal with 23%
moisture and 20% ash.
Process Goals
The liquid solvent extraction process is being developed as part of a
two-stage coal liquefaction process which involves extraction and product
upgrading for the production of transportation fuels and chemical feedstocks.
Testing in a 66 lb/hr PDU has been performed to determine operating parameters
and performance. Continued research is also being conducted in the extract
upgrading stage. The overall goal of this program is to establish the poten-
tial of this process to operate on a commercial basis in an integrated self-
sufficient mode. Future plans call for the construction of a 25 tonne/day
pilot plant. However, construction of this pilot plant has recently been
cancelled by the National Coal Board.
Relationship to Prior Technology
Development of the Liquid Solvent Extraction (LSE) process started in the
early '60's at the National Coal Board's (NCB) Coal Research Establishment
(formerly called the British Coal Utilization Research Laboratories) in Stoke
Orchard, near Cheltenham. This work centered on the extraction of coal
liquids by anthracene and other tar oils to produce carbon products and elec-
trode coke. A 66 lb/hr process development unit was constructed in the early
'70's to make electrode cokes to be used in the steel and aluminum smelting
furnaces. This PDU is the basis for the proposed construction of a 25 tonne/
day pilot plant. In general this process is similar to the Solvent Refined
Coal (SRC) process which also utilizes a coal derived solvent to dissolve coal
at high temperatures and moderate pressures.
Operating Facilities
All research on the Liquid Solvent Extraction process has been carried
out at the NCB's Coal Research Establishment facilities in Stoke Orchard near
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Table 2. EXAMPLE OF INTEGRATED LIQUEFACTION PLANT AND REFINERY
PROCESSING A BITUMINOUS COAL
Basic
Pitch used
SNG used for
Both pitch
case
as.gasifier
process gas
and SNG used
feedstock
in process
Product Slats (ton/100
tom d.a.
f. coal)
Gasoline
9
91..5
10
11
Diesel and jet fuel
23,
25
27
30
Pitch
6
0
T
0
LPG
4
5
5
6
Substitute natural gas
9-
9-
0
Coal input (ton d.a.f./ba
rrel
0
transport fue
ls prod
uced)
- to process
0.25
0'.25
0.25
0.25
- to gasifier
0.14-
0.10
0.08.
0.04.
- to boiler
0.03.
0.03
0.03.
0.03
Total:
0.4Z
0.3&
0.36
0.32
Barrels liquid fuels/
ton, d.a.f. coal
2.4
2.6?
2.&
3.1
Overall thermal
efficiency (Z)
6&
66-
65
63
Table 3. EXAMPLE OF INTEGRATED REFINERY PROCESSING A LIGNITIC COAL
(23% moisture, 20% ash)
Product Slate (ton/ton d.a.f. coal)
Gasoline 16
Diesel and jet-fuel 19
Pitch 0
Substitute natural gas 0
LPG 2.3
Coal input (ton d.a.f/barrel
transport fuels produced)
to process 0.27
to gasifier 0.07
to boiler 0.05
Barrels liquid fuels/ton d.a.f.
coal 2.6
Overall thermal efficiency 60.2
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Cheltenham. This research has progressed from bench scale studies in the
'60's and early '70's to the construction of a 66 lb/hr PDU in the mid-
'70's. Photographs of the 1st stage PDU liquefaction unit and the 2nd stage
hydrotreating section are shown in Figures 4 and 5, respectively. Future
plans called for the construction of a 25 tonne/day pilot plant to be located
at Point of Ayr, North Wales This project has recently been cancelled.
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Figure 4. SOLVENT EXTRACTION PLANT AT THE COAL RESEARCH ESTABLISHMENT
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Figure 5. HYDROCRACKING PILOT PLANT OF THE COAL RESEARCH ESTABLISHMENT
Major Funding Agencies
The National Coal Board has sponsored all funding for this process to
date. Funding for the $100 million 25 tonne/day pilot plant would have come
from the National Coal Board, Commission of the European Communities
($10 million) and Phillips Petroleum. Recently British Petroleum which was
also a co-sponsor withdrew their support to the project. If pilot plant
testing had been successful, future plans called for the construction of a
1,000 ton/day semi-commerical plant in an overseas market in the late '80's.
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Technical Problems
The two areas of most concern are the solid/liquid separation section and
the catalytic hydrocracking/extract upgrading section. In the filtration
section hot batch filtration has been tested in a 1100 lb/day unit used for
electrode coke production. However, a continuous unit has not been tested.
The batch unit was able to reduce ash content in the liquid extract to below
0.1 percent. Although this type of hot filtration is more expensive than
other separator techniques it is anticipated that the filtration section will
still account for less than 10% of the capital costs. In the PDU a glass
fibre cloth with a particle size retention of 0.5 um was used successfully,
but its weakness and lack of rigidity make it unsuitable for the continuous
high temperature operation required in a commercial facility. An alternative
precoated wire screen method has shown promise in experimental runs but screen
size and precoat material selection has yet to be optimized. Alternative
precoat materials to Celite are being investigated which are made of carbon-
aceous materials. The spent precoat containing ash can then be sent to the
gasifier or combustion section for disposal, thereby decreasing plant waste.
Research on the hydrocracking section has focused on catalyst selec-
tion. Over 50 different catalysts have been tested in a stirred autoclave for
their lifetime, resistance to contaminates, and kinetic properties. Alter-
natives to catalyst poisoning include designing an extract cleaning system to
remove deactivating agents before they reach the catalysts.
Capital Costs
Capital costs for a commercial scale LSE process plant have not been
published. The capital cost of the 25 tonne/day pilot plant was expected to
be $100 million.
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BRITISH DIRECT COAL COMBUSTION PROCESS
19. National Coal Board Pressurized Fluidized Bed Coal Combustion Process
Process Description
The Pressurized Fluidized Bed Coal Combustion (PFBC) process is being
developed for combined gas/steam cycle utility power generation. Research in
this area is being conducted by the National Coal Board (NCB) in a 5 MW test
facility at Leatherhead. A typical flow diagram of a PFBC combined cycle
utility power plant is shown in Figure 1. The concept involves burning coal
in a fluidized-bed of calcium-containing mineral such as limestone to absorb
the polluting sulfur in the coal. The sulfated limestone is removed from the
combustor and discarded for land fill applications or regenerated for elemen-
tal sulfur recovery and reuse in the combustor. In the pressurized combustion
concept, the generated hot flue gases are expanded through a turbine to
generate electricity.
The fluidized-bed boiler furnace section consists of an enclosed space
with a perforated base to admit combustion air. Part of the enclosed space
directly above the air distribution plate is occupied by a layer of granular
material such as sand or limestone. The air, which is forced through the
supply plenum is sufficient to lift (fluidize) the bed and suspend it in the
airstream. This promotes violent boiling (mixing) and agitation of the bed
material. After fluidization has occured the combustible fuel is introduced
and ignited. The bed material absorbs the heat of reaction and transfers as
much as 50% of it to immersed water tubes and waterwalls. The waterwall
tubes, which surround the inner fluidized-bed walls, maintains the wall
temperatures within a safe operating range. The optimum operating temperature
is between 1500? and 1700?F. When a high moisture content fuel such as wood
wastes or municipal refuse is burned, the heat required to evaporate the
moisture in fuel maintains the bed temperature within acceptable limits
without the use of waterwall cooling tubes.
Because of the enhanced heat transfer characteristics of the fluidized
bed, the unit is smaller than a conventional boiler of the same output. In
order to take advantage of the heat transfer mechanism and to limit the tem-
perature, the bed may be divided into a number of cells each surrounded by
waterwalls. The heat of the gas leaving the combustion zone is removed by
conventional convection heat-recovery equipment.
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Figure 1. SUPERCHARGED BOILER COMBINED CYCLE
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a 1000 hour testing phase. This testing
in Leatherhead has recently completed
program also involved Stal-Laval,
Combustion Systems Ltd., and Babcock International. The fluidized bed com-
bustor configuration used in the Leatherhead facilities is shown in Figure 2.
This unit incorporates a range of superheater and reheater tube alloys
testing at bed temperatures typical
conventional boiler. The bed has a
can operate at pressures of up to 6 bars
to those that would be encountered
cross-sectional area of up to 0.84
in a
m2 and
at heat inputs of 5 MW.
1. Compressed air inlet
2. Water inlets and outlets to tube
bank circuits
3. Baffle tubes in freeboard
4. "High level" bed offtake
5. Tube bank
6. Start-up burners
7. Coal nozzle
8. Air distributor
9. Bed offtake
10. Corrosion probes
11. Air cooling supply to corrosion
probes
12. Gas offtake baffle
13. Mixing baffle
14. Gas splitter
15. Corrosion indicator probes
16. Zirconia-cell oxygen probes
I
Figure 2. ARRANGEMENT OF FLUID BED COMBUSTOR Mk VI
1
1
The combustion gases leaving the 5 MW PDU are split into two streams as
shown in Figure 3. In the initial testing period each stream had 3 cyclone
dust collectors in series for particulate removal. Each stream then proceeded
to a cascade of turbine blades that serves as a turbine blade test unit. One
cascade was supplied by Stal-Laval and the other by GE to test gas turbine
alloys exposed to combustion gases with relative velocities of up to 525 m/s.
The gas cleanup cyclones were effective at removing most particles larger than
10 microns. The particles that did not reach the cascade blades were rela-
tively soft. After 650 hours of operation the Stal-Laval stream was modified
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by removing one of the cyclones to increase particulate loading. After 1000
operating hours the cascade systems were dismantled to check for wear and
corrosion. With the exception of minor ash buildup which did not affect
performance, none of the coated blades or specimens in either streams showed
signs of corrosion attack.
Stream 1
2? cyclone
1 ? cyclone
iet-gown--' -
hopper]
2? cyclone
1 ? cyclone
1. Dust sampling (CURL)
2. lkor probes
3. PMS laser and impactor
4. Spectron laser
5. Dust sampling (NYSERDA)
and ionization alkali monitor
6. Gas analysis (O,, CO,, CO, SO2, NO,)
7. SO3 analysis
8. Spectral alkali photometer
9. Electrostatic charging
10. Electrostatic charge sampler
11. Propane injection (reheat)
Figure 3. TEST FACILITY
The 3 foot by 2 foot cross-section bed is usualy operated with an 8 foot
deep fluidized-bed. The bed normally operates in the temperature range of
700?C to 950?C and at about 6 bars pressures. Combustion efficiencies as high
as 99% have been achieved when operating with 30% excess air and a fluidizing
velocity of 5 feet per second. When operating on 3% sulfur coal over 85% of
the sulfur was retained in the bed. This bed contained 1.5 times the stoich-
iometric quantity of dolomite required to retain the sulfur.
Due to the small size of the Leatherhead facility scale up to a 500 MW
pilot plant was deemed too risky. Therefore, an 80 MW experimental pilot
plant was constructed at Grimethorpe for futher testing. The overall flow
diagram for this facility is shown in Figure 4. The combustor (Figure 5) is
designed to operate at 12 bars and has a 14 meter height and 4 meter outside
diameter. When operating at maximum capacity this 2 m x 2 m cross-sectional
area combustor operates with a fluidizing velocity of 2.5 m/s, a bed temper-
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ature of 850?C, and an air mass flow of 31 kg/s. This unit can handle
10 tons/hr of coal with a top size of 6.4 mm. Coal mixed with the limestone
or dolomite S02 sorbent is fed into the bed via a Petrocarb feeder at nine
injection nozzles. The calcium to sulfur mole ratio to achieve 90 removal of
the SO2 is 2.6.
Figure 4. OVERALL FLOW DIAGRAM
t
The fluidized bed has a maximum depth of 4.5 meters and contains a bank
of evaporator tubes. The water-cooled walls extend an additional 4.5 meters
above the expanded bed height. This 4.5 meter freeboard is designed to pro-
vide space for larger entrained particles to settle out and return to the bed.
The freeboard area also contains a small band of heat exchanger tubes for com-
bustion gas cooling in the event of excessive above bed burning. The combus-
tion product leaving the top of the PFBC unit is split into four streams for
feeding into four primary cyclones arranged radially around the combustor.
Each primary cyclone is connected to a secondary cyclone where the cleaned gas
finally exits into a common header. From here the gas is cooled to below
300?C before pressure letdown, silencing and exiting into a 90 meter tall flue
stack.
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Air inlet
Furnace hood
Freeboard cooler inlet
Support bracket
Furnace support steelwork
Freeboard cooler bypass_
Pressure shell-30mm thick
Internal insulation
rated at 150 kW
Internal access galleries
Coal feed nozzles
Distributor plate
Furnace water wall
inlet Distributor
Pressure shell bottom flange-
Pressure vessel bottom dome
Coal inlets
C
d
e
g
3270
h
i
4880
Figure 5. COMBUSTOR PRESSURE VESSEL
6000
During testing tube bundle configurations and dust collection equipment
will be modified to record system performance. Part of this testing will
involve passing combustion gas through a tertiary cleaning system and then
expanding it through a cascade of gas turbine blades to obtain additional
performance data. These gases leaving the cascade will be cooled and depres-
surized in a separate system. The first phase of testing is envisioned to
require 2900 hours and take 24 to 30 months to complete. Testing started in
April 1981. A technology fact sheet for this process is shown in Table 1.
1
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I N S T I T U T E O F G A S T E C H N O L O G Y
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Process Goals
Testing in the 5 MW PFBC reactor at Leatherhead has established the
steady state operating performance of this technology at 6 bars pressure. The
80 MW pilot plant at Grimethorpe is designed to provide a wider range of
operating conditions for pressures up to 12 bars. This testing program is
designed to:
? Determine combustor performance and pollution emission levels of a wide
range of operating conditions at elevated pressures
? Establish system performance for different coal feedstocks; two different
U.S. and West German coals will be tested in addition to U.K. coals
? Provide part-load data and dynamic response data for systems control
design.
Assess system performance effects when design changes are made, for
example, elimination of some of the coal feed nozzles or changes in bed or
freeboard heights.
1
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? Determine corrosion behavior of tubes submerged in the fluidized bed as
well as those subjected to particulates in the freeboard.
Relationship to Prior Technology
Though fluidized-bed combustion is often divided into the two distinct
areas of atmospheric and pressurized technology, the two technologies have
common development roots and several common research concerns. Because of the
similarities of the technologies, in this report they will be treated in an
intermingled manner.
The concept of the fluidized-bed was invented in the 1920's as a means to
promote chemical reactions. By the early 1940's, the fluidized-bed concept
was in commercial use for petroleum cracking. By the late 1950's the
fluidized-bed technique was commercially successfully for metallurgical heat
treatment and ore roasting. Hundreds of these units have been sold. In the
late 1940's, several U.S., British, German, and French companies began
development of fluidized-bed combustion sytems. A French design, which
features two-staged coal combustion, was not a market success.
In the late 1950's, the British National Coal Board's (NCB) Coal Research
Establishment operated by the British Coal Utilization Research Association
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(BCURA) continued earlier work by Douglas Elliott of Britain's Central
Electricity General Board (CEBG). General lack of interest in coal due to low
cost oil and gas slowed research. By the late 1960's, however, interest in
fluidized-bed coal combustion was prompted in Britain by the NCB's desire to
sell coal worldwide and by the emergence of future oil supply problems. In
1968, Elliott developed the concept of pressurized fluidized-bed combustion.
Under his direction, a 5 MW (thermal) unit was commissioned in 1969. Though
the NCB proposed a 20 MW pilot facility in 1971, the British Government and
the CEGB refused funding.
In the United States, Michael Pope of Pope, Evans, and Robbins (PER) won
support from the U.S. Office of Coal Research (OCR) in 1965 to build three
atmospheric fluidized-bed test units. The largest of these units, a 0.5 MW
(thermal) unit, was first operated in 1965 at Alexandria, Virginia. Foster
Wheeler Corporation and Combustion Power Company were carrying out limited
fluidized-bed combustion research by 1970. The phenomenon of greatly reduced
sulfur emissions was first reported as a result of PER's work with OCR in 1968
using the Alexandria, Virginia test facility,. In 1972, OCR provided funds
for PER to build three protoype atmospheric fluidized-bed boilers. The first
was built at Alexandria, Virginia. Environmental concern and the EPA led to a
joint EPA/NCB program aimed at sulfur dioxide and nitrogen oxides control
using fluidized-bed combustion. The program indicated that up to 95% of the
sulfur in high sulfur coal could be captured by fluidized-bed combustion.
With OCR funding, PER built a 30 MW unit at Rivesville, West Virginia in
1972. The formation of ERDA after the OPEC oil embargo resulted in
$19.8 million in U.S. Government funding for fluidized-bed combustion in 1975.
Recent development in Europe has had a more commercial orientation than
in the United States. In 1972, Douglas Elliott of the NCB and a partner
founded Fluidfire Development, Ltd. The company has developed and success-
fully marketed heat treatment furnaces, refuse burning boilers, and flue heat
recovery units using fluidized-bed designs. The company has been aggressive
in concept development; one development has been a fluidized-bed boiler with a
bed depth of only a few inches. In one configuration, a novel method of con-
trolling the boiler was developed by varying the position of the fluidized-bed
around the heat exchangers.
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Similarly, with commercialization intended, the NCB sponsored converting
the boiler at the Babcock and Wilcox, Ltd. factory in Renfrew, Scotland to
fluidized-bed firing in 1974. By August 1975, the 40,000 lb/hr unit was in
operation. The 10 ft2 bed operates at 400 psi and 560?F. The unit which is
started up by oil-fired overhead burners now operates on oil. However, its
original testing program included coal firing; up to 98% sulfur retention was
noted with 3.6% sulfur coal and a limestone bed. After testing, and to date,
no loss in bed tube material has been observed. The unit's turndown ratio is
4:1. In 1976, Babcock and Wilcox, Ltd. announced that units would be made
available with a warranty for production of up to 500,000 lb/hr of steam.
The NCB has sponsored other small conversions as well as including the
conversion of an 80,000 lb/hr boiler operated by British Steel Corporation.
This unit was to be commissioned in 1978.
Operating Facilities
The National Coal Board has operated a PFBC unit with a 5 MW rated capa-
city at its Coal Utilization Research Association Laboratories (BCURA, now
CURL) in Leatherhead. Over 3000 testing hours have been logged on this unit
with operating pressures of up to 6 bars. Initial bench scale studies at the
NCB's Coal Research Establishment (CRE) labs provided the data for the
Leatherhead facility.
In 1977 the International Energy Agency started construction of the 80 MW
Grimethorpe facility which is based on the research data from the Leatherhead
facility. Cold flow commissioning began in October 1979 and hot commissioning
started in September 1980. Cold flow tested required an additional two month
period due to mechanical problems in the ancillary equipment. The first phase
of the experimental program at Grimethorpe started in April 1981 which will
involve 2900 hours of testing over a 2 to 2-1/4 year period.
Major Funding Agencies
PFBC studies which started in 1969 and much of the laboratory work since
1972 has been sponsored by the U.S. DOE in cooperation with other U.S. organi-
zation. These include the Electric Power Research Institute, American
Electric Power (AEP) and General Electric. Funding for the 80 MW PFBC
facility at Grimethorpe in Yorkshire was provided by the International Energy
Agency.
 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
Technical Problems
Although much of the experimental work to date has concentrated on the
reactor design, material problems and turbine blade performance, little
research has been done on coal feeding problems. In the current approach,
coal is dried and crushed to a compatible size dependent on the pressurized
fluidizing gas velocity. For example, for a fluidizing velocity of 1.2 m/s a
coal and dolomite mixture with a top size of 3 mm would be required as feed.
Based on the current use of nozzles to inject the coal into the PFBC reactor,
one nozzle is required for each 1.5 m2 of bed area. Therefore, for a 200 to
250 MW(e) plant over 50 coal nozzles would be required to feed the coal into
the reactor. Larger nozzles or a less complicated feed system is required to
make large commercial plant operation feasible.
In addition to reducing the number of nozzles required, it would be
advantageous to increase the maximum allowable coal size range to 30 mm.
Large coal of this size has been experimentally combusted in a fluidized bed
with no problems. However, if noncombustible materials accumulate within the
bed, performance can be degraded. A larger coal size distribution will
require new nozzle design, especially when handling wet run of mine coals.
This large size distribution will also require better fluidizing bed con-
trol. In may be desirable to use a circulating bed which will recycle the
dust overflow back into the bed.
The final area of concern is increases to overall plant efficiency for
the PFBC combined cycle plant. Current efficiency is estimated to be in the
range of 39 to 40 percent with reheat of the steam system and a gas turbine
inlet temperature of 800?C. This performance is 5 percent better than a
conventional coal burning power plant with flue gas desulfurization to reduce
S02 while burning a 3 to 4 percent sulfur coal. Higher efficiencies of
45 percent can be obtained by increasing gas turbine inlet temperatures to
1100?C. However, slag formation which will bond to turbine blades is likely
at this temperature. A secondary combustor may be required to reach this
temperature by heating the PFBC reactor output. Fuel for this after burner
stage may be produced in a preliminary pyrolysis/gasifier stage of the PFBC
reactor.
t
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1
1
1
Capital Costs
Reduction of capital costs of 12 percent and a decrease of operating
costs when compared with conventional coal combustion power generation have
been predicted. The 80 MW pilot plant at Grimethorpe was constructed at an
estimated cost of $10 x 106 in 1978.
 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
BRITISH COAL GASIFICATION PROCESS
20. Esso Chemically Active Fluidized Bed Coal Gasification Process
Process Description
The Esso Chemically Active Fluidized Bed gasification process (CAFB) has
been tested in a 150 kg/hr reactor. A schematic of this pilot plant is shown
in Figure 1. The pilot plant consists of a fluidized bed reactor where fuel
is partially oxidized at 900?C within a bed of lime particles fluidized by
air. A top coal size of 3000 microns can be fed into the gasifier. Over 80%
of the sulfur in the fuel, either oil, gas or coal, is retained by the lime
when it is converted to CaS. In addition to sulfur, metals within the coal
structure such as V, N, and Na are also retained in the lime bed. The spent
lime material is then transferred from the gasification reactor to the ad-
joining regeneration fluidized bed reactor. In this reactor the CaS which was
formed in the gasification reactor is oxidized back to lime and SO2 at 1050?C.
Cyclone
Coal -?
Air -
Lime
900 ?C
Gasifier
-~ Raw Gas
1050 ?C
Regenerator
1
Bed Drain Air Bed Drain
Cyclone
Figure 1. FLOW DIAGRAM OF ESSO CHEMICALLY ACTIVE FLUIDIZED BED
GASIFICATION PROCESS
During the reaction in the gasification vessel the coal undergoes
pyrolysis to release large concentrations of condensible hydrocarbons. In
actuality only partial gasification occurs and large quantities of low sulfur
char is produced as a by-product. Tests results using Illinois No. 6 and the
Texas lignite coals shown in Table 1 are presented in Table 2. A summary of
the sulfur removal affectiveness of this process when operating on the
Illinois and Texas coals is shown in Table 3. A technology fact sheet
summarizing this process is presented in Table 4.
B-162
G A S
 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
Table 1. ANALYSIS OF ILLINOIS NO. 6 AND TEXAS LIGNITE
FED TO CONTINUOUS UNIT
Illinois No. 6 Texas Lignite
Actual Dry Basis Actual Dry Basis
x x x x
1
Moisture
4.63
-
14.02
-
Ash
7.36
7.80
18.50
21.50
Carbon (corrected)
70.3
73.64
51.20
59.50
Hydrogen (corrected)
4.61
4.83
3.60
4.20
Sulphur (total)
1.78
1.90
0.82
0.95
Nitrogen
1.44
1.50
1.00
1.20
Oxygen+errors (by
9
85
10
40
10
90
12
60
difference)
Gross calcries/GM
.
6936
.
7260
.
5004
.
5817
Gross Btu/lb
12486
13068
9007
10470
CO2 (x)
0.16
-
0.51
-
Table 2. COMPARISON OF GASIFIER GAS QUALITY FROM FUEL OIL AND COAL
Gasifier Fuel
Heavy Fuel
Oil
Illinois No.6
Coal
Texas
Lignite
Nitrogen + inerts
58.4
59.2
59.0
Carbon monoxide
10.2
12.2
12.2
Carbon dioxide
10.2
9.9
12.1
Methane
7.7
4.2
2.2
Ethylene
5.0
0.8
0.7
Ethane
0.1
0.1
0.1
Hydrogen
8.4
13.6
13.7
 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
Table 3. SUMMARY OF DESULPHURIZING PERFORMANCE DURING
MINI-RUN ON COAL
Coal
(type and
size)
Mean
% SRE
on Coal
Expected Result
on Oil (X SRE)
SO Emission
(16502 per
10 Btu)
Illinois No. 6
(1405p. down)
Texas Lignite
74.8
73.8'
0.73
(800 down &
64.0
68.5
0.76
1405 down)
Texas Lignite
(1/8" down)
82.4
67.2
Process Goals
The objective of this development is to develop a commercial coal or oil
burning gasifier which can remove sulfur from pollutants from the fuel and can
be easily integrated with existing gas-fired boiler plants. Research on the
operating range and performance of this gasifier is underway. Also being
examined in the long term affect of coal ash on the sulfur absorbing proper-
ties of the line bed material.
