CN115557829A - Method for co-producing methanol and ethanol from synthesis gas - Google Patents

Method for co-producing methanol and ethanol from synthesis gas Download PDF

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CN115557829A
CN115557829A CN202110742682.3A CN202110742682A CN115557829A CN 115557829 A CN115557829 A CN 115557829A CN 202110742682 A CN202110742682 A CN 202110742682A CN 115557829 A CN115557829 A CN 115557829A
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reactor
methanol
synthesis gas
molecular sieve
ethanol
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CN115557829B (en
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袁兴东
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Highchem Co Ltd
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    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J23/00—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00
    • B01J23/70—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00 of the iron group metals or copper
    • B01J23/76—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00 of the iron group metals or copper combined with metals, oxides or hydroxides provided for in groups B01J23/02 - B01J23/36
    • B01J23/80—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00 of the iron group metals or copper combined with metals, oxides or hydroxides provided for in groups B01J23/02 - B01J23/36 with zinc, cadmium or mercury
    • 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
    • C07C29/132—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 an oxygen containing functional group
    • C07C29/136—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 an oxygen containing functional group of >C=O containing groups, e.g. —COOH
    • C07C29/147—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 an oxygen containing functional group of >C=O containing groups, e.g. —COOH of carboxylic acids or derivatives thereof
    • C07C29/149—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 an oxygen containing functional group of >C=O containing groups, e.g. —COOH of carboxylic acids or derivatives thereof with hydrogen or hydrogen-containing gases
    • 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
    • C07C29/74—Separation; Purification; Use of additives, e.g. for stabilisation
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C31/00—Saturated compounds having hydroxy or O-metal groups bound to acyclic carbon atoms
    • C07C31/02—Monohydroxylic acyclic alcohols
    • C07C31/04—Methanol
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C31/00—Saturated compounds having hydroxy or O-metal groups bound to acyclic carbon atoms
    • C07C31/02—Monohydroxylic acyclic alcohols
    • C07C31/08—Ethanol
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C67/00—Preparation of carboxylic acid esters
    • C07C67/36—Preparation of carboxylic acid esters by reaction with carbon monoxide or formates
    • C07C67/37—Preparation of carboxylic acid esters by reaction with carbon monoxide or formates by reaction of ethers with carbon monoxide
    • 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
    • Y02E50/10—Biofuels, e.g. bio-diesel

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  • Organic Chemistry (AREA)
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  • Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)

Abstract

The invention relates to a method for the co-production of methanol and ethanol from synthesis gas, wherein the reaction process is carried out in three reactors, the method comprising: a) The synthesis gas and dimethyl ether as raw materials enter a first reactor to contact with a solid acid catalyst in the first reactor and react to obtain an effluent containing methyl acetate and/or acetic acid; b) Separating and purifying the effluent from the first reactor, and allowing the purified MA and the synthesis gas to enter a second reactor to contact and react with a hydrogenation catalyst in the second reactor to obtain an effluent containing methanol and ethanol; c) Separating the effluent from the second reactor to obtain methanol and ethanol products; d) Optionally, the methanol from step c) is passed to a third reactor for dehydration to obtain dimethyl ether, and the obtained dimethyl ether and unconverted synthesis gas are passed to the first reactor for recycle reaction. The process has the flexibility of generating methanol and ethanol from the synthesis gas, simultaneously solves the stability problem of the MA hydrogenation catalyst, has no step of preparing the methanol from the synthesis gas, does not need to separate carbon monoxide and hydrogen from the synthesis gas, and can realize the co-production of the methanol and the ethanol from the synthesis gas compared with an indirect method.

Description

Method for co-producing methanol and ethanol from synthesis gas
Technical Field
The invention relates to a method for co-producing methanol and ethanol from synthesis gas.
Background
Methanol and ethanol are important basic chemical raw materials. At present, methanol is mainly prepared from coal synthesis gas by high-pressure, medium-pressure and low-pressure synthesis methods, generally adopts a medium-low pressure process, mainly adopts a copper-based metal catalyst, and has high activity and good selectivity. The ethanol can be used for manufacturing acetaldehyde, ethylene, ethylamine, chloroethane and the like, is a basic raw material of products such as medicines, dyes, coatings, synthetic rubber, detergents and the like, has the characteristics of low heat value, high gasification potential, good anti-explosion performance and the like in the application of fuel energy industry, can be directly used as liquid fuel or mixed with gasoline for use, reduces the emission of carbon monoxide, hydrocarbons, particles, oxynitride and benzene series harmful substances in automobile exhaust, effectively improves the environmental quality, has important significance for solving the problem of atmospheric pollution and realizing sustainable development. The existing ethanol production processes are mainly a carbohydrate or cellulose fermentation process based on a biomass route and an ethylene hydration process based on a petroleum route. In recent years, the fuel ethanol production and sale amount in China is rapidly increased, and the fuel ethanol production and sale amount in China becomes the third largest fuel ethanol production country in the world after the United states and Brazil. Based on the current situation of relatively rich energy structure of Chinese coal resources, the development of a new process for synthesizing ethanol by using coal or biomass-based synthesis gas has become an important development direction of fuel energy industry.
