Note: Descriptions are shown in the official language in which they were submitted.
1041SS3
The present invention relates to the concurrent
production of methanol and synthetic natural gas from a
carbonaceous solid or liquid material via an integrated process.
The prior art considered in the preparation of
this specification is as follows: U.S. 1,735,925; U.S.
1,741,307; U.S. 1,741,308; U.S. 1,824,896; U.S. 1,831,179;
U.S. 2,276,343; U.S. 3,598,527; and U.S. 3,666,682.
Synthetic natural gas can be prepared from the
gasification of a carbonaceous solid or liquid materials
such as coal or high boiling petroleum residues; however,
the present costs are relatively high.
The conventional manufacture of methanol is based
on the catalytic conversion of carbon oxides and hydrogen
over a catalyst at elevated pressure. The carbon oxides/
- hydrogen synthesis gas mixture is usually prepared by steam -;
reforming of a natural or refinery gas stream high in methane
content. Since this leads to a synthesis gas steam too rich
in hydrogen for the stoichiometry of the methanol synthesis,
extraneous carbon dioxide can be compressed and fed into the
methanol plant to achieve a proper hydrogen/carbon oxides balance.
Synthesis of methanol from natural gas is be-
coming increasingly unattractive as natural gas availability
decreases and gas supplies are supplemented by synthetic
substitutes derived from other fossil fuel sources; i.e.,
naphtha or heavy carbonaceous materials such as coal,
coke, and petroleum residue. Production of the methanol
synthesis gas directly from these heavy carbonaceous materials
can be accomplished by high temperature partial
.. . .
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;i 1041SS3
1 oxidation with essentially pure oxygen, a very expensive
- 2 process on the scale required for reasonably sized methanol - 3 plants.
4 The synthesis of methanol from carbon oxides and
~ 5 hydrogen is an equilibrium reaction and requires exten-
- 6 sive recycle of unconverted reactants through the high
7 pressure reactor in order to achieve the high degree of
8 conversion and full utilization of the expensive synthesis
9 gas constituents required for economic operation. Build-
up of inerts in this recycle and imbalance in the ratio
11 of hydrogen to carbon oxides leads to high purge rates
i 12 of synthesis gas and expensive recycle compressor as well
13 as reactor requirements in the conventional methanol
14 plant.
However, methanol can be produced by the use of
16 carbon monoxide and hydrogen in a molar ratio respectiv-
17 ely of about 1:2. The present invention overcomes the
18 high costs of the synthetic natural gas production and
19 the methanol synthesis's use of expensive feed stock,
processing equipment, recycle facilities and reciproca- - -.
21 ting compressors.
22 Accordingly, it is an ob~ect of this invention -
23 to provide a process for integrating the production of `-
24 methanol and synthetic natural gas.
It has been unexpectedly found that the
26 above ob~ects can be accomplished by an integrated process
27 for the concurrent production of methanol and synthetic
28 natural gas (which is substantially methane) from a car-
29 bonaceous solid or liquid material. In general, this inte-
grated process comprises the gasification of a carbonaceous .
."
- 3 ~
'.' ' .'' ."
1041SS3
solid or liquid material to form a crude synthesls gas whlch ls then sub~ected
to a water gas shift reaction under controlled conditions to provide the
necessary stoichiometry for the subsequent productlon of methanol and addltional
quantities of methane. The "shifted" or converted synthesis gas is then passed
to a carbon dioxide and sulfur compound removal zone, followed in series by a
methanol synthesis zone and a methanation zone. The desired end products, i.e.
methanol and synthetic natural gas are concurrently and respectively withdrawn
from these latter two zone~.
More particularly, the present invention provides an integrated
process for the production of methanol and a synthetic natural gas, which is
substantially methane, from a carbonaceous material which comprises the steps of:
a. gasifying the carbonaceous material in a first reaction zone
at sufficient pressure aDd temperature to produce a synthesis gas
stream comprising methane, carbon monoxide, carbon dioxide, hydrogen,
and steam, in which the synthesis gas contains about 5-30 mole %
methane on a dry basis;
b. passing said synthesis gas stream to a second reaction zone
wherein at least a portion of said carbon monoxide is reacted with
at least a portion of the steam present therein and is converted at
a temperature between about 200 and 950F and a pressure between
about 15 psia and 3000 psia and in the presence of a water gas shift
conversion catalyst, to carbon dioxide and hydrogen, thereby produc-
ing a converted gas stream with the proviso that the shift conversion
is controlled to such an extent as to provide the necessary stoich-
iometry for the subsequent production of methanol and additional
quantities of methane;
c. passing said converted stream to a third zone wherein the sulfur
compounds and a ma~or portion of the carbon dioxide in said converted
stream are removed therefrom to produce a purified stream;
d. passing said purified stream into a fourth reaction zone wherein
a portion of the carbon oxides and hydrogen in said purified stream
-
' ' '' ~ ,, ` ' `
1041SS3
are converted, ln the presence of a methanol converslon catalyst and
at a sufficient temperature and at pressures ranging from 400 to
about 1500 p8ig, to methanol;
e. recovering methanol from said fourth reaction zone and separately
recovering the unreacted portion of said purified stream which contains
a reduced amount of carbon oxides;
; f. passing the unreacted portion of said purified stream containing
the remaining carbon oxides, hydrogen, and the methane produced in
the first reaction zone from ~aid fourth reaction zone to a fifth
reaction zone wherein said remaining carbon oxides and hydrogen are
converted, in the presence of a methanation catalyst and at a
temperature between about 400 and 800 F and a pressure between about
250 psig and 1500 psig, to methane.
