CN111253213A - Process method and system for preparing ethanol by acetate hydrogenation - Google Patents

Process method and system for preparing ethanol by acetate hydrogenation Download PDF

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CN111253213A
CN111253213A CN201811456961.8A CN201811456961A CN111253213A CN 111253213 A CN111253213 A CN 111253213A CN 201811456961 A CN201811456961 A CN 201811456961A CN 111253213 A CN111253213 A CN 111253213A
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gas
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liquid
tower
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CN111253213B (en
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田靖
赵娜
丁干红
吕建宁
王宏涛
尹佳子
张佳楠
张磊
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Tianjin University
Yangquan Coal Industry Group Co Ltd
Wison Engineering Ltd
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Tianjin University
Yangquan Coal Industry Group Co Ltd
Wison Engineering Ltd
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    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C29/00—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring
    • C07C29/132—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
    • C07C29/76—Separation; Purification; Use of additives, e.g. for stabilisation by physical treatment
    • C07C29/80—Separation; Purification; Use of additives, e.g. for stabilisation by physical treatment by distillation
    • 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
    • C07C29/88—Separation; Purification; Use of additives, e.g. for stabilisation by treatment giving rise to a chemical modification of at least one compound
    • C07C29/90—Separation; Purification; Use of additives, e.g. for stabilisation by treatment giving rise to a chemical modification of at least one compound using hydrogen only

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  • Chemical Kinetics & Catalysis (AREA)
  • General Chemical & Material Sciences (AREA)
  • Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)

Abstract

The invention relates to a process method and a system for preparing ethanol by acetate hydrogenation, wherein the process method comprises a gas phase hydrogenation reaction, a product separation and a liquid phase hydrogenation reaction, the system comprises a gas phase reaction device, a liquid phase reaction device and a product separation device, the system is coupled with the product separation device through the liquid phase reaction device, and a small amount of unconverted acetate and other reaction byproducts containing carbonyl in a gas phase reaction liquid phase product are subjected to liquid phase hydrogenation conversion removal by using hydrogen of the gas phase reaction device, so that the circulation amount of the gas phase reaction device is reduced, the product yield is increased, the total purge gas discharge amount is reduced, the product separation difficulty is reduced, and the system energy consumption is saved. Compared with the prior art, the method has the advantages of simple flow, low energy consumption of unit product, high product yield, low investment, convenient operation and easy industrial large-scale production.

Description

Process method and system for preparing ethanol by acetate hydrogenation
Technical Field
The invention relates to an ethanol production and separation process, in particular to a process method and a system for preparing ethanol by acetate hydrogenation, which adopt a liquid phase hydrogenation method to optimize reaction and separation processes, and belong to the technical field of chemical industry.
Background
Ethanol (CH)3CH2OH) is an important basic chemical raw material and is widely applied to various fields of food, chemical industry, military industry, medicine and the like. The ethanol is clean high-octane fuel, has the characteristics of high octane number, good antiknock property, small air pollution caused by products after combustion and the like, is a world-recognized environment-friendly clean fuel and an oil quality improver, can be used as vehicle fuel to be added into gasoline, and is a novel clean fuel. Fuel ethanol is an important direction for the development and utilization of renewable energy, the usage ratio thereof is gradually increasing, ethanol gasoline has been used for over 30 years in the united states and brazil, and the use of ethanol gasoline is also actively promoted in europe and south-east asia.
The gap of the ethanol gasoline is huge, according to the 'implementation scheme about expanding the production of the biofuel ethanol and popularizing and using the ethanol gasoline for the vehicle' promoted by the country in 2017, the ethanol gasoline for the vehicle is to be popularized and used nationwide by 2020. The fuel ethanol at the present stage is mainly obtained by fermenting aged grains, but the current production capacity can not achieve the domestic purpose. If only aging grain to produce ethanol is emphasized, there is a limit to the yield of ethanol. At present, a large amount of methyl acetate is produced as a byproduct in the production of polyvinyl alcohol (PVA), and the industrial application of the methyl acetate is limited; in addition, the dimethyl ether production capacity in China is greatly surplus, and dimethyl ether can also be carbonylated with carbon monoxide to generate methyl acetate, so that the ethanol preparation by acetate hydrogenation is developed, the shortage of the ethanol preparation yield from grains can be supplemented, and the problems of the existing byproduct for PVA production, the release of the surplus dimethyl ether production capacity and the like can be solved.
Patent CN 102976892B discloses a method for preparing ethanol by acetate hydrogenation, which is characterized in that in a fixed bed reactor filled with a copper-based catalyst, the reaction temperature is 220 ℃, the reaction pressure is 3MPa, the molar hydrogen-ester ratio is 30, and the mass space velocity of acetate is 2h-1The conversion rate of acetate is 98.5%, the selectivity of ethanol is up to 99.6%, and the catalyst has extremely high hydrogenation activity, selectivity and stability. But a small amount of unconverted acetic ester or by-product acetaldehyde and other carbonyl-containing compounds still exist in the reaction product, so that the separation difficulty of subsequent products is increased.
Patent CN 103265402B discloses a method for reducing energy consumption in a process of preparing ethanol by acetate hydrogenation, which adopts a binary azeotropic mixture generated by unreacted acetate and alcohol to directly return to a hydrogenation section for catalytic hydrogenation treatment instead of subsequent special rectification separation, thus reducing energy consumption for subsequent special separation. But the circulation quantity of an ester hydrogenation reaction loop is increased, and the purity and the recovery rate of an ethanol product are difficult to guarantee.
Patent CN 105439816B discloses a process for producing ethanol by acetate hydrogenation, which mainly utilizes secondary condensation, pressure difference in the existing system and a hydrogen recovery device to purify and recycle part of hydrogen in purge gas, thereby reducing reactant loss and hydrogen consumption. However, the hydrogen recovery device is complex, and the system cannot solve the problem that acetic ester cannot be completely converted, so that the subsequent product separation is influenced.
The patent CN 105367385B discloses a separation method for preparing ethanol and coproducing methanol by methyl acetate hydrogenation, which considers light hydrocarbon, acetaldehyde and other impurities possibly existing in methyl acetate hydrogenation reaction products, adopts a technical scheme that a three-tower sequential separation process is coupled with a methanol refining and dealdehyding process, and greatly reduces separation energy consumption while meeting product quality. However, the amount of the produced acetaldehyde, methyl acetate and other substances at the top of the first distillation tower is small, and acetaldehyde, methyl acetate and other substances cannot be completely produced and returned, so that acetaldehyde, methyl acetate and the like still exist in a downstream tower system, further dealdehyding treatment is needed to ensure the quality of a methanol product, the methanol product is obtained at the top of the second distillation tower and the fourth distillation tower, the repeated evaporation and condensation of the methanol at the tops of the two towers inevitably increases the separation energy consumption, and the production of an acetal reaction product at the bottom of the fourth distillation tower can cause the loss of the methanol product. In addition, the stream S2 extracted from the top of the first distillation column and returned to the reaction system will increase the load of the reaction system.
In conclusion, in the acetate hydrogenation reaction, the incompletely converted reactant acetate and the product ethanol can form a binary azeotropic mixture, and the binary azeotrope is difficult to separate; the adoption of the method for returning the binary azeotrope to the reactor to continue the catalytic hydrogenation increases the energy consumption, and the product purity of the ethanol is difficult to ensure; the invention provides a method for continuously hydrogenating and converting unconverted acetic ester in a liquid-phase product into ethanol by utilizing a liquid-phase hydrogenation reaction, which avoids the circulation of the acetic ester, reduces the amount of circulating hydrogen in a gas-phase reaction device and reduces the energy consumption of a system.
Disclosure of Invention
The invention aims to overcome the defects of the prior art and provide a process method and a device for preparing ethanol by acetate hydrogenation.
The purpose of the invention can be realized by the following technical scheme:
a process method for preparing ethanol by acetate hydrogenation comprises the following steps:
1) gas phase hydrogenation
Mixing an acetate raw material and hydrogen, gasifying and preheating to form a steam feed, and adding the steam feed into a gas phase hydrogenation reactor for hydrogenation reaction; condensing a reaction product output by the gas-phase hydrogenation reactor, and performing gas-liquid separation in a gas-liquid separator to obtain a gas-phase product A and a liquid-phase product A, wherein part of the gas-phase product A is recycled to the gas-phase hydrogenation reactor for use;
2) product separation
Sending the liquid-phase product A in the gas-phase hydrogenation reaction into a product separation device for separation to obtain an absolute ethyl alcohol product, a liquid-phase product B containing acetic ester and carbonyl byproducts and heavy component materials in a tower kettle, and sending the absolute ethyl alcohol product and the heavy component materials in the tower kettle out of a battery limit;
3) liquid phase hydrogenation reaction
Preheating a liquid-phase product B obtained by separating hydrogen and a product, adding the liquid-phase product B into a liquid-phase hydrogenation reactor for hydrogenation reaction, condensing a reaction product output by the liquid-phase hydrogenation reactor, carrying out gas-liquid separation in a gas-liquid separator to obtain a gas-phase product C and a liquid-phase product C, and separating the liquid-phase product C in a product separation device.
The gas-phase hydrogenation reaction is specifically that an acetate raw material is added into a vaporizer, fresh hydrogen and circulating hydrogen are mixed to form a hydrogen raw material, the hydrogen raw material is preheated by a feeding and discharging heat exchanger a and then is introduced into the vaporizer, a mixture of hydrogen and acetate is obtained at an outlet of the vaporizer, the mixture of hydrogen and acetate is heated by a heat exchanger and then enters a gas-phase hydrogenation reactor to undergo a gas-phase hydrogenation reaction to obtain a gas-phase hydrogenation reaction product, the gas-phase hydrogenation reaction product is cooled by the feeding and discharging heat exchanger a and a cooler a and then enters a gas-liquid separator a to obtain a gas-phase product A and a liquid-phase product A, and the gas-phase product A is fed into a compressor and discharged from a gas-; the feed of the compressor is pressurized by the compressor and then is used as a circulating hydrogen raw material to continuously participate in the gas-phase hydrogenation reaction; and the purge gas of the gas phase reaction device is subjected to liquid phase hydrogenation reaction.
The composition of the acetate raw material comprises methyl acetate, ethyl acetate or a mixture of the methyl acetate and the ethyl acetate; the feed acetate to the vaporizer is a high pressure liquid;
the acetic ester raw material and the hydrogen raw material are contacted, mixed and gasified in a vaporizer, and the gasification pressure is 1.0-6.0 MPaG; the gasification heat source of the acetic ester in the vaporizer comes from the sensible heat of the circulating material flow at the outlet of the feeding and discharging heat exchanger a, and no external independent heat supply is needed.
The temperature of the material at the outlet of the superheater is 200-300 ℃; the superheated acetic ester gas does not contain liquid; the heating source adopts high-pressure saturated steam or superheated steam or an electric heater for heat supply;
the catalyst in the gas phase hydrogenation reactor is a copper catalyst, the temperature in the reactor is 200-300 ℃, the pressure in the reactor is 1.0-6.0 MPaG, the molar ratio of hydrogen to acetic ester at the inlet of the gas phase hydrogenation reactor is 2-50, the gas phase hydrogenation reactor is an isothermal tubular reactor, the tube side is filled with the catalyst, the shell side is filled with boiler water serving as a heat-taking medium, and heat generated by reaction is removed through water gasification of the shell side of the reactor;
the cooling medium of the cooler a is circulating water or chilled water; the material temperature of the gas-phase hydrogenation reaction product at the outlet of the cooler a is 20-50 ℃;
the fresh hydrogen is mixed with the circulating hydrogen at the inlet or the outlet of the compressor; the purity of the fresh hydrogen is more than or equal to 99.9mole percent, and the rest is inert components; the fresh hydrogen can enter the inlet or the outlet of the compressor according to the actual boundary area pressure and the pressure requirement of the reaction system; the compressor is a gas phase circulating compressor, the feeding of the compressor is mainly gas phase material flow at the outlet of the gas-liquid separator a, and the outlet pressure is 1.0-6.0 MpaG.
