CN113968829B

CN113968829B - Propylene epoxidation method and epoxidation system

Info

Publication number: CN113968829B
Application number: CN202010725355.2A
Authority: CN
Inventors: 李红波; 王皓; 王瑾; 丁晖殿; 林民; 罗一斌; 朱斌
Original assignee: Sinopec Research Institute of Petroleum Processing; China Petroleum and Chemical Corp
Current assignee: Sinopec Research Institute of Petroleum Processing; China Petroleum and Chemical Corp
Priority date: 2020-07-24
Filing date: 2020-07-24
Publication date: 2024-02-09
Anticipated expiration: 2040-07-24
Also published as: CN113968829A

Abstract

The invention discloses a propylene epoxidation method and an epoxidation system, wherein the epoxidation method comprises an epoxidation reaction process, an epoxidation reaction product separation process and a solvent recovery and circulation process, the solvent recovery and circulation process comprises a hydrogenation step, a methanol refining step, an ethanol removal step and a methanol circulation step, and at least part of ethanol in a methanol refined stream is removed in the ethanol removal step to obtain recovered methanol. According to the propylene epoxidation method and the epoxidation system, before the methanol which is refined and recovered by the material flow obtained by hydrogenation reaction is recycled to the epoxidation reaction, the operation of removing the ethanol is added, at least part of the ethanol in the recovered methanol is removed, the problem of reduction of the selectivity of the epoxidation reaction product can be effectively inhibited in the long-time continuous reaction process, the selectivity of the epoxidation reaction product is kept at a higher level for a longer time, and the quality stability of the propylene oxide product is kept.

Description

Propylene epoxidation method and epoxidation system

Technical Field

The present invention relates to a propylene epoxidation process and an epoxidation system.

Background

Propylene Oxide (PO) is the third largest Propylene-derived organic compound raw material next to polypropylene, acrylonitrile, and is mainly used for producing polyether, propylene glycol, and the like. It is also the main raw material of fourth generation detergent nonionic surfactant, oil field demulsifier and pesticide emulsifier. Propylene oxide derivatives are widely used in the industries of automobiles, buildings, foods, tobacco, medicines, cosmetics and the like.

The propylene oxide production process mainly comprises a chlorohydrin method, a co-oxidation method (also called an indirect oxidation method) and a direct oxidation method. The chlorohydrin method has long production history, and has the advantages of mature process, high operation elasticity, good selectivity, low requirement on the purity of raw material propylene, low fixed investment and the like; however, the chlorohydrin process produces a large amount of wastewater and waste residues. Every 1 ton of propylene oxide is produced, 40-50 tons of chlorine-containing wastewater and 2 tons of waste residues are produced, and hypochlorous acid produced in the production process has serious corrosion to equipment. The co-oxidation method is mainly an ethylbenzene co-oxidation method and an isobutane co-oxidation method, overcomes the defects of corrosion equipment and more sewage in the chlorohydrin method, has little pollution to the environment and has lower cost; the defects are long process flow, multiple raw material varieties, high propylene purity requirement, large investment, and the like, and the co-production products must be considered.

The hydrogen peroxide direct oxidation (HPPO) is prepared by directly reacting hydrogen peroxide and propylene to only generate propylene oxide and water, has the advantages of simple process flow, high product yield, no co-production product, basically no pollution and environmental friendliness, and is considered as a development trend of propylene oxide synthesis technology.

CN1649858A discloses a process for the epoxidation of olefins by:

i) Reacting an olefin with hydrogen peroxide in the presence of an epoxidation catalyst and an alcohol solvent;

ii) separating product olefin oxide and unreacted olefin from the reaction product of step i);

iii) Recovering a stream comprising an alcohol solvent;

iv) subjecting the stream recovered in step iii) to a hydrogenation reaction.

In this process, the alcohol solvent recovered by further distillation of the stream obtained by the hydrogenation reaction of step iv) can be recycled for use in step i).

Although the hydrogenation of the recovered stream containing the alcohol solvent can remove impurities that deactivate the epoxidation catalyst, there is still the problem of reduced selectivity of the product of the epoxidation reaction over the course of the reaction time during actual operation.

