CN203033918U - Process system for indirectly producing ethanol by using synthesis gases - Google Patents
Process system for indirectly producing ethanol by using synthesis gases Download PDFInfo
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- CN203033918U CN203033918U CN2012207125324U CN201220712532U CN203033918U CN 203033918 U CN203033918 U CN 203033918U CN 2012207125324 U CN2012207125324 U CN 2012207125324U CN 201220712532 U CN201220712532 U CN 201220712532U CN 203033918 U CN203033918 U CN 203033918U
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- LFQSCWFLJHTTHZ-UHFFFAOYSA-N Ethanol Chemical compound CCO LFQSCWFLJHTTHZ-UHFFFAOYSA-N 0.000 title claims abstract description 204
- 239000007789 gas Substances 0.000 title claims abstract description 96
- 238000000034 method Methods 0.000 title claims abstract description 80
- 230000015572 biosynthetic process Effects 0.000 title claims abstract description 37
- 238000003786 synthesis reaction Methods 0.000 title claims abstract description 37
- OKKJLVBELUTLKV-UHFFFAOYSA-N Methanol Chemical compound OC OKKJLVBELUTLKV-UHFFFAOYSA-N 0.000 claims abstract description 435
- 238000005810 carbonylation reaction Methods 0.000 claims abstract description 107
- 238000005984 hydrogenation reaction Methods 0.000 claims abstract description 104
- 230000006315 carbonylation Effects 0.000 claims abstract description 100
- LCGLNKUTAGEVQW-UHFFFAOYSA-N Dimethyl ether Chemical compound COC LCGLNKUTAGEVQW-UHFFFAOYSA-N 0.000 claims abstract description 91
- 238000006243 chemical reaction Methods 0.000 claims abstract description 67
- XBDQKXXYIPTUBI-UHFFFAOYSA-M Propionate Chemical compound CCC([O-])=O XBDQKXXYIPTUBI-UHFFFAOYSA-M 0.000 claims abstract description 61
- KXKVLQRXCPHEJC-UHFFFAOYSA-N acetic acid trimethyl ester Natural products COC(C)=O KXKVLQRXCPHEJC-UHFFFAOYSA-N 0.000 claims abstract description 61
- 238000011084 recovery Methods 0.000 claims abstract description 61
- 229910052739 hydrogen Inorganic materials 0.000 claims abstract description 34
- 239000001257 hydrogen Substances 0.000 claims abstract description 31
- UFHFLCQGNIYNRP-UHFFFAOYSA-N Hydrogen Chemical compound [H][H] UFHFLCQGNIYNRP-UHFFFAOYSA-N 0.000 claims abstract description 28
- 229910002091 carbon monoxide Inorganic materials 0.000 claims abstract description 17
- UGFAIRIUMAVXCW-UHFFFAOYSA-N Carbon monoxide Chemical compound [O+]#[C-] UGFAIRIUMAVXCW-UHFFFAOYSA-N 0.000 claims abstract description 13
- 239000007788 liquid Substances 0.000 claims description 94
- 239000000463 material Substances 0.000 claims description 85
- 239000012071 phase Substances 0.000 claims description 46
- 239000007791 liquid phase Substances 0.000 claims description 33
- XLYOFNOQVPJJNP-UHFFFAOYSA-N water Substances O XLYOFNOQVPJJNP-UHFFFAOYSA-N 0.000 claims description 29
- 238000000066 reactive distillation Methods 0.000 claims description 27
- 239000002994 raw material Substances 0.000 claims description 20
- 238000009833 condensation Methods 0.000 claims description 11
- 230000005494 condensation Effects 0.000 claims description 11
- 238000010992 reflux Methods 0.000 claims description 10
- CURLTUGMZLYLDI-UHFFFAOYSA-N Carbon dioxide Chemical compound O=C=O CURLTUGMZLYLDI-UHFFFAOYSA-N 0.000 claims description 8
- 229910002092 carbon dioxide Inorganic materials 0.000 claims description 6
- 239000001569 carbon dioxide Substances 0.000 claims description 4
- 239000003054 catalyst Substances 0.000 abstract description 44
- 238000004519 manufacturing process Methods 0.000 abstract description 16
- 238000000926 separation method Methods 0.000 abstract description 9
- 238000005265 energy consumption Methods 0.000 abstract description 7
- 238000006297 dehydration reaction Methods 0.000 abstract description 5
- 235000019441 ethanol Nutrition 0.000 description 61
- QTBSBXVTEAMEQO-UHFFFAOYSA-N Acetic acid Chemical compound CC(O)=O QTBSBXVTEAMEQO-UHFFFAOYSA-N 0.000 description 48
- 239000000047 product Substances 0.000 description 23
- 239000000945 filler Substances 0.000 description 17
- 239000007787 solid Substances 0.000 description 15
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- 238000002156 mixing Methods 0.000 description 12
- 239000007795 chemical reaction product Substances 0.000 description 11
- 230000002378 acidificating effect Effects 0.000 description 9
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- 239000003930 superacid Substances 0.000 description 8
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- IJGRMHOSHXDMSA-UHFFFAOYSA-N Atomic nitrogen Chemical compound N#N IJGRMHOSHXDMSA-UHFFFAOYSA-N 0.000 description 6
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- 229910001870 ammonium persulfate Inorganic materials 0.000 description 3
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- 238000000855 fermentation Methods 0.000 description 3
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- 229910052703 rhodium Inorganic materials 0.000 description 3
- 239000010948 rhodium Substances 0.000 description 3
- MHOVAHRLVXNVSD-UHFFFAOYSA-N rhodium atom Chemical compound [Rh] MHOVAHRLVXNVSD-UHFFFAOYSA-N 0.000 description 3
- RYGMFSIKBFXOCR-UHFFFAOYSA-N Copper Chemical compound [Cu] RYGMFSIKBFXOCR-UHFFFAOYSA-N 0.000 description 2
- 229910017773 Cu-Zn-Al Inorganic materials 0.000 description 2
- QUSNBJAOOMFDIB-UHFFFAOYSA-N Ethylamine Chemical compound CCN QUSNBJAOOMFDIB-UHFFFAOYSA-N 0.000 description 2
- LYCAIKOWRPUZTN-UHFFFAOYSA-N Ethylene glycol Chemical compound OCCO LYCAIKOWRPUZTN-UHFFFAOYSA-N 0.000 description 2
- 244000275012 Sesbania cannabina Species 0.000 description 2
- QAOWNCQODCNURD-UHFFFAOYSA-N Sulfuric acid Chemical compound OS(O)(=O)=O QAOWNCQODCNURD-UHFFFAOYSA-N 0.000 description 2
- QCWXUUIWCKQGHC-UHFFFAOYSA-N Zirconium Chemical compound [Zr] QCWXUUIWCKQGHC-UHFFFAOYSA-N 0.000 description 2
- 229910052681 coesite Inorganic materials 0.000 description 2
- 238000002485 combustion reaction Methods 0.000 description 2
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- GKOZUEZYRPOHIO-UHFFFAOYSA-N iridium atom Chemical compound [Ir] GKOZUEZYRPOHIO-UHFFFAOYSA-N 0.000 description 2
- 238000004898 kneading Methods 0.000 description 2
- 238000000465 moulding Methods 0.000 description 2
- 239000011858 nanopowder Substances 0.000 description 2
- 239000003208 petroleum Substances 0.000 description 2
- BASFCYQUMIYNBI-UHFFFAOYSA-N platinum Chemical compound [Pt] BASFCYQUMIYNBI-UHFFFAOYSA-N 0.000 description 2
- 238000004321 preservation Methods 0.000 description 2
- 230000000630 rising effect Effects 0.000 description 2
- 239000000377 silicon dioxide Substances 0.000 description 2
- VYPSYNLAJGMNEJ-UHFFFAOYSA-N silicon dioxide Inorganic materials O=[Si]=O VYPSYNLAJGMNEJ-UHFFFAOYSA-N 0.000 description 2
- 229910052682 stishovite Inorganic materials 0.000 description 2
- 238000003860 storage Methods 0.000 description 2
- 230000002194 synthesizing effect Effects 0.000 description 2
- 229910052905 tridymite Inorganic materials 0.000 description 2
- 229910052726 zirconium Inorganic materials 0.000 description 2
- QTBSBXVTEAMEQO-UHFFFAOYSA-M Acetate Chemical compound CC([O-])=O QTBSBXVTEAMEQO-UHFFFAOYSA-M 0.000 description 1
- 229910000975 Carbon steel Inorganic materials 0.000 description 1
- 229920002134 Carboxymethyl cellulose Polymers 0.000 description 1
- VGGSQFUCUMXWEO-UHFFFAOYSA-N Ethene Chemical compound C=C VGGSQFUCUMXWEO-UHFFFAOYSA-N 0.000 description 1
- 239000005977 Ethylene Substances 0.000 description 1
- 240000008042 Zea mays Species 0.000 description 1
- 235000005824 Zea mays ssp. parviglumis Nutrition 0.000 description 1
- 235000002017 Zea mays subsp mays Nutrition 0.000 description 1
- IKHGUXGNUITLKF-XPULMUKRSA-N acetaldehyde Chemical compound [14CH]([14CH3])=O IKHGUXGNUITLKF-XPULMUKRSA-N 0.000 description 1
- 239000004480 active ingredient Substances 0.000 description 1
- PNEYBMLMFCGWSK-UHFFFAOYSA-N aluminium oxide Inorganic materials [O-2].[O-2].[O-2].[Al+3].[Al+3] PNEYBMLMFCGWSK-UHFFFAOYSA-N 0.000 description 1
- QVGXLLKOCUKJST-UHFFFAOYSA-N atomic oxygen Chemical compound [O] QVGXLLKOCUKJST-UHFFFAOYSA-N 0.000 description 1
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- 238000009835 boiling Methods 0.000 description 1
- 239000006227 byproduct Substances 0.000 description 1
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- 238000006555 catalytic reaction Methods 0.000 description 1
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- HRYZWHHZPQKTII-UHFFFAOYSA-N chloroethane Chemical compound CCCl HRYZWHHZPQKTII-UHFFFAOYSA-N 0.000 description 1
- 239000003245 coal Substances 0.000 description 1
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- WGCNASOHLSPBMP-UHFFFAOYSA-N hydroxyacetaldehyde Natural products OCC=O WGCNASOHLSPBMP-UHFFFAOYSA-N 0.000 description 1
- 239000001866 hydroxypropyl methyl cellulose Substances 0.000 description 1
- 235000010979 hydroxypropyl methyl cellulose Nutrition 0.000 description 1
- UFVKGYZPFZQRLF-UHFFFAOYSA-N hydroxypropyl methyl cellulose Chemical compound OC1C(O)C(OC)OC(CO)C1OC1C(O)C(O)C(OC2C(C(O)C(OC3C(C(O)C(O)C(CO)O3)O)C(CO)O2)O)C(CO)O1 UFVKGYZPFZQRLF-UHFFFAOYSA-N 0.000 description 1