For energy saving, the feed-discharge heat exchanger a utilizes the outlet reactant flow of the reactor to heat the outlet recycle flow of the compressor, thereby reducing the temperature of the reactor.
The product separation specifically comprises the steps that a liquid-phase product A enters a No. 1 rectifying tower for rectification separation, a gas phase at the top of the tower is condensed by a condenser to obtain a liquid-phase product B and a gas-phase product B, the gas-phase product is extracted from purge gas, and the liquid-phase product B enters a liquid-phase reaction device;
and the liquid phase material in the tower bottom of the No. 1 rectifying tower enters a product rectifying tower to be separated to obtain an absolute ethyl alcohol product.
A separation method of a product rectifying tower is arranged according to the composition of the acetic ester raw material;
when the acetic ester raw material component is ethyl acetate, the product rectifying tower comprises a 3# rectifying tower, and an absolute ethyl alcohol product is extracted from the top of the 3# rectifying tower;
when the acetic ester raw material component is methyl acetate or a mixture of ethyl acetate and methyl acetate, the product rectifying tower comprises a 2# rectifying tower and a 3# rectifying tower, a tower kettle liquid-phase material of the 1# rectifying tower enters the 2# rectifying tower, a methanol-containing product is extracted from the top of the 2# rectifying tower, the tower kettle liquid-phase material enters the 3# rectifying tower, an absolute ethyl alcohol product is extracted from the top of the 3# rectifying tower, and tower kettle liquid is sent out of a boundary region.
The tower top operating pressure of the No. 1 rectifying tower is 0-500 kPaG, the reflux ratio is 0.01-100, and the content of acetic ester in a tower bottom liquid phase is less than or equal to 5 ppm-wt; the feeding of the reactor is the liquid phase component at the outlet of the gas-liquid separator a and the liquid phase product of the gas-liquid separator b in the liquid phase reaction device, part of purge gas is discharged from the tower top, and the liquid phase extracted from the tower top is pumped and pressurized to be used as the feeding of the liquid phase reaction device; the liquid phase in the tower bottom is used as the feed of the No. 2 rectifying tower.
The top operating pressure of the 2# rectifying tower is 0-500 kPaG, the reflux ratio is 0.01-100, the purity of the methanol product is more than or equal to 99.8 wt%, and the tower kettle liquid phase is used as the feeding material of the 3# rectifying tower;
the operation pressure of the top of the 3# rectifying tower is 0-500 kPaG, the reflux ratio is 0.01-100, an absolute ethyl alcohol product is extracted from a liquid phase at the top of the tower, the purity is more than or equal to 99.7 wt.%, and a liquid phase at the bottom of the tower is taken as a heavy component and is sent out of the system.
The liquid-phase hydrogenation reaction is specifically that hydrogen and a liquid-phase product B are mixed to form a liquid-phase hydrogenation reaction raw material, the liquid-phase hydrogenation reaction raw material is preheated by a feeding and discharging heat exchanger B and then enters a liquid-phase hydrogenation reactor to carry out liquid-phase hydrogenation reaction, and the obtained reaction product is cooled by the feeding and discharging heat exchanger B and a cooler B and then enters a gas-liquid separator B to form a liquid-phase product C and a gas-phase product C; the liquid-phase product C enters a product separation device for separation; the gas phase product C is withdrawn as purge gas.
The catalyst in the liquid phase hydrogenation reactor is a nickel catalyst, the reaction pressure is 1.0-5.0 MPaG, the reaction temperature is 80-250 ℃, and the molar ratio of hydrogen to acetic ester in the liquid phase hydrogenation raw material is 2-100;
the cooling medium of the cooler b is circulating water or chilled water; and the material temperature of the gas-phase hydrogenation reaction product at the outlet of the cooler b is 10-50 ℃.
The total purge gas is a mixture of gas obtained after a gas-phase product A in a gas-phase reaction device is treated by a liquid-phase reaction device and the overhead gas of the No. 1 rectifying tower;
the hydrogen in the liquid phase hydrogenation reaction is from a gas phase product A or fresh hydrogen in the gas phase hydrogenation reaction.
The gas-phase product part of the gas-liquid separator a is used as the purge gas of the gas-phase reaction device to enter the liquid-phase reaction device for reaction, and the hydrogen content of the purge gas of the gas-phase reaction device can be indirectly supplemented by fresh hydrogen through a gas-phase reaction circulation loop or directly supplemented by fresh hydrogen.
The invention provides a process system for preparing ethanol by acetate hydrogenation, which comprises a gas-phase reaction device, a liquid-phase reaction device and a product separation device;
wherein the gas phase reaction device is respectively connected with the liquid phase reaction device and the product separation device, and the liquid phase reaction device is connected with the product separation device. The system utilizes hydrogen of a gas phase reaction device to carry out liquid phase hydrogenation conversion removal on a small amount of unconverted acetate and other reaction byproducts containing carbonyl in a liquid phase product of the gas phase reaction device, and the liquid phase product after liquid phase hydrogenation is returned to a product separation device for separation;
acetic ester and hydrogen in the gas phase reaction device are subjected to gas phase hydrogenation reaction, and liquid phase products generated after condensation and gas-liquid separation of reaction products are sent to the product separation device;
the product separation device comprises a rough separation tower and a product rectifying tower, wherein a liquid-phase product containing acetic ester and carbonyl byproducts is obtained at the tower top of the rough separation tower, and an absolute ethyl alcohol product is extracted at the tower top of the product rectifying tower;
the liquid phase reaction device utilizes hydrogen or fresh hydrogen from the gas phase reaction device to carry out liquid phase hydrogenation reaction with the liquid phase product containing the acetic ester and the carbonyl byproducts to obtain a liquid phase hydrogenation reaction product, and the liquid phase product obtained after the condensation and separation of the liquid phase hydrogenation reaction product is sent to the product separation device.
Specifically, the gas phase reaction device comprises a vaporizer, a superheater, a gas phase hydrogenation reactor, a feeding and discharging heat exchanger a, a cooler a, a gas-liquid separator a and a compressor; the acetate raw material enters from a first raw material inlet of a vaporizer, the hydrogen raw material enters from a second raw material inlet of the vaporizer, and an outlet of the vaporizer and the superheater are sequentially connected with the reactor through pipelines; the feeding and discharging heat exchanger a heats hydrogen raw materials by using materials at the outlet of the reactor, the outlet of the reactor is connected with the hot material inlet of the feeding and discharging heat exchanger a, the hot material outlet of the feeding and discharging heat exchanger a is connected with the inlet of the cooler a, the outlet of the cooler a is connected with the inlet of the gas-liquid separator a, and the gas-phase outlet of the gas-liquid separator a is respectively connected with the inlet of the compressor and the liquid-phase reaction device; the liquid phase outlet of the gas-liquid separator a is connected with the product separation device; and the outlet of the compressor is connected with the cold material inlet of the feeding and discharging heat exchanger a, and the cold material outlet of the feeding and discharging heat exchanger a is connected with the second raw material inlet of the vaporizer.
The product separation device comprises a rough separation tower and a product separation tower, the rough separation tower comprises a No. 1 rectifying tower, a liquid-phase product outlet of a condenser at the top of the No. 1 rectifying tower is connected with the liquid-phase reaction device, and a liquid-phase outlet of a tower kettle of the No. 1 rectifying tower is connected with the product separation tower;
when the acetic ester raw material component does not contain methyl acetate, the product separation tower comprises a 3# rectifying tower, a tower kettle liquid phase outlet of the 1# rectifying tower is connected with an inlet of the 3# rectifying tower, and anhydrous ethanol is extracted from the tower top of the 3# rectifying tower;
when the acetic ester raw material component contains methyl acetate, the product separation tower comprises a 2# rectifying tower and a 3# rectifying tower, a tower kettle liquid phase outlet of the 1# rectifying tower is connected with an inlet of the 2# rectifying tower, a methanol product is collected from the tower top of the 2# rectifying tower, a tower kettle liquid phase outlet is connected with an inlet of the 3# rectifying tower, and an ethanol product is collected from the tower top of the 3# rectifying tower;
the liquid phase reaction device comprises a liquid phase hydrogenation reactor, a feeding and discharging heat exchanger b, a cooler b and a gas-liquid separator b; and a cold material inlet of the feeding and discharging heat exchanger b is respectively connected with the gas-phase reaction device and the product separation device, a cold material outlet of the feeding and discharging heat exchanger b is connected with an inlet of the liquid-phase hydrogenation reactor, an outlet of the liquid-phase hydrogenation reactor is connected with a hot material inlet of the feeding and discharging heat exchanger b, and a hot material outlet of the feeding and discharging heat exchanger b is sequentially connected with the cooler b and the gas-liquid separator b.
And a cold material inlet of the feeding and discharging heat exchanger b is respectively connected with a gas-phase outlet of a gas-liquid separator a in the gas-phase reaction device and a liquid-phase product outlet of a 1# rectifying tower top condenser in the product separation device.
The system is coupled with the product separation device through the liquid phase reaction device, and a small amount of unconverted acetic ester and other carbonyl-containing reaction byproducts in the gas phase reaction liquid phase product are subjected to liquid phase hydrogenation conversion removal by utilizing the hydrogen of the gas phase reaction device, so that the circulation quantity of the gas phase reaction device is reduced, the product yield is increased, the total purge gas discharge quantity is reduced, the product separation difficulty is reduced, and the system energy consumption is saved.
Compared with the prior art, the beneficial effects of the invention are embodied in the following aspects:
(1) according to the invention, through liquid-phase hydrogenation, the content of aldehydes and esters compounds in a gas-phase hydrogenation reaction product is reduced or eliminated, the utilization rate of raw materials is improved, the separation difficulty of the product is reduced, and the product quality is easier to guarantee;
(2) according to the invention, by adopting a mode of combining product separation and liquid phase hydrogenation, the conversion rate and the product yield of the raw material acetic ester are improved, and the gas phase circulation amount of a gas phase reaction device circulation loop is reduced, so that the system energy consumption is reduced;
(3) the invention adopts a mode of direct contact gasification of reactor feeding, does not need to additionally supply a vaporization heat source, and has high heat efficiency;
(4) the method utilizes the purge gas of the circulation loop of the gas phase reaction device as the hydrogen source of the liquid phase hydrogenation reaction, does not have a special hydrogen purification device, and has simple flow and low investment.
In a word, the invention has the advantages of simple flow, low energy consumption of unit product, high product yield, low investment, convenient operation, easy industrial upsizing and the like.
Drawings
FIG. 1 is a process flow diagram of the present invention in examples 1 and 2;
FIG. 2 is a process flow diagram in comparative example 1 and comparative example 2;
FIG. 3 is a process flow diagram of the present invention in example 3;
FIG. 4 is a process flow diagram in comparative example 3;
in the figure, 1 is a vaporizer, 2 is a superheater, 3 is a gas phase hydrogenation reactor, 4 is a feed and discharge heat exchanger a, 5 is a cooler a, 6 is a gas-liquid separator a, 7 is a compressor, 8 is a liquid phase hydrogenation reactor, 9 is a rectification column # 1, 10 is a rectification column # 2, 11 is a rectification column # 3, 12 is a feed and discharge heat exchanger b, 13 is a cooler b, and 14 is a gas-liquid separator b; s1 is an acetic ester raw material, S2 is fresh hydrogen, S3 is superheater feeding, S4 is compressor feeding, S5 is cooler feeding, S6 is purge gas of a gas phase reaction device, S7 is 1# rectifying tower top purge gas, O1 is total purge gas, O2 is a methanol product, O3 is an anhydrous ethanol product, and O4 is a heavy component.
