CA2213025C - Process of producing methanol from natural gas - Google Patents
Process of producing methanol from natural gas Download PDFInfo
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- CA2213025C CA2213025C CA002213025A CA2213025A CA2213025C CA 2213025 C CA2213025 C CA 2213025C CA 002213025 A CA002213025 A CA 002213025A CA 2213025 A CA2213025 A CA 2213025A CA 2213025 C CA2213025 C CA 2213025C
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- C—CHEMISTRY; METALLURGY
- C07—ORGANIC CHEMISTRY
- C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
- C07C29/00—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring
- C07C29/15—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by reduction of oxides of carbon exclusively
- C07C29/151—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by reduction of oxides of carbon exclusively with hydrogen or hydrogen-containing gases
- C07C29/1516—Multisteps
- C07C29/1518—Multisteps one step being the formation of initial mixture of carbon oxides and hydrogen for synthesis
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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
- 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
- Y02P20/52—Improvements relating to the production of bulk chemicals using catalysts, e.g. selective catalysts
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Abstract
Natural gas is converted with steam and oxygen in an autothermally operated reactor, and from the reactor a raw synthesis gas with a temperature in the range from 800 to 1200°C is withdrawn. To the cooled, raw synthesis gas a gas rich in H2 is admixed. The synthesis gas with an increased H2 content is converted in a methanol synthesis reactor on a copper catalyst at a pressure in the range from 10 to 150 bar and temperatures from 180 to 350°C. From the synthesis reactor a gas mixture containing methanol vapor is withdrawn, methanol is condensed out, and a cooled residual gas is separated. The residual gas is divided into a first partial stream (recycle gas) and a second partial stream (purge gas).
The recycle gas is admixed to the synthesis gas before the methanol synthesis reactor. The purge gas consists of 30 to 80 vol-% H2, 7 to 35 vol-% CO2 and 3 to 25 vol-% CO. The hourly quantity of purge gas is 8 to 25 % (calculated dry) of the hourly quantity of raw synthesis gas. The purge gas is charged into a separating means, where the gas rich in H2 is recovered, which is admixed to the raw synthesis gas. The hourly quantity of carbon dioxide supplied to the separating means in the purge gas is 10 to 50 % of the hourly quantity of CO2 in the raw synthesis gas.
The recycle gas is admixed to the synthesis gas before the methanol synthesis reactor. The purge gas consists of 30 to 80 vol-% H2, 7 to 35 vol-% CO2 and 3 to 25 vol-% CO. The hourly quantity of purge gas is 8 to 25 % (calculated dry) of the hourly quantity of raw synthesis gas. The purge gas is charged into a separating means, where the gas rich in H2 is recovered, which is admixed to the raw synthesis gas. The hourly quantity of carbon dioxide supplied to the separating means in the purge gas is 10 to 50 % of the hourly quantity of CO2 in the raw synthesis gas.
Description
Process of Producing Methanol from Natural Gas This invention relates to a process of producing methanol from natural gas, the natural gas is catalytically converted with steam and molecular oxygen, and a raw synthesis gas is withdrawn from the catalytic conversion, which rw synthesis g~
contains H2, CO and C02, a gas rich in H2 and recycle gas are admixedto the raw synthesis gas and a synthesis gas mixture is pro-duced, which is passed through at least one methanol synthe-sis reactor which contains a copper catalyst and is operated at a pressure in the range from 10 to 150 bar and at tempera-tures in the range from 180 to 350°C, from the methanol syn-thesis reactor a gas mixture containing methanol vapor is withdrawn, which is cooled to such an extent that methanol condenses out, a cooled residual gas is separately withdrawn from the methanol and divided into the recycle gas and a sec-ond partial stream (purge gas), from the purge gas, which has a HZ content of at least 30 vol-% (calculated dry), a gas rich in H2 is separated in a separating means and admixed to 3D the raw synthesis gas.
Such process is known from EP-B-0 233 076. The natural gas or a partial stream of the natural gas is first of all passed through a steam reforming, and the cracking gas thus produced together with the residual natural gas is then charged into an autothermally operated reactor which contains a catalyst.
. _ 2 _ Steam reforming is effected as usual in a tubular furnace, where a nickel catalyst is provided in a plurality of tubes, which are disposed in a combustion chamber and are heated from the outside by hot combustion gas. The steam reforming in accordance with the known process serves to increase the hydrogen content in the succeeding autothermal reactor and also to provide for an increased hydrogen content in the raw synthesis gas produced.