- 230000002401 inhibitory effect Effects 0.000 description 1
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- 229910052750 molybdenum Inorganic materials 0.000 description 1
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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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- Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)
Abstract
The utility model provides a process system for indirectly producing ethanol by using synthesis gases. The process system comprises a methanol reactor, a flash tank, a reaction rectifying tower, a carbonylation reactor, a methyl ether recovery tower, a methyl acetate rectifying tower, a hydrogenation reactor, a methyl acetate recovery tower and an ethanol product tower. Hydrogen and carbon monoxide in the methanol reactor are catalyzed to synthesize methanol, coarse methanol products are roughly sorted through the flash tank and then enter the reaction rectifying tower to have a dehydration reaction to generate methyl ether, the methyl ether, the carbon monoxide and the hydrogen are mixed and then filled into the carbonylation reactor to have a carbonylation reaction to generate methyl acetate, coarse methyl acetate products are separated and purified and then are filled into the hydrogenation reactor for hydrogenation to generate ethanol, and coarse ethanol products are purified to obtain ethanol products. Catalysts are adopted, the process and devices of the process system have the advantages of high single-pass conversion and large reaction heat effective use ratio, coarse product separation load is greatly reduced, the production flow is shortened, and production energy consumption is greatly reduced.
Description
Technical Field
The utility model relates to a process for indirectly producing ethanol by synthesis gas, which belongs to the field of ethanol production by synthesis gas.
Background
Ethanol is an important basic chemical raw material, is used for manufacturing acetaldehyde, ethylene, ethylamine, ethyl acetate, acetic acid, chloroethane and the like, derives a plurality of intermediates of products such as medicines, dyes, coatings, spices, synthetic rubber, detergents, pesticides and the like, and produces more than 300 products. Meanwhile, in the application of fuel energy industry, ethanol has the characteristics of low heat value, high vaporization latent heat, good anti-explosion performance, high oxygen content, easy phase separation in the presence of a small amount of water, more complete combustion, low CO emission, similar combustion performance and the like compared with common gasoline, so that the ethanol is called as 'green energy' in the 21 st century. Currently, due to non-renewable petroleum and unstable petroleum producing areas, fuel energy safety issues are drawing more and more attention on the global scale, and the popularization and application of fuel ethanol are actively carried out in all countries of the world, and 60% of ethanol production in the world is used as vehicle fuel. Therefore, the development of energy industries using alcohol as a raw material, such as ethanol and ethanol diesel, has been a major issue in the international fuel energy industry.
At present, grains are mainly used as raw materials for producing fuel ethanol, corn fuel ethanol becomes an important energy supply source in many countries, but due to the rising of the price of grains, the price of ethanol is continuously rising. At present, the price of absolute ethyl alcohol is nearly 8000 yuan/ton, so that a non-food process route is a very intelligent choice.
Patent CN1122567C discloses a method for directly producing ethanol by hydrogen and carbon monoxide synthesis. The method has the advantages of simple process route, high energy efficiency and easy scale production from the technical principle, and is an ideal synthetic route. However, 1 ton of ethanol is produced from 1.4 ton of synthesis gas in the process, the conversion rate of the synthesis gas is only 71%, the byproduct of ethanol production from the synthesis gas is water, and about 1/4 of hydrogen in the synthesis gas is lost. The product components are complex, and the crude product also needs hydrogenation and refining treatment. The development of the catalyst for directly preparing ethanol from the synthesis gas still stays in a small test stage at present, and the catalyst is expensive noble metal (rhodium and the like) catalyst, so that the problems of low reaction rate and the like exist, and the problems are technical bottlenecks which cannot be broken through at present and for a long time in the future.
Patent CN102317412 discloses a method for preparing ethanol by syngas biological method, i.e. a microbial fermentation technology is utilized to produce ethanol from syngas (including waste gas containing carbon monoxide and hydrogen), because fermentation needs a period of time, continuous production is difficult to realize, the cost is high at present, and the method is only in the concept design stage, and large-scale industrialization needs to be verified.
Patent CN101952231 discloses a method for preparing ethanol from synthesis gas by acetic acid esterification and hydrogenation. The process comprises the steps of firstly converting synthesis gas into methanol, then carbonylating the methanol and the synthesis gas to produce acetic acid, then producing acetic ester with the methanol or ethanol, and finally hydrogenating the acetic ester to produce ethanol. In the process, noble metals such as rhodium, iridium and the like are adopted in the methanol carbonylation process, expensive zirconium materials or Hastelloy are adopted as equipment, the one-time investment of the equipment is large, the reaction route is long, and the energy consumption of the device is high.
The process route of preparing ethanol from synthesis gas by acetic acid hydrogenation is the mainstream process route of preparing ethanol from coal at present. The process synthesizes CO and CO2And H2First methanol is prepared, then methanol is carbonylatedProducing acetic acid, and finally hydrogenating the acetic acid to prepare the ethanol. By adopting the process, 1.3034t of acetic acid and 972.4Nm3 of hydrogen are theoretically consumed per ton of ethanol, and 0.391 ton of water is generated. At present, the price of acetic acid is 2600-3600 yuan/ton, the price is relatively cheap, and the production technology of the acetic acid is mature, so that the ethanol produced by acetic acid hydrogenation is expected to be large-scale and large-scale. However, in the process route, noble metals such as rhodium, iridium and the like are adopted in the methanol carbonylation process; in the technology for preparing ethanol by hydrogenating acetic acid, the platinum/tin catalyst is utilized to directly and selectively produce ethanol from acetic acid (patent USP7863489), the hydrogenation process needs to be completed in two steps, and the economical efficiency needs to be verified. Meanwhile, the process has higher requirements on materials and larger investment, for example, expensive zirconium materials or hastelloy are adopted in a reactor for preparing acetic acid by methanol carbonylation and equipment of an acetic acid hydrogenation reactor.
SUMMERY OF THE UTILITY MODEL
The utility model aims at overcoming the defects in the prior art and providing a process system for indirectly producing ethanol by synthesis gas.
The utility model adopts the following technical scheme:
a process system for indirectly producing ethanol by synthesis gas comprises a methanol reactor, a flash tank, a reaction rectifying tower, a carbonylation reactor, a dimethyl ether recovery tower, a methyl acetate rectifying tower, a hydrogenation reactor, a methyl acetate recovery tower and an ethanol product tower; wherein: the bottom outlet of the methanol reactor is connected with the inlet of the flash tank through a pipeline; the liquid phase outlet of the flash tank is connected with the lower inlet of the reactive distillation column through a pipeline; the top outlet of the reaction rectifying tower is connected with the top inlet of the carbonylation reactor through a pipeline; an outlet at the bottom of the carbonylation reactor is connected with an inlet of the dimethyl ether recovery tower through a pipeline; the bottom outlet of the dimethyl ether recovery tower is connected with the inlet of the methyl acetate rectifying tower through a pipeline; the top outlet of the methyl acetate rectifying tower is connected with the top inlet of the hydrogenation reactor through a pipeline; the bottom outlet of the hydrogenation reactor is connected with the inlet of the methyl acetate recovery tower through a pipeline, and the bottom outlet of the methyl acetate recovery tower is connected with the inlet of the ethanol product tower through a pipeline; and an outlet at the bottom of the ethanol product tower is a product ethanol discharge hole.
The methanol reactor is a fixed bed reactor or a fluidized bed reactor; preferably a fixed bed reactor; particularly preferred is a plate-type fixed bed reactor.
