CA1041553A - Methanol and synthetic natural gas concurrent production - Google Patents
Methanol and synthetic natural gas concurrent productionInfo
- Publication number
- CA1041553A CA1041553A CA203,979A CA203979A CA1041553A CA 1041553 A CA1041553 A CA 1041553A CA 203979 A CA203979 A CA 203979A CA 1041553 A CA1041553 A CA 1041553A
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- CA
- Canada
- Prior art keywords
- methanol
- gas
- reaction zone
- methane
- zone
- Prior art date
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- Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
- Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
- Y02E—REDUCTION OF GREENHOUSE GAS [GHG] EMISSIONS, RELATED TO ENERGY GENERATION, TRANSMISSION OR DISTRIBUTION
- Y02E50/00—Technologies for the production of fuel of non-fossil origin
- Y02E50/30—Fuel from waste, e.g. synthetic alcohol or diesel
-
- Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
- Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
- Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
- Y02P20/00—Technologies relating to chemical industry
- Y02P20/50—Improvements relating to the production of bulk chemicals
- Y02P20/582—Recycling of unreacted starting or intermediate materials
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- Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)
Abstract
ABSTRACT OF THE DISCLOSURE
Methanol and synthetic natural gas are concurrently produced by operating sequentially a carbonaceous material gasification zone, in series with a (water gas) shift conversion zone, in series with a sulfur compound and carbon dioxide removal zone, in series with a methanol synthesis zone and in series with a methanation zone.
Methanol and synthetic natural gas are concurrently produced by operating sequentially a carbonaceous material gasification zone, in series with a (water gas) shift conversion zone, in series with a sulfur compound and carbon dioxide removal zone, in series with a methanol synthesis zone and in series with a methanation zone.
Description
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 .. . .
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 .. . .
-2- ~
, ~ . .
- ~ , t ;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 .
."
, ~ . .
- ~ , t ;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 ~
'.' ' .'' ."
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 -' ' '' ~ ,, ` ' `
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 :
'.' ' .'' ."
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 -' ' '' ~ ,, ` ' `
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 ~ ~
~ .
~ . ,,:
..'' ' '' -4a-,,~'~ ' ' . .
" . , .. . : ,. j.~ ., ,. ,, ., .,, . ,., ~- :. . ..
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 ..
. .
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, _ ~ _ . . .
', ~,,1 , .~i .; . .
: . ., ,,~ ,~. ~ , . .
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 -. ~ . . . . ., . . ~ . , , 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 - :
,.: ` - . .. - . ~ :
. . . . . .
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 - : -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.
29 Utilizing the process conditions set forth in Tables I and II for bituminous coal set out below and .
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: ;
': ~' . .
:, , ': .
. .
~, - .
, ~ ~
!' "'` ' ' "`"~
- ` ' -- 21 ~
~ .
Lr~o L~ co ......
Lr~ O ~1 ~ Lt~ l C`J ~ 0 ~ ~ o cu m ~ - .
, ~CQ ~ ~ , , ¢ ~ ~ c, o ~! P; ~ o o E~ ~ O O oc~ LS~ o h ~ o~ .Lr~
¢ ~ ~r~ ~ )O L~ L ~ O O ', . .
~3 0 ~ ~I I I I ' I I r-l ~N::I'~ N
.) ~t ~3 O OLr~) L~ L~ I I I I I I
O ~ L~O ~O O L~O O O
P; ~~D _l O O O
methanol ~ ~
~ .
~ . ,,:
..'' ' '' -4a-,,~'~ ' ' . .
" . , .. . : ,. j.~ ., ,. ,, ., .,, . ,., ~- :. . ..
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 ..
. .
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, _ ~ _ . . .
', ~,,1 , .~i .; . .
: . ., ,,~ ,~. ~ , . .
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 -. ~ . . . . ., . . ~ . , , 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 - :
,.: ` - . .. - . ~ :
. . . . . .
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 - : -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.
