EPA-AA-SDSB-82-05
Technical Report
Indirect Coal Liquefaction Processes
by
David Fletcher
and
John McGuckin
February 1982
NOTICE
Technical Reports do not necessarily represent final EPA decisions
or positions. They are intended to present technical analysis of
issues using data which are currently available. The purpose in
the release of such reports is to facilitate the exchange of tech-
nical information and to inform the public of technical develop-
ments which may form the basis for a final EPA decision, position
or regulatory action.
Standards Development and Support Branch
Emission Control Technology Division
Office of Mobile Source Air Pollution Control
Office of Air, Noise and Radiation
U.S. Environmental Protection Agency
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Table of Contents
Page
I. Introduction/Summary 1
H. History of Methanol Production 7
III. The Methanol Production Process 9
IV. Gasification Technology 11
A. Fixed or Slow Moving Bed 12
B. Fluidized Bed 18
C. Entrained Bed 19
V. Synthesis Technology 22
A. ICI Low-Pressure Methanol Synthesis 26
B. Lurgi Low-Pressure Methanol Synthesis 27
C. Haldor Topsoe Methanol Synthesis 27
D. Mitsubishi Gas Chemical Methanol Synthesis .... 28
E. Vulcan-Cincinnati High-Pressure 29
Methanol Synthesis
F. Wentworth Brothers' Methyl Fuel Process 29
G. Chem Systems' Liquid Phase Methanol Synthesis. . . 30
H. Mobil Methanol-to-Gasoline Process 31
I. Fischer-Tropsch Process 32
VI. Comparison of Indirect Liquefaction Design Studies ... 33
A. Methanol from Bituminous Coal. 34
B. Methanol from Subbituminous Coal ......... 44
C» Methanol from Lignite 48
D. Production of Gasoline from Coal via 61
Fischer-Tropsch and Mobil MTG Technology
References 73
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I. Introduction/Summary
The purpose of this paper is to assess the coal to methanol,
methanol to gasoline, and Fiseher-Tropsch. technologies, and to
estimate the capital investment and product cost of these indirect
liquefaction processes. Over the past five years many study
designs have been performed on the production of methanol and
other indirect liquids from coal. Some of these are original
designs, while others are secondary studies, taking one or more
original designs and adjusting economic parameters, etc. Figure 1
shows the chronology of the available indirect coal liquefaction
studies and their interrelationships. Since the secondary studies
only modified the economic basis of the original studies
referenced and not the basic design and, since each secondary
study used a different basis, preventing intercomparison, this
study will restrict itself to the original studies and attempt to
place them all on one single comparable basis. The following is a
list of these original studies:
"Screening Evaluations: Synthetic Liquid Fuels Manufacture,"
Ralph M. Parsons Company for EPRI, August, 1977, EPRI AF-523.[1]
(This report estimates the cost of methanol from four different
gasification technologies, Foster-Wheeler, BGC-Lurgi, Koppers-
Totzek, and Texaco, with Chem Systems methanol synthesis. The
study also looks at the Fiseher-Tropsch process following
BGC-Lurgi gasification.)
"Coal to Methanol Via New Processes Under Development: An
Engineering and Economic Evaluation," C.F. Braun and Company for
EPRI, October, 1979, EPRI AF-1227.[2] (This report covers two
coal to methanol processess Illinois No. ,6 coal to. methanol via
Texaco gasification and Chem Systems methanol synthesis, and
Wyodak coal to distillate fuel and vacuum residual oil via a non-
catalytic hydroliquefaction process in which the residual oil is
processed into methanol by the same process as the coal.)
"Economic Feasibility Study, Fuel Grade Methanol From Coal
For Office of Commercialization of the Energy Research and
Development Administration," McGeorge, Arthur, DuPont Company, for
ERDA, 1976, TID-27606.[3] (Eastern coal to methanol via Texaco
gasification with ICI synthesis.)
"Conceptual Design of a Coal-To-Methanol Commercial Plant"
(Vols. I-IV), Badger Plants, Inc., for DOE, February, 1978,
FE-2416-24.[4] (Eastern coal-to-methanol via Lurgi "slag-bath"
gasification and Lurgi low pressure methanol synthesis technology.)
"Production Economics for Hydrogen, Ammonia, and Methanol
During the 1980-2000 Period," Cornell, H.G., Heinzelmann, F.J.,
and Nicholson, E.W.S., Exxon Research and Engineering Co., April,
1977.[5] (Eastern coal to methanol via Koppers-Totzek and
Shell-Koppers gasification with ICI synthesis.)
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Year
76
77
78
79
80
81
Figure 1
Methanol Report "Tree'
1JEPRI (Parsons)
Screening
Exxon (Chem
Systems)
AlBadger Methanol
Mobil
Badger
Gasoline
Methanol
Use Options
)Wentworth
C.F. Braun
ro
I
C_) Original studies*
/~\Secondary studies-
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"Methanol From Coal, An Adaptation from the Past," E.E.
Bailey, (Davy McKee), presented at The Sixth Annual International
Conference; Coal Gasification, Liquefaction and Conversion to
Electricity, University of Pittsburgh, 1979.[6] (Subbituminous
coal to methanol via Winkler gasification and ICI synthesis.)
"Research Guidance Studies to Assess Gasoline From Coal By
Methanol-To-Gasoline and Sasol-Type Fischer-Tropsch Technologies -
Final Report," Schreiner, Max, Mobile R&D Company, for DOE,
August, 1978, FE-2447-13.[7] (Comparison of eastern coal to
methanol and SNG, and gasoline and SNG by Lurgi gasification/Lurgi
synthesis/Mobil MTG with gasoline from Lurgi gasification and
Fischer-Tropsch synthesis.)
"Lignite-to-Methanol: An Engineering Evaluation of Winkler
Gasification and ICI Methanol Synthesis Route," DM International,
Inc. for EPRI, October 1980, EPRI AP-1592, Project 832-3.[8]
(Lignite to methanol via modified Winkler and ICI synthesis.)
"Production of Methanol from Lignite," Wentworth Bros., Inc.,
and C.F. Braun and Co., for EPRI, September 1979, EPRI AF-1161,
TPS-77-729.[9] (Lignite to methanol via Texaco gasification and
WBI synthesis.)
"Conceptual Design of a Coal-to-Methanol-to-Gasoline
Commercial Plant," for DOE, March 1979, FE-2416-43.[10] (Adds
Mobil process to methanol design of study no. 4 above.)
Methanol; To estimate the cost of producing methanol, all of
the design studies were: 1) normalized to a production yield of
50,000 fuel oil equivalent barrels per calendar day (FOEB/CD) and
a common financial basis and 2) inflated to $1981, as discussed in
a previous report.[11] Of the thirteen designs contained in the
above ten studies, nine used bituminous coal, two used
subbituminous coal and two used lignite as a feedstock. The
studies included eight different coal gasification technologies
(Foster-Wheeler (1), BGC/Lurgi (1), Koppers-Totzek (2), Texaco
(4), Lurgi (1), "Slag-Bath" (1), modified Winkler (2) and
Koppers-Shell(l)) and four different types of methanol synthesis
processes (Lurgi (2), ICI (5), Chem Systems (5), and Wentworth
Bros. (1)).
The Winkler, Lurgi and Koppers-Totzek gasifiers have been
proven on a commercial scale and the Texaco process is very close
to commercialization. The rest of the gasifiers are still
advanced technologies. The Winkler, a fluidized bed reactor, and
Lurgi, a fixed bed reactor, are best suited for the non-caking
western lignite and subbituminous coals. Koppers-Totzek and
Texaco are examples of the entrained bed gasifier which can handle
all types of coal, but may be the only type of gasifier that can
economically utilize the caking eastern coals.
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Of the synthesis units, ICI. and Lurgi are used extensively
today. Wentworth Bros, claim that their process is commercial and
Chem Systems is a new process which is still being tested. Lurgi
and ICI have been competing for the last ten years and both have
highly developed processes, good efficiencies and, according to
Parsons,[1] room for further improvement is small. In addition,
Parsons states that the Chem Systems process only shows a slightly
higher thermal efficiency and lower capital cost than the ICI
system. Since the costs of the proven ICI and Lurgi synthesis
processes are indistinguishable and it appears that the cost for
the Chem Systems process is only slightly lower, it has been
decided to place most of the emphasis here on comparing the costs
of the various gasification technologies, which appear to be more
significant.
The original ranges of product costs and capital costs
reported by the thirteen studies are very large due at least in
part to the large range in plant size ($3.74-12.55 per mBtu for
product cost and $0.401-$5.05 billion for capital, $1981, for
plants ranging from 2,000-58,000 ton per day of methanol). With
such a wide spread of data it is very difficult to estimate the
actual cost of methanol, let alone compare it with any other coal
technologies. After normalizing the costs for the thirteen
studies the ranges of costs were much smaller.
For bituminous coals the product cost ranged from $4.65-9.05
per mBtu for the low capital charge rate (CCR) and $8.14-12.54 per
mBtu for the high CCR. The gasifiers used in these studies are
Foster-Wheeler, BGC-Lurgi, Koppers-Totzek, Lurgi Slag Bath, and
Texaco(2).
Of these gasifiers the Koppers-Totzek is proven, and the
remainder represent advanced technology. The cost of methanol
from these gasifiers are presented in Table 1. When using the
Koppers-Totzek gasifier the cost ranges from $7.23-12.42/mBtu
depending on the capital charge rate; for the Texaco gasifier the
cost ranges from $5.90-6.48 and $9.44-10.41/mBtu; for the other
advanced technology the cost ranges from $5.30-6.08 to
$8.74-9.78/mBtu.
The range of instantaneous plant investment for the nine
cases was $1.93-$2.92 billion (50,000 FOEB/CD plant). As shown in
Table 1, the instantaneous plant investment for the methanol plant
using bituminous coal ranged from $1.99 to $2.21 billion when the
Texaco gasifier was used, $2.92 when the Koppers-Totzek gasifier
was used, and ranged from $1.93-2.22 billion when the other
advanced technology gasifiers were used.
The range of product and capital costs for methanol from
subbituminous coals and lignite are smaller than that of
bituminous. Of the two studies using subbituminous coals, one
uses proven gasification and synthesis technology, Lurgi/Lurgi
[7], while the other uses a gasification technology which the
manufacturer claims is "here now," and a proven synthesis process,
modified Winkler/ICI.[6] The average product cost range is fairly
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Table 1
Product and Capital Costs of Selected
Coal Liquefaction Processes(1981 Dollars)
Texaco
(Bituminous)[2,3]
Koppers (Bitum.)[l]
Advanced Technology
(Bi tuminous)[1,4]
Lurgi (Subbit.)[7]
Modified Winkler
(Lignite)[8]
Texaco (Lignite)
Lurgi Mobil MTG
(Subbit.)[7]
Mobil MTG
Incremental Cost
Fischer Tropsch[7]
Product Cost
($/mBtu)
Product Mix
100% MeOH*
100% MeOH*
100% MeOH*
47.9% MeOH*
49.7% SNG
2.4% Gasoline
Average
100% MeOH*
100% MeOH*
41.2% Reg. Gasoline
53.3% SNG
5.5% LPG
Average
85-90% Reg. Gasoline
10-15% LPG
1.8% LPG
64.5% SNG
2.6% Alcohols
25.3% Gasoline
4.6% Diesel Fuel
1.3% Heavy Fuel Oil
Average
11.5%
CCR
5.90-6.48
7.23 -
5.30-6.08
7.04
5.63
7.04
6.34
5.70
6.92
8.01
6.41
6.25
7.06
1.45
6.56
6.82
8.52
8.52
7.67
6.56
7.60
30%
CCR
9.44-10.41
12.42
8.74-9.78
12.48
9.98
12.48
11.24
9.56
12.24
14.35
11.48
11.20
12.65
2.87
11.36
. 11.80
14.75
14.75
13.28
11.36
13.38
Capital
Cost**
(Billions
of Dollars)
1.99-2.21
2.92
1.93-2.22
2.59
2.17
3.00
2.95
0.68
3.00
* MeOH = 95-98% methanol, 1-3% water, and the remainder higher alcohols.
** Instantaneous capital costs.
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small, $6.16-$6.34 per mBtu for the low CCR and $10.26-$11.24 per
mBtu for the high CCR. The instantaneous plant investment range
is $2.10-$2.59 billion. Although the costs seem to compare
favorably, only the Lurgi/Mobil prices are shown in Table 1. This
is because the modified Winkler/ICI plant size had to be scaled up
significantly where as the Lurgi/Mobil plant size was much closer
to the selected 50,000 FOEB/CD and was therefore considered more
accurate.
Four lignite cases were studied. However, two of these cases
used Texaco gasifiers with coal slurry concentrations which are
still in a developmental stage. The other cases involved the
Texaco gasifier (at an appropriate coal slurry concentration) and
the Winkler gasifier. At this slurry concentration the Texaco
gasifier appeared to have a large economic disadvantage relative
to the Winkler gasifier, so the Winkler was chosen as the best
design. The resulting product costs for the low and high CCRs are
$5.70 and $9.56/mBtu, respectively. The instantaneous investment
plant cost is $2.17 billion. These costs are shown in Table 1.
In summary, the prices which have been chosen for this study
represent two commercially proven gasification technologies,
Koppers-Totzek and Lurgi, a modified Winkler, which the
manufacturer will back financially, and the near commercial Texaco
gasifier. For bituminous coals, the Koppers-Totzek prices are
higher than Texaco's because the former operates at atmospheric
pressure.
MTG; To evaluate the cost of producing gasoline from coal
utilizing Mobil's methanol-to-gasoline (MTG) process, two
different studies by Mobil and Badger were analyzed in the same
manner as the methanol studies.[7,10] Gasoline costs from these
two studies varied widely. Therefore, it was presumed that
incremental product and capital costs for Mobil's MTG gasoline
relative to methanol could be determined from both studies and be
more accurate since methanol costs (capital and product) were
available for the same technology by the same designers.[7,4]
When the cost of gasoline was compared to that of methanol, the
incremental cost of gasoline for both studies was very close.
Since the MTG process is a patent of Mobil's, it is believed that
their study is more reliable; therefore their costs were used in
preference to Badger's, which were slightly higher.
The Mobil study analyzed a few different cases with respect
to the Mobil MTG process. The most economical was the case which
produced gasoline and SNG as the major products. For this case,
the average product cost ranged from $7.06-12.65 depending on the
CCR. The total instantaneous plant investment was $2.95 billion.
These costs are shown in Table 1. By comparing this case with
Mobil's other case (methanol from Lurgi gasification of
subbituminous coal) an incremental cost of gasoline relative to
methanol was determined. Based on a 50,000 FOEB/CD MTG unit, the
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incremental cost of gasoline over methanol was found to range from
$1.45-2.87 per mBtu depending on the CCR. The incremental
instantaneous investment was found to be approximately $680
million for a plant producing all methanol and then gasoline.
These costs are also shown in Table 1.
Fischer-Tropsch; There were two studies which investigated
Fischer-Tropsch synthesis technology, Parsons and Mobil.[1,7]
Since the Mobil study was based on a more thorough design than the
Parsons study, its costs were used in preference. The instantan-
eous plant investment cost for the Mobil case was $3.00 billion.
Its average product cost ranged from $7.60-13.38 per mBtu depend-
ing on the CCR. The costs of the products from this case are pre-
sented in Table 1. The Mobil study was also used to determine the
average product cost difference between Fischer-Tropsch synthesis
and methanol synthesis plants. The instantaneous plant investment
difference is $355 million and the operating cost difference is
$67 million with the Fischer-Tropsch case costing more. The
figures translate into an average product cost difference of
$1.00/mBtu.