Detailed Description
The present invention will be described in detail with reference to specific examples. The following examples will assist those skilled in the art in further understanding the invention, but are not intended to limit the invention in any way. It should be noted that variations and modifications can be made by persons skilled in the art without departing from the spirit of the invention. All falling within the scope of the present invention.
Example 1
A process system for preparing ethanol by acetate hydrogenation is shown in figure 1, and is suitable for a reaction system in which an acetate raw material is methyl acetate, and specifically comprises a gas phase reaction device, a liquid phase reaction device and a product separation device; the specific equipment comprises a vaporizer 1, a superheater 2, a gas-phase hydrogenation reactor 3, a feed and discharge heat exchanger a4, a cooler a5, a gas-liquid separator a6, a compressor 7, a liquid-phase hydrogenation reactor 8, a 1# rectifying tower 9, a 2# rectifying tower 10, a 3# rectifying tower 11, a feed and discharge heat exchanger b12, a cooler b13 and a gas-liquid separator b 14.
The specific connection relationship among the devices is as follows:
the gas phase reaction device comprises a vaporizer 1, a superheater 2, a gas phase hydrogenation reactor 3, a feed and discharge heat exchanger a4, a cooler a5, a gas-liquid separator a6 and a compressor 7; the acetate S1 enters from a first raw material inlet of a vaporizer, the hydrogen raw material enters from a second raw material inlet of the vaporizer, and an outlet of the vaporizer, the superheater 2 and the reactor 3 are sequentially connected through pipelines; the feed and discharge heat exchanger a4 heats hydrogen raw material by using material at the outlet of the reactor 3, the outlet of the reactor 3 is connected with the hot material inlet of the feed and discharge heat exchanger a4, the hot material outlet of the feed and discharge heat exchanger a4 is connected with the inlet of the cooler a5, the outlet of the cooler a5 is connected with the inlet of the gas-liquid separator a6, and the gas-phase outlet of the gas-liquid separator a6 is connected with the inlet of the compressor 7 and the liquid-phase reaction device; the liquid phase outlet of the gas-liquid separator a6 is connected with the product separation device; the outlet of the compressor 7 is connected with the cold material inlet of the feed/discharge heat exchanger a4, and the cold material outlet of the feed/discharge heat exchanger a4 is connected with the second raw material inlet of the vaporizer 1.
The product separation device comprises a rough separation tower and a product separation tower, the rough separation tower comprises a No. 1 rectifying tower 9, a liquid phase material outlet of a condenser at the top of the No. 1 rectifying tower 9 is connected with the liquid phase reaction device, and a liquid phase outlet of a tower kettle 9 of the No. 1 rectifying tower is connected with the product separation tower;
the product separation tower comprises a 2# rectifying tower 10 and a 3# rectifying tower 11, a tower kettle liquid phase outlet of the 1# rectifying tower 9 is connected with an inlet of the 2# rectifying tower 10, a methanol product O2 is extracted from the tower top of the 2# rectifying tower 10, a tower kettle liquid phase outlet is connected with an inlet of the 3# rectifying tower 11, and an ethanol product O3 is extracted from the tower top of the 3# rectifying tower 11;
the liquid phase reaction device comprises a liquid phase hydrogenation reactor 8, a feeding and discharging heat exchanger b12, a cooler b13 and a gas-liquid separator b 14; the cold material inlet of the feeding and discharging heat exchanger b12 is respectively connected with the gas-phase outlet of the gas-liquid separator a6 and the liquid-phase product outlet of the condenser at the top of the 1# rectifying tower 9, the cold material outlet of the feeding and discharging heat exchanger b12 is connected with the inlet of the liquid-phase hydrogenation reactor 8, the outlet of the liquid-phase hydrogenation reactor 8 is connected with the hot material inlet of the feeding and discharging heat exchanger b12, and the hot material outlet of the feeding and discharging heat exchanger b12 is sequentially connected with the cooler b13 and the gas-liquid separator b 14.
The process method for preparing ethanol by acetate hydrogenation in the embodiment comprises the following steps:
(i) the liquid phase acetic ester raw material S1 is methyl acetate, is pressurized to 5.0MPaG by a pump, and is directly contacted with hot flow from a feeding and discharging heat exchanger a4 in a vaporizer 1 for mixing and gasification, and the temperature after gasification is 144 ℃;
(ii) (ii) heating the outlet gas phase of the vaporizer 1 in the step (i) to 235 ℃ through a superheater 2, then feeding the outlet gas phase into a gas phase hydrogenation reactor 3, adopting a copper-based catalyst, wherein the reaction pressure is 4.885MPaG, the reaction temperature is 240 ℃, the molar ratio of hydrogen to acetic ester is 20, the once-through conversion rate of methyl acetate is 97%, and the selectivity of ethanol is 96%; the gas phase hydrogenation reactor 3 is an isothermal tubular catalytic reactor, an acetate hydrogenation reaction occurs in a tube pass catalyst, and heat generated by the reaction is removed through partial gasification of boiler water on the shell pass of the gas phase hydrogenation reactor 3;
(iii) (iii) cooling the outlet gas phase of the gas-phase hydrogenation reactor 3 in the step (ii) at about 240 ℃ by a feed-discharge heat exchanger a4 and a cooler a5 in sequence, wherein the cooled temperature is 111 ℃ and 45 ℃ in sequence, performing gas-liquid phase separation on reaction products in a gas-liquid separator a6, feeding liquid phase products into a No. 1 rectifying tower 9 for pre-separation, pressurizing most of the feed S4 of a gas-phase product by a compressor 7, mixing with fresh hydrogen S2, heating by a feed-discharge heat exchanger a4, and feeding the mixture into a vaporizer 1; the purge gas S6 of the rest gas phase reaction devices is used as the feed of the liquid phase hydrogenation reactor 8; the number of theoretical plates of the No. 1 rectifying tower 9 is 70, the operating pressure at the top of the tower is 69kPaG, and the reflux ratio is 3;
(iv) the liquid phase at the top of the rectifying tower 1# 9 in the step (iii) is pressurized by a pump and then mixed with part of purge gas of a gas-liquid separator a6, namely the purge gas S6 of the gas-phase reaction device, then the mixture is heated by a feeding and discharging heat exchanger b12 and then enters a liquid-phase hydrogenation reactor 8 for reaction, a nickel-based catalyst is adopted, the reaction pressure is 1.0MPaG, the reaction temperature is 80 ℃, the molar ratio of hydrogen to acetate is 3, the product after the reaction is sequentially cooled to 45 ℃ through a feeding and discharging heat exchanger b12 and a cooler b13, then the product is separated in a gas-liquid separator b14, the separated liquid-phase product is returned to the rectifying tower 1# 9, and the separated gas is mixed with the purge gas S7 at the top of the rectifying tower 1# 9 and then discharged as total purge gas O1;
(v) and (3) feeding the tower bottom liquid of the 1# rectifying tower 9 in the step (iv) into a 2# rectifying tower 10 for separation, extracting a methanol product O2 from the tower top liquid phase of the 2# rectifying tower 10, feeding the tower bottom liquid phase of the 2# rectifying tower 10 into a 3# rectifying tower 11 for separation, extracting an absolute ethanol product O3 from the tower top liquid phase of the 3# rectifying tower 11, and extracting a heavy component O4 from the tower bottom liquid phase of the 3# rectifying tower 11 to be sent out of a boundary region. The number of theoretical plates of the 2# rectifying tower 10 is 80, the operation pressure at the top of the tower is 29kPaG, and the reflux ratio is 3.2; the number of theoretical plates of the 3# rectifying tower 11 is 30, the operating pressure at the top of the tower is 29kPaG, and the reflux ratio is 1.
Finally, the purity of the methanol product is more than or equal to 99.8 wt.% and the purity of the absolute ethyl alcohol product is more than or equal to 99.7 wt.% through the embodiment. Table 1-1 shows the calculation results of example 1:
table 1-1 material balance for example 1
Logistics S1 S2 O1 O2 O3 O4
Phase state Liquid phase Gas phase Gas phase Liquid phase Liquid phase Liquid phase
Composition of wt％ wt％ wt％ wt％ wt％ wt％
Hydrogen gas 0.0000 99.0564 15.9898 0.0000 0.0000 0.0000
Nitrogen gas 0.0000 0.2755 1.7640 0.0000 0.0000 0.0000
Argon gas 0.0000 0.5892 3.7766 0.0000 0.0000 0.0000
Methane 0.0000 0.0789 0.5065 0.0000 0.0000 0.0000
Ethane (III) 0.0000 0.0000 9.5264 0.0000 0.0000 0.0000
Ether compounds 0.0000 0.0000 31.6982 0.0000 0.0000 0.0000
Acetaldehyde 0.0000 0.0000 0.0829 0.0000 0.0000 0.0000
Acetic acid methyl ester 100.0000 0.0000 4.7934 0.0000 0.0000 0.0000
Methanol 0.0000 0.0000 27.4295 99.8988 0.0182 0.0000
Acetic acid ethyl ester 0.0000 0.0000 1.9095 0.0012 0.0000 0.0000
Ethanol 0.0000 0.0000 2.5214 0.1000 99.7390 90.0000
Heavy alcohols 0.0000 0.0000 0.0000 0.0000 0.0099 9.9677
Water (W) 0.0000 0.0000 0.0017 0.0000 0.2329 0.0322
Acetic acid 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000
Total mass flow rate, kg/hr 20742 1169 182 8910 12682 138
Comparative example 1
The comparative example is a process flow diagram of the cooperation of a traditional gas phase reaction device and a product separation device, and as shown in fig. 2, the specific process steps are as follows:
(i) the liquid phase acetic ester raw material S1 is methyl acetate, is pressurized to 5.0MPaG by a pump, is directly contacted with hot flow from a feeding and discharging heat exchanger a4 in a vaporizer 1 for mixing and gasification, and the temperature after gasification is 107 ℃;
(ii) (ii) heating the outlet gas phase of the vaporizer 1 in the step (i) to 235 ℃ through a superheater 2, then feeding the outlet gas phase into a gas phase hydrogenation reactor 3, adopting a copper-based catalyst, wherein the reaction pressure is 4.885MPaG, the reaction temperature is 240 ℃, the molar ratio of hydrogen to acetic ester is 20, the once-through conversion rate of methyl acetate is 97%, and the selectivity of ethanol is 96%; the gas phase hydrogenation reactor 3 is an isothermal tubular catalytic reactor, an acetate hydrogenation reaction occurs in a tube pass catalyst, and heat generated by the reaction is removed through partial gasification of boiler water on the shell pass of the gas phase hydrogenation reactor 3;
(iii) (iii) cooling the outlet gas phase of the gas-phase hydrogenation reactor 3 in the step (ii) at about 240 ℃ by sequentially passing through a feed-discharge heat exchanger a4 and a cooler a5, wherein the cooled temperature is 119 ℃ and 45 ℃ sequentially, performing gas-liquid phase separation on reaction products in a gas-liquid separator a6, separating liquid phase products in a 1# rectifying tower 9, pressurizing most of the feed S4 of a gas-phase flow compressor in the gas-phase products by a compressor 7, mixing with fresh hydrogen S2, heating by a feed-discharge heat exchanger a4, passing into a vaporizer 1, mixing the rest gas phase with the purge gas S6 of the gas-phase reaction device and the purge gas S7 at the top of the 1# rectifying tower 9, and discharging as total purge gas O1; the number of theoretical plates of the No. 1 rectifying tower 9 is 70, the operating pressure at the top of the tower is 69kPaG, and the reflux ratio is 3;
(iv) the overhead liquid phase of the rectifying tower 9 # 1 in the step (iii) is pressurized by a pump and then returned to the vaporizer 1, and is mixed with the acetic ester raw material S1 for gasification circulation, and the steps (i), (ii) and (iii) are repeated;
(v) and (3) the tower bottom liquid of the 1# rectifying tower 9 in the step (iv) enters a 2# rectifying tower 10 for separation, the top liquid phase of the 2# rectifying tower 10 is separated into a methanol product O2, the tower bottom liquid phase of the 2# rectifying tower 10 enters a 3# rectifying tower 11 for separation, the top liquid phase of the 3# rectifying tower 11 is separated into an absolute ethyl alcohol product O3, and the tower bottom liquid phase of the 3# rectifying tower 11 is separated into a heavy component O4 and is sent out of a boundary region. The number of theoretical plates of the 2# rectifying tower 10 is 80, the operation pressure at the top of the tower is 29kPaG, and the reflux ratio is 3.7; the number of theoretical plates of the 3# rectifying tower 11 is 30, the operating pressure at the top of the tower is 29kPaG, and the reflux ratio is 1.