29 Utilizing the process conditions set forth in Tables I and II for bituminous coal set out below and .
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
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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 '` .
~ 3~
: ,- . ~ - ~ . .
~ . . . . . ..
.. - .. - . . .. - ... . .,. ~ . ,. . : : .. , , ,: .
'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 - .. - . ; . - ~ :. . ... . .
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 '~
.': ~ .
_ ~4 -.. -., . . , . ~ ~ ~ .. .. . .
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. :
~.~
~' ' ~ ' "~ '"'.
. ' .
' ~ .
. - 35 -~' , , .
. - . .. . . .. . .. .
- . . . . . , .. - .. - ..
. . . . . . . .
- ., ~
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 '` .
~ 3~
: ,- . ~ - ~ . .
~ . . . . . ..
.. - .. - . . .. - ... . .,. ~ . ,. . : : .. , , ,: .
'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 - .. - . ; . - ~ :. . ... . .
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 '~
.': ~ .
_ ~4 -.. -., . . , . ~ ~ ~ .. .. . .
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. :
~.~
~' ' ~ ' "~ '"'.
. ' .
' ~ .
. - 35 -~' , , .
. - . .. . . .. . .. .
- . . . . . , .. - .. - ..
. . . . . . . .
- ., ~
Claims (19)
IS CLAIMED ARE DEFINED AS FOLLOWS:
1. 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 and 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 950°F 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 producing a converted gas stream with the proviso that the shift conversion is controlled to such an extent as to provide the necessary stoichiometry 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 major 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 are converted, in the presence of a methanol conversion catalyst and at a sufficient temperature and at pressures ranging from 400 to about 1500 psig, 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 contain-ing the remaining carbon oxides, hydrogen, and the methane produced in the first reaction zone from said 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.
a. gasifying the carbonaceous material in a first reaction zone at sufficient pressure and 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 950°F 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 producing a converted gas stream with the proviso that the shift conversion is controlled to such an extent as to provide the necessary stoichiometry 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 major 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 are converted, in the presence of a methanol conversion catalyst and at a sufficient temperature and at pressures ranging from 400 to about 1500 psig, 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 contain-ing the remaining carbon oxides, hydrogen, and the methane produced in the first reaction zone from said 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.
2. The process of claim 1 wherein said fourth reaction zone comprises at least two methanol converters in series, each converter having a methanol removal means thereafter in series.
3. The process of claim 1 wherein the carbonaceous material is atmospheric residuum, vacuum residuum, or mixtures thereof.
4. The process of claim 1 wherein the carbonaceous material is coke.
5. The process of claim 1 wherein the carbonaceous material is cellulosic material.
6. The process of claim 1 wherein the carbonaceous material is municipal waste.
7. The process of claim l wherein the gasification in the first reaction zone is a partial oxidation of said carbonaceous material.
8. The process of claim 1 wherein the gasification in the first reaction zone is the gasification of said carbonaceous material in the presence of steam, an oxygen-containing gas and a particulate catalyst containing an alkali metal component, a-solid support and an in situ formed carbonaceous deposit on said support.
9. The process of claim 1 wherein the reaction conditions in said fourth reaction zone are such as to provide that not more than approximately one-fifth of the carbon oxides reacted are converted to methane and approximately four-fifths of said carbon oxides are converted to methanol.
10. The process of claim 1 wherein the molar ratio of carbon monoxide to hydrogen in the synthesis gas discharged from the gasification zone is at least 1:2.3 and said gas is passed directly to the acid gas removal zone.