II. History of Methanol Production
The primary source of all methanol prior to the 1920's was
the destructive distillation of wood. In this pyrolysis process
air was excluded while the wood was heated to a temperature of
160-400 degrees Centigrade. As the components of the wood heated
they volatilized and thermally decomposed.: The: products were
separated into gases and a condensed liquid called pyroligneous
acid. Upon further distillation this liquid could be separated
into acetic acid, acetone and rather impure methanol. Since the
yield was three to six gallons per ton of wood, the product was
very expensive.[12]
During the pre-World War I period, the development of a syn-
thetic methanol process began in Germany and France. Between 1910
and 1916 there were several patents issued in Europe describing
the chemical reaction of carbon monoxide (CO) and hydrogen (H2)
to form alcohols, ketones, aldehydes, etc. The reaction was
carried out at temperatures of 300 to 400 degrees Centigrade and
at pressures at or above 1500 psi. Catalysts containing chromium,
zinc, manganese and cobalt, or their oxides were used to help the
conversion of the carbon monoxide and hydrogen to methanol.[12]
In 1923, BASF in Germany became the first company to produce
commercial-scale synthetic methanol. The U.S. started importing
synthetic methanol produced from coal or coke in 1924. Soon
Commercial Solvents Corporation and DuPont became interested and
by 1928 each had a commercial plant producing methanol in the
U.S.[12]
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When- coal, and coke derived synthetic methanol hit the U.S.
market there was an enormous difference between the price of
natural and synthetic methanol. Natural methanol cost 68 cents
per gallon while synthetic methanol could be made for 36 cents per
gallon. The price competition was so great that the natural wood
distillers united and managed to persuade the tariff commission to
increase the import tariff to 18 cents per gallon. They also were
able to get legislation passed which mandated the use of natural
methanol to denature ethanol, thus securing a third of the total
methanol market.[12]
The wood distillers managed to keep the price of natural
methanol competitive for a number of years through consolidation
and larger, more efficient plants. However, with the large
discoveries of petroleum and natural gas and the mass production
of high-purity methanol, the synthetic manufacturers were soon
able to lower the price of methanol beyond reach, leaving natural
methanol producers to their captive denaturant market.[13]
The first plants were built in conjunction with other plants
to make use of carbon dioxide or hydrogen by-products. However,
as the demand for methanol grew, plants were built specifically
for methanol production. The first feedstock to be gasified to
carbon monoxide and hydrogen was coal. Later the feedstock was
shifted to oil and then natural gas as large discoveries of these
sources were made and their prices dropped. Natural gas was an
ideal feedstock because it contained very little, if any, sulfur
and its .price was very low.i Thus by the 1960's,. synthetic
methanol in the U.S. was almost entirely produced from natural gas
utilizing the high-pressure methanol synthesis process.[13]
By 1967 the combination of a common feedstock, comparative
technology and a competitive market had stabilized the price of
methanol at 27 cents per gallon. However, in 1967, Imperial
Chemical Industries (ICI) introduced a newly developed
low-pressure synthesis process based on a copper-zinc-chromium
catalyst in place of the zinc-chromium catalysts previously used.
Since these copper catalysts were more reactive than the others,
lower operating pressures and temperatures could be used. In
fact, by the latter part of 1971, the selling price of methanol
had dropped to 11 cents per gallon.[13]
In the 70's the increasing cost of production, the demand for
low-sulfur natural gas and the OPEC oil embargo of 1973 brought
into focus the energy crisis and the finite supply of fossil
fuels. The tripling of oil prices and doubling of the cost of
natural gas caused the price of methanol to triple between 1973
and 1975 (14 cents/gal to 42 cents/gal). Since 1975 the price of
methanol has continued to increase with that of natural gas. The
current price for methanol is between 72 and 80 cents per
gallon.[13]
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Between the time it was first introduced into the U.S. and
today, methanol has exhibited a dramatic growth. For the first 45
years there was a 13.7 percent average annual growth rate.[12] In
the 1930's plant sizes ranged from 20 to 40 tons of methanol per
day. By the early 50's the size has risen to 150-200 tons/day.
In the 70's the capacity has gone from 1,500 to 2,000 tons/day to
single trains of 5,000 tons/day.[13]
Methanol production in the United States is now near 4
million tons per year or about 70,000 barrels per day (BPD).
Virtually all of this is produced from natural gas. [12] The
natural gas (essentially pure methane) is reformed with steam to
produce a synthesis gas consisting mainly of carbon monoxide and
hydrogen. After purification, the synthesis gas is compressed and
combined in a catalytic converter to produce methanol. The
reaction is highly exothermic while the conversion per pass is
relatively small (2-10 percent). Large volumes of unconverted gas
are recycled through the converter in order to achieve high
overall conversion and to assist in removing the exothermic heat
of reaction. Overall CO conversions of 96 to 99 percent can be
obtained.[14]
III. The Methanol Production Process
The basis of all processes for manufacturing synthetic
methanol is the catalytic reaction of carbon monoxide and carbon
dioxide with hydrogen to produce methanol.[12] These reactions
are shown-below.
Carbon Monoxide -f Hydrogen = Methanol
CO + 2H2 = CH3OH (1)
Carbon Dioxide + Hydrogen = Methanol + Water
C02 + 3H2 = CH3OH + H20 (2)
The source of carbon monoxide or carbon dioxide is usually
derived from the partial combustion of a hydrocarbon fuel such as
coal, coke, natural gas, naptha, or a heavy petroleum fraction.
The primary source of hydrogen is water and the hydrogen
contained in the feedstock, which is the case of coal is very low
(3-6 percent). The reactions shown in Equations (1) and (2) are
carried out at pressures between 750 and 4500 psi at a temperature
of 250 to 350 degrees Centigrade in the presence of a metal or
metal oxide catalyst. The metals used depend upon the process,
and are usually proprietary. Catalysts may contain zinc,
chromium, or copper-based compounds or oxides.
A description of a typical coal to methanol process
follows.[15]
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Goal Receiving and Preparation; Vibrating feeders transport
the coal to the sizing equipment, ring mill crushers and rod melts
where the coal is sized for the specific gasifier in which it will
be processed.
Coal Gasification; The coal is heated to very high
temperatures and partially-oxidized to carbon monoxide and
hydrogen in the presence of oxygen (or air) and steam. The
majority of the sulfur is converted to hydrogen sulfide with some
production of carbonyl sulfide. The nitrogen in the coal is
converted to free nitrogen combined with some traces of ammonia
and hydrogen cyanide. The ash is removed from the bottom in a dry
or molten slag depending on the temperature and gasification
technique used.
Gas Cooling; The hot raw gas is cooled and scrubbed with
recycle gas liquor or sour water from the shift converter. Then
the gas is cooled further in a heat exchanger where steam is
produced by the waste heat.
Gas Shift; Here the ratio of hydrogen to carbon monoxide is
increased by adding steam and pushing the following water-gas
shift reaction to the right; CO + H20 = C02 + H2.
Acid Gas Removal; In this process the sulfur is removed from
the synthesis gas to prevent poisoning of the methanol synthesis
catalyst. In the Selexol process hydrogen sulfide is removed
first, and then carbon dioxide and carbonyl sulfide are removed.
In the following Rectisol process, naptha, HCN and water are
removed by washing the gas with a small quantity of methanol.
Methanol Synthesis; In this stage the clean shifted
synthesis gas is catalytically converted into crude methanol by
the following two reactions;
CO + 2H2 = CH3OH and C02 + 3H2 = CH3OH + H20.
Auxiliary Facilities; The functional relationships of the
auxiliary facilities to the major process areas are as follows;
Water Supply - provides for treatment, storage and
distributionofprocess water requirements, including makeup
cooling water.
Water Cooling - provides for treatment, storage and
distribution of process cooling water.
Oxygen Production - cryogenically separates air into oxygen
and nitrogen. Oxygen is used in coal gasification. Some of the
nitrogen is used in carbon dioxide removal, the remainder being
vented to the atmosphere.
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Slag Removal/Char Recovery - separates dissolved gases from
the raw gasifier slag. Also separates char from the slurry
produced in gas cooling. Overhead gases and char are used to
generate steam. Wet slag is sent to slag/ash disposal.
Slag/Ash Disposal - combines wet slag, bottom ash from steam
generation and dusty liquor from flue gas cleanup. Dewatered
slag/ash mixture is suitable for landfill disposal. Wastewater is
used for preparation of coal slurry for coal gasification.
Sulfur Recovery - processes the acid gas stream produced
during hydrogen sulfide removal, converting H2S to elemental
sulfur. Process technology employed is usually Glaus (bypass type
configuration).
Steam Production - uses char, purge and overhead gases, along
with supplemental coal to provide plant steam/power requirements.
Flue Gas Clean-up - renders steam generation product gases
environmentally acceptable for stack discharge to atmosphere.
IV. Gasification Technology
The gasification of coal began in the early 1850's when it
was discovered that the gas could be burned more efficiently than
solid coal and it was cleaner and easier to use. The technology
developed fast and by the 1850's gas light for streets in London
was commonplace:* Between- 1935 and 1960 there were close to 1,200
municipal "gasworks" serving larger towns and cities in the U.S.
However, the introduction of natural gas pipelines in the 1930's
initiated the decline and almost disappearance of coal
gasification in the U.S. With the increased cost of natural gas,
interest in coal gasification has been renewed.
Numerous processes are now being developed to gasify coal,
the most abundant hydrocarbon resource in the U. S., to low-,
medium- and high-BTU gas. In the gasification process coal is
reacted with a mixture of steam and air or steam and oxygen. With
the former, a low-BTU gas is produced with a heating value between
100 to 200 BTU/scf.[16] This low-BTU gas is made up of nitrogen,
carbon monoxide, hydrogen, carbon dioxide and water.[17] This gas
has a low-BTU content because it contains a large, portion of
nitrogen (since air contains 80 percent nitrogen) which dilutes
the energy content of the carbon monoxide and hydrogen produced.
If the coal is mixed with steam and oxygen a medium-BTU gas is
produced consisting of carbon monoxide, hydrogen, carbon dioxide
and methane, which has a heating value between 250 and 400
BTU/scf.[16] High-BTU gas (or synthetic natural gas (SNG)) with a
heating value of 970 BTU/scf can be produced from medium-BTU gas
by methanation or hydrogen removal.[18,19]
There are several factors such as thermal efficiency,
reliability, capital investment, coal flexibility and product
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spectrum which are important and should be considered when
comparing gasifiers. Table 2 shows some of these comparisons for
different gasification systems. For instance thermal efficiency
is important from a processing view. To achieve maximum
efficiency a gasifier should have low oxygen and steam demands,
low unburned carbon and heat losses, and should operate at high
temperatures.[20] However, since some of these factors are not
compatible with others, it is almost impossible to obtain a
gasifier which optimizes each factor. For example high oxygen
requirements (lower efficiency) are necessary to obtain high
carbon conversions (higher efficiency) and to avoid large
by-product formations (lower capital investment).[20] In
addition, elevated pressures (high efficiency) produce more
by-products[20] and interfere with reliability[20] but reduce raw
gas compression requirements. Other desirable factors are high
temperatures which reduce by-products and increase coal
flexibility and capacity.[20] In general then, it is apparent
that there is a trade off between efficiency and some of these
other factors. The best gasifier will be that process unit which
optimizes the majority of these factors while achieving the
highest efficiency.
Before discussing the individual gasifier types it is helpful
to examine the properties of coals used in the U.S. There are
four properties of coal which are important in the process
selection of gasifiers: 1) ash fusion temperature, 2) free
swelling index (FSI), 3) moisture, and 4) sulfur content. The ash
fusion temperature is that temperature at which the ash becomes
fluid. FSI is a measure of a coal's tendency to agglomerate or
cake when heated; the higher the FSI, the greater the
agglomeration.[15]
Eastern coals (predominantly bituminous) typically have low
fusion temperatures (1990-2200°F), moderate to high FSI, low
moisture (4-10 percent by weight as received) and high sulfur
(averaging 2.0 weight percent). Western coals (mainly
subbituminous) exhibit high fusion temperatures (2300-2400°F), low
FSI, high moisture (28 weight percent) and low sulfur (averaging
0.7 weight percent). Lignite, which is found predominantly in
North Dakota, has an even higher moisture content (35 weight
percent) and also a low percentage of sulfur (averaging 0.8 weight
percent).[15,21]
Coal gasifiers are classified according to the way coal is
fed to them. The three main gasifier categories are the fixed or
slow moving bed, the fluidized bed and the entrained bed. Tables
3 and 4 show a list of coal gasifiers by type and Table 5
summarizes the advantages and disadvantages of the three different
types of gasifiers.
A. Fixed or Slow Moving Bed
Fixed or slow moving bed gasifiers consist of beds that carry
or move the coal vertically downward through the zone where it is
heated and decomposed. These gasifiers can be further divided
-------
-13-
Name
Bed Type
Commercial
Coal
Flexibility
Table 2
Comparison of Gasification Systems
Koppers-
Lurgi BGC-Lurgi Winkler Totzek Texaco
Shell-
Koppers
Fixed Fixed
Yes Near
Western Western
By-Product Yes
Efficiency(%) 64
Capacity 500
(STPD Coal)
Fluid
Yes Yes
Western All
Entrained Entrained Entrained
Near Near
All All
Yes
72
1,250
No
57
1,000
No
58
850
No
68-72
2,000
No
75
1,000
Source: [20]"Coal Can Be Gasoline", (Kellogg Co.). Hydrocarbon
Processing, LeBlanc, J.R., Moore, D.O., and Cover, A.E., pp.133-137, June
1981.
-------
-14-
Table 3
Coal Gasifiers
Pressure
Oxygen or Air
Agglomeration
Prevention
FIXED OR SLOW
MOVING BED
Gegas
Lurgi
Merc
Riley-Morgan
Wellman-Galusha
Wilputte
ATC/Wellman
FW/Stoic
Ruhr-100
Woodall-Duckham
BGC/Lurgi
Gf ere
FLUIDIZED BED
Winkler
Rheinbraun
C02 acceptor
Hygas
Synthane
Westinghouse
U-gas
Cogas
EDS (Exxon)
ENTRAINED FLOW
Bell HMF
Koppers-Totzek
Mountain fuel
Shell-Koppers
Texaco
Bi-gas
C-E
Foster Wheeler
Peatgas
Rockwell Int'l
Dry Ash, Single Stage
To 500 psig Air
To 450 psig Oxygen or air
To 105 psig Air
40 in water Air
10 in. water Air
atm Air
Dry Ash, 2-Stage
Air
atm
atm
1,500 psig
40 in. water
Oxygen
Air or oxygen
Slagging
To 400 psig Oxygen
To 400 psig Oxygen
Stirrer paddles
Rotating blades
Spiraling stirrer
Agitator in
rotating bed
Spiraling arms
Rotating arm
None
None
Stirrer blades
None
Stirrer
Stirrer
atm
150 psig
150 psig
1,200 psig
1,000 psig
225 psig
350 psig
10 psig
500 psig
Oxygen or air
Oxygen
Air
Oxygen or air
Oxygen
Air
Oxygen or air
Air
None
Single-Stage
to 225 psig Air
10-12 psig Oxygen and steam
to 150 psi Oxygen and steam
to 450 psi Oxygen or air
to 1,200 psig Oxygen or air
To 1,500 psig Syngas and steam
atm Syn gas and air
atm Syn gas and steam
To 500 psig Syngas
To 1,500 psig Hydrogen
Oxygen and steam
air
Air and steam
Air and steam
Oxygen and steam
Not known
Source: [18] Institute of Gas Technology, Oil and Gas Journal,
7/16/81, pg. 57
-------
-15-
Table 4
Coal Gasification Process Technology Status
Gasifier
Fixed Bed
Lurgi, Dry Ash
British Gas/Lurgi
Slagging
Wellman Galusha
KilnGas
Fluidized Bed
Winkler
U-GAS
Westinghouse
Entrained Flow
Koppers-Totzek
Texaco
Combustion
Engineering
Shell-Koppers
Technology
Process Status
Location
1st Generation Commercial
Worldwide
2nd Generation Semicommercial Westfield, Scotland
1st Generation Commercial
14 operating in U.S.
others outside U.S.
2nd Generation Pilot (1971-) Oak Creek, Wis.
1st Generation Commercial Worldwide
2nd Generation Pilot (1974-) Chicago, IL
2nd Generation PDU (1975-) Waltz Mill, PA
1st Generation Commercial
2nd Generation Pilot
2nd Generation PDU (1974-)
2nd Generation Pilot
Worldwide
Montebello, CA
Mussel Shoals, Ala.
& outside the U.S.
Windsor, Conn.
W. Germany
Source: [16] Oil & Gas Journal; June 29, 1981, pg. 106.