The object underlying the invention is to modify the known process such that the process can be performed with low in-vestment costs, and at the same time a favorable consumption of process means is achieved. It should also be possible to work without steam reforming. In accordance with the inven-tion this is accomplished in the above-mentioned process in that (a) the catalytic conversion of natural gas with steam and oxygen is performed only in the autothermally operated reactor, which comprises at least one burner fed with the natural gas and the oxygen, and below the burner a bed of granular catalyst;
(b) the raw synthesis gas is withdrawn from the autothermally operated reactor at a temperature in the range from 800 to 1200°C, is cooled, steam is condensed out, and conden-sate produced is separated;
(c) the raw, cooled synthesis gas has a stoichiometric number S of 1.3 to 1.9 prior to admixing the gas rich in H2, where S is calculated from the molar concentrations of H2, CO and Co2 in accordance with S = (H2 - C02) . (CO + C02);
_ 3 _ (d) the purge gas consists of 30 to 80 vol-% H2, 7 to 35 vol-%, and preferably at least 10 vol-% COZ and 3 to 25 vol-% CO, (e) the hourly quantity of purge gas is 8 to 25 % vol-%, and preferably 10 to 15 vol-% (calculated dry) of the hourly quantity of cooled, raw synthesis gas;
(f) the hourly quantity of CO2 supplied to the separating means in the purge gas is 10 to 50 %, and preferably 15 to 35 % of the hourly quantity of C02 in the raw synthe-sis gas; and (g) the hourly quantity of Co supplied to the separating means in the purge gas is 2 to 10 % of the hourly quan-tity of CO in the raw synthesis gas;
(h) the gas rich in H2, which was withdrawn from the separat-ing means and admixed to the raw synthesis gas, has a H2 content of at least 80 vol-%, and by admixing H2 a stoi-chiometric number S of 1.95 to 2.2 is achieved for the synthesis gas.
In the process in accordance with the invention it is ensured that the purge gas represents a relatively large gas stream.
This is achieved in that the methanol synthesis reactor is operated such that apart from the methanol vapor a large re-sidual gas stream can continuously be withdrawn. It is delib-erately omitted to achieve an optimally high consumption of H2, CO and C02 in the synthesis reactor. In particular the consumption of C02 is reduced, so that the consumption of H2 is also reduced considerably. In the process in accordance with the invention it is ensured that the purge gas quantity contains at least the amount of hydrogen required for the necessary increase of the stoichiometric number of the raw ' - 4 -synthesis gas. In the methanol synthesis reactor the neces-sary amount of catalyst decreases, so that the methanol syn-thesis reactor can be designed smaller. It is also possible to operate the reactor at a relatively low pressure or with a low recycle ratio. Apart from these advantages considerably reducing the costs of the plant it should be noted that the process is also conducted without CO conversion and without C02 removal in the raw synthesis gas. Since steam reforming is omitted in accordance with the invention, it is definitely possible to employ a higher pressure in the autothermally op-erated reactor. Since the pressure in the methanol synthesis reactor can at the same time be relatively low, it is thus possiblelto largely adjust the pressures in the gas produc-tion and in the synthesis.
The autothermally operated reactor, which operates without indirect heating of the catalyst, is known and described for instance in Ullmann's Encyclopedia of Industrial Chemistry, 5th edition, Vol. A 12, pp. 202 to 204. The natural gas is supplied to the reactor preferably preheated. Usually, tech-nically pure oxygen is supplied to the burner of the reactor, in order to keep the content of inert gas in the raw synthe-sis gas as low as possible. The supply of steam usually lies in the range from 0.5 to 3.0 mol, based on the molar carbon content of the natural gas. The autothermally operated reac-tor can have a one-part or a multistage design. Instead of natural gas, the autothermally operated reactor can also be charged with other charges, e.g. with liquefied gas or refin-ery gas.
For the methanol synthesis reactor known constructions are considered, in particular the water-cooled tubular reactor or the adiabatically operated fixed-bed reactor.
The separating means for producing the gas rich in H2 is likewise designed in a manner known per se, e.g, as pressure-' - 5 -swing adsorption plant or low-temperature gas separation plant.
Embodiments of the process will now be explained with refer-ence to the drawing. The drawing represents a flow diagram of the process.
The autothermally operated reactor 1 comprises a bed 2 of granular nickel catalyst. Above the bed, a burner 3 is dis-posed, to which desulfurized natural gas is supplied through line 4, gas rich in 02 is supplied through line 5, and steam is supplied through line 6. As gas rich in o2 there is nor-mally used technically pure oxygen. The heat required for the conversion in the reactor 1 is only applied through partial oxidation. The temperatures at the outlet 8 of the reactor 1 lie in the range from 800 to 1200°C, and usually are at least 900°C.
Hot, raw synthesis gas leaves the reactor 1 and flows through a cooling 9, which may have a multistage design, which is not represented in detail in the drawing. Upon cooling, an aque-ous condensate is produced, which is withdrawn via line 10.
The cooled, raw synthesis gas, which flows through line 11, usually has temperatures in the range from 20 to 80°C. This gas is compressed by a compressor 12. Through line 14, gas rich in H2 is supplied, which comes from a separating means 15. The H2 content of the gas in line 14 is at least 80 vol-%, and preferably at least 90 vol-%.
The stoichiometric number S of the gas in line 11 lies in the range from 1.3 to 1.9, and mostly is not more than 1.8. By admixing the gas rich in H2 of line 14, the stoichiometric number of the synthesis gas in line 16 is increased to 1.95 to 2.2, as this is necessary for methanol synthesis. S is calculated in the known manner from the molar concentrations of H2, CO and C02 in accordance with S = (H2 - C02) . (CO + C02)~
To the gas in line 16, recycle gas is admixed via line 18, and the synthesis gas mixture thus formed is first of all supplied to a heat exchanger 20 via line 19, before it enters the methanol synthesis reactor 22 through line 21.