Finally, the purity of the methanol product is more than or equal to 99.8 wt.% and the purity of the absolute ethyl alcohol product is more than or equal to 99.7 wt.% which have the same quality specification with the example 1 are obtained through the comparative example 1.
Tables 1-2 are the calculated results for comparative example 1:
logistics S1 S2 O1 O2 O3 O4
Phase state Liquid phase Gas phase Gas phase Liquid phase Liquid phase Liquid phase
Composition of wt％ wt％ wt％ wt％ wt％ wt％
Hydrogen gas 0.0000 99.0564 14.8817 0.0000 0.0000 0.0000
Nitrogen gas 0.0000 0.2755 1.6147 0.0000 0.0000 0.0000
Argon gas 0.0000 0.5892 3.4541 0.0000 0.0000 0.0000
Methane 0.0000 0.0789 0.4605 0.0000 0.0000 0.0000
Ethane (III) 0.0000 0.0000 9.0064 0.0000 0.0000 0.0000
Ether compounds 0.0000 0.0000 29.4624 0.0000 0.0000 0.0000
Acetaldehyde 0.0000 0.0000 3.6828 0.0000 0.0000 0.0000
Acetic acid methyl ester 100.0000 0.0000 7.1892 0.0000 0.0000 0.0000
Methanol 0.0000 0.0000 25.3024 99.8988 0.0182 0.0001
Acetic acid ethyl ester 0.0000 0.0000 2.8637 0.0012 0.0000 0.0000
Ethanol 0.0000 0.0000 2.0787 0.1000 99.7330 90.0000
Heavy alcohols 0.0000 0.0000 0.0004 0.0000 0.0059 9.9618
Water (W) 0.0000 0.0000 0.0029 0.0000 0.2428 0.0381
Acetic acid 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000
Total mass flow rate, kg/hr 20742 1169 200 8907 12647 158
Comparing the results calculated in example 1 with those calculated in comparative example 1, as shown in tables 1 to 3 and tables 1 to 4,
tables 1-3 recycle stream data
Figure BDA0001887936010000121
TABLE 1-4 comparison of plant loads
Figure BDA0001887936010000122
Compared with the comparative example 1, under the condition that the feeding S1 and S2 are the same and the hydrogen-ester ratio of the reactor feeding is the same, the amount of purge gas in the example 1 is reduced by 18kg/hr, the amount of methanol products is increased by 4kg/hr, the amount of ethanol products is increased by 35kg/hr and the heavy components are reduced by 21 kg/hr; the mass flow of superheater feed S3 of example 1 was reduced by 24%, and its temperature increased by 24 ℃; the mass flow of the compressor feed S4 was reduced by 18%, the temperature of the cooler feed S5 was reduced by 8 ℃, and therefore the superheater heat load was reduced by 36%, the cooler a load was reduced by 24%, the feed and discharge heat exchanger a heat load was reduced by 5%, and the compressor load was reduced by 4%. The invention can save the public engineering consumption of each ton of ethanol products as follows: circulating cooling water 21t, electricity 2KWh, medium pressure steam 4(MPa.G)0.5t and low pressure steam 0.5(MPa.G)0.1 t.
The comparison result shows that the invention increases the product yield, reduces the discharge of purge gas, reduces the circulation volume of a gas phase reaction device, reduces the separation difficulty of subsequent products and saves the energy consumption of the system.
Example 2
The process flow diagram of the embodiment is the same as that of the embodiment 1, and the specific process steps are as follows:
(i) the liquid phase acetic ester raw material S1 is a mixture of 50mole percent of methyl acetate and 50mole percent of ethyl acetate, is pressurized to 5.0MPaG by a pump, and is directly contacted with hot flow from a feed-discharge heat exchanger a4 to be mixed and gasified in a vaporizer 1, and the gasification temperature is 140 ℃;
(ii) heating the outlet gas phase of the vaporizer 1 in the step (i) to 235 ℃ through a superheater 2, then feeding the outlet gas phase into a gas phase hydrogenation reactor 3, and adopting a copper-based catalyst, wherein the reaction pressure is 4.885MPaG, the reaction temperature is 240 ℃, the molar ratio of hydrogen to acetic ester is 20, the once-through conversion rate of methyl acetate is 97%, the selectivity of ethanol is 96%, the once-through conversion rate of ethyl acetate is 98%, and the selectivity of ethanol is 99%; the gas phase hydrogenation reactor 3 is an isothermal tubular catalytic reactor, an acetate hydrogenation reaction occurs in a tube pass catalyst, and heat generated by the reaction is removed through partial gasification of boiler water on the shell pass of the gas phase hydrogenation reactor 3;
(iii) (iii) cooling the outlet gas phase of the gas-phase hydrogenation reactor 3 in the step (ii) at about 240 ℃ by a feed-discharge heat exchanger a4 and a cooler a5 in sequence, wherein the cooled temperature is 114 ℃ and 45 ℃ in sequence, performing gas-liquid phase separation on reaction products in a gas-liquid separator a6, feeding the liquid phase products into a No. 1 rectifying tower 9 for pre-separation, pressurizing most of the gas phase products, namely compressor feed S4 by a compressor 7, mixing with fresh hydrogen S2, heating by a feed-discharge heat exchanger a4, and feeding into a vaporizer 1; the rest gas phase is the purge gas S6 of the gas phase reaction device and is used as the feed of the liquid phase hydrogenation reactor 8; the number of theoretical plates of the No. 1 rectifying tower 9 is 70, the operating pressure at the top of the tower is 69kPaG, and the reflux ratio is 3;
(iv) the liquid phase at the top of the rectifying tower 9 in the step (iii) is pressurized by a pump and then mixed with part of purge gas of a gas-liquid separator a6, namely purge gas S6 of a gas-phase reaction device, and then the mixture is heated by a feeding and discharging heat exchanger b12 and then enters a liquid-phase hydrogenation reactor 8 for reaction, a nickel catalyst is adopted, the reaction pressure is 1.0MPaG, the reaction temperature is 80 ℃, the molar ratio of hydrogen to acetate is 3.5, the product after the reaction is sequentially cooled to 45 ℃ through a feeding and discharging heat exchanger b12 and a cooler b13, then the product is separated in a gas-liquid separator b14, the separated liquid phase product is returned to the rectifying tower 9 No. 1, and the separated gas is mixed with the purge gas S7 at the top of the rectifying tower 9 No. 1 and then discharged as total purge gas O1;
(v) and (3) introducing the tower bottom liquid of the 1# rectifying tower 9 in the step (iv) into a 2# rectifying tower 10 for separation, extracting a methanol product O2 from the tower top liquid phase of the 2# rectifying tower 10), introducing the tower bottom liquid phase of the 2# rectifying tower 10 into a 3# rectifying tower 11 for separation, extracting an absolute ethanol product O3 from the tower top liquid phase of the 3# rectifying tower 11, and extracting a heavy component O4 from the tower bottom liquid phase of the 3# rectifying tower 11 to be sent out of a boundary region. The number of theoretical plates of the 2# rectifying tower 10 is 80, the operation pressure at the top of the tower is 29kPaG, and the reflux ratio is 7.9; the number of theoretical plates of the 3# rectifying tower 11 is 30, the operating pressure at the top of the tower is 29kPaG, and the reflux ratio is 1.5.
Finally, the purity of the methanol product is more than or equal to 99.8 wt.% and the purity of the absolute ethyl alcohol product is more than or equal to 99.7 wt.% through the embodiment. Table 2-1 shows the calculation results of example 2:
table 2-1 material balance for example 2
Figure BDA0001887936010000131
Figure BDA0001887936010000141
Comparative example 2
The process flow diagram of comparative example 2 is the same as that of comparative example 1, and the specific process steps are as follows:
(i) the liquid phase acetic ester raw material S1 is a mixture of 50mole percent of methyl acetate and 50mole percent of ethyl acetate, is pressurized to 5.0MPaG by a pump, and is directly contacted with hot flow from a feed-discharge heat exchanger a4 to be mixed and gasified in a vaporizer 1, and the gasification temperature is 108 ℃;
(ii) heating the outlet gas phase of the vaporizer 1 in the step (i) to 235 ℃ through a superheater 2, then feeding the outlet gas phase into a gas phase hydrogenation reactor 3, and adopting a copper-based catalyst, wherein the reaction pressure is 4.885MPaG, the reaction temperature is 240 ℃, the molar ratio of hydrogen to acetic ester is 20, the once-through conversion rate of methyl acetate is 97%, the selectivity of ethanol is 96%, the once-through conversion rate of ethyl acetate is 98%, and the selectivity of ethanol is 99%; the gas phase hydrogenation reactor 3 is an isothermal tubular catalytic reactor, an acetate hydrogenation reaction occurs in a tube pass catalyst, and heat generated by the reaction is removed through partial gasification of boiler water on the shell pass of the gas phase hydrogenation reactor 3;
(iii) (iii) cooling the outlet gas phase of the gas-phase hydrogenation reactor 3 in the step (ii) at about 240 ℃ by sequentially passing through a feed-discharge heat exchanger a4 and a cooler a5, wherein the cooled temperature is 121 ℃ and 45 ℃ sequentially, performing gas-liquid phase separation on reaction products in a gas-liquid separator a6, separating the liquid phase products in a 1# rectifying tower 9, pressurizing most of the gas phase products, namely a compressor feed S4 by a compressor 7, mixing with fresh hydrogen S2, heating by a feed-discharge heat exchanger a4, passing into a vaporizer 1, mixing the rest gas phase with purge gas S6 of a gas-phase reaction device, and discharging as total purge gas O1 after mixing with a top purge gas S7 of a 1# rectifying tower 9; the number of theoretical plates of the No. 1 rectifying tower 9 is 70, the operating pressure at the top of the tower is 69kPaG, and the reflux ratio is 3;
(iv) the overhead liquid phase of the rectifying tower 9 # 1 in the step (iii) is pressurized by a pump and then returned to the vaporizer 1, and is mixed with the acetic ester raw material S1 for gasification circulation, and the steps (i), (ii) and (iii) are repeated;
(v) and (3) the tower bottom liquid of the 1# rectifying tower 9 in the step (iv) enters a 2# rectifying tower 10 for separation, the top liquid phase of the 2# rectifying tower 10 is separated into a methanol product O2, the tower bottom liquid phase of the 2# rectifying tower 10 enters a 3# rectifying tower 11 for separation, the top liquid phase of the 3# rectifying tower 11 is separated into an absolute ethyl alcohol product O3, and the tower bottom liquid phase of the 3# rectifying tower 11 is separated into a heavy component O4 and is sent out of a boundary region. The number of theoretical plates of the 2# rectifying tower 10 is 80, the operation pressure at the top of the tower is 29kPaG, and the reflux ratio is 9; the number of theoretical plates of the 3# rectifying tower 11 is 30, the operating pressure at the top of the tower is 29kPaG, and the reflux ratio is 1.5.