-------
-16-
Table 5
Gasifier Characteristics
Fixed Bed Gasifier Fluidized Bed Gasifier
Advantages
Extensive practical
experience
High carbon conver-
sion efficiency
Low temperature
tars
operation
Large fuel inven-
tory provides
safety, relia-
bility and
stability
Limitations
Sized Coal re-
quired
Coal fines must
disposed of or
handled separ-
ately
Product gas con-
inventory;
tains tars and
heavier hydro-
carbons
Lowest capacity
due to limited
gas-flow rates
Internal moving
parts with high-
er degree of
mechanical com-
plexity
Caking coal tech-
nology not commer-
cially proven
Uniform temperature and
compositions throughout
fluidized zone
Excellent solid-gas con-
tact
No internal moving parts
Can handle wide variety
of coals
Large fuel inventory
provides safety, relia-
bility, and stability
Distributor plate design
is critical
Requires dry coal for
feeding
Entrained Bed Gasifier
Highest capacity per
unit volume
Produces inert slagged
ash with low carbon
content
Product gas free of
and phenols
Handles all types of
coal
No moving parts and
has simple geometry
Less developed than
fixed bed
Critical design areas
include combustor
nozzles and heat re-
covery in presence of
molten slag
Removal of fines re-
quired to prevent
elutriation or flow
instability
Fluidization requirements
sensitive to coal charac-
teristics
Smallest
fuel
requires advanced con-
trol techniques to en-
sure safe reliable
operation
Source: [16] Oil and Gas Journal, June 29, 1981, pg. 101.
-------
-17-
into two types which describe the flow of air in them: updraft
and downdraft.
The simplest air gasifier is the updraft or countercurrent
gasifier which introduces air at the bottom of the furnace where
it first comes into contact with the hottest temperatures of the
reactor. Since the combustion gases immediately enter a zone of
excess char, any carbon dioxide or water present is reduced to
carbon monoxide and hydrogen by the excess carbon. In addition to
producing the desired products, carbon monoxide and hydrogen,
these hot gases contain large amounts of tars, phenols, cresols
and other oxygen containing organic compounds. As the gas rises
its temperature decreases as heat is transferred from the hot gas
to the cooler incoming coal. This low temperature hinders the
oxidation of the coal and is the major cause of the by-products
produced.[17]
One of the problems caused by these by-products (chemicals,
oils and tars), is that they condense in the cooler regions,
causing maintenance problems. In addition, these components
contribute to the majority of environmental problems associated
with fixed bed gasification systems.[22]
The downdraft gasifer is specifically designed to eliminate
the tars and oils associated with the updraft gasifer. Tars and
oils are formed near the middle of the bed (where air is injected)
and carried by the airflow through a relatively large hot zone in
which they have"time to further decompose or be cracked to simpler
gases or char. One of the important results of this cracking is
that an effect called "flame stabilization" occurs which maintains
the temperature range between 800°C to 1000°C. When the
temperature rises, endothermic reactions predominate, causing the
gas to cool; when the temperature drops, the exothermic reactions
predominate, thus heating the gas.[17]
The tars and oils are reduced to less than 10 percent of the
amount produced in updraft gasifiers thereby making gas clean-up
easier and less expensive. Since the gas velocities are low in
both updraft and downdraft gasifiers, the ash settles through the
grate so that very little is carried with the gas.[17]
One example of a moving bed gasifier is the Lurgi gasifier
which is commercially available through Lurgi Kohle and
Mineraloeltechnik. In the Lurgi process, coal is fed into the
gasifier via automatically operated coal locks. As the bed of
coal moves from the top to the bottom of the gasifier it comes in
contact with a counter-current hot gaseous mixture of oxygen and
steam introduced at the bottom which successively dries,
devolatilizes and gasifies the coal. The partial oxidation of the
coal with oxygen supplies the necessary heat for the coal
gasification while the addition of steam prevents the temperature
from rising above the ash fusion (or melting) point. The ash left
after gasification is removed by a rotating grate at the bottom of
the gasifier.
-------
-18-
As shown in Table 3, the Lurgi "dry-ash" fixed bed gasifier
is a first generation unit which has been commercially proven and
is used worldwide.[16,20] The Sasol I plant in South Africa which
has been operating for over 25 years utilizes the Lurgi gasifier
(and also the Fischer-Tropsch synthesis unit) to produce 10,000
barrels per day of fuel. (It and Sasol II are the only existing
commercial-size coal-liquefaction plants in the world.) The Dunn
Nakota project, which is scheduled to produce 85,000 barrels per
day of methanol by 1987 via Lurgi gasification, could be the
largest commercial-scale coal gasification process built in the
U.S.[23] The main disadvantages of the Lurgi gasifier are that it
1) has problems with the caking of eastern coals, 2) produces
byproducts, 3) has high steam requirements and 4) has a low
capacity per volume of gasifier (i.e., high capital cost).
The BGC/Lurgi slagging gasifier is a second generation
reactor which is now being tested in Scotland by Lurgi and British
Gas Corp with support from 13 U.S. companies and DOE.[19] In this
gasifier, coal is fed into the top of the unit by a distribution
system. As the coal descends in a moving bed, it is successively
dried, devolatized and gasified. At the bottom of the gasifier
oxygen and steam are fed and slag is withdrawn. The operating
pressure is 300-350 psig with a gas temperature of 800-1100°F and
an ash temperature over 2000°F so the slag can be removed in a
molten form.[l] Because it does operate in the slagging mode it
can tolerate a higher throughput of coal and oxygen without
entraining coal dust in the product gas.[24]
The latest papers describe this technology as near
commercial.[16,20] Its improvements over the older Lurgi dry-ash
gasifier are a higher efficiency and a reduction in steam use.
However, it still has problems with caking eastern coals and still
produces by-products.
B. Fluidized Beds
Over the last 60 years fluidized beds have been developed to
provide uniform temperatures and efficient contact between gases
and solids. This is accomplished by blowing gas upward through a
bed of solid coal so rapidly that the bed becomes suspended and
churns as if it were a fluid. Fluidized reactors are more compact
because they have a higher throughput (due to higher reaction
rates), but the higher velocity of the gas carries out ash and
char with it that must be removed by cleaning the product gas.
The fluid bed often contains limestone to react with and remove
the sulfur from the coal. Fluidized bed reactors have a
considerably faster heating rate than moving bed gasifiers and,
therefore, the reactor temperature must be held below the
softening or initial deformation temperature of the coal ash which
is typically well below 1040°C. However, at this temperature many
undesirable by-products are stable and the churning of the bed
enables materials at all stages of decomposition to be found
-------
-19-
throughout the bed. Because of this contact, tars and oils have a
tendency to escape from the heating zone before they can be fully
decomposed. This removal and disposal of these by-products can
pose a number of environmental problems for the fluidized
reactor. Claims that they can produce very low tars and char with
recirculation still remain to be proven.[17] Since typical
operating temperatures are low with respect to the ash melting
temperature of coals, the fluid-bed gasifier also has problems
with eastern coals.
The Winkler fluid-bed gasifier is a first generation unit
which is commercially proven and used around the world.[16]
According to DM International over 70 Winkler gasifiers have been
built.[8] The two main disadvantages with the Winkler are that it
operates at atmospheric pressure (large volume per throughput) and
that it has a tendency to clog when using eastern coals. A
pressurized modification of the Winkler is now under development
which should improve its efficiency. [14,20]'
In the two designs of lignite gasification that will be
reviewed in this study, modified Winkler gasifiers have been
used. In both cases the modified Winkler operates at a higher
pressure (65 psig) than the established Winkler which operates at
atmospheric pressure (14.7 psig). The lignite is dried from 35
percent moisture to 8 percent and is then continuously fed by a
pressure lock and screw conveyor system into the Winkler gasifier
where it is maintained as a fluid bed at 65 psig. Steam is
injected near the bottom of the reactor to fluidize the coal and
to cool the larger ash particles discharging from the gasifier
bottom while steam and oxygen are injected at several points
within the bed to gasify the coal. Since the gasifier operates at
high temperatures (1800-1900°F), tars, oils, gaseous hydrocarbons
and carbon present are converted to carbon oxides and hydrogen.
Only a small percentage of methane is left in the raw gas
product. In the fluidized bed, heavier particles such as ash fall
down through the bed into the char discharge, while lighter
particles are carried out of the bed by the product gas. In the
Winkler gasifier approximately 70 percent of the total char is
entrained in the hot product gases leaving the top of the reactor.
This modified Winkler is still being tested and therefore it
cannot be considered to be commercially proven. However, since
Davy McKee believes that this design contains equipment similar to
other high pressure units, they feel that the gasifier is feasible
and are therefore prepared to offer commercial guarantees.
C. Entrained Bed
The entrained bed gasifier, which dates back to the 1950's,
is the most recently developed gasifier. In this gasifier fine
particles of coal are suspended in a stream of oxygen which moves
rapidly into and through the decomposition zone. The entrained
-------
-20-
bed gasifier is typically operated at a temperature above the
melting point of the coal ash. At this temperature, which is
typically 1260-1316°C, the gasification reaction rates are much
faster and many of the undesirable by-products associated with the
fixed bed and the fluid bed systems are unstable and are
destroyed. When the entrained bed gasifier is operated at
pressures substantially above atmospheric, high throughput and
high single pass conversion can be obtained. One drawback is that
the feedstock must be reduced to a relatively small size which
would add to the total preparation cost. However, there is a
tradeoff since the smaller particles are more efficiently gasified.
These gasifiers are also called "slagging" because they
remove the ash in a molten, slag form. One of the big advantages
of entrained bed gasifiers is that they can utilize any type of
coal. As shown in Table 4, Koppers-Totzek, Texaco and
Shell-Koppers are all entrained-bed gasifiers.
In the Koppers—Totzek gasifier pulverized coal is
horizontally injected with steam and oxygen into the reactor which
is essentially operating at atmospheric pressure. The
gasification temperature is around 2700°F. At this high
temperature, the ash is in a molten slag form which drops into a
quench tank and is removed.[1]
The Koppers-Totzek gasifier is a first generation technology
which, like the Winkler and Lurgi, has had extensive commercial
experience, and therefore is considered proven and available
technology.[16,20] Five of the 24 proposed projects submitted to
the Synthetic Fuels Corporation plan to use Koppers-Totzek
gasifiers which would seem to confirm its reliability. It will
handle all types of coal but does require large raw gas
compressors since it operates at atmospheric pressure.
The Texaco gasifier is a coal-slurry fed, high-capacity
gasifier which handles all types of coals and produces very little
by-product. The slurry which is composed of pulverized coal and
water is pumped with oxygen into the top of the high—pressure
(600-700 psig) gasifier and fired downwards. The product gas is
withdrawn through a side nozzle at a temperature around 2500°F.
The molten slag is removed through a slag hopper beneath the
quench chamber.[1]
Since the coal is fed in a water slurry the coal does not
have to be dried. This can be a big advantage over gasifiers
(predominantly for western coals) which use part of their coal to
dry the rest of the feedstock. Drying is expensive, it reduces
efficiency and it raises operating costs.
The high operating pressure is also an advantage since the
synthesis gas must be fed at even higher pressures to the methanol
unit. Although the operating cost for high pressure may be
-------
-21-
higher, it more than makes up for the high cost of compressors
needed with low pressure gasifiers.
However, in order to have good efficiency the solid content
of the slurry feed must be high, 50-60 percent. When lignite is
slurried with water, the highest untreated solid concentration is
about 43 percent because lignite naturally contains up to 35
percent moisture. If the lignite is pretreated, the moisture
content can be lowered to more efficient levels.[9] The drawback
is that pretreatment is an added cost to production (although not
too large).
Another disadvantage of the Texaco process is that it
requires more oxygen than most of the other processes.
Although the Texaco gasifier has . not yet been used on a
commercial scale it has been extensively tested at a pilot plant
in Montebello, California[l] and at three demonstration plants:
the Ruhrchemie/Ruhrkohle plant in Oberhausen, West Germany; 2)
Tennessee Valley Authority's ammonia-from-coal plant in Muscle
Shoals, Alabama; and 3) an air blown gasification plant at a
chemical facility in the USA.[25] Texaco appears to be the
leading second generation technology and is being planned for two
projects already underway: Tennessee Eastman's project in
Kingsport, Tennessee to produce acetic anhydride and other
chemicals from methanol made from coal, and Southern California's
Cool-Water power generation station in Daggett, California.[16]
The Shell-Koppers gasifier is very similar to the Texaco
gasifier in that it can also use any coal and produces very little
byproduct. However, it is likely that the process will not be
commercialized for a couple of years since only limited data is
available on a 150 ton per day demonstration plant.[20]
One of the gasifiers that was used in the design studies to
be reviewed later was a Foster-Wheeler entrained bed gasifier.
This gasifier unit consists of two stages. In the second, which
is an entrained gasifier operating at 300-400 psig and 1700°F,
transport gas from stage one and pulverized raw coal are
introduced yielding slag and the product gas. The char which is
removed from the product gas is then sent with steam and oxygen to
the first stage producing the transport gas which is recycled to
the second stage.[1] As of 1977 the Foster-Wheeler gasifier was
in the early stages of pilot plant development.[1]
The gasifier used by Badger was a version of an entrained bed
gasifier.[4] According to Badger this gasifier is operated in an
oxygen-blown mode with a molten slag-bath at the bottom. The
gasifier has a total of 14 feed nozzles; 6 for coal and lime, 6
for oxygen and steam, and 2 for recycled char. The nozzles, which
are distributed around the periphery of the vessel, fire
tangentially and at a 45 degree angle toward the surface of the
-------
-22-
slag to make it rotate. Dense-phase pulverized coal and lime are
pneumatically fed into the lower section of the gasifier which
operates at 500 psig. The lime is a fluxing agent which is added
to obtain a slag viscosity of 10 poise. The oxygen and
superheated steam are added as gasifying agents. The coal, which
is partially pyrolized in the reaction, is gasified at 3000°F.
The advantages that are claimed in the literature for this
gasifier are that it can handle any type of coal and that the raw
gas is free of tar and high boiling hydrocarbons. When Badger
compared dry and wet (slurry) feeding they found that dry feeding
was economically superior to the slurry feed because the slurry
feed required 29 percent higher coal feed and a 73 percent higher
oxygen feed for a given synthesis gas, and therefore methanol,
rate. Badger also found that a steam-oxygen gasification medium
produces the highest thermal efficiencies ([4] pg. 64,65).
According to Badger "this single shaft high pressure
slag-bath gasifier is based on published information for
entrainment and other types of gasifiers and for the Rummel/Otto
Gasifier which is proven at atmospheric pressure. It is a new
concept and further development work may be necessary. Similar
gasification principles have been studied and pilot plant tests
have been conducted at lower pressures. Mechanical problems are
recognized and are believed to be solvable".[4]
Until recently, industry has been very sluggish in its
progress to reimplement coal gasifiers in the U.S. However, the
increasing cost of natural gas has sparked a new interest in coal
gasification and the majority of the coal or shale-based synthetic
fuel projects currently being planned use coal gasification.[23]
Table 6 lists some of the current projects which are now planned
or proposed.
One example is the previously-mentioned Cool Water
combined-cycle power-generation demonstration plant, to be located
in Daggett, California. It will gasify 1000 tons per day of coal
to produce 100 MW of electricity. The facility, which will use
the "proven" Texaco Coal Gasification Process[25], is currently
under construction and initial production is estimated for
1984.[27]
V. Synthesis Technology
The purpose of this section is to review available indirect
liquefaction processes with the emphasis being placed on
commercial feasibility, process description (reactor
configuration, operating conditions, etc.), product quality, and a
comparison of technological advantages and limitations. The
processes that have been reviewed include seven methanol synthesis
technologies (ICI, Lurgi, Haldor Topsoe, Mitsubishi Gas Chemical,
Vulcan-Cincinnati, Wentworth Brothers' and Chem Systems) and two
gasoline/petroleum synthesis technologies (Mobil's
Methanol-to-Gasoline and Fischer-Tropsch).
-------
-23-
Project Name
Table 6
Coal to Methanol Projects
Plant Size (Barrels Construction
Methanol/day) Date
1. Great Plains Coal 125
Gasification Project
Mercer County, ND
2. Coal-to-Methanol-to 4,200
Acetic Anhydride
Tennessee Eastman
Kingsport, TN
3. *Beluga Methanol 54,000
Project, Granite
Point, AK
4. Grants Project
**(ETCO), Grants, NM
5. Mapco Synfuels
Carmi, IL
6. Peat-to-Methanol
**(ETCO), Creswell, NC
7. Keystone Project
Cambria and Somer-
set Counties, PA
8. Dunn Nokota 85,000
Lignite-to-Methanol
Dunn County, ND
9. Chokecherry 3,608
**(ETCO), Moffat
County, CO
10. North Alabama Coal 25,000
Gasification Project
Murphy Hill, AL
11. New England Energy 18,000
Park Project
Fall River, MA
July 1980
1980
On Stream
Date
1984
1983
3,608
18,000
3,714
13,300
1982
1982
1982
1984
1985
1982
1982
1983
1989
1984-1985
1987
1984
1987
1989
1984-1985
1986
1988
* Feedstock is 60 percent natural gas, 40 percent coal
** Energy Transition Corporation (ETCO)
Sources: [23,27]
-------
-24-
The results of this section are briefly summarized in Table
7. Of the seven methanol synthesis processes that were examined,
the ICI, Lurgi, Haldor Topsoe, Mitsubishi Gas Chemical and
Vulcan-Cincinnati technologies have several commercial scale
processes in operation today. The Wentworth Brothers' methyl fuel
process is adapted from proven technologies and may be close to
commercialization.[14] The Chem Systems process is not
commercially feasible at this time since it is only at the pilot
plant stage. The latest report on the Mobil MTG process[28, 3/80]
was that the 4 barrel per day (BPD) pilot plant was the biggest
operating unit to date, but that plans for 100 BPD and 13,000 BPD
plants were proceeding. Mobil states that MTG is ready for
commercialization's], but at this time the MTG process is not
commercially proven. The Fischer-Tropsch process, which has been
operating for 25 years in South Africa, is unquestionably proven.