In the drawing, the reactor 22 is represented as known tubu-lar reactor, where the copper catalyst is provided in a plu-rality of tubes 23. For the indirect cooling of the interior of the tubes 23, cooling water is supplied to the reactor through line 24, and heated and partially evaporated cooling medium is withdrawn via line 25. In the tubes 23, the tem-peratures lie in the range from 180 to 350°C, and mostly in the range from 200 to 300°C. The pressure in the tubes is 10 to 150 bar, and mostly lies in the range from 20 to 100 bar.
From the reactor 22 a gas mixture containing methanol vapor is withdrawn via line 26, which gas mixture is first of all cooled in the heat exchanger 20 and is then supplied to an indirect cooler 28 via line 27. Via line 29, the mixture is supplied to the condenser 30, from which a condensate con-taining methanol and water is withdrawn via line 31. The fur-ther distillative treatment of this condensate, which is known per se, need not be explained in detail.
From the condenser 30, cooled residual gas is withdrawn through line 32 and is divided into two gas streams. The first stream is recirculated as recycle gas through line 18.
The second stream, here referred to as purge gas, is supplied to the separating means 15 via line 33, which separating means operates for instance according to the principle of pressure-change adsorption. In the separating means 15 an ex-haust gas containing methane, CO and C02 is produced in addi-tion to the gas rich in H2 of line 14, which exhaust gas is withdrawn via line 34 and can be used as combustion gas or be supplied to some other use.
For admixing the gases of lines 14 and 18 to the synthesis gas, a compressor will be required, which was, however, omit-ted in the drawing for simplification. The reactor 1 can be operated at a pressure above 40 bar, where the compressor 12 can be omitted.
Examples:
The following examples, which are calculated in part, are di-rected to the production of 1000 tons methanol per day. The nickel catalyst in the reactor 1 consists of 15 wt-% nickel on a carrier containing A12o3. The copper catalyst for the methanol synthesis contains 60 wt-o copper, 30 wt-o zinc ox-ide and 10 wt-o aluminium oxide and operates at a pressure of 80 bar. The separating means 15 is a pressure-swing adsorp-tion plant, for which a H2 yield of 85% is assumed.
Example 1:
There is used a plant corresponding to the drawing; the natu-ral gas in line 4 has been preheated to 500°C, just as the technically pure oxygen in line 5 and the steam in line 6.
Per hour, 1472 kmol natural gas, 1946 kmol steam, and 850 kmol 02 are supplied to the burner 3. In addition to methane, the natural gas contains 1.5 mol-o C2H6 and 0.4 mol-% C3Hg +
C4Hlo. Further details are shown in the following Table I, where in different lines the components of the respective mixture are indicated in mol-% and kmol/h:
_ 8 Table I
Line I 8 I 14 I 16 I 33 I 34 mol-% kmol/hmol-% mol-% mol-% kmol/hkmol/h COZ 6.1 383 - 8 12 60 60 CO 17 1070 - 22.4 4.9 24 24 H2 47.3 2982 100 69 73.4 365 56 CH4 0.8 50 - 1 8.9 44 44 H20 28.8 1821 - _ _ - _ CH30H - - - - 0.3 2 2 Total amo~lnt kmol/h6310 310 4800 498 188 Tem erature C 1000 35 35 35 35 Pressure bar 35 83 82 7g 5 The stoichiometric number S of the raw synthesis gas in line 8 is 1.79, and for the gas in line 16 S is 2Ø
Example 2_ The pressure in the reactor 1 is now 50 bar, so that the com-pressor 12 can be omitted. There is used the same natural gas as in Example 1, the temperature in lines 4, 5 and 6 is each 600°C. Per hour, 1498 kmol natural gas, 1982 kmol steam and 800 kmol technically pure oxygen are supplied to the burner 3. Further data are listed in Table II:
TnlW n, T T
Line 8 14 16 33 34 mol-% kmol/h mol-% mol-% mol-% kmol/h kmol/h COZ 5.9 372 -- 7.8 10.7 50 50 CO 16.7 1057 -- 22.2 6.3 30 30 HZ 47.3 2993 100 67.8 61.6 289 54 CH4 1.7 103 -- 2.2 20.8 95 95 H20 28.4 1794 -- -- 0.1 -- --CH30H __ __ __ __ 0.5 2 2 Total amount 6233 235 4764 470 234 kmol/h Tem erature 1000 35 35 35 35 C
Pressure bar 50 53 52 49 5 S 1.83 2 1st Comparative Exam le (calculated):
For the production of 1000 tons methanol per day there is used the natural gas of Example 1, preheated to 500°C, whereof 360 kmol/h are delivered through a commercial tubular furnace for steam reforming to a commonly used nickel cata-lyst. Thus, a H2-containing primary gas with about 40 vol-H2 is produced, to which 1030 kmol/h natural gas are admixed.