CN 109574798A provides a method for directly producing ethanol from synthesis gas. The method takes the synthesis gas as a raw material, integrates the processes of methanol synthesis, dimethyl ether (DME) preparation by methanol, methyl Acetate (MA) preparation by dimethyl ether carbonylation and ethanol preparation by methyl acetate hydrogenation, and realizes the direct production of ethanol by the synthesis gas. The invention reduces a methanol synthesis unit and a corresponding separation unit, and also reduces a separation unit for preparing methyl acetate by dimethyl ether carbonylation. But byproducts such as acetic acid and the like can be generated in the carbonylation process of the dimethyl ether, the acetic acid has stronger corrosivity, particularly corrosivity on a rear-section hydrogenation reactor, and influence on the stability of a hydrogenation catalyst; meanwhile, the hydrogen/methyl acetate ratio (hydrogen-ester ratio) in the second reactor is low, which affects the stability of the hydrogenation catalyst. The invention does not address the stability problem of the process.
CN 103012062B discloses a method for indirectly preparing ethanol from syngas. The method comprises the steps of firstly, synthesizing methanol by using synthesis gas formed by mixing hydrogen and carbon monoxide as a raw material, dehydrating the methanol to prepare dimethyl ether, then mixing the dimethyl ether with the carbon monoxide and the hydrogen to carry out carbonylation reaction to prepare methyl acetate, purifying the methyl acetate to carry out hydrogenation, and purifying a hydrogenation product to obtain the ethanol. The hydrogenation part of the process uses pure hydrogen, needs to be separated from synthesis gas firstly, and needs a device for preparing methanol from the synthesis gas, thereby having economic defects.
Disclosure of Invention
The inventor of the invention researches and considers that the direct method has the problems of corrosivity and influence on the stability of the hydrogenation catalyst although the process is simple and the flow is short; the indirect method has the problems of long process flow and unreasonable economy. The invention combines the two methods, and provides a new method for preparing ethanol from synthesis gas, namely, firstly, in a first reactor, the synthesis gas is adopted to react with dimethyl ether to generate methyl acetate and trace acetic acid; removing by-products influencing the usability of the device and the stability of a hydrogenation catalyst through separation, optionally removing unconverted DME, and feeding the purified MA and the unreacted synthesis gas of the first reactor into a second reactor; in order to maintain high hydrogen-ester ratio, the separated gas of the effluent at the outlet of the second reactor is circularly pumped into the second reactor, and MA and the synthesis gas react in the second reactor to simultaneously generate methanol and ethanol; the hydrogenation catalysts with different compositions are selected to generate methanol and ethanol mixtures with different proportions according to the requirements of subsequent processes, particularly the mixture with the molar ratio of the methanol to the ethanol of 1.1-5.0, and pure methanol and pure ethanol are respectively obtained after separation. Methanol and ethanol can be directly obtained as products, and the methanol can also enter a dehydration reactor to generate DME required by the first carbonylation reactor. The process has the flexibility of generating methanol and ethanol from the synthesis gas, and simultaneously, the second reactor can adopt high hydrogen-ester ratio, thereby solving the stability problem of the MA hydrogenation catalyst.
The object of the present invention is to overcome some of the problems of the prior art and to provide a process for the conversion of synthesis gas with which the co-production of methanol and ethanol from synthesis gas can be achieved.
The embodiments for achieving the above objects of the present invention can be summarized as follows:
1. a process for co-producing methanol and ethanol from syngas, comprising:
a) The synthesis gas and dimethyl ether as raw materials enter a first reactor to contact with a solid acid catalyst in the first reactor and react to obtain an effluent containing methyl acetate and/or acetic acid;
b) Separating and purifying the effluent from the first reactor, and allowing the purified MA and the synthesis gas to enter a second reactor to contact and react with a hydrogenation catalyst in the second reactor to obtain an effluent containing methanol and ethanol;
c) Separating the effluent from the second reactor to obtain methanol and ethanol as products, preferably at a methanol to ethanol molar ratio of 1.1 to 5.0, more preferably 1.5 to 3.0;
d) Optionally, the methanol from step c) is passed to a third reactor for dehydration reaction to obtain dimethyl ether, and the obtained dimethyl ether and unconverted synthesis gas are passed to the first reactor for recycle reaction.