Finally, the purity of the methanol product is more than or equal to 99.8 wt.% and the purity of the absolute ethyl alcohol product is more than or equal to 99.7 wt.% which have the same quality specification with the example 1 are obtained through the comparative example 1.
Tables 2 to 2 are the calculation results of comparative example 2, and tables 2 to 3 and tables 2 to 4 are the comparison of the calculation results of example 2 and comparative example 2.
Tables 2-2 Material balance of comparative example 2
Logistics S1 S2 O1 O2 O3 O4
Phase state Liquid phase Gas phase Gas phase Liquid phase Liquid phase Liquid phase
Composition of wt％ wt％ wt％ wt％ wt％ wt％
Hydrogen gas 0.0000 99.0564 18.1081 0.0000 0.0000 0.0000
Nitrogen gas 0.0000 0.2755 1.9534 0.0000 0.0000 0.0000
Argon gas 0.0000 0.5892 4.2609 0.0000 0.0000 0.0000
Methane 0.0000 0.0789 0.5719 0.0000 0.0000 0.0000
Ethane (III) 0.0000 0.0000 13.5219 0.0000 0.0000 0.0000
Ether compounds 0.0000 0.0000 18.4546 0.0000 0.0000 0.0000
Acetaldehyde 0.0000 0.0000 2.3154 0.0000 0.0000 0.0000
Acetic acid methyl ester 45.6757 0.0000 4.5817 0.0000 0.0000 0.0000
Methanol 0.0000 0.0000 25.8834 99.8973 0.0184 0.0000
Acetic acid ethyl ester 54.3243 0.0000 4.2732 0.0027 0.0000 0.0000
Ethanol 0.0000 0.0000 6.0675 0.1000 99.8191 90.0001
Heavy alcohols 0.0000 0.0000 0.0014 0.0000 0.0032 9.9776
Water (W) 0.0000 0.0000 0.0064 0.0000 0.1593 0.0222
Acetic acid 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000
Total mass flow rate, kg/hr 22706 1169 160 4435 18823 458
Tables 2-3 recycle stream data
Figure BDA0001887936010000151
Figure BDA0001887936010000161
TABLE 2-4 comparison of plant loads
Figure BDA0001887936010000162
Compared with the comparative example 2, under the condition that the feeding S1 and S2 are the same and the hydrogen-ester ratio of the reactor feeding is the same, the amount of purge gas in the example 2 is reduced by 7kg/hr, the amount of ethanol products is increased by 23kg/hr, and the heavy components are reduced by 15 kg/hr; the mass flow of superheater feed S3 of example 2 was reduced by 20%, its temperature increased by 32 ℃; the mass flow of the recycle hydrogen compressor feed S4 was reduced by 14%, the temperature of the cooler a feed S5 was reduced by 7 ℃, and therefore the superheater heat load was reduced by 32%, the cooler a load was reduced by 21%, the feed heat exchanger a heat load was reduced by 4%, and the compressor load was reduced by 4%. The invention can save the public engineering consumption of each ton of ethanol products as follows: 11t of circulating cooling water, 1KWh of electricity, 0.3t of medium-pressure steam 4(MPa.G) and 0.003t of low-pressure steam 0.5 (MPa.G).
The comparison result shows that the invention increases the product yield, reduces the discharge of purge gas, reduces the circulation volume of a gas phase reaction device, reduces the separation difficulty of subsequent products and saves the energy consumption of the system.
Example 3
This example is applicable to the case where the acetate raw material is ethyl acetate, and the specific process flow diagram is shown in fig. 3, and the specific process steps are as follows:
(i) the liquid phase acetic ester raw material S1 is ethyl acetate, is pressurized to 5.0MPaG by a pump, is directly contacted with hot flow from a feeding and discharging heat exchanger a4 in a vaporizer 1 for mixing and gasification, and the temperature after gasification is 140 ℃;
(ii) (ii) heating the outlet gas phase of the vaporizer 1 in the step (i) to 235 ℃ through a superheater 2, then feeding the outlet gas phase into a gas phase hydrogenation reactor 3, and adopting a copper-based catalyst, wherein the reaction pressure is 4.885MPaG, the reaction temperature is 240 ℃, the molar ratio of hydrogen to acetic ester is 20.1, the once-through conversion rate of ethyl acetate is 98%, and the selectivity of ethanol is 99%; the gas phase hydrogenation reactor 3 is an isothermal tubular catalytic reactor, an acetate hydrogenation reaction occurs in a tube pass catalyst, and heat generated by the reaction is removed through partial gasification of boiler water on the shell pass of the gas phase hydrogenation reactor 3;
(iii) (iii) cooling the outlet gas phase of the gas-phase hydrogenation reactor 3 in the step (ii) at about 240 ℃ by a feed-discharge heat exchanger a4 and a cooler a5 in sequence, wherein the cooled temperature is 116 ℃ and 45 ℃ in sequence, performing gas-liquid phase separation on the reaction product in a gas-liquid separator a6, feeding the liquid phase product into a No. 1 rectifying tower 9 for pre-separation, pressurizing most of the gas phase product, namely the compressor feed S4 by a compressor 7, mixing with fresh hydrogen S2, heating by a feed-discharge heat exchanger a4, and feeding the gas phase product into a vaporizer 1; the rest gas phase is the purge gas S6 of the gas phase reaction device and is used as the feed of the liquid phase hydrogenation reactor 8; the number of theoretical plates of the No. 1 rectifying tower 9 is 70, the operating pressure at the top of the tower is 69kPaG, and the reflux ratio is 3;
(iv) the liquid phase at the top of the rectifying tower 9 in the step (iii) is pressurized by a pump and then mixed with part of purge gas of a gas-liquid separator a6, namely purge gas S6 of a gas-phase reaction device, and then the mixture is heated by a feeding and discharging heat exchanger b12 and then enters a liquid-phase hydrogenation reactor 8 for reaction, a nickel-based catalyst is adopted, the reaction pressure is 1.0MPaG, the reaction temperature is 80 ℃, the molar ratio of hydrogen to acetate is 4.4, the product after the reaction is sequentially cooled to 45 ℃ through a feeding and discharging heat exchanger b12 and a cooler b13, then the product is separated in a gas-liquid separator b14, the separated liquid-phase product is returned to the rectifying tower 9 at the top, and the separated gas is mixed with the purge gas S7 at the top of the rectifying tower 9 at the 1# and then discharged as total purge gas O1;
(v) and (3) feeding the tower bottom liquid of the 1# rectifying tower 9 in the step (iv) into a 3# rectifying tower 11 for separation, extracting an absolute ethanol product O3 from the tower top liquid phase of the 3# rectifying tower 11, and extracting a heavy component O4 from the tower bottom liquid phase of the 3# rectifying tower 11 and sending the heavy component out of a battery limit. The number of theoretical plates of the 3# rectifying tower 11 is 30, the operating pressure at the top of the tower is 29kPaG, and the reflux ratio is 1.6.
Finally, the purity of the absolute ethyl alcohol product obtained by the embodiment is more than or equal to 99.7 wt.%. Table 3-1 shows the calculation results of example 3:
table 3-1 material balance for example 3
Figure BDA0001887936010000171
Figure BDA0001887936010000181
Comparative example 3
The process flow diagram of the comparative example is shown in fig. 4, and is suitable for the case that the acetate raw material is ethyl acetate, and the process steps are as follows:
(i) the liquid phase acetic ester raw material S1 is ethyl acetate, is pressurized to 5.0MPaG by a pump, is directly contacted with hot flow from a feeding and discharging heat exchanger a4 in a vaporizer 1 for mixing and gasification, and the temperature after gasification is 112 ℃;
(ii) (ii) heating the outlet gas phase of the vaporizer 1 in the step (i) to 235 ℃ through a superheater 2, then feeding the outlet gas phase into a gas phase hydrogenation reactor 3, and adopting a copper-based catalyst, wherein the reaction pressure is 4.885MPaG, the reaction temperature is 240 ℃, the molar ratio of hydrogen to acetic ester is 20.1, the once-through conversion rate of ethyl acetate is 98%, and the selectivity of ethanol is 99%; the gas phase hydrogenation reactor 3 is an isothermal tubular catalytic reactor, an acetate hydrogenation reaction occurs in a tube pass catalyst, and heat generated by the reaction is removed through partial gasification of boiler water on the shell pass of the gas phase hydrogenation reactor 3;
(iii) (iii) cooling the outlet gas phase of the gas-phase hydrogenation reactor 3 in the step (ii) at about 240 ℃ by sequentially passing through a feed-discharge heat exchanger a4 and a cooler a5, wherein the cooled temperature is 123 ℃ and 45 ℃, performing gas-liquid phase separation on reaction products in a gas-liquid separator a6, separating the liquid phase products by passing through a 1# rectifying tower 9, pressurizing most of the gas phase products, namely a compressor feed S4 by a compressor 7, mixing with fresh hydrogen S2, heating by a feed-discharge heat exchanger a4, passing through a vaporizer 1, mixing the rest gas phase with a purge gas S6 of a gas-phase reaction device and a purge gas S7 at the top of the 1# rectifying tower 9, and discharging as a total purge gas O1; the number of theoretical plates of the No. 1 rectifying tower 9 is 70, the operating pressure at the top of the tower is 69kPaG, and the reflux ratio is 3;
(iv) the overhead liquid phase of the rectifying tower 9 # 1 in the step (iii) is pressurized by a pump and then returned to the vaporizer 1, and is mixed with the acetic ester raw material S1 for gasification circulation, and the steps (i), (ii) and (iii) are repeated;
(v) and (3) introducing the tower bottom liquid of the 1# rectifying tower 9 in the step (iv) into a 3# rectifying tower 11 for separation, extracting an absolute ethanol product O3 from the tower top liquid phase of the 3# rectifying tower 11, and extracting a heavy component O4 from the tower bottom liquid phase of the 3# rectifying tower 11 and sending the heavy component out of a battery limit. The number of theoretical plates of the 3# rectifying tower 11 is 30, the operating pressure at the top of the tower is 29kPaG, and the reflux ratio is 1.6.
Finally, the purity of the absolute ethyl alcohol product with the same quality specification as that of the absolute ethyl alcohol product obtained in the example 3 is more than or equal to 99.7 wt% through the comparative example 3.