~, Referring now to the drawing, Figure 1 is a schematic flowsheet of
the process of this invention and represents a preferred embodiment but is -
......... ................................................................................... .... ... ..... ... ... ,~ .
not to be considered as limiting. ~ -
A carbonaceous solid or heavy liquid material is in~ected, line 1,
into a first reaction zone, such as a coal gasification zone, 2, wherein it is
contacted with in~ected steam. The reaction therebetween results in the
production of a crude synthesis gas (1000-2000F and 50-1500 psig) which exits,
line 3, and is iniected into the water gas shift conversion zone 4. In the
water gas shift conversion zone 4, the carbon monoxide contained in said -~
synthesis gas is reacted (15-3000 psig and 200-1000F) with steam to form
carbon dioxide ant hydrogen. The "shifted" or converted gas then exits zone
:
4, line 5, and is in~ected into the acid gas removal zone 6 wherein the sulfur
compounds such as hydrogen sulfide and carbon dioxide are removed. The sub-
stantially purified synthesis gas then exits, line 7, and is injected into a
- .
methanol ~ ~
~ .
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10~5S3
1 synthesis zone which comprlses, for example, one methanol
reactor (400-600F and 250-1500 psig) and a methanol re-
3 covery zone in series. First, the purlfied eas is in~ect-
4 ed into methanol reactor 8, line 7, where it is catalytical-
ly sub~ected to the conversion of carbon monoxide plus hy-
; 6 drogen to methanol. The molar ratio of carbon monoxide to
; 7 hydrogen reached therein is about 1:2. Carbon dioxide will
8 also react with hydrogen to form methanol, but the relative
9 comsumptions will be 1:3 C02:H2. The gases then exit, line
9, and are passed through a cooler and knock-out device 10
11 wherein a portion thereof which is substantially methanol
12 is recovered, line 11. The methanol, via line 11, is then
,A 13 sub~ected to a purification in methanol purification zone
14 12. One reactor-condensation zone has been shown in this
figurR as an illu~tration of the operation o~ the methanol
16 synthesis zone. Depending upon the proceæs conditions and
17 the degree of conversion to methanol derived, the number
18 of such zones required may vary from a single zone to as
19 many as 3 or 4 in series. Thereafter, the substantially
purified methanol is recovered, line 13, and which can be -
- 21 commercially utilized as such or sub~ected to further puri-
22 fication steps and/or reactions. The remaining synthesis
23 gas which has not undergone methanQl synthesis in methanol
24 reactor 8 then proceeds, line 14, intothe methanation
zone 15 wherein the carbon oxides remaining in the gas are
26 catalytically converted to methane via the reaction (400-
27 800F and 250-1500 psig) with hydrogen. The methane so
28 produced in methanation zone 15 adds to the other methane
29 already present in the synthesis gas and exits, line 16,
a8 a complete synthetic natural gas which is substantially
. - : s
., .: . :
104~553
1 methane and which can be in~ected directly (for example,
2 at about 1000 psig) into a pipeline for commercial use
3 thereof. Appropriate compression facilities, not shown
4 in the attached flowplan, can be provided anywhere in the
process sequence, but preferably after zone 4.
6 Although the invention has been illustrated by
7 means of a flowsheet and specific process conditions and
8 equipment, the following discussion more fully describes
9 the conditlons and equipment set forth in the fbwsheet.
The carbonaceous solid or liquid material for ;~
11 use in the gasification 70ne i8 any material whIch con-
12 tains carbon and which will ultimately, when gasified,
13 produce a crude synthesis gas containing methane, carbon
14 monoxide, carbon dioxide, sulfur compounds, hydrogen, and
water. The preferred carbonaceous material for the present ~ -
16 gasification process is a bituminous, sub-bituminous, lig-
17 nite or brown coal, although other essentially organic
18 sources such as coke, char, tar sands, shale oil, solidi-
19 fied petroleum or petroleum coke, heavy residuums, such as
from the vacuum distillation of petroleum, or the like may
21 be employed.