The mixture is charged into an autothermally operated reac-tor, as it is solely used in the inventive process in accor-dance with the drawing. By adding 02 and steam, a raw synthe-sis gas is produced in a total amount of 5900 kmol/h, with a temperature of 1000°C and a pressure of 35 bar, where S =
1.95. Through line 14, 66 kmol/h hydrogen are admixed to this gas, and S is thus increased to 2Ø The amount of purge gas (line 33) is only 147 kmol/h, which is charged into a pres-sure-change adsorption plant. Through line 34, 81 kmol/h ex-haust gas are discharged, which is composed of (in kmol/h) 8.2 C02, 4.1 CO, 21.1 H2, 0.5 methanol, 2.4 N2, and 44.7 CH4.
It should be noted that in particular the quantity and compo-sition of the purge gas are considerably different from Exam-ples 1 and 2.
2nd Comparative Example (calculated):
There is used the natural gas of Example 1 and the process in accordance with the drawing. From the raw synthesis gas of line 11, 200 of the Co2 are removed, so that the stoichiomet-ric number is increased to 1.93. The different lines contain the following gas quantities and gas compositions:
Line 11 14 16 33 34 mol-% kmol/hmol-% mol-% mol-% kmol/h kmol/h COz 7.1 307 -- 6.9 6.1 9 9 CO 24.2 1050 -- 23.7 3 5 5 HZ 67.5 2924 100 68.2 65.8 98 3 CH4 1.1 49 -- 1.1 23.1 35 35 CH30H + Nz -- -- -- -- 2 3 3 Total amount 4330 95 4425 15 55 kmol/h Tem erature 35 35 35 35 35 C
Pressure bar 34 83 82 7g 5 S 1.93 2 Here as well, the composition and quantity of the purge gas in line 33 is considerably different from Examples 1 and 2.
contains H2, CO and C02, a gas rich in H2 and recycle gas are admixedto the raw synthesis gas and a synthesis gas mixture is pro-duced, which is passed through at least one methanol synthe-sis reactor which contains a copper catalyst and is operated at a pressure in the range from 10 to 150 bar and at tempera-tures in the range from 180 to 350°C, from the methanol syn-thesis reactor a gas mixture containing methanol vapor is withdrawn, which is cooled to such an extent that methanol condenses out, a cooled residual gas is separately withdrawn from the methanol and divided into the recycle gas and a sec-ond partial stream (purge gas), from the purge gas, which has a HZ content of at least 30 vol-% (calculated dry), a gas rich in H2 is separated in a separating means and admixed to 3D the raw synthesis gas.
Such process is known from EP-B-0 233 076. The natural gas or a partial stream of the natural gas is first of all passed through a steam reforming, and the cracking gas thus produced together with the residual natural gas is then charged into an autothermally operated reactor which contains a catalyst.
. _ 2 _ Steam reforming is effected as usual in a tubular furnace, where a nickel catalyst is provided in a plurality of tubes, which are disposed in a combustion chamber and are heated from the outside by hot combustion gas. The steam reforming in accordance with the known process serves to increase the hydrogen content in the succeeding autothermal reactor and also to provide for an increased hydrogen content in the raw synthesis gas produced.
The object underlying the invention is to modify the known process such that the process can be performed with low in-vestment costs, and at the same time a favorable consumption of process means is achieved. It should also be possible to work without steam reforming. In accordance with the inven-tion this is accomplished in the above-mentioned process in that (a) the catalytic conversion of natural gas with steam and oxygen is performed only in the autothermally operated reactor, which comprises at least one burner fed with the natural gas and the oxygen, and below the burner a bed of granular catalyst;
(b) the raw synthesis gas is withdrawn from the autothermally operated reactor at a temperature in the range from 800 to 1200°C, is cooled, steam is condensed out, and conden-sate produced is separated;
(c) the raw, cooled synthesis gas has a stoichiometric number S of 1.3 to 1.9 prior to admixing the gas rich in H2, where S is calculated from the molar concentrations of H2, CO and Co2 in accordance with S = (H2 - C02) . (CO + C02);
_ 3 _ (d) the purge gas consists of 30 to 80 vol-% H2, 7 to 35 vol-%, and preferably at least 10 vol-% COZ and 3 to 25 vol-% CO, (e) the hourly quantity of purge gas is 8 to 25 % vol-%, and preferably 10 to 15 vol-% (calculated dry) of the hourly quantity of cooled, raw synthesis gas;
(f) the hourly quantity of CO2 supplied to the separating means in the purge gas is 10 to 50 %, and preferably 15 to 35 % of the hourly quantity of C02 in the raw synthe-sis gas; and (g) the hourly quantity of Co supplied to the separating means in the purge gas is 2 to 10 % of the hourly quan-tity of CO in the raw synthesis gas;
(h) the gas rich in H2, which was withdrawn from the separat-ing means and admixed to the raw synthesis gas, has a H2 content of at least 80 vol-%, and by admixing H2 a stoi-chiometric number S of 1.95 to 2.2 is achieved for the synthesis gas.
In the process in accordance with the invention it is ensured that the purge gas represents a relatively large gas stream.