The methanol reactor is filled with a methanol synthesis catalyst, the methanol synthesis catalyst is a Cu-Zn-Al catalyst, and the carrier of the catalyst is Al2O3: 10-30 wt%, and the active ingredients and contents of the catalyst are respectively CuO: 5-30 wt% and ZnO: 10 to 70 wt%.
Reacting technical grade carbon monoxide and hydrogen in the presence of the methanol synthesis catalyst in the methanol reactor to produce methanol.
Further, a methanol reactor outlet heat exchanger and a methanol reactor inlet heat exchanger are arranged between a bottom outlet and a top inlet of the methanol reactor; the bottom outlet of the methanol reactor is connected with the hot material flow inlet of the outlet heat exchanger of the methanol reactor through a pipeline; a cold flow outlet of the methanol reactor outlet heat exchanger is connected with a cold flow inlet of the methanol reactor inlet heat exchanger through a pipeline; a cold flow outlet of the inlet heat exchanger of the methanol reactor is connected with a top inlet of the methanol reactor through a pipeline; and a cold material flow inlet of the outlet heat exchanger of the methanol reactor is connected with a synthesis gas raw material pipeline through a pipeline.
The synthesis gas raw material pipeline is connected with the hydrogen main pipe and the carbon monoxide main pipe through pipelines.
And carrying out heat exchange on a reaction product led out from the bottom of the methanol reactor and the synthesis gas raw material in an outlet heat exchanger of the methanol reactor, heating the synthesis gas raw material to the inlet temperature of a catalyst bed layer of the methanol reactor through an inlet heat exchanger of the methanol reactor, and then entering the methanol reactor to carry out catalytic synthesis methanol reaction.
Further, a hot liquid outlet of the outlet heat exchanger of the methanol reactor is connected with a first gas-liquid separator through a pipeline; the bottom liquid phase outlet of the first gas-liquid separator is connected with the inlet of the flash tank through a pipeline; and a top gas-phase outlet of the first gas-liquid separator is connected with a cold material flow inlet of the methanol reactor outlet heat exchanger through a pipeline.
Further, a first compressor is connected between the top gas-phase outlet of the first gas-liquid separator and the cold material flow inlet of the methanol reactor outlet heat exchanger through a pipeline.
Furthermore, a top gas-phase outlet of the first gas-liquid separator is also connected with a first non-condensable gas discharge pipeline.
And gas phase discharged from the top of the first gas-liquid separator passes through a first non-condensable gas discharge pipeline, discharges partial non-condensable gas, is pressurized by the first compressor, exchanges heat by the methanol reactor outlet heat exchanger and the methanol reactor inlet heat exchanger, and circulates to the methanol reactor for continuous reaction.
Furthermore, a first water cooler is arranged between a hot material outlet of the methanol reactor outlet heat exchanger and an inlet of the first gas-liquid separator.
Further, a pressure reducing valve is connected between the bottom liquid phase outlet of the first gas-liquid separator and the inlet of the flash tank through a pipeline.
Further, a preheater is connected between the liquid phase outlet of the flash tank and the lower inlet of the reactive distillation column through a pipeline. And the liquid phase led out from the flash tank is preheated by the preheater to the bubble point temperature of the methanol and then enters the reaction rectifying tower for reaction.
Furthermore, a heat preservation tank is connected between the cold material flow outlet of the preheater and the lower inlet of the reactive distillation column through a pipeline.
Further, the outlet of the tower top of the reaction rectifying tower is connected with a condensation reflux device through a pipeline; the outlet of the condensation reflux device is connected with the top inlet of the carbonylation reactor through a pipeline; the condensation reflux device is also provided with a non-condensable gas discharge pipeline.
Furthermore, a material leading-out port is arranged on the side part of the top of the reactive distillation column, and the material leading-out port is connected with a connecting pipeline between the preheater and the lower inlet of the reactive distillation column through a pipeline.
Furthermore, the number of theoretical plates of the reactive distillation column is 9-50, preferably 30-50, and particularly preferably 40-50. The material extraction port is positioned at the 1 st to 5 th tower plates of the reaction rectifying tower, preferably the 1 st to 2 nd tower plates.
Solid super acidic filler is filled in the reaction rectifying tower; the solid super acidic filler is a ceramic filler loaded with solid super acid; the solid super acid is an acid type catalyst with the highest acid strength in the solid acid, and does not corrode equipment. The Hammett function is: -16.02<H0<14.520, corresponding to 10000 times of 100% liquid sulfuric acid. The solid super acid is selected from SO4 2-The preparation method of the/beta molecular sieve composite catalyst comprises the following steps:
(1) will (NH4)2S2O8Dipping on a nanometer beta molecular sieve to prepare nanometer catalyst powder;
(2) adding ceramic pug and a bonding auxiliary agent for mixing, kneading and molding;
(3) and drying and roasting to obtain the solid super acidic filler.
Wherein,
（NH4)2S2O8after being dipped on the nano beta molecular sieve, SO is formed after being roasted4 2-The reason for the/beta molecular sieve composite catalyst is (NH)4)2S2O8Decomposition at calcination temperature of NH4 +Comes out in a hydration mode, and the residual SO4 2-A/beta molecular sieve.
Preferably, the particle size of the nano beta molecular sieve in the step (1) is 68-80 nm.
Preferably, the mixing in step (2) is spraying or mechanical mixing, preferably mechanical mixing.
Preferably, the binding aid in step (2) is selected from hydroxypropyl methylcellulose, sesbania powder and carboxymethyl cellulose.
Preferably, the mixing mass ratio of the catalyst powder, the ceramic mud and the bonding auxiliary agent is 45-89.8: 10-50: 0.2 to 5, preferably, 50 to 85: 10-49: 1-5.
Preferably, the roasting temperature in the step (3) is 500-600 ℃, and the roasting time is 3-6 h.
Preferably, the size of the solid super acid filler is 3-5 mm multiplied by 3-5 mm, and preferably 5mm multiplied by 5 mm.
Preferably, the filler particles are loaded with 0.58-0.62 wt% of super acid active centers on average.
The solid super acidic packing is random packing or regular packing; the shape of the random packing comprises a saddle shape, a Raschig ring, a pall ring, a wheel shape, a rectangular saddle ring, a spherical shape or a columnar shape.
The shape of the structured packing comprises a corrugated plate shape or a honeycomb shape.
Preferably, the solid super acid filler is a random saddle type filler.
And in the reaction rectifying tower, the methanol liquid and the steam are in continuous countercurrent contact on the solid super acidic filler to carry out the reaction rectifying process for preparing the dimethyl ether by dehydrating the methanol.
Further, a tower bottom outlet of the reaction rectifying tower is connected with a methanol recovery tower through a pipeline; and the outlet at the top of the methanol recovery tower is connected with a connecting pipeline between the flash tank and the preheater through a pipeline.
Furthermore, a carbonylation reactor outlet heat exchanger and a carbonylation reactor inlet heat exchanger are arranged between the bottom outlet and the top inlet of the carbonylation reactor; the bottom outlet of the carbonylation reactor is connected with the hot material flow inlet of the outlet heat exchanger of the carbonylation reactor through a pipeline; the cold flow outlet of the outlet heat exchanger of the carbonylation reactor is connected with the cold flow inlet of the inlet heat exchanger of the carbonylation reactor through a pipeline; the cold flow outlet of the carbonylation reactor inlet heat exchanger is connected with the top inlet of the carbonylation reactor through a pipeline; and a cold material flow inlet of the outlet heat exchanger of the carbonylation reactor is connected with an outlet at the top of the reactive distillation column, a hydrogen main pipe and a carbon dioxide main pipe through pipelines.
Dimethyl ether from the top outlet of the reaction rectifying tower, hydrogen and dioxide from a hydrogen main pipe and a carbon dioxide main pipe and a reaction product led out from the bottom outlet of the carbonylation reactor exchange heat in an outlet heat exchanger of the carbonylation reactor, and then the dimethyl ether, the hydrogen and the dioxide and the reaction product are heated to the inlet temperature of a catalyst bed layer of the carbonylation reactor through an inlet heat exchanger of the carbonylation reactor and enter the carbonylation reactor to carry out the reaction of catalytically synthesizing methyl acetate.
The carbonylation reactor is a fixed bed reactor or a fluidized bed reactor, preferably a fixed bed reactor, and particularly preferably a plate-type fixed bed reactor.
The carbonylation reactor is filled with a carbonylation catalyst which is H-MOR, the contents of active auxiliary agent and CuO 0-30 wt%, CoO0.1-10 wt%, MoO0.1-10 wt%, preferably CuO 5-15 wt%, CoO0.5-6 wt%, Mo2O30.5～8wt%。
In the carbonylation reactor, in the presence of the carbonylation catalyst, the carbonylation reaction between dimethyl ether and hydrogen and carbon monoxide is carried out to prepare methyl acetate.
Further, a hot material flow outlet of the outlet heat exchanger of the carbonylation reactor is connected with a second gas-liquid separator through a pipeline; a bottom liquid phase outlet of the second gas-liquid separator is connected with an inlet of the dimethyl ether recovery tower through a pipeline; and the top gas-phase outlet of the second gas-liquid separator is connected with the cold material flow inlet of the outlet heat exchanger of the carbonylation reactor through a pipeline.