22 Other meterials which are suitable for feed to
23 the gasification zone are tar and heavy crudes found indi-
24 ginous in certain locations, vegetable matter such as wood, ;
wood chips, charcoal, paper, agricultural wastes, and col- ~ -
26 lected organic products and organic waste products in
27 general.
28 In the gasification reaction zone, the carbona-
29 ceous solid or liquid material is sub~ected to gasifica-
tion in order to prepare the crude synthesis gas containing
..
. .
1041553
1 the aforementioned components. The ~asificatlon of
2 a carbonaceous ~aterial can be carried out in any manner
3 as long as the desired end result, i.e., a crude synthesis
; 4 gas, is produced. It is usually desirable in the process
to produce as much methane as possible in the gasification
6 process directly, minimizing the heat duty in the gasifier
7 and minimizing the amount of shifting, acid gas removal
8 and methanation which have to be carried out subsequently.
9 Present coal gasification processes generate up to about
one-half of their total eventual methane product directly
11 in the gasifier, while the remaining half is produced in
12 the down-stream methanator. This leads to about 10-30~
13 methane in the crude (dry) synthesis gas. Other gasifi-
14 cation processes, especially those based on partial oxida-
tion, lead to crude synthesis gases containing between
16 0.5-5% methane. It is a typical embodiment herein that
17 the gasification be carried out, particularly utilizing
18 coal as the carbonaceous solld material, in accordance
19 with known processes.
A typi¢al carbonaceous solids hydrogasification
21 process especially useful for coal hydrogasification in-
22 volves heating subdivided carbonaceous feed solids contain-
23 lng volatilizable hydrocarbonaceous matter to at least min-
24 imum hydrogasification temperature by dilute phase suspen-
sion in a hydrogasifying gas in contact with subdivided
26 hot solids having a temperature greater than minimum hydro-
27 gasification temperature. The feed and heating solids are
28 co-currently passed with the hydrogasifying gas through
29 a transfer line hydrogasiflcation zone having a length
which, for the velocity of the solids passage therethough,
_ ~ _
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: . ., ,,~ ,~. ~ , . .
10~15S3
1 limits the residence of the solids therein to only a time
2 necessary for devolitilization of the carbonaceous ~olids
3 material and for conversion of a predetermined proportion
4 of the carbon of the solids to methane. Suitably from
about 1 to about 50 mol percent of the carbon of the
6 carbonaceous feed solids is converted to methane. Pre-
7 ferably the hydrogasifying gas is produced in a fluid bed
8 stream/char reaction zone into which carbonaceous solids
9 from the transfer line hydrogasification zone are charged
after separation therefrom of product gases containing
11 methane.
12 In a typical embodiment of part of the present
13 invention, the transfer line hydrogasification reactor ls
14 coupled with a fluid bed steam-char reactor for generating
a hydrogen-containing gas, and subdivided coal feed solids
16 are suspended with hot char solids withdrawn from the fluid
17 bed o~ the steam/char reactor. The solids are transported - -
18 in dilute phase through the transfer line hydrogasifica- -~
19 tion zone by a hot hydrogen-containing gas separately
withdrawn from the steam/char reactor. The fluid bed
21 steam/char reactor is operated at a temperature in the
22 range from about 1500 to about 2000F. and a pressure in
23 the range from about 50 to about 1000 psia. Thus, the
24 char solids and the hydrogen-containing gas withdrawn
from the fluid bed reactor have a temperature in the
26 range from about 1500 to about 2000F. Because the char
27 solids from the fluid bed reactor have more heat capacity
~8 than the hydrogen-containing gas, the feed coal solids
29 transported through the transfer line hydrogasificr are
heated principally by contacting the hot char co-currently
- - . . - , - ; ~ . . ~ ~; .
- . . .
.. . - . ~ . . ' :-
i
104~S53
1 transported therewith under the flow conditions existing
2 in the transfer line. The heated coal undergoes coal de-
3 volatilization, and the reactive fresh carbon of the feed
4 coal reacts rapidly with the hydrogen in the hydrogen-
containing gas to produce methane. The produced gas (crude
6 synthesis gas) is recovered from the transfer line hydro-
7 gasifier and sent to the water shift conversion zone
8 (hereinafter described), and the residual coal solids and
9 the char solids are charged to the fluid bed steam/char
reactor.
11 For the steam/char reactor of the preferred em-
12 bodiment, there i6 employed the fluid solids technique
13 of maintaining highly divided solids in the form of a
14 dense turbulent mass fluidized by an upwardly flowing
gasifoFm reaction material, here including super heated
16 steam, to resemble and have the hydrostatic and hydro-
17 dynamic characteristics of a boiling liquid. This tech-
18 nique is of special advantage, particularly in a con-
19 tinuous operation, for it provides larger solid reaction
sur~aces, better mixing, greatly improved temperature
21 control and generally higher yields of hydrogen than oc-
22 cur when using fixed or gravitating beds. Furthermore,
23 it facilitates the handling of solids, for they may be
24 treated in a manner analogous to that used for liquids;
this enables particulate solids to be withdrawn and con-
26 veyed as if liquid-like into the transfer line hydrogasi-
27 fication reactor.