From a process point of view the high pressure methanol
synthesis technologies (Vulcan-Cincinnati and Wentworth Bros.) are
better suited for large scale production plants whereas the low
pressure methanol synthesis technologies (ICI, Lurgi, Haldor
Topsoe, Mitsubishi and Chem Systems, which operate between 30 and
130 atm) can be used with any size plant.[14] This is because the
high pressure plants have high throughputs which tend to
compensate for the higher cost for compression, especially for
very large plants. Although individual efficiencies have not been
reported for the methanol synthesis processes it is probable that
many of the technologies have comparable efficiencies since they
are highly developed- and very competitive. Two big factors that
affect efficiency are the extent of heat recovery and the percent
conversion of carbon monoxide and hydrogen to methanol per pass in
the converter. Of the methanol technologies listed most have a
conversion per pass of about 5 percent (e.g., ICI) while the Chem
Systems process claims up to 20 percent.[14] Concerning heat
recovery Lurgi claims to be more efficient than ICI because it
uses a heat exchanger type reactor versus the quench type used by
ICI.[1] Since Chem Systems uses a liquid phase process it should
get an even higher recovery of heat than Lurgi. Over the past ten
years ICI and (more recently) Lurgi have dominated the methanol
synthesis market.[1,29] Since these two processes are so
competitive it would seem logical that their economics would be
the same, and compared to other commercial processes, be
comparable if not less, expensive.
Parsons has stated that the Chem Systems process shows a
slightly higher thermal efficiency and slightly lower capital cost
than Lurgi and ICI synthesis; however, they believe room for
improvement over these synthesis units is small.[1] A comparison
of the Wentworth Brothers' process (using available published
information) with other processes did not show any inherent
economic advantage for the WBI process.[14]
-------
-25-
Vendor Catalyst
ICI* Cu/Zn/Al
Lurgi*
Topsoe'
Supported
Cu
Cu/Zn/Cr
Vulcan- Zn/Cr
Cincinnati
Mitsu-
bishi
Gas Chem-
ical*
Wentworth
Bros.*-
Chem
Systems
Cu/Zn/Cr
Multi-
Catalyst
Cu/Zn
Table 7
Methanol Synthesis Processes
Pressure Temperature
(atm) (°C) Reactor Type
30-50
34-68
215-250
Mobil MTG
Fischer-
Tropsch*
Zeolite
Cobalt or
Iron
2.7
0-25
330-400
200-325
50-100 220-290 Single fixed-
bed
235-280 Tube in shell
50-100 220-350 Radial flow
300-350 300-400 Multiple bed
50-130 240-310
up to 400 200-400
Liquid en-
trained and
liquid fluid-
ized
Fixed or fluid
Cooling
Multiple gas
quench
Steam genera-
tion
Boiler-feed-
water heating
Cold-shot
quench, plus
external gas
cooling
Multiple gas
quench
Recirculated
inert hydro-
carbon liquid
Fixed and fluid Steam genera-
with cooling ation
tubes
* Proven on a commercial scale.
Source: [13]
-------
-26-
A. ICI Low-Pressure Methanol Synthesis
Status; The ICI low pressure (50-100 atm) methanol synthesis
process is commercially proven worldwide.[6,13,14,17,30,31]
Process; Feed gas consisting of an approximately 2 to 1
ratio of hydrogen to carbon monoxide is fed into the synthesis
loop. The methanol conversion is highest when the hydrogen to
carbon monoxide ratio is 2 to 1 and the carbon monoxide to carbon
dioxide ratio is as high as possible.
The first part of the synthesis consists of desulfurizing the
feed gas when necessary to prevent the highly sensitive copper-
based catalyst from being poisoned. This is accomplished by
passing the gas through sulfur guard beds, which are typically
made of zinc oxide (or, less commonly, activated carbon) to
achieve sulfur levels below 1 ppmv.
The feed gas is then compressed to the recycle loop pressure,
mixed with the recycle gas and then compressed to reactor pressure
as it enters the methanol converter. The converter is a pressure
vessel containing a bed of catalyst.
The temperature of the bed is controlled by the extent of the
exothermic reaction and the quenching of the reaction by cold feed
gas. The pressure range is 50 to 100 atm while the temperature
must be kept below 300°C (210 to 300°C) since the catalyst becomes
deactivated at higher temperatures. The exit gas is passed
through heat recovery units for initial cooling and then sent to
the methanol separation unit where a crude methanol product is
produced (95 percent methanol by weight). Conversion of CO to
methanol per pass is about 5 percent.[14]
Catalyst life at pressures of 50 to 60 atm is 4 years while
maximum catalyst life at 100 atm is 2.5 years (average is 1-2
years).
Advantages/Disadvantages; Compared with high-pressure
processes, the ICI process is more adaptable to both large and
small plants (55 to 2750 TPD) whereas high pressure processes are
limited to large plants (1,000 to 2,500 TPD). Compression costs
are lower because of reduced pressure.
The disadvantages are that the high-pressure processes allow
a higher throughput of gas for the same size reactor and that the
catalyst cost for low pressure technology is five times as
expensive (50 cents/ton vs. 10 cents/ton).[14]
B. Lurgi Low-Pressure Methanol Synthesis
Status; The Lurgi low-pressure synthesis process (30-50 atm)
is commercially proven.[13,14,17,29,31]
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-27-
Process; The Lurgi methanol synthesis process uses a
shell-and-tube reactor. The copper-zinc catalyst is packed in
vertical tubes contained within a reactor shell which is filled
with boiling water. The exothermic heat of reaction is removed by
the generation of steam, thereby controlling the temperature of
the reactor.
The hydrogen to carbon monoxide ratio of the feed gas is
normally between 2 and 3, whereas the ratio of (hydrogen minus
carbon dioxide) to (carbon monoxide plus carbon dioxide) is held
around 2.2. After desulfurization the feed gas is compressed,
combined with recycle gas and preheated before being fed into the
reactor at one specific location. The Lurgi reactor has an
operating range of 30 to 100 atm and 200 to 300°C but is typically
operated at 70 atm and 260 to 270°C. The exit gas contains about
4-6 percent methanol and is sent to condensors to recover the
crude methanol product which is generally sent on for
purification. [14]
Advantages/Disadvantages; Like other low-pressure processes
the Lurgi process has an economic advantage over high-pressure
processes due to decreased compression costs at lower pressure.
The reactor design also permits direct recovery of the exothermic
heat of reaction by steam generation rather than a partial quench
of the reaction to control heat build up.[14]
According to Lurgi, a natural gas to methanol plant using
their synthesis technology is: more efficient[17] and consumes. 3-5
percent less natural gas per ton of pure methanol than competing
technology. They estimate the annual savings for a 2,500 stpd
plant is $4.2 - 5.0 million (U.S. dollars).[29]
When Badger compared the low pressure (1500 and 1400 psig)
methanol synthesis processes employing quench type converters and
licensed by Imperial Chemical Industries and Mitsubishi Gas
Chemicals Corporation with that of Lurgi's tubular type low
pressure (750 and 1200 psig) process they selected Lurgi's process
for two reasons:
- Lower investment and operating costs for syn-gas compression
- Maximizes medium pressure steam production; thus minimizing
overall utility costs and pay out time.
C. Haldor Topsoe Methanol Synthesis
Status; The Haldor Topsoe Methanol Synthesis process is
commercially proven.[14,17,31]
-------
-28-
Process; The Haldor Topsoe process is similar to other
low- and intermediate-pressure methanol synthesis processes. The
synthesis utilizes a copper-zinc-chromium catalyst in two or three
radial flow converters operated in series. After being
desulfurized (20 ppbv sulfur level) with zinc oxide guard beds,
the feed gas is mixed with recycle gas and passed through the
reactors flowing radially outward through each catalyst bed.
Operating pressure and temperature ranges are, respectively, 48 to
144 atm and 220 to 350°C. The exothermic heat from each reactor
is recovered by heating boiler feedwater with the hot exit gases.
The gases are then condensed and sent to a separator where crude
methanol is separated from uncondensed gases and later sent to
product upgrading.[14]
Advantages/Disadvantages; The Haldor Topsoe process can
operate at intermediate pressures for higher throughputs. It can
also operate at higher temperatures which increases the activity
of the catalyst provided it can retain its active sites and
structural integrity.
D. Mitsubishi Gas Chemical Methanol Synthesis
Status; The Mitsubishi Gas Chemical (MGC) methanol synthesis
process is commercially proven.[14,31]
Process; The MGC process appears to be very similar to Id's
intermediate-pressure process since both designs use a quench
converter with a ternary-- copper-based catalyst operated at. low
temperature and intermediate pressure (240-310°C and 50-130 atm).
The feed gas is split into a feed stream which is heated and fed
into the converter, and a quench stream which is injected at
several bed levels to control the buildup of the exothermic heat
of reaction. After being used to preheat the feed gas the exit
gas is condensed and sent on to distillation for a product purity
in excess of 99 percent methanol by weight. Part of the recycle
gas is used for fuel. The catalyst has an expected life of just
over 1 year since it is very sensitive to sulfur.
Advantages/Disadvantages; The MGC intermediate pressure
process has the advantage of accomodating moderately higher
throughputs than lower pressure processes while keeping
compression costs down as compared to high pressure processes.
The ICI catalyst appears to have a longer catalyst life (2
years for ICI vs. greater than 1 year for MGC). The MGC process
typically uses a higher hydrogen to carbon oxides ratio in the
feed gas than other processes (3.1 compared with 2.2 for other
processes) but this is because it has only been used when natural
gas is the feedstock.[14]
-------
-29-
E. Vulcan-Cincinnati High Pressure Methanol Synthesis
Status; The Vulcan-Cincinnati high-pressure process is
commercially proven. [14,17 ,31] However, the company had to stop
operation in 1973 when the Middle East war forced the cancellation
of a very large methanol plant in Saudi Arabia in which Vulcan had
heavily invested (see Wentworth Brothers' process below).
Process: The feed gas ratio for H2/(CO + 1.5 C02) should
be adjusted to a value of 2 for optimum conversion after
desulfurization. The feed gas is then compressed and fed to the
converter which is usually operated in the range of 340 to 400°C
and 200 - 300 atm. The converter operates adiabatically with
considerable temperature rise due to the exothermic heat of
reaction, which is controlled by quenching the reaction with cold
feed gas at several levels. After conversion the crude methanol
product is condensed for removal yielding a product containing up
to 97 wt. percent methanol (depending on feed gas composition).
There is also an option of producing up to 20 wt. percent of
higher alcohols by changing operating conditions which would be
helpful if used for blending with gasoline.
The catalyst, which is poisoned by t^S levels greater than
3 to 5 ppm, has a typical life of about 4 years and can be
regenerated. Conversion of CO to methanol per pass is
approximately 5 percent.
Advantages/Disadvantages ; The high pressure process is well
suited for large methanol synthesis trains due to the high
throughputs occuring. Catalyst costs are also less than the
low-pressure copper-based catalysts, are not as sensitive to
sulfur and can be regenerated. The process can produce a wider
range of fuel products (3 - 20 percent higher alcohols).
Some disadvantages are: 1) that the high-pressure process
has greater compression costs, 2) that the catalyst requires
higher temperatures and pressures because it is not as active as
the copper-based catalysts, and 3) that it may not be suited for
small methanol plants. [14]
F. Wentworth Brothers Methyl Fuel Process
Status; The term methyl fuel, copyrighted by
Vulcan-Cincinnati, represents the product of a methanol synthesis
process which is focused on producing methanol for fuel rather
than chemical uses. After Vulcan-Cincinnati stopped operation in
1973, the Wentworth Brothers and other engineers from Vulcan
formed a new corporation in May 1975. Based on Vulcan experience
and technology and relying on catalyst improvements and a reactor
design adapted from proven petroleum technology, Wentworth
Brothers, Inc. (WBI), is now marketing what they believe to be a
much improved process for the production of large quantities of
-------
-30-
fuel-grade methanol. No commercial plants are in operation, but
short-term tests of the catalyst at 300 TPD in a commercial
methanol train have reportedly verified the basic operating
parameters for the WBI methyl fuel process.[14]
Process; Details of the process operation, catalyst
formulation, and reactor configuration are considered proprietary
and are not available.
Advantages/Disadvantages; What is known is that the new
catalyst is reportedly more active and durable than conventional
low pressure catalysts although at the expense of the selectivity
for methanol. The more active catalyst allows operation of the
process at increased space velocities (throughput per reactor
volume), and at higher temperatures and pressures which maximize
fuel production per reactor train. The range of operating
conditions for the methyl fuel process includes pressures up to
4,000 psi (270 atm) and temperatures from 200 to 400°C. The
catalyst is claimed to be effective at C02 concentrations
ranging from 20 percent to essentially zero versus conventional
copper-based catalysts which require some C02« Oak Ridge
National Laboratory states that from available published
information it is not possible to ascertain whether the WBI Methyl
Fuel Process has any inherent economic advantage over conventional
copper-based methanol synthesis processes.[14]
G. Chem Systems Synthesis
Status; .As of August 1980, the Chem Systems process was
ready to move to the pilot plant stage;[32]
Process; The major difference between the Chem Systems
synthesis and the other synthesis processes is that an inert
hydrocarbon liquid is used as the medium for the catalyst instead
of a gaseous phase. This liquid phase allows high conversions of
carbon monoxide and hydrogen to methanol in addition to maximum
recovery of reaction heat.[32]
In the process synthesis gas containing carbon monoxide,
carbon dioxide and hydrogen is passed upward into the reactor
concurrent with the inert hydrocarbon liquid, which is recovered
in the separation plant and recycled back to reactor with the
unconverted synthesis gas.[1,32] The fuel grade methanol product
is 95-96 percent methanol by wt.[14]
Advantages/Disadvantages; Chem Systems claims that their
conversion to methanol per pass is about 4 times as great as other
processes (20 percent vs. 5 percent). However, Parsons believes
that this process only has slight cost advantages over the
existing processes.[1][14]
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-31-
Parsons states that a new catalyst formulation with superior
mechanical strength still needs to be developed to make the liquid
phase methanol synthesis viable.[1] Breakdown of catalyst,
inhibition of catalyst by fluid and insufficient solubility of the
synthesis gas in the fluid are other possible problem areas with
this design.[17]
H. Mobil Methanol-To-Gasoline (MTG) Process
Status; Mobil has conducted developmental studies of this
process in fixed- and fluid-bed bench-scale units with two
reactors being used in the fixed bed unit. The fixed bed reactor
achieved over 200 days of successful operation. The single
reactor fluid-bed unit has undergone two months of testing. Since
the fluid bed reactor had a number of advantages over the fixed
bed reactor, a 4-BPD fluid-bed pilot plant was designed, built and
operated under a follow-on DOE contract in 1976-78. Startup and
operation of this fluid—bed unit were reported to be very
successful. Plans are currently under way for a 100 BPD fluid bed
pilot plant sponsored by DOE, the Federal Republic of Germany,
German industrial participants, and Mobil.[14,22] Of interest
also is the reported news that, since November 1979, the
government of New Zealand has been pursuing the Mobil
methanol-to-gasoline process, with negotiations proceeding for a
13,000 BPD fixed-bed unit for installation almost
immediately.[22,28]
Process; The conversion of methanol to hydrocarbons and
water is a very exothermic reaction giving off 740 Btu/lb of
methanol. Heat removal is therefore the principal problem in
designing a reactor system.[7] For the fixed bed reactor the
problem is minimized by dividing the reaction into two steps and
using two reactors in series. In the first reactor crude methanol
is partially dehydrated to an equilibrium mixture of methanol,
dimethyl ether and water over a dehydration catalyst[20,28]
releasing about 20 percent of the reaction heat.[7] In the second
reactor the new shape-selective zeolite catalyst is used to
convert both methanol and dimethyl ether to a liquid hydrocarbon
product. This hydrocarbon liquid product is then sent to a
fractionation unit where a deethanizer sends the ethane rich
overhead product to the SNG train and the bottoms are sent to a
stabilizer. The overhead product of the stabilizer is composed of
isobutane and butene/propene which are sent to an alkylator to
produce more gasoline and commercial grade propane and butane.