Relationship to Prior Technology
The concept of the Catalytically Active Fluidized Bed process was first
developed in 1967 and was originally applied to the desulfurization of fuel
oils. In 1975 this concept was extended to the use of coals in a batch test
at Abingdon, U.K. However, the concept of removing sulfur compounds from gas
using a limestone absorbent is not new. This approach has been applied to
sulfur removal from stack gases as a pollution abatement process.
Operating Facilities
Esso Petroleum Company Ltd., Abingdon, Oxon, U.K. has operated a 150
kg/hr process development unit at the Abingdon test facilities since 1975. A
20 MW pilot plant unit has been proposed for operation at the Texas Central
Power and Light Company which will operate on Texas lignite.
1
 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
1
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
Major Funding Agencies
Esso Petroleum Company Ltd. and the U.S. Environmental Protection Agency
have provided the funding for this process.
Technical Problems
No technical problems have been reported in the literature.
Capital Costs
Capital costs for this process have yet to be published.
32(3)/process/ER
t
I N S T I T U T E O F G A S T E C H N O L O G Y
Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
AUSTRALIAN COAL CONVERSION PROCESS
21. CSIRO Flash Pyrolysis Coal Liquefaction Process
Process Description
The Commonwealth Scientific and Industrial Research Organization -
Division of Fossil Fuels (CSIRO) high speed pyrolysis process utilizes a
fluidized bed reactor to convert pulverized coal (0.2 mm in size) to tar, gas
and reactive char. The reactions occur in an oxygen-free atmosphere at about
600?C and atmospheric pressure. Reaction times of about 1 second are main-
tained followed by rapid quenching of the product. The tar, which can amount
to more than 30% of the m.a.f. coal feed, is then hydrogenated to lighter
grade liquid fuels. Some of the product gas and char are used to produce
hydrogen for the hydrogeneration reaction. A plant layout chart and flow
diagram of the flash pyrolysis process are shown in Figures 1 and 2, respec-
tively.
The goal of this program is to determine the technical and economic fea-
sibility of the flash pyrolysis process on Australian coals for the production
of liquid fuels. The CSIRO process is also being tested for the pyrolysis of
wood wastes, the production of olefins from coal pyrolysis, and the production
of carbon anodes from coked flash pyrolysis tars.
Relationship to Prior Technology
The flash pyrolysis process has been studied by various groups for the
last 30 years or more. Experiments have been performed in laboratory bench-
scale units to nearly commercial scale-pilot plants. Electric power has been
produced by electric utilities in the U.S. and West Germany using char pro-
duced by the pyrolysis process.
Operating Facilities
The CSIRO research effort on flash pyrolysis started in 1974. A 1 to
3 gram/hr fluidized-bed pyrolysis reactor and a 100 gram/hr entrained bed
pyrolysis reactor have been in operation since 1975.
In 1977 a 20 kg/hr process development unit capable of operating in
either the fluidized bed or entrained bed mode was commissioned. This unit
was converted to an integrated facility in 1980/81 incorporating the pyro-
lyzer, a char-burning heat generator and an on-line tar hydrogenator.
B-168
I N S T I T U T E O F G A S T E C H N O L O G Y
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I
 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
1
I
COAL
PREPARATION
HEAT
GENERATOR
TAR
COLLECTION
1
TAR
HYDROGENATOR
HYDROGEN
PRODUCTION
I GAS I I
SYNTHETIC CRUDE OIL
Figure 1. LAYOUT OF COUPLED PYROLYSER - POWER PLANT INSTALLATION
PROBE
DISTRIBUTOR
PLATE
Figure 2. FLOW DIAGRAM OF 20 kg h f FLASH PYROLYSER
I N S T I T U T E O F G A S T E C H N O L. O G Y
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
Operation of this facility, which started in late 1981, will provide the
necessary data for the construction of a 1 tonne/hr pilot plant.
The CSIRO work has been carried out within the Organization mainly by the
Division of Fossil Fuels and and Division Applied Organic Chemistry. Outside
CSIRO, support has come from the Australian Mineral Development Laboratories,
and the Universities of Sydney and Queensland. A version of the CSIRO
laboratory-scale reactor has been built at the University of Waterloo,
Ontario, for the pyrolysis of wood-waste. Another version has been built by
DuPont at Wilmington, Delaware, to study the production of olefins by flash
pyrolysis of coals at temperatures around 900?C. In the CSIRO laboratories a
second PDU-scale rig has been operated for the Aluminum Development Council,
by COMALCO, to determine if flash pyrolysis tars can be coked to produce a
substitute for petroleum coke used in carbon anodes for aluminum smelting.
COMALCO are also building a version of the laboratory-scale rig as part of the
carbon anode project.
Major Funding Agencies
Funding on all synthetic fuels programs in Australia amounted to about
$10 x 106 in 1979-80. The CSIRO spent $5.07 X 106 in 1979-80 in synthethic
liquids fuels research. Part of this money ($420,000) came from the National
Energy Research, Development and Demonstration Council (NERDDC) which itself
total budget that
research.
Technical Problems
Agglomeration of
small scale reactor.
pyrolysis. Oxidative
year of $5.05 X 106 for synthetic liquid fuels
coal in the fluidized bed reactor was a problem in the
Nearly all coals show some plastic behavior during flash
pre-treatment will reduce agglomeration but will also
reduce tar yields. However, agglomeration did not appear to be a problem in
the larger size reactors. In addition to caking tendencies, the coals cation
content effects yields. In general, for a given coal, as the cation content
increases, the tar yield decreases. Acid washing of the coal can be used to
decrease cation content.
Capital Costs
The capital costs and operating costs for the CSIRO flash pyrolysis
process are shown in Table 1. All costs are in late 1980 dollars. An
Australian Millmerran coal was assumed to be the feedstock.
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I N S T I T U T E O F G A S T E C H N O L O G Y
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
Table 1. CAPITAL AND OPERATING COSTS FOR 40,000 BARRELS PER DAY
CSIRO FLASH PYROLYSIS COAL LIQUEFACTION PLANTS
Capital Costs ($106)*
Feed preparation 60
Pyrolysis and product recovery 274
Primary liquid hydrogenation (1) 448
1
1
1
2. Offsites, utilities, engineering and design 312
3. Estimating contingency, work capital,
initial catalysts and chemicals, and
start up costs
4. Total capital investment 1483
Operating Costs ($106 p.a.)*
1. Raw materials
Feedstock 137 (2)
Water 4
2. Other
Maintenance 44
Taxes, insurance and overheads 33
Operating labor and supervision 11
Catalysts and chemicals 12
Purchased fuel gas and electricity - (3)
1
3. Credits for ammonia and sulphur byproducts -4
4. Net operating costs 237
Notes.
(1) Includes H2 production
(2) Cost at $1/GJ, based on $28/tonne dry, opportunity cost.
Cost shown net after allowing byproduct char credit at $1/GJ.
(3) All requirements produced from pyrolysis gas and coal char.
 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
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Approved For Release 2007/02/20: 6IASRDP83Mb0514cRC61a06b6b0'P8-
 POLISH COAL LIQUEFACTION PROCESS
22. Polish Central Mining Institute Catalytic Hydrogenation Coal Liquefaction
Process
Process Description
The Central Mining Institute, Division of Carbochemistry, has developed a
catalytic hydrogenation coal liquefaction process. In this process, coal is
first crushed and dried before entering a ball mill where it is ground to less
than 0.5 mm. The prepared coal is then fed to a mixing tank where recycled
process solvent is
form a slurry. The slurry contains
approximately 28.6 wt % coal and 71.4 wt % recycled solvent. The recycled
solvent is actually a mixture of .44 vol. % anthracene oil and 56 vol % high
boiling range recycled oil. The properties of this solvent are presented in
Table 1.
From the slurry preparation tank the mixture is heated and sent to the
extract mixer column where the coal is reacted with the solvent. Coal reac-
tion time in this reactor is about 50 minutes. The reaction proceeds at 410?C
and at 45 atm. The mixture exiting this reactor is then hot filtered to
remove mineral matter and unreacted coal residue. After filtration the coal
oil enters a vacuum distillation unit where the more volatile oils are
separated. Bottoms from the distillation unit and filtrate cake are sent to a
carbonize r.
1
I
I
Table 1. PROPERTIES OF RECYCLED OIL SOLVENTS (Hydrogen Donor)
Specific gravity at 20?C, g/ml
Ultimate analysis (maf),wt%
1.02
Carbon
90.92
Hydrogen
8.64
Sulfur
First drop boiling/initial
boiling point ?C
Fraction boiling to 360?C, vol%
0.31
Fraction boiling to 400?C, vol%
97.8
T E C H N O L O G Y
I NA~$prbv6dFoP RIlease 2007/02/20: CIA-RDP83M00914RO01000060018-0
 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
t
The low-ash extract produced in the vacuum distillation unit is then
processed in a catalytic hydrogenation unit. The coal extract entering this
unit contains about 50% liquids with a boiling range above 400?C. Hydrogena-
tion occurs in the presence of a nickel-molybdenum catalyst at 240 atm and at
400?C. The hydrogenated liquids are then sent to an atmospheric distillation
unit where the low and middle boiling range liquids are separated from the
high boiling range (>400?C) liquids. The high-boiling range hydrogenation
product is then vacuum distilled. Recovered oil from the vacuum distillation
unit is recycled to the slurry preparation section for use as the hydrogen
donor solvent, and the bottom residue is carbonized. The property of the coal
used in pilot plant testing is shown in Table 2. When operating on this coal
the properties of the low-ash extract and liquid products from the hydrogena-
tion unit that can be expected from this process are shown in Table 3. These
properties reflect operation of the process under the conditions in Table 4.
The technology fact sheet for this process is in Table 5.
The light liquids products are reported to be suitable for further
refining into motor fuel products.
Process Goals
This process is being developed in Poland by the Central Mining
Institute's division of Carbochemistry at Tychy-Wyry. Research is being per-
formed to determine the operating parameters of the process under various
operating conditions. Particular attention has been focused on the selection
of a sulfur-tolerent catalyst for use in the hydrogenation reactor. The
following process steps are also being investigated.
? Coal/solvent slurry preparation
? Coal extraction with and without hydrogen
? Extract residue separation
? Extract solvent recovery/distillation
? Extract residue low-temperature carbonization
? Extract hydrogenation
? Hydrogenation product distillation.
Research has also been conducted on the production of electrode coke, binders,
and carbon absorbents from the coal liquids products.
B-175
I N S T I T U T E O F G A S T E C H N O L O G Y
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
Moisture, wt%
Ash (dry basis), wt%
Volatile matter (maf basis), wt%
Elements (maf basis), wt%
Carbon
Hydrogen
Sulfur
Petrographical composition, wt%
vitrynite
Egzynite
Inertynite
Mineral matter
Heating value, kcal/kg
79.33
5.52
0.58
73.0
9.0
13.0
1
I N S T I T U T E O F G A S T E C H N O L O G Y
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
Table 3. THE PROPERTIES OF THE COAL EXTRACT AND LIQUID PRODUCTS AT
THE 1200 kg COAL/DAY PLANT
Products
Property
Coal
Extract
Light
Liquid
Product
Heavy
Liquid
Product
Water content, wt%
trace
0.3
trace
Ash (dry basis) wt%
0.10
trace
0.16
Specific gravity at
20?C, g/ml
1.22
1.08
1.13
Ultimate anlysis
(maf basis), wt%
Carbon
Hydrogen
6.50
7.56
6.95
Sulfur
Benzene insolubles, wt%
14.6
1.2
9.5
Asphaltene, wt%
5.6
1.3
12.3
Oils, wt%
79.8
97.5
78.2
Initial b.p., ?C
262
93
247
Fraction boiling to
200?C, vol%
--
3.0
--
Fraction boiling to
320?C, vol%
1.5
46.4
7.5
I N S T I T U T E O F G A S T E C H N O L O G Y
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
Table 4. PROCESS CONDITIONS OF THE CENTRAL MINING INSTITUTE
CATALYTIC HYDROGENATION COAL LIQUEFACTION PROCESS
1. Extraction
1.1. Coal content in the slurry wt % 28.6
1.2. Feed rate of the slurry 1/hr 14.8
1.3. Reaction time min. 50
1.4. Pressure atm. 45
1.5. Temperature ?C 400
2.1. Temperature ?C 233
2.2. Filtration rate kg/m2/hr 287
3. Carbonization
3.1. Max. temperature ?C 550
4. Filtrat Destillation
4.1. Pressure mm 40
4.2. Max. temperature column top ?C 265
5. Hydrogenation
5.1. Catalyst Ni-Mo
5.2. Pressure atm 250
5.3. Reactor temperature ?C 430
5.4. Through-put kg feed/
liter of 0.37
catalyst/
hour
Relationship to Prior Technology
This process is based on the Bergius-Piers process (IG Farben Process)
which was developed in Germany. However, hydrogenation of the coal without
the addition of a catalyst or hydrogen in the first reaction step represents a
significant deviation from the original concept. It was not stated directly
in the literature, but one can conclude that hydrogen generation for the cat-
alytic hydrogenation step of this process will be produced by the gasification
of coal residue from the process. This has not been technically demonstrated
and is another different modification of this process.
1
1
1
I N S T I T U T E O F G A S T? E C H N O L 0 G Y
Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
 1
Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
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 Operating Facilities
A bench-scale unit with a coal capacity of 120 kg/day was constructed at
the Carbochemistry Institute of Tychy Wyry in the mid-70's. This unit was
used to gain operating information for the construction of a 1200 kg/day pilot
plant in 1977. This unit is also located at the Carbochemistry Institute.
Major Funding Agencies
This process is supported by the Polish Government.
Technical Problems
No technical problems have been reported in the literature.
Capital Costs
The capital costs for this process have not been published.
T E C H N O L O G Y
I N S T T U T E O F G A S
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 Approved For Release 2007/02/20: CIA-RDP83MOO914ROO
SOUTH AFRICAN COAL LIQUEFACTION PROCESS
23. SASOL (Lur i/Fischer-Tropsch) Coal Liquefaction Process
Process Description
Sasol I
SASOL (South African Coal, Oil, and Gas Corporation, Ltd.) was founded in
1950 for the production of synthetic liquids from coal. Their first synthetic
liquids plant was completed in 1955 and is located at Sasolburg, 50 miles
southwest of Johannesburg, South Africa. In 1979 Sasol was reoganized into
Sasol Limited with 70% of the government owned company sold to the public.
The new corporate structure which includes the two new synthetic liquids
plants and the South African government's Industrial Development Corps.'
involvement is presented in Figure 1.
OWNERSHIP (SHAREHOLOING)
- - - LOANS
Figure 1. SASOL GROUP STRUCTURE
I N S T I T U T E O F G A S
T E C H N O !_ G
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
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South Africa's first coal to liquids plant, Sasol I, utilizes Lurgi/
Fischer-Tropsch technology to produce a broad range of synthetic products. A
block flow diagram of this facility as it exists today, is shown in Figure 2.
This facility contains 17 Lurgi gasifiers; 13 Mark III, 3 Mark IV units, and
one experimental Mark V gasifier. The Mark III, IV, and V Lurgi gasifiers
have a 3.68, 3.85, and 4.7 m internal diameter, respectively. The Mark IV
gasifiers were installed in 1977/8 and the Mark V unit was commissioned in
1980.
In the Sasol facility coal is delivered from the adjacent Sigma mine in
two coal sizes, -12 mm and 12-50 mm. This low grade bituminous coal, analysis
shown in Table 1, is split about 40% fines and 60% coarse material. The fines
are fed to two power plants and the remainder is gasified. The coal is deliv-
ered from the mine by conveyors and stored in bunkers before entering the
gasifier lockhoppers. These lockhoppers located at the top of the gasifiers
hold between 6.5 and 10 tonnes of coal depending on their age. After pressur-
ization with raw gas the coal is fed by gravity through conical valves into
the distribution grate of the Lurgi gasifier (Figure 3). Gasification is
carried out at approximately 375 psig and 1200?C. Total facility gasifier
inputs include 9,000 to 11,000 tonne/day of coal, 65,00 to 95,000 m3/hr of
oxygen, and 329,000 to 438,000 kg/hr of high pressure steam. Total raw gas
production ranges from 450,000 m3/hr to 600,000 m3/hr. Overall availability
of the gasifier plant is nearly 85%.
The raw gas leaving the gasifier at about 500?C is scrubbed and cooled to
20 to 25?C to remove higher boiling point tars, oils and entrained ash dust.
All tars and oils are filtered before undergoing further processing. The
cooled raw gas is then sent to one of three Rectisol trains where carbon
dioxide, hydrogen sulfide and organic sulfur are removed. Rectisol utility
consumption averages 35.7 kWhr of electricity, mostly used for methanol
recirculation pumps, and 79.2 kg of steam per 1000 m3 of pure gas produced.
The pure gas leaving the Rectisol units is sent primarily to the Fischer-
Tropsch sections; a Ruhrchemie-Lurgi Arge fixed bed design and the other and M.
W. Kellog Co. Synthol entrained bed design. A small fraction of the pure gas
is also sent to a methane-reformer where the methane content is reduced before
entering the Synthol reactors. The synthesis gas composition entering the two
types of Fischer-Tropsch reactors is shown in Table 2.
I N S T I T U T E O F G A S T E C H N O L 0 G. Y
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
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Type
Bituminous
a
Moisture (as received)
10.7 wt %
I
Volatiles (dry basis)
22.3 wt %
Ultimate Analysis (dry basis), wt %,
Ash
35.9
I
Sulfur
0.5
Nitrogen
1.2
Carbon
50.8
Hydrogen
2.8
Oxygen
8.8
Ash fusion Temperature, ?C:
Softening point
1340
Melting point
1430
Fluid point
1475
Energy Content, MJ/kg (dry)
O FEED COAL
Figure 3. LURGI MARK IV GASIFIER
I N S T I T U T E O F G A S T E C H N O L O G Y
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Table 2. RAW GAS COMPOSITIONS, VOLUME %
Component
Arge
Synthol
H2
39
40
CO
22
18
C02
28
31
CH4
9
9
N2
H2S
1
1
The Synthol reactor system has been developed to the commercial stage by
Sasol, which today owns all the rights. This process produces predominantly
lower-boiling point hydrocarbons in the gasoline and diesel oil range and the
Arge process produces high-boiling point hydrocarbons, including a range of
solid waxes. A breakdown of the product mix from these two reactors is shown
in Figure 2. Both processes use an iron-based catalyst activated with certain
promoters. These catalysts were manufactured in West Germany prior to 1969
when Sasol commissioned its own catalyst preparation plant. Production compo-
sition can be varied somewhat by charging catalyst properties and reactor con-
ditions.
Part of the synthesis gas can be blended with industrial gas which is
distributed by the Gascor subsidiary in a pipeline system. Some of the
Synthol tail gas is also blended with synthesis gas as feed for an ammonia
synthesis unit. Before entering the ammonia synthesis section the feed gas is
shifted in a water-gas shift reactor and the carbon dioxide and hydrocarbons
are removed. Nitrogen for the ammmonia synthesis reaction is supplied as a
byproduct of the gasification section air separation units. Ammonia produced
can be further converted to ammonium nitrate, nitric acid and limestone
ammonium nitrate.
Also linked to the Sasol One plant is an olefin section. Fuel gas, LPG
and gasoline from Sasol I are the primary feedstock. Here, two naphtha
crackers produce nearly 125,000 tonne/yr of 98% pure ethylene. In addition to
r
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Approximately 25,000 tonne/yr of styrene and 30,000 tonne/yr of 1,3-butadiene
are produced in this facility.
Sasol II and III
After the Arab Oil embargo South Africa re-evaluated its synthetic-oil
production potential. This re-evaluation resulted in the contruction of two
new synthetic oil production facilities located at Secunda, 140 km northeast
of Sasolburg and 130 km east of Johannesburg. Like Sasol I, Sasol II and
Sasol III utilize Lurgi/Fischer-Tropsch technology and are mine-mouth
facilities. Figure 4 is a simplified flow diagram of Sasol II which was
commissioned in 1981. Sasol III is located adjacent to Sasol II and is a
mirror image of this facility. Commissioning of Sasol III is scheduled for
mid-1982. A view of Sasol II in the construction phase is shown in Figure 5.
Figure 5. A VIEW OF SASOL II PLANT. TOGETHER, SASOL II AND III WILL OCCUPY
6 SQUARE MILES
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
Coal for the Sasol II and III facilities is being produced at the
Bosjesspruit mine just south of the facilities. Total production of the four
shaft mine which started operation in 1979 is estimated to be 27.5 million
tonnes/yr. The mine, operated by Sasol I, is part of the Highweld coalfield
and employs 6,600 people. The life of this mine at anticipated production
levels has been estimated to be over 70 years. The typical coal properties
for this relatively high quality bituminous coal are shown in Table 3. Coal
from the mine is delivered to the synthetic fuels facility by conveyor.
Type Bituminous
Moisture (as received) 5.5% wt
Ash (moisture free) 22.5%
Volatiles (moisture free) 24.8%
Fixed carbon (moisture free) 52.7%
r
Heating value 23.9-24.5 Mj /kg
Carbon (daf) 76.9%
Hydrogen (daf) 4.3%
Sulfur (daf) 1.3%
Nitrogen (daf) 2.0%
Oxygen (daf) 13.6%
Ash properties
Softening point 1290?C
Melting point 1330?C
Fluid point 1360?C
Sasol II is equipped with 36 Lurgi Mark IV high pressure gasifiers; 30 on
line and 6 standby units. These gasifiers, which were tested for 3 years at
Sasol I, weighed 140 tonnes each, have a 3.85 m internal diameter and operate
at 27 bars pressure. A total of 1.65 million m3/hr of raw gas will be
produced by these gasifiers with the composition shown in Table 4. Total
gasifier section coal consumption will be from 25,000 to 30,000 tonne/day., In
addition to coal these gasifiers will also require 30,000 to 36,000 tonne/day
of high pressure steam and 8,000 to 9,000 tonne/day of 98.5% pure oxygen.
Total oxygen requirements for Sasol II is nearly 12,000 tonne/day at
approximately 500 psig (8,600 to the gasifiers and 3,400 to the reformers).
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Table 4. RAW GAS FROM GASIFIERS AND PURE GAS FROM RECTISOL UNIT COMPOSITIONS
Raw Gas Composition
Pure Gas Composition
r
I
H2 38.1 volume H2 56% by volume
N2 0.3 CO 28%
CO 19.0% CO2 1.5%
CO2 32.0% CH4 13.5%
CH4 9.4% C2H4 0.02%
C2+ 0.5% C2H6 0.20%
H2S 0.7% N2 0.08%
A 0.46
S 0.07 ppm
Gas produced in the gasification section will be scrubbed in a
Phenosolvan plant to recover approximately 120,000 tonne/yr of anhydrous
ammonia. Tars and oils separated from the raw gas stream before di-isopropyl
treatment in the Phenosolvan plant (about 200,000 tonne/yr) will be used to
produce creasotes, road tars, pitch and coal tar fuels. Naphthas will be
recovered for further processing to high octane gasoline components.
After the raw gas has been scrubbed it is sent to the Rectisol section
which consists of 4 process trains. These units produce about 1,100,000 m3 of
pure gas with the composition shown in Table 4. Gases absorbed by the
Rectisol units -70?C methanol wash are sent to a Stretford section where
90,000 tonnes/yr of elemental sulfur are recovered. Pure gas from the
Rectisol section and reformed tail gas from the Synthol section are blended
from the fresh feed stream to the Synthol section.
Based on Synthol and Arge operating experience at Sasol I it was decided
that the Synthol Fischer-Tropsch technology was most appropriate for Sasol II
and III. The Badger Companies scale up and designs were used to construct
eight Synthol units at the Sasol II facility. Each unit has a capacity of
about 300,000 to 350,000 m3/hr of raw gas. The Synthol reactors operate at
340?C and approximately 340 psia. It is estimated that these Synthol reactors
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will consume 1.5 million m3/hr of pure gas. Catalyst for these reactors is
prepared on site. The product slate produced by the Synthol section is
presented in Table 5. Unreacted pure gas and tail gases from the Synthol unit
are recycled to a cryogenic separation section before sending the 225,000
m3/hr of 90% pure methane output to eight methane reformer units. Methane
reformer output is recycled into the Synthol reactor section.
Table 5. PRODUCT SELECTIVITY FOR THE SASOL TWO SYNTHOLS
(mass % basis)
Methane
Ethane/Ethylene
Propane/Propylene
Butanes/Butylenes
C5 to 195?C fraction
190?C to 400?C fraction
400?C to 520?C fraction
Heavier than 520?C fraction
Chemicals
0.5
6.0
100.0
The liquid output of the Synthol reactors is sent to an oil upgrading
section after oil stabilization. This section, shown in Figure 6, consists of
a fractionation, vacuum distillation, naphtha hydrotreating, catalytic
reforming (Platforming), catalytic condensation/polymerization, polymer
gasoline hydrogenation distillate finishing and gasoline blending section.