28 The heat required to maintain the desired temper-
29 atures in the steam/char fluid bed reactor, which requires
operating temperatures of from about 1500F to about
,
_ 9 _
'`. .
.' , ', ~ . ~~ ., ~ ' ,
lO~lS~3
1 2000F., may be supplied in any manner known in the art
2 o~ hydrogen-containing gas generation. For example, ~uf-
3 ficient amounts Or an oxidizing gas may be supplied to
4 the steam/char reaction zone to generate by partial com-
bustion within the zone the heat required by the steam/
6 char reaction to produce the hydrogen-containing gas.
7 Preferably, however, the necessary heat is supplied as
8 sensible heat of hot solids by burning solid carbonaceous
9 gasification residue with air in a separate combustion
zone and circulating highly heated combustion residue
11 from this combustion zone to the steam/char reactor zone.
12 The latter method has the advantage of avoiding dilution
13 of the product gases with noncombustible fluid gases.
14 While the gasification of a carbonaceous solid
material has been described above, it is to be understood
16 that such is merely a typLcal and not a critical li~ita-
17 tion. Other gasification processes can be utilized as
18 long as the desired crude synthesis gas containing meth- -` -
19 ane, carbon monoxide and hydrogen is produced.
The crude synthesis gas which contain~ carbon -~
21 monoxide is then sub~ect to a water shift conversion in
22 a second reaction zone under controlled conditions in
23 order to convert a portion of the carbon monoxide to car-
24 bon dioxide and hydrogen. The amounts Or steam employed
in the "shirt converter" are those amounts which will
26 equilibrate with carbon monoxide to provide the stoich-
27 iometry rOr the subsequent production Or methanol and
28 additional quantities Or methane. Specirically, the crude
29 syntheæis gas iB fed, after the introduction Or water or
steam, if necessary, to, for example, a high temperature ;?'.... ~ ,:
lO~lS~
1 shift converter which achieves the desired conversion of
2 carbon monoxide to carbon dioxide and hydrogen.
3 It may even be necessary to remove some of the
4 water present in the crude synthesis gas before entering
the shift conversion zone in order to avoid converting
6 too much CO into hydrogen, which could result in an un-
7 satisfactory stoichiometric ratio of H2 to carbon oxides
8 for methanol and methane synthesis. This ad~ustment of
9 water content can be achieved, if so desired, by a hot
water scrub of all or part of the crude synthesis gas at
11 a controlled temperature.
12 The high temperature shift conversion is accom-
13 plished at elevated temperature and pressure usi~g a cata-
14 lyst.
The catalyst used for a high temperature shift
16 conversion may be any polyvalent metal or oxide thereof
17 capable of converting carbon monoxide to carbon dioxide.
18 Among the catalysts which may be used are iron oxide, co-
19 balt oxide, chromia, molybdena and tungsten oxide and the
like. Other catalysts may also be used and such other
21 catalysts will be obvious to one skilled in the art.
22 The space velocity during the high temperature
23 shift converslon may range from about 1000 to about 5000
24 volumes of dry gas at standard conditions (60F. and at-
mospheric pressure) per hour per volume of catalyst and
26 preferably from about 2000 to about 3000.
27 The inlet temperature for the high temperature
28 shift converter will vary from about 600F. to about 800F.
29 Because the shift conversion reaction is exothemic, the `~
temperature of the gas in the course of its passage over
~ 31 -
. ~ . . . . ., . . ~ . , ,
10415S3
1 the catalyst will rise beyond that of the inlet tempera-
2 ture. Therefore, the outlet temperature from the con-
3 verter will vary from about 700F. to about 900F. Pres-
4 æure will range from 50-1500 psig.
The major portion or all of the desired conversion
6 of carbon monoxide is completed in the high temperature
7 shift converter. The gases are then cooled in a heat ex-
8 changer and the recovered heat may be used to generate
- 9 steam. The cooled gases may then be fed to a low temper- -
ature shift converter which substantially completes the
11 extent o~ conversion of carbon monoxide to carbon dioxide
12 with the production of a corresponding amount of hydrogen.
13 If desired, a low temperature shift converter
14 need not be employed if conversion of the carbon monoxide
to hydrogen to the desired extent will take place in the
16 high temperature shlft converter. -
17 The pressure used in the low temperature shift
18 converter will be substantially the same as that employed
19 in the high temperature shift converter, i.e., between
about 50 to about 1500 psig, dependlng upon the pressure
21 used for methanol synthesis.
22 The low temperature shift conversion i9 accompli-
23 shed using an inlet temperature of from about 350F. to -~
24 about 600F. and an outlet temperature of from about
. . .