Table 3-2 is the calculation result of comparative example 3:
tables 3-2 Material balance of comparative example 3
Logistics S1 S2 O1 O3 O4
Phase state Liquid phase Gas phase Gas phase Liquid phase Liquid phase
Composition of wt％ wt％ wt％ wt％ wt％
Hydrogen gas 0.0000 99.0564 25.7947 0.0000 0.0000
Nitrogen gas 0.0000 0.2755 2.8662 0.0000 0.0000
Argon gas 0.0000 0.5892 6.2458 0.0000 0.0000
Methane 0.0000 0.0789 0.8424 0.0000 0.0000
Ethane (III) 0.0000 0.0000 23.0680 0.0000 0.0000
Ether compounds 0.0000 0.0000 0.0000 0.0000 0.0000
Acetaldehyde 0.0000 0.0000 0.0000 0.0000 0.0000
Acetic acid methyl ester 0.0000 0.0000 0.0000 0.0000 0.0000
Methanol 0.0000 0.0000 0.0000 0.0000 0.0000
Acetic acid ethyl ester 100.0000 0.0000 7.3994 0.0005 0.0000
Ethanol 0.0000 0.0000 33.7010 99.8797 90.0000
Heavy alcohols 0.0000 0.0000 0.0056 0.0031 9.9752
Water (W) 0.0000 0.0000 0.0769 0.1167 0.0248
Acetic acid 0.0000 0.0000 0.0000 0.0000 0.0000
Total mass flow rate, kg/hr 24670 1169 109 24976 754
TABLE 3-3 recycle stream data
Figure BDA0001887936010000191
TABLE 3-4 comparison of plant loads
Figure BDA0001887936010000192
Compared with the comparative example 3, under the condition that the feeding S1 and S2 are the same and the hydrogen-ester ratio of the reactor feeding is the same, the amount of purge gas in the example 3 is reduced by 1kg/hr, the amount of ethanol products is increased by 14kg/hr, and the heavy components are reduced by 13 kg/hr; the mass flow of superheater feed S3 of example 3 was reduced by 17%, and its temperature increased by 28 ℃; the mass flow of the recycle hydrogen compressor feed S4 was reduced by 8%, the temperature of the cooler a feed S5 was reduced by 7 ℃, and therefore the superheater heat load was reduced by 28%, the cooler a load was reduced by 19%, the feed heat exchanger a heat load was reduced by 2%, and the compressor load was reduced by 3%. The invention can save the public engineering consumption of each ton of ethanol products as follows: 8t of circulating cooling water, 1KWh of electricity, 0.2t of medium-pressure steam 4(MPa.G) and 0.003t of low-pressure steam 0.5 (MPa.G).
The comparison result shows that the invention increases the product yield, reduces the discharge of purge gas, reduces the circulation volume of a gas phase reaction device, reduces the separation difficulty of subsequent products and saves the energy consumption of the system.
Example 4
The process flow diagram of the embodiment is shown in fig. 1, and the specific process steps are as follows:
(i) the liquid phase acetic ester raw material S1 is methyl acetate, is pressurized to 1.5MPaG by a pump, is directly contacted with hot flow from a feeding and discharging heat exchanger a4 in a vaporizer 1 for mixing and gasification, and the temperature after gasification is 67 ℃;
(ii) heating the outlet gas phase of the vaporizer 1 in the step i to 200 ℃ through a superheater 2, then entering a gas phase hydrogenation reactor 3, adopting a copper-based catalyst, wherein the reaction pressure is 1.385MPaG, the reaction temperature is 200 ℃, the molar ratio of hydrogen to acetic ester is 10, the once-through conversion rate of methyl acetate is 97%, and the selectivity of ethanol is 96%; the gas-phase hydrogenation reactor 3 is an isothermal tubular catalytic reactor, an acetate hydrogenation reaction occurs in a tube-side catalyst of the gas-phase hydrogenation reactor, and heat generated by the reaction is removed through partial gasification of boiler water on the shell side of the reactor 3;
(iii) (iii) cooling the outlet gas phase of the gas-phase hydrogenation reactor 3 in the step (ii) at about 200 ℃ by a feed-discharge heat exchanger a4 and a cooler a5 in sequence, wherein the cooled temperature is 96 ℃ and 20 ℃ in sequence, performing gas-liquid phase separation on reaction products in a gas-liquid separator a6, feeding liquid phase products into a No. 1 rectifying tower 9 for pre-separation, pressurizing most of the gas phase products, namely compressor feed S4 by a compressor 7, mixing with fresh hydrogen S2, heating by a feed-discharge heat exchanger a4, and feeding the gas phase products into a vaporizer 1; the rest gas phase is purge gas S6 of the gas phase reaction device and is used as the feeding material of the liquid phase hydrogenation reactor 8; the number of theoretical plates of the No. 1 rectifying tower 9 is 70, the operating pressure at the top of the tower is 69kPaG, and the reflux ratio is 3;
(iv) the liquid phase at the top of the rectifying tower 9 # in the step (iii) is pressurized by a pump and then mixed with part of purge gas of a gas-liquid separator a6, namely purge gas S6 of a gas-phase reaction device, and then the mixture is heated by a feeding and discharging heat exchanger b12 and then enters a liquid-phase hydrogenation reactor 8 for reaction, a nickel-based catalyst is adopted, the reaction pressure is 1.0MPaG, the reaction temperature is 150 ℃, the molar ratio of hydrogen to acetic ester is 5, the product after the reaction is sequentially cooled to 20 ℃ through a feeding and discharging heat exchanger b12 and a cooler b13, then the product is separated in a gas-liquid separator b14, the separated liquid-phase product is returned to the rectifying tower 9 # 1, and the separated gas is mixed with the purge gas S7 at the top of the rectifying tower 9 # 1 and then discharged as total purge gas O1;
(v) and (3) feeding the tower bottom liquid of the 1# rectifying tower 9 in the step (iv) into a 2# rectifying tower 10 for separation, extracting a methanol product O2 from the tower top liquid phase of the 2# rectifying tower 10, feeding the tower bottom liquid phase of the 2# rectifying tower 10 into a 3# rectifying tower 11 for separation, extracting an absolute ethanol product O3 from the tower top liquid phase of the 3# rectifying tower 11, and extracting a heavy component O4 from the tower bottom liquid phase of the 3# rectifying tower 11 to be sent out of a boundary region. The number of theoretical plates of the 2# rectifying tower 10 is 80, the operation pressure at the top of the tower is 29kPaG, and the reflux ratio is 3.2; the number of theoretical plates of the 3# rectifying tower 11 is 30, the operating pressure at the top of the tower is 29kPaG, and the reflux ratio is 1.
Finally, the purity of the methanol product is more than or equal to 99.8 wt.% and the purity of the absolute ethyl alcohol product is more than or equal to 99.7 wt.% through the embodiment. Tables 4-1, 4-2, and 4-3 are the calculation results of example 4, respectively:
table 4-1 material balance for example 4
Logistics S1 S2 O1 O2 O3 O4
Phase state Liquid phase Gas phase Gas phase Liquid phase Liquid phase Liquid phase
Composition of wt％ wt％ wt％ wt％ wt％ wt％
Hydrogen gas 0.0000 99.0564 23.3701 0.0000 0.0000 0.0000
Nitrogen gas 0.0000 0.2755 2.6347 0.0000 0.0000 0.0000
Argon gas 0.0000 0.5892 5.6406 0.0000 0.0000 0.0000
Methane 0.0000 0.0789 0.7564 0.0000 0.0000 0.0000
Ethane (III) 0.0000 0.0000 14.1765 0.0000 0.0000 0.0000
Ether compounds 0.0000 0.0000 46.9958 0.0000 0.0000 0.0000
Acetaldehyde 0.0000 0.0000 0.0064 0.0000 0.0000 0.0000
Acetic acid methyl ester 100.0000 0.0000 0.2853 0.0000 0.0000 0.0000
Methanol 0.0000 0.0000 5.4706 99.8988 0.0182 0.0000
Acetic acid ethyl ester 0.0000 0.0000 0.0980 0.0012 0.0000 0.0000
Ethanol 0.0000 0.0000 0.5649 0.1000 99.7396 90.0000
Heavy alcohols 0.0000 0.0000 0.0000 0.0000 0.0103 9.9678
Water (W) 0.0000 0.0000 0.0005 0.0000 0.2319 0.0322
Acetic acid 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000
Total mass flow rate, kg/hr 20742 1169 122 8957 12697 136
TABLE 4-2 recycle stream data
Figure BDA0001887936010000211
TABLE 4-3 Equipment loads
Figure BDA0001887936010000221
Example 5
The process flow diagram of the embodiment is shown in fig. 1, and the specific process steps are as follows:
(i) the liquid phase acetic ester raw material S1 is methyl acetate, is pressurized to 6MPaG by a pump, is directly contacted with hot flow from a feeding and discharging heat exchanger a4 in a vaporizer 1 for mixing and gasification, and the temperature after gasification is 237 ℃;
(ii) heating the outlet gas phase of the vaporizer 1 in the step (i) to 300 ℃ through a heater 2, then feeding the outlet gas phase into a gas phase hydrogenation reactor 3, adopting a copper-based catalyst, wherein the reaction pressure is 5.885MPaG, the reaction temperature is 300 ℃, the molar ratio of hydrogen to acetic ester is 50, the once-through conversion rate of methyl acetate is 97%, and the selectivity of ethanol is 96%; the gas phase hydrogenation reactor 3 is an isothermal tubular catalytic reactor, an acetate hydrogenation reaction occurs in a tube pass catalyst, and heat generated by the reaction is removed through partial gasification of boiler water on the shell pass of the gas phase hydrogenation reactor 3;
(iii) (iii) cooling the outlet gas phase of the reactor 3 in the step (ii) at about 300 ℃ by a feed-discharge heat exchanger a4 and a cooler a5 in sequence, wherein the cooled temperature is 99 ℃ and 50 ℃ in sequence, performing gas-liquid phase separation on reaction products in a gas-liquid separator a6, feeding liquid phase products into a No. 1 rectifying tower 9 for pre-separation, pressurizing most of material flow in the gas phase products, namely compressor feed S4 by a compressor 7, mixing with fresh hydrogen S2, heating by a feed-discharge heat exchanger a4, and feeding the mixture into a vaporizer 1; the rest gas phase is purge gas S6 of the gas phase reaction device and is used as the feeding material of the liquid phase hydrogenation reactor 8; the number of theoretical plates of the No. 1 rectifying tower 9 is 70, the operating pressure at the top of the tower is 69kPaG, and the reflux ratio is 3;
(iv) the liquid phase at the top of the rectifying tower 9 # in the step (iii) is pressurized by a pump and then mixed with partial purge gas of a gas-liquid separator a6, namely the purge gas S6 of the gas-phase reaction device, and then the mixture is heated by a feeding and discharging heat exchanger b12 and then enters a liquid-phase hydrogenation reactor 8 for reaction, a nickel-based catalyst is adopted, the reaction pressure is 5.0MPaG, the reaction temperature is 250 ℃, the molar ratio of hydrogen to acetic ester is 6, the product after the reaction is sequentially cooled to 50 ℃ by a feeding and discharging heat exchanger b12 and a cooler b13, then the product is separated in a gas-liquid separator b14, the separated liquid-phase product is returned to the rectifying tower 9 # 1, and the separated gas is mixed with the purge gas S7 at the top of the rectifying tower 9 # 1 and then discharged as the total purge gas O1;
(v) and (3) feeding the tower bottom liquid of the 1# rectifying tower 9 in the step (iv) into a 2# rectifying tower 10 for separation, extracting a methanol product O2 from the tower top liquid phase of the 2# rectifying tower 10, feeding the tower bottom liquid phase of the 2# rectifying tower 10 into a 3# rectifying tower 11 for separation, extracting an absolute ethanol product O3 from the tower top liquid phase of the 3# rectifying tower 11, and extracting a heavy component O4 from the tower bottom liquid phase of the 3# rectifying tower 11 to be sent out of a boundary region. The number of theoretical plates of the 2# rectifying tower 10 is 80, the operation pressure at the top of the tower is 29kPaG, and the reflux ratio is 3.2; the number of theoretical plates of the 3# rectifying tower 11 is 30, the operating pressure at the top of the tower is 29kPaG, and the reflux ratio is 1.