This is achieved in that the methanol synthesis reactor is operated such that apart from the methanol vapor a large re-sidual gas stream can continuously be withdrawn. It is delib-erately omitted to achieve an optimally high consumption of H2, CO and C02 in the synthesis reactor. In particular the consumption of C02 is reduced, so that the consumption of H2 is also reduced considerably. In the process in accordance with the invention it is ensured that the purge gas quantity contains at least the amount of hydrogen required for the necessary increase of the stoichiometric number of the raw ' - 4 -synthesis gas. In the methanol synthesis reactor the neces-sary amount of catalyst decreases, so that the methanol syn-thesis reactor can be designed smaller. It is also possible to operate the reactor at a relatively low pressure or with a low recycle ratio. Apart from these advantages considerably reducing the costs of the plant it should be noted that the process is also conducted without CO conversion and without C02 removal in the raw synthesis gas. Since steam reforming is omitted in accordance with the invention, it is definitely possible to employ a higher pressure in the autothermally op-erated reactor. Since the pressure in the methanol synthesis reactor can at the same time be relatively low, it is thus possiblelto largely adjust the pressures in the gas produc-tion and in the synthesis.
The autothermally operated reactor, which operates without indirect heating of the catalyst, is known and described for instance in Ullmann's Encyclopedia of Industrial Chemistry, 5th edition, Vol. A 12, pp. 202 to 204. The natural gas is supplied to the reactor preferably preheated. Usually, tech-nically pure oxygen is supplied to the burner of the reactor, in order to keep the content of inert gas in the raw synthe-sis gas as low as possible. The supply of steam usually lies in the range from 0.5 to 3.0 mol, based on the molar carbon content of the natural gas. The autothermally operated reac-tor can have a one-part or a multistage design. Instead of natural gas, the autothermally operated reactor can also be charged with other charges, e.g. with liquefied gas or refin-ery gas.
For the methanol synthesis reactor known constructions are considered, in particular the water-cooled tubular reactor or the adiabatically operated fixed-bed reactor.
The separating means for producing the gas rich in H2 is likewise designed in a manner known per se, e.g, as pressure-' - 5 -swing adsorption plant or low-temperature gas separation plant.
Embodiments of the process will now be explained with refer-ence to the drawing. The drawing represents a flow diagram of the process.
The autothermally operated reactor 1 comprises a bed 2 of granular nickel catalyst. Above the bed, a burner 3 is dis-posed, to which desulfurized natural gas is supplied through line 4, gas rich in 02 is supplied through line 5, and steam is supplied through line 6. As gas rich in o2 there is nor-mally used technically pure oxygen. The heat required for the conversion in the reactor 1 is only applied through partial oxidation. The temperatures at the outlet 8 of the reactor 1 lie in the range from 800 to 1200°C, and usually are at least 900°C.
Hot, raw synthesis gas leaves the reactor 1 and flows through a cooling 9, which may have a multistage design, which is not represented in detail in the drawing. Upon cooling, an aque-ous condensate is produced, which is withdrawn via line 10.
The cooled, raw synthesis gas, which flows through line 11, usually has temperatures in the range from 20 to 80°C. This gas is compressed by a compressor 12. Through line 14, gas rich in H2 is supplied, which comes from a separating means 15. The H2 content of the gas in line 14 is at least 80 vol-%, and preferably at least 90 vol-%.
The stoichiometric number S of the gas in line 11 lies in the range from 1.3 to 1.9, and mostly is not more than 1.8. By admixing the gas rich in H2 of line 14, the stoichiometric number of the synthesis gas in line 16 is increased to 1.95 to 2.2, as this is necessary for methanol synthesis. S is calculated in the known manner from the molar concentrations of H2, CO and C02 in accordance with S = (H2 - C02) . (CO + C02)~
To the gas in line 16, recycle gas is admixed via line 18, and the synthesis gas mixture thus formed is first of all supplied to a heat exchanger 20 via line 19, before it enters the methanol synthesis reactor 22 through line 21.
In the drawing, the reactor 22 is represented as known tubu-lar reactor, where the copper catalyst is provided in a plu-rality of tubes 23. For the indirect cooling of the interior of the tubes 23, cooling water is supplied to the reactor through line 24, and heated and partially evaporated cooling medium is withdrawn via line 25. In the tubes 23, the tem-peratures lie in the range from 180 to 350°C, and mostly in the range from 200 to 300°C. The pressure in the tubes is 10 to 150 bar, and mostly lies in the range from 20 to 100 bar.
From the reactor 22 a gas mixture containing methanol vapor is withdrawn via line 26, which gas mixture is first of all cooled in the heat exchanger 20 and is then supplied to an indirect cooler 28 via line 27. Via line 29, the mixture is supplied to the condenser 30, from which a condensate con-taining methanol and water is withdrawn via line 31. The fur-ther distillative treatment of this condensate, which is known per se, need not be explained in detail.
From the condenser 30, cooled residual gas is withdrawn through line 32 and is divided into two gas streams. The first stream is recirculated as recycle gas through line 18.
The second stream, here referred to as purge gas, is supplied to the separating means 15 via line 33, which separating means operates for instance according to the principle of pressure-change adsorption. In the separating means 15 an ex-haust gas containing methane, CO and C02 is produced in addi-tion to the gas rich in H2 of line 14, which exhaust gas is withdrawn via line 34 and can be used as combustion gas or be supplied to some other use.