2. The process of embodiment 1, wherein the volume content of syngas in the feedstock is from 10 to 100%, the volume content of dimethyl ether is from 0 to 90%, and the volume ratio of carbon monoxide to hydrogen in the syngas is from 0.1 to 10; preferably, the volume content of carbon monoxide and hydrogen in the raw material is 50-100%; the volume content of any one or more of nitrogen, helium, argon and carbon dioxide in the raw material of the synthesis gas is 0-50%.
3. The process according to embodiment 1 or 2, wherein the reaction temperature of the first reactor is from 100 to 300 ℃, preferably from 120 to 250 ℃, and the reaction pressure is from 0.5 to 20MPa, preferably from 1 to 15MPa; the reaction temperature of the second reactor is 100-300 ℃, preferably 150-175 ℃, the reaction pressure is 0.5-20MPa, preferably 1.0-15.0MPa, and the hydrogen-ester ratio is 30-300, preferably 50-150; the reaction temperature of the third reactor is 180-420 ℃, preferably 200-400 ℃, and the reaction pressure is 0.1-4MPa, preferably 0.2-3MPa.
4. The process of any of embodiments 1-3, wherein the solid acid catalyst in the first reactor comprises at least one molecular sieve in the following: FER zeolite molecular sieve, MFI zeolite molecular sieve, MOR zeolite molecular sieve, ETL zeolite molecular sieve, MFS zeolite molecular sieve and molecular sieve products obtained by modifying elements except framework composition elements or pyridine, preferably pyridine modified MOR zeolite molecular sieve.
5. The process according to any of embodiments 1-4, wherein the solid acid catalyst is the hydrogen form product of the zeolitic molecular sieve, or consists of 10-95 wt.% of the hydrogen form product with the balance being a matrix, or is a molecular sieve product obtained by pyridine modification of the hydrogen form product, wherein the matrix is at least one selected from alumina, silica, kaolin, and magnesia.
6. The process of any of embodiments 1-5, wherein the hydrogenation catalyst in the second reactor is a copper-based catalyst.
7. The process according to embodiment 6, wherein the copper-based catalyst comprises from 15 to 90% by weight of SiO as support 2 5-50 wt% of Cu and 0-50 wt% of at least one of transition metal, preferably at least one of Fe, co, W, mo, ti, V, ni, zr, pt, pd and Au, alkaline earth metal, preferably at least one of Ca, mg and Ba, and zinc group element, preferably Zn.
8. The process according to any of embodiments 1-7, wherein the catalyst in the third reactor is a methanol to dimethyl ether solid acid catalyst, preferably a molecular sieve based and oxide based catalyst, more preferably HZSM-5, beta molecular sieve, silicoaluminophosphate molecular sieve and active Al 2 O 3 And mixtures thereof.
9. The process according to any of embodiments 1-8, wherein the first reactor and/or the second reactor and/or the third reactor is a fixed bed or a fluidized bed reactor, preferably a fixed bed tubular reactor.
10. A copper-based catalyst comprising 15 to 90% by weight of SiO as a carrier 2 5-50 wt% of Cu and 0-50 wt% of at least one of transition metal, preferably at least one of Fe, co, W, mo, ti, V, ni, zr, pt, pd and Au, alkaline earth metal, preferably at least one of Ca, mg and Ba, and zinc group element, preferably Zn.
The present invention includes but is not limited to the following benefits:
1. the method simplifies a device for preparing methanol from synthesis gas, realizes the co-production of methanol and ethanol on a hydrogenation reactor by adopting the synthesis gas as a gas source, and can flexibly adjust the molar ratio of the methanol to the ethanol between 1.1 and 5.0. Because the effluent of the first reactor is purified and the high hydrogen-ester ratio is adopted in the second reactor, the service life of the catalyst of the second reactor is ensured.
2. The process flow of the invention always adopts the synthesis gas as the raw material, thereby reducing the separation process of the synthesis gas and improving the economy.
Detailed Description
The method of the invention comprises the following steps: dimethyl ether and synthesis gas contact with a solid acid catalyst in a first reactor to react to obtain an oxygen-containing compound MA; the liquid phase by-products are then removed, optionally with unconverted DME, by separation.
The synthesis gas and the purified methyl acetate contact and react with a hydrogenation catalyst in a second reactor to generate methanol and ethanol; and then, ethanol is separated out as a product, methanol can be separated out as a product, and can also be dehydrated in a third reactor to generate dimethyl ether, and the obtained dimethyl ether is circularly fed into a reaction system and fed into the first reactor for carbonylation with synthetic gas as a reaction raw material. The method can realize the coproduction of methanol and ethanol by the synthesis gas, the ratio of the methanol to the ethanol can be controlled by selecting the hydrogenation catalysts with different compositions, the invention reduces the independent methanol synthesis unit, adopts the synthesis gas in the whole process flow, reduces the equipment investment and the energy consumption, and has simple whole process flow. In addition, MA generated in the first reactor is purified, so that the corrosion resistance of the second reactor and the stability of the catalyst are improved, and meanwhile, the stability of the catalyst is improved by adopting a high hydrogen-ester ratio, so that the process has a good application prospect.