Further, a second compressor is connected between the top gas-phase outlet of the second gas-liquid separator and the cold flow inlet of the outlet heat exchanger of the carbonylation reactor through a pipeline.
Further, a top gas-phase outlet of the second gas-liquid separator is also connected with a second non-condensable gas discharge pipeline.
And the gas phase discharged from the top of the second gas-liquid separator discharges part of non-condensable gas through the second non-condensable gas discharge pipeline, then the gas phase is pressurized by the second compressor, and then the gas phase is circulated to the carbonylation reactor for continuous reaction after heat exchange is carried out by the outlet heat exchanger of the carbonylation reactor and the inlet heat exchanger of the carbonylation reactor.
Furthermore, a second water cooler is arranged between the hot material flow outlet of the outlet heat exchanger of the carbonylation reactor and the inlet of the second gas-liquid separator.
Further, a top gas-phase outlet of the second gas-liquid separator is connected with a non-condensable gas discharge pipeline.
Furthermore, the top outlet of the dimethyl ether recovery tower is connected with a connecting pipeline between the top outlet of the reactive distillation tower and the cold material flow inlet of the outlet heat exchanger of the carbonylation reactor through a pipeline.
Further, the hydrogenation reactor is a fixed bed reactor or a fluidized bed reactor, preferably a fixed bed reactor, and particularly preferably a plate-type fixed bed reactor.
Further, a hydrogenation catalyst is filled in the hydrogenation reactor; the carrier of the hydrogenation catalyst is SiO2The active component and the content of the hydrogenation catalyst are respectively CuO: 5-50 wt%, ZnO: 0 to 20wt% and Mo2O3：0～5wt%。
Further, a hydrogenation reactor outlet heat exchanger and a hydrogenation reactor inlet heat exchanger are arranged between the bottom outlet and the top inlet of the hydrogenation reactor; the bottom outlet of the hydrogenation reactor is connected with the hot material flow inlet of the outlet heat exchanger of the hydrogenation reactor through a pipeline; a cold flow outlet of the outlet heat exchanger of the hydrogenation reactor is connected with a cold flow inlet of the inlet heat exchanger of the hydrogenation reactor through a pipeline; a cold material flow outlet of the hydrogenation reactor inlet heat exchanger is connected with a top inlet of the hydrogenation reactor through a pipeline; and a cold material flow inlet of the outlet heat exchanger of the hydrogenation reactor is connected with a tower top outlet of the methyl acetate rectifying tower and a hydrogen main pipe through pipelines.
And the methyl acetate from the top outlet of the methyl acetate rectifying tower, the hydrogen from the hydrogen main pipe and the reaction product led out from the bottom outlet of the hydrogenation reactor exchange heat in the outlet heat exchanger of the hydrogenation reactor, and then the reaction product is heated to the inlet temperature of the catalyst bed layer of the hydrogenation reactor through the inlet heat exchanger of the hydrogenation reactor and enters the hydrogenation reactor to perform the reaction of catalytically synthesizing ethanol.
Further, a hot material outlet of the outlet heat exchanger of the hydrogenation reactor is connected with a third gas-liquid separator through a pipeline; a bottom liquid phase outlet of the third gas-liquid separator is connected with an inlet of the methyl acetate recovery tower through a pipeline; and a top gas-phase outlet of the third gas-liquid separator is connected with a cold material flow inlet of the outlet heat exchanger of the hydrogenation reactor through a pipeline.
Further, a third compressor is connected between a top gas-phase outlet of the third gas-liquid separator and a cold material flow inlet of the outlet heat exchanger of the hydrogenation reactor through a pipeline.
Further, a top gas-phase outlet of the third gas-liquid separator is also connected with a third noncondensable gas discharge pipeline.
And discharging gas phase at the top of the third gas-liquid separator, discharging partial non-condensable gas through a third non-condensable gas discharge pipeline, then pressurizing the gas phase by the third compressor, exchanging heat by the outlet heat exchanger of the hydrogenation reactor and the inlet heat exchanger of the hydrogenation reactor, and circulating the gas phase to the hydrogenation reactor for continuous reaction.
Further, a third water cooler is arranged between the hot material outlet of the outlet heat exchanger of the hydrogenation reactor and the third gas-liquid separator.
Further, the outlet of the tower top of the methyl acetate recovery tower is connected with a one-stage or multi-stage condensing device through a pipeline; the liquid phase outlet of the condensing device is connected with the cold material flow inlet of the outlet heat exchanger of the hydrogenation reactor through a pipeline; and a gas phase outlet of the condensing device is connected with a cold material inlet of the outlet heat exchanger of the carbonylation reactor through a pipeline.
Further, the outlet of the top of the ethanol product tower is connected with the cold material flow inlet of the preheater through a pipeline.
Further, the flash tank is replaced by two rectifying towers or three rectifying towers which are connected in sequence through pipelines.
Compared with the prior art, the beneficial effects of the utility model are that:
the invention relates to a process system for indirectly producing ethanol by using synthesis gas, which comprises the steps of firstly preparing methanol from the synthesis gas in a methanol reactor of the system, then catalytically dehydrating, catalytically rectifying and generating dimethyl ether by using methanol, then carbonylating the dimethyl ether to generate Methyl Acetate (MA) by using hydrogen mordenite (H-MOR), further hydrogenating the generated methyl acetate on a copper-based catalyst to generate ethanol and methanol, and recycling the methanol. The dimethyl ether process unit adopts a catalytic rectification method, realizes the combination of catalytic reaction and separation coupling mode, saves equipment investment and separation energy consumption, simplifies the flow, effectively utilizes reaction heat, and timely separates products, thereby being beneficial to the balance of methanol dehydration reaction and inhibiting the occurrence of side reaction; due to the existence of methanol dehydration reaction, the method is more beneficial to the separation of azeotrope in a boiling point approaching or reaction system, has high conversion per pass, greatly lightens the load of the separation of crude products, shortens the production flow and greatly reduces the production energy consumption; the process unit for preparing methyl acetate by dimethyl ether carbonylation has the advantages that the catalyst is a conventional non-noble metal catalyst, the reaction condition is mild, the corrosivity of reaction materials is low, the equipment investment is saved, the process mode is novel, and the production efficiency is high; the process unit for preparing the ethanol by acetate hydrogenation has the advantages of mild reaction conditions, easy separation, low energy consumption, less investment, low energy consumption, conventional copper-based composite catalyst and low cost. Meanwhile, the corrosion of the raw materials and the products is weak, so that the carbon steel can be adopted, and the investment amount is greatly reduced. The operation cost of the process system is low, and particularly, the catalysts involved in the process are common non-noble metal catalysts, so that the process system has obvious advantages compared with the common noble metal catalysts used in other process routes. According to the estimation, the investment of the process equipment is only 1/2 or even lower when acetic acid is directly hydrogenated to prepare ethanol, the production cost is lower than that of the conventional technology for preparing ethanol by a biological fermentation method, and the process is the most possible industrialized process route at present. The method has important significance for solving the problems that the dimethyl ether productivity is seriously excessive and the production and sale prices are seriously hung upside down at present.
The catalyst adopted by the invention simplifies the intermediate fractionation process due to high conversion per pass and good selectivity. The process flow setting of the system fully considers the coupling of the heat release unit and the heat utilization unit, and can greatly reduce the comprehensive energy consumption. The method has the advantages of less equipment investment, environmental protection, simple preparation process, economy and practicality.
Drawings
FIG. 1 is a simplified schematic of a process system according to one embodiment of FIG. 1;
FIG. 2 is a simplified schematic of a process system according to one embodiment of FIG. 2;
FIG. 3 is a simplified schematic diagram of the process system of one embodiment of FIG. 3;
reference numerals: a, a methanol reactor; b, a methanol reactor inlet heat exchanger; c, a methanol reactor outlet heat exchanger; d, a first water cooler; e a first gas-liquid separator; f, flash evaporation tank; g, a preheater; h reaction rectifying tower; i, a methanol recovery tower; a J carbonylation reactor; an inlet heat exchanger of the K carbonylation reactor; an outlet heat exchanger of the L carbonylation reactor; m, a second water cooler; n second gas-liquid separator; an O dimethyl ether recovery tower; a P methyl acetate rectifying tower; q, a hydrogenation reactor; an inlet heat exchanger of the hydrogenation reactor; s, an outlet heat exchanger of a hydrogenation reactor; a third water cooler; a U third gas-liquid separator; v, a methyl acetate recovery tower; w ethanol product tower.
Detailed Description
The present invention will be further described with reference to the following specific examples. It should be understood that these examples are for illustrative purposes only and are not intended to limit the scope of the present invention. Furthermore, it should be understood that any changes or modifications of the present invention may be made by those skilled in the art after reading the teachings of the present invention, and such equivalents also fall within the scope of the appended claims.
Experimental procedures without specific conditions noted in the examples below, generally following conventional conditions, such as: a manual for chemical operations, or according to the conditions recommended by the manufacturer.