For admixing the gases of lines 14 and 18 to the synthesis gas, a compressor will be required, which was, however, omit-ted in the drawing for simplification. The reactor 1 can be operated at a pressure above 40 bar, where the compressor 12 can be omitted.
Examples:
The following examples, which are calculated in part, are di-rected to the production of 1000 tons methanol per day. The nickel catalyst in the reactor 1 consists of 15 wt-% nickel on a carrier containing A12o3. The copper catalyst for the methanol synthesis contains 60 wt-o copper, 30 wt-o zinc ox-ide and 10 wt-o aluminium oxide and operates at a pressure of 80 bar. The separating means 15 is a pressure-swing adsorp-tion plant, for which a H2 yield of 85% is assumed.
Example 1:
There is used a plant corresponding to the drawing; the natu-ral gas in line 4 has been preheated to 500°C, just as the technically pure oxygen in line 5 and the steam in line 6.
Per hour, 1472 kmol natural gas, 1946 kmol steam, and 850 kmol 02 are supplied to the burner 3. In addition to methane, the natural gas contains 1.5 mol-o C2H6 and 0.4 mol-% C3Hg +
C4Hlo. Further details are shown in the following Table I, where in different lines the components of the respective mixture are indicated in mol-% and kmol/h:
_ 8 Table I
Line I 8 I 14 I 16 I 33 I 34 mol-% kmol/hmol-% mol-% mol-% kmol/hkmol/h COZ 6.1 383 - 8 12 60 60 CO 17 1070 - 22.4 4.9 24 24 H2 47.3 2982 100 69 73.4 365 56 CH4 0.8 50 - 1 8.9 44 44 H20 28.8 1821 - _ _ - _ CH30H - - - - 0.3 2 2 Total amo~lnt kmol/h6310 310 4800 498 188 Tem erature C 1000 35 35 35 35 Pressure bar 35 83 82 7g 5 The stoichiometric number S of the raw synthesis gas in line 8 is 1.79, and for the gas in line 16 S is 2Ø
Example 2_ The pressure in the reactor 1 is now 50 bar, so that the com-pressor 12 can be omitted. There is used the same natural gas as in Example 1, the temperature in lines 4, 5 and 6 is each 600°C. Per hour, 1498 kmol natural gas, 1982 kmol steam and 800 kmol technically pure oxygen are supplied to the burner 3. Further data are listed in Table II:
TnlW n, T T
Line 8 14 16 33 34 mol-% kmol/h mol-% mol-% mol-% kmol/h kmol/h COZ 5.9 372 -- 7.8 10.7 50 50 CO 16.7 1057 -- 22.2 6.3 30 30 HZ 47.3 2993 100 67.8 61.6 289 54 CH4 1.7 103 -- 2.2 20.8 95 95 H20 28.4 1794 -- -- 0.1 -- --CH30H __ __ __ __ 0.5 2 2 Total amount 6233 235 4764 470 234 kmol/h Tem erature 1000 35 35 35 35 C
Pressure bar 50 53 52 49 5 S 1.83 2 1st Comparative Exam le (calculated):
For the production of 1000 tons methanol per day there is used the natural gas of Example 1, preheated to 500°C, whereof 360 kmol/h are delivered through a commercial tubular furnace for steam reforming to a commonly used nickel cata-lyst. Thus, a H2-containing primary gas with about 40 vol-H2 is produced, to which 1030 kmol/h natural gas are admixed.
The mixture is charged into an autothermally operated reac-tor, as it is solely used in the inventive process in accor-dance with the drawing. By adding 02 and steam, a raw synthe-sis gas is produced in a total amount of 5900 kmol/h, with a temperature of 1000°C and a pressure of 35 bar, where S =
1.95. Through line 14, 66 kmol/h hydrogen are admixed to this gas, and S is thus increased to 2Ø The amount of purge gas (line 33) is only 147 kmol/h, which is charged into a pres-sure-change adsorption plant. Through line 34, 81 kmol/h ex-haust gas are discharged, which is composed of (in kmol/h) 8.2 C02, 4.1 CO, 21.1 H2, 0.5 methanol, 2.4 N2, and 44.7 CH4.
It should be noted that in particular the quantity and compo-sition of the purge gas are considerably different from Exam-ples 1 and 2.
2nd Comparative Example (calculated):
There is used the natural gas of Example 1 and the process in accordance with the drawing. From the raw synthesis gas of line 11, 200 of the Co2 are removed, so that the stoichiomet-ric number is increased to 1.93. The different lines contain the following gas quantities and gas compositions:
Line 11 14 16 33 34 mol-% kmol/hmol-% mol-% mol-% kmol/h kmol/h COz 7.1 307 -- 6.9 6.1 9 9 CO 24.2 1050 -- 23.7 3 5 5 HZ 67.5 2924 100 68.2 65.8 98 3 CH4 1.1 49 -- 1.1 23.1 35 35 CH30H + Nz -- -- -- -- 2 3 3 Total amount 4330 95 4425 15 55 kmol/h Tem erature 35 35 35 35 35 C
Pressure bar 34 83 82 7g 5 S 1.93 2 Here as well, the composition and quantity of the purge gas in line 33 is considerably different from Examples 1 and 2.