More specifically, in the process for co-producing methanol and ethanol from syngas of the present invention, the reaction is accomplished in a first reactor, a second reactor, and optionally a third reactor, the process comprising:
a) Allowing synthesis gas and dimethyl ether as raw materials to enter a first reactor to contact with a solid acid catalyst in the first reactor and react to obtain an effluent containing methyl acetate and/or acetic acid;
b) Separating and purifying the effluent from the first reactor, and allowing the purified MA and the synthesis gas to enter a second reactor to contact and react with a hydrogenation catalyst in the second reactor to obtain an effluent containing methanol and ethanol;
c) Separating the effluent from the second reactor to yield methanol and ethanol as products, preferably in a molar ratio of methanol to ethanol of from 1.1 to 5.0, for example 1.5,2.0,2.5,3.0,3.5,4.0,4.5, more preferably from 1.5 to 3.0;
d) Optionally, the methanol from step c) is passed to a third reactor for dehydration to obtain dimethyl ether, and the obtained dimethyl ether and unconverted synthesis gas are passed to the first reactor for recycle reaction.
In one embodiment of the process of the present invention, the synthesis gas content of the feedstock is in the range of 10 to 100% by volume, such as 20%,30%,40%,50%,60%,70%,80%,90%.
In one embodiment of the process of the present invention the dimethyl ether content of the feed is from 0 to 90% by volume, for example 5%,10%,20%,30%,40%,50%,60%,70%,80%.
In one embodiment of the process of the present invention, the volume ratio of carbon monoxide to hydrogen in the synthesis gas in the feed is in the range of from 0.1 to 10, for example 0.5,1.0,1.5,2.0,2.5,3.0,3.5,4.0,4.5,5.0,5.5,6.0,6.5,7.0,7.5,8.0,8.5,9.0,9.5; the molar ratio of syngas to DME is in the range of 10 to 100, for example 15, 20, 25, 30, 35, 40, 45, 50, 55, 60, 65, 70, 75, 80, 85, 90, 95.
In one embodiment of the process of the present invention, the synthesis gas may also contain, in addition to carbon monoxide and hydrogen, any one or more inert gases of nitrogen, helium, argon and carbon dioxide. Preferably, the carbon monoxide and hydrogen content is 50-100% by volume, for example 60%,70%,80%,90%; the content of any one or more of nitrogen, helium, argon and carbon dioxide in the synthesis gas is 0-50% by volume, such as 10%,20%,30%,40%.
In one embodiment of the process of the invention, the reaction temperature of the first reactor is in the range of from 100 to 300 ℃ such as 110 ℃,120 ℃,130 ℃,140 ℃,150 ℃,160 ℃,170 ℃,180 ℃,190 ℃,200 ℃,210 ℃,220 ℃,230 ℃,240 ℃,250 ℃,260 ℃,270 ℃,280 ℃,290 ℃ and the reaction pressure is in the range of from 0.5 to 20MPa, such as 1MPa,2MPa,3MPa,4MPa,5MPa,6MPa,7MPa,8MPa,9MPa,10MPa, 11112MPa, 13MPa,14MPa,15MPa, 10MPa, 18MPa; the CO/DME ratio is in the range of 1 to 20, for example 1,2,3,4,5,6,7,8,9, 10, 11, 12, 13, 14, 15, 16, 17, 18, 19.
In one embodiment of the process of the invention, the reaction temperature of the second reactor is in the range of from 100 to 300 ℃ such as 110 ℃,120 ℃,130 ℃,140 ℃,150 ℃,160 ℃,170 ℃,180 ℃,190 ℃,200 ℃,210 ℃,220 ℃,230 ℃,240 ℃,250 ℃,260 ℃,270 ℃,280 ℃,290 ℃ and the reaction pressure is in the range of from 0.5 to 20MPa, such as 1MPa,2MPa,3MPa,4MPa,5MPa,6MPa,7MPa,8MPa,9MPa,10MPa, 11112MPa, 13MPa,14MPa,15MPa, 10MPa, 18MPa; the hydrogen to ester ratio is 30 to 300, for example 40, 50, 60, 70, 80, 90, 100, 110, 120, 130, 140, 150, 160, 170, 180, 190, 200, 210, 220, 230, 240, 250, 260, 270, 280, 290.