Process design for this section was done by Mobil, UOP and Linde. Total
liquids production from Sasol II is estimated to be 2.1 million tonne/yr.
This breaks down to 1.5 million tonne/yr of motor fuels, 160,000 tonnes/yr of
ethylene, 200,000 tonnes/yr of tar product, 100,000 tonnes/yr of ammonia and
90,000 tonne/yr of elemental sulfur. Sasol III production will be similar,
but emphasis will be placed on maximizing motor fuel production and minimizing
other byproducts. A summary fact sheet of the Sasol technology is presented
in Table 6.
I N S T I T U T E O F G A S T E C H N O L O G Y
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Decanted oil Vacuum
Distillate
distillation
Figure 6. SASOL II OIL UPGRADING SECTION
The South African government has recognized the country's dependence on
imports of transportation fuels as far back as 1927. Research on German
indirect liquefaction technologies in the '30's and '40's culminated in the
government support of Sasol I. This facility has been in operation since 1955
producing a broad range of chemical and transportation fuels. It is the gov-
ernment's goal to increase transportation fuel production from coal in South
Africa to lessen the countries dependence on foreign imports. Sasol I, II,
and III are expected to supply an estimated 40 to 50% of South Africa's liquid
fuel needs by 1985.
Relationship to Prior Technology
Sasol Technology is based on Lurgi gasification and Fischer-Tropsch syn-
thesis. The Lurgi gasification process was developed in Germany in the 1930's
and the Fischer-Tropsch synthesis reaction was discovered in 1925. Within
10 years of the Fischer-Tropsch discovery the first plant was in operation.
By 1944 nine F-T plants were producing 560,000 tonnes/yr of fuel in Germany.
Total world F-T capacity at that time is estimated to be 1.1 million tonnes/yr.
South Africa has an abundance of minerals but in all the exploration
activities no discoveries of important oil deposits have been found. In 1927,
a South African White Paper was published discussing the available processes
for production of oil from coal. Developments in Germany were closely fol-
lowed with particular interest in the Fischer-Tropsch process. One of the
'hydrotreating ,?,'.,.... j
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South African mining corporations, the Anglo Transvaal Consolidated Investment
Co., better known as Anglo Vaal, acquired in 1935 the South African rights to
the German Ruhrchemie/Lurgi Arge Fischer-Tropsch process and the American
Hydrocarbon Research Inc. Hydrocol Fischer-Tropsch process.
During the next few years Anglo Vaal devoted much attention to the devel-
opment of a scheme for the production of oil-from-coal. During the war and in
the post-war years, Anglo Vaal remained in close contact with developments.
In 1943, negotiations were held in America which led to the procurement of the
rights of the American M.W. Kellogg Synthol variation of the Fischer-Tropsch
process. In 1946 a new study was made and an application made to the govern-
ment to create a suitable framework within which a long-term industry could be
established. During 1947 the Liquid Fuel and Oil Act was passed and in agree-
ment was reached between the South African government and Anglo Vaal in which
the Anglo Vaal rights were taken over by the government. The South African
Coal, Oil and Gas Corporation Ltd. was formed and incorporated under the
companies' act as an ordinary public company.
Though it was clear that the plant would be based upon the synthesis of
hydrogen and carbon monoxide as invented and developed by Fischer and Tropsch,
it still had to be decided which processes to choose for the individual steps
in this integrated complex. For gasification the Lurgi pressure gasification
with steam and oxygen was selected because this process had already been dem-
onstrated in gasifiers of a smaller size. It had the advantage of being able
to work on the rather low grade, high ash coal available to Sasol. The fact
that it operated at a pressure of approximately 350 psi which is also the
desired operating pressure for the Fischer-Tropsch plant, was an additional
advantage. This avoids cumbersome compression of large volumes of gas arising
from low pressure gasification.
The raw gas from such a gasification system contains, of course, apart
from the hydrogen and carbon monoxide, appreciable quantities of undesired
components such as unsaturated hydrocarbons, sulphur compounds, etc. More-
over, the raw gas contains a large percentage of carbon dioxide which has to
be brought down to a lower level. A number of possibilities to purify the gas
existed, all involving at least two or three different process steps. How-
ever, in the late '40's Lurgi in cooperation with Linde of Germany, had devel-
oped a combined gas purification process (Rectisol) which used one solvent,
I N S T I T U T E O F G A S T E C H N O L O G Y
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methanol, at low temperatures as the single absorption agent to remove all
undesired components from the gas. A full size plant did not exist, but the
pilot plant work was convincing enough to justify its selection for the Sasol
plant.
The gasification process produces as a side-stream a gas liquor in which
components such as ammonia and phenols are dissolved. Obviously, such a gas
liquor cannot be disposed of before treatment and for this treatment another
"first" was chosen, the Lurgi Phenosolvan process in which phenols are
extracted from the water with a solvent such as butyl acetate. The ammonia
can then be recovered by stripping with steam and converted into, for
instance, ammonium sulphate. An additional advantage of having the gasifi-
cation, gas purification and gas liquor treatment all from one process know-
how supplier, was that the responsibility for the performance of these plants
which are all to a certain extent inter-related, was concentrated with one
company.
On the Fischer-Tropsch process itself the choice was not easy. On the
one hand there was the German Arge design which used a fixed bed reactor
system which was developed by Lurgi in Germany and known to work. The reac-
tion took place in long tubes surrounded by a bath of boiling water for tem-
perature control and the only difference between the small demonstration
reactors and the proposed reactors for the Sasol plant, was in the number of
tubes in one shell. This was not expected to give scale-up problems. On the
other hand there was the American Synthol developed moving bed reactor type
using a fluidized catalyst on which only pilot plant data was available but
which offered the opportunity of building reactor units with a much higher
capacity. Though the basic chemistry for both reactor types is the same, the
fixed bed reactor produces in general straight chain hydrocarbons with a high
average molecular weight and most of the production is in the range of diesel
oil and paraffin waxes. The fluid bed process produces branched olefins of a
low average molecular weight and the production is mainly in the range of LPG
and gasoline. In view of the uncertainties the wise decision was taken to
build two synthesis plants in parallel using both systems.
The original flow sheet made provision to send approximately two thirds
of the pure synthesis gas to the fixed bed reactors and to send the tail gas
of that system with the increased methanol content, to the reforming plant
1
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I N S T I T U T E O F G A S T E C H N O L O G Y
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where it was converted together with the remaining one third of the pure gas
into feedgas of the right composition for the fluid bed plant. The tailgas of
this plant is also recycled to the reforming units after removal of carbon
dioxide. Detail design of the plants was done by M. W. Kellogg. Construction
started towards the middle of 1952 and the first units were put into operation
towards the end of 1954. By the end of 1955 all the main construction work
was completed.
Operating Facilities
Sasol technology is being commercially utilized at Sasol I in Sasolburg
and at Sasol II and III in Secunda, South Africa. Sasol I produces a host of
products ranging from synthesis gas, chemicals, motor fuels to waxes, tars,
and pitch. Total output is in the range of 14,600 to 18,000 bbl/day of liquid
fuels and chemicals. Output for each of the Sasol II and III facilities is
estimated to be 30,000 bbl/day of motor fuels 160,000 tonnes/yr of ethylene,
50,000 tonnes/yr of other chemicals, 200,000 tonnes/yr of tar products,
100,000 tonnes/yr of ammonia, and 100,000 tonnes/yr of elemental sulfur.
Major Funding Agencies
Anglo Transvaal Consolidated Investment Ltd. (Anglovaal) turned over the
rights of the Sasol technology to the South African government in 1950. The
government in turn formed the South African Coal, Oil and Gas Corporation
Limited with funding coming from the treasury via the Industrial-Development
Corporation (IDC). The IDC is the state corporation that has provided much of
the venture capital to such high risk projects. The first Sasol facility at
Sasolburg was constructed with this government funding at an estimated cost of
$230 million in 1952-1955. The IDC also provided the loan capital for expan-
sion activities at Sasol I throughout the years. As sole owner of Sasol I the
IDC received all profits.
The financing of Sasol II will come from the State Oil Fund (1.711 billion
Rands), export credits (492 million Rands), and parliamentary grants
(300 million Rands). Export credits will provide 20% of the financing of
Sasol III; the balance will be provided by the State Oil Fund, parlimentary
grants, and the proceeds of two stock issues in a newly formed company called
Sasol Ltd. The issues were made in September and October of 1980.
245 million shares were offered to institutional investors and 17 million
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shares to the public. Cost per share was 2 Rands. The response was tremen-
dous, with more than $1 billion offered for each of the two issues. These
shares account for 70% of the stock in the new company Sasol Ltd.; the
remaining 30% is assigned to the Industrial Development Corp. (IDC) and its
subsidiary Konoil Ltd. Sasol Ltd. will have 100% ownership in Sasol I and 50%
ownership in Sasol II and III; the balance will be controlled by the state
through IDC/Konoil. Once Sasol II and III are operative, further shares will
be offered to the public from time to time. This dilution of government
involvement reflects policies aimed at limiting the government's role in the
economy.
In 1979 when Sasol was restructured, Sasol Limited was created as the new
holding company. The new corporate structure is shown in Figure 1 along with
ownership percentages. The old Sasol is now known as Sasol I and is a wholly
owned subsidiary of Sasol Limited. The Sasol I subsidiaries that have
developed throughout the years to give the company its diversified energy
business are shown in Table 7. This table presents the Sasol I subsidiaries
and their principal activities. In addition to the companies presented in
Figure 1 and Table 7, Sasol also oversees South Africa's strategic petroleum
reserve which has been stored in underground coal mines.
Sasol I has been commercially proven with over 25 years of operation.
Sasol II which started full operation in 1981 did not experience any major
startup problems. Sasol III which will by fully commissioned in 1982 is not
expected to have any startup or operating difficulties since it is based
entirely on technology used in Sasol II.
Capital Costs
See Major Funding Agencies section.
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Table 7. SASOL ONE SUBSIDIARY COMPANIES
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1. Sasol Marketing Company (SMC) - markets petroleum products of the Sasol
group excluding road binder material, tar and bitumen.
2. South African Gas Distribution Corporation (Gascor) - Distributes
industrial gas by pipeline from Sasol One to the industrial complex of the
Witwatersrand and Vaal Triangle. It has also been designated as the
distributor of any natural gas that may be discovered in South Africa.
3. National Petroleum, Refiners of South Africa (Natref) - Refines crude oil
which is pumped by pipeline from the coast 600 km away for Sasol One into
LPG, gasoline, diesel oil, kerosene, jet fuel, bitumen and other
products. Sasol One owns 52.5%. The other partners are the National
Iranian Oil Company (NIOC), 17.5% and Compagnie Francaise des Petroles
(Total), 30%.
4. Sasol Dorpsgebiede (SDB) - Undertakes township development at Sasolburg
and provides housing for the Sasol's group's employees.
5. Inspan Beleggings - Holds the major portion of the coal rights of
Bosjesspruit's coal fields.
6. Leslie Coal Development Company - Holds Sasol One's longer-term coal
rights.
7. Allied Tar Acid Refiners (Atar) - Refines tar acids in one of Sasol One's
factories. Atar also markets phenols and cresylic acid through SMC.
8. Tosas - Holds a 50% interest in FTS Binders which in turn markets road
binder material, tar, and bitumen.
9. Southern Oil Exploration Corporation (Soeker) - Sasol One has a 50% share
of Soekor which is to lead and coordinate the search for oil in South
Africa on behalf of the government with state funds.
10. Fedgas - Markets industrial gases such as oxygen nitrogen and argon some
of which are supplied by Sasol One. Sasol One has a 20% interest in
Fedgas.
11. Inspan Bedryf, Sasol Konstruksiemaatskappy and naftachem are dormant
companies.
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SASOL ONE AND
SUBSIDIARIES
INDUSTRIAL
DEVELOPMENT CORP.
I
5oz 50%
OWNERSHIP (SHAREHOLDING)
LOANS
Figure 7. SASOL GROUP ORGANIZATION
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24. Modderfontein (African Explosives and Chemicals Industries Ltd.)
Coal-to-Methanol Process
Process Description
The AECI Ltd. has proposed to produce 800,000 metric tons/yr of methanol
from coal for blending into gasoline stocks. The approach now being employed
at their Modderfontein facility to produce methanol uses commercially avail-
able Koppers-Totzek (K-T) and ICI technology. Current capacity of the Modder-
fontein facility is 1,000 metric tons/day of ammonia and approximately
90 metric tons/day of methanol.
The K-T gasifiers currently used at the Modderfontein facility are of the
two-head design, but future expansion plans will probably employ the four-head
design. Both designs are based on high-temperature, atmospheric-pressure,
entrained-bed concepts which utilize pulverized coal. This process was first
developed in Germany in the 1930's. The gasifier's characteristics enable it
to use most types of coal to produce a clean synthesis gas consisting of
chiefly CO and H2, with few if any hydrocarbon contaminants.
The coal feedstock for the Modderfontein facility is delivered by rail
from a mine 90 km away. A typical analysis of the feed coal is shown is
Table 1. The coal is unloaded by one person from the incoming coal cars to a
conveyor belt which is also operated by one person. Coal is stored in bunkers
which is typically sufficient for two weeks of operation. During winter
months the coal storage is larger due to increased regional demand. From coal
storage the coal is pulverized and simultaneously dried to about 1.5% moisture
in two ring and ball mills. The resulting coal dust particles are typically
sized 90% less than 90pm. The pulverized coal is transported to a network of
bunkers, and fed to the gasifier via screw feeders.
The prepared coal entering the gasifier is entrained into a stream of
premixed oxygen and steam and the reaction mixture enters the gasifier via
burner nozzles located in the gasifier heads. Very rapid exothermic reactions
occur causing the temperature in the flame to exceed 2000?C. Subsequent
endothermic reactions and cooling by the steam jacket gasifier wall reduces
the overall temperature within the gasifier to 1600?C. Residue time of the
coal in the gasifier is typically about 0.5 to 1 second. The Modderfontein
facility currently utilizes six two-headed K-T gasifiers. A diagram of the
K-T two headed gasifier is shown in Figure 1.
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Ultimate analysis (dry basis) % m/m
C
H
N
0
S
Ash
Inherent moisture, % m/m
Volatile combustible matter, %
(air dry basis)
64.3
3.7
2.3
8.6
0.6
20.5
1.9
t
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Figure 1. K-T GASIFIER
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A portion of the mineral matter in the coal is slagged in the gasifier
and impinges on the walls where it accumulates and flows down to a slag outlet
in the base. The molten slag is then quenched and granulated in a water bath
and removed by a continuous scraper system. The remaining mineral matter
leaves the gasifier as fly ash with associated unreacted carbon. The exit
temperature of 1600?C requires that the still molten fly ash be quenched with
direct water injection to about 900?C. This avoids fouling in the waste heat
boiler system where steam is raised at 55 bars. The gas is further contacted
with water in a washing tower where most of the solids are removed. The
synthesis gas is subjected to further dust removal before passing to a raw gas
recompression section prior to gas purification. A typical composition of the
raw gas at this point is shown in Table 2.
Typical Analysis of Raw Gas by Volume I (Dry Basis)
CO 58%
H2 27%
C02 12%
CH4 100 ppm
H2 0.5%
COS 0.04%
S02 0.1 ppm
HCN 100 ppm
NO 30 ppm
NH3 15 ppm
N2 0.9%
Ar 0.6%
02 100 ppm
A block flow diagram of the whole coal to methanol process is shown in
Figure 2. The dust-free raw gas from the gasification plant is compressed to
30 bars in 2 parallel stream turbine-driven raw gas compressors. The com-
pressed gas is then sent to the first stage of the two-stage Rectisol gas
purification unit. Gas entering this stage is first treated by a water
absorber to remove HCN and then scrubbed with cold methanol to remove H2S and
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COS to less than 1 ppm. Significant amounts of CO2 are not removed in this
unit.
Oxygen
plant
Co., prep,
drying
and
pulverizing
Coal
Raw
gas
comp.
Gasification
and heat
recovery
T
Agh
r.
ASh
Shift
conversion
Sulfur plant
and tail gas
treatment
Acid
gas
removal
Flue gas
Sulfur
Figure 2. COAL-TO-METHANOL PROCESSING USING KOPPERS-TOTZEK
GASIFICATION PROCESSING
MEOH
synthesis
loop
Crude
MEOH
The gas is then recompressed to 50 bars before injection into the water-
gas shift unit where the CO-to-H2 ratio is adjusted for the methanol synthesis
reaction. The methanol synthesis process used in the Modderfontein plant is
of the ICI design. This unit, shown in Figure 3, consists of compressing the
makeup gas, mixing this makeup gas with recycle gas, and then feeding the
mixture to the methanol converter. The two overall reactions in the catalyzed
bed occur as follows:
CO + 2H2 + CH3OH
C02+H2;CO+H20
The hot effluent is cooled by a waste heat recovery unit, heat exchange
with incoming feed, and cooling water. The liquid-rich stream is finally
flashed to remove gases, which are then recycled to the reactor. The cata-
lytic methanol reactor uses a copper based catalyst to operate at a tempera-
ture of 400? to 575?F and a pressure of 750 to 1,500 psig. The crude methanol
requires further distillation to produce a pure product for more fuel blending
or can be sent to a Mobil MTG reactor to produce gasoline.
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In the Mobil process, shown in Figure 4, methanol is partially dehydrated
to an equilibrium mixture of methanol, dimethyl ether, and water over a
dehydration catalyst in a limited reactor. In a second reactor, a zeolite
conversion catalyst is used to convert both methanol and dimethyl ether to
high-octane gasoline. During this operation recycled gas is used as a heat
sink to remove the exothermic reaction heat. In the overall process the
reaction can be described as follows:
2CH3OH + CH3 -O-CH3 + H2O
CH3 -O--CH3 + 2(-CH2) + H2O
Typical yields of the raw product from this process are presented in
t
Table 3. The hydrocarbon product is primarily gasoline which must undergo
further treating to add butanes and alkylate the butenes and propylenes. The
finished gasoline product in Table 4 typically has a Rvp of 9 psig and
unleaded RON of 93.
Figure 3. METHANOL SYNTHESIS
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Yields. wt % of charge
Methanol + ether
Hydrocarbons
Water
Co. C06
Coke, other
Hydrocarbon product, wt %
Light gas
Propane
Propane
i-Butane
n-Butane
Butanes
C. + gasoline
Gasoline (including
alkylate)
LP-gas
Fuel gas
TYPICAL YIELDS FROM CRUDE METHANOL USING
FIXED BED PROCESS
Components, wt. %
0.0 Butanes
39.2 Alkylate
93.2 C. + syethasizsd gasoline
0.3
0.2
100.0 Composition, vol %
Paraffins
Olefins
1.4 Naph hen.
6.6 Aromatics
0.2
8.9
3.3
1.1
70.9
Research octane
Clear
Leaded, 3 cc TEL/U.S. gal
100.0 Reid vapor pressure, psig
Specific gravity
Sulfur. wt %
85.0 Nitrogen. wt %
93 RON clear, 9 Rvp) Corrosion, copper strip
13.0 ASTM distillation, 'F
1.4 10%
About 3000 stpd tnatltanol (100% basis) are raquind b produce 10,000 bpd
total liquid products (90so1Mts plus LPG)
30%
60%
g0%
2.T
a2
11a.1
gS .
100
9.0
0.730
Nil
Nil
1A
114
146
198
330
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Crude
methanol
Hydrocarbon
liquid product
t
Conversion Cooler Separator
reactor
Figure 4. MOBIL METHANOL-TO-GASOLINE PROCESS -
FIXED BED OPERATION
In the spring of 1980 AECI Ltd. announced plans to construct an
800,000 metric ton/yr methanol facility based on coal gasification technology.
This decision stemmed from a South African government announcement to grant
private-enterprise fuel-from-coal tax breaks similar to those granted SASOL.
The methanol produced in the AECI Ltd. facility would be used as a blending
feedstock (up to 10%) to stretch existing gasoline and diesel supplies. How-
ever, reports from the AECI Ltd. in 1981 indicated that methanol may not be
more desirable. Plans to go ahead with this development have yet to be
announced.
Relationship to Prior Technologies
The technology used to produce methanol has already been commercially
developed. This includes the K-T gasifier which was developed in Germany in
the 1930's and the ICI methanol process developed in England in the late 50's
and early 60's. If gasoline is to be the final product, the Mobil MTG process
will also be required. This process is currently in the pilot plant phase,
but, a large commercial facility is being built in New Zealand using the Mobil
MTG process.
Operating Facilites
AECI Ltd. currently operates a 1000 metric ton/day ammonia plant and a
90 metric ton/day methanol plant which uses coal as the feedstock. This
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facility at Modderfontein, South Africa, started operation in 1972. This
complex, which is the largest ammonia plant in South Africa, took 30 months to
construct.
Major Funding Agencies
Privately financed by AECI Ltd.
Technical Problems
The Modderfontein plant can exceed its 1000 metric tons/day design capa-
city. The plant consists of six gasifiers, but five gasifiers are sufficient
to reach the ammonia plant design capacity. The gasifiers are arranged in a
pair of three. At start-up, the gasifiers ran at 70% of capacity, and even-
tually reached 130% of design capacity. They expect greater turndowns are
possible with the four-headed gasifier. AECI personnel felt that a spare
gasifier would have been a desirable feature. Of course, this is easier to
justify on large plants. Gasifiers can easily be taken on and off stream for
brief periods of time (say, one hour); they are kept hot during standby by
means of integral burners. They said the gasification efficiency varies with
the type of coal. They strongly recommend full-scale testing of coals prior
to selection system. They found the gasifiers sensitive to the type of coal
fed.
They estimate that the No. 4 plant has twice the amount of equipment a
conventional natural gas or oil based ammonia plant would have. Thus one
would expect more difficulty in maintaining high reliability. (They did not
want to discuss on-stream factors.) A chart on the meeting room wall during a
trip there in 1981 indicated that actual production was nearly equal to
scheduled production for the previous few months.
Their main problems are associated with the stoker-type boiler. No
redundancy was built into the system, hence they often have difficulty
supplying adequate steam. There are also considerable time delays when addi-
tional steam is demanded, further complicating operations. Cycling the steam
output is hard on the boiler components and contributes to increased mainten-
ance. They reminded us that the pulverized coal alternative to a stoker is
more difficult to operate. Another plant problem mentioned is erosion of
process surfaces. These repairs are made during their overhaul period, which
occurs every two years (standard practice for the chemical process industry).
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They have solved the refractory lining problems that were encountered earlier.
They stated that in general their plant experiences more corrosive contaminant
than a natural gas-based ammonia plant. They have been able to slowly
increase the intervals between certain maintenance procedures through the
years as they became more familiar with the plant. In their opinion gas is
easy to make compared to the remaining processes in the plant.
They strongly felt that all portions of the process should be highly
instrumented because of the inherent possibility of fire, explosions, and
toxicity dangers. They are currently in compliance with all environmental
regulations, although they were having some recent dust problems from the coal
milling section (that had been previously reported in the literature as well,
indicating a possible chronic problem). They must obtain an operating permit
from the Air Pollution Control Officer, who has the power to shut down the
facility. They enjoy amiable relations with the environmental authorities.
South African law requires that pressure vessels be tested every four years;
this is done during one of the overhauls scheduled every two years. They
currently burn the hydrogen sulfide removed from the process. They envision
two alternatives to burning: production of sulfur from the gas via a Claus
plant, or production of sulfuric acid. They warn that heavy metals in coal
ash may be a problem in the U.S. (they are not experiencing problems). They
see no reason not to recycle about 20% of their process water.
They store a minimum coal supply of two weeks. Turing winter, when
demand is high, they prefer to have a larger supply. Coal supply problems
have not occurred.
The plant has not suffered from labor problems. Labor is represented by
unions according to race. The plant is a closed shop, but not all labor is
unionized. The plant uses a four-shift system with three shifts per day.
Gasification labor requires a part-time foreman, one supervisor, one operator
in the control room, two outside laborers, and less than one auxilliary
laborer.
The coal supply section is less sophisticated and requires one laborer to
unload the coal cars, one laborer to oversee conveyor operations and one
laborer in the coal mill. They estimate skilled labor receives a net salary
of R12,000 ($15,000) and unskilled, R6,000 ($7,500). If labor were more
expensive they postulated that they might increase automation in the coal
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handling section to eliminate one or two positions, or simply require each
worker to be responsible for more process operations. They feel that their
labor productivity is typical for South Africa.
The major process industry differences they see between the U.S. and
South Africa are coal and electricity prices. Typical coal prices are R6 and
R9 per ton (this is about $0.50 to $0.70 per million Btu). Electricity in the
vicinity of the plant costs about R0.01 per kWh (13 U.S. mills/kWh) and in the
Cape about R0.021 per kWh (26 U.S. mills/kWh).
Capital Cost
Construction of the 800,000 metric ton/yr methanol plant was estimated to
be $549 million in 1980.