Disclosure of Invention

The invention aims to solve the problem that the selectivity of the products of the epoxidation reaction is reduced in the long-time continuous operation process of an epoxidation reaction system.

According to a first aspect of the present invention, there is provided a propylene epoxidation process comprising an epoxidation reaction process, an epoxidation reaction product separation process and a solvent recovery recycle process,

in the epoxidation reaction process, propylene and hydrogen peroxide are contacted with an epoxidation reaction catalyst in the presence of methanol and water under epoxidation reaction conditions to obtain an epoxidation reaction product stream;

in an epoxidation reaction product separation process, separating the epoxidation reaction product stream to obtain a propylene oxide product and a methanol stream comprising methanol and water;

the solvent recovery and recycling process comprises a hydrogenation step, a methanol refining step, an ethanol removing step and a methanol recycling step,

in the hydrogenation step, at least part of methanol stream and hydrogen are contacted with a catalyst with hydrogenation catalysis to carry out hydrogenation treatment under hydrogenation reaction conditions to obtain a hydrogenation product stream, the hydrogenation product stream is separated to obtain a gas-phase hydrogenation stream containing hydrogen and a liquid-phase hydrogenation stream containing methanol,

in the methanol refining step, the liquid-phase hydrogenation material flow is separated to obtain a methanol refined material flow, or the liquid-phase hydrogenation material flow and the rest methanol material flow are separated to obtain a methanol refined material flow,

In the ethanol removal step, removing at least a portion of the ethanol in at least a portion of the methanol refinery stream to obtain a purified methanol refinery stream;

in the methanol recycling step, at least a portion of the recovered methanol, which is the purified methanol refinery stream or the purified methanol refinery stream and the remaining portion of the methanol refinery stream, is recycled to the epoxidation reaction process.

According to a second aspect of the present invention, there is provided an epoxidation system comprising an epoxidation reaction unit, an epoxidation reaction product separation unit, and a solvent recovery circulation unit, wherein the epoxidation product stream output port of the epoxidation reaction unit is in communication with the feed input port of the epoxidation reaction product separation unit to be separated, the feed input port of the solvent recovery circulation unit is in communication with the recovered methanol output port of the epoxidation reaction product separation unit, the solvent recovery circulation unit comprises a hydrogenation subunit, a methanol refining subunit and a methanol circulation subunit, wherein the solvent recovery circulation unit further comprises an ethanol removal subunit for removing at least part of the ethanol in the methanol refining stream output by the methanol refining subunit to obtain recovered methanol, and feeding the recovered methanol into the epoxidation reaction unit through the methanol circulation subunit.

According to the propylene epoxidation method and the epoxidation system, before methanol is recycled to an epoxidation reaction, the operation of removing ethanol is added, at least part of ethanol in the recycled methanol is removed, the problem of reduction of the selectivity of an epoxidation reaction product can be effectively inhibited in a long-time continuous reaction process, the selectivity of the epoxidation reaction product is kept at a higher level for a longer time, and the quality stability of a propylene oxide product is maintained. The reason why the method and system of the present invention can effectively suppress the problem of the decrease in the selectivity of the epoxidation reaction product during the continuous reaction for a long period of time may be that: the recovered methanol-containing stream is subjected to hydrotreating, and impurities such as acetaldehyde in the methanol-containing stream are converted to ethanol by hydrotreating, and during subsequent recovery of methanol by distillation, a portion of the ethanol remains in the recovered methanol and is circulated with the circulating methanol in the epoxidation reaction system, resulting in a gradual increase in the ethanol content in the circulating methanol. Propylene is dispersed in methanol as a solvent in an epoxidation reaction system to contact and react with an epoxidation reaction catalyst to form propylene oxide, the epoxidation reaction system is a complex system comprising a plurality of substances with different physicochemical properties such as methanol, water, propylene oxide and the like, and the reaction system is required to be kept uniform and stable, however, accumulated ethanol has a tendency to cause the stability of the epoxidation reaction system to be reduced, and further the selectivity of products of the epoxidation reaction to be reduced.