400F. to about 650F. ; ~ ~
26 The steam-gas ratio ior both the high temperature ~ -
27 and low temperature shift conversion will be from about
28 0.2 to 1 to about 1 to 1. However, lower or higher ratios
29 of steam to dry gas may also be used depending on the
design requirements of the manufacturing facility.
' :
- 12- ;:
104~S~3
1 The space velocity utillzed in low temperature
2 shift converslon depends upon the degree of carbon mono-
3 xide conversion desired and the steam-dry gas ratio em-
4 ployed. Generally, the space velocity may vary from
about 2000 to about 4000 volumes of dry gas at standard
6 conditions (60F. and atmospheric pressure) per hour per
7 volume o~ catalyst. It is pre~erred to utilize space
8 velocities of from about 2500 to about 3500.
9 The catalyst for water shift conversion may be
in any suitable form such as granules, pellets, tablets,
11 and the like.
12 In addition to the above described water shift
13 conversion process, other processes can be used where
14 one so desires.
The converted synthesis gas is then passed to a
16 third reaction zone wherein sulfur compounds such as hy-
17 drogen sulfide and carbon dioxide are removed therefrom.
18 While this zone has been named the acid gas removal zone,
19 i~ i8 to be understood that both carbon dioxide and sul-
fur compounds are removed therein; however, more than
21 one chamber may be utilized to carry out the desired end
22 result. For example, carbon dioxide can be removed from ~ -
23 the converted synthesis gas by passing said gas through a
24 vessel in which is regenerative solvent capable of remov~
ing carbon dioxide. Among the solvents which may be used ~
26 to remove carbon dioxide are monoethanolamine, diethanol- ~ -
27 amine, hot potassium carbonate and an additive such as
28 arsenic oxide dlethanolamine and the like.
29 Hydrogen sulfide can be remoYed by the same re-
generative solvents and by also passing said gas, for
- 13 ~ ;
.` ~:
-
10415~3
1 example, through a solution such as ammonlum thioarsenate.
2 Organic sulfur compounds can also be removed from the
3 converted synthesls gas by passing sald gas over hot lime
4 or activated aluminum oxide which converts any organic
sulfur components into hydrogen sulfide. Furthermore,
6 the newly ~ormed hydrogen sulfide can then be eliminated
7 by passing the said ~as containing the same through a
8 caustic scrubber which removes all traces of hydrogen sul- -
9 fide and carbon dioxide. It is to be understood that the
foregoing processes outlined for the removal of sulfur
11 compounds and substantially all of the carbon dioxlde are
12 exemplary only and other processes can be utilized as
13 long as the desired end results are achieved.
14 Any undesirable nitrogen compounds in the synthe-
sis g~s are also removed in the scrubbing operations des-
16 cribed above.
17 The purified, converted synthesis gas is then
18 sub~ected to methanol synthesis in a fourth reaction zone.
19 The composition of the synthesis gas, prior to
the conversion of a portion of said synthesis gas in the
21 methanol synthesis zone, will vary depending upon the
22 type of carbonaceous solid material processed ln the
23 gasification facility and the desired methane and methan-
24 nol production rates.
There are three basic overall reactions which
26 take place inthe methanol synthesis zone although there
27 are other minor reactions which also take place in the
28 converter. These reactions are as follows:
29 (a)CO+2H2 CH30H -
(b)C02+3H2 CH30H+H20
~4 -
, .~. - :.
~ . : '.' ~ ~ . ,
~0415S3
( c ) C2~H2 CO~H20
2 The reactions in the methanol synthesis zone oc-
3 cur under the influence of elevated temperature and pres-
4 sure and in the presence of a catalyst.
The temperature for methanol synthesis may vary
6 from about 400F. to about 600F. and preferably from
7 about 450F. to about 520F., although temperatures in
8 excess of 600F. may be used with caution where one so de-
9 sires as long as the desired end result is obtained.
However, there is danger of excessive methane formation
11 and temperature run-aways.
12 The synthesis pressure also may vary from about
13 4 to about 1500 psig. -~
14 The catalyst used for the methanol synthesis
may be either a slngle catalyst or a mixture Or catalysts.
16 It may be finely ground, pelleted, granular in nature, an
17 extrusion using a binding agent, or any other suitable
18 form.
19 Among the catalysts which may be used are parti~
ally reduced oxides of copper, zinc and chromium as a
21 catalyst system, zinc oxide and chromium oxide, zinc ox- -
22 ide and copper, copper and aluminum oxide or cerium ox-
23 ide, zinc oxide and ferric hydroxide, zinc oxide and
24 cupric oxide, a copper zinc alloy, and oxides of zinc,
magnesium, cadmium, chromium, vanadium and/or tungsten,
26 with oxides of copper, silver, lron and/or cobalt, and
27 the like. Other catalysts which are well known in the ;;
28 art may also be used and the lnvention is not to be con- ~ ~ -
29 strued as limited to any particular catalyst or catalyst ` ~ -
system.