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SOUTH AFRICAN COAL LIQUEFACTION PROCESS
25. Sasol Direct Coal Liquefaction Process
Process Description
Little, if any, information has been published on Sasol's direct lique-
faction technology development program. Information published by the Nippon
Brown Coal Liquefaction group of Japan has indicated that Sasol has developed
an SRC-II type pilot plant process which utilizes catalytic hydrogenation
technology. However, published verification of this facility has not been
found. In addition, Sasol officials indicated in mid-1981 that direct coal
liquefaction technologies were not superior to their indirect process, espe-
cially when taking South Africa's poor coal quality into consideration.
However, bench-scale experiments have been conducted by the Fuel Research
Institute of South Africa on catalytic hydroliquefaction of South African
bituminous coals. In these experiments four different liquefaction reactors
were tested on one South African coal type. The proximate and ultimate
analysis of the bituminous coal used in these experiments is shown in Table 1.
Table 1. ANALYSIS OF WATERBERG TRANSVAAL COAL USEDa
Moisture (wt % air dried basis)
3.4
Ash (wt % air dried basis)
12.7
Volatile Matter (wt %, air dried basis)
34.8
C (wt % daf)
80.7
H (wt% daf)
5.5
N (wt % daf)
1.5
S (wt % daf)
1.0
0 (wt %, daf) (by difference)
11.3
Vitrinite (vol %)
83.2
Exinite (vol %)
4.2
Inertinite (vol %)
5.7
Ro (mean maximum reflectance of vitrinite)
0.720
a 0.5 - 0.25 pm (30 - 60 mesh)
The first process is hydropyrolysis (dry hydrogenation) with a "hot-rod"
reactor, 2-5 operated in a semi-continuous mode. In this reactor the coal is
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heated at about 200?C min-' while secondary reactions are minimized by
removing and quenching the products. The second process was hydropyrolysis in
a rotating autoclave with a heating rate of about 7?C min 7l and retention of
the product at reaction conditions for 1 hr. The other processes were super-
critical gas extraction using toluene with and without hydrogen. These
extractions were carried out in an autoclave with a lengthy heating-up period
and in a modified "hot-rod" reactor with reasonably rapid heating. Identical
temperatures and pressures were used either with the same concentration of
catalyst or without catalyst. In all the experiments, sand was mixed with the
coal to limit agglomeration.6 Thermolysis of supercritical toluene is a
problem; however, the products of the thermal breakdown are known.7
Liquefaction Procedures
Method A (Runs 1 and 2). Hydrogenation was carried out in a "hot-rod"
reactor.2-5 The coal (25 g), impregnated with stannous choloride catalyst
(tin 1 wt % of the coal) for the catalyst run, was mixed with sand (1:2 by
weight). The reactor was heated (about 200?C min-') to 450?C and maintained
at this temperature for 15 min. Hydrogen (20 mPa; 221 min-') was passed
through the fixed bed of coal/sand/catalyst. The products were condensed in
high-pressure cold traps.
Method B (Runs 3 and 4). The reactor was a 1 liter rotating autoclave
fitted with a glass liner. Coal (50 g) sand catalyst preparation procedures
were the same as in Method A. The charge was heated (about 7?C min-') under
hydrogen to 450?C and maintaned at this temperature for 1 h (pressure 20 MPa).
The product was washed from the cooled reactor system with toluene. The
solid residue was extracted with boiling toluene (250 ml) in a Soxhlet
extractor for 12 h. The toluene solutions were combined and the toluene
removed under reduced pressure. Hexane (250 ml) was added to the extract and
it was allowed to stand for 24 hr with occasional shaking. The solution was
filtered to leave a residue (asphaltene) and the hexane was removed from the
filtrate under reduced pressure to give the oil. The residue in the Soxhlet
thimble after toluene extraction was extracted with pyridine in a Soxhlet
extractor to give pre-asphaltenes (toluene-insoluble pyridine-soluble
product).
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Method C (Runs 5-8). Supercritical gas extraction was carried out in a
stirred 1 liter autoclave. Coal (100 g) sand catalyst mixtures were the same
as that for Method A. In the runs without hydrogen, the autoclave was charged
with coal/sand/catalyst and 600 ml of toluene and heated. It took about 1 h
to reach the required temperature and pressure. On reaching the desired
temperature (450?C) and pressure (20 MPa), toluene (about 2 liter l i-1) was
pumped into the reactor through the coal bed via a dip tube. The toluene
extract was cooled at atmospheric pressure in a water-cooled condenser. The
runs with hydrogen were similar except that only 300 ml of toluene was loaded
into the autoclave, which was then flushed and pressurized with hydrogen to a
cold pressure of 5 MPa. After attaining the reaction temperature and pressure
(450?C and 20 MPa), the autoclave was maintained at these conditions for 1 h
before starting to pump the toluene.
The extraction condensate was filtered to remove material which precipi-
tated on cooling. This residue was soluble in pyridine and designated as pre-
asphaltene. The toluene was removed from the filtrate under reduced pressure,
and this product was then treated with hexane as described to give the
asphaltene and oil fractions.
Method D (Runs 9-12). The apparatus used was the same as for Method A
except that provision was made for the introduction of toluene as well as
hydrogen. Coal/sand/catalyst preparation procedures were as for Method A. In
the runs without hydrogen, toluene (4 liter h 1; 20 MPa) was passed through
the reactor for 15 min on attaining the reaction temperature. In the runs
with hydrogen, hydrogen (8 liter min-') was simultaneously passed through the
reactor. The pressure was maintained at 20 MPa. The toluene extract was
collected in the condenser system.
The conversions and product distributions are shown in Table 2. In the
case of supercritical gas extractions without hydrogen (Runs 5, 6 and 9), the
yield of extract was greater than the conversion due to the thermal breakdown
of toluene. The gas yields (including water) were obtained by difference and
in the experiments where supercritical toluene was used, these will be low
owing to errors caused by formation of toluene pyrolysis products. G.I.C.
analysis of the oils from supercritical amounts of bibenzyl and the other
toluene pyrolysis products.
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Molecular weights, viscosities, and sulphur analysis are given in
Table 3. Sulphur analysis was not carried out on the supercritical gas oils
(Runs 5-8) because of the large amount of toluene pyrolysis product in the
oils. Thermogravimetry in nitrogen of the four hydrogenation oils (Samples
1-4) were very similar. About 60% weight loss had occurred at 200?C and
essentially 100% weight loss had occurred by 400?C. Thermogravimetric
analysis in nitrogen of the asphaltenes showed that the asphaltenes from the
supercritical gas extraction without hydrogen assistance were less volatile
than the other asphaltenes.
The fractions from the silica gel chromatographic separation were grouped
as aliphatic hydrocarbons, aromatic hydrocarbons, and polar compounds. The
results for the oils from the hydrogenation samples are summarized in Table 4.
The samples from the rotating autoclave contain a lower percentage of polar
compounds than those obtained from the equivalent "hot-rod" experiment.
Because of the large quantity of toluene pyrolysis product in the oils from
the supercritical extractions, these oils have not been grouped as aliphatic,
aromatic, and polar compounds.
Table 2. PROCESS. CONDITIONS. CONVERSION AND PRODUCT DISTRIBUTION
Conversion
Extract yield
oil
Asphaltene
Pre-asphaltene
Gas
1
HR
HR
H2
38.7
-
13.2
13.2
6.8
5
5
2
RA
H2, SnCl2
90.6
-
29.4
10.9
5.9
.
44
4
3
H2
56.1
-
13.0
3.5
3.1
.
36
5
4
RA
H2, SnCI2
79.6
-
24.5
6.8
1.8
.
46
5
5
SCGE
32.4
39.6
14.7
13.8
10
2
.
6
7
SCGE
SCGE
SnCI2
H2
29.0
41
7
36.0
4
32
14.4
16
8
11.6
.
8.2
8
SCGE
H2, SnCI2
.
69.5
.
52.1
.
19.9
13.1
25.1
2.7
4
4
9.3
17
4
9
10
SCGE/HR
SCGE/HR
SnCl
31.2
32
6
32.6
4
31
8.6
15.4
.
8.6
.
11
SCGE/HR
2
H2
.
34.0
.
24.0
7.3
7.7
15.6
13.2
8.5
1
3
1.2
10
0
12
SCGE/HR
H2. SnCI2
49.2
44.5
12.0
26.4
.
6.1
.
4.7
a HR, Hot-rod reactor; RA, rotating autoclave; SCGE, supercritical gas extraction using autoclave; SCGE/HR, supercritical gas extraction using
hot-rod reactor
Table 3. MOLECULAR WEIGHT, VISCOSITY AN SULPHUR ANALYSIS OR PRODUCT
Molecular wt. Viscosity
of oil wt % S
Run Oil Asphaltene (mPa s) in oil
1 221 398 62 0.40
2 197 330 40 0.33
3 194 352 33 0.30
4 194 351 17 0.24
5 214 487 15
6 215 49. 13
7 222 427 34
8 210 428 22
9 237 0.52
10 235 0.51
11 248 0.46
12 230 0.32
1
1
1
t
1
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1
1
Run
Alphatics
(wt % oil)
Aromatics
(wt % oil)
Polar compounds
(wt % oil)
1
8
48
44
2
6
55
39
3
5
59
36
4
4
70
26
The conversion of coal to liquid and gaseous products was highest for the
"hot-rod" reactor when a catalyst was used (Table 2). The conversion was much
lower without a catalyst. Using the rotating autocalve the effect of the
catalyst was far less pronounced presumably owing to the much longer residence
time. As expected, when the product is held in the autoclave, gas production
at the expense of liquid product increases. The amount of asphaltenes and
pre-asphaltenes22 (or asphaltols23) is higher for the "hot-rod" reactor
experiments than for the autoclave experiments. With time, asphaltenes and
pre-asphaltenes are converted to oil and gas. The higher conversion obtained
with the "hot-rod" reactor with a catalyst compared to the equivalent rotating
autoclave run may indicate that polymerization of part of the product does
occur to give mainly a pyridine-insoluble char. However, it is possible that
the higher rate of heating in the "hot-rod" reactor may produce more intense
thermal fragmentation and thus higher conversion.
Different hydrogenation processes even when carried out at the same
pressure and temperature on the same coal, give considerably different
products. The liquid product from hydrogenation in a rotating autoclave
contains less hetero-atoms and is more aromatic than the product from a short
residence time semicontinuous reactor working at the same pressure and
temperature. These differences may have important implications for the
further processing of the product to transportation fuels and petro-
chemicals. The long residence time of batch autoclaves tends to "mask" the
effect of a catalyst compared to short residence time reactors. Therefore,
catalyst screening in batch autoclaves may only be of limited value when
applying these result to short residence time reactors. Supercritical gas
extraction gives a more aliphatic liquid product, and even in the presence of
hydrogen and catalyst is a milder process than hydropyrolysis at the same
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temperature and pressure. The work with supercritical gas extraction has
shown the problem that solvent breakdown can have on the yield and the nature
of the liquid obtained in a batch process.
Relationship to Prior Technology
Based on catalytic hydrogenation and other technologies; however, a
process configuration has not been established.
0 erating Facilities
Bench-scale studies are being carried out at the Fuel Research Institute
Laboratories at Pertoria, South Africa.
Technical Problems
Technology is not mature enough to project technical problems.
Capital Costs
Technology is not mature enough to project captial costs.
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t
26. Central Fuels Research Institute (Lurgi) Coal Gasification Process
Process Description
The conventional Lurgi process pilot plant was installed and commissioned
in 1962. The plant has a capacity to gasify 1 tonne of coal per hour. The
gasifier is a conventional rotary grate type gas producer designed to operate
up to a pressure of 32 kg/cm2. The gasifier shell, 1000 mm dia. and 5800 mm
length, is a double-walled alloy steel vessel of welded construction. In the
annular space between the walls, steam is raised at gasification pressure.
The inside of the gasifier is partly lined with refractory bricks in the com-
bustion zone. A rotating coal distribution (speed 7-45 rpm) is fitted inside
the gasifier at its top for maintaining a fuel bed of constant height. The
grate at the base of the generator can be rotated at speeds varying from 4 to
24 rpm. A coal lockhopper chamber for feeding the coal and an ash lockhopper
for ash discharge are provided at the top and bottom of the gasifier, respec-
tively. The tar-laden hot gases emerge out of the gasifier at a temperature
of 400?-500?C. The plant is equipped with gas cleaning system, tar recovery
system, and safety devices with adequate instrumentation wherever necessary.
The flow diagram of the system is given in Figure 1. The performance of this
gasifier on a number of Indian coals is presented in Table 1.
The goal of this research effort is to study the gasification character-
istics of noncaking Indian coals with steam and oxygen. The primary appli-
cation of this conventional technology is the production of ammonia. The
gasification characteristics of the following noncaking coals were studied in
detail:
? Talcher, bottom seam of Talcher coalfields in Orissa.
? King seam coal of Singareni (Kothagudem) field in A.P.
? Dobrana seam coal of Raniganj field in West Bengal.
e Burhar seam coal of Shahagpur in M.P.
Relationship to Prior Technology
This project is based on the commercially available Lurgi gasification
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Table 1. SOME TYPICAL OPERATION DATA OF PRESSURE GASIFICATION TESTS
Gasifier pressure, kg/cm'(g)
Dry coal feed rate, Kg/hr
Oxygen/coal (dry), Nm'/kg
Steam/coal (dry), kg/kg
Dry coal input on cross-sectional
area of Gasifier, kg/m'/hr.
Gas (raw) make rate, Nm'/hr.
Raw gas compo-ition, % by volume
CO5
CnHm
Surkachar Lower
Kenda
10.6 14.0 21.1 21.1 21.1
363 486 518 453 556
0.201 0.217 0.278 0.281 0.257
1.276 1.306 1.834 1.668 1.552
462 619 660 577 708
617 648 830 709 917
27.8
0.6
O s 0.0
CO 19.2
H, 42.7
CH4
N,
8.8
0.9
Calorific value of purified
gas .CO2=2%) K. cal/Nm, 4207
Thermal efficiency cold gas, % 73.8
Carbon gasified, % 76.4
Consumption per 1000 Nm' of raw gas
Coal (dry), kg 587
Steam, kg 749
Oxygen, Nm' 118
Consumption per 1000 Nm'
of purified gas (COs=2%)
Coal (dry), kg 799
Steam, kg. 1016
Oxygen, Nm' 160
Yield of tar/1000 kg of dry coal, litre -
Coal
Steam
Gasifier
Oxygen
Ash
Dobrana
Sample 'A' Sample 'B'
21.1 21.1 25.0
390 359 450
0.346 0.322 0.284
1.720 1.995 1.747
497 457 573
635 582 794
27.3 30.4 31.2 29.3 29.9 31.2 30.9
1.2 0.4 08 0.8 0.4 0.9 0,8
0.0 0.0 0.0 - 0.0 0.0 0.0
17.4 16.0 15.2 19.8 19.1 15.4 17.2
41.5 39.6 40.5 37.9 38.1 38.9 38.7
12.2 13.1 11.9 11.7 11.8 11.9 11.7
0.4 0.5 0.4 0.5 0.7 1.7 0.7
4242 4228 4218 4159 4074 4178 4176
75.5 75.1 78.3 78.1 75.7 76.0 74.4
75.4 79.7 81.7 83.1 86.3 82.9 82.8
750 624 639 606 615 617 567
980 1171 1066 941 1057 1231 991
163 173 180 156 212 199 161
1011 879 910 840 860 879 804
1321 1649 1519 1305 1478 1734 1406
220 244 256 216 296 283 228
- 48.0 44.0 35.0 - 53.0 57.5
Par Removal
Hot Potash
Wash Plant
Figure 1. LURGI PRESSURE GASIFICATION PILOT PLANT
- Raw Gas
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Operating Facility
The 1 tonne/hr Lurgi gasifier has been in operation since 1962 at the
CFRI in Dhambad.
Major Funding Agencies
Council of Scientific and Industrial Research, Government of India.
Technical Problems
None.
Capital Costs
The project is not expected to be scaled up; therefore no large scale
capital cost projections have been made.
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27. Central Fuels Research Institute (Bergius) Coal Liquefaction Process
Process Description
CFRI claims that they have completed experiments to liquefy coal by the
Bergius process in a bench-scale unit at a capacity of 18 kg/hr. They are
looking for foreign cooperation to scale up the process to 1 TPH capacity.
Solvent extraction of -5 mesh coal with aromatic oils was attempted in
small batch reactors. The extraction was carried out for about 3 hours at 300
to 360?C (572? to 680?F) at the vapor pressure of the solvent (10 to 15 kg/cm2
or 140 to 210 psi). No hydrogen or catalyst was added, and about 50% of the
coal was extracted. CFRI claims that the coarse coal feed enables easier
filtration of the extract. The extract could be used as a substitute
petroleum feed stock. Delayed coking of the extract produces a low ash coke
suitable for anodes (less than 0.5% ash) in electrometallurgy or a petroleum
coke substitute.
Feedstock requirements for this process include lignites, subbituminous,
and bituminous coals with high vitrian and fusain content, but with low ash
content. Higher sulfur coals are an advantage for this process since the
sulfur promotes catalyst activation. The primary output of this process is
middle distallates, particularly diesel oil, which is the preferred product in
India. In addition, hydrocarbon gases, naptha, and kerosene are also
produced. The estimated process efficiency is in the range of 65% to 70%.
Process Goals
The goals of this program are to develop indigenous coal liquefaction
technology within India to offset foreign oil imports.
Relationship to Prior Technology
This development is an offshoot of the Bergius-Piers process (I.G. Farben
Process) developed in Germany prior to World War II.
Operating Facility
The CFRI has conducted tests in an 18 kg/hr bench scale unit in Dhambad
since the mid-70's. A larger facility has not been built due to lack of
funding support.
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Major Funding Agencies
This small-scale development work is being supported by the Central Fuel
Research Institute.
Technical Problems
Materials of construction, especially in vessel lining material, and
valve life are considered the major problems. Hesser problems include
separation of heavy oils from ash solids and the production of an economical
source of hydrogen, which is required as a feedstock to this process. Oil
loss from the process has also been indicated as a problem.
Capital Costs
Capital costs for a 1 million tonne/yr liquid product facility has been
estimated to be Rs 11,000 million (Base year 1980-Rs 8.00 = $ 1 U.S.). Oper-
ating costs for such a facility are estimated to be roughly Rs 1150 million/yr,
excluding feedstock costs. Coal feedstock costs were estimated at Rs
675 million/yr.
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INDIAN COAL GASIFICATION PROCESS
28. Bharat Heavy Electricals Ltd. (BHEL) Combined Cycle Coal Gasification
Process
Process Description
The BHEL process utilizes a high-pressure, Lurgi-type fixed-bed reactor
in a combined-cycle layout to generate electric power. In this process lumpy
coal of graded size is gasified with air and steam at a pressure of 10 bars
(106 N/m2) to yield a low calorific value gas of 4960 KJ/Nm3. In addition to
the gas, tar, oil, and liquor are obtained as by-products.
The process flow diagram, shown in Figure 1, comprises the gasifier,
mixing vessel, gas cooler, and scrubber. Coal of graded size is fed to the
gasifier through a suitable feeding system. Air and steam mixed in the mixing
vessel are led into the gasifier through the tuyeres located in the rotating
grate of the gasifier. Coal coming in contact with air and steam undergoes
successive processes of carbonization, gasification, and combustion yielding a
low calorific value gas and ash. The ash is continuously withdrawn from the
gasifier through the rotating grate and ash discharge chamber. The gas leaves
the gasifier through the gas exit line.
As the gasifier operates continuously under pressure it is necessary to
have some means of introducing coal and discharging the resultant ash from the
process at atmospheric pressure. The coal feeding system and ash discharge
system provide these means, respectively.
The grate and coal distributors are two other parts in the gasifier which
regulate a uniform distribution of gas, air, steam, and coal in the gasifier.
The gasifier is also provided with a feedwater jacket where part of the steam
needed for gasification is produced.
The raw gas leaving the gasifier at about 540?C traverses through the gas
cooler and drops down to a temperature of 350?C. In the process it superheats
the steam circulated from the main waste heat boiler by indirect contact, and
in addition dust particles and part of the tar vapors are condensed and col-
lected. Part of the superheated steam is led to the mixing vessel where it
admixes with air fed from a booster and saturated steam from the gasifier.
The rest of the steam is led to the steam turbine for power generation.
B-222
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Coal Bunker Flare
Coal -
1 Steam
~
i
1
Steam
Stack
Ash
Lock
t
Waste
Water
Ash
Pump
+ AL
Dust,Tar, Dust,Tat Make up
Liquor Liquor
Figure 1. HIGH PRESSURE FIXED-BED GASIFICATION
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The gas enters the scrubber where it is contacted by either cooling water
or/and gas liquor spray. Tar and the other liquor products get condensed here
and are withdrawn from the bottom of the scrubber. The liquor is recycled
back to the scrubber. The final gas leaving the scrubber is free of tar and
alkali vapours.
The clean gas from the gasification plant is the starting material for
the gas/steam turbine combined cycle operation. This gas has a calorific
value of around 4960 KJ/Nm3. This process has been tested using Singareni
coal with the properties shown in Table 1. The cleaned synthesis gas which is
fed to the combined cycle process has the expected properties shown in
Table 2.
Process Goals
The program has the following objectives:
- To undertake technical and techno-economic evaluation of advanced power
cycle concepts.
- To identify the more promising cycle configurations, define know-how gaps
regarding hardware and systems, and initiate R&D Projects.
- To design, develop and install a combined cycel demonstration power plant
based on coal gasification to establish engineering, product nad systems
designs, operating conditions and generate reliable cost data. A combined
cycle demonstration plant of about 5 MW capacity is being installed at
BHEL, Tiruchirapalli, for which the system design has been completed and
hardware design is underway. The plant evisages a new fixed bed pres-
surized gasifier, a gas turbine of about 3 MW capacity, a waste heat
recovery boiler, and a conventional steam turbogenerator of 2 MW in a
BHEL-patented power cycle.
Relationship to Prior Technologies
The gasifier used in this process is based on the Lurgi fixed-bed design.
BHEL used the operating experience of the 1 ton/day Lurgi pilot plant gasifier
at the Central Fuels Research Institute to design a modified version of the
moving-bed gasifier using indigenous materials and components. Joint efforts
with Lurgi of West Germany were abandoned due to high costs.
Operating Facilities
A 5-MW combined cycle facility using the high-pressure Lurgi type gasi-
fier was to have been commissioned in 1980. This gasifier is located at
I
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Table 1.
SINGARENI COAL FEEDSTOCK CHARACTERISTICS AND MATERIAL INPUTS
Coal: Singareni coal
Particle size
6 t/h
6-30 mm
Ash fusion temperature:
1280?C
Net CV (KJ/kg):
17,640
Proximate analysis (Wt X):
Moisture
4.80
Ash
36.90
V.M.
25.44
F.C.
33.48
Ultimate analysis (Wt 7.):
42.82
2.82
S
0.62
N
1.00
0
5.97
MM
40.59
Moisture
4.18
7.2 t/h; 14 bar (14x105N/m2);
100?C
2.05 t/h; 14 bar; 380?C
Net CV
4930 KJ/Nm)
Steam
1.8 t/h; 10 bar (106Nm2)saturated
Tar
0.29 t/h; 9 bar (9x105Nm2); 110?C
14.76 t/h; 9 bar (9x105Nm2); 110?C
Gas analysis (Vol. %):
CO2
15.48
CO
9.20
H2
19.48
H2O
19.42
CH4
4.58
N2
31.81
H2S
0.02
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Trichy, India. BHEL is also developing a single-stage fluidized-bed gasifi-
cation unit at Hyderabad and a Koller-type single-stage gasifier at Trichy.
These two developments are summarized in Table 3.
Major Funding Agencies
Bharst Heavy Electrical Ltd. is funding the above mention gasification
development programs.
Technical Problems
The technial problems involved with the this process include hot gas
cleanup, mechanical problems associated with gasifier grid rotation, scale-up,
and lock-hopper coal feeding. In addition, ball valves are being substituted
for cone-type valves used by Lurgi the gasifier. These new valves have yet to
be proven.
Capital Costs
Capital costs for the BHEL process combined cycle facility have not been
published. However, because of the relatively dirty gas produced by the
Lurgi-type gasifier it is believed that this process will be more expensive
that other combined cycle projects such as the Cool Water Texaco project in
California.
1
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i. Single-Stage fluidized bed gasification unit (Hyderabad):
Coal throughput: 750 kg/h; particle size: 0-6 mm
Max. operating pressure: 10 bar
Operating temperature range: 1000-1300?C
Gasification media: Steam + air
Gas cleaning: cyclone, quenching pipe and venturi scrubber
Product gas: low-heating valve gas with a net heating value of
4200 kcal/kg [17.58 MJ/kg)
Status: Conceptual design finalized; detailed engineering and design in
progress.
ii. Koller-type Single-Stage Gasifier (Trichy):
Capacity: 4000 N3/h
Product gas temp.: cooled from 350-450?C to 100-110?C in a water-sprayed
pre-cooler
C.V.?f product gas: 1384-1694 kcal/Nm3 at full load
operating pressure: atmospheric
Air supply: 2600 Nm3/h
Max. hot gas efficiency: 89%
cold gas efficiency: 79% at 100% full load
Steam generated by jacket boiler: 330 kg/h at 0.5 ata
Additional details: Water-cooled jacket; automatic charging and
weighing equipment; rotatng ash pan, at 0-2.5 rev/
Status: pilot plant under performance tests.