As shown in fig. 1-2, the utility model provides a process system for indirectly producing ethanol by synthesis gas, which comprises a methanol reactor A, a flash tank F, a reaction rectifying tower H, a carbonylation reactor J, a dimethyl ether recovery tower O, a methyl acetate rectifying tower P, a hydrogenation reactor Q, a methyl acetate recovery tower V and an ethanol product tower W; wherein: the bottom outlet of the methanol reactor A is connected with the inlet of the flash tank F through a pipeline; the liquid phase outlet of the flash tank F is connected with the lower inlet of the reaction rectifying tower H through a pipeline; the top outlet of the reaction rectifying tower H is connected with the top inlet of the carbonylation reactor J through a pipeline; an outlet at the bottom of the carbonylation reactor J is connected with an inlet of the dimethyl ether recovery tower O through a pipeline; the bottom outlet of the dimethyl ether recovery tower O is connected with the inlet of the methyl acetate rectifying tower P through a pipeline; the top outlet of the methyl acetate rectifying tower P is connected with the top inlet of the hydrogenation reactor Q through a pipeline; the bottom outlet of the hydrogenation reactor Q is connected with the inlet of a methyl acetate recovery tower V through a pipeline, and the bottom outlet of the methyl acetate recovery tower V is connected with the inlet of an ethanol product tower W through a pipeline; the outlet at the bottom of the ethanol product tower W is a product ethanol discharge hole.
The methanol reactor A is a fixed bed reactor or a fluidized bed reactor, preferably a fixed bed reactor; particularly preferred is a plate-type fixed bed reactor. The methanol reactor A is filled with a methanol synthesis catalyst.
And filling solid super acidic filler in the reaction rectifying tower H.
The carbonylation reactor J is a fixed bed reactor or a fluidized bed reactor, preferably a fixed bed reactor, particularly preferably a plate-type fixed bed reactor. The carbonylation reactor J is filled with a carbonylation catalyst.
The hydrogenation reactor Q is a fixed bed reactor or a fluidized bed reactor, preferably a fixed bed reactor, and particularly preferably a plate-type fixed bed reactor. The hydrogenation reactor Q is filled with a hydrogenation catalyst.
As a preferred embodiment: a methanol reactor outlet heat exchanger C and a methanol reactor inlet heat exchanger B are arranged between the bottom outlet and the top inlet of the methanol reactor A; the bottom outlet of the methanol reactor A is connected with the hot material flow inlet of the outlet heat exchanger C of the methanol reactor through a pipeline; a cold material flow outlet of the methanol reactor outlet heat exchanger C is connected with a cold material flow inlet of the methanol reactor inlet heat exchanger B through a pipeline; a cold material flow outlet of the inlet heat exchanger B of the methanol reactor is connected with a top inlet of the methanol reactor A through a pipeline; and a cold material flow inlet of the methanol reactor outlet heat exchanger C is connected with a synthesis gas raw material pipeline through a pipeline.
The synthesis gas raw material pipeline is connected with the hydrogen main pipe and the carbon monoxide main pipe through pipelines.
As another preferred embodiment, the hot liquid outlet of the outlet heat exchanger B of the methanol reactor is connected with a first gas-liquid separator E through a pipeline; a bottom liquid phase outlet of the first gas-liquid separator E is connected with an inlet of the flash tank F through a pipeline; the top gas phase outlet of the first gas-liquid separator E is connected with the cold material flow inlet of the methanol reactor outlet heat exchanger C through a pipeline.
As another preferred embodiment, a first compressor is connected between the top gas phase outlet of the first gas-liquid separator E and the C cold material flow inlet of the outlet heat exchanger of the methanol reactor through a pipeline; and a top gas-phase outlet of the first gas-liquid separator E is also connected with a first non-condensable gas discharge pipeline.
As another preferred embodiment, a first water cooler D is arranged between the hot material outlet of the outlet heat exchanger C of the methanol reactor and the first gas-liquid separator E.
As another preferable embodiment, a pressure reducing valve is connected between the bottom liquid phase outlet of the first gas-liquid separator E and the inlet of the flash drum F through a line.
As another preferred embodiment, the top outlet of the reactive distillation column H is connected with a condensing reflux device through a pipeline; the outlet of the condensation reflux device is connected with the inlet at the top of the carbonylation reactor J through a pipeline; the condensation reflux device is also provided with a non-condensable gas discharge pipeline.
As another preferred embodiment, a preheater G is connected between the liquid phase outlet of the flash drum F and the lower inlet of the reactive distillation column H via a line.
As another preferred embodiment, a material leading-out port is arranged on the side part of the top of the reactive distillation column H, and the material leading-out port is connected with a connecting pipeline between the preheater G and the lower inlet of the reactive distillation column H through a pipeline.
As another preferred embodiment, the bottom outlet of the reactive distillation tower H is connected with a methanol recovery tower I through a pipeline; the top outlet of the methanol recovery tower I is connected with a connecting pipeline between the flash tank F and the preheater G through a pipeline.
In another preferred embodiment, a carbonylation reactor outlet heat exchanger L and a carbonylation reactor inlet heat exchanger K are arranged between a bottom outlet and a top inlet of the carbonylation reactor J; the bottom outlet of the carbonylation reactor J is connected with the hot material flow inlet of the outlet heat exchanger L of the carbonylation reactor through a pipeline; a cold material flow outlet of the outlet heat exchanger L of the carbonylation reactor is connected with a cold material flow inlet of the inlet heat exchanger K of the carbonylation reactor through a pipeline; the cold material flow outlet of the inlet heat exchanger K of the carbonylation reactor is connected with the top inlet of the carbonylation reactor J through a pipeline; and a cold material flow inlet of the outlet heat exchanger L of the carbonylation reactor is connected with the top outlet of the reactive distillation column H, a hydrogen main pipe and a carbon dioxide main pipe through pipelines.
In another preferred embodiment, the hot material flow outlet of the outlet heat exchanger L of the carbonylation reactor is connected with a second gas-liquid separator N through a pipeline; a bottom liquid phase outlet of the second gas-liquid separator N is connected with an inlet of the dimethyl ether recovery tower O through a pipeline; the top gas phase outlet of the second gas-liquid separator N is connected with the cold material flow inlet of the outlet heat exchanger L of the carbonylation reactor through a pipeline.
In another preferred embodiment, a second compressor is connected between the top gas phase outlet of the second gas-liquid separator N and the cold material flow inlet of the carbonylation reactor outlet heat exchanger L through a pipeline.
In another preferred embodiment, the top gas-phase outlet of the second gas-liquid separator N is further connected to a second non-condensable gas discharge line.
As another preferred embodiment, the top outlet of the dimethyl ether recovery column O is connected with a connecting pipeline between the top outlet of the reactive distillation column H and the cold material flow inlet of the carbonylation reactor outlet heat exchanger L through a pipeline.
As another preferred embodiment, a hydrogenation reactor outlet heat exchanger S and a hydrogenation reactor inlet heat exchanger R are arranged between the bottom outlet and the top inlet of the hydrogenation reactor Q; the bottom outlet of the hydrogenation reactor Q is connected with the hot material flow inlet of the outlet heat exchanger S of the hydrogenation reactor through a pipeline; a cold material flow outlet of the outlet heat exchanger S of the hydrogenation reactor is connected with a cold material flow inlet of the inlet heat exchanger R of the hydrogenation reactor through a pipeline; a cold material flow outlet of the hydrogenation reactor inlet heat exchanger R is connected with a top inlet of the hydrogenation reactor Q through a pipeline; and a cold material flow inlet of the outlet heat exchanger S of the hydrogenation reactor is connected with a tower top outlet of the methyl acetate rectifying tower P and a hydrogen main pipe through pipelines.
In another preferred embodiment, the hot material flow outlet of the outlet heat exchanger S of the hydrogenation reactor is connected with a third gas-liquid separator U through a pipeline; a bottom liquid phase outlet of the third gas-liquid separator U is connected with an inlet of the methyl acetate recovery tower V through a pipeline; and a top gas-phase outlet of the third gas-liquid separator U is connected with a cold material flow inlet of the outlet heat exchanger S of the hydrogenation reactor through a pipeline.
In another preferred embodiment, a third compressor is connected between the top gas-phase outlet of the third gas-liquid separator U and the cold stream inlet of the outlet heat exchanger S of the hydrogenation reactor through a pipeline.
In another preferred embodiment, the top gas-phase outlet of the third gas-liquid separator U is further connected to a third noncondensable gas discharge line.
In another preferred embodiment, a third water cooler T is arranged between the hot material flow outlet of the outlet heat exchanger S of the hydrogenation reactor and the third gas-liquid separator U.
As another preferred embodiment, the outlet of the top of the methyl acetate recovery tower V is connected with one-stage or multi-stage condensing equipment through a pipeline; a liquid phase outlet of the condensing device is connected with a cold material flow inlet of a hydrogenation reactor outlet heat exchanger S through a pipeline; the gas phase outlet of the condensing device is connected with the cold material flow inlet of the outlet heat exchanger L of the carbonylation reactor through a pipeline.
As another preferred embodiment: the aforementioned flash tank F may be replaced with a refined methanol preparation unit, i.e., two rectification columns or three rectification columns connected in series by a pipeline, as shown in fig. 3.