Claims (3)
1. ~A process of producing methanol from natural gas, where the natural gas is catalytically converted with steam and molecular oxygen, and a raw synthesis gas is withdrawn from the catalytic conversion, which raw synthesis gas con-tains H2, CO and CO2, a gas rich in H2 and recycle gas are admixed to the raw synthesis gas and a synthesis gas mixture is pro-duced, which is passed through at least one methanol syn-thesis reactor which contains a copper catalyst and is operated at a pressure in the range from 10 to 150 bar and at temperatures in the range from 180 to 350°C, from the methanol synthesis reactor a gas mixture containing methanol vapor is withdrawn, which is cooled to such an extent that methanol condenses out, a cooled residual gas is separately withdrawn from the methanol and divided into the recycle gas and a second partial stream (purge gas), from the purge gas, which has a H2 content of at least 30 vol-% (calculated dry), a gas rich in H2 is separated in a separating means and admixed to the raw synthesis gas, characterized in that (a) ~the catalytic conversion of natural gas with steam and oxygen is performed only in the autothermally op-erated reactor, which comprises at least one burner fed with the natural gas and the oxygen, and below the burner a bed of granular catalyst;
(b) ~the raw synthesis gas is withdrawn from the autother-mally operated reactor at a temperature in the range from 800 to 1200°C, is cooled, steam is condensed out, and condensate produced is separated;
(c)~the raw, cooled synthesis gas has a stoichiometric number S of 1.3 to 1.9 prior to admixing the gas rich in H2, where S is calculated from the molar concen-trations of H2, CO and CO2 in accordance with S = (H2 - CO2): (CO + CO2);
(d) ~the purge gas consists of ~30 to 80 vol-% H2, 7 to 35 vol-% CO2, and 3 to 25 vol-% CO, (e) ~the hourly quantity of purge gas is 8 to 25 %
(calculated dry) of the hourly quantity of cooled, raw synthesis gas;
(f) ~the hourly quantity of carbon dioxide supplied to the separating means in the purge gas is 10 to 50 % of the hourly quantity of CO2 in the cooled, raw synthe-sis gas; and (g) ~the hourly quantity of carbon monoxide supplied to the separating means in the purge gas is 2 to 10 % of the hourly quantity of CO in the cooled, raw synthe-sis gas;
(h) ~the gas rich in H2, which was withdrawn from the separating means and admixed to the raw synthesis gas, has a H2 content of at least 80 vol-%, and by admixing H2 a stoichiometric number of 1.95 to 2.2 is achieved for the synthesis gas.
(b) ~the raw synthesis gas is withdrawn from the autother-mally operated reactor at a temperature in the range from 800 to 1200°C, is cooled, steam is condensed out, and condensate produced is separated;
(c)~the raw, cooled synthesis gas has a stoichiometric number S of 1.3 to 1.9 prior to admixing the gas rich in H2, where S is calculated from the molar concen-trations of H2, CO and CO2 in accordance with S = (H2 - CO2): (CO + CO2);
(d) ~the purge gas consists of ~30 to 80 vol-% H2, 7 to 35 vol-% CO2, and 3 to 25 vol-% CO, (e) ~the hourly quantity of purge gas is 8 to 25 %
(calculated dry) of the hourly quantity of cooled, raw synthesis gas;
(f) ~the hourly quantity of carbon dioxide supplied to the separating means in the purge gas is 10 to 50 % of the hourly quantity of CO2 in the cooled, raw synthe-sis gas; and (g) ~the hourly quantity of carbon monoxide supplied to the separating means in the purge gas is 2 to 10 % of the hourly quantity of CO in the cooled, raw synthe-sis gas;
(h) ~the gas rich in H2, which was withdrawn from the separating means and admixed to the raw synthesis gas, has a H2 content of at least 80 vol-%, and by admixing H2 a stoichiometric number of 1.95 to 2.2 is achieved for the synthesis gas.
2. ~The process as claimed in claim 1, characterized in that the gas rich in H2 withdrawn from the separating means and admixed to the raw synthesis gas has a H2 content of at least 90 vol-%.
3. The process as claimed in claim 1 or 2, characterized in that the autothermally operated reactor is operated at a pressure above at least 40 bar, and the raw synthesis gas is supplied to the methanol synthesis without use of a gas compressor.