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JAPANESE COAL LIQUEFACTION PROCESS
29. Mitsui Solvent Refined Coal Process
Process Description
The Mitsui Solvent Refined Coal (SRC) process was developed to produce a
clean burning solid fuel from coal. This solid fuel, which has low melting
characteristics, is low in ash and sulfur content. The characteristics of
this solid fuel product are presented in Table 1. Testing of this process has
been performed in a 5 ton/day pilot plant since 1977 by the Mitsui Coke Co.
In this facility coal is sized to less than 25 mm for storage in a coal
bin. From storage the coal is sent to a dry grind pulverizer and reduced to a
maximum of 100 mesh before storage in a coal weighing hopper. This hopper is
capable of feeding 5 tons/day of coal to a slurry preparation tank at a con-
stant rate. In the slurry tank a uniform slurry containing one part coal and
three parts recycled process solvent is produced using an agitator and a cir-
culation pump.
Ash
Sulfur
Carbon
Hydrogen
Low Heating Value
Specific Gravity
Softening Point
Grindability (Hardgrove Index)
0.1 wt. % max.
0.3 wt. % max.
88 - 90 wt. %
5 - 6 wt. %
ca. 9,000 kcal/kg
1.25
approx. 150?C
150
The slurry is continuously pumped from the slurry mixing tank to a slurry
preheat section prior to entering the dissolver reactor. Pressurized hydrogen
is added to the slurry stream at the inlet of the preheat section. The slurry
entering the dissolver unit is reacted at 420? to 430?C at 1075 psi for ap-
proximately 1 hour. Most of the coal entering the dissolver is decomposed to
a liquid product with the exception the coal's ash and other unreactive con-
stituents. The coal extract is then sent to a separation system, which
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1
1
t
1
t
consists of a high- and medium-pressure separation process. The coal extrac-
tion is separated into vapor phase products and liquid phase products, which
include residue. The vapor phase products are then condensed to separate the
light hydrocarbon and water from the gaseous products. The liquid condensate
is stored and the recovered gases are recycled to the preheat section. Before
recycling these gases are purified in a caustic soda and water wash system. A
part of the washed gases are purged to maintain the hydrogen content in the
gases at the dissolver. The remaining gases are compressed and mixed with
make-up hydrogen before entering the dissolver preheat section.
The liquid phase products are depressurized, cooled and sent to the
filter feed tank. The stored liquid products are put into a pressurized leaf
filter. The operation is a repetition of precoating, filtering, rinsing, and
drying, to process 6 tons of liquid products per one cycle operation. Solid
free filtrate which is obtained from the filter is sent to the light end
column feed tank. The stored filtrate is charged to the light end column and
distilled under atmospheric pressure, then separated into naphtha, wash
solvent, and SRC plus process solvent.
The naphtha and the wash solvent are transferred and stored in tanks.
The SRC plus the process solvent are heated in the vacuum flash preheater and
sent to the vacuum flash column where the SRC is separated from the process
solvent. The SRC is sent by gravity to the solidification unit where the SRC
is quenched by water and solidified. The solidified SRC is conveyed and
stored as product. The process solvent is stored and used as recovered
process solvent.
A process flow diagram for this process is presented in Figure 1 for the
5 ton/day pilot plant. This facility has been operated on Miike Coal
(Japanese) as well as Millmerran and Victorian Australian Coals. The pilot
plant test results for these coals are presented in Tables 2, 3, and 4.
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a
z
Q J
0
N
? Y S
? ZI U
J HI
N
{
.
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(1) Ultimate and Ash Analysis of Coal (dry coal basis)
Carbon
Hydrogen
Nitrogen
Sulfur
Ash
(2) Test Result
Solubility 97% (moisture ash free basis)
Chemical Hydrogen Consumption 2 wt. % (dry coal basis)
Distillate Yield 10 wt. % (dry coal basis)
SRC Yield 63 wt. % (dry coal basis)
(1)
Ultimate and Ash Analysis of Coal (dry coal basis)
Carbon
64.0 wt. %
Hydrogen
5.5
Nitrogen
1.1
Sulfur
0.5
Ash
19.7
(2)
Test Result
Solubility
Chemical Hydrogen Consumption
Distillate Yield
SRC Yield
91% (moisture ash free basis)
3 wt. % (dry coal basis)
10 wt. % (dry coal basis)
50 wt. % (dry coal basis)
(1) Ultimate and Ash Analysis of Coal* (dry coal basis)
Carbon
Hydrogen
Nitrogen
Sulfur
Ash
* Used as Baiquettes with 15 wt. % water.
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Process Goals
The goals of the Mitsui SRC Development Co. Ltd., are to develop this
process for the production of a clean fuel source as well as a source of car-
bon material for electrode production and as a source of coking additives.
As an energy source, even coals of high sulfur content, if turned into
SRC, can be converted to a clean energy source of low ash and sulfur content.
This is an important feature of SRC. In comparison with other coal lique-
faction methods the SRC process is less costly because it consumes less hydro-
gen and in addition, SRC, if heated, can be used as a liquid fuel. By pro-
cessing through hydrocracking SRC can be converted to light fuels such as
gasoline.
Light and middle distillates produced from SRC process contain some
chemical industry feedstocks such as benzene, toluene, xylenes, and phenols.
Various raw materials for chemical industry can also be produced by
hydrocracking of SRC.
As a carbon source, needle coke can be produced, with high yields through
delayed coking and calcining of SRC. This high-quality calcined coke can be
raw material of ultra-high-power electrodes (UHP).
In producing blast furnace coke, the raw material blended with SRC can
produce coke which has a high strength at elevated temperature. Therefore,
from the viewpoint of utilizing unused coal resources (noncaking coal, brown
coal) to counter the future shortage of supplies of strongly coking coal, this
application of SRC has great significance.
Relationship to Prior Technology
In 1971, perceiving the promising future of coal liquefaction by a new
solvent extraction method, namely, the Gulf Oil Corporations Solvent Refined
Coal (SRC) process, the Mitsui Mining Company was quick to proceed with a study
of the new technology. In June, 1972, the Mitsui SRC Research Consortium was
organized by four enterprises, Mitsui Mining Co., Ltd., Mitsui Coke Co., Ltd.,
and Mitsui & Co., Ltd., to establish an SRC process research and development
organization.
The consortium started liquefaction tests with more than fifty kinds of
domestic and foreign coal by the autoclave method at Mitsui Coke Company's
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Ohmuta Works, followed by the completion of a bench plant for continuous
operation in 1973 for further study of the process. In December 1974, for the
utilization of overseas leading technology, the Consortium made a contract for
joint SRC development with Gulf Oil Corporation, the parent company of the
Pittsburg and Midway Coal Mining Co., the originator of the SRC process.
In 1975 the consortium developed its own SRC process for the purpose of
producing not only SRC as a clean coal but SRC as a substitute for strongly
coking coal and coking additive for use in the production of iron-manufacturing
coke. In 1973 the consortium built Japan's first coal liquefaction plant
using the SRC process (coal feed 5 T/SD) at Mitsui Coke Company's Ohmuta Works
with technical assistance from Gulf.
While the products of Gulf's SRC process and Mitsui process are both in
the form of pitch-like solid, the raw material can be all kinds of coal, lig-
nite, and brown coal, except anthracite. In addition, the product contains
little sulfur with its heating value reaching as high as 9,000 kcal/kg, which
proves its usability not only as a clean energy but also as a substitute for
coking coal and high grade carbon material. Development is also under way to
use SRC as a substitute for strongly coking coal and coking additive under a
joint research contract made with Nippon Steel Corp. which started in April,
1979.
In the meantime, Gulf Oil Corp., the Consortium's partner, upgraded the
SRC-I process to the SRC-II process capable of producing a clean liquid fuel.
Gulf further experimented with the improved process at a 50 T/SD pilot plant
in the suburbs of Tacoma. Plans called for the construction an operation of a
6,000 T/SD plant in West Virginia for large-scale demonstration of the process
with financial support of the U.S. Department of Energy (DOE). This plan
called for the involvement of the United States, West Germany, and Japan (cost
borne by the governments of these countries). Gulf invited Ruhrkohle AG of
West Germany and Mitsui SRC Research Consortium to organize a private joint
venture corporation to carry out this plan. On the understanding that it is
Japan's only private enterprise qualified to join the venture and at the
request of Gulf and Ruhrkohle, the Mitsui Consortium decided to participate in
the project. The private joint venture corporation, SRC International, Inc.
was formally established in July, 1980. But the governments of the three
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countries decided to terminate this SRC-II project because of financial' pro-
blems in June, 1981. The SRC International, Inc., however, will be continued
to develop further coal liquefaction studies as a private joint venture cor-
poration.
Operating Facilities
A 5 ton/day pilot plant has been in operation since March 1978 at the
Mitsui Coke Company's Ohmuta Works, Goseimachi, Ohmuta-City, Fukuoka. Con-
struction, which started in September 1976, was completed in December 1977 at
a cost of 1.8 billion Yen. Operating costs have been estimated to be
400 million Yen per year. This facility has a 5 ton/day coal input and, will
produce approximately 3 tons/day of SRC. The plant is operated by four
groups, each of five persons, on three shifts. The operation is controlled
and observed from the control room except for a part of the coal handling.
Major Funding Agencies
To cope with the domestic and overseas situation, the four consortium
members established Mitsui SRC Development Co., Ltd. in February 1980 to take
over the consortium's business for the further development of SRC under a
substantially complete system. In addition, the seven companies which took
part in building the 5 T/SD SRC plant held shares in the new company to
strengthen the arrangements for further SRC commercialization. The organiza-
tional outline of this new company is presented in Table 6.
In preparation for SRC commercialization, Mitsui conducted basic tests of
Victoria Brown Coal samples and developed a special dehydration technique for
brown coal and a process of treating woody tissue. This development confirmed
that brown coal is both technically and economically advantageous as an,SRC
feedstock. The State Government of Victoria, which evaluated Mitsui's techni-
cal level, has agreed to perform a feasibility study by Mitsui after June 1980
for the construction of a commercial SRC plant in the vast brown coal field
area.
Capital Costs
Australia's CSR and Japan's Mitsui SRC Development Co. have agreed to a
50-50 joint venture to evaluate commercial production of solvent-refined coal
and liquid fuels from lignite in Australia's Latrobe Valley in Victoria. They
will spend more than $3.5 million in studying a 6,000 ton/day SRC plant using
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The Mitsui process. The study is expected to be completed by June of 1982,
and a $1.7 billion plant could be onstream as early as 1987. However, the
most recent information indicated that MITSUI/CSR has postponed the construction
of this facility. In addition MITSUI/CSR has expanded their development
goals to include coal gasification technology through Germany's Ruhrkohle.
The Mitsui SRC process is similar to the Gulf SRC development and there-
fore these processes should share common problems. Most of the major problems
have been resolved in pilot plant research.
Table 6. ORGANIZATIONAL OUTLINE FOR THE MITSUI SRC DEVELOPMENT CO., LTD.
Founded: February 20, 1980
Head Office: Mitsui Main Building, 1-1, 2-chome, Nihonbashi-Muromachi,
Chuo-du, Tokyo 103, Japan
Plant Location: 1-banchi, Goseimachi, Ohmuta-shi, Fukuoka-ken 836, Japan
Chairman: Toshikuni Yahiro
(President of Mitsui & Co., Ltd.)
President: Shingo Ariyoshi
(Chairman of Mitsui & Co., Ltd.)
Number of Employees: 72
Capital: Y500,000,000.
Shareholders: %
Mitsui & Co., Ltd. 28.5
Mitsui Mining Co., Ltd. 23.5
Mitsui Coke Co. , Ltd. 5.0
Toyo Engineering Corporation 12.0
Mitsui Engineering & Shipbuilding
Co., Ltd. 8.0
The Japan Steel Works, Ltd. 8.0
Ishikawajima-Harima
Heavy Industries, Co., Ltd. 6.0
Mitsui Construction Co., Ltd. 3.2
Mitsui Toatsu Chemicals, Inc. 3.0
Mitsui Miike Machinery Co., Ltd. 2.0
Yamatake-Honeywell Co., Ltd. 0.8
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30. Nippon Brown Coal "Kominic" Hydrogenation Direct Coal Liquefaction
Process
Process Description
In 1972 the Japanese firms of Kobe Steel, Ltd., Mitsubishi Chemical
Industries, Ltd., and Hissho-Iwai Co., Ltd., formed the Kominic group. The
purpose of this enterprise was to exploit the vast brown coal (lignite)'
deposits in Victoria State, Australia, to produce a solid solvent-refined coal
(SRC-I) fuel. With this in mind, Komonic group entered into a research
relationship with the South African Coal, Oil and Gas Corp. (SASOL) to
investigate the liquefaction characteristics of brown coal in SASOL's SRC-I
liquefaction process. In 1977 Kobe Steel constructed 500 kilogram/day process
development unit for further SRC-I studies on Victoria brown coals at the
Iwaya Works west of Osaka.5 During the late '70's Japan's as well as the
Kominic group's goals shifted away from solids production (for metallurical
coke) toward liquid fuels production. To meet these goals, Kobe Steel
developed an additional high-pressure hydrotreating stage which when added to
the original SRC-I type process produces a liquid with approximately 70% heavy
oil and 30% light oi1.2 During this development period (1979) the Komi;nic
group ran tests on Victoria brown coal in SASOL's SRC-II liquefaction
process. After these tests, Kominic modified their SRC-I process at Osaka to
an SRC-II type process using the group's new high pressure hydrotreating
stage.
It is believed that the Kominic process has been developed by the joint
partner company of Kobe Steel, Ltd. In this process (Figure 1) coal is
pulverized, preferably 200 to 300 mesh, and mixed with recycled pulverized
catalyst (cobalt-molybdenum and/or iron, iron-sulfur) of the same consis-
tency. The recycled catalyst is presumably recycled ash residue, depending on
the coal properties. However, the catalytic effects of this ash residue are
highly dependent on the coal type and coal deposit location. In addition to
pulverized coal and catalyst, a hydrocarbon solvent such as anthracene oil,
having a boiling point of over 150?C, is added and slurried in a slurry
tank. This slurry is then pumped, at a flow velocity of from 10 to 400 m/hr
to a preheater.. The slurry is mixed with a high pressure, hydrogen-rich gas,
and this mixture is then heated to a temperature of 420? to 440?C.
1
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The hot slurry mixture is then introduced into the base of the hydro-
genation reactor at a pressure of 100 to 150 atm. The stream leaving the top
of the reactor consists of dissolved gases in a light oil/solid residue
slurry. This mixture is then sent to a solid-liquid separator.
The solid-liquid separator consists of a liquid cyclone and a solid
accumulating tank. Connected to the top of the liquid cyclone is a gas-liquid
outlet pipe. The solid-liquid separator is operated in a batch mode and
therefore a second solid-liquid separator is utilized in an alternating
pattern. The solid residue discharge from the separator is collected in a
slurry tank where part of it is recycled. The liquid that is withdrawn from
the separator is then sent to a dehydrogenation cyclopolycondensation reactor
for further processing.
The dehydrogenation-cyclopolycondensation reactor is operated at ~20? to
440?C and 100 to 150 atmospheres. In this reactor the liquid is subject to
treatment under noncatalytic conditions in the presence of a small amount of
hydrogen gas with a total partial pressure of 7 to 70%. The reaction time
within the reactor is 5 to 90 minutes, as was the case in the hydrogenation
reactor. With this treatment the higher oil liquid entering the reactor is
converted from an oil having naphthenic or paraffinic-rich properties to a
heavier oil having aromatic-rich properties. This oil is then sent to a gas-
liquid separator for further processing. The aromatic rich oil produced in
the NBCLP process can then be sent to a refinery for further processing.
To achieve the atomic ratio of hydrogen to carbon of liquid fuels, direct
liquefaction processes rely on doubling or tripling the ratio from that of
coal. Unlike gasification processes, which produce hydrogen internally by the
water-gas shift reaction, direct liquefaction processes require the outiside
addition of hydrogen. About 1000 SCF of hydrogen is needed per 100 pounds of
coal (MAF) feedstock. Therefore, for a 25,000 ton/day facility, about
5 X 108 CF/day of hydrogen is required. To meet this demand, hydrogen must be
produced by steam reforming of methane or light oil feedstock, partial
oxidation of heavy crude, or coal gasification. About 26,000 bbl/day of
naptha would be required for the process to produce a nominal 60,000 bbl/day
of products. Because naptha is not likely to be available in adequate supply
in Japan, hydrogen production via coal gasification is more likely.
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A U.S. patent that as recently been awarded to Kobe Steel for this
process is shown as an Exhibit. A technology fact sheet for this process is
presented in Table 1.
To further promote to commercialization of this direct liquefaction
process, the Komonic group was reorganized with the addition of Idemitsu Kosan
and Asia Oil on August 12, 1980.6 The new company, Nippon Brown Coal
Liquefaction Co., is capitalized at $2.1 million with equal 20% investments
from each of the five participating companies. This new company has recently
entered into negotiations for a joint Japan Australia brown coal liquefaction
project with the Victoria state government and the Australian government.
Future plans call for the construction of a 50 metric ton/day pilot plant
for which construction was started in 1981 with operation expected in 1983.
This plant will be constructed near the coal mines in Victoria state at an
estimated cost of $150 to $175 million (U.S.). About 90% of this plant will
be funded by the Japanese government as part of its national Sunshine
program. In return for government support all technologies and know-how
gained through the operation of the pilot plant will be made available to the
government. The Australian government will supply the basic infrastructure,
plant site, electric power, industrial water, and feedstock coal.7
A second phase is currently envisioned that calls for the construction
startup of a 5,000 ton/day demonstration plant in 1985 at a cost of between
$833 million and about $1 billion (U.S.). This demonstration plant would then
be expanded into a commercial size facility by 1990 with the addition of five,
5,000 ton/day units, The Victoria state government has pledged to the
Japanese full coperation in supplying the feedstock brown coal for the 30,000
ton/day commercial plant. The total project cost for the first three phases
is estimated at over $4 billion (U.S.).6
Relationship to Prior Technology
This development is not directly related to any commercially developed
process. However, the general principals of the NBCL process are similar to
the SRC-II process developed in the U.S.
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Operating Facilities
In 1977 Kobe Steel constructed a 500 kg/day process development u
the Iwaya Works west of Osaka to test Victoria brown coals.
Major Funding Agencies
The NBCL liquefaction process has attained the status of a nation
project and as such is being partially funded by the Japanese government to a
commercial level.
The Nippon Brown Coal process is currently in the pilot plant sta e of
development. Due to the lack of available data and operating experience,
critical problems for this process have yet to surface. However, due to the
similarity of this process to other coal liquefaction processes, in particular
the SRC process, parallel problems might be expected to occur. In the SRC
process corrosion/erosion problems occur in four main process areas that are
similar to those found in the Nippon Brown Coal liquefaction process. These
areas include coal receiving and preparation, preheating and dissolvin
filtration and mineral residue drying, and solvent recovery.
In the area of coal receiving and preparation, typical problems
associated with similar facilities occurred in coal slurry centrifugal pumping
equipment and high-pressure plunger-type slurry preheater charge pumps These
problems centered on packings and seal leakage, as well as check valve and
plunger erosion. In the preheating and dissolving sections few materials
problems developed. These minor problems involve nozzle sleeve lining
corrosion and scale formation. However, in the area of high- and
intermediate-pressure separator equipment, severe stress corrosion cracking of
Type 304 stainless steel linings occurred.
attributed to polythionic acid and chloride stress corrosion cracking 'at weld
sites. Corrosion was not experienced in any of the other materials inithe
separators, including carbon steel.
Corrosion/erosion material problems also occurred in the filtration and
mineral residue drying areas. Most problems occurred in the heat exc1anger
material where severe erosion of metal tubes was encountered as well as tube
bulging due to carburization in the 3/8 inch thick 304 SS dryer shell.
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Severe corrosion/erosion problems have also occurred in the reboilers,
air coolers, fractionation equipment, piping, and tower walls of the solvent
recovery section of the SRC process. These problems have caused several major
shutdowns. The most severe corrosion occurred in moving parts and materials
that came in direct contact with the liquid solvent. Most of the problems
required a change in material specifications to eliminate the situation.
However, acceptable corrosion rates have yet to be achieved in areas such as
the stainless-steel trays used in the separator towers.
Capital Costs
The investment cost of the Nippon Brown Coal Liquefaction process
developed by Kobe Steel, Ltd., has not been published. However, based on a
process flow sheet comparison with other coal liquefaction processes, the
Nippon Brown Coal process does not differ significantly from the SRC-II
process which is being developed by the Gulf Oil Corp. in the U.S. Based on
this comparison, the two processes should have similar capital cost
requirements. The total plant investment costs for the SRC-II process are
shown in Table 2. These costs include engineering, land, plant, and G&A
costs. Also included is a 20% contingency factor based on the state of the
technology readiness. Make-up hydrogen production will cost about
$7.50/million Btu based on capital costs and coal requirements in addition to
residues. This amounts to a product cost content for liquids of about
$3.50/million Btu. Hydrogen at $10.00/million Btu would result in a product
cost content of about $4.50/million Btu.
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Table 2. TOTAL PLANT INVESTMENT, NIPPON BROWN COAL LIQUEFACTIOJ
1978 $ 1980 $
(Millions) (Millions)
Major Liquefaction
Equipment 675 810
Other Equipment and
245 295
Labor and Erection 530 635
Hydrogen Productirn
Plant (Coal Gasi`ic;ation) 825 1000
Analysis cost basis: $2.7 billion (1980 $)
Coal Requirement: 25,000 tons/day @ 12,500 Btu/lb
(including 2500 tons/day added to
7500 tons/day residues for
hydrogen make-up)
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United States Patent [)9)
Nakako et al.
[54] COAL LIQI ! F ~CT1ON PROCESS AND
APPARATUS THEREFOR
Inventors: Yukio Nakako, Nishinomiya; Shizuo
Yokota, Kobe. both of Japan
[73] Assignee: Kobe Steel. Ltd., Kobe, Japan
[21) A ppl.N o.: 915,575
[22] Filed: Jun. 14, 1978
Related U.S. Application Data
(63) Continuation-in-part of Ser. No. 801,920, May 31.
1977. abandoned.
(30] Foreign Application Priority Data
May 28. 1976 [JP] Japan .................................. 51-62811
May 28. 1976 [JP] Japan .................................. 51-62812
May 28, 1976 [JP) Japan .................................. 51-62813
May 28, 1976 [JP) Japan .................................. 51-62814
May 28, 1976 [JP] Japan .................................. 51-62815
(51] Int. C1.1 ................................................ C10G 1/06
[52] U.S. CI ...................................... 208/10; 208/8 R
[58] Field of Search .......................... 208/8, 10
(56] References Cited
U.S. PATENT DOCUMENTS
Re. 25,770 6/1961 Johanson ................................ 208/10
2.913.397 11/1959 Murray et al ............................ 208/8
3.018.242 1/1962 Gonn ....................................... 209/8
3.117,921 1/1964 Gorin ....................................... 208/8
3.594.303 7/1971 Kirk et a! ................................. 208/8
3.644,192 2/1972 Li et al ..................................... 208/8
3.932.266 1/1976 Sze et al ................................... 208/8
4,219,403
Aug. 26, 1980
FOREIGN PATENT DOCUMENTS
20957 of 1929 Australia
8303 '/1932 Australia
Primarv Examiner-C Davis
Attorney, 4gent. or Firm-Oblon. F sher. Spi%ak.
McClelland & Mater
(57] ABSTRACT
A coal liquefaction apparatus whit. comprises a siurrn
mixing tank. a preheater, a hydrogenation reactor. anc
a gas-liquid-solid separator or separators in series anc a
gas-liquid separator and at least one solid-liquic separa-
tor are interposed between the hydrogenation reactor
and a dehydrogenation cyclopolyc ndensation reactor
which is positioned upstream of t e finai gas-liquid-
solid separator.
The coal liquefaction process corn nses the steps of
heat treating a slurry prepared by mi ing coal fines with
a hydrocarbon based solvent having a boiling point
greater than 150' C. in the present of hydrogen at a
temperature of 300' to 500' C. and pressure of 50 to
700 atms. thereby forming a gas-liquid-solid mixture:
separating and removing solids from said gas-liquid-
solid mixture as a reaction product: separating and re-
moving a residuum liquid fraction rom said mixture:
and heat treating said residuum liq id fraction in the
presence of hydrogen at a low partial pressure at a
temperature of 300' to 500' C. and pressure of 50 to
700 atms.