.'
- 15 - :
,.: ` - . .. - . ~ :
. . . . . .
10415S3
1 The methanol synthesis zone i5 a pressure vessel
2 containing a charge of catalyst arranged in the vessel as
3 a continuous bed or alternatively, as several independent- -
4 ly supported catalyst beds. Facilities are provided in
the converter to permit heat removal, or, alternatively,
6 the in~ection of cold synthesis gas into the catalyst
7 bed or between the catalyst beds in order to control the
8 reaction temperature. The quantity of catalyst provided
9 in the converter will depend on the methanol synthesis
pressure employed, the synthesis gas composition, and
11 the desired degree of conversion of synthesis gas to
12 methanol in each catalyst bed.
13 The space v~locity employed in the methanol syn-
14 thesis converter is from about 5000 to about 50,000
volumes of dry gas at standard conditions (60F. and at~
16 mospheric pressure) per hour per volume of catalyst and
17 preferably from about 7000 to about 25,000.
18 Prior to recovering the methanol product, and ~;
19 where one so desires, the effluent from the methanol con- ~ `
verter undergoes a heat exchange with a portion of the
21 incoming synthesis gas in order to preheat the incoming
22 synthesis gas to the initiation temperature of the meth-
23 anol synthesis reactlon.
24 The methanol converter effluent may then be water
cooled to condense the methanol as well as any water
26 formed in the methanol synthesis converter. The amount
27 of water contained in the crude methanol will depend upon
28 the amount of carbon dioxide left after the acid gas -
29 scrubbing and then reacting in the methanol synthesis
zone and the amount of water present in the methanol
,
104~5S3
1 synthesis gas feed.
2 Any suitable system may be used if it is desired
3 to purify the crude methanol. One such suitable ~ystem
4 comprises a topping column operating at low pressure of
up to about 20 psig which consists of a plurality of
6 bubble trays designed to remove light components contained
7 in the crude methanol. The number of bubble trays em-
8 ployed will vary depending upon the desired purity of
9 the refined methanol, the pressure used in the column,
the amount of heat supplied to the column and other fac-
11 tors.
12 The partially refined methanol may then be sent
13 to a refining column. This column operates at low pres-
14 sure, for example, about 30 psig and separates methanol
from water and high boiling organic compounds. The num-
16 ber of bubble trays employed in the refining column may
17 also vary.
18 The degree of purity required forthe product
19 methanol obviously depends on the final product disposi-
tion. If the methanol is to be used merely as a fuel, ;
21 very little, if any, purification may be required. It
22 may be desirable to stabilize the product to eliminate
23 highly volatile dimethyl ether, but a rerunning operation ;~
24 to fractionate out water is usually not required in the -
process of this invention. This i8 one of the advantages
26 of this particular combination since the acid gas scrub-
27 bing eliminated the bulk of the C02 ln the methanol feed - ~ ;
28 gas. Since it is C02 which gives rise to water in the
methanol synthesis reactor according to equation (b) shown
:` .'
~7 - : -
10415S3
1 above, very little water is formed in the converter.
2 Consequently, the product will be sat~sfactory as a fuel
3 without the expensive and heat-consuming methanol-water
4 distlllation.
In conventional methanol plants based on synthesis
6 gas from steam reforming, the C02 content of the synthe-
7 sis gas is always high enough to yeild 10-20~ water or
8 methanol in the converter. Acid gas removal is not re-
9 quired and not done in these conventional methanol plants.
However, it is required in synthetic natural gas plants `~
11 in order to make a high BTU product. This is one of the
12 advantages of the present combination where a process
13 step and requirement in one product helps and influences
14 advantageously the production of the other.
The synthesis gas which exits the methanol syn-
16 thesis reaction zone contains hydrogen, carbon monoxide,
17 methane and a small amount of carbon dioxide, and this
18 gas is then sub~ected to methanation in order to convert -
19 the remaining carbon oxide to methane. This conve~sion
involves the reaction of carbon monoxlde and dioxide with
21 hydrogen in a molar ratio of about 1:3 (or 1:4) respect-
22 ively. The temperatures for such conversion will vary; -
23 however, these temperatures range from about 400F. to `
24 about 800F. overall. The inlet temperature in the metha-
nation chamber will vary, for example, from about 400F.,
26 to about 600F. and the outlet temperature from about
27 600F. to about 750F. The pressure at which methanation
28 i8 accomplished ranges from about 250 to about 1500, pre- ~
29 ferably 300 to about 1000, psig. The catalyst which may ,
be used in carrying out the methanation may be a partlal-
;
~ , .
1041S~3
1 ly reduced nickel oxide catalyst or any other suitable
2 catalyst. Such catalysts are well known in the art and
3 the invention is not to be construed as limited to any
4 particular methanation catalyst or any particular temper-
atures and pressures.