Detailed Description

The endpoints and any values of the ranges disclosed herein are not limited to the precise range or value, and are understood to encompass values approaching those ranges or values. For numerical ranges, one or more new numerical ranges may be found between the endpoints of each range, between the endpoint of each range and the individual point value, and between the individual point value, in combination with each other, and are to be considered as specifically disclosed herein.

in the ethanol removal step, removing at least a portion of the ethanol in at least a portion of the methanol refinery stream to obtain a purified methanol refinery stream,

In the epoxidation reaction step, the epoxidation catalyst is preferably a titanium-containing molecular sieve, more preferably a titanium silicalite molecular sieve. The titanium-silicon molecular sieve is a generic term for a class of zeolite in which titanium atoms replace a portion of silicon atoms in the lattice framework, and can be represented by the chemical formula xTiO ₂ ·SiO ₂ And (3) representing. The content of titanium atoms in the titanium-silicon molecular sieve is not particularly limited in the present invention, and may be selected conventionally in the art. Specifically, x may be 0.0001 to 0.05, preferably 0.01 to 0.03, more preferably 0.015 to 0.025.

The titanium silicalite molecular sieve may be a conventional titanium silicalite molecular sieve having various topologies, such as: the titanium silicalite molecular sieve may be selected from titanium silicalite molecular sieves with MFI structure (such as TS-1), MEL structure (such as TS-2), BEA structure (such as Ti-Beta), MWW structure (such as Ti-MCM-22), MOR structure (such as Ti-MOR), TUN structure (such as Ti-TUN), two-dimensional hexagonal structure (such as Ti-MCM-41, ti-SBA-15), and other structure (such as Ti-ZSM-48). The titanium-containing molecular sieve is preferably selected from titanium silicalite molecular sieves of MFI structure.

In a preferred embodiment, the titanium-containing molecular sieve is a hollow titanium silicalite molecular sieve, the grains of the hollow titanium silicalite molecular sieve are hollow structures, the hollow nodesThe radial length of the cavity part is 5-300nm, and the titanium silicalite molecular sieve has a temperature of 25 ℃ and a P/P ratio ₀ The benzene adsorption amount measured under the conditions of=0.10 and adsorption time of 1h is at least 70mg/g, and a hysteresis loop exists between an adsorption isotherm and a desorption isotherm of low-temperature nitrogen adsorption of the titanium silicalite molecular sieve.

The titanium-containing molecular sieve can be titanium-containing molecular sieve raw powder, or can be formed titanium-containing molecular sieve, preferably formed titanium-containing molecular sieve.

In the epoxidation reaction step, the hydrogen peroxide is supplied in the form of an aqueous solution, preferably an aqueous hydrogen peroxide solution having a hydrogen peroxide concentration of 40 to 80% by weight, more preferably an aqueous hydrogen peroxide solution having a hydrogen peroxide concentration of 45 to 65% by weight.

In the epoxidation reaction step, propylene is preferably used in excess of hydrogen peroxide. Specifically, the molar ratio of propylene to hydrogen peroxide may be 1.2-10:1, preferably 1.2-5:1, more preferably 1.5-4:1, further preferably 2-3:1. in the epoxidation reaction step, the molar ratio of methanol to hydrogen peroxide is preferably 4 to 20:1, more preferably from 6 to 12:1, further preferably 8 to 10:1. in a preferred embodiment, methanol: propylene: the molar ratio of the hydrogen peroxide is 4-20:1.2-10:1. in a more preferred embodiment, methanol: propylene: the molar ratio of the hydrogen peroxide is 6-12:1.2-5:1, preferably 8-10:1.5-4:1.

in the epoxidation reaction step, propylene and hydrogen peroxide are preferably contacted with an epoxidation reaction catalyst in the presence of at least one basic substance in methanol and water to further increase the product selectivity of the epoxidation reaction. Specific examples of the alkaline substance may include, but are not limited to: ammonia (i.e. NH) ₃ ) Amine, quaternary ammonium base and M ¹ (OH) _n (wherein M ¹ Is an alkali or alkaline earth metal such as sodium, potassium, magnesium or calcium; n is M and ¹ an integer having the same valence). The alkaline material is preferably used in an amount such that the liquid mixture in contact with the epoxidation catalyst has a pH of from 6.5 to 9.