The bottoms product of the stabilizer is a stabilized gasoline
which is mixed with the gasoline product from the alkylation unit
and sent to the gasoline blending unit to yield a high octane (93
research octane) gasoline.[7]
The inlet temperature of the second reactor is about 625°F.
The adiabatic fixed bed process operates at essentially 100
percent conversion of methanol to hydrocarbons and water until the
-------
-32-
catalyst deactivates by carbon formation to an activity level
where only partial conversion of methanol is achieved.[28] The
zeolite catalyst must be regenerated once every 14 days.[7]
In the fluid bed process one reactor is used, operating at
750°F and 40 psig. The hydrocarbon product is generally treated
in the same manner as the fixed bed product with the exception of
a few changes. A deethanizer absorber is used in place of the
high pressure deethanizer tower to provide a recycle stream to the
reaction for increased propane-plus yield. A rich oil tower is
also required.[7]
Advantages/Disadvantages; For the fluid-bed reactor the
methanol conversion is greater than 95 percent, producing about 44
percent hydrocarbons and 56 percent water. The pentane-plus
gasoline fraction of the hydrocarbons is about 60 percent. The
propene, butene, and isobutane produced are approximately the
right proportions for alkylation, bringing the total yield of 9
Ib. Reid Vapor Pressure gasoline (96 unleaded RON) up to 88
percent of the total hydrocarbon yield. The thermal efficiency
for the methanol conversion is quoted at 95 percent.[14,33]
One potential problem with the gasoline produced from both of
these processes is the presence of durene which boils in the
gasoline range but has a freezing point of 175°F. This could
cause engine problems since the durene could crystallize out in an
engine's carburetor. Durene is present in conventional gasoline
in. very small amounts but could be present in relatively large
amounts (3-6 percent) in Mobil MTG-gasoline. Durene levels of 5
percent in gasoline did cause some unsatisfactory engine
operations during tests but at 4 percent levels effects were
minimal. Since durene levels can be maintained to acceptable
levels by proper process controls and it could always be mixed
with conventional gasoline[34] the presence of durene may not pose
too much of a problem.
I. Fischer-Tropsch
Status; The Fischer-Tropsch process is proven technology
which has been producing liquid hydrocarbons at SASOL in
Sasolberg, South Africa since 1955.[7] This is the only
commercial scale coal liquefaction plant operating in the world
today. The current SASOL plant uses two reactor schemes, a
fixed-bed and a fluid bed. The current SASOL expansion to 50,000
BPD is based on a fluid-bed design.
Process; With the fluid bed reactor, purified synthesis gas
is compressed and charged into the reactor. After mixing with the
circulating hot iron catalyst, the reaction takes place as the
mixture flows up the reactor through tube bundles in which oil is
pumped for heat removal. At the top of the reactor, the mixture
enters a large vessel in which cyclones separate the iron and
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-33-
vapor. The hot oil is circulated to a steam generator where 200
psig steam is produced. The overhead vapor is condensed and the
vapor split into a recycle and purge stream, the latter being sent
to hydrocarbon recovery. The condensed liquid is split into a
cold recycle liquid and a light oil product. Based on a 2.13
molar feed ratio of H2SCO to the F-T synthesis units the final
product yield consists of 24 percent SNG, 54 percent 10 RVP
gasoline, 9.8 percent diesel fuel, 5.5 percent alcohols, 3.8
percent LPG and 2.8 percent heavy fuel oil (percentages on a HHV
BTU basis).[7]
The commercial catalysts include cobalt, for the fixed-bed
reactor, and iron for both the fixed- and fluid-bed reactors.
Operating conditions range from 200-325°C and from atmospheric
pressure to 25 atm depending on the desired products. Because all
of the reactions in the F-T process are exothermic, heat removal
is important. For the fixed bed reactor, designs include a heat
exchanger being cooled by boiling water or circulating oil. Fluid
bed reactors use internal tube bundles for reaction heat
removal.[7]
Advantages/Disadvantages; According to a study done by
Mobil[7] which compares the F-T process to the Mobil MTG process,
the MTG process has a number of advantages over F-T. For example,
the Mobil technology gives a . higher liquid product/SNG ratio
(energy basis), 47/53 vs. 35/65. The F-T route has 18 processing
steps compared with nine for MTG. The MTG gasoline had a higher
RON (93 vs. 91) and a lower olefin content (11 vs. 20 percent).
The economics show that . the MTG gasoline cost is moderately
cheaper than F-T gasoline ($.60 '"- $1.00 per gal. vs. $.70 - $1.35
per gal.). The overall efficiency for MTG was 62 percent vs. 58
percent for F-T.[7]
A recent article in Hydrocarbon Processing advocating F-T
stated that the 40-45 percent efficiency reported for the
operating F-T process is really depressed because the SNG produced
is reformed to make additional H2 and CO instead of being sold
as a product (SNG is not marketable in South Africa). If the
process was brought to the U.S. the author believes that SNG would
be a viable product as a pipeline quality methane gas, thus
increasing the F-T efficiency to 60 percent (close to 58 percent
which was previously stated by Mobil). In addition the article
points out that coal consumption for the F-T process is high only
because they use a low quality coal with 30 percent ash.[35]
VII. Comparison of Indirect Liquefaction Design Studies
As discussed earlier, there have been a number of studies
evaluating the indirect liquefaction of coal. This chapter is
divided into four main sections, the production of methanol from:
1) bituminous coals, 2), subbituminous coals, and 3) lignite and
4) the production of gasoline from methanol using the Mobil MTG
-------
-34-
technology; however, this fourth section also includes the
production of gasoline and other products via Fischer-Tropsch
technology.
Each section begins with a brief introduction and a
comparison of three aspects of the various studies: level of
engineering design, feedstock analysis, and material balances and
efficiencies. Then capital, operating and product costs for each
study will be presented on a consistent economic basis and then
compared to reconcile as many of the differences as possible.
A. Methanol from Bituminous Coal
There were five original studies available which investigated
the technical feasibility of producing methanol from bituminous
coals:
1. R.M. Parsons, Co., for EPRI, "Screening Evaluations:
Synthetic Liquid Fuel Manufacture,"[1]
2. C.F. Braun for EPRI, "Coal to Methanol Via New
Processes Under Development: An Engineering and Economic
Evaluation,"[2]
3. Dupont Co. for U.S. ERDA, "Economic Feasibility Study,
Fuel Grade Methanol from Coal for Office of Commercialization of
ERDA,"[3]
4. Badger Plants, Inc., "Conceptual Design of a Coal-to-
Methanol Commercial Plant,"[4] and,
5. Exxon Research and Engineering Co., "Production
Economics for Hydrogen, Ammonia, and Methanol During the 1980-2000
Period."[5]
The studies differ in depth of design and use different
assumptions with respect to key economic parameters; also plant
sizes vary widely. Since methanol-from-coal technology is well
developed and much of it is common to all of the studies, the
large cost differences between the studies should be reconcilable
by placing them on a consistent economic basis.
Depth of Design: The level of engineering detail of the five
studies varies. The most detailed studies where found to be those
performed by Badger and EPRI/C. F. Braun. These studies include
complete material balance information for a detailed flowsheet and
extensive design details, including drawings. The studies with
the next highest level of engineering design are the EPRI/Ralph M.
Parsons and Du Pont studies. These studies are screening
evaluations and their level of detail is not sufficient to allow
comparison with more detailed studies. The EPRI/Ralph M. Parsons
study includes evaluations of four gasification processes in
combination with Chem Systems Methanol synthesis; these four
processes are denoted as Cases 1, 2, 3 and 4 of the Parsons
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-35-
study. The study found to be based on the least detail is the
Exxon/Chem Systems study which consists of summary economic
information for two gasification processes, but provides no
details for material and energy balances. These two processes are
denoted as Cases 1 and 2 of the Exxon study.
Ultimate Analysis for Bituminous Coals; Ultimate analyses of
the bituminous coals used in the various studies are presented in
Table 8. Also listed in this table are the heating values of the
coals. The Exxon/Chem Systems study did not report an analysis
for their coal; however, they used an Illinois high sulfur coal
which is probably similar to the Illinois No. 6 bituminous coals
reported in the two EPRI studies. The Du Pont coal is also
similar to the coals presented in the EPRI studies. However, the
coal considered in the Badger report is a low sulfur coal which
would meet the sulfur dioxide standard for large power
plants. [12] It is unlikely that a coal of this quality would be
used to produce methanol.
Material Balance and Efficiencies; Feedstock and product
rates for each study are presented in Table 9. Methanol and coal
are presented on both a short ton per calendar day (tpd) basis and
on an energy basis. Other products include fuel gas in Cases 1
and 2 of the Parsons study; the Badger study includes chemical
grade methanol along with the fuel grade. All rates are based on
50,000 FOEB/CD of total products.
Table 9 shows that the process efficiencies for the studies
vary from 49.3 percent for the Koppers-Totzek/ICI case prepared by
Exxon/Chem Systems to 58.2 percent for the Texaco/Chem System case
prepared by EPRI/Parsons. An investigation of these efficiencies
shows that the two Koppers-Totzek cases are amongst those with the
lowest efficiencies. This is not surprising since the Koppers-
Totzek gasifier is the only first generation gasifier listed in
this table, and since the gasifier operates at near atmospheric
pressure so that the product synthesis gas must be compressed,
thus resulting in an efficiency penalty. When comparing the three
Texaco gasifier studies it can be seen that there is a substantial
difference (7.5 percent) between the highest and lowest
efficiency. The C.F. Braun study reports an efficiency of 55.7
percent whereas the DuPont study reports an efficiency of 50.6
percent. A key variable in the Texaco process which affects the
efficiency is the coal/water slurry concentration. The greater
the coal concentration in the slurry, the higher the efficiency of
the Texaco process.[9] The C.F. Braun case utilized a 59 percent
coal concentration in the slurry whereas the DuPont study utilized
a 54 percent coal concentration. This difference will account for
part of the efficiency discrepancy. The Parsons/EPRI study did
not report any processing information for the Texaco gasifier and
therefore the high efficiency reported for this process could not
be investigated. The remainder of the processes reported in Table
9 utilize other advanced technology gasifiers, and therefore are
expected to have higher efficiencies than first generation
gasifiers, e.g., Koppers-Totzek.
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-36-
Table 8
Ultimate Analysis of Bituminous Coal Feedstocks
Parsons
Exxon[5]
Study: Cases 1,2,3,4[1]
and 2
Coal Type
High
HHV
(Btu/lb.)
LHV
(Btu/lb.)
Ultimate Analysis
Wt % Dry Coal
C
H
0
N
S
Ash
Total
Wt% Moisture
Bituminous
111. No. 6
12,235
(wet)
11,709
»
69.5
5.3
10.0
1.3
3.9
10.0
100.0
4.2
C. F. Braun[2]
Bituminous
111. No. 6
12,150
(dry)
—
68.25
5.00
11.23
0.81
3.88
10.83
100.0
10
DuPont[3]
Badger [4]
Bituminous S. Appal.
12,113
(dry)
10,874
66.89
4.47
8.41
1.28
4.47
14.48
100.0
6.38
12,840
(dry)
-
73.8
4.8
6.4
1.6
1.1
12.3
100.0
2.4
Case 1
111.
Sulfur
11,390
(wet)
-
(As Recieved)
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Table 9
Methanol from Bituminous Coal:
Feedstock and Product Rates
(Normalized to 50,000 FOEB/CD Product)
Study:
Mass Basis
Feedstocks
Coal, tpd
Products
Methanol, tpd
HHV Btu/lb
Parsons [1]
Case 1 Case 2 Case 3
21,722 21,150 23,006
14,500 15,132 15,349
9,610 9,610 9,610
Case 4 C.F. Braun[2]
20,714 21,795
15,349 15,172
9,610 9,722
DuPont[3]
25,685
16,223
9,092
Badger [4 J
20,048
14,570
9,407
Exxon[5]
Case 1 Case 2
24,447 26
15,227 15
9,687 9
,188
,227
,687
Fuel Gas,
mscf/CD
Chemical Grade -
Methanol, tpd
Energy Basis (HHV), inBtu/CD
Feedstocks
Coal 531,531 517,544 562,977 506,873
Electricity -
(energy equiv-
alent)
Products
iQ75
529,623
582,543 514,834 556,902 596,565
1810 1812
Methanol 278,682 290,845 295,000 295,000
(Fuel Grade)
Fuel Gas 16,317 4,155
Methanol -
(Chemical Grade)
295,000
Thermal Eff'cy, % 55.5
57
52.4*
58.2
55.7
295,000 274,121 295,000 295,000
20,879
50.64 57.3 52.8 49.3
i
w
* 95% conversion assumed for gasifier as opposed to 100 percent for the other cases.
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-38-
Economics; In this sub-section the capital and operating
costs for each study will be presented in order to obtain the
desired product costs. These costs have been placed on a
consistent economic basis as reported in a previous report.[11]
Table 10 presents all of the investment costs (normalized to
50,000 FOEB/CD) broken down into individual process unit costs.
This table shows that the total instantaneous investments range
from $1.62 billion for the Shell-Koppers/ICI case of the Exxon
study to $2.56 billion for the Koppers-Totzek case of the Parsons
study, which represents a $940 million instantaneous investment
difference.
Operating costs are presented in Table 11. Some of the
studies did not itemize the operating costs which makes it
difficult to compare these costs between each study. The net
annual operating costs range from about $315 - 480 million. Most
of the operating cost estimates lie in the $340 - 440 million per
year range.
Tables 12 and 13 present economic summaries of methanol costs
for capital charge rates of 11.5 and 30 percent. For the lower
capital charge rates product costs vary from $5.30 for the
Parsons/BGC Lurgi study to $7.23/mBtu for the
Parsons/Kopper-Totzek study. For the higher capital charge rate
product costs vary from $8.74 to $12.42/mBtu.
Now some of the capital and operating cost differences will
be reconciled. First, one would expect the Koppers-Totzek cases
to have a high capital cost because it is a low pressure process;
this results in higher compression and gasification costs than for
high pressure gasification technologies. Also, since the
Koppers-Totzek gasifier is a first generation gasifier, its
capital cost is expected to be higher than the more advanced
technology cases. The instantaneous plant investment for the
Parsons/Koppers-Totzek study is $2.56 billion while that of the
Exxon study is $1.9 billion, which represents about a $650 billion
investment difference. The investment estimate of the Exxon study
is similar to those of the advanced technology gasification cases
which seems unreasonable. The only difference between the two
Koppers-Totzek cases is the methanol synthesis technology used.
Since there is very little difference expected in the capital
costs of the Chem Systems and ICI synthesis units, it is
surprising to see such a large capital cost difference between
these two cases, and the difference may be attributed to lack of
design detail for the Exxon study which was discussed earlier.
Therefore, it is believed that the Parsons study is more
representative for the Kopper-Totzek case.
Three studies investigated the use of Texaco gasification
technology: Parsons, C.F. Braun, and DuPont. The respective
plant investment costs for these studies are $2.05, $1.75 and
-------
Table 10
Methanol From Bituminous Coals: Capital Costs Summary*
(Millions of First Quarter 1981 Dollars)
Study;
Parsons [1]
Case 1
Technology:
Gasification/ Foster
Synthesis Chem.
Coal Preparation
Tar and Phenol
Recovery
Gasification
Shift Conversion
Acid Gas Removal
Sulfur Removal
Synthesis Gas
Compression
Methanol Synthesis
Cryogenic Recovery
Sulfur Recovery/
Methanol Drying
Oxygen Production
Steam and Power
Generation
Subtotal
General Facilities
and Offsites
Contingency
Contractor's Fee
Total Instantaneous
Plant Investment
Case 2
BGC Lurgi/
Wheeler/
Systems
83
-
186
42
281
14
71
170
29
28
348
68
1320
198
228
44
1790
Chem.
Systems
40
99
145
56
241
14
76 '
175
30
29
263
80
1248
187
215
42
1692
Case 3
Koppers
Totzek/
Chem. Systems
87
-
566
51
266
15
125
177
-
29
374
197
1887
283
325
63
2558
Case 4
Texaco/
Chem.