In one embodiment of the process of the invention, the reaction temperature of the third reactor is in the range of from 180 to 420 ℃, such as 190 ℃,200 ℃,210 ℃,220 ℃,230 ℃,240 ℃,250 ℃,260 ℃,270 ℃,280 ℃,290 ℃,300 ℃,310 ℃,320 ℃,330 ℃,340 ℃,350 ℃,360 ℃,370 ℃,380 ℃,390 ℃,400 ℃,410 ℃ and the reaction pressure is in the range of from 0.1 to 4MPa, such as 0.5MPa,1MPa,1.5MPa,2MPa,2.5MPa,3MPa,3.5MPa.
In one embodiment of the process of the present invention, the solid acid catalyst in the first reactor comprises at least one molecular sieve of: FER zeolite molecular sieve, MFI zeolite molecular sieve, MOR zeolite molecular sieve, ETL zeolite molecular sieve, MFS zeolite molecular sieve and molecular sieve products obtained by modifying elements except framework composition elements or pyridine, preferably pyridine modified MOR zeolite molecular sieve.
Preferably, the solid acid catalyst in the first reactor is a pyridine modified MOR zeolite molecular sieve. Preferably, the MOR zeolite molecular sieve is modified with 5 to 95 wt%, e.g., 10 wt%, 15 wt%, 20 wt%, 25 wt%, 30 wt%, 35 wt%, 40 wt%, 50 wt%, 55 wt%, 60 wt%, 65 wt%, 70 wt%, 75 wt%, 80 wt%, 85 wt%, 90 wt% pyridine.
Preferably, the solid acid catalyst is the hydrogen form product of the zeolitic molecular sieve, or is comprised of 10 to 95 wt.%, e.g., 15 wt.%, 20 wt.%, 25 wt.%, 30 wt.%, 35 wt.%, 40 wt.%, 45 wt.%, 50 wt.%, 55 wt.%, 60 wt.%, 65 wt.%, 70 wt.%, 75 wt.%, 80 wt.%, 85 wt.%, 90 wt.% of the hydrogen form product and the balance of the matrix, or is the molecular sieve product resulting from pyridine modification of the hydrogen form product.
Preferably, the matrix is at least one selected from the group consisting of alumina, silica, kaolin, and magnesia.
In one embodiment of the process of the present invention, the catalyst in the second reactor is a hydrogenation catalyst having methanol synthesis as well as ester hydrogenation properties.
Preferably, the hydrogenation catalyst in the second reactor is a copper-based catalyst.
More preferably, the carrier of the copper-based catalyst is SiO 2 The copper-based catalyst comprises 15 to 90 wt%, e.g., 20 wt%, 25 wt%, 30 wt%, 35 wt%, 40 wt%, 45 wt%, 50 wt%, 55 wt%, 60 wt%, 65 wt%, 70 wt%, 75 wt%, 80 wt%, 85 wt% of SiO 2 5-50 wt%, such as 10 wt%, 15 wt%, 20 wt%, 25 wt%, 30 wt%, 35 wt%, 40 wt%, 45 wt% of Cu and 0-50 wt%, such as 5 wt%, 10 wt%, 15 wt%, 20 wt%, 25 wt%, 30 wt%, 35 wt%, 40 wt%, 45 wt% of at least one of a transition metal, preferably at least one of Fe, co, W, mo, ti, V, ni, zr, pt, pd and Au, an alkaline earth metal, preferably at least one of Ca, mg and Ba, and a zinc group element, preferably Zn.
In one embodiment of step b) of the process according to the invention, the effluent from said first reactor is subjected to separation to remove by-products affecting the stability of the hydrogenation catalyst, optionally to remove unconverted DME.
In one embodiment of the process of the present invention, the catalyst in the third reactor is a solid acid catalyst for the preparation of dimethyl ether from methanol, preferably a molecular sieve-based and oxide-based catalyst, more preferably HZSM-5, a beta molecular sieve, a silicoaluminophosphate molecular sieve and active Al 2 O 3 And mixtures thereof.
In one embodiment of the process of the present invention, the first reactor and/or the second reactor and/or the third reactor is a fixed bed or a fluidized bed reactor, preferably a fixed bed shell and tube reactor.
In a further preferred embodiment, the reaction conditions of the first reactor are: the reaction temperature is 120-250 ℃, and the reaction pressure is 1-15MPa; reaction conditions of the second reactor: the reaction temperature is 120-175 ℃, and the reaction pressure is 1-15MPa; the hydrogen-ester ratio is 50-150; reaction conditions of the third reactor: the reaction temperature is 200-400 ℃, and the reaction pressure is 0.2-3MPa.
The present invention is specifically illustrated by the following examples, but the present invention is not limited to these examples.