Finally, the purity of the methanol product is more than or equal to 99.8 wt.% and the purity of the absolute ethyl alcohol product is more than or equal to 99.7 wt.% through the embodiment. Tables 5-1, 5-2, and 5-3 are the calculation results of example 5, respectively:
TABLE 5-1 Material balance for example 5
Logistics S1 S2 O1 O2 O3 O4
Phase state Liquid phase Gas phase Gas phase Liquid phase Liquid phase Liquid phase
Composition of wt％ wt％ wt％ wt％ wt％ wt％
Hydrogen gas 0.0000 99.0564 15.1645 0.0000 0.0000 0.0000
Nitrogen gas 0.0000 0.2755 1.6366 0.0000 0.0000 0.0000
Argon gas 0.0000 0.5892 3.5302 0.0000 0.0000 0.0000
Methane 0.0000 0.0789 0.4782 0.0000 0.0000 0.0000
Ethane (III) 0.0000 0.0000 9.0932 0.0000 0.0000 0.0000
Ether compounds 0.0000 0.0000 30.0905 0.0000 0.0000 0.0000
Acetaldehyde 0.0000 0.0000 0.1330 0.0000 0.0000 0.0000
Acetic acid methyl ester 100.0000 0.0000 5.7720 0.0000 0.0000 0.0000
Methanol 0.0000 0.0000 29.5215 99.8988 0.0182 0.0000
Acetic acid ethyl ester 0.0000 0.0000 2.4517 0.0012 0.0000 0.0000
Ethanol 0.0000 0.0000 2.1269 0.1000 99.7378 90.0000
Heavy alcohols 0.0000 0.0000 0.0000 0.0000 0.0087 9.9676
Water (W) 0.0000 0.0000 0.0018 0.0000 0.2353 0.0324
Acetic acid 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000
Total mass flow rate, kg/hr 20742 1169 193 8902 12674 143
TABLE 5-2 recycle stream data
Figure BDA0001887936010000231
TABLE 5-3 plant loads
Figure BDA0001887936010000232
Figure BDA0001887936010000241
Example 6
The process flow diagram of the embodiment is shown in fig. 1, and the specific process steps are as follows:
(i) the liquid phase acetic ester raw material S1 is methyl acetate, is pressurized to 6MPaG by a pump, is directly contacted with hot flow from a feeding and discharging heat exchanger a4 in a vaporizer 1 for mixing and gasification, and the temperature after gasification is 150 ℃;
(ii) (ii) heating the outlet gas phase of the vaporizer 1 in the step (i) to 200 ℃ through a superheater 2, then feeding the outlet gas phase into a gas phase hydrogenation reactor 3, adopting a copper-based catalyst, wherein the reaction pressure is 1MPaG, the reaction temperature is 200 ℃, the molar ratio of hydrogen to acetic ester is 2, the one-way conversion rate of methyl acetate is 90%, and the selectivity of ethanol is 89%; the gas phase hydrogenation reactor 3 is an isothermal tubular catalytic reactor, an acetate hydrogenation reaction occurs in a tube pass catalyst, and heat generated by the reaction is removed through partial gasification of boiler water on the shell pass of the gas phase hydrogenation reactor 3;
(iii) cooling the outlet gas phase of the gas-phase hydrogenation reactor 3 in the step ii at the temperature of about 200 ℃ by a feeding and discharging heat exchanger a4 and a cooler a5 in sequence, wherein the cooled temperature is 99 ℃ and 20 ℃ in sequence, performing gas-liquid phase separation on reaction products in a gas-liquid separator a6, feeding the liquid phase products into a No. 1 rectifying tower 9 for pre-separation, pressurizing most of the gas phase products, namely compressor feeding S4 by a compressor 7, mixing with fresh hydrogen S2, heating by a feeding and discharging heat exchanger a4, and feeding the gas phase products into a vaporizer 1; the rest gas phase is purge gas S6 of the gas phase reaction device and is used as the feeding material of the liquid phase hydrogenation reactor 8; the number of theoretical plates of the No. 1 rectifying tower 9 is 70, the operating pressure at the top of the tower is 1kPaG, and the reflux ratio is 0.01;
(iv) the liquid phase at the top of the rectifying tower 9 # in the step (iii) is pressurized by a pump and then mixed with part of purge gas of a gas-liquid separator a6, namely purge gas S6 of a gas-phase reaction device, and then the mixture is heated by a feeding and discharging heat exchanger b12 and enters a liquid-phase hydrogenation reactor 8 for reaction, a nickel-based catalyst is adopted, the reaction pressure is 1MPaG, the reaction temperature is 200 ℃, the molar ratio of hydrogen and acetic ester in the reaction material is 2, the product after the reaction is sequentially cooled to 10 ℃ by a feeding and discharging heat exchanger b12 and a cooler b13, then the product is separated in a gas-liquid separator b14, the separated liquid-phase product is returned to the rectifying tower 9 # 1, and the separated gas is mixed with the purge gas S7 at the top of the rectifying tower 9 # 1 and then discharged as total purge gas O1;
(v) and (3) feeding the tower bottom liquid of the 1# rectifying tower 9 in the step (iv) into a 2# rectifying tower 10 for separation, extracting a methanol product O2 from the tower top liquid phase of the 2# rectifying tower 10, feeding the tower bottom liquid phase of the 2# rectifying tower 10 into a 3# rectifying tower 11 for separation, extracting an absolute ethanol product O3 from the tower top liquid phase of the 3# rectifying tower 11, and extracting a heavy component O4 from the tower bottom liquid phase of the 3# rectifying tower 11 to be sent out of a boundary region. The number of theoretical plates of the 2# rectifying tower 10 is 80, the operation pressure at the top of the tower is 0kPaG, and the reflux ratio is 3; the number of theoretical plates of the 3# rectifying tower 11 is 30, the operating pressure at the top of the tower is 0kPaG, and the reflux ratio is 0.7.
Finally, the purity of the methanol product is more than or equal to 99.8 wt.% and the purity of the absolute ethyl alcohol product is more than or equal to 99.7 wt.% through the embodiment.
Example 7
The process flow diagram of the embodiment is shown in fig. 1, and the specific process steps are as follows:
(i) the liquid phase acetic ester raw material S1 is methyl acetate, is pressurized to 6MPaG by a pump, is directly contacted with hot flow from a feeding and discharging heat exchanger a4 in a vaporizer 1 for mixing and gasification, and the temperature after gasification is 250 ℃;
(ii) heating the outlet gas phase of the vaporizer 1 in the step (i) to 300 ℃ through a heater 2, then feeding the outlet gas phase into a gas phase hydrogenation reactor 3, adopting a copper-based catalyst, wherein the reaction pressure is 6MPaG, the reaction temperature is 300 ℃, the molar ratio of hydrogen to acetic ester is 50, the once-through conversion rate of methyl acetate is 97%, and the selectivity of ethanol is 96%; the gas phase hydrogenation reactor 3 is an isothermal tubular catalytic reactor, an acetate hydrogenation reaction occurs in a tube pass catalyst, and heat generated by the reaction is removed through partial gasification of boiler water on the shell pass of the gas phase hydrogenation reactor 3;
(iii) (iii) cooling the outlet gas phase of the gas-phase hydrogenation reactor 3 in the step (ii) at about 300 ℃ by a feed-discharge heat exchanger a4 and a cooler a5 in sequence, wherein the cooled temperature is 99 ℃ and 50 ℃ in sequence, performing gas-liquid phase separation on reaction products in a gas-liquid separator a6, feeding liquid phase products into a No. 1 rectifying tower 9 for pre-separation, pressurizing most of the gas phase products, namely compressor feed S4 by a compressor 7, mixing with fresh hydrogen S2, heating by a feed-discharge heat exchanger a4, and feeding the gas phase products into a vaporizer 1; the rest gas phase is purge gas S6 of the gas phase reaction device and is used as the feeding material of the liquid phase hydrogenation reactor 8; the number of theoretical plates of the No. 1 rectifying tower 9 is 70, the operating pressure at the top of the tower is 500kPaG, and the reflux ratio is 100;
(iv) the liquid phase at the top of the rectifying tower 9 # in the step (iii) is pressurized by a pump and then mixed with partial purge gas of a gas-liquid separator a6, namely the purge gas S6 of the gas-phase reaction device, and then the mixture is heated by a feeding and discharging heat exchanger b12 and then enters a liquid-phase hydrogenation reactor 8 for reaction, a nickel-based catalyst is adopted, the reaction pressure is 5MPaG, the reaction temperature is 250 ℃, the molar ratio of hydrogen and acetic ester in the reaction material is 100, the product after the reaction is sequentially cooled to 50 ℃ through a feeding and discharging heat exchanger b12 and a cooler b13, then the product is separated in a gas-liquid separator b14, the separated liquid-phase product is returned to the rectifying tower 9 # 1, and the separated gas is mixed with the purge gas S7 at the top of the rectifying tower 9 # 1 and then discharged as the total purge gas O1;
(v) and (3) feeding the tower bottom liquid of the 1# rectifying tower 9 in the step (iv) into a 2# rectifying tower 10 for separation, extracting a methanol product O2 from the tower top liquid phase of the 2# rectifying tower 10, feeding the tower bottom liquid phase of the 2# rectifying tower 10 into a 3# rectifying tower 11 for separation, extracting an absolute ethanol product O3 from the tower top liquid phase of the 3# rectifying tower 11, and extracting a heavy component O4 from the tower bottom liquid phase of the 3# rectifying tower 11 to be sent out of a boundary region. The number of theoretical plates of the 2# rectifying tower 10 is 80, the operation pressure at the top of the tower is 500kPaG, and the reflux ratio is 100; the number of theoretical plates of the 3# rectifying column 11 was 30, the operating pressure at the top of the column was 500kPaG, and the reflux ratio was 100.
Finally, the purity of the methanol product is more than or equal to 99.8 wt.% and the purity of the absolute ethyl alcohol product is more than or equal to 99.7 wt.% through the embodiment.
The foregoing description of specific embodiments of the present invention has been presented. It is to be understood that the present invention is not limited to the specific embodiments described above, and that various changes and modifications may be made by one skilled in the art within the scope of the appended claims without departing from the spirit of the invention.

Claims (15)

1. A process method for preparing ethanol by acetate hydrogenation comprises the following steps:
gas phase hydrogenation
Mixing and gasifying an acetate raw material (S1) and a hydrogen raw material to obtain a steam feed, preheating the steam feed, and adding the preheated steam feed into a gas-phase hydrogenation reactor (3) for hydrogenation reaction; condensing a reaction product output by the gas-phase hydrogenation reactor, and performing gas-liquid separation in a gas-liquid separator to obtain a gas-phase product A and a liquid-phase product A, wherein at least part of the gas-phase product A is recycled to the gas-phase hydrogenation reactor (3);
product separation
Sending the liquid phase product A in the gas phase hydrogenation reaction into a product separation device for separation to obtain a liquid phase product B containing acetic ester and carbonyl byproducts and an absolute ethyl alcohol product (O3);
the method is characterized by further comprising the following steps:
liquid phase hydrogenation reaction
Mixing and preheating hydrogen and the liquid phase product B, adding the mixture into a liquid phase hydrogenation reactor (8) for hydrogenation reaction, condensing a reaction product output by the liquid phase hydrogenation reactor (8), performing gas-liquid separation in a gas-liquid separator to obtain a gas phase product C and a liquid phase product C, and separating the liquid phase product C in a product separation device.