Systems
73
-
385
51
283
15
-'
177
-
29
430
106
1549
232
217
52
2050
C.F. Braun[2] DuPont[3] Badger [4]
Texaco/
Chem.
Systems
110
-
167
97
187
-
-
164
-
26
300
123
1174
306
222
43
1745
Texaxo/
ICI
112
-
205
75
175
-
-
190
-
-
373
302
1432
209
246
48
1935
Slag Bath/
Lurgi
37
33
187
66
193
16
33
197
16
24
313
90
1205
453
249
42
1949
Exxon [5 J
Case 1
Shell-
Koppers/
ICI
_
-
-
-
-
-
-
-
-
-
-
-
775
597
206
40
1618
Case 2
Kopper
Totzek/
ICI
_
-
-
-
-
-
-
- i
(_^
— VC
-
-
-
951
666
243
47
1907
Investment costs are based on 50,000 FOEB/CD of product.
-------
Table 11
Methanol From Bituminous Coals: Operating Cost Summary
(Millions of First Quarter 1981 Dollars Per Year)
Study:
Parsons [1]
Case 1
Technology:
Gasification/ Foster
Synthesis Chem.
Raw Materials:
111. No. 6 Coal
Catalysts and
Chemicals
Utilities:
Power
Process Water
Stack Gas Clean-up
Labor and Related:
Labor
Supervision
Plant Overhead
Capital Related:
Maintenance
General Plant
Overhead
Insurance and
Property Tax
Interest on Working
Capital
Other Operating
Costs
Gross Annual
Operating Cost
By-Product Credit
Net Annual
Wheeler/
Systems
218
10.6
7.5
124
360
(14.7)
345
Case 2
BGC Lurgi/
Chem.
Systems
212
10.6
7.1
120
350
(13.9)
336
Case 3
Koppers
Totzek/
Chem. Systems
231
12.1
10.7
181
435
(12)
423
Case 4
Texaco/
Chem.
Systems
208
10.6
8.8
149
387
(13.1)
374
C.F. Braun[2]
Texaco/
Chem.
Systems
219
20.7
36.3
10.9
40.1
23
47.4
7.3
1.2
406
(13.2)
393
DuPont[3]
Texaxo/
ICI
258
15
12
97
41
8.1
17
458
(19.5)
429
Badger [4]
Slag Bath/
Lurgi
199
3.4
23
29.9
1.5
2.9
7.2
52
319
(3.8)
315
Exxon[5]
Case 1
Shell-
Koppers/
ICI
245
4.1
6.8
19.4
8.7
1.3
4.1
1.7
63.4
41.2
42.1
6.8
445
(15.0)
430
Case 2
Kopper
Totzek/
ICI
263
4.1
6.8
20.6 i
12. Ig
1.3
4.1
1.7
74.5
48.4
47.5
8.0
492
(16.0)
476
Operating Cost
-------
Table 12
Stud}
Total Instantaneous
Investment
Total Adjusted
Capital Investment
Economic Summary of Methanol from Bituminous Coal, CCR = 11.5%
(Millions of First Quarter 1981 Dollars)
Parsons[1]
Exxon[5j
Case 1 Case 2 Case 3 Case 4 C.F. Braun[2] DuPont[3] Badger[4] Case 1 Case 2
1790 1692 2558 2050 1745 1935 1949 1618
2030
1919
2901
2325
1979
2194
2210
1835
1907
2163
Start-up Costs
Pre-paid Royalties
Total Capital
Investment
Working Capital
Total Capital
Requirement
Annual Capital Charge
Annual Operating Cost
Total Annual Charge
Product Cost
$/FOEB of Product
$/mBtu of Product
126
8
2164
125
2289
249
345
594
32.55
5.52
119
9
2047
118
2165
235
336
571
31.29
5.30
179
13
3093
179
3106
356
423
779
42.68
7.23
147
10
2482
147
2629
302
374
676
37.00
6.28
122
10
2111
122
2233
243
393
636
34.85
5.90
136
10
2340
135
2475
269
429
698
38.25
6.48
137
10
2357
137
2494
287
315
602
32.99
5.59
114
10
1959
113
2072
225
430
655
35.89
6.08
134
10
2207
133
2440
265
476
741
40.60
6.88
-------
Table 13
Economic Summary of Methanol from Bituminous Coal, CCR = 30%
(Millions of First Quarter 1981 Dollars)
Parsons[l]
Stud}
Total Instantaneous
Investment
Total Adjusted
Capital Investment
Exxon[5J
Case 1 Case 2 Case 3 Case 4 C.F. Braun[2] DuPont[3]
1790
1998
1692
1888
2558
2855
2050
2287
1745
1947
1935
2159
Badger[4] Case 1 Case 2
1949 1618 1907
2176
1806
2128
Start-up Costs
Pre-paid Royalties
Total Capital
Investment
Working Capital
Total Capital
Requirement
Annual Capital Charge
Annual Operating Cost
Total Annual Charge
Product Cost
$/FOEB of Product*
j/mBtu of Product
126
8
2132
125
2257
640
345
985
53.97
9.15
119
9
2016
118
2134
605
336
941
51.56
8.74
179
13
3047
179
3226
914
423
1337
73.26
12.42
147
10
2444
147
2591
777
374
1151
63.07
10.69
122
10
2079
122
2201
624
393
1017
55.73
9.44
136
10
2305
135
2440
692
429
1121
61.42
10.41
137
10
2323
135
2460
738
315
1053
57.70
9.78
114
10
1930
113
2043
579
430
1009
55.29
9.37
134
10
2272
133
2405
682
476
1158
63.29
10.73
i
-e-
ro
1
One FOEB =5.9 mBtu
-------
-43-
$1.94 billion. The total gasification costs for the C.F. Braun
and DuPont studies are almost identical (about $565 million)
whereas the total gasification cost from the Parsons study is
about $240 million higher. The main difference in the C.F. Braun
and DuPont studies lie in differences in the cost of the
non-gasification equipment (oxygen production, offsites, etc.).
It is difficult to determine which of the studies is most
representative; based on level of design, the C.F. Braun study
should be considered most representative. However, to be somewhat
conservative in comparing methanol costs to direct liquefaction
cost both the C.F. Braun and the DuPont studies will be chosen and
a range of costs used. The product costs for these Texaco studies
lie in about the same range as those for the other advanced
technology cases, which seems reasonable.
In an effort to determine the representativeness of the
Badger costs, DOE commissioned Oak Ridge National Laboratory in
June, 1978 to make an independent assessment of the Badger
report.[36] ORNL reported that Badger's design is based on
equipment sizes well beyond the present state-of-the-art in order
to take advantage of the projected economies of .scale. Therefore,
ORNL believes that the Badger design is more representative of an
Nth plant design rather than a first plant design. For a first
plant design the Badger capital and operating costs appeared to be
unreasonable to ORNL. The operating costs for the Badger study
listed in Table 11 are lower than any of the others listed in the
table and perhaps are suspect. However, the Badger study is based
on an advanced technology "slag bath" gasifier, and the capital
cost based on this design is a bit higher than those of the other
studies based on advanced technology gasifiers listed in Table
10. It is expected that this design would have a lower capital
cost than those based on technology commercially available or near
commercially available. The degree to which this Badger estimate
may represent Nth plant designs is uncertain.
Even though the Texaco gasifiation process is advanced, its
costs will be presented separately from the other advanced
gasification costs. Based on the above discussion, the DuPont and
C.F. Braun studies will be used to represent the range of product
costs for methanol from Texaco gasification technology, and the
Parsons study for Kopper-Totzek technology. The remainder of the
studies represent methanol costs for the other advanced
gasification technologies. These costs are as follows:
$/mBtu
Capital Charge Rate
11.5% 30%
Koppers-Totzek 7.23 12.42
Texaco 5.90-6.48 9.44-10.41
Advanced Technology 5.30-6.08 8.74-9.78
-------
-44-
B. Methanol from Subbituminous Coal
There are two original studies available which investigated
the technical and economic feasibility of producing methanol from
subbituminous coals. These studies are:
1. "Methanol from Coal: An Adaptation from the Past,"
Bailey, Davy McKee Corp., 1979.[6]
2. "Research Guidance Studies to Assess Gasoline from Coal
by Methanol-to-Gasoline and Sasol-Type Fischer-Tropsch
Technologies, Schreiner, Mobil Research and Development Corp.,
August, 1978.[7]
The Davy McKee study investigated the use of a Davy McKee
fluidized bed gasifier, which is a modified Winkler gasifier which
has not been demonstrated on a commercial scale. ICI technology
was used for methanol synthesis. The Mobil study utilized BGC
Lurgi technology for gasification and Lurgi methanol synthesis.
Depth of Design; Neither study seems to have been based on a
high level of engineering design. However, since the Davy McKee
study utilized modified Winkler/ICI technology and since Davy
McKee has designed and built commercial processes using Winkler
and ICI technology, their study is probably based on processing
and cost correlations associated with plants they have
constructed. For the Mobil study, process information was based
on either published or licensor data, while investment estimates
were principally derived from in-house data. For offsite units
vendor quotes were used where obtainable.
Ultimate Analyses of Subbituminous Coal Feedstocks; Ultimate
analyses for the subbituminous coals are presented in Table 14.
The higher and lower heating values are also shown. Both coals
are from Wyoming and have very similar compositions.
Material Balance and Efficiencies; Feedstock and product
rates for both studies are presented in Table 15. Methanol and
coal are presented on both a short ton per calendar day (tpd) and
an energy basis. The Davy McKee study produces 100 percent
methanol while the Mobil study produces about 48 percent methanol,
50 percent SNG, and 2 percent naptha. Sulfur, ammonia and coal
fines are produced as by-products from the Mobil study. Coal
fines are also produced since the Lurgi gasifier cannot process
them. By-products were not reported for the Davy McKee study, but
are produced; therefore, for economic purposes the sulfur and
ammonia yields from the other study were assumed for it.
Product qualities for the Davy McKee study are not reported,
but product qualities for the Mobil case are presented in Table
16. The methanol is 99.66 percent pure, but it is still
considered to be of fuel grade quality. The SNG is about 96
percent methane. The naptha product has an octane ((R+M)/2) of
88.8, and is a suitable gasoline blending stock.
-------
-45-
Table 14
Ultimate Analysis of Subbituminous Coal Feedstocks
Study;
Coal Type;
HHV, Dry, Btu/lb
LHV, Dry, Btu/lb
Ultimate Analysis of Dry Coal, Wt %
C
H
0
N
S
Ash
Total
Wt % Moisture (as recieved)
Davy McKee[6]
Wyoming
11,818
10,963
Mobil[7]
Wyoming
11,818
10,963
69.2
4.7
17.9
0.7
0.4
7.1
100
28
70.8
4.9
18.3
0.7
0.4
5.1
100
28
-------
-46-
Table 15
Methanol from Subbituminous Coal:
Feedstock and Product Rates
(50,000 FOEB/CD of Product)
Davy McKee[6] Mobil[7]
Mass Basis
Feedstock
Dry Coal, tpd 26,820 19,063
Product
Methanol, tpd 15,227 7,270
Synthetic Natural
Gas, mscf/CD - 150
Naptha, bbl/CD - 1,351
By-products, tpd
Sulfur . - ;63
Ammonia - 103
Coal Fines - 1,501
Energy Basis, mBtu/CD, (HHV)
Feedstocks
Coal 639,918 450,563
Electricity 3,448 1,198
Products
Fuel Grade Methanol 295,000 141,388
Synthetic Natural Gas - 146,588
Naptha - 7,024
Coal Fines
Thermal Efficiency, % 45.9 65.3
-------
-47-
Table 16
Methanol from Subbituminous Coal; Product Qualities
Davy McKee[6]
The quality of the fuel grade methanol was not reported in
the Davy McKee study.
Mobil[7]
1.
HHV:
LHV:
2.
SNG
Composition
Hydrogen
Methane
Carbon Dioxide
Inerts (N£ and Ar)
975 Btu/scf
878 Btu/scf
Methanol
Light Boiling Compounds
Heavy Boiling Compounds
Water
3. Naptha
Gravity, "API
(R+M)/2 (unleaded)
Reid Vapor Pressure, Ib.
Weight
1.7
95.9
0.5
1.9
100.00
Weight %
99.66
0.12
0.07
0.15
Weight %
43.5
88.8
3.5
-------
-48-
The thermal efficiencies (based on higher heating values) for
the Davy McKee and Mobil cases are 45.9 and 65.3 percent,
respectively. This is a very significant difference. The Davy
McKee efficiency is a bit lower than the lowest efficiency
reported for methanol from bituminous coals (49.3 percent) in
Table 9. However, the efficiency for the Mobil case is
significantly greater than any of the efficiencies reported for
methanol from bituminous coals. One reason for this high
efficiency is that the raw syngas from the Lurgi gasifier is high
in methane content and the simple isolation of this as product is
more efficient than converting it to carbon monoxide and hydrogen
and then to methanol. Less processing of the raw syngas is
required, and, therefore, a greater efficiency is the result.
Economics; Both studies have been placed on a consistent
economic basis as discussed in a previous report.[11] Table 17
presents the investment costs broken down as much as possible into
individual process unit costs. An inspection of Table 17 shows
that the total instantaneous investments are $1.84 billion for the
Davy McKee case and $2.26 billion for the Mobil case.
Operating costs are presented in Table 18. The difference in
operating cost between both cases is mainly due to annual coal
feedstock cost differences which primarily is a function of
process efficiency. As noted earlier, by-product credit for Case
1 is based on the ammonia and sulfur yields of Case 2.
Table 19 and 20 present economic summaries and product costs
when using capital charge rates of 11.5 and 30 percent. The
methanol product cost for the Davy McKee case ranges from
$6.16-10.26/mBtu while the average product costs for the Mobil
case range from $6.34-$11.24, depending on the capital charge
rate. In addition to average product costs, Tables 19 and 20
present product costs for the methanol, SNG, and gasoline produced
in Mobil study which are based on the product value technique
discussed in another report.[11]
It is possible that the capital cost from the Mobil study is
more accurate than that from the Davy McKee study; the reason for
this is that the original Davy McKee Plant has to be scaled up
significantly whereas the other was much closer to the selected
50,000 FOEB/CD. Therefore, the Mobil study's costs will be used
in preference.
C. Methanol from Lignite
The following two original studies investigated the technical
feasibility of producing methanol from lignite:
-------
-49-
Table 17
Methanol from Subbituminous Coals;
Millions of First Quarter
Capital Cost Summary
1981 Dollars
Davy McKee[6] Mobil [7
Technology
Gasif ication/Methanol Synthesis
Investment Costs
Coal Preparation and Handling
Gasification and Gas Cleaning
Shift Conversion
Acid Gas, Sulfur Recovery, Sulfur
Guard
Syngas Compression
Total: Coal Preparation, Gasification,
Processing
SNG Production
Methanol Synthesis and Distillation
Oxygen Production
Offsites and Product Storage
Infrastructure
Engineering and Design
Environmental Studies, etc.
Other Project Costs
Contingency
Total Instantaneous Plant Investment
Modified
Winkler/ICI
87
98
36
175
66
462
N/A
153
262
302
17
119
-
284
240
1840
Lurgi/
Lurgi
-
628
38
102
161
451
67
184
3
331
295
2260
-------
-50-
Table 18
Methanol from Subbituminous Coal; Operating Cost Summary
(Millions of First Quarter 1981 Dollars Per Year)
Technology
Gasification/Methanol
Synthesis
Raw Materals
Coal
Catalysts and Chemicals
Utilities
Power
Water
Labor and Related
Labor
Supplies
Capital Related
Administration and General Overhead
Local Taxes and Insurance
Interest on Working Capital
Gross Annual Operating Cost
By-product Credit
Net Annual Operating Costs
Mobil[7]
Davy McKee[6] Case 2
Modified Lurgi/Lurgi
Winkler
231
8.4
4.7
173
6.9
2.1
32.5
33.3
39.4
59.2
7.9
416.4
(9.3)
407
49.0
29.0
31.4
62.6
9.8
366
(18.3)
348
-------
-51-
Table 19
Economic Summary of Methanol from
Subbituminous Coal, CCR = 11.5 Percent
(Millions of First Quarter 1981 Dollars)
Total Instantaneous Plant Investment
Total Adjusted Capital Investment
Start-up Cost*
Pre-paid Royalties
Total Capital Investment
Initial Catalyst and Chemicals and
Working Capital***
Total Capital Requirement
Annual Capital Charge
Annual Operating Costs
Total Annual Charge
Product Cost
$/FOEB of Product****
$/mBtu of Product
Methanol, $/mBtu
SNG, $/mBtu
Gasoline, $/mBtu
Davy McKee
1,840
2,087
131
10**
2,228
131
2,359
256
407
663
36.33
6.16
6.16
-
.