Examples
Preparation of solid acid catalyst-pyridine modified Hydrogen form sample
10g of hydrogen MOR (SiO) 2 /Al 2 O 3 =25, supplied by the company tokyo) samples were loaded into a reaction tube, and the temperature was gradually raised to 350 ℃ under a nitrogen atmosphere of 100mL/min and maintained for 4 hours; then carrying pyridine with nitrogen, and treating for 8 hours at normal temperature; then using pure N 2 The gas, treated at 300 ℃ for 8 hours, yielded a pyridine-modified sample having a pyridine content of 7% by weight, marked Py-MOR.
Preparation of MA hydrogenation catalyst-copper-based catalyst
In a beaker, cu (NO) was added in an amount shown in Table 1 3 ) 2 ·3H 2 O and Zn (NO) 3 ) 2 ·6H 2 O was dissolved in 300g of deionized water to obtain an aqueous mixed metal nitrate solution. In a separate beaker 91.7g of concentrated ammonia (28%) diluted with 100g of deionized water was added with the amount of SiO indicated in Table 1 2 (A380, supplied by EVONIK, germany) and vigorously stirring an aqueous ammonia solution at room temperature, then slowly adding the resulting aqueous mixed metal nitrate solution to the aqueous ammonia solution for about 30min, and evaporating off ammonia and water. Then, the precipitate was washed to neutrality with deionized water and centrifuged. And drying the obtained precipitate in a 120 ℃ oven for 24 hours, grinding, tabletting, crushing and screening the dried sample, then placing the sample in a muffle furnace, heating to 450 ℃ at the heating rate of 1 ℃/min, and roasting for 5 hours to obtain a roasted sample. The composition of each sample is shown in table 1.
TABLE 1
Figure BDA0003143250530000091
Preparation of catalyst for preparing dimethyl ether from methanol
The catalyst for preparing dimethyl ether from methanol is hydrogen type ZSM-5 (SiO) 2 /Al 2 O 3 =50, supplied by Tosoh) molecular sieves and gamma-alumina in a ratio of 50: 50 and is labeled as a dehydration catalyst.
Example 1
A first reactor with a stainless steel column having an inner diameter of 14.00mm and a length of 60cm was charged with 10g of Py-MOR catalyst, and was purged with a gas containing CO and H 2 The synthesis gas (composition 32% CO/64% 2 Ar/4%) at a flow rate of 400mL/min. After being pressurized and liquefied, the dimethyl ether is pumped into a first reactor by a micro pump at the speed of 0.038g/min, the reaction temperature is 160 ℃, the reaction pressure is 4.5MPa, and the reaction time is 4 hours. The reaction effluent was subjected to side-line sampling and full-component analysis by using Shimadzu 2014 gas chromatography, and the detector was FID. The analysis conditions are as follows: the column length is 30m, the column temperature adopts programmed temperature rise, the initial temperature is 30 ℃, the temperature is kept for 3min, then the temperature is raised to 120 ℃ at 4 ℃/min, and the temperature of the detector is 220 ℃. From the chromatographic quantification, DME conversion and product distribution were calculated. The conversion of DME was 72.40%, the MA selectivity was 95.28%, the acetic acid selectivity was 0.22%, and the others were 4.50%. After gas-liquid separation of the reaction effluent, the gas phase part enters a second reactor; the liquid phase part of DME is recycled into the first reactor, and MA is used as the raw material of the second reactor.
A second reactor, which had the same size as the first reactor, was charged with 10.0g of hydrogenation catalyst B-1; MA flow rate of 0.04g/min, gas as syngas (composition 32% CO/64% 2 Ar content: 1525mL/min, reaction temperature of 170 ℃, reaction pressure of 3.0MPa, hydrogen-ester ratio of 80, and reaction time of 4 hours. Separating, and analyzing a gas phase part by TCD (TCD chromatography) under the analysis conditions that an analysis column is a 4m long activated carbon stainless steel column, the column temperature is 80 ℃, the detector temperature is 120 ℃, and calculating the conversion rate of CO; the liquid phase portion was analyzed by the same method as in the first reactor. The conversion of MA and the product composition were calculated. It was calculated that 1.48g of ethanol and 1.37g of methanol were obtained, the molar ratio of methanol to ethanol being 1.32. The results of the reaction at the outlet of the second reactor are shown in Table 2.
The third reactor adopts 5.0g of dehydration catalyst, and the methanol separated from the second reactor is pumped by a high-pressure liquid pump, the flow rate is 3.5g/h, the reaction temperature is 270 ℃, and the pressure is normal pressure. The amount of DME formed after the reaction was 2.3g/h.
Example 2
The difference from example 1 is that the second reactor is charged with 10.0g of catalyst B-2. A sample was taken after 4 hours of reaction and analyzed, and the reaction results at the outlet of the second reactor are shown in Table 2.
Example 3
The difference from example 1 is that the second reactor is charged with 10.0g of catalyst B-3. A sample was taken after 4 hours of reaction time and the reaction results at the outlet of the second reactor are shown in Table 2.