12 Claims. 12 Drawing Figures
HIGH PRESSURE
HYDROGEN-RICH GAS
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U.S. Patent Aug. 26, 1980 Sheet I of 5 4,219,403
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U.S. Patent Aug. 26, 1980
Sheet 2 of 5 4,2 ~ 9,403 1
I
1
FIG.2
i.
HIGH PRESSURE
HYDROGEN-RICH
COAL F1NEI f GAS
FIG.3
24
2I-.
I
HIGH PRESSURE
HYDROGEN RICH GAS
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U.S. Patent Aug. 26, 1980 Sheet 3 of 5 4,219,403
HIGH PRESSURE
REDUCTIVE GAS
COAL FWE
A
F_b~
KX I GS 106"
~05
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U. S. Patent Aug. 26, 1980 Sheet 4 of 5 4,~ 19,403 1
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U.S. Patent Aug. 26, 1980 Sheet 5 of 5 4,219,403
HIGH PRESSURE 313
HYDROGEN-RICH GAS
1
I
COAL FINE
0
301
303 312
FIG. 10
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COAL LIQUEFACTION' PROCESS AND
APPARATUS THEREFOR
This application is a continuation-in-pan application
of co-pending application Set. No. 801,920. filed May
31, 1977. and now abandoned.
BACKGROUND OF THE INVENTION
I Field of the Invention
The present invention relates to a coal liquefaction
process and an apparatus therefor. and more particu-
larl% to a coal liquefaction process which can be per-
formed efficiently to improve the yield of reaction
products. particularly, the heavy oil product which is
well suited as a metallurgical carbonaceous carbon ma-
terial.
2. Description of the Prior Art
A coal liquefaction process is known in which coal
fines ara treated in the presence of hydrogen to liquify
the coal. The coal fines used in the coal liquefaction
process include low grade coals such as bituminous,
semi-bituminous, and sub-bituminous coals and lignite
as well as similar solid carbonaceous materials such as
shale. According to the conventional process of the
type described. coal fines, a hydrocarbon solvent hav-
ing a boiling point of over 150' C., and a suitable cata-
lyst such as a ferro-sulfuric system catalyst, if desired.
are mixed in a slurry, (The use of a catalyst may not be
necessary or essential because the ash in coal functions
as a catalyst), and then the slurry is preheated in a pre-
heater. A high pressure hydrogen-rich gas is added
thereto preferably prior to the preheating of the slurry.
The preheated slurry and a high pressure hydrogen-rich
gas are passed into a reactor where a hydrogenation
reaction is conducted at a high temperature and pres-
sure, e.g., 300' to 500' C., 50 to 700 atms. Then. a mix-
ture of reaction products of reactor effluent is intro-
duced into two or more separators connected through
pressure-reducing valves to each other, wherein the
pressure is progressively reduced and gas. liquid and
solid are flash distilled.
At the present time, the objective in the liquefaction
of coal is to form a heavy oil product having a high
boiling point for use as a metallurgical carbonaceous
material for use, for instance, in the manufacture of
steel-making coke or carbon electrodes for alumina
electrolysis. The liquid product or effluent. generally,
includes solids such as ash, unreacted coal, catalysts.
and insoluble reaction products. Accordingly, the re-
moval of these elements would improve the quality of
the heavy oil product for its intended use. In general. a
metallurgical carbonaceous material should have an ash
content of less than 10%.
Coal liquefaction process hitherto has been beset with
many formidable problems. which will be described as
follows: -
Problem I
Because of excessive hydrogenation. the yield of a
heavy oil fraction in the liquid reaction product is not
high enough. Moreover. solids condense along with a
heavy oil fraction. in the final stage separator. where
solids and heavy oil are to be separated. However. in
the conventional method. a mixture having a high sis-
cosity results at this stage. so that much time and effort
must be devoted to filtering in the ,cpantion stage to
sroarate solids from the ,il. Far !hi' reaso n..i ueb:ml s
2
added to lower the viscosity of the mixtu e. and if re-
quired, the mixture is heated. followed b centrifugal
separation, sedimentation separation, or s paration by
means of separators such as liquid cycl nes. In any
5 event, a light oil in the case should be added to the oil
in a considerable amount. and this resul s in an un-
wanted increase in the amount of the mixture to be
treated, which causes an increase of the n mber of ap-
paratus for separating the solid and liquid and deterio-
10 rates an economic effect In addition, upon flash distilla-
tion, a solid fraction and a heavy oil fracti n both pass
through pressure reauc:ng vaives. so that i the pressure
is instantaneeusiy reduced to a considerably lower
level, then wear of the pressure reducing v lues occurs.
15 To avoid this, many separators and pressure reducing
valves have to be used in order to gradually reduce the
pressure of the system. The use of many such separators
and reducing valves increases the expense of capital
equipment.
20
Problem 2
In the coal hydrogenation reactor, a m xture of hy-
drogen gas or a high pressure reductive gas such as
CO+H20, CO-H2O-H2, CO-- or H rich gas and
25 the coal slurry which has to be preheated is subjected to
a liquefaction reaction at a high temperature and pres-
sure, followed by flash distillation to separate the prod-
uct obtained into gas. liquid, and solid p oducts. It is
advantageous to introduce the slurry and t e high pres-
30 sure reductive gas into the reactor from it bottom and
expel the products from the top of the re ctor. In this
case, the viscosity of the solvent is decreas d because of
the reaction at high pressure and temperat re, so that a
tendency arises for the settling of solids s ch as unre-
35 acted coal fines, catalysts and ash from t e liquid. To
avoid this problem. the upward rate of flow of the mix-
ture is increased relative to the settling r to of solids
during reaction. However, in order to achieve this ob-
jective, it is necessary to reduce the cross ctional area
40 of the reactor to some extent, and the nur ber of reac-
tors connected in series should be increas to achieve
sufficiently long residence times of the mix ure for reac-
tion in the reactors. This is uneconomical t cause many
pieces of apparatus such as gas-liquid sep rators. pipe.
45 and couplings must be used. Moreover, ore mainte-
nance problems arise because of the more xtensive use
of equipment. One of the attempts to solve his problem
has been to reduce the number of reacto s while the
liquid effluent from one reactor is recycle to another.
50 thereby extending the residence time o the slurry
within the reactors. Alternatively, a reductive gas in
great amounts is injected into the reactor to retard the
settling of solids in the liquid reactant. Ho ever. in this
technique, the concentration of unreacte coal in the
55 reactor is equalized both at the entrance a d exit of the
reactor. so that the reactor itself changes i type from a
piston flow reactor to a complete mixing reactor. with
the result that the reaction efficiency decre es substan-
tially relative to the reaction space or v lume of the
00 reactor.
Problem
A high boiling point and high viscosity reaction prod-
uct is obtained from the bottom of the sep rator in the
0! final stage of the multipie stage flash distillation. Ac-
cordingt.. the .tezree of condensation of lids is not
,ufficientt% high. thereb% requirine further eparation
..situ. 'ram :he iiQut,, Hov e%er. because of the hizh
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viscosity of the reaction product, satisfactory separation
of solids cannot be attained by a filtering process. For
this reason, as has been described earlier, a light oil is
added to the liquid product to decrease the viscosity of
the mixture or heat is applied thereto, followed by cen-
trifugal separation, sedimentation separation or separa-
tion in a liquid cyclone. Accordingly, the amount of the
mixture to be treated is increased, thus failing to meet
practicability requirements. It is therefore evident that
no satisfactory separation process for solids has yet been
found.
SUMMARY OF THE INVENTION
Accordingly, one object of the present invention is to
provide a coal liquefaction process and an apparatus 15
therefor, which improves the yield of liquified product
suitable for use as a metallurgical carbonaceous mate-
rial, while avoiding the wear of pressure reducing
valves, and dispensing with multiple stage separators
and pressure reducing valves. 20
Another object of the present invention is to provide
a coal liquefaction process and an apparatus therefor,
which provides an improved reaction efficiency rela-
tive to the space within the reactor, without using many
reactors and couplings. 25
Still a further object of the present invention is to
provide a coal liquefaction process and an apparatus
therefor which improves the separating efficiency of
solids from the liquid in separators after the hydrogena-
tion reaction. 30
Yet a further object of the present invention is to
provide a coal liquefaction process and an apparatus
which eliminates the public nuisance problem caused by
the disposal of catalysts.
Still a further object of the present invention is to 35
provide a coal liquefaction process and an apparatus. in
which solids may be efficiently separated from a high
boiling point, high viscosity reaction product obtained
from the bottom of a final stage separator, in a reason-
able manner. 40
According to the first aspect of the present invention,
solids are separated from a reaction mixture of low
viscosity and at high temperature immediately after the
hydrogenation reaction, and the reaction mixture from
which the solids have been removed is then subjected to 45
a dehydrogenation-cyclopolycondensation reaction
under a low partial pressure of hydrogen at a high tem-
perature under non-catalytic conditions. The dehy-
drogenation-cyclopolycondensation reaction is a reac-
tion in which a light oil is dehydrogenated under non- 50
catalytic conditions at a low hydrogen partial pressure,
thereby being converted into a heavy oil, while the
fraction of the reaction product which has been given a
naphthenic or paraffinic-rich property because of the
addition of an excessive amount of hydrogen, is dehy- 55
drogenated and cyclicpolycondensed. More particu-
larly. the reaction mixture from the hydrogenation reac-
tor is introduced as it is or after passing through a gas-
liquid separator, into a solid-liquid separation system
consisting of solid-liquid separators having pressure 60
reducing valves, with the lower portions of the separa-
tors being connected to solid accumulating tanks, and
with the top portions thereof connected to gas-lined
outlet pipes. The liquid fraction separated therein :s
subjected to a non-catalytic heat treatment in the pres- e5
ence of hydrogen at a low partial pressure. Suitable
solid-liquid separators employable in the present in-, en-
tion are :,.:.ones. sand cones. and the like
4
The non-catalvn: heat treatment .s ,u:ii trur :he
reaction product is maintained at a given temperature
for a given period of time in the presence of hydrogen
at a low partial pressure. Any type apparatus may he
used, as long as the above described conditions an he
maintained. For instance. a device having the same
construction as that of the reactor, or heating Besse:
which is used for preheating may be used as the non-
catalytic heat treatment vessel. More specifically. the
duced as it is, or after passing it through gas-;iyuid ,er.,i?
rators into solid-liquid separators at a temperature euu.i.
to or less than the temperature at the ex;r :,f a reactor
but, in any case. a temperature 100? C no less than the
latter. In the solid-liquid separators. solids accumulate
in the lower solid-accumulating tank. while liquid and
gas. if any, overflow and are withdrawn through oser-
head gas-liquid outlet pipes. The liquid fraction thus
withdrawn is mixed with a hydrogen-rich gas. as re-
quired, and then introduced into a dehydrogenation-
cyclopolycondensation reactor. Meanwhile, the reac-
tion product from the hydrogenation reactor contains
an excessive amount of a high pressure hydrogen-rich
gas, so that hydrogen need not be added in this stage.
However, when the reaction product passes through a
gas-liquid separator. the addition of hydrogen is re-
quired, or a small amount of high pressure hydrogen-
rich gas should preferably be introduced into the dehy-
drogenation reactor. In the dehydrogenation reactor, a
reaction mixture devoid of solids is maintained at a high
temperature in the presence of a small amount of hydro-
gen or at a low partial pressure under non-catalytic
conditions so that the portion of the product which
possesses naphthenic or paraffinic properties. is dehy-
drogenated and cyclopolycondensed, thereby being
converted into a heavy oil fraction which imparts an
aromatic-rich property to the, oil which in turn yields a
heavy oil well suited as a metallurgical carbonaceous
material. In this respect. the presence of a small amount
of hydrogen or a low partial pressure of hydrogen is
mandatory for preventing an excessive amount of dehy-
drogenation-cyclopolycondensation. The reaction mix-
ture subjected to the dehydrogenation reaction is with-
drawn from the top of the dehydrogenation-
cyclopolycondensation reactor, then passed through
separators and then flash-distilled by reducing the pres-
sure through pressure-reducing valves. However, be-
cause the reaction mixture is devoid of solids in this
stage, the pressure-reducing valves are not damaged
and there is no longer the need to separate solids from
the liquid in the separator.
Meanwhile, in the solid-liquid separating system.
when one solid accumulating tank becomes filled with
solids, then the solid-liquid separating system therefor
shut off from the reaction-mixture-inlet passage. s here-
upon the pressure in the separator is reduced to atmo-
spheric pressure by means of a pressure-reducing s aive.
and then, the accumulated solids are discharged
through a bottom outlet port, as required. The solids
thus discharged contain materials having a catalytic
function, and thus may be used again in the coal slut-r
At least two solid-liquid separating devices in naralle!
are provided for one reaction system so that two-sohd-
liquid separating devices may be used aiternate!,. i.e
according to the so-called batch system operation
More particulariy. the reaction mixture from the hvcr:-
genauon reactor ,s first introduced .ender high ^ressur
:nto one ?oiid-'ieuid separating JCs ice. and w n.e- '`.e
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device is tilled with solids, then the connection is
switched from the filled device to the other solid-liquid
separating device in order to introduce the reaction
mixture into the latter, while the pressure in the first
solid-liquid separating device is reduced to atmospheric
pressure in order to discharge solids therefrom. This
cycle of operation is repeated for an efficient continuous
separation of solids from liquid.
In the second aspect of the present invention, the
diameter of the reactor is increased and the number of
reactors is reduced, while retaining the desired level of
efficiency required for liquefaction or the hydrogena-
tion reaction. In other words, the upward flow velocity
of the reaction mixture in the reactor is adjusted in
order to accelerate the settling of solids therein, and
solids thus settled are discharged from the bottom of the
reactor, while a fresh catalyst is supplied, as required,
thereby maintaining the desired hydrogenation reac-
tion.
More specifically, in the present invention, at least
two reactors each having a solid outlet port in the bot-
tom of the reactors are connected in series, and a pre-
heated mixture of a coal slurry consisting of coal fines,
catalyst and a high pressure reductive gas is introduced
into the first reactor through its bottom port so that it
passes through the reactor at such a flow velocity that
solids may settle in the reactor. In this case, the reaction
mixture is separated into a relatively solid-rich layer and
a relatively solid-lean layer. The solids which settle are
discharged from the solid outlet port provided in the
bottom portion of the reactor. In this respect, one or
two solid accumulators are connected to the bottom of
the reactor, so that solids may be stored therein in a
sufficient amount, followed by flash distillation, and
then the withdrawal of the solids. At the same time the
solids present in the reaction mixture cannot be com-
pletely separated in the first reactor, and hence, over-
flow of solids occurs along with the reaction liquid,
which are separated in the succeeding reactor in the
same manner.
In the second embodiment of the present invention.
the catalyst substantially separates from the liquid and is
removed in the first reactor, so that fresh catalyst
should be supplied to the second reactor and thereafter
through pipes leading to the catalyst accumulating tank
to promote the hydrogenation reaction. Accordingly,
the reaction is conducted in an efficient manner because
of the supply of fresh catalyst. In addition, different
kinds of catalysts may be used in reactors. For instance.
a catalyst of the cobalt-molybdenum system, iron or
iron-sulfur which possesses a high activity in the lique-
faction reaction, is used in the first reactor for a highly
efficient reaction, while a catalyst of a low activity is
used for the second reaction and also thereafter when
the reaction medium contains a relatively small amount
of unreacted coal. Furthermore, no catalyst is supplied
to the final reactor, so that a product possessing a naph-
thenic or paraffinic property, because of excessive hy-
drogenation is heated in the presence of a low partial
pressure of hydrogen under non-catalytic conditions for
the dehydrogenation-cyclopolycondensation reaction.
thereby converting the liquid product into a heavy oil
product having aromatic characteristics, which is well
adapted for use as a metallurgical carbonaceous mate-
rial.
The flow velocity of the reaction mixture of the pres-
ent invention depends on the kinds ana grain sizes of
coal fines and :atalvsts used In short, the flow velocity
6
should be selected such that the solids in t
mixture may settle. thus leaving a solid-rich
solid-lean layer therein. For instance, w h
oxide catalyst is used, and the grain sizes of
5 and the coal fines are 200 mesh, then the 1
velocity of the slurry stream should be about
to prevent settling of the solids. i.e.. 360 m/
the flow velocity of the reaction mixture
fluidize the same is about 1.5 m/hour In a
10 type of reactor, the flow velocity should
about 1.2 m/hour to 360 m/hour. If the flow
excessively low, then the liquefaction reacts
proceed satisfactorily, but instead. coking oc
the flow velocity should preferably be over
15 On the other hand. if the flow velocity is g
3600 m/hour, then the undesirable excessiv
e reaction
ayer and a
!n an iron
he catalyst
)west flow
10 cm/sec
our. while
n order to
ebullated
ange from
velocit, , is
n does not
urs. Thus.
0 m/hour.
,eater than
overflow
of solids takes place. The grain sizes of the coal fines
and the catalyst particles should range fro 50 to 400
mesh, preferably from 200 to 300 mesh. Fo the grain
20 sizes in this range, the flow velocity of the lurry may
range from I to 3600 m/hour, preferably from 10 to 400
m/hour.
In the third aspect of the present invention, the reac-
tion mixture is separated into a solid-rich layer and a
25 solid-lean layer, with an interface betwee the two
layers being maintained at a given equilibria level. In
the solid-rich layer of a given volume, ash and unre-
acted coal fines are present which promote he hydro-
genation reaction. On the other hand, in the solid-lean
30 layer, the dehydrogenation-cyclopolvconde cation re-
action occurs which results in the yield of heavy oil
product having an improved aromatic property, which
is preferable from the viewpoint of a desirable metallur-
gical carbonaceous material. In addition, the formation
35 of two layers permits the separation of solids of a lower
ash content in an increased amount. Furthermore, the
solid-rich layer thus separated may be wit drawn, as
required. so that solids may be added to the slurry for
reuse as a catalyst, thus saving the amount of catalyst to
40 be used. More particularly, in the present invention. in
the hydrogenation reaction of coal fines, a t be having
an opening tip is inserted into the hydrogenation reac-
tor, while the other end thereof is connecte to an ash
accumulator which is maintained substanti ly at the
45 same pressure level as that of the hydrogenation reac-
tor. Then, the pressure in the accumulator is djusted so
that a solid-rich layer may be introduced into the accu-
mulator in order to maintain the interface between the
two layers at a given equilibrium level. such that the
50 volume ratio of the solid-lean layer to the solid-rich
layer falls between 1/6 to 2.
In still another feature of the present iven ion. a tube
having an open tip is inserted into the reactor through
the base of the reactor, while the other end of the tube
55 is connected to ash accumulators, which have a solid
withdrawing means at the base of the react r. The ash
accumulators have gas pressure, flow rate control
means and gas injection means in the tops o the accu-
mulators. As a mixture of slurry and high p essure hy-
60 drogen-rich gas is introduced into the reactor, only the
solid-lean laver is withdrawn from the top of the reac-
tor, so that the interface between the twol layers as-
cends. When the interface between the tw o layers
passes over the open tip of the tube to a desired height
05 therefrom. which depends on the reaction conditions.
the size of the reactor and the like. the soud-rich layer
is introduced into an ash accumulator in an amount
proportionai to the amount of the react:. n mixture
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being fed therein. Upon the introduction of the solid-
rich layer into the ash accumulator, a high pressure
hydrogen-rich gas or hydrogen is charged into the ash
accumulator substantially at the existing pressure level
in the reactor, and then the pressure in the accumulator 5
is adjusted to a level somewhat lower than the pressure
in the reactor so as to allow the introduction of a solid-
rich layer into an ash accumulator. i.e., by continuously
bleeding a gas at a given rate therefrom. As a result, the
interface between the solid-rich layer and the solid-lean 10
layer may be maintained at a given equilibrium level.
The solid-rich layer introduced into the ash accumula-
tor is flash-distilled and added to the slurry for reuse. In
the ash accumulator system, two ash accumulators may
be used in an alternate embodiment. 15
According to the fourth aspect of the present inven-
tion, the interface between the solid-rich layer and the
solid-lean laver is maintained in close vicinity to the
open tip of a tube which is inserted in the reactor by
withdrawing the solid-rich laver through the open tip of 20
a tube. thereby providing an equilibrium between the
solid-rich layer and the solid-lean layer.
The tube, as used herein, may be fixedly or movably
inserted into the reactor, with the end thereof being
connected via a pressure reducing valve to a slurry tank 25
or a solid-liquid separator, such as a liquid cyclone. In
this case, as well, the volume ratio of the solid-lean layer
to the solid-rich layer should preferably range from 1/6
to 2.
If ash, catalyst and unreacted coal fines are separated 30
from the solid-rich layer, then the hydrogenation reac-
tion efficiency decreases, and unreacted coal undergoes
a coking reaction, thereby adversely affecting the yield
of an intended product.
Upon adjustment of the level of the interface between 35
the solid-rich layer and the solid-lean layer to a vicinity
close to the open tip of the tube in the reactor, when a
mixture of slurry and a high pressure hydrogen-rich gas
is continuously introduced into the reactor, the solid-
lean layer alone is withdrawn from the top of the reac- 40
tor. so that the interface between the two lavers ascends
to the open tip of the tube. In this stage, the solid-rich
laver is withdrawn through the tube in order to main-
tain the interface between the two layers at an equilib-
rium level which is close to the open tip of the tube. The 45
solid-rich layer thus withdrawn is flash-distilled as it is.
and then added to the slurry for reuse as a catalyst. or
otherwise separated into liquid and solids, while the
liquid fraction is added to the solid-lean laver again, and
the solid fraction is recovered so that it can be added to 50
the slurry for reuse. In this case, the solid-rich layer thus
withdrawn is of low viscosity, thus facilitating the sepa-
ration into liquid and solid phases.
According to the fifth aspect of the present invention.
the reaction mixture from the hydrogenation reactor is 55
introduced as it is, or via a gas-liquid separator. into a
solid-liquid separator having a solid accumulator con-
nected to the bottom thereof. In this respect. the reac-
tion mixture contains a solvent or a light oil and is of a
low viscosity because the reaction mixture is preheated. 60
thus providing ease of separation. In addition, a pres-
sure-reducing valve is provided on the gas-liquid with-
drawing pipe connected to the top of the solid-liquid
separator. so that upon pressure reduction for flash
distillation. solids will not pass through the pressure- t,5
reducing valve. thus avoiding errosion of the valve.
This permits pressure reduction at a rapid rate. In this
respect. part of the gas withdrawn from the solid-liquid
8
separator may be cooled for liquefaction for further
distillation in a distilling column. When the solid-liquid
separator is filled with solids. then a pressure-reducing
valve on a gas-liquid withdrawing pipe is opened in
order to reduce the pressure to atmospheric pressure
instantaneously, for flash distillation. The cycle of oper.
ation can be repeated for efficient solid-liquid separa-
tion.
BRIEF DESCRIPTION OF THE DRAWINGS
FIG. I is a flow sheet illustrative of a prior art lique-
faction process for coal fines.
FIG. 2 is a diagrammatic view of a solid-liquid sepa-
rating device within the scope of the present invention
FIG 3 is a flow sheet illustrating a liquefaction pro-
cess according to the present invention, which employs
two solid-liquid separating devices:
FIG. 4 is a flow sheet illustrative of one embodiment
of the liquefaction process according to the present
invention;
FIG. S is a view illustrative of one embodiment of a
reactor of the present invention;
FIG. 6 is a view illustrative of another embodiment of
the reactor of the present invention:
FIG. 7 is still another embodiment of a reactor of the
present invention:
FIG. 8 is a flow sheet of the hydrogenation process of
the present invention which employs the reactor of
FIG. 7;
FIG. 9 is a yet another embodiment of the reactor of
the present invention;
FIG. 10 is a flow sheet illustrative of one embodiment
of the liquefaction process of the present invention
which employs the reactor of FIG. 9;
FIG. 11 is a diagrammatic view of another enibodi-
ment of the solid-liquid separating device of the present
invention: and
FIG. 12 is a flow sheet illustrative of the liquefaction
process of the present invention which employs two
solid-liquid separating devices.
DESCRIPTION OF THE PREFERRED
EMBODIMENTS
FIG. 1 illustrates a prior art liquefaction process.
Coal fines and a solvent such as a hydrocarbon having
a boiling point of over about 150' C.. and a catalyst. if
required. are slurried in a slurry tank; and then the
slurry thus prepared is delivered by a slurry pump 2 to
a preheater 3. Before the slurry is passed into the pre-
heater it is mixed with a high pressure hydrogen-rich
gas. The mixture of slurry and a hydrogen-rich gas.
which have been preheated to about 300' to 500? C.. is
introduced under pressure into a hydrogenation reactor
4 through its base for reaction at a temperature of about
300? to 500' C., and a pressure of about 50 to "00 arms.