Mobil
2,260
2,563
163
25
2,751
163
2,914
335
348
683
37.43
6.34
7.04
5.63
7.04
* Start-up cost =6.3 percent of Total Adjusted Capital Investment.
** Royalties were assumed equal to $10 million unless reported by
study.
*** Working Capital and Initial Catalyst and Chemical = 6.3 percent
of Total Adjusted Capital Investment.
**** One FOEB =5.9 mBtu.
-------
-52-
Table 20
Economic Summary of Methanol from
Subbitumlnous Coal, CCR = 30 Percent
(Millions of First Quarter 1981 Dollars)
Start-up Cost
Pre-paid Royalties
Total Capital I;
Working Capital
Product Cost
Davy McKee
meous Plant Investment 1,840
I Capital Investment 2,053
131
.ties 10
Investment 2,194
il 131
Requirement 2,325
. Charge 698
ng Costs 407
lharge 1,105
Toduct* 60.55
'roduct 10.26
., $/mBtu 10.26
iBtu
! . SmBtu
Mobil
2,260
2,522
163
25
2,710
163
2,873
862
348
1,210
66.30
11.24
12.48
9.98
12.48
One FOEB =5.9 mBtu.
-------
-53-
1« Produ.ction of Methanol from Lignite, prepared by
Wentworth Brothers Incorporated (WBI), and C.F. Braun and Company
for EPRI.[9]
2. Lignite-to-Methanol, an Engineering Evaluation of
Winkler Gasification and ICI Methanol Synthesis Route, prepared ~by
Davy McKee International, Inc.[8]
Both these studies represent approximately the same amount of
engineering design. The WBI/C.F. Braun study uses Texaco
gasification and WBI methanol synthesis technology. Three cases
from this study are presented. Case 1 was prepared by WBI and was
designed based on a 55 percent lignite/45 percent water slurry
concentration. Gasification of a lignite concentration this high
has not been commercially demonstrated. Case 2 represents a C.F.
Braun modification of the WBI design still using the 55 percent
slurry concentration. Since the 55 percent lignite slurry
concentration has not been commercially demonstrated, C.F. Braun
also analyzed a methanol from lignite case based on a 43 percent
lignite slurry concentration which has been suscessfully gasified
(Case 3).
The DMI study is based on Winkler gasification and ICI
methanol synthesis. Both of these technologies have been
commercially proven.
Ultimate Analysis of Lignite; All four cases were based on
the same lignite, and the ultimate analysis for this lignite is
presented in Table 21. Gasifier yields and oxygen requirements
for all four cases are based on this analysis.
Material Balance and Efficiencies; Feedstock and product
rates for each case are presented in Table 22. Methanol and
lignite are presented on both a short ton per calender day (tpd)
and an energy basis. The methanol produced is of fuel grade
quality, even though in Cases 1, 2, and 3 the methanol product
rates are reported on a dry equivalent basis. All rates are based
on 50,000 FOEB/CD of liquid products. Sulfur is the only
by-product reported in Table 22.
Thermal efficiencies vary from 43.9 percent for Case 3 to
51.2 percent for Case 1. The 43.9 percent efficiency results from
the low lignite concentration in the slurry. The vaporization of
the additional water in lower lignite concentration slurries
consumes energy in the gasifier and produces larger quantities of
synthesis gas. The result is a lower thermal efficiency and an
increase of the capacities of all process units except methanol
synthesis.
Economics; Table 23 presents all of the investment costs
broken down into individual process units. These costs are based
on a plant size of 50,000 FOEB/CD of product. An inspection of
-------
-54-
Table 21
Ultimate Analysis of Lignite Feedstock
Heating Values
HHV, dry, Btu/lb 10,179
HHV, wet, Btu/lb 9,765
LHV, approximated, dry, Btu/lb 6,460
Ultimate Analysis, Dry Coal, Wt%
C 58.98
H 4.55
0 19.05
N 0.77
S 1.40
Ash 15.25
Total 100.00
Wt.% Moisture (as received) 35
-------
-55-
Table 22
Methanol from Lignite: Feedstock and Product Rates
(Normalized to 50,000 FOEB/CD of Product)
WBI[9] C.F. Braun[9]
Study; Case 1 Case 2 Case 3 Davy McKee[8;
Mass Basis
Feedstocks
Lignite, tpd (wet) 44,250 44,596 52,071 48,171
Products
Methanol (tpd) 15,063* 15,063* 15,063* 15,226
By-Products
Electricity, energy 8,756 10,107 13,721
equivalent per day
Sulfur, tpd 324 324 384 312
Energy Basis, mBtu/CD, (HHV)
Feedstocks
Lignite 571,705 576,172 671,982 622,363
Products
Methanol 295,000 295,000 295,000 295,000
Thermal Efficiency, % 51.6 51.2 43.9 47.4
* Methanol on a dry equivalent basis.
-------
-56-
Study:
Table 23
Methanol from Lignite: Capital Cost Summary
(Millions of First Quarter 1981 Dollars)
WBI[9] C.F. Braun[9]
Case 1 Case 2 Case 3 DavyMcKee[8]
Technology
Gasification/Methanol
Synthesis
Slurry Concentration, %
Plant Investment Costs
Lignite Storage and
Preparation
Syngas Generation, Gas
Adjustment, and Puri-
fication
Gasification, Com-
pression and Shift
Conversion
Acid Gas Removal,
Chlorideand Sulfur
Guard, Compression
Methanol Synthesis,
Distillation and
Hydrogen Recovery
Methanol Synthesis
Gas Desulfurization,
and Sulfur Recovery
Air Separation
Utility System
Utilities and Offsites
General Facilities
Engineering Fees, Home
Office Cost, and
License Fees
Contingency
Total Instantaneous Plant
Investment
Texaco/ Texaco/ Texaco/ Winkler/
WBI WBI WBI ICI
55 55 43 N/A
41.1 43.5 49.3
637.8 779.4 932.8
155.7
23.0
154.6
91.4
263
2110
164.7
24.4
457.2 483.6
285.96 314.0
96.0
92
286
2283
164.7
27.5
677
314
122.9
92.5
331
2628
114.2
278
368.7
35
661.1
225.8
219
1901
-------
-57-
Table 23 shows that the plant investment estimates vary from
$1.90-2.63 billion. Cases 1, 2, and 3 are based on the same
technology. Case 1 was prepared by WBI; whereas Case 2 is C.F.
Braun's analysis of the WBI design. Both are based on the same
lignite slurry concentration (55 percent).
Braun evaluated the capital costs for adjustments they
thought necessary to appraise the WBI work. The necessary
adjustments were the addition of one spare gasifier/exchanger set
per train and operation with the start-up boiler continuously on
the line thus increasing export power. This equipment was added
as insurance to maintain production levels and to provide
flexibility to the complex. Thus, Case 2 is more conservative;
its capital cost is $170 million more than that for Case 1. It
must be noted that to obtain a 55 percent slurry concentration,
feed pretreatment is necessary and the cost of pretreatment
equipment was not included in either of the estimates.
The instantaneous investment for the C.F. Braun case which
utilizes the 43 percent lignite slurry concentration is $2.63
billion, which is $350-500 million more than the Case 1 and 2
investments, respectively. This higher investment results from
increased capacities of all process units (except methanol
synthesis) needed to accommodate the larger amounts of water (and
steam) present with the 43 percent lignite slurry.
The total instantaneous investment for the Davy McKee case is
$1.901 billion. This case is based on Winkler/ICI technology.
Operating costs are presented in Table 24. Net annual
operating costs range from $237 million for Case 1 to $380 million
for Case 3. Reasons for the low operating cost estimates for Case
1 are that: 1) general and administration cost have not been
included, 2) 1 percent of the total instantaneous plant investment
was used for property taxes and insurance as opposed to 2.5
percent for the other studies, and 3) labor costs were reported to
be less than those for the other studies.
Operating costs for Case 3 are expected to be higher than
those for Case 1 and 2 because of its higher capital investment
and higher feedrate of lignite.
Cases 1 and 2 are based on the same technology and lignite
slurry concentration. Since Case 2 is a further analysis of Case
1 and is more conservative, it is expected that the operating and
capital cost for Case 2 are more representative. Therefore, Case
2 will be used in preference to Case 1 for developing methanol
product costs.
Tables 25 and 26 present economic summaries of methanol costs
for capital charge rates of 11.5 and 30 percent. For the lower
capital charge rate, product costs vary from $5.70 to $6.92/mBtu.
-------
-58-
Table 24
Methanol from Lignite: Operating Cost Summary
(Millions of First Quarter 1981 Dollars Per Year)
Technology
Gasification/Me thanol
Synthesis
Slurry Concentration, %
Annual Operating Costs
Raw Materials
Coal
Fuel
Catalysts and
Chemicals
Utilities
Water
Labor and Related
Operating
Maintenance
Administration and
Support
Capital Related
Ash Disposal
Maintenance
Materials
General and
Administration
Property Taxes and
Insurance
Interest on Working
Capital
Gross Annual Operating
Cost
By-product Credit
Sulfur
Electric Power
WBI[9]
Case 1
Texaco/
WBI
55
160.7
—
3.6
—
24.5
— .
—
—
2.1
52.8
—
21.1
8.9
273.7
(5.8)
(30.5)
C.F. Braun[9]
Case 2
Texaco/
WBI
55
160.7
12.2
3.6
— —
11.2
37.9
14.7
2.1
46.3
24.9
52.8
9.6
376
(5.8)
(35.2)
Case 3
Texaco/
WBI
43
. 190
—
3.6
— —
11.2
45.4
16.9
2.5
58.1
30
65.7
11.0
434.4
(6.9)
(47.8)
Davy McKee
Winkler/
ICI
N/A
175.7
—
17.9
0.2
19. -9-
18.4
11.5
2.4
27.9
25.1
47.5
8.0
354.5
(5.7)
—
Export,
(3.5^/kw-hr)
Net Annual Operating
Cost
237
335
380
349
-------
-59-
Table 25
Economic Summary of Methanol from
Lignite, CCR = 11.5 Percent
(Millions of First Quarter 1981 Dollars)
Total Instantaneous Plant
Investment
Total Adjusted Capital
Investment
Start-up Cost
Pre-paid Royalties
Total Capital Investment
Working Capital
Total Capital Requirement
Annual Capital Charge
Total Annual Charge
Product Cost
$/FOEB of Methanol*
$/mBtu of Methanol
WBI
Case 1
2,110
2,393
148
10
2,551
148
2,699
293
530.5
29.07
4.93
C.
Case
2,283
2, .589
160
10
2,759
160
2,919
317
652
35
6
F . Braun
2 Case 3
2,628
2,980
184
10
3,174
184
3,358
365
745
.73 40.83
.06 6.92
Davy McKee
1,901
2,156
133
10
2,299
133
2,432
264
613
33.61
5.70
One FOEB =5.9 mBtu.
-------
-60-
Table 26
Economic Summary of Methanol
from Lignite, CCR = 30%
(Millions of First Quarter 1981 Dollars)
Total Instantaneous Plant
Investment
Total Adjusted Capital
Investment
Start-up Cost
Pre-paid Royalties
Total Capital Investment
Working Capital
Total Capital Requirement
Annual Capital Charge
Annual Operating Costs
Total Annual Charge
Product Cost
$/FOEB of Methanol*
$/mBtu of Methanol
WBI
Case 1
2,110
2,355
148
10
2,513
148
2,661
754
237
992
54.36
9.21
C.F.
Case 2
2,283
2,548
160
10
2,718
160
2,878
815
335
1,150
63.
10.
Braun
Case 3
2,628
2,933
184
10
3,127
184
3,311
938
380
1,318
01 72.23
68 12.24
Davy McKee
1,901
2,122
133
10
2,265
133
2,398
680
349
1,029
56.38
9.56
One FOEB = 5.9 mBtu.
-------
-61-
For the higher capital charge rate, methanol costs vary from $9.56
to $12.24/mBtu.
Since the present state-of-the-art gasification of lignite
using the Texaco gasifier is costly due to the effect of the
lignite water slurry concentration, the Davy McKee costs will be
used in preference. The Davy McKee study utilizes the proven
Winkler gasification which does not require a coal slurry feed.
Thus, the product cost of methanol varies from $5.70/mBtu for the
low CCR to $9.56/mBtu for the high CCR.
D. Production of Gasoline from Coal via Fischer-Tropsch
and Mobil's MTG Technology
There are three original studies available which investigate
the technical feasibility of producing gasoline from coal-derived
methanol. These studies are:
1. "Coal-to-Methanol-to-Gasoline Commercial Plant," Badger
Plants, Incorporated, Cambridge Massachusetts, FE-2416-43-Vl,2,
March 1979.[10]
2. "Research Guidance Studies to Assess Gasoline from Coal
by Methanol-to-Gasoline and Sasol-Type Fischer-Tropsch
Technologies," Max Schreimer, Mobil Research and Development
Corporation, FE-2447-13, August 1978.[7]
3. "Screening Evaluation: Synthetic Liquid Fuels
Manufacture," Prepared by the Ralph Parsons Co. for EPRI.[1]
The Badger study is based on a "slag bath" gasifier which is
a new concept and may still require developmental work. (See a
more detailed discussion above in Section IV.) Lurgi technology
is used for methanol synthesis, and Mobil fixed bed technology is
used for methanol-to-gasoline conversion.
The Mobil study actually includes three cases, designated
Cases 1, 2, and 3. Cases 1 and 2 utilize Lurgi technology for
coal gasification and methanol synthesis, and Mobil fixed bed
technology for methanol-to-gasoline conversion. These cases
differ in that Case 1 produces approximately 50 percent gasoline
and SNG, whereas Case 2 produces approximately 100 percent
gasoline. Case 3 uses Lurgi gasification technology but employs
Fischer-Tropsch technology for product synthesis.
The Parsons study is based on the BGC/Lurgi gasifier which
still needs to be commercially demonstrated. Fischer-Tropsch
technology is used for product synthesis.
Depth of Design; Both the Badger and the Mobil studies are
based on a comparable level of engineering design. The investment
estimates are of budget or scoping quality. The Badger study is
-------
-62-
based on process licensor's economic data for proprietary
processes and on vendor quotes derived from in-house equipment
specifications for non-proprietary processes. Badger states their
cost estimate represents an accuracy of minus 5 percent, plus 20
percent. The Mobil study is of the same order of accuracy as the
Badger study. The Parsons study is a screening evaluation, and
could be expected to be less accurate than the other two studies.
Ultimate Analysis of Coal-to-Gasoline Feedstock; Ultimate
analysis for the coal feedstocks are presented in Table 27. The
Badger study uses a Southern Appalachian bituminous coal; whereas
Mobil uses a Wyoming subbituminous coal, and Parsons uses an
Illinois No. 6 bituminous coal. The coal considered in the Badger
study is a low sulfur coal which would meet the sulfur dioxide
emissions standard for large power plants.[12] It is unlikely
that a coal of this quality would be used for synfuels production.
For the Badger study, the coal, free of debris, cleaned,
sized, and washed is delivered to the plant site. Thus, this case
excludes coal preparation costs which has been included in the
other studies.
Material Balance and Efficiencies; Feedstock and product
rates for each case are presented in Table 28. All rates are
based on 50,000 FOEB/CD of products (excluding by-products). For
the Badger study gasoline represents almost 100 percent of the
product slate. The efficiency for this case is 49 percent. For
Case 1 of the Mobil study the major products are SNG and gasoline
and the overall process efficiency is 63.2 percent.. The
production of SNG increases the overall process efficiency over
the all-gasoline cases since the isolation of SNG produced in the
Lurgi gasifier requires less energy than gasoline-production.
Gasoline is the main product from Case 2; the efficiency of this
case is 46.6 percent which is comparable to the Badger plant
efficiency. A variety of products are produced from Case 3
(Fischer-Tropsch) with the main products being SNG and gasoline.
The efficiency of this case is 57 percent. The efficiency of the
Parsons case is 56 percent. While both the Parsons and Mobil
(Case 3) studies are based on Fischer-Tropsch technology, their
product slates vary widely. For the Parsons case 15,000 FOEB/CD
of heavy fuel oil is produced compared to 700 for the Mobil case;
whereas 190 mscf/CD of SNG is produced for the Mobil case and only
112 mscf/CD for the Parsons case.