The process flow for producing the ethanol by adopting the system is as follows:
CO from line 1 and H from line 22After being merged, enters the pipeline 3 and is combined with the pipelineThe recycle gas from line 8 is mixed and enters line 4. The mixed gas in the pipeline 4 is subjected to continuous heat exchange through a methanol reactor outlet heat exchanger C and a methanol reactor inlet heat exchanger B in sequence, then is introduced from the top of the methanol reactor A through a pipeline 5 to perform catalytic synthesis methanol reaction, and a reaction product enters the methanol reactor outlet heat exchanger C through a pipeline 6 to perform heat exchange with the mixed gas from the pipeline 4, then enters a water cooler D through a pipeline 7 and is cooled to a specified temperature, and then enters a gas-liquid separator E to perform gas-liquid separation. Discharging partial non-condensable gas from the top of the gas-liquid separator E through a pipeline 9, pressurizing, circulating to a pipeline 4 through a pipeline 8, reducing the pressure of bottom liquid-phase crude methanol through a pressure reducing valve through a pipeline 10, and performing gas-liquid phase separation in a flash tank F; the non-condensable gas such as methane, hydrogen and the like separated by the flash tank F is discharged to the outside for recycling through a pipeline 48; the liquid phase crude methanol flows out through a pipeline 11, joins with methanol recovered from a pipeline 45 and a pipeline 18 and then enters a pipeline 12, is preheated to the bubble point temperature by a preheater G, joins with methanol led out from a tower top lateral line 14 of a reaction rectifying tower H, and then is introduced into the lower part of the reaction rectifying tower H through a pipeline 13, and is subjected to a reaction rectifying process for preparing dimethyl ether by dehydrating the methanol in the reaction rectifying tower H filled with solid super strong acid filler; the light components such as dimethyl ether, a small amount of methanol and water at the top of the reactive distillation column H are subjected to three-stage condensation, most of the light components flow back to the column, the non-condensable gas is subjected to water seal at the top of the column and then is sent to an external recovery system through a pipeline 15, and the rest of the non-condensable gas enters a pipeline 16. Unreacted methanol is led out from the top side line 14 of the reactive distillation column H. Heavy components such as methanol, water and the like in the tower bottom of the reaction rectifying tower H enter a methanol recovery tower I through a pipeline 17 for methanol recovery. The methanol extracted from the top of the methanol recovery tower I is condensed and extracted and then is converged into a pipeline 12 through a pipeline 18, and impurities such as water and the like discharged from the tower kettle of the methanol recovery tower I are sent to the outside for treatment. CO from line 49 and H from line 472The mixture is introduced into a line 21, and the dimethyl ether from the line 16 and the recovered dimethyl ether from the line 43 and the line 29 are merged and introduced into a line 20, and are mixed with CO and H from the line 212The mixed raw material gas and the circulating gas from the pipeline 26 are mixed, then enter the outlet heat exchanger L of the carbonylation reactor through the pipeline 22, enter the inlet heat exchanger K of the carbonylation reactor from the pipeline 23 for further heat exchange, and achieve the aim ofThe temperature of the catalyst bed layer inlet of the carbonylation reactor J enters the carbonylation reactor J for carbonylation reaction. The carbonylation reaction product enters a heat exchanger L at the outlet of the carbonylation reactor through a pipeline 24 to exchange heat with the carbonylation reaction raw material, then enters a water cooler M through a pipeline 25 to be further cooled, and enters a gas-liquid separator N to carry out gas-liquid separation. The gas phase discharged from the top of the gas-liquid separator N is partially non-condensable gas discharged from a pipeline 27, then pressurized, circulated to a pipeline 22 through a pipeline 26, and the bottom liquid phase crude ester enters a dimethyl ether recovery tower O through a pipeline 28 for purification. Dimethyl ether and a small amount of light components such as methyl acetate and the like at the top of the dimethyl ether recovery tower O are extracted and pressurized and then are circulated to a pipeline 20 through a pipeline 29, and heavy components containing a small amount of mixed acid and the like extracted at the bottom of the tower enter a methyl acetate rectifying tower P through a pipeline 30 for further purification. Heavy components containing heteropolyacids such as acetic acid and the like extracted from the tower bottom of a methyl acetate rectifying tower P are discharged to the outside through a pipeline 31 for recovery treatment, methyl acetate light components containing a small amount of methanol extracted from the tower top enter a pipeline 32, then are mixed with methyl acetate circulating gas from a pipeline 42 and enter a pipeline 33, then are mixed with hydrogen from a pipeline 34 and circulating gas from a pipeline 39 to form hydrogenation reaction raw material gas, enter a pipeline 35, and are subjected to two-stage continuous heat exchange through a hydrogenation reactor outlet heat exchanger S and a hydrogenation reactor inlet heat exchanger R to reach the inlet temperature of a hydrogenation reactor catalyst bed layer, and then are fed from the top of a hydrogenation reactor Q for hydrogenation reaction. The hydrogenation reaction product is led out from the bottom of the hydrogenation reactor Q, enters a heat exchanger S at the outlet of the hydrogenation reactor through a pipeline 37 for heat exchange, enters a water cooler T through a pipeline 38 for cooling to a specified temperature, and enters a gas-liquid separator U for gas-liquid separation. The gas phase at the top of the gas-liquid separator U is discharged by a pipeline 40, and then is pressurized by a compressor, and then is circulated to a pipeline 35 through a pipeline 39; the liquid phase at the bottom of the gas-liquid separator U enters a methyl acetate recovery tower V through a pipeline 41 for rectification. Light components extracted from the top of the methyl acetate recovery tower V are condensed, non-condensable gas containing dimethyl ether, a small amount of methyl acetate and the like is circulated to the pipeline 20 through a pipeline 43, and condensed liquid is extracted and then circulated to the pipeline 33 through a pipeline 42. Heavy components in a V-tower kettle of the methyl acetate recovery tower are extracted and then enter an ethanol product tower W through a pipeline 44 for further purification; light components such as methanol extracted from the top of the ethanol product tower W are circulated to the pipeline 12 through the pipeline 45, and the mixture is discharged from the bottom of the towerThe ethanol product is sent to an ethanol product storage tank.
Examples of industrial applications for ethanol production using the above process systems and procedures are as follows:
the methanol reactor A is a plate-type fixed bed reactor with an inner diameter of 40mm and a height of 1800mm, and is filled with a Cu-Zn-Al series methanol synthesis catalyst (consisting of CuO15wt%, ZnO65wt% and Al)2O320 wt%); the temperature of the reaction zone is 225-265 ℃, and the reaction pressure is 5 MPa;
commercial carbon monoxide (purity 98V%) and H2(purity 99.9V%) 2.7: 1, mixing the raw materials into synthesis gas, continuously exchanging heat through a methanol reactor outlet heat exchanger C and a methanol reactor inlet heat exchanger B to reach the inlet temperature of a catalyst bed layer of the methanol reactor A of 195 ℃, feeding the raw materials from the top of the methanol reactor A to perform catalytic synthesis methanol reaction, discharging a bottom reaction product, performing heat exchange on the discharged material and the raw materials of the synthesis gas through the methanol reactor outlet heat exchanger C, feeding the discharged material into a water cooler D, cooling the discharged material to 30-50 ℃, and feeding the cooled material into a gas-liquid separator E; gas phase at the top of the gas-liquid separator E (composition: Hydrogen 77.49V%, CO9V%, CO)21.74V%, 9.99V% of nitrogen and 1.78V% of other gases) and then pressurizing and mixing with the raw material of the synthesis gas for recycling after discharging a small amount of non-condensable gas; the composition of the condensate coming out of the bottom of the gas-liquid separator E was: 90.77wt% methanol, 8.78wt% water, CO20.4wt%, others 0.05 wt%.
And (3) decompressing the condensate at the bottom of the gas-liquid separator E by a decompression valve, introducing the condensate into a flash tank F (with the inner diameter of 32mm and the height of 1500 mm) for gas-liquid phase separation, discharging the gas-phase noncondensable gas to the outside for uniform discharge, heating the liquid phase (comprising 90.77wt% of methanol, 8.78wt% of water and 0.45wt% of other components) to the bubble point temperature of the methanol (below 0.8 MPa) by a preheater G, entering a heat preservation tank for storage, and feeding the liquid phase from a 43 th tower plate of a reaction rectifying tower H for methanol dehydration to prepare dimethyl ether.
The reaction rectifying tower H (the inner diameter is 40mm, the height is 4000 mm) is a pressurized rectifying tower, and 50 random saddle type solid super acidic fillers with the height of a theoretical plate are filled in the pressurized rectifying tower H; the solid super acidic filler isSelf-made filler: the first is to (NH4)2S2O8Dipping the mixture on a nano beta molecular sieve with the particle size of 60-80nm to prepare catalyst nano powder (with the particle size of 68-80 nm), and mechanically mixing the catalyst nano powder, the ceramic pug and a bonding auxiliary agent (hydroxypropyl methyl cellulose or sesbania powder) according to the mass ratio of 60: 35: 5, then kneading, molding and drying, and roasting at 550 ℃ for 4 hours to obtain the super acid active center composite material, wherein the size of the super acid active center composite material is 5mm multiplied by 5mm, and the average load of filler particles is 0.6wt percent.