11. An integrated process for the production of methanol and a synthetic natural gas, which is substantially methane, from a carbonaceous material selected from the group consisting of coal, residue or mixtures thereof which comprises the steps of:
a. gasifying the carbonaceous materials in a first reaction zone at sufficient pressure and temperature to produce a synthesis gas stream comprising methane, carbon monoxide, carbon dioxide, hydro-gen, and steam, in which the synthesis gas contains about 10-30 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 of between about 200° and 950° F 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 producing a converted gas stream with the proviso that the shift conversion is controlled to such an extent as to provide the necessary stoichiometry 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 substantially all 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 are converted, in a once-through process, in the presence of a methanol conversion catalyst and at a sufficient temperature and at pressures ranging from 400 to about 1500 psig, 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 methane product in the first reaction zone from said 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 sufficient temperature and pressure, to methane.
a. gasifying the carbonaceous materials in a first reaction zone at sufficient pressure and temperature to produce a synthesis gas stream comprising methane, carbon monoxide, carbon dioxide, hydro-gen, and steam, in which the synthesis gas contains about 10-30 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 of between about 200° and 950° F 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 producing a converted gas stream with the proviso that the shift conversion is controlled to such an extent as to provide the necessary stoichiometry 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 substantially all 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 are converted, in a once-through process, in the presence of a methanol conversion catalyst and at a sufficient temperature and at pressures ranging from 400 to about 1500 psig, 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 methane product in the first reaction zone from said 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 sufficient temperature and pressure, to methane.
12. The process of claim 1 wherein the carbonaceous material is selected from the group consisting of coal, residue and mixtures thereof.
13. The process of claim 1 wherein the carbonaceous material is coal.
14. The process of claim 2 wherein the first reaction zone comprises two sections, the first section wherein steam and char are reacted together to provide a synthesis gas and heated solids, and the second section wherein coal is concurrently passed with the synthesis gas and heated solids through said second section which is a transfer line hydrogasification zone.
15. The process of claim 11 wherein the carbonaceous material is coal.
16. The process of claim 15 wherein the first reaction zone comprises two sections, the first section wherein steam and char are reacted together to provide a synthesis gas and heated solids, and the second section wherein coal is concurrently passed with the synthesis gas and heated solids through salt second section which is a transfer line hydrogasification zone.
17. The process of claim 11 wherein said fourth reaction zone comprises at least two methanol converters in a series, each converter having a methanol removal means thereafter in series.
18. The process of claim 11 wherein the gasification in the first reaction zone is the gasification of said carbonaceous material in the presence of steam, an oxygen-containing gas and a particulate catalyst containing an alkali metal component, a solid support and an in situ formed carbonaceous deposit on said support.
19. The process of claim 11 wherein the molar ratio of said carbon monoxide to hydrogen in a synthesis gas discharged from the gasification zone is at least 1:2.3 and said gas is passed directly to the acid gas removal zone.
Applications Claiming Priority (1)
| Application Number | Priority Date | Filing Date | Title |
|---|---|---|---|
| US38406173A | 1973-07-30 | 1973-07-30 |
Publications (1)
| Publication Number | Publication Date |
|---|---|
| CA1041553A true CA1041553A (en) | 1978-10-31 |
Family
ID=23515872
Family Applications (1)
| Application Number | Title | Priority Date | Filing Date |
|---|---|---|---|
| CA203,979A Expired CA1041553A (en) | 1973-07-30 | 1974-07-03 | Methanol and synthetic natural gas concurrent production |
Country Status (3)
| Country | Link |
|---|---|
| JP (1) | JPS5041803A (en) |
| CA (1) | CA1041553A (en) |
| ZA (1) | ZA744415B (en) |
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| US4534772A (en) * | 1982-04-28 | 1985-08-13 | Conoco Inc. | Process of ether synthesis |
| US4751248A (en) * | 1987-07-02 | 1988-06-14 | Phillips Petroleum Company | Preparation of alcohols from synthesis gas |
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-
1974
- 1974-07-03 CA CA203,979A patent/CA1041553A/en not_active Expired
- 1974-07-09 ZA ZA00744415A patent/ZA744415B/en unknown
- 1974-07-30 JP JP49086689A patent/JPS5041803A/ja active Pending
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Also Published As
| Publication number | Publication date |
|---|---|
| AU7153274A (en) | 1976-01-29 |
| ZA744415B (en) | 1975-07-30 |
| JPS5041803A (en) | 1975-04-16 |
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