Both of the Fischer-Tropsch synthesis cases produce
significant quantities of SNG and/or residual oil. For a
transportation fuels oriented synthetic fuels industry, their
product slates would be unacceptable. Both cases produce
approximately 33 percent transportation fuels (gasoline and diesel
fuel). However, currently transportation fuels (jet fuel, diesel
fuel, and gasoline) account for 51 percent of the refined
petroleum products and this percentage is expected to increase to
nearly 55 percent by the year 2000 (see Table 7 [37]). At the
expense of an increased product cost, both of these plants could
be altered to meet a more desirable product slate.
-------
-63-
Table 27
Coal to Methanol to Gasoline; Ultimate Analysis of Coals
Study; Badger[10] Mobil[7] Parsons[1]
Coal Type; Southern Appalachian Wyoming
Sub-Bituminous
HHV, Dry, Btu/lb 12,840 11,818 12,771
LHV, Dry, Btu/lb - 10,963 11,709
Ultimate Analysis of
Dry Coal, Wt. Percent
70.84 69.5
4.85 5.3
18.32 10.0
0.71 1.3
0.43 3.9
4.85 10.0
c
H
0
N
S
Ash
Total
-73.8
4.8
6.4
1.6
1.1
12.3
100.0
100.0 100.0
% Moisture 2.4 28.0 4.2
(as received)
-------
-64-
Table 28
Coal-to-Gasoline: Feedstock and Product Rates
(Normalized to 50,000 FOEB/CD of Product)
Mobil[7]
Feedstocks
Coal (tpd)
Electricity
(mBtu/day)
Methanol
Products
Propane bpd
Isobutane, bpd
Butane, bpd
SNG, mSCF/d
Alcohols, tpd
Gasoline, bpd
Diesel Fuel, bpd
Heavy Fuel Oil, bpd
By-Products
Power, mBtu/d
Coal Fines, mBtu/d
Sulfur, tpd
Ammonia , tpd
Energy Basis mBtu/d
Feedstocks
Coal
Electricity
Methanol
Products
LPG
Butane
SNG
Alcohols
Gasoline
Diesel Fuel
Heavy Fuel Oil
Total
Thermal Efficiency, %
Badger [12]
23,147
7,572
—
2,911
4,544
-
-
-
46,729
-
—
-
-
207
—
593,497
7,572
—
10,897
—
-
-
264,855
-
-
294,552
49
Case 1
29,467
-
-
1,678
-
2,380
160
-
23,795
-
—
470
22,043
66
111
501,471
-
—
6,405
9,975
157,139
-
121,474
-
-
295,000
63.2
Case 2
37,219
-
-
3,606
-
5,162
-
-
50,843
-
—
153
-
83
140
633,392
-
-
13,814
21,633
-
-
259,555
-
-
295,000
46.6
Case 3
30,406
-
4.4
1,211
-
160
190
255
14,853
2,523
680
296
-
67
113
517,454
-
85
4,619
690
190,129
7,603
74,607
13,487
3,865
295,360
57.1
Parsons [1]
21,528
-
-
2,005*
112
951*
10,160
6,014*
14,849
-
-
762
-
526,790
-
-
-
11,829
94,524
5,613
59,942
35,484
87,608
295,000
56.0
FOEB
-------
-65-
Product Qualities; The main products from the Badger and
Mobil studies are SNG and gasoline. Chemical and physical
analyses of these products are presented in Table 29. Analyses of
the products from the Parsons study were not available. The SNG
from Cases 1 and 3 of the Mobil study is of satisfactory quality
and is compatible with natural gas. The unleaded gasolines
presented in Table 29 meet all 1976 ASTM specifications. Compared
to typical present-day gasolines these are slightly lower in API0
gravity. It would be preferable if the durene content of the
gasoline were less than 4 wt. percent since durene contents of 5
wt. percent in conventional gasolines have caused carburetor icing
and stalling. The olefinic concentration of the Case 3 gasoline
(20 vol. %) is higher than that of conventional gasolines which
could possibly cause problems with gum formation in storage,
although experience with higher olefinic gasolines is still
limited. Consequently, marketing such a gasoline would require
further testing.
All of the propane and butane products are satisfactory
fuels. The isobutane from the Badger study is of high purity and
may be used as a petrochemical or as a refinery feedstock. The
diesel fuel from Case 3 of the Mobil study could be marketed as a
premium diesel fuel, No. 1-D. The heavy fuel oil from this case
contains no sulfur or metals and thus could be marketed as a
premium gas turbine fuel. The alcohols from these case are a
mixture of C2~C() alcohols, and are essentially free of acids,
aldehydes, ketones and water.
MTG Process Economics; Table 30 presents capital investment
costs broken down into individual process unit costs. An
inspection of this table shows that the estimates of the total
instantaneous plant investment for the MTG process range from
about $2.6 billion for the Badger study and Case 1 of the Mobil
study, to about $3.6 billion for Case 2 of the Mobil study. The
Badger study and Case 2 of the Mobil study are both designed to
produce gasoline as the major product. Since both of these cases
are based on Mobil's methanol-to-gasoline technology and produce
similar product slates, it is expected that their investment costs
would be comparable, but this is not the case. Mobil's capital
estimate is nearly $1 billion more than Badger's. Even though the
capital cost of a subbituminous coal plant is expected to be
greater than a bituminous coal plant, this difference is much too
large. Table 30 shows that the "gasification, et al" costs for
these cases are almost identical, even though the Mobil case is
slightly less efficient and operates with subbituminous coal as a
feedstock as opposed to bituminous coal for the Badger study. On
this basis one would expect Mobil's gasification costs to be
greater than Badger's, which would tend to make the investment
difference between the two studies even greater.
-------
-66-
Table 29
Product Qualities; Coal-to-Gasoline
Study:
1) SNG
Composition, %
Hydrogen
Me thane
Ethene
Ethane
Propene
Propane
Butane
Carbon Dioxide
Inerts (N2 + Ar)
Total
Heat of Combustion, Btu/scf
Water
Sulfur
Carbon Monoxide (0.1% Max)
Study:
2) Gasoline
Gravity, "API
Research Octane Number
Motor Octane Number
Volatility
Reid Vapor Pressure, lb.
Distillation, °F
IBP
10%
30%
50%
70%
90%
EP
Sulfur, Wt.%
Composition, Vol. %
Paraffins
Olef ins
Napthenes
Aromatics
Case 1
1.7
95.5
-
0.2
-
0.1
0.1
0.5
1.9
100
980
0.01%
None
0.02%
Badger
62.7
92.7
82.7
10.0
79
106
140
187
259
339
390
Nil
54
12
7
27
Mobil [7]
Mobil[7]
Case 1,2
61.4
93
83
10.0
85
110
146
200
262
336
388
Nil
51
11
9
29
Case 3
3.8
89.7
1.0
2.3
1.0
0.1
-
0.5
1.6
100
1000
0.01%
None
0.07%
Case 3
67.2
91
83
10.0
86
108
137
186
249
335
420
Nil
60
20
3
17
Durene Content
4 Vol.5
4.6 Wt.%
-------
-67-
Table 30
Coal-to-Gasoline: Capital Cost Summary
(Million of First Quarter 1981 Dollars
Study;
Mobil[7]
Badger[10] Case 1
Case 2
Case 3 Parsons[1]
Technology (Gasification/
Synthesis)
Investment Costs
Slag Bath/
Lurgi/MTG
Coal and Lime Preparation 78
Coal Gasification 210
Shift Conversion 63
Acid Gas Removal 216
Sulfur Recovery 18
Hydrocarbon Recovery
By-product Separation
and Recovery
Syngas Compression 36
Methanol Synthesis 220
Cyrogenic Hydrogen Recovery 18
Total Gasification, et al 859
SNG Production
Gasoline Production 216
F-T Synthesis and F-T
Product Processing -
Oxygen Production 351
Steam and Power Generation 121
Cooling Water and
Make-up, WWT 90
Environmental 63
Waste Disposal -
Storage and Shipping 11
General Facilities 118
Infrastructure
Other Project Costs
Engineering and Design 246
Miscellaneous 138
Sub-Total 2197
Contingency 330
Contractor's Fee 56
Total Instantaneous Plant
Investment 2583
Lurgi/ Lurgi/Lurgi/ Lurgi/
Lurgi/MTG MTG Fischer-
Tropsch
304
49
109
89
78
82
4
715
38
106
167
231
69
104
34
70
376
189
107
2206
331
55
2592
840
370
261
266
128
286
83
510
288
3022
455
76
3563
307
49
110
90
30
19
614
37
218
168
274
77
109
13
71
270
208
169
2228
343
57
2628
BGC Lurgi/
Fischer-
Tropsch
41
147
57
244
44
100
70
31
733
342
266
11
212
1624
244
41
1909
-------
-68-
Operating costs are presented in Table 31. Operating costs
for the Mobil Case 2 study are $150 million more than those for
the Badger study. Unfortunately, not enough information was
available to reconcile these capital and operating cost
differences.
Tables 32 and 33 present economic summaries and average
product costs when using capital charge rates (CCR) of 11.5 and 30
percent. The average product costs for the MTG processes range
from $7.37-9.75/mBtu for the low CCR and from $12.94-17.43 for the
high CCR. In addition to average product costs, product costs for
the various studies based on the product value method discussed in
a previous report are also presented in these tables.[11]
While the cost estimates of these two studies are difficult
to reconcile, the incremental product cost to produce gasoline
from methanol using the MTG process can be determined and may be
more consistent. To accomplish this the incremental investment
and operating costs will be determined between: 1) the Badger
methanol plant (from the bituminous coals section) and the Badger
gasoline from methanol plant, and 2) the Mobil methanol plant
(from the subbituminous coals section) and the Mobil (Case 1)
gasoline from methanol plant. Then the incremental product costs
for each case will be compared. Mobil's Case 1 MTG unit is sized
to produce 20,600 FOEB/CD of gasoline while Badger's was sized to
produce 45,000 FOEB/CD. Therefore, in the economics to be
presented below the Mobil study MTG unit has been scaled up to
45,000 FOEB/CD.
The incremental instantaneous plant investments ares
1. $634 million - Badger Study
2. $596 million - Mobil Study
The incremental operating costs are:
1. $97 million - Badger Study
2. $53 million - Mobil Study
After determining the total annual charge per the procedure
discussed in a previous report,[11] the incremental charge to
produce 45,000 FOEB/CD of gasoline (50,000 FOEB/CD of total
product) from methanol via the Mobil MTG process is:
$/mBtu
Capital Charge Rate
11.5% 30%
Badger 1.76 3.12
Mobil 1.45 2.87
-------
-69-
Table 31
Coal-to-Gasoline: Operating Cost Summary
(Millions of First Quarter 1981 Dollars Per Year)
Raw Materials
Coal
Limestone
Catalysts and
Chemicals
Utilities
Power
Water
Labor and Related
Operations
Maintenance
Supervision
General Services
Capital Related
Operating
Maintenance
Administration and
General Overhead
Local Taxes and
Insurance
Interest on Working
Capital
Other Operating Cost
Gross Annual
Operating Cost
By-product Credit
Net Annual Operating Cost
Badger [10]
232
8.4
27.3
28.3
19.5
13.4
1.7
3.1
38
7.5
5.2
32
416
(3.8)
412
Mobil[7]
Case 1 Case 2
184 232
7.9
2.2
9.6
46.5
2.4
31.1
34.6
71.9
5.3 8.2
295
396 535
(17.4) (11.0)
378 524
Case 3 Parsons
186 216
9.4 11
2.2
14.2
53.1
3.5
35.5
41.7
80
5.5 7.0
135
431
(9.6) (13.9)
422 355
-------
-70-
Table 32
Coal-to-Gasoline: Economic Summary, CCR = 11.5%
(Millions of First Quarter 1981 Dollars)
Total Instantaneous
Plant Investment
Total Adjusted Capital
Investment
Start-up Cost
Pre-paid Royalties
Total Capital Investment
Working Capital
Total Capital Requirement
Annual Capital Charge
Annual Operating Costs
Total Annual Charge
Average Product Cost
$/FOEB of Product
$/mBtu of Product
Product Costs, $/mBtu
LPG
Butane
SNG
Alcohols
Gasoline
Diesel Fuel
Heavy Fuel Oil
Badger
2583
2929
182
26
3136
182
3319
382
412
794
43.51
7.37
5.82
5.82
-
-
7.55
-
-
Case 1
2592
2939
181
26
3146
181
3327
383
378
761
41.68
7.06
6.17
6.17
6.41
•
8.01
-
-
Mobil
Case 2
3563
4040
249
34
4323
249
4572
526
524
1050
57.52
9.75
7.72
7.72
-
-
10.03
-
-
Case 3
2688
3048
188
20
3256
188
3444
396
422
818
44.83
7.60
6.56
6.56
6.82
8.52
8.52
7.67
6.56
Parsons
1904
2159
133
9
2301
133
2434
280
355
635
34.79
5.90
5.31
5.31
5.52
6.90
6.90
6.21
5.31
-------
-71-
Table 33
Coal-to-Gasoline: Economic Summary, CCR = 30%
(Millions of First Quarter 1981 Dollars)
Mobil
Total Instantaneous Plant
Investment
Total Adjusted Capital
Investment
Startup Cost
Pre-paid Royalities
Total Capital Investment
Working Capital
Total Capital Requirement
Annual Capital Charge
Annual Operating Costs
Total Annual Charge
Average Product Cost
$/FOEB of Product
$/mBtu of Product
Product Costs, $/mBtu
LPG
Butane
SNG
Alcohols
Gasoline
Diesel Fuel
Heavy Fuel Oil
Badger Case 1 Case 2 Case 3 Parsons
2,583 2,592 3,563 2,688 1,904
2,883 2,893 3,976 3,000
182 181 249 188
26
26
34
20
3,091
182
3,273
982
412
1,394
76.37
12.94
10.19
10.19
-
-
13.25
-
_
3,100
181
3,281
984
378
1,362
74.65
12.65
11.06
11.06
11.49
-
14.36
-
-
4,259
249
4,508
1,352
524 .
1,876
102.82
17.43
13.80
13.80
-
-
17.93
-
-
3,208
188
3,396
1,019
422
1,441
78.95
13.38
11.36
11.36
11.80
14.75
14.75
13.28
11.36
2,125
133
20
2,278
133
2,411
723
355
1,078
59.09
10.01
9.04
9.04
9.39
11.74
11.74
10.57
9.04
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The costs from the Badger study are slightly higher than those
from Mobil. Since Mobil has researched and developed the MTG
process, it is believed that their study is more reliable;
therefore, their costs will be used in preference to Badger's.
Fischer-Tropsch Process Economics; This section examines the
investment cost differences between the two Fischer-Tropsch
studies. The instantaneous plant investment of the Mobil/Fischer-
Tropsch case is $719 million more than that of the Parsons case.
When inspecting onsite process equipment costs, the Parsons case
cost $343 million more. However, the cost of offsite type
equipment is $908 million more for the Mobil case. Therefore,
even though there is a large onsite investment cost difference,
the major differences between the two studies appears to be in
offsite investment costs. The Mobil study is probably more
accurate since it is based on a more thorough design. Operating
costs from the Mobil/Fischer-Tropsch study are $75 million greater
than those from the Parsons study. Unfortunately not enough
information was available to reconcile these differences.
Tables 32 and 33 present economic summaries and average
product costs for both CCR's. The average product costs for the
two Fischer-Tropsch studies ranged from $5.90-7.60/mBtu for the
low CCR to $10.01 to 13.38/mBtu for the high CCR. Product costs
based on the product value method are also presented in these
tables. Since the Parsons' study is based on a less thorough
design than the Mobil study, the Parsons' study will not be
further investigated.
To determine the average product cost difference between a
Fischer-Tropsch synthesis plant and a methanol synthesis plant,
the Mobil study (Lurgi gasification/Fischer-Tropsch synthesis) can
be compared with Mobil's Lurgi gasification/Lurgi methanol
synthesis case from Table 17. Differences in investment and
operating costs between these two cases reflect the differences in
synthesis technology. The instantaneous plant investment
difference is $355 million and the operating cost difference is
$67 million with the Fischer-Tropsch case being greater. These
figures translate into an average product cost difference of
$1.00/mBtu.
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-73-
References
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-74-
12. "The Potential for Methanol from Coal: Kentucky's
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-75-
27. "Projects Applying to SFC For Aid Under Second
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