The reaction mixture from the reactor 4 is passed
through separators 5.6 and 7 which are connected in
series in the indicated order, and then pressure-reducing
valves 8. 9. provided on pipes which connect the sepa-
rators in series, are opened so as to reduce the pressure
gradually for flash distillation of the slurry into solids
and liquid The gas effluent withdrawn from the top of
the first separator 5 is cooled for liquefaction. as desires.
while a light oil fraction is distilled in a distilling l-
umn A mixture of light and medium oils. and s, ivent
withdrawn from the top% of separators 6 and 7 is .ii-
tilled in a distilling column. and then the solvent tt-,u>
recovered is recycled for use as a Murry term rc crl-
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'est. Meanwhile, the heavy oil fraction withdrawn
from the bottom of separator 7 contains a considerable
amount of solids, which generally should be separated
from the heavy oil. This is referred to as a de-ash opera-
tion.
In the first embodiment of the liquefaction process of
the invention, as shown in FIGS. 2 and 3, a solid-liquid
separating device 10 is positioned downstream of the
reactor 4. so that the reaction mixture from the reactor
4 may be separated efficiently
The solid-liquid separating device 10 consisting es-
sentially of a liquid cyclone 11 which is a type of solid-
liquid separator, and a solid accumulating tank 12 con-
nected to the bottom of the cyclone 11. Connected to
10
terms of liquid cycline 11). while the liquid is with-
drawn through outlet pipe 13 by overflow into reactor
21.
The solids thus separated accumulate in solid-
accumulating tank 12. When solid-accts ulating tank 12
is filled with solids, then stop valves 6' and 14 are
opened. The stop valves 16 and 14 are closed. so that
the introduction of the solid-liquid mtxt re is switched
from the first separating device 10 to the second sepa-
rating device 10'. for the separation of lids and liquid
as well as for the accumulation of solid On the otner
hand, the pressure reducing valve 17 oft the first separat-
ing device 10 is opened to reduce the pressure inside to
atmospheric pressure, and stop valve I is opened so
that the solids which accumulate therein are withdrawn
through outlet pipe 19. The solids thus thdraAn are
delivered to slurry tank 1 for reuse en. all stop
valves and pressure reducing valves in separating de-
vice 10 are closed. When the second separating device
10' is filled to capacity with solids, then the introduction
of the solid-liquid mixture is switched fr m the second
separating device 10' to the first separa ing device 10
This cycle of operation can be repeated for a continuous
operation.
The liquid to be delivered to reactor 2 is introduced
into reactor 21 which is maintained subs antiaily at the
same temperature and pressure as that of reactor 4.
wherein the liquid is subjected to treatment under non-
catalytic conditions in the presence of a small amount of
hydrogen which is fed into reactor 21 through gas inlet
pipe 23. The treatment conditions deper. on the size of
the apparatus. the quality of the desired liquefaction
product, and the like. In order to prods e a heavy oil
product which is well adapted for use as metallurgical
carbonaceous material, preferably th temperature
ranges from 400' to 500' C., a total pressure of from 70
to 150 atms in the presence of hydrogen f a low partial
pressure, the hydrogen partial pressur is preferably
from "r to 701% of the total pressure. and t e reason nme
should be as long as that of the hydrogenation reaction.
for instance. 5 to 90 minutes. With this treatment. a
further lighter oil fraction or a reaction product having
naphthenic or paraffinic-rich properties. whicn is pro-
duced by the addition of an excessive am unt of hydro-
gen. may be subjected to a dehydr geration cy-
clopolycondensation reaction and con cried into a
heavy oil fraction. which has the desired aromatic-rich
property at an increased yield of I to 30 'c in compari-
son to the amount of starting coals (tit. F or med,u:n
abrasion furnace black). The liquid thus treated 's with-
drawn through outlet pipe 24 which is connected to the
top of reactor 21 and is delivered to th set arator for
further processing. which is well known
It :s apparent from the above discuss., n ccncerning
the liquefaction process of the in%enti n. a reac;:cr.
mixture devoid of solids is heat-treated i . the presence
of hydrogen. which results in an improved vie:1 of a
heavy oil fraction. while solids may be separated under
low viscosity conditions at high temperature inc pres-
sure. thereby providing improved ~,eoarai:nz efficiency
and minimizing the ash content of the i utd procuct
The liquefaction process achieved by tydrog_naticn
the top of the liquid cyclone 11 is a gas-liquid outlet pipe 15
13. while a stop valve 14 is provided on pipe 13. A
reaction mixture inlet pipe 15 is connected to the upper
portion of liquid cyclone 11 at a position lower than the
joint of the gas-liquid outlet pipe 13. while a stop valve
16 is also provided on pipe 15. In addition, a pressure 20
reducing valve 17 is connected to the upper portion of
solid accumulating tank 12, while a solid outlet pipe 19,
having a stop valve 18, is connected td 'a bottom portion
of tank 12.
In the liquefaction process of the invention, a non- 25
catalytic heat treating device is positioned downstream
of the solid-liquid separating device to reform the lique-
faction products, thereby improving the yield of the
heavy oil fraction which is suitable for use as a metallur-
gical carbonaceous material.
In the first embodiment of the invention, as shown in
FIG. 3, two or more solid-liquid separating devices 10
and 10' are provided directly or through a gas-liquid
separator 20 downstream of the reactor 4. In FIG. 3 the
primed reference numerals are used to distinguish the 35
second solid-liquid separating device and parts associ-
ated therewith for common use with those of the first
device from the first device.
Gas-liquid outlet pipes 13 and 13', which are attached
to solid-liquid separating devices 10 and 10'. are con- 40
nected to a gas-liquid inlet pipe 22, which is connected
to the bottom portion of the non-catalytic heat treating
device, or reactor 21. A high pressure, hydrogen rich.
gas injection pipe 23 is connected to reactor 21. while
an effluent outlet pipe 24 is attached to the top of reac- 45
tor 21. which leads in turn to separator S.
In the operation of the apparatus for the liquefaction
process of the present invention, as shown in FIGS. 2
and 3, the reaction mixture from the reactor 4 is passed
through the gas-liquid separator 20 at a temperature of 50
about 300' to 500' C. and a pressure of about 50 to '00
arms. and then gas is withdrawn from the top of separa-
tor 20. while a solid mixture is withdrawn from the
bottom thereof. which is then introduced into the first
solid-liquid separating device 10. The solid-liquid mix- 55
ture is subjected to a somewhat lower temperature and
pressure than the reaction mixture prior to its introduc-
tion into gs-liquid separator 20. All stop valves and
pressure reducing valves in the solid-liquid separating
devices 10 and 10', are maintained in their zlosed posi-
tions at first. and then stop valve 16 on inlet pipe 15,
which leads to the inlet of separating device 10. and
stop vah e 14 on inlet pipe 15. which leads to separating
device 10 are opened to allow the introduction of the
in the present process includes:
i : i a high degree of hydrogenation of ; a: ants :n 'he
presence Jt htdroeen anti :atatst J 7 ac::.
such as a .ai ilvst of the cohait-mri. DC?^:T
tem. :ron ; r iron,uitur s.sieni at hi
m^eratur:
effluent from reactor 4 into device 10. The effluent is o5
separated :nto a liquid-rich phase this will be referred
to simriy as a liquid), and a solid-rich phase The ;oiid-
iieuia in:,.cure will be referred ::-) 4b a sot:d 'A hen .(std n
t
1
1
1
I
11
t
1
r
 I
1
t
t
f
t
t
Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
(2) a relatively low degree of hydrogenation in the
presence of an iron system catalyst or in the ab-
sence of a catalyst in the presence of hydrogen; and
(3) liquefaction at a high temperature and high pres-
sure in a hydrogen donor solvent having aromatic
characteristics such as anthracene oil, without or in
the presence of a small amount of hydrogen.
The term "hydrogenation reaction" is used herein in
association with the above-described processes in-
cluded in the present invention.
FIG. 4 illustrates the second embodiment of the lique-
faction process of the present invention. Coal fines,
solvent and catalyst are slurried in slurry tank 101 and
then the slurry thus prepared is delivered by slurry
pump 102 to preheater 103. Prior to passage of the
slurry into the preheater a highly reductive gas is mixed
with the slurry. The mixture of slurry and high pressure
reductive gas, which has been preheated to about 300'
to 500' C., is fed under pressure into the first reactor 104
through its base, wherein the mixture is passed from the
bottom to the top at a flow rate (preferably 10 to 400
m/hour) such that the solids in the reaction mixture
may settle against the upward flow of the mixture for
reaction at a temperature of about 300' to 500' C. and a
pressure of about 50 to 700 arms. The reaction mixture
effluent, which overflows the top of reactor 104 is intro-
duced into the second reactor 104' through its base, and
then the reaction mixture effluent, which overflows the
top of reactor 104', is introduced into the third reactor
104" through its base. At this time, fresh catalyst from
the catalyst accumulating tank 105 is slurried in a suit-
able solvent and then the slurry is delivered by means of
pumps 106. 106' and 106" to reactors 104, 104' and 104",
respectively. The solids which settle in the various
reactors are discharged through solid outlet portions
107, 107', and 107", positioned at the bases of the reac-
tors. The reaction mixture effluent from the final reac-
tor 104" is introduced into a gas-liquid separator 108,
and then part of the gas effluent from the top of the
gas-liquid separator 108 is cooled for liquefaction, while
the liquid residuum is further distilled in a distilling
column. The liquid effluent from the base of gas-liquid
separator 108 (in this case, the liquid may contain some
amount of solids) is subjected to flash distillation under
a reduced pressure into gas, liquid, and solids, followed
by further distillation. The solids obtained from the
distillation contain unreacted coal fines, catalyst and the
like, and may be used repeatedly. Fresh catalyst can
also be combined with the recovered catalyst for reuse.
In the reactor of the invention, the reaction mixture
tends to separate into a solid-rich lower layer and a
solid-lean upper layer. Accordingly, it is preferable that
the flow velocity of the reaction mixture be adjusted by
an appropriate means such as a tube which may be
inserted into the reactor to withdraw the solid-rich
laver, and that one or more solid accumulators having
the same pressure as that of the reactor can be con-
nected to the bottom of the reactor. Thus, gas is bled
through the gas outlet pipes which are connected to the
solid accumulators by opening the gas pressure and
flow rate control valves provided on the gas outlet
pipes. at a discharge rate which is commensurate with
the rate a solid-rich liquid is introduced into the solid
accumulators under pressure. so that an interface be-
tween the solid-rich laver and the solid-lean laver may
be maintained at a given level. 41n general. the volume
ratio of the solid-lean layer to the solid rich laver should
preferably be adjusted to 1/ o to 2 i In addition, the
12
solid-rich liquid and the solid-iean iuuid arc ~itndr3.
through the open tip of the tube inserted into the reac-
tor, so that the interface between the solid-rich laser
and the solid-lean laver may be maintained in the vie in-
5 ity close to the open tip of the tube, thereby maintaining
an equilibrium level within the reactor. The separation
of the solid-rich layer and the solid-lean layer permits
the desired dehydrogenation-cvclopolycondensatior.
reaction to progress in the solid-lean layer, as has been
10 described earlier. thereby increasing the yield of a
heavy oil fraction having aromatic property character -
istics
A description in greater detail of the reactor!. A i11 ;e
presented.
IS Referring to FIG. 5. reactor 110 is shown whose base
is provided with an inlet port 113, which is adapted to
introduce a mixture of a slurry and high pressure reduc-
tive gas therein, and whose top portion is provided with
an outlet port 114. which is adapted to withdraw the
20 solid-lean layer therethrough. Reactor 110 is connecter:
via pipe 111 and valve 115 to a solid accumulator 112.
The solid accumulator 112 has its top portion connected
to a gas injection pipe 117 which is provided with a gas
injection valve 116. and a gas outlet pipe 119. which is
25 provided with a gas pressure flow rate control valve
118. A solid outlet pipe 121 having a stop valve 120
thereon for withdrawing solids therethrough is con-
nected to the base of accumulator 112. In the reactor
shown in FIG. 5. pipe 111 branches into two pipes
30 which are connected to two solid accumulators 112 and
112', which are arranged in parallel with each other. In
this respect, like parts in the second solid accumulator
are designated with like primed reference numerals
which are in common use with the corresponding parts
35 of the first solid accumulator.
In operation of the reactor shown in FIG. 5. the solid
accumulators 112 and 112' are isolated from communi-
cation with the reactor by closing the valves 115 and
115', and the gas pressure flow rate control valves 118
40 and 118' as well as stop valves 120 and 120' are close.:
for the first time. Then, a high pressure reductive gas is
introduced through the gas injection velves 116 and
116' substantially at the same pressure as the pressure in
reactor 110, after which injection valves 116 and 116'
45 are maintained in a closed position.
A mixture of slurry and a high pressure reductive zas
which has been preheated to about 300' to 500' C. is
introduced through the inlet port 113 into reactor 110 at
a slurry flow rate of I to 3600 m/hour. preferably IC'; to
50 400 m/hour. In this case, the reactor 110 is maintained
at a temperature of about 300' to 500' C. and a pressure
of about 50 to 700 arms. The mixture thus introduced
under pressure is separated into a solid-lean laver A
(this will be referred to as laver A) and a solid-rich lave-
55 B containing ash. catalyst. and unreacted coal fines in .i
uniformly or thoroughly mixed condition. This wiii be
referred to as a laver B.) In laver B. ash and catalyst are
condensed and accumulate so that the li iuefaction reac-
tion is promoted. On the other hand. in layer A. w nick
b0 is heated in the presence of hydrogen at a cos partial
pressure or a small amount of hydrogen almost under
catalyst-free conditions, a light oil fraction or a react:on
product. which possesses naphthenic or paraffinic-etch
properties and which results from exce'sr'.e hydr; zcra-
e5 tion, is subjected to a dehvdrogenar..on-cycioror?.c.-r.-
densation reaction, thereby nemz zonyerte.] i:,i,-
heavy oil fraction which has ar;'manic nroreru:..k *i:c-
is best suited as a netatlur icaj .aroonac )u' ^:a::r:a.
I B-257
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-
13
4,219,403
Layer A is continuously withdrawn through outlet
port 114. while a mixture of a slurry and a high pressure
reductive gas is fed under pressure through inlet port
113 into reactor 110, so that the interface between layer
A and layer B ascends beyond the tip of the tube 111.
At this stage, valve 11S is opened to bring the first
solid accumulator 112 into communication with reactor
110. Since accumulator 112 and reactor 110 are main-
tained substantially at the same pressure level, layer B is
not introduced into the accumulator 112. Then, the gas
pressure flow rate control valve 118 is opened, so that
gas is discharged from the accumulator 112 at a rate
proportional to the rate at which layer B is being intro-
duced therein. (For instance, when the solids are pres-
ent in the slurry in an amount of 25 to 40%, when the
high-pressure-reductive-gas-feed rate is 14 to 30
Nm3/hour, when the feed rate of slurry is 50 to 100
kg/hour, when the volume of the reactor is 100 liters,
when the reaction temperature is 400' to 450' C., and
when the reaction pressure is 70 to 150 atms, the feed
rate of layer B is 3 to 20 kg/hour.) As a result, layer B
is introduced at a given flow rate into accumulator 112
so that the interface between layer A and layer B
reaches an equilibrium at a given level with the result
that the volume ratio of layer A to layer B may be
maintained at 1/6 to 2, as shown in FIG. 5. The above
ratio is well suited for hydrogenation in layer B and
dehydrogenation-cyclopolycondensation in layer A.
When a sufficient amount of layer B has been intro-
duced into the solid accumulator 112, valve 115 is
opened, valve 115 is closed, and the first accumulator
112 is shut off from the reactor 110, so that layer B may
be introduced into the second accumulator 112. The
layer B, which accumulates in the fast accumulator 112,
is subjected to flash-distillation by opening valve 118,
while residuum solids are discharged through valve
120, which is maintained in its open position. Subse-
quently, accumulator 112 is pressurized to the same
pressure level as that in reactor 110. This cycle of opera-
tion is repeated by alternately using the accumulators
112 and 112'.
Referring to FIG. 6, reactor 121 has a tube 122 which
is inserted therein through its base and opens into the
reactor through its open tip, in addition to an inlet por-
tion 123 adapted to introduce a mixture of slurry and a
high pressure reductive gas, and an outlet port 124
adapted to withdraw a solid-lean layer therethrough.
In the operation of the reactor 121 shown in FIG. 6,
a mixture of slurry and a high pressure reductive gas,
which has been preheated to about 300' to 500' C. is
introduced via inlet port 123 into the reactor 121, which
is maintained at a temperature of about 300' to 500' C.
and a pressure of about SO to 700 atms. The mixture thus
introduced is separated into layer A (a solid-lean layer)
and layer B. which includes ash, catalyst, unreacted
coal fines and the like in a uniformly or thoroughly
mixed condition, i.e., a solid-rich layer. In layer B, ash
and catalysts are condensed and accumulate, thereby
promoting the hydrogenation reaction. On the other
hand. in layer A, as in the case of FIG. 5, the dehy-
drogenation-cyclopolvcondensation reaction takes
place, so that the product is convened into a heavy oil
fraction.
Laver A is continuously withdrawn through outlet
port 123. while a mixture of slurry and a high pressure
reductive gas is continuously fed through inlet port 123
under pressure so that the interface between the laver A
and the layer B ascends.
14
On the other hand, the open tip of tube 12
position of 6/7 to 3 of the height of reactor
the interface reaches the open tip of the to
t is set at a
121. When
122, layer
rough the
of a mix-
the slurry
f the high
when the
when the
e reaction
e reaction
f laver B,
.s a result,
inity close
ie ratio of
nge of 1/6
B (as well as the layer A) is withdrawn t
open tip at a rate proportional to a feed rat
ture. (For instance, when the solid content o
is 25 to 40% by weight, when the feed rate
pressure reductive gas is 14 to 30 Nm3/hour
feed rate of the slurry is 50 to 100 kg/hour
volume of the reactor is 100 liters, when t
temperature is 400' to 450' C. and when t
pressure is 70 to 150 aims, then the rate
which is withdrawn, is 3 to 20 kg/hour.).
the interface reaches an equilibrium in the vi
to the open tip of tube 122, so that the volu
layer A to layer B may be maintained in the r
to 2. (See FIG. 6)
tube 122 is flash-distilled into solid and liqui
acted coal fines, catalysts, and the like.
As is apparent from the above description,
reactors is reduced, while the flow velocity
tion mixture within the reactor is lowered,
vides the same advantages as those of a pisto
reactor.
Attention will now be turned to the thi
and S. A hydrogenation reactor 210 is equip
other end of the tube.
adapted to introduce a mixture of a slurry
pressure hydrogen-rich gas, and an outle
reactor 210 is connected via a pipe 212 and v
equipped with a gas injection valve 216. an
charge pipe 219 equipped with a gas pressur
control valve 218 are connected to the top
accumulator 212, while a solid withdrawin
bottom portion of the ash accumulator 212.
two lines which are connected to two ash ac
212 and 212', which are arranged in parallel
bottom of
fractions.
rain unre-
he diame-
number of
f the reac-
td the set-
actor pro-
flow type
i embodi-
o FIGS. 7
d with an
d therein.
212 at the
port 213
nd a high
port 214
top. The
lye 215 to
pipe 217
a gas dis-
flow rate
of the ash
pipe 221
ed to the
n the em-
ched into
umulators
with each
other, respectively. As in the previous embodiment, like
parts in the second ash accumulator are designated by
like primed reference numerals, which are us d in com-
mon with those of the first ash accumulator 212.
As shown in FIG. 8, a pipe 213 leading from pre-
heater 203 is connected to reactor 210 andithe outlet
port of reactor 210 is connected to a separator 205 via
the ash accumulators 212 and 212', while
201.
accumulators 212 and 212' are isolated from
pressure hydrogen-rich gas is introduced t
as supply
d 217' for
olid-with-
lurry tank
tus of the
8. the ash
eactor 210
s pressure
top valves
yen. a high
rough gas
t
t
t
1
R
I
Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-
 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-0
t
t
1
I
t
t
I
1
I
-1.219,403
15 16
iniection .gives 216 and 216 into ash accumulators 212 Lion reaction, while the dehydrogenation-
and 212' in order to bung the pressures therein to the cyclopolycondensation reaction is promoted in the soi-
level of the pressure in reactor 210. after which the id-lean laver so that the yield of the heavy oil fraction.
injection valves 216 and 216' are maintained closed. suitable for use as a metallurgical carbonaceous mate-
A mixture of slurry and a high pressure hydrogen- 5 rial, is increased. In addition, solids may be separated in
rich gas. which have been preheated to about 300' to the reactor so that the ash content of the mixture mas be
500' C. is introduced at a slurry flow rate of I to 3600 reduced and the catalyst may be reused, which provides
m/sec. preferably 10 to 400 m /sec. into reactor 210 a considerable economic advantage as well as avoiding
which is maintained at a temperature of about 300? to the public nuisance problem of the disposal of the cata-
500' C. and a pressure of about 50 to 700 atms The 10 lyst wastes. The conditions of the operation are the
mixture thus introduced under pressure is separated into same as those of the preceding embodiment. i.e . the
a solid-lean layer A and a solid-rich laver B containing withdrawal rate of laver B should preferably he in the
ash, catalyst, and unreacted coal fines in a uniformly ranee of 3 to 20 kg under the same conditions as those
mixed condition. A hydrogenation reaction is promoted of the preceding embodiment.
in layer B because ash and catalyst are condensed and 1! The fourth embodiment of the liquefaction process of
accumulate therein. In layer A. the mixture is heated in the present invention will be described with reference
the presence of hydrogen at a low partial pressure or a to FIGS. 9 and 10.
small amount of hydrogen almost under catalyst-free FIG. 9 shows a reactor 310 of the present invention
conditions, so that a light oil or part of a product which The reactor 310 is provided with tube 311 which is
possesses naphthenic or paraffinic-nch properties, be- -0 inserted through the base of the reactor with its open tic
cause of the addition of an excessive amount of hydro- positioned therein. The reactor 310 further is provided
gen is converted into a heavy oil fraction which has with an inlet portion 312 in its base which is adapted to
aromatic properties and is therefore suitable as a metal- introduce a mixture of a slurry and a high pressure
lugrical carbonaceous material, as prepared from the
dehydrogenation-cyclopolycondensation reaction. 25 hydrogen-rich gas, as well as an outlet port 313 at the
Layer A is withdrawn through outlet port 214 into top which is used to withdraw a solid-lean laver there-
separator 205, while a mixture of the slurry and a high from. As shown in FIG. 10, a pipe leading from pre-
pressure hydrogen-rich gas is continuously fed through heater 303 is connected to inlet port 312 of reactor 310.
inlet port 213 into the reactor, so that the interface while outlet port 313 is connected to separator 305. The
between layer A and layer B ascends beyond the open 30 lower end of tube 311 is connected to solid-liquid sepa-
tip of tube 211. In this stage. valve 215 is opened in rator 314.
order to bring the first ash accumulator 212 into com- In the operation of the liquefaction apparatus of the
munication with reactor 210. Accumulator 212 and present invention, a mixture of a slurry and a high pres-
reactor 210 are maintained almost at the same pressure sure hydrogen-rich gas, which has been preheated to a
level so that layer B is not fed into accumulator 212. 35 temperature of about 300' to 500' C. is introduced at a
Then, the gas pressure flow-rate control valve 218 is slurry flow rate of I to 3600 m/hour, preferably 10 to
opened so that gas may be discharged from accumulator 400 m/hour through inlet port 312 into reactor 310.