Applications Claiming Priority (2)
| Application Number | Priority Date | Filing Date | Title |
|---|---|---|---|
| DE19644216.8 | 1996-10-24 | ||
| DE19644216A DE19644216A1 (en) | 1996-10-24 | 1996-10-24 | Process for producing methanol from natural gas |
Publications (2)
| Publication Number | Publication Date |
|---|---|
| CA2213025A1 CA2213025A1 (en) | 1998-04-24 |
| CA2213025C true CA2213025C (en) | 2006-06-13 |
Family
ID=7809871
Family Applications (1)
| Application Number | Title | Priority Date | Filing Date |
|---|---|---|---|
| CA002213025A Expired - Lifetime CA2213025C (en) | 1996-10-24 | 1997-08-25 | Process of producing methanol from natural gas |
Country Status (7)
| Country | Link |
|---|---|
| EP (1) | EP0839786B1 (en) |
| CA (1) | CA2213025C (en) |
| DE (2) | DE19644216A1 (en) |
| DK (1) | DK0839786T3 (en) |
| MY (1) | MY125560A (en) |
| NO (1) | NO315938B1 (en) |
| SA (1) | SA98180958B1 (en) |
Cited By (2)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| US20210130272A1 (en) * | 2017-02-23 | 2021-05-06 | Casale Sa | Process for methanol production |
| US11001547B2 (en) * | 2015-12-04 | 2021-05-11 | Mitsubishi Heavy Industries Engineering, Ltd. | System and method for producing methanol |
Families Citing this family (10)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| NO314989B1 (en) | 1999-12-28 | 2003-06-23 | Statoil Asa | Process for increasing production in an existing processing plant for conversion of natural gas to a product, as well as processing plants |
| CA2357527C (en) | 2001-10-01 | 2009-12-01 | Technology Convergence Inc. | Methanol recycle stream |
| US7205376B2 (en) | 2004-05-04 | 2007-04-17 | General Electric Company | Processes for the production of cumene, polyetherimides, and polycarbonates |
| DE102004028200B3 (en) * | 2004-05-28 | 2005-12-15 | Hippweb E.K. | Method for carrying out heterogeneous catalytic exothermic gas phase reactions for the synthesis of methanol |
| US7019039B1 (en) * | 2005-07-14 | 2006-03-28 | Starchem Technologies, Inc. | High efficiency process for producing methanol from a synthesis gas |
| EP2399894B1 (en) | 2006-05-30 | 2014-03-12 | Starchem Technologies, Inc. | Methanol production process and system |
| EP3016924A1 (en) * | 2013-04-26 | 2016-05-11 | Silicon Fire AG | Process and reactor system for synthesis of methanol with cycle gas and purge gas recycling |
| EP3181541B1 (en) * | 2015-12-18 | 2018-08-29 | L'Air Liquide Société Anonyme pour l'Etude et l'Exploitation des Procédés Georges Claude | Method and device for the recovery of methanol |
| EP3181540B1 (en) * | 2015-12-18 | 2019-07-24 | L'Air Liquide Société Anonyme pour l'Etude et l'Exploitation des Procédés Georges Claude | Process for the separation of methanol from gas mixtures |
| RU2616919C1 (en) * | 2016-06-02 | 2017-04-18 | Андрей Владиславович Курочкин | Unit of complex non-waste gas treatment with methanol production |
Family Cites Families (4)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| US4546111A (en) * | 1983-04-22 | 1985-10-08 | Foster Wheeler Energy Corporation | Process for the production of oxygenated organic compounds such as methanol |
| GB2142331A (en) * | 1983-06-29 | 1985-01-16 | Toyo Engineering Corp | A process for producing methanol |
| CA1263671A (en) * | 1986-02-10 | 1989-12-05 | David Leon Banquy | Process for the production of synthesis gas |
| DE4130718A1 (en) * | 1991-09-14 | 1993-03-18 | Metallgesellschaft Ag | PROCESS FOR GENERATING A SYNTHESIS GAS FOR METHANOL SYNTHESIS |
-
1996
- 1996-10-24 DE DE19644216A patent/DE19644216A1/en not_active Withdrawn
-
1997
- 1997-03-27 MY MYPI97001311A patent/MY125560A/en unknown
- 1997-07-30 DE DE59708052T patent/DE59708052D1/en not_active Expired - Lifetime
- 1997-07-30 EP EP97113105A patent/EP0839786B1/en not_active Expired - Lifetime
- 1997-07-30 DK DK97113105T patent/DK0839786T3/en active
- 1997-08-25 CA CA002213025A patent/CA2213025C/en not_active Expired - Lifetime
- 1997-08-29 NO NO19973970A patent/NO315938B1/en not_active IP Right Cessation
-
1998
- 1998-03-10 SA SA98180958A patent/SA98180958B1/en unknown
Cited By (3)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| US11001547B2 (en) * | 2015-12-04 | 2021-05-11 | Mitsubishi Heavy Industries Engineering, Ltd. | System and method for producing methanol |
| US20210130272A1 (en) * | 2017-02-23 | 2021-05-06 | Casale Sa | Process for methanol production |
| US11746076B2 (en) * | 2017-02-23 | 2023-09-05 | Casale Sa | Process for methanol production |
Also Published As
| Publication number | Publication date |
|---|---|
| NO315938B1 (en) | 2003-11-17 |
| NO973970D0 (en) | 1997-08-29 |
| EP0839786B1 (en) | 2002-08-28 |
| CA2213025A1 (en) | 1998-04-24 |
| NO973970L (en) | 1998-04-27 |
| DK0839786T3 (en) | 2002-12-30 |
| DE19644216A1 (en) | 1998-04-30 |
| SA98180958B1 (en) | 2006-05-20 |
| EP0839786A3 (en) | 1999-10-06 |
| EP0839786A2 (en) | 1998-05-06 |
| MY125560A (en) | 2006-08-30 |
| DE59708052D1 (en) | 2002-10-02 |
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