Example 4
The difference from example 1 is that the second reactor is charged with 10.0g of catalyst B-4. A sample was taken after 4 hours of reaction time and the reaction results at the outlet of the second reactor are shown in Table 2.
Example 5
The difference from example 1 is that the second reactor is charged with 10.0g of catalyst B-5. A sample was taken after 4 hours of reaction and analyzed, and the reaction results at the outlet of the second reactor are shown in Table 2.
Example 6
The difference from example 1 is that the second reactor is charged with 10.0g of catalyst B-6. A sample was taken after 4 hours of reaction time and the reaction results at the outlet of the second reactor are shown in Table 2.
Example 7
The difference from example 1 is that the second reactor is charged with 10.0g of catalyst B-7. A sample was taken after 4 hours of reaction and analyzed, and the reaction results at the outlet of the second reactor are shown in Table 2.
Example 8
The difference from example 1 is that the second reactor is charged with 4.0g of catalyst B-6. A sample was taken after 4 hours of reaction and analyzed, and the reaction results at the outlet of the second reactor are shown in Table 2.
Example 9
The difference from example 1 is that the second reactor is charged with 7.0g of catalyst B-6. A sample was taken after 4 hours of reaction and analyzed, and the reaction results at the outlet of the second reactor are shown in Table 2.
Example 10
The difference from example 1 is that the second reactor is charged with 13.0g of catalyst B-6. A sample was taken after 4 hours of reaction time and the reaction results at the outlet of the second reactor are shown in Table 2.
Comparative example 1
The difference from example 6 was that the first reactor was charged with 10.0g of Py-MOR catalyst in the upper section and 10.0g of B-6 catalyst in the lower section, and the second reactor was removed. After 4 hours of reaction, a sample was taken for analysis, and the reaction results at the outlet of the first reactor were as follows: DME conversion was 70.5%, CO conversion was 24.60%, and the product composition was 1.23% MA, 60.00% methanol, 34.97% ethanol and 3.80% others.
Comparative example 2
The difference from example 6 is that the gaseous feed to the second reactor is pure hydrogen. Samples were taken after 4 hours of reaction and analyzed, and the results at the outlet of the second reactor are shown in Table 2.
Comparative example 3
The difference from example 6 is that the effluent from the outlet of the first reactor, after separation to remove unconverted DME, is passed directly to the second reactor without further separation. Samples were taken after 4 hours of reaction and analyzed, and the results at the outlet of the second reactor are shown in Table 2.
Comparative example 4
The difference from example 1 is that the second reactor is charged with 10.0g of catalyst B-1'. A sample was taken after 4 hours of reaction time and the reaction results at the outlet of the second reactor are shown in Table 2.
Example 11
The difference from example 6 is that the reaction temperature of the second reactor was 160 ℃. A sample was taken after 4 hours of reaction and analyzed, and the reaction results at the outlet of the second reactor are shown in Table 2.
Example 12
The difference from example 6 is that the reaction temperature of the second reactor was 165 ℃. A sample was taken after 4 hours of reaction time and the reaction results at the outlet of the second reactor are shown in Table 2.
Example 13
The difference from example 6 is that the reaction temperature in the second reactor was 175 ℃. A sample was taken after 4 hours of reaction and analyzed, and the reaction results at the outlet of the second reactor are shown in Table 2.
Example 14
The difference from example 6 is that the flow rate of the synthesis gas in the second reactor was 762.5mL/min and the hydrogen-ester ratio was 40. A sample was taken after 4 hours of reaction time and the reaction results at the outlet of the second reactor are shown in Table 2.
Example 15
The difference from example 6 is that the synthesis gas flow rate of the second reactor was 1144mL/min, and the hydrogen-ester ratio was 60. A sample was taken after 4 hours of reaction time and the reaction results at the outlet of the second reactor are shown in Table 2.
Example 16
The difference from example 6 is that the flow rate of the synthesis gas in the second reactor was 1906mL/min and the hydrogen-ester ratio was 100. A sample was taken after 4 hours of reaction time and the reaction results at the outlet of the second reactor are shown in Table 2.
Example 17
The difference from example 6 is that the synthesis gas flow rate in the second reactor was 2288mL/min and the hydrogen-to-ester ratio was 120. A sample was taken after 4 hours of reaction time and the reaction results at the outlet of the second reactor are shown in Table 2.
Example 18
The difference from example 6 is that a sample was taken after 100 hours of reaction and analyzed, and the reaction results at the outlet of the second reactor are shown in Table 2.
Example 19
The difference from example 6 is that a sample was taken after 500 hours of reaction and analyzed, and the reaction results at the outlet of the second reactor are shown in Table 2.