6 The various chambers or zones are connected to
7 one another via a suitable arrangement of pipes and
8 valves interspersed in the proper places.
9 Moreover, the ratio for a given methanol produc-
rion rate to a given synthetic natural gas production
11 rate may be varied to produce more methanol and less
12 synthetic natural gas or more synthetic natural gas and
13 less methanol. It is even possible to produce methanol
14 or synthetic natural gas beyond the rated capacity of -~
the methanol or synthetic natural gas facilities by elimi- --16 nating or reducing the production of methanol or minimi~
17 zing the production of synthetic natural gas.
18 If it is desired to produce no methanol and only
19 synthetic natural gas from the integrated facility, then
the process is modified by not introducing the purified
21 converted gas to the methanol synthesis zone. This is --
22 easily done by a system of valving so that gas from the ~ ~ ;
23 acid gas removal zone is discharged to the methanation
24 zone. On the other hand, methanol production can be maxi-
mized by providing recycle for gas around the methanol
26 converters or by providing additional reactor/condensing
27 stages.
28 Such versatility in the integrated methanol syn- `
29 thetic natural gas manufacturing facility of this inven- ~
30. tion is particularly advantageous because the manufactur- - -
: ,, ,
- 19 _ ,~ . ,
. ,~. ,~ . . . .
-- -- . . - . .,, - . - ~. . . . .
~O~lSS3
1 ing facility and process are responsive to chanees ln
2 economic factors. Thus, during period~ of peak gas de-
3 mand, it may not be deslrable to make methanol, which is
4 easily achieved by by-passing the methanol reactors.
Since they only constitute a small portion of the total
6 coal-gasification plant, very little debit is inc~rred
7 in operating the methanol synthesis portion only part-
8 time.
9 Conversely, during periods of low gas demand,
conversion to methanol can be maximized without decreasing
11 the synthesls gas generation portion ofthe plant. Some
12 gas recycle around the methanol converters, or providing
13 spare methanol synthesis reactors for such a purpose can
14 be done.
. In line with the a~ove flexibility, methanol syn- -
16 thesis can be considered a fly-wheel to permit steady
17 operation in a synthetic gas plant, indepentent of the
18 gas demand. If so desired, conversion of stored methanol
19 product to gas during periods of exceedingly high gas de-
mand can be considered as another peak-showing advantage
21 of the present combination. Additionally because of the
22 unlty of the manuracturing facility of this invention,
23 costly equipment, such as duplicate steam generating fac-
24 ilities, a carbon dioxide compressor, methanol synthesis
loop or recycle, and associated equipment which normally
26 would have to be included in separate manufacturing fac- ~;
27 llities are elimlnated.
28 EXAMPLE
29 Utilizing the process conditions set forth in
Tables I and II for bituminous coal set out below and
.
10415S3
l with reference to the present invention drawing and
2 process conditions, a typicaI analysis of the ga~ at
3 each step of the present invention i8 shown in Table
4 III: ;
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1041~53
1 Reference to Table I shows that the amount of
2 methanol made t5,200 mols/hr) represents about 20% of the
3 total carbon value out of the coal easification plant
4 (22,000 mols/hr of methane plus 5,200 mols/hr of methanol).
The amount of C02 which has to be removed in the C02
6 scrubber is 14,600 mols/hr, and the shift converter duty
7 is 4,200 mols/hr of C0 shi~ted to H2and C02.
8 Now, if methanol were not co-produced in this
9 particular gasi~ication plant, the amount of shifting
would have to be increased to 5,900 mols/hr of C0 and the
11 C02 scrubbing load would increase to 16,300 mols/hr in
12 order to achieve the 3/1 H2/C0 ratio ~or the methanation
13 step.
14 Therefore, it is clear from the above that as a
result.of the methanol/synthetic gas combination both the
16 shift duty and the C02 scrubbing duties are decreased over
:: . . .
17 those required when methane is the only product. ` - -
?8 The decrease in methane production between the -
19 synthetic gas only plant and the combination process o~
the present invention i6 about 3,600 mols/hr. If the
21 5,200 mols/hr o~ methanol produced instead are utilized
22 as ~uel, there is a considerable increase (15% approxi-
23 mately) in fuel value of the incremental product.
24 A further improvement in the synthetic gas process,
brought about by the incorporation of methanol manufacture `
26 into the overall process scheme, can be found in the meth-
27 anation section. As i8 apparent from Table I, in the com-
28 bination procesæ, the methanator ha~ to convert 6,200
29 mols/hr of carbon oxides (mostly C0) into methane, in the
presence o~ a diluent of about 16,000 mols/hr of already
'` .
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'104i~3
1 existing methane which will act as a heat sink ln the
2 highly exothermic methanation reaction. On the other
3 hand, when there is no co-production of methanol, the
4 6ame amount of already existing methane, i.e., 16,000
mols/hr, will have to act as heat sink for the methanation
6 of 9,800 mols/hr of carbon oxides (again mostly CO), a
7 much more æevere heat load. Consequently, the methanol
8 co~production will have a favorable effect on the design
9 of the methanation section of the plant.