In the reaction rectifying tower H, the liquid and the steam of the methanol are in countercurrent contact on the random saddle type solid super acidic filler, and the dimethyl ether preparation reaction is carried out on the filler by dehydration, wherein the temperature of a reaction zone is 120-220 ℃, the pressure is 1.5MPa, and the hourly space velocity of the methanol liquid is 4.5H-1The reflux ratio is 2.5; unreacted methanol is led out from the side line of the top of the tower positioned at the 2 nd tower plate, is mixed with the liquid phase from the flash tank F, and circulates to the reaction rectifying tower H for continuous reaction; the gas phase at the top of the tower is subjected to three-stage condensation (the first-stage condensation temperature is 40 ℃, the second-stage condensation temperature is normal temperature, the third stage is cooled to-26 ℃ by using glycol refrigerant), the non-condensable gas is subjected to water seal at the top of the tower and then is sent to an external recovery system, and the discharged material at the top of the tower (the composition of the dimethyl ether is 99.9wt percent, and the other is 0.1wt percent) enters a carbonylation reactor J after being discharged from the tower; the material discharged from the tower bottom of the reaction rectifying tower H (the composition is 71.8wt% of methanol, and 28.2% of other heavy components such as water generated in the reaction) is extracted and then enters a methanol recovery tower I (the normal pressure rectifying tower has the inner diameter of 32mm and the height of 3000 mm); and (3) sending impurities such as the tower kettle discharge water of the methanol recovery tower I to the outside for treatment, mixing the methanol extracted from the tower top of the methanol recovery tower I with the liquid phase from the flash tank F, and circulating to the reaction rectifying tower H for continuous reaction.
The carbonylation reactor J is a plate-type fixed bed reactor with the inner diameter of 40mm and the height of 1800mm, and is filled with a carbonylation catalyst (H-MOR 90wt%, active auxiliary agent components CuO9wt%, CoO0.5wt%, Mo)2O30.5 wt%), the temperature of the reaction zone is 230-300 ℃, and the reaction pressure is 5 MPa.
Pressurizing the discharged material at the top of the reactive distillation column H to 5MPa, and then leading the discharged material to come from a carbon monoxide main pipe(purity 98 v%) of CO and H of the Hydrogen header2(the purity is 99.9 v%) and continuously exchanges heat through an outlet heat exchanger L of the carbonylation reactor and an inlet heat exchanger K of the carbonylation reactor, and the mixture reaches the inlet temperature of a catalyst bed layer and enters a carbonylation reactor J for carbonylation reaction of dimethyl ether and CO; dimethyl ether, CO and H in the column2The feeding molar ratio is 1: 60: 15, the mass hourly space velocity of the liquid in the tower is 0.2 Kg/Kg.h; leading out a carbonylation reaction product from the bottom of the tower, carrying out heat exchange with a cold raw material through a carbonylation reactor outlet heat exchanger L, then, entering a water cooler M, cooling to 30-50 ℃, and then, entering a gas-liquid separator N; discharging 0.05% of non-condensable gas from a gas phase (with the composition of 76V% of carbon monoxide, 19.5V% of hydrogen, 2.5V% of nitrogen, 1.2V% of dimethyl ether, 0.5V% of methane and 0.3V% of others) from the top of a gas-liquid separator N, pressurizing, and mixing with a discharge material from the top of a reactive distillation column H to circulate V to a carbonylation reactor J for continuous reaction; the condensate discharged from the bottom of the gas-liquid separator N consists of: 78.2wt% of methyl acetate, 16.3wt% of dimethyl ether, 1.2wt% of methanol, 3.4wt% of heteropolyacid and 0.9wt% of the rest.
Introducing the condensate from the bottom of the gas-liquid separator N into a dimethyl ether recovery tower O (the inner diameter is 32mm, the height is 3000mm, the tower top temperature is about 65 ℃, the tower bottom temperature is 120 ℃, and the tower kettle pressure is 0.6 MPa); pressurizing light components such as dimethyl ether and the like evaporated from the top of the dimethyl ether recovery tower O, circulating the pressurized light components to the carbonylation reactor J for continuous reaction, and rectifying heavy components (comprising 94.6wt% of methyl acetate and 5.4wt% of heteropolyacid) at the bottom of the methyl acetate rectifying tower P (the inner diameter is 32mm, the height is 3000mm, the top temperature is 65 ℃, the bottom temperature is 140 ℃, and the pressure is normal pressure); the bottom of the methyl acetate rectifying tower P is led out polybasic heteropolyacid containing heavy components and is sent to the outside for recycling treatment, and the discharging composition at the top of the tower is as follows: 99.5wt% of methyl acetate and 0.5wt% of the rest.
Pressurizing the discharged material at the top of a methyl acetate rectifying tower P to 3MPa, mixing the material with hydrogen (the purity is 99.9%) from a hydrogen main pipe, continuously exchanging heat through a hydrogenation reactor outlet heat exchanger S and a hydrogenation reactor inlet heat exchanger R, feeding the material from the top of a hydrogenation reactor Q (a plate type fixed bed reactor, the inner diameter is 40mm, and the height is 1800 mm), and adding a hydrogenation catalyst (the composition is CuO32wt%, ZnO15wt%, Mo2O33wt%，SiO250 wt%) of the catalyst, and carrying out hydrogenation reaction at the reaction zone temperature of 190-250 ℃, the reaction pressure of 3MPa and the liquid hourly space velocity of 3Kg/Kg.
Leading out a hydrogenation reaction product from the bottom of a hydrogenation reactor Q, carrying out heat exchange with a cold raw material through a hydrogenation reactor outlet heat exchanger S, then entering a water cooler T, cooling to 30-50 ℃, and then entering a gas-liquid separator U; discharging 0.05% of non-condensable gas from a gas phase (volume composition: 99v% of hydrogen, 0.05v% of methane, 0.02v% of nitrogen, 0.02v% of carbon monoxide and 0.91v% of the rest) from the top of the gas-liquid separator U, pressurizing, mixing with a hydrogenation reaction raw gas, circulating to the hydrogenation reactor Q, and continuously reacting; the mass of the condensate coming out from the bottom of the gas-liquid separator U comprises: 59.3wt% of ethanol, 38.2wt% of methanol, 1.1wt% of methyl acetate, 0.9wt% of ethyl acetate and 0.5wt% of the rest.
Introducing the condensate from the bottom of the gas-liquid separator U into a methyl acetate recovery tower V (a normal pressure rectifying tower with the inner diameter of 32mm and the height of 3000mm, the tower top temperature of about 65 ℃ and the tower bottom temperature of 120 ℃); condensing the gas phase discharged from the top of the methyl acetate recovery tower V, circulating non-condensable gases such as dimethyl ether to the carbonylation reactor J for continuous utilization, and circulating the methyl acetate condensate to the hydrogenation reactor Q for continuous reaction after pressurization; the liquid phase at the bottom of the methyl acetate recovery tower V (the weight composition is 38.9wt% of methanol, 60.8wt% of ethanol and 0.3wt% of the rest) enters an ethanol product tower W (an atmospheric distillation tower, the inner diameter is 32mm, the height is 3000mm, the temperature at the top of the tower is 64 ℃, and the temperature at the bottom of the tower is 85 ℃) for rectification; the material discharged from the top of the ethanol product tower W (the composition is 98.5wt% of methanol and the other 1.5 wt%) is circulated to the reaction rectifying tower H for continuous reaction; the discharge from the bottom of the ethanol product tower W is a fuel ethanol product with the purity of more than or equal to 92.1 percent (V/V).
The above embodiments are merely illustrative of the principles and effects of the present invention, and are not intended to limit the present invention. Modifications and variations can be made to the above-described embodiments by those skilled in the art without departing from the spirit and scope of the present invention. Accordingly, it is intended that all equivalent modifications or changes which may be made by those skilled in the art without departing from the spirit and technical spirit of the present invention be covered by the claims of the present invention.
Claims (27)
1. A process system for indirectly producing ethanol by synthesis gas is characterized by comprising a methanol reactor, a flash tank, a reaction rectifying tower, a carbonylation reactor, a dimethyl ether recovery tower, a methyl acetate rectifying tower, a hydrogenation reactor, a methyl acetate recovery tower and an ethanol product tower; wherein: the bottom outlet of the methanol reactor is connected with the inlet of the flash tank through a pipeline; the liquid phase outlet of the flash tank is connected with the lower inlet of the reactive distillation column through a pipeline; the top outlet of the reaction rectifying tower is connected with the top inlet of the carbonylation reactor through a pipeline; an outlet at the bottom of the carbonylation reactor is connected with an inlet of the dimethyl ether recovery tower through a pipeline; the bottom outlet of the dimethyl ether recovery tower is connected with the inlet of the methyl acetate rectifying tower through a pipeline; the top outlet of the methyl acetate rectifying tower is connected with the top inlet of the hydrogenation reactor through a pipeline; the bottom outlet of the hydrogenation reactor is connected with the inlet of the methyl acetate recovery tower through a pipeline, and the bottom outlet of the methyl acetate recovery tower is connected with the inlet of the ethanol product tower through a pipeline; and an outlet at the bottom of the ethanol product tower is a product ethanol discharge hole.