212 at a rate proportional to the feed rate of laver B. which is maintained at a temperature of about 300' to
Layer B is fed into accumulator 212 at a given feed rate 500' C. and a pressure of about 50 to 700 arms. The
so that the interface between layer A and layer B 40 mixture thus introduced under pressure into reactor 310
reaches a given equilibrium level above the open tip of is separated into a solid-lean layer A and a solid-nch
tube 211, with the result that the volume ratio of layer layer B including ash, catalyst, and unreacted coal fines
A to layer B may be maintained over a range of 1/6 to in a uniformly or thoroughly mixed condition. In layer
2. (FIG. 7) The above ratios are well suited for the B. since ash and the catalyst settle and accumulate, the
hydrogenation reaction in layer B, and the dehy- 45 hydrogenation reaction may be promoted in the reac-
drogenation-cyclopolycondensation reaction in laver tor. Laver A is heated in the presence of hydrogen at a
A. low partial pressure or a small amount of hydrogen.
Valve 215' is opened when layer B is introduced into almost in a catalyst-free condition, and a light oil or a
ash accumulator 212 in a sufficient amount. Thereafter, portion of the reaction product which has achieved a
valve 215 is closed so that the first accumulator 212 is 50 naphthenic or paraffinic-rich property by the addition
isolated from reactor 210, thereby introducing layer B of an excessive amount of hydrogen is subjected to a
into the second accumulator 212'. in the same manner as dehydrogenation-cyc!opolycondensation reaction so
that of the first accumulator. The pressure reducing that the reaction product is converted into a heavy oil
valve 218 is opened and the mixture is flash-distilled of an aromatic-rich property. and is therefore well
from the First accumulator 212. After the pressure in the 55 suited as a metallurgical carbonaceous material. thereby
accumulator 212 has been returned to atmospheric pres- improving the yield of the heavy oil product.
sure. stop valve 220 is opened so that solids are with- Layer A is withdrawn through outlet line 313 into
drawn through the solid withdrawing or outlet pipe 221 separator 305. while a mixture of a slurry and a high
and fed to slurry tank 201 for reuse. Subsequently, accu- pressure hydrogen-nch gas is continuously introduced
mulator 212 is pressurized to the same pressure level as 0 through inlet port 312 so that the interface between
that in reactor 210. The above cycle of operation is laver A and laver B ascends.
repeated for the alternate use of accumulators 212 and On the other hand, the open tip of the tube 311 is set
212'. to a height of b/' to J of the height of reactor 310
As is apparent from the above-described liquefaction When the interface between the two layers reaches the
process )f :he present invention. a mixture is separated e' open tip of the tube in reactor 310. layer B is withdraw ti
tnt; a ,o!id-lean !aver and a solid-rich layer for different through the open tip at a rare which is zommensurate
rtes ,f reactions, so :hat ash anu . awty,t contents may with :he fend rate )f the mixture Al a re,wt. the rte'
7e to settle :n order to promote the h%ar,oeena- "ate i, maintained :n ?he ?.:c:nit' ; t,,,e t,, the 'reel ':r"
B-259
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 Approved For Release 2007/02/20: CIA-RDP83M00914R001000060018-C'
17 4,219,403 18
tube 311 at all times, so that the volume ratio of layer A the introduction of the solid-liquid mixture from the
to layer B may be maintained from 1/6 to 2. (FiG. 9) first solid-liquid separating device 410 to the second
The layer B thus withdrawn is separated into solid solid-liquid separating device 410'. for the separation of
and liquid fractions in solid-liquid separator 314, while solid and liquid and the accumulation of solids. the
solids are delivered for reuse to slurry tank 1, and the S other hand, after the switching operation, the pressure
liquid fraction is fed to separator 305 for further pro- reducing valve 414 is opened while the stop vale 417
cessing by conventional prior are procedures. and 415 are closed so as to isolate the aforesaid sol d-liq-
The advantages and conditions of withdrawal of uid separating device from the other system in order
layer B are the same as those in the preceding embodi- that the pressure in the device may be reduced to tmo-
ment. 10 spheric pressure instantaneously for flash-distil aeon,
The fifth embodiment of the liquefaction apparatus of thereby separating the same into gas-liquid and lids.
the invention will be described with reference to FIGS. The gas and liquid are withdrawn through the g s-liq-
11 and 12. uid outlet pipe 413 and line 421. The solids whit con-
As shown in FIG. 11, a solid-liquid separating device dense are withdrawn through the solid outlet po 418
410 is positioned downstream of reactor 404, thereby 15 in the base of the solid accumulating tank by o ning
efficiently separating solids from the reaction mixture stop valve 419. When the first solid-liquid sepa acing
which is introduced from reactor 404. The solid-liquid device 410 becomes empty and the second solid liquid
separating device 410 consists essentially of a liquid separating device 410' is filled with solids, the int oduc-
cyclone 411 which is one type of a solid-liquid separa- tion of the solid-liquid mixture is switched from the
tor, and a solid accumulator 412, which is connected to 20 second solid-liquid separating device 410' to th first
the bottom portion of cyclone 411. A gas-liquid with- solid-liquid separating device 410. Likewise, flash istil-
drawing or outlet pipe 413 is connected to the top por- lation is conducted therein for separation of the mixture
tion of liquid cyclone 411, and a pressure reducing into gas, liquid and solids. In this manner, two solid-liq-
valve 414 is provided on a branch line of pipe 413, while uid separating devices are used alternately fora effi-
a stop valve 415 is provided in another branch line of 25 cient operation by a so-called batch system operation.
pipe 413. A reaction mixture inlet pipe 416 is connected The gas and liquid effluents withdrawn through lines
to the top portion of liquid cyclone 411 at a position 421 and 421' pass through a condensor, as requited, so
lower than the point of juncture of the gas-liquid with- that a portion of the gas may be cooled and liq efied.
drawing pipe 413 with cyclone 411. A stop valve 417 is The liquid is further distilled in a distilling colts in. On
provided on pipe 416. In addition, a stop valve 419 is 30 the other hand, the liquid effluent withdrawn t rough
provided at the solid outlet port 418 at the base of solid- lines 422 and 422' is further distilled in a distil W g col-
accumulating tank 411 umn so that the solvent which is recovered is ret sed as
Two or more solid-liquid separating devices 410 and a slurry solvent. The gas product withdrawn from the
410' are provided as shown in FIG. 12, directly or via a top portion of the gas-liquid separator 420 is cool and
gas-liquid separator 420 downstream of the reactor 404. 35 liquefied in a condensor as required.
(Two solid-liquid separating devices 410 and 410' are As is apparent from the foregoing discussion, a lique-
provided in FIG. 12.) Like parts in the second solid-liq- feed reaction product may be separated into solids and
uid separating device in FIG. 12 are designated with liquid under considerably low viscosity conditions
like primed reference numerals for common use with thereby dispensing with the prior art necessity of add-
the corresponding parts of the first solid-liquid separat- 40 ing a light oil to the liquid to lower the vi osity
ins device 410. The operation of the apparatus of the thereof, thus allowing for the separation and rem val of
present invention for separating and removing solids solids in an efficient manner with the accomp ying
from a liquified reaction product will be described with improvement in quality. In addition, the size of a appa-
reference to FIG. 12. A mixture from the top of reactor ratus may be reduced to a considerable extent i com-
404, which is maintained at a temperature of about 300' 45 parison to the size of the prior art apparatus thus hiev-
to 500' C. and a pressure of about 50 to 700 arms, is ing the desired saving in equipment investmen . Fur-
passed through the gas-liquid separator 420 so that gas thermore, upon flash distillation. by reduction o pres-
may be withdrawn from the top of the separator 420, sure in the system, solids do not pass through the pres-
while a solid-liquid mixture is introduced into the first sure reducing valves, thus preventing valve corrosion
solid-liquid separating device 410 through its base. The 50 problems. This further permits the reduction f the
solid-liquid mixture introduced into the solid-liquid pressure to atmospheric pressure instantan ously,
separating device is somewhat low in temperature and thereby avoiding the need to provide many sepa ators.
pressure in comparison to the temperature and pressure Still furthermore. in the de-ashing operation of the pnor
of the reaction mixture prior to introduction into a gas- art, heat is needed to lower the viscosity or the mixture'
liquid separator. When a solid-liquid mixture is intro- 55 while the apparatus according to the present inv~enuon
duced into the solid-liquid separating device 410, the requires no such heating, thus saving energy.
stop valve 417 on the inlet pipe 416 is opened. while the Having now fully described the invention. it will be
stop valve 417' on the inlet pipe 416' to the second apparent to one of ordinary skill in the art that; many
solid-liquid separator 416' is closed. changes and modifications can be made thereto ~~yy~ ithout
The solid-liquid mixture thus introduced is separated 60 departing from the spirit or scope of the invent;odi as set
into a solid-lean phase and a solid-nch phase in the forth herein.
liquid cyclone 411. The liquid overflows through the What is claimed as new and intended to be secured by
gas-liquid withdrawing pipe 413, while stop valve 415, Letters Patent is:
pressure reducing valve 414 and stop wive 419 are 1. A coal liquefaction process which camp s:
closed. The solids accumulate in the solid accumulating 65 admixing coal fines with a hydrocarbon solvent hav-
tank 412. When the solids have accumulated in the solid ing a boiling point greater than 150' C. to orm a
accumulating tank 412, the stop valve 417 is dosed, coal slurry:
while the stop valve 417' is opened in order to itch admixing with said coal slt:rr% a hvdr!,ern-n. 'n _as.
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hydrogenating said coal slurry by heating said hydro-
gen-containing admixture at a temperature of from
300' to 500' C. and a pressure of from 50 to 700
atms whereby said coal fines are liquified and a
solid-liquid admixture is formed;
separating said liquid-solid admixture into liquid and
solid fractions; and
dehydrogenating and cyclopolycondensing said liq-
uid fraction at a temperature of from 400' to 500'
C. and a total pressure of 70-150 atms in the pres-
ence of hydrogen at a low partial pressure wherein
20
introducing said coal slurry-gas mixture into the
lower portion of a reactor at an upward flow rate
such that a solid rich layer is formed in the lower
portion of said reactor and a solid lean layer is
formed in the upper portion of said reactor;
maintaining said reactor at a temperature of from
300' to 500' C. and at a pressure of from 50 to 700
atms whereby said coal slurry is hydrogenated in
said solid rich-layer and the resulting liquid reac-
tion product is dehydrogenated and cyclopolycon-
densed in said solid lean layer; and
withdrawing a portion of said solid lean layer to
thereby recover a heavy oil product.
7. The process of claim 6, wherein the volume ra::c o;
said solid lean layer to said solid rich layer is from 1.6 to
2:1.
8. The process of claim 6, wherein said coal slurry
contains a hydrogenation catalyst.
9. The method of claim S. wherein said solid lean
layer is essentially free of said hydrogenanon catalyst.
10. The process of claim 6, wherein a portion of said
solid rich layer is withdrawn; the solid and liquid com-
ponents of said layer separated; and the solid compo-
nents admixed with said coal slurry.
11. The process of claim 6, wherein said upward flow
rate is from 1 to 3,600 m/hour.
12. The process of claim 11, wherein said upward
flow rate is from 10 to 400m/hour.
. . * S
said low partial pressure of hydrogen is in the
range of 7 to 70% of said total pressure,
to produce an aromatic-rich heavy oil product. 15
2. The process of claim 1, wherein said coal slurry is
hydrogenated in the presence of a hydrogenation cata-
lyst.
3. The process of claim 2, wherein said hydrogena-
tion catalyst comprises cobalt and molybdenum. 20
4. The process of claim 2, wherein said hydrogena-
tion catalyst comprises iron.
S. The process of claim 4, wherein said hydrogena-
tion catalyst comprises iron and sulfur.
6. A coal liquefaction process which comprises: 25
admixing coal fines with a hydrocarbon solvent hav-
ing a boiling point greater than 150' C. to form a
coal slurry;
admixing with said coal slurry a hydrogen-rich gas;
30
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JAPANESE COAL LIQUEFACTION TECHNOLOGY
31. Mitsui Engineering and Shipbuilding Direct Coal Liquefaction Process
Process Description
This process, which is being developed by the Mitsui Engineering and
Shipbuilding, Co., a Mitsui Group Company, is based on the direct hydr gen a-
tion process flow diagram for their newly constructed 2.4 ton/day proc ss
development unit. The key operation steps that are different from other
processes include the following: (1) preheated coal slurry/H2 gas str am
containing disposable FeS catalyst is introduced into the tubular liquefaction
reactor where additional quantities of H2 gas are injected at different stages
in the reactor system at about 750? to 930?F and 2840 to 4270 psia, (2 the
slurry phase in the high-pressure separator is separated into solids
(unreacted coal, catalyst, ash, etc.) and oil by means of a centrifuge or a
hydroclone system, (3) oil thus separated is fractionated into naphtha light
oil, heavy oil, and bottoms via the atmospheric and vacuum towers, and 4)
single stage operation.
Although this development is somewhat different than the original IG
Farben process it closely parallels research in West Germany to modify the
original IG Farben process.
The direct hydrogenation process for coal liquefaction is the third of
the three major processes under the Japanese Sunshine project. The design of
the process emphasizes improving the rates of coal liquefaction reaction to
achieve high liquid hourly space velocities (LHSV) (i.e., of at least bout
10). The design feature to increase both the mass transfer and chemic 1
reaction rates consist of (a) the use of the tubular reactor to obtain
turbulent flow, (b) promotion of the "bubble flow" mode, (c) maintenan a of
high hydrogen concentration in the process stream by use of multistage
injection of fresh and recycle hydrogen streams into the reactor syste#n, and
(d) application of high reaction temperature (400-500?C or 750-930?F) and
pressure (19.6-29.4 MPa or 2840-4270 psia). Another feature of the process is
use of the disposable FeS catalyst.
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Relationship to Prior Technology
The Mitsui process is based on the IG Farben Process and has been further
refined by the government industry development laboratory, Hokkaido, in
Sapporo City. The laboratory became the first Japanese institution to have
completed a "continuous" test facility back in March 1979. It is still being
run to seek the optimal reaction conditions for the direct liquefaction opera-
tion at the temperature range of 400 to 450?C and pressure ranging from 200 to
700 atm. The bench-scale one- and five-liter reactors are utilized in a six--
week cycle: one week for test preparation, two weeks for continuous run, one
week for waste disposal, and the last two weeks for analysis of coal liquids.
The Sapporo, Hokkaido, laboratory is confident that the direct liquefac-
tion route will eventually prove the most reasonable and practical technology
largely because of its simplicity. This confidence stems partly from the
technical know-how accumulated over the past 16 years. The lab's basic work,
aimed at utilizing the northern-most island's coal resources, started in
1965. While still providing technology to the Mitsui engineering group, the
laboratory is about to launch its study to pursue the so-called coal
chemistry - an avenue of research intended to make best use of coal liquids
for production of chemicals.
The 0.1 tonne/day bench-scale unit is continuing its operation to conduct
the support R&D for the process development unit. Main subject areas of
studies include (a) evaluation of PDU operating conditions, (b) investigation
of liquefaction characteristics of different coal types, and (c) determination
of properties and material balance of the recycle solvent. Basic studies
conducted by Hokkaido and Yamagata Universities, Asahi Chemical Industry
Company, Hitachi Ltd., Japan Steel Works, and the National Chemical Laboratory
for Industry are (a) catalyst development, (b) recovery systems for slurry
feed and power, (c) reactor materials, and (d) coal pretreatment methods.
Operating Facilities
A 2.4 metric ton/day facility has been completed at the Nippon Kokan
Kawasaki shipyard. This direct coal liquefaction PDU was completed at a cost
of $19 million.
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Major Funding Agency
This process is one of four coal liquefaction efforts which is being
funded by the Japanese New Energy Development Organization (NEDO). This pro-
cess will compete for funding with the Sumitomo and Mitsubishi efforts or
funding which will probably be used to construct one 500 metric ton/day
demonstration facility in the 1985-86 timeframe. The Mitsui headed gro
developing this direct coal liquefaction process consists of the Mitsui
Engineering and Shipbuilding Co., Nippon Kokan, Asahi Chemical Industry Co.,
Hitachi Ltd., and the government controlled Electric Power Development o.
Capital Costs
Capital costs for a commercial size facility have not been publish d.
However, this process should be similar in cost to those being develope in
West Germany which are based on the IG Farben process.
Technical Problems
No major technical problems have been reported.
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JAPANESE COAL CONVERSION PROCESS
32. Sumitomo Solvent Extraction Coal Liquefaction Process
Process Description
Because of the limited amount of published information and the lack of
pilot plant data, little is known of the Sumitomo coal liquefaction process.
A rough process flow diagram is presented in Figure 1. As shown, coal is
dehydrated and pulverized before being mixed with a hydrogen-donating solvent
(tetralin) in the ratio of 2 to 5 times the coal volume to solvent.24 The
slurry is then mixed and heated before entering the high-pressure, high-tem-
perature liquefaction reactor. The liquefaction reactions occur at 400 to
450?C and 100 to 150 atmospheres. The lighter products are cooled and subse-
quently separated into gaseous and liquid streams. The bottoms from the high
pressure and high temperature liquefaction reaction are sent to a solid/liquid
separator where solid solvent treated coal is produced. The liquids produced
in this separator tower are cooled and sent to a distillation tower. During
cooling, the heavier liquid products are collected; lighter liquids and spent
solvent are separated in the distillation tower. The liquids product yield is
reported to be 3.3 bbls/ton of coal.26 The spent solvent is further cooled
before being sent to storage. The lighter oils are also cooled and separated
before being sent to storage.
The spent solvent (naphthalene) is then converted back to the hydrogen
donating solvent tetralin. This is accomplished by heating and mixing hydro-
gen gas with the naphthalene prior to entering a series of reactor towers
where the actual transformation occurs. The resulting donor solvent tetralin
is then sent to storage for reuse.
The key to the process is the availability of hydrogen for the solvent
reclaiming step. Although some hydrogen is reclaimed from the process, most
hydrogen requirements must come from an external source. For this purpose SMI
has developed a molten iron bath coal gasification process to generate the
necessary hydrogen.
In the combined coal liquefaction/coal gasification mode, the hard-to-
liquefy portion of the incoming coal is separated and fed to the molten iron
coal gasification reactor for the production of hydrogen. The primary design
feedstock choice for the liquefaction process is subbituminous coal. With
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this combined process and feedstock the overall process efficiency is claimed
as less than 60%.21 A technology fact sheet for the this process is presented
in Table 1.
In 1970 the Sumitomo Coal Mining Co., Ltd, (SCM) a member of the Sumitomo
Group, started research in the area of coal liquefaction in anticipation of an
era of increased coal utilization. In 1976 SCM constructed a continuous
operation 3 metric ton/day solvent refined coal liquefaction pilot plant at a
coal mine near Akahera in central Hokkaido. This pilot plant produced a low
ash, sulfur, and melting point solid product. In the process, a mixture of
pulverized coal and hydrogen doner coal-derived solvent were heated to 400?C
under high pressure. The pilot plant achieved more than 100 hrs of continuous
operation in 1978.2,3
In 1975 another Sumitomo Group member, Sumitomo Metal Industries, Ltd.
(SMI) of Osaka, initiated research on ways of converting steam coal to binder
pitch for metallurgical coke production. This investigation was prompted by a
shortage of metallurgical coal in Japan at that time. Because of SCM's
experience in coal conversion technology, SCM joined SMI in studying binder
.pitch production. This effort was soon curtailed by a decline in steel
production which lessened the need for metallurgical coke. In 1978 this
change in direction was illustrated by the support given by the Ministry of
International Trade and Industry (MITI) to SMI and SCM for the development of
the coal liquefaction technology. As part of the Sunshine Program this
project has received full funding by MITI.
As an aside to coal liquefaction research, SMI has developed a ferrous
bath coal gasification process. This gasification process was developed to
help the steel industry toward an "oilless steel plant" capability.5 In addi-
tion, the ferrous bath gasification can also be used to produce hydrogen for
the coal liquefaction process.
Relationship to Prior Technology
The Sumitomo liquefaction process is an extension of research conducted
in the mid 70's for the production of binder pitch. This process is similar
to the Exxon Donor Solvent process developed in the U.S.
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Operating Facilities
With experience gained from SCI's binder pitch research and SCM's solvent
refined coal solid production investigation,3 SCM and SMI constructed a
1 metric ton/day pilot plant at the Hasaki Research Center in Ibaraki
Prefecture. This pilot plant, which produces a heavy, medium and lightweight
liquefied coal, started preliminary operation in February of 1980. Full
operation of the pilot plant began in March of 1981.4
Major Funding Agencies
Funding for this process comes from the Japanese Ministry of Interna-
tional Trade and Industry (MITI). A discussion on the continuation of this
process will be made by MITI and the New Energy Development Organization
(NEDO) in mid 1982.
Technical Problems
The Sumitomo coal liquefaction process is currently in the pilot plant
stage of development. Due to the lack of available data and operating experi-
ence critical problems for this process have yet to surface. But because this
process is similar to other coal liquefaction processes - in particular the
SRC and EDS process - parallel problems might be expected to occur. In the
SRC process corrosion/erosion problems occur in four main process areas that
are similar to those found in the Sumitomo liquefaction process. These areas
include coal receiving and preparation, preheating and dissolving, filtration
and mineral residue drying, and solvent recovery.
In the area of coal receiving and preparation typical problems associated
with similar facilities occurred in coal slurry centrifugal pumping equipment
and high-pressure, plunger-type slurry preheater charge pumps. These problems
centered on packings and seal leakage, as well as check valves and plunger
erosion. In the preheating and dissolving sections a few material problems
developed. These minor problems involve nozzle sleeve lining corrosion and
scale formation. But severe stress corrosion cracking of Type 304 stainless-
steel lining occured in the high- and intermediate-pressure separator
equipment. This stress corrosion cracking was attributed to polythionic acid
and chloride stress corrosion cracking at weld sites. Corrosion was not
experienced in any of the other materials in the separators, including carbon
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Corrosion/erosion material problems also occurred in the filtration and
mineral residue drying areas. Most problems occurred in the heat excharger
material where severe erosion of metal tubes was encountered as well as tube
bulging due to carburization in the 3/8 inch thick 304 SS drying shell.
Severe corrosion/erosion problems have also occurred in the reboil rs,
air coolers, fractionation equipment, piping, and tower walls of the so vent
recovery section of the SRC process. These problems have caused severs major
shutdowns. The most severe corrosion occurred in moving parts and materials
that came in direct contact with liquid solvent. Most of the problems
required a change in material specifications to alleviate the situation.
However, acceptable corrosion rates have yet to be achieved in areas such as
the stainless-steel trays used in the separator towers.
For the most part, the problems discussed above are applicable to most
coal liquefaction processes and not just the SRC processes or the Sumitomo
liquefaction processes. For this reason these generic problems will probably
not hinder any one process toward commercial development.
Capital Costs
The investment cost of the Sumitomo coal liquefaction process developed
by Sumitomo Coal Mining Co. Ltd. and Sumitomo Metals INd., Ltd. has no been
published. However, based on a process flow sheet comparison with oth r coal
liquefaction processes, the Sumitomo process does not differ significantly
from the EDS (Exxon Donor Solvent) process which is being developed by the
Exxon Research and Engineering Co. in the U.S. (see Figures 2 and 3.29 Based
on this comparison, the two processes should have similar capital cost
requirements. The total plant investment costs for the EDS process (i 1975
dollars) are shown in Table 2.30 These costs include engineering, lad,
plant, and G&A costs. Coal gasification facilities for hydrogen prod ction
have been substituted for the naptha reformers of the 1975 EDS design. Also
included is a 20% contigency factor based on the state of the technol gy
readiness. All costs are based on a facility that uses 25,000 tons/d y of
coal and produces 60,000 barrels/day of liquid products.
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T
SLUMP, RCTCL( SOLVENT
FIL'.AYE a(CVCL( SCLV(N!
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N-0.YNT USE
U K?
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Figure 2. EDS CONCEPTUAL BLOCK FLOW DIAGRAM
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Figure 3. SUMITOMO LIQUEFACTION PROCESS CONCEPTUAL
BLOCK FLOW DIAGRAM
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Table 2. TOTAL PLANT INVESTMENT FOR 25,000 TON/DAY
LIQUEFACTION FACILITY
Offsites (including coal preparation,
handling, product tankage, loading,
utilities)
Liquefaction Facilities
Solvent Hydrogenation
Hydrogen Recovery and Gasification
Gas and Water Treating
Total (including 20% contingency)
Analysis Cost Basis: $2.7 billion (1980)
1975 Dollars 1980 Dollars
(millions) (millions)
300 660
300 660
100 220
450 1000
60 130
Coal Feed: 25,000 tons/day @ 12,500 Btu/lb (including
2500 tons/day for make-up hydrogen required
in addition to 7500 tons/day coal residue)
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JAPANESE COAL GASIFICATION PROCESS
33. Sumitomo Molten Iron Coal Gasification Process
Process Description
The Sumitomo process consists of four main systems: coal preparation,
coal injection, molten iron gasification, and gas cleanup. Coal is initially
sized 20 to 50 mm for storage. From storage the coal is sent to a hot-gas-
swept pulverizer where the coal is sized 70% to 80% through 200 mesh and sim-
ultaneously dried to 5% to 15% moisture. The dried, pulverized coal is then
sent to storage before injection into the gasifier.
One of the most important elements of this process is the coal injection
system. This system consists of a specially designed lance that injects coal,
steam, and oxygen simultaneously into the molten iron bath. Unlike other sub-
merged lance designs, this non submerged top-blowing lance utilizes a ultra
high speed blowing technique to form a cavity in the molten slag as shown in
Figure 1.
As the feed is injected into the gasifier a 2000?C flame is generated,
and the coal injected at the molten iron surface is cracked in a matter of
seconds by the extremely high temperatures. The molten bath assures complete
conversion of the reactants while absorbing unwanted elements such as ash and
sulfur. The molten ash that forms on the surface of the molten iron bath is
skimmed off and withdrawn in the molten state. Part of the slag is composed of
CaS, which is formed when the sulfur in the coal reacts with CaO that is added
to the molten bath to promote desulfurization. From the literature it is
unclear whether the CaO is injected with the pulverized coal or added in a
separate manner. The molten iron bath also acts as a buffer when excess coal
is dissolved in the bath as carbon. This helps to maintain a steady gas com-
position.
The possibility also exists to inject scrap metal or iron ore into the
bath instead of steam. The steam, which acts as a heat moderator, can there-
fore be replaced to allow hot metal production as well as synthesis gas pro-
duction. A typical synthesis gas composition for steam or scrap injection is
presented in Table 1.
The pilot plant gasifier has a fire brick lining which contains the 1500?
to 1600?C molten iron bath. This pilot plant gasifier has a maximum inner
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Injection of Steam
Use of Scrap
Generated Gas
2150-2220 Nm3/T - Coal
20002050 Nm3/T - Co
CO
60 - 64%
64 - 68%
C02