Clearly, it can be seen from the above that the
11 co-production of methanol results in ma~or improvements
12 in the shifting, acid gas scrubbing and methanation sec-
13 tion of the synthetic natural gas plant. In addition,
14 as was explained earlier, the co-production of a gas and
liquid fuel provides flexibility in the overall plant op-
16 eration and therefore simplifies the design of all sec-
17 tions so as to meet fluctuating demands of the gas product.
18 . In addition, the combination process possesses
19 distinct advantages for the methanol process over a plant
having methanol as the sole product. The advantages of
-21 minimizing the water produced in the methanol reactor,
22 and the resultant simplification in product purification
23 equipment was described earlier.
24 A further and significant improvement is the
elimination of the recycle compressor slnce a high degree
26 of conversion of the carbon oxides and hydrogen to metha-
27 nol i8 not required. Any unconverted CO, C02 and H2
28 passes through the subsequent methanator with the result
29 of converting it to the other, highly useful co-product
of the plant; methane. Also, this once through rather
- 32 -
- .. - . ; . - ~ :. . ... . .
1041S~3
1 than recycle methanol reactor scheme can tolerate a ~uch
2 higher content o~ inerts, specifically methane, ethane
3 and nitrogen, in the feed gas than a recycle process.
4 In a conventional methanol plant with recycle, the inert
level builds up rapidly and would quickly choke the re-
6 action unless a high, and uneconomical, purge rate were
7 imposed.
8 A further advantage in the methanol synthesis
9 of the combination process over a plant having methanol
as the sole product is the required selectivity of the
11 catalyst of methanol versus methane formation. In a meth-
12 anol-only plant, methane formation is highly undesirable -
13 since it rapidly builds up as an inert in the recycle gas
14 leading to uneconomically high purge rates from the synthe-
6iS gas loop. Thus,,valuable feed gas is lost not only
16 by conversion to methane, but also in the excess purge
..
17 which must be removed from the recycle loop to keep inert
18 level in the reactor at a predetermined level. In con-
19 trast to this, a modest amount of methane made in the
methanol reactors, coupled with adequate heat removal for
21 the high heat of reaction of methanation, i9 not undesir-
22 able in the combination process and may actually be helpful.
23 It will unload the subsequent methanation reactor. Since
?4 most methanol catalysts exhibit increasing methanatlon
activity with increasing temperature and catalyst age,
26 the production o~ up to 20%, on a molar basis, of methane
27 compared to methanol in the methanol reactor allows longer
28 catalyst life runs and a wider temperature range than
29 conventional operation.
The methanation activity of methanol catalysts
~ 33 ~
,. - , . . . .. , , : . .. : . . : ~.
10415~3
1 at increased temperature can be u~ed advantageously in
2 the operation of this combinatlon plant during high ga~
3 demand periods. By increasing the temperature in the
4 methanol reactors from 400-600F. up to 700-800F., they
can be used as pre-methanators to take some of the meth-
6 anation load ofr the subsequent methanation zone.
7 In view of the analyses set forth above, the
8 present invention integrated process uniquely provides
9 for the concurrent production Or methanol and synthetic
lo natural gas. Furthermore, the costs of this integrated
11 process are ~ubstantially less than the combination Or
12 two separate process racilities.
13 The present invention has been described herein
14 with reference to particular embodi~ents thereof. It
will be appreciated by those skilled in the art, however,
16 that various changes and modifications can be made there-
17 in without departing from the scope Or the invention as
18 presented.
19 While the process description Or the specific
embodiment described herein covered the hydrogasification
21 of coal, and the conversion of the resultant gas into
22 both methane, for synthetic natural gas, and methanol,
23 the present invention 6hould in no way be limited to
24 such a combination. Other suitable gasirication process-
es Or coal or heavy residue include electric heating in
26 the preæence Or 6team, partial oxidation with air, oxygen
27 or both, coking Or heavy petroleum liquids with oxidative
28 heating, and gasirication of the produced coke, catalytic
29 gasification of heavy petroleum stocks in the presence Or
steam and an oxygen-containing gas, partial oxidation Or
'~
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.. -., . . , . ~ ~ ~ .. .. . .
lO~lS~
1 coke, and any processes which produce gas mixtures con-
2 taining substantial amounts of carbon monoxide, methane
3 and hydrogen. In any of these processes the co-production
4 of methanol and methane by the instant invention from the
crude synthesis gas has many advantages over the manufac- -
6 ture of either one or the other. The process is a~so
7 suitable when operated on the regenerator gas from many
8 petroleum processes where a circulating solid or catalyst
9 is heated and residual carbon removed by combustion with
an oxygen-containing gas. :
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