Example 20
The difference from example 6 is that a sample was taken after 1000 hours of reaction and analyzed, and the reaction results at the outlet of the second reactor are shown in Table 2.
Example 21
The difference from example 6 is that a sample was taken after 1500 hours of reaction and analyzed, and the reaction results at the outlet of the second reactor are shown in Table 2.
Example 22
The difference from example 6 is that samples were taken after 2000 hours of reaction and analyzed, and the reaction results at the outlet of the second reactor are shown in Table 2.
Table 2: reaction results at the outlet of the second reactor
Figure BDA0003143250530000141
Figure BDA0003143250530000151

Claims (10)

1. A process for co-producing methanol and ethanol from syngas, comprising:
a) The synthesis gas and dimethyl ether as raw materials enter a first reactor to contact with a solid acid catalyst in the first reactor and react to obtain an effluent containing methyl acetate and/or acetic acid;
b) Separating and purifying the effluent from the first reactor, and allowing the purified MA and the synthesis gas to enter a second reactor to contact and react with a hydrogenation catalyst in the second reactor to obtain an effluent containing methanol and ethanol;
c) Separating the effluent from the second reactor to obtain methanol and ethanol as products, preferably at a methanol to ethanol molar ratio of 1.1 to 5.0, more preferably 1.5 to 3.0;
d) Optionally, the methanol from step c) is passed to a third reactor for dehydration to obtain dimethyl ether, and the obtained dimethyl ether and unconverted synthesis gas are passed to the first reactor for recycle reaction.
2. The process of claim 1 wherein the feedstock has a volume content of syngas of 10-100%, a volume content of dimethyl ether of 0-90%, and a volume ratio of carbon monoxide to hydrogen of 0.1-10 in the syngas; preferably, the volume content of carbon monoxide and hydrogen in the raw material is 50-100%; the volume content of any one or more gases of nitrogen, helium, argon and carbon dioxide in the synthesis gas raw material is 0-50%.
3. The process according to claim 1 or 2, wherein the reaction temperature of the first reactor is 100-300 ℃, preferably 120-250 ℃, the reaction pressure is 0.5-20MPa, preferably 1-15MPa; the reaction temperature of the second reactor is 100-300 ℃, preferably 150-175 ℃, the reaction pressure is 0.5-20MPa, preferably 1.0-15.0MPa, and the hydrogen-ester ratio is 30-300, preferably 50-150; the reaction temperature of the third reactor is 180-420 ℃, preferably 200-400 ℃, and the reaction pressure is 0.1-4MPa, preferably 0.2-3MPa.
4. The process of any of claims 1-3, wherein the solid acid catalyst in the first reactor comprises at least one molecular sieve of: FER zeolite molecular sieve, MFI zeolite molecular sieve, MOR zeolite molecular sieve, ETL zeolite molecular sieve, MFS zeolite molecular sieve and molecular sieve products obtained by modifying elements except framework composition elements or pyridine, preferably pyridine modified MOR zeolite molecular sieve.
5. The process of any one of claims 1-4, wherein the solid acid catalyst is the hydrogen form product of the zeolite molecular sieve, or consists of 10-95 wt% of the hydrogen form product with the balance being a matrix, or is a molecular sieve product obtained by modifying the hydrogen form product with pyridine, wherein the matrix is at least one selected from alumina, silica, kaolin, and magnesia.
6. The process of any one of claims 1-5, wherein the hydrogenation catalyst in the second reactor is a copper-based catalyst.
7. The process according to claim 6, wherein the copper-based catalyst comprises from 15 to 90% by weight of SiO as support 2 5-50 wt% of Cu and 0-50 wt% of at least one of transition metal, alkaline earth metal and zinc group element, the transition metal is preferably Fe, co, W,At least one of Mo, ti, V, ni, zr, pt, pd and Au, the alkaline earth metal is preferably at least one of Ca, mg and Ba, and the zinc group element is preferably Zn.
8. The process of any one of claims 1-7, wherein the catalyst in the third reactor is a methanol to dimethyl ether solid acid catalyst, preferably a molecular sieve based and oxide based catalyst, more preferably HZSM-5, beta molecular sieve, silicoaluminophosphate molecular sieve and active Al 2 O 3 And mixtures thereof.
9. The process according to any one of claims 1 to 8, wherein the first reactor and/or the second reactor and/or the third reactor is a fixed bed or a fluidized bed reactor, preferably a fixed bed shell and tube reactor.
10. A copper-based catalyst comprising 15 to 90% by weight of SiO as a carrier 2 5-50 wt% of Cu and 0-50 wt% of at least one of transition metal, preferably at least one of Fe, co, W, mo, ti, V, ni, zr, pt, pd and Au, alkaline earth metal, preferably at least one of Ca, mg and Ba, and zinc group element, preferably Zn.
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