2. The process method for preparing ethanol by hydrogenating acetic ester according to claim 1, wherein the liquid phase hydrogenation reaction specifically comprises the following steps:
mixing hydrogen and the liquid-phase product B to form a liquid-phase hydrogenation reaction raw material, preheating the liquid-phase hydrogenation reaction raw material by a feeding and discharging heat exchanger B (12), then feeding the liquid-phase hydrogenation reaction raw material into a liquid-phase hydrogenation reactor (8) for liquid-phase hydrogenation reaction, cooling the obtained reaction product by the feeding and discharging heat exchanger B (12) and a cooler B (13), and then feeding the cooled reaction product into a gas-liquid separator B (14) to form a liquid-phase product C and a gas-phase product C; the liquid-phase product C enters a product separation device for separation; the gas phase product C is withdrawn as purge gas.
3. The process method for preparing ethanol by hydrogenating acetate according to claim 2, wherein the catalyst in the liquid phase hydrogenation reactor (8) is a nickel-based catalyst, the reaction pressure is 1.0-5.0 MPaG, the reaction temperature is 80-250 ℃, and the molar ratio of hydrogen to acetate in the liquid phase hydrogenation raw material is 2-100;
the cooling medium of the cooler b (13) is circulating water or chilled water; and the material temperature of the gas-phase hydrogenation reaction product at the outlet of the cooler b (13) is 10-50 ℃.
4. The process of claim 2, wherein the hydrogen in the liquid phase hydrogenation reaction is from the gas phase product A or fresh hydrogen in the gas phase hydrogenation reaction (S2).
5. The process method for preparing ethanol by hydrogenating acetic ester according to claim 1, wherein the gas phase hydrogenation reaction specifically comprises the following steps:
adding an acetate raw material (S1) into a vaporizer (1), mixing fresh hydrogen (S2) and circulating hydrogen to form a hydrogen raw material, preheating the hydrogen raw material by a feeding and discharging heat exchanger a (4), then introducing the preheated hydrogen raw material into the vaporizer, obtaining a hydrogen and acetate mixture at an outlet of the vaporizer, heating the hydrogen and acetate mixture by a heater (2), introducing the heated hydrogen and acetate mixture into a gas phase hydrogenation reactor (3) for gas phase hydrogenation reaction to obtain a gas phase hydrogenation reaction product, cooling the gas phase hydrogenation reaction product by the feeding and discharging heat exchanger a (4) and a cooler a (5), and introducing the cooled gas phase hydrogenation reaction product into a gas-liquid separator a (6) to obtain a gas phase product A and a liquid phase product A, wherein the gas phase product A is divided into a compressor feeding (S4) and purge gas (S6) of a gas phase reaction device; and the compressor feed (S4) is pressurized by the compressor (7) and then is used as a circulating hydrogen raw material to continuously participate in the gas-phase hydrogenation reaction, and the purge gas (S6) of the gas-phase reaction device is used for carrying out the liquid-phase hydrogenation reaction.
6. The process method for preparing ethanol by hydrogenating acetic ester according to claim 5, wherein the acetic ester raw material (S1) and the hydrogen raw material are contacted, mixed and gasified in a vaporizer (1), and the gasification pressure is 1.0-6.0 MPaG; the temperature of the material at the outlet of the superheater (2) is 200-300 ℃;
the acetic ester gasification heat source in the superheater is the sensible heat of the circulating material flow at the outlet of the feed-discharge heat exchanger a (4).
7. The process of preparing ethanol by hydrogenation of acetic ester according to claim 5, wherein the step of adding hydrogen into acetic ester,
the composition of the acetate feedstock (S1) comprises methyl acetate, ethyl acetate or a mixture of the two;
the catalyst in the gas phase hydrogenation reactor (3) is a copper catalyst, the temperature in the reactor (3) is 200-300 ℃, the pressure is 1.0-6.0 MPaG, the molar ratio of hydrogen to acetic ester at the inlet of the gas phase hydrogenation reactor (3) is 2-50, the gas phase hydrogenation reactor (3) is an isothermal tubular reactor, the tube side is filled with the catalyst, and the shell side is filled with boiler water as a heat-taking medium;
the cooling medium of the cooler a (5) is circulating water or chilled water; the material temperature of a gas-phase hydrogenation reaction product at the outlet of the cooler a (5) is 20-50 ℃;
the fresh hydrogen (S2) is mixed with the recycled hydrogen at the inlet or outlet of the compressor (7).
8. The process method for preparing ethanol by hydrogenating acetic ester according to claim 1, wherein the product separation specifically comprises the following steps:
the liquid phase product A enters a No. 1 rectifying tower (9) for rectification separation, the gas phase at the top of the tower is condensed by a condenser to obtain a liquid phase product B and a gas phase product B, the gas phase product B is taken out as purge gas, and the liquid phase product B enters a liquid phase reaction device;
and the liquid material in the tower bottom of the No. 1 rectifying tower (9) enters a product rectifying tower to be separated to obtain an absolute ethyl alcohol product (O3).
9. The process method for preparing ethanol by hydrogenating acetic ester according to claim 8, wherein a product rectifying tower is arranged according to the composition of acetic ester raw materials;
when the acetic ester raw material component is ethyl acetate, the product rectifying tower comprises a 3# rectifying tower (11), and an absolute ethyl alcohol product (O3) is extracted from the top of the 3# rectifying tower (11);
when acetate raw materials component is methyl acetate or ethyl acetate and methyl acetate mixture, the product rectifying tower includes 2# rectifying tower (10) and 3# rectifying tower (11), the tower bottom liquid phase material of 1# rectifying tower (9) gets into 2# rectifying tower (10), the top of 2# rectifying tower (10) is extracted and is contained methyl alcohol product (O2), and tower bottom liquid phase material gets into 3# rectifying tower (11), anhydrous alcohol product (O3) are extracted to the top of 3# rectifying tower (11), and the tower bottom liquid is sent out the boundary region.
10. The process of hydrogenation of acetic ester to ethanol as claimed in claim 9, wherein the step of hydrogenation comprises the steps of,
the operation pressure of the top of the 1# rectifying tower (9) is 0-500 kPaG, the reflux ratio is 0.01-100, and the content of acetic ester in a tower bottom liquid phase is less than or equal to 5 ppm-wt;
the operation pressure of the top of the 2# rectifying tower (10) is 0-500 kPaG, the reflux ratio is 0.01-100, and the purity of the methanol product (O2) is more than or equal to 99.8 wt%;
the operation pressure of the top of the 3# rectifying tower (11) is 0-500 kPaG, the reflux ratio is 0.01-100, and an absolute ethyl alcohol product (O3) is extracted from the liquid phase at the top of the tower, and the purity is more than or equal to 99.7 wt.%.
11. The system for applying the process method for preparing the ethanol by hydrogenating the acetic ester according to claim 1, which comprises a gas phase reaction device and a product separation device;
acetic ester and hydrogen in the gas phase reaction device are subjected to gas phase hydrogenation reaction, and liquid phase products generated after condensation and gas-liquid separation of reaction products are sent to the product separation device;
the product separation device comprises a rough separation tower and a product rectifying tower, wherein a liquid phase product containing acetic ester and carbonyl byproducts is obtained at the top of the rough separation tower, and an absolute ethyl alcohol product (O3) is obtained at the top of the product rectifying tower;
it is characterized in that;
the liquid phase reaction device performs liquid phase hydrogenation reaction on the hydrogen or fresh hydrogen from the gas phase reaction device and the liquid phase product containing the acetic ester and the carbonyl byproducts to obtain a liquid phase hydrogenation reaction product, and the liquid phase product obtained after the liquid phase hydrogenation reaction product is condensed and separated is sent to the product separation device.
12. The process system for preparing ethanol by acetate hydrogenation according to claim 11,
the liquid phase reaction device comprises a liquid phase hydrogenation reactor (8), a feeding and discharging heat exchanger b (12), a cooler b (13) and a gas-liquid separator b (14); the cold material inlet of the feeding and discharging heat exchanger b (12) is respectively connected with the gas phase reaction device and the product separation device, the cold material outlet of the feeding and discharging heat exchanger b (12) is connected with the inlet of the liquid phase hydrogenation reactor (8), the outlet of the liquid phase hydrogenation reactor (8) is connected with the hot material inlet of the feeding and discharging heat exchanger b (12), and the hot material outlet of the feeding and discharging heat exchanger b (12) is sequentially connected with the cooler b (13) and the gas-liquid separator b (14).
13. The process system for preparing ethanol by acetate hydrogenation according to claim 11,
the gas-phase reaction device comprises a vaporizer (1), a superheater (2), a gas-phase hydrogenation reactor (3), a feeding and discharging heat exchanger a (4), a cooler a (5), a gas-liquid separator a (6) and a compressor (7); the acetate raw material (S1) enters from a first raw material inlet of a vaporizer, the hydrogen raw material enters from a second raw material inlet of the vaporizer, and an outlet of the vaporizer and the superheater (2) are sequentially connected with the reactor (3) through pipelines; the feeding and discharging heat exchanger a (4) heats hydrogen raw materials by using materials at the outlet of the reactor (3), the outlet of the reactor (3) is connected with the hot material inlet of the feeding and discharging heat exchanger a (4), the hot material outlet of the feeding and discharging heat exchanger a (4) is connected with the inlet of the cooler a (5), the outlet of the cooler a (5) is connected with the inlet of the gas-liquid separator a (6), and the gas-phase outlet of the gas-liquid separator a (6) is respectively connected with the inlet of the compressor (7) and the liquid-phase reaction device; the liquid phase outlet of the gas-liquid separator a (6) is connected with the product separation device; the outlet of the compressor (7) is connected with the cold material inlet of the feeding and discharging heat exchanger a (4), and the cold material outlet of the feeding and discharging heat exchanger a (4) is connected with the second raw material inlet of the vaporizer (1).
14. The process system for preparing ethanol by hydrogenating acetic ester according to claim 11, wherein the product separation device comprises a rough separation tower and a product separation tower, the rough separation tower comprises a # 1 rectifying tower (9), a liquid-phase product outlet of an overhead condenser of the # 1 rectifying tower (9) is connected with the liquid-phase reaction device, and a liquid-phase product outlet of a tower kettle of the # 1 rectifying tower (9) is connected with the product separation tower;
the product separation tower is provided with a product separation tower according to the composition of the acetic ester;
when the acetic ester raw material component does not contain methyl acetate, the product separation tower comprises a 3# rectifying tower (11), a tower kettle liquid phase outlet of the 1# rectifying tower (9) is connected with an inlet of the 3# rectifying tower (11), and anhydrous ethanol (O3) is extracted from the tower top of the 3# rectifying tower (11);
when acetate raw materials component contains methyl acetate, the product knockout tower includes 2# rectifying column (10) and 3# rectifying column (11), the tower cauldron liquid phase export of 1# rectifying column (9) with the import of 2# rectifying column (10) links to each other, methanol product (O2) are extracted to the top of 2# rectifying column (10), tower cauldron liquid phase export with the import of 3# rectifying column (11) links to each other, ethanol product (O3) are extracted to the top of 3# rectifying column (11).
15. The process system for preparing ethanol by hydrogenating acetic ester according to claim 12, wherein a cold material inlet of the feed and discharge heat exchanger b (12) is respectively connected with a gas phase outlet of a gas-liquid separator a (6) in the gas phase reaction device and a liquid phase product outlet of an overhead condenser of a # 1 rectifying tower (9) in the product separation device.
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