2. The process system of claim 1, wherein the methanol reactor is a fixed bed reactor or a fluidized bed reactor.
3. The process system of claim 1, wherein a methanol reactor outlet heat exchanger and a methanol reactor inlet heat exchanger are disposed between the bottom outlet and the top inlet of the methanol reactor; the bottom outlet of the methanol reactor is connected with the hot material flow inlet of the outlet heat exchanger of the methanol reactor through a pipeline; a cold flow outlet of the methanol reactor outlet heat exchanger is connected with a cold flow inlet of the methanol reactor inlet heat exchanger through a pipeline; a cold flow outlet of the inlet heat exchanger of the methanol reactor is connected with a top inlet of the methanol reactor through a pipeline; a cold material flow inlet of the methanol reactor outlet heat exchanger is connected with a synthesis gas raw material pipeline through a pipeline; the synthesis gas raw material pipeline is connected with the hydrogen main pipe and the carbon monoxide main pipe through pipelines.
4. The process system of claim 3, wherein the hot stream outlet of the methanol reactor outlet heat exchanger is connected to a first gas-liquid separator via a line; the bottom liquid phase outlet of the first gas-liquid separator is connected with the inlet of the flash tank through a pipeline; and a top gas-phase outlet of the first gas-liquid separator is connected with a cold material flow inlet of the methanol reactor outlet heat exchanger through a pipeline.
5. The process system of claim 4, wherein a first compressor is connected between the top gas phase outlet of the first gas-liquid separator and the cold stream inlet of the methanol reactor outlet heat exchanger via a line; and a top gas-phase outlet of the first gas-liquid separator is also connected with a first non-condensable gas discharge pipeline.
6. The process system of claim 4, wherein a first water cooler is disposed between the hot stream outlet of the methanol reactor outlet heat exchanger and the first gas-liquid separator.
7. The process system according to claim 4, wherein a pressure reducing valve is connected between the bottom liquid phase outlet of the first gas-liquid separator and the inlet of the flash drum via a line.
8. The process system as claimed in claim 1, wherein the top outlet of the reactive distillation column is connected with a condensing reflux device through a pipeline; the outlet of the condensation reflux device is connected with the top inlet of the carbonylation reactor through a pipeline; the condensation reflux device is also provided with a non-condensable gas discharge pipeline.
9. The process system of claim 1, wherein a preheater is connected between the liquid phase outlet of the flash drum and the lower inlet of the reactive distillation column via a line.
10. The process system of claim 9, wherein a holding tank is connected between the cold stream outlet of the preheater and the lower inlet of the reactive distillation column via a line.
11. The process system of claim 9, wherein a material outlet is arranged on the side part of the top of the reactive distillation column, and the material outlet is connected with a connecting pipeline between the preheater and the lower inlet of the reactive distillation column through a pipeline.
12. The process system as claimed in claim 9, wherein the bottom outlet of the reactive distillation column is connected with a methanol recovery column through a pipeline; and the outlet at the top of the methanol recovery tower is connected with a connecting pipeline between the flash tank and the preheater through a pipeline.
13. The process system of claim 1, wherein the carbonylation reactor is a fixed bed reactor or a fluidized bed reactor.
14. The process system of claim 1, wherein the hydrogenation reactor is a fixed bed reactor or a fluidized bed reactor.
15. The process system of claim 1, wherein a carbonylation reactor outlet heat exchanger and a carbonylation reactor inlet heat exchanger are disposed between the bottom outlet and the top inlet of the carbonylation reactor; the bottom outlet of the carbonylation reactor is connected with the hot material flow inlet of the outlet heat exchanger of the carbonylation reactor through a pipeline; the cold flow outlet of the outlet heat exchanger of the carbonylation reactor is connected with the cold flow inlet of the inlet heat exchanger of the carbonylation reactor through a pipeline; the cold flow outlet of the carbonylation reactor inlet heat exchanger is connected with the top inlet of the carbonylation reactor through a pipeline; and a cold material flow inlet of the outlet heat exchanger of the carbonylation reactor is connected with an outlet at the top of the reactive distillation column, a hydrogen main pipe and a carbon dioxide main pipe through pipelines.
16. The process system of claim 15, wherein the hot stream outlet of the carbonylation reactor outlet heat exchanger is piped to a second vapor-liquid separator; a bottom liquid phase outlet of the second gas-liquid separator is connected with an inlet of the dimethyl ether recovery tower through a pipeline; and the top gas-phase outlet of the second gas-liquid separator is connected with the cold material flow inlet of the outlet heat exchanger of the carbonylation reactor through a pipeline.
17. The process system of claim 16, wherein a second compressor is connected between the top gas phase outlet of the second gas-liquid separator and the cold stream inlet of the carbonylation reactor outlet heat exchanger via a line.
18. The process system of claim 16, wherein a second non-condensable gas discharge line is further connected to the top gas-phase outlet of the second gas-liquid separator.
19. The process system of claim 16, wherein a second water cooler is disposed between the hot stream outlet of the carbonylation reactor outlet heat exchanger and the second vapor-liquid separator.
20. The process system of claim 15, wherein the top outlet of the dimethyl ether recovery column is connected via a line to the connecting line between the top outlet of the reactive rectification column and the cold stream inlet of the carbonylation reactor outlet heat exchanger.
21. The process system of claim 15, wherein a hydrogenation reactor outlet heat exchanger and a hydrogenation reactor inlet heat exchanger are disposed between the bottom outlet and the top inlet of the hydrogenation reactor; the bottom outlet of the hydrogenation reactor is connected with the hot material flow inlet of the outlet heat exchanger of the hydrogenation reactor through a pipeline; a cold flow outlet of the outlet heat exchanger of the hydrogenation reactor is connected with a cold flow inlet of the inlet heat exchanger of the hydrogenation reactor through a pipeline; a cold material flow outlet of the hydrogenation reactor inlet heat exchanger is connected with a top inlet of the hydrogenation reactor through a pipeline; and a cold material flow inlet of the outlet heat exchanger of the hydrogenation reactor is connected with a tower top outlet of the methyl acetate rectifying tower and a hydrogen main pipe through pipelines.
22. The process system of claim 21, wherein the hot stream outlet of the hydrogenation reactor outlet heat exchanger is piped to a third gas-liquid separator; a bottom liquid phase outlet of the third gas-liquid separator is connected with an inlet of the methyl acetate recovery tower through a pipeline; and a top gas-phase outlet of the third gas-liquid separator is connected with a cold material flow inlet of the outlet heat exchanger of the hydrogenation reactor through a pipeline.
23. The process system of claim 22, wherein a third compressor is connected via a line between the top gas phase outlet of the third gas-liquid separator and the cold stream inlet of the hydrogenation reactor outlet heat exchanger.
24. The process system of claim 22, wherein a third non-condensable gas discharge line is further connected to the top gas-phase outlet of the third gas-liquid separator.
25. The process system of claim 22, wherein a third water cooler is disposed between the hot stream outlet of the hydrogenation reactor outlet heat exchanger and the third gas-liquid separator.
26. The process system as claimed in claim 21, wherein the top outlet of the methyl acetate recovery tower is connected with one or more stages of condensing units through a pipeline; the liquid phase outlet of the condensing device is connected with the cold material flow inlet of the outlet heat exchanger of the hydrogenation reactor through a pipeline; the gas phase outlet of the condensing device is connected with the cold material flow inlet of the outlet heat exchanger of the carbonylation reactor through a pipeline.
27. The process system of any one of claims 1 to 26, wherein the flash drum is replaced by two rectification columns or three rectification columns connected in series by a pipeline.
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| CN2012207125324U CN203033918U (en) | 2012-12-20 | 2012-12-20 | Process system for indirectly producing ethanol by using synthesis gases |
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| CN109503326A (en) * | 2018-11-30 | 2019-03-22 | 西南化工研究设计院有限公司 | A kind of technique that dimethyl ether produces ethyl alcohol indirectly |
| CN111777491A (en) * | 2020-07-16 | 2020-10-16 | 南京延长反应技术研究院有限公司 | Micro-interface reaction system and method for preparing ethanol from coal |
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| CN109503326A (en) * | 2018-11-30 | 2019-03-22 | 西南化工研究设计院有限公司 | A kind of technique that dimethyl ether produces ethyl alcohol indirectly |
| CN109503326B (en) * | 2018-11-30 | 2020-10-09 | 西南化工研究设计院有限公司 | Process for indirectly producing ethanol by dimethyl ether |
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| CN114230450A (en) * | 2021-12-15 | 2022-03-25 | 江苏湖大化工科技有限公司 | Methyl propionate synthesis process device utilizing coupling hydrogenation type reaction rectifying tower |
| CN114230450B (en) * | 2021-12-15 | 2023-12-12 | 江苏湖大化工科技有限公司 | Methyl propionate synthesis process device utilizing coupling hydrogenation reaction rectifying tower |
| CN114605235A (en) * | 2022-04-15 | 2022-06-10 | 南京工业大学 | CO (carbon monoxide)2Method for preparing dimethyl ether by hydrogenation |
| CN116082120A (en) * | 2022-11-25 | 2023-05-09 | 中国科学院大连化学物理研究所 | Technological method and device for preparing ethanol by continuous reaction of methanol |
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