In the epoxidation reaction step, the epoxidation reaction is preferably carried out in a fixed bed reactor, the epoxidation reaction catalyst is filled in a catalyst bed of the fixed bed reactor, a feed stream containing propylene, hydrogen peroxide, methanol and water flows through the catalyst bed, and the feed stream is contacted with the epoxidation reaction catalyst to carry out the epoxidation reaction, thereby obtaining an epoxidation reaction product stream containing propylene oxide, methanol, water and unreacted propylene. The feed stream may flow through the catalyst bed from top to bottom or from bottom to top. Preferably, the feed stream flows through the catalyst bed from bottom to top, for example: the feed stream may enter the fixed bed reactor from the bottom of the fixed bed reactor, flow through a catalyst bed, and recover an epoxidation reaction product stream from the top of the fixed bed reactor. The fixed bed reactor is preferably a tubular fixed bed reactor in which the ratio of the inner diameter of the tube array to the length of the tube array (simply referred to as "aspect ratio") is preferably 50 to 500, more preferably 100 to 250, still more preferably 150 to 200. The number of fixed bed reactors may be one or more than two, for example 2 to 10 fixed bed reactors. In a preferred embodiment, the number of fixed bed reactors is more than two fixed bed reactors connected in series. In the preferred embodiment, propylene and methanol are preferably fed into the first fixed bed reactor, and hydrogen peroxide may be fed into the first fixed bed reactor in its entirety, or may be divided into n parts, and fed into the first fixed bed reactor and n reactors downstream of the first fixed bed reactor, the number of fixed bed reactors being m, n being an integer located in the interval [2, m ].

In the epoxidation reaction step, the epoxidation reaction may be carried out at a temperature of 20 to 80 ℃, preferably at a temperature of 30 to 60 ℃, more preferably at a temperature of 40 to 50 ℃.

According to the process of the present invention, in an epoxidation reaction product separation process, the epoxidation reaction product stream is separated to yield a propylene oxide product, and a methanol stream comprising methanol and water.

In one embodiment, the epoxidation reaction product stream may be rectified in a rectification column, with a column overhead containing unreacted propylene being obtained from the top of the rectification column, a column bottoms stream containing propylene oxide, alcohol and water being obtained from the bottom of the rectification column, which column bottoms stream may be subjected to extractive distillation in an extractive distillation column, with a column overhead containing propylene oxide being obtained from the top of the extractive distillation column, and a column bottoms stream containing alcohol and water being obtained from the bottom of the extractive distillation column.

In a preferred embodiment, the epoxidation reaction product stream is subjected to a first separation to obtain a light stream comprising propylene oxide, propylene and a portion of methanol and a heavy stream comprising water and a remaining portion of methanol, and the light stream is subjected to a second separation to obtain a gas phase stream comprising propylene, a propylene oxide product and a methanol liquid phase stream comprising methanol.

In this preferred embodiment, the epoxidation reaction product may be subjected to a first separation in a first distillation column to yield a light stream and a heavy stream. The top pressure of the first distillation column may be 0.01 to 0.5MPa, preferably 0.05 to 0.2MPa, the pressure being gauge pressure. The overhead temperature of the first distillation column may be from 60 to 110 ℃, preferably from 65 to 90 ℃, more preferably from 65 to 80 ℃. The theoretical plate number of the first distillation column may be 10 to 50, preferably 15 to 45, more preferably 20 to 40. In the invention, the theoretical plate numbers are all theoretical plate numbers counted from top to bottom.

The light stream may be subjected to a second separation in a second distillation column to obtain a vapor purge stream comprising propylene and a propylene oxide stream comprising propylene oxide and methanol, and the vapor purge stream is contacted with an absorbent to obtain a vapor stream comprising propylene and a propylene oxide absorbent stream comprising absorbent and propylene oxide.

The second separation is based on being able to substantially separate propylene from the light stream. Preferably, the propylene oxide stream containing propylene oxide and methanol obtained by the second separation of the light stream has a propylene content of 0.1 wt% or less. The top pressure of the second distillation column may be 0.01 to 0.5MPa, preferably 0.05 to 0.2MPa, the pressure being gauge pressure. The top temperature of the second distillation column may be from 35 to 80 ℃, preferably from 35 to 60 ℃, more preferably from 35 to 50 ℃. The theoretical plate number of the second distillation column is preferably 20 to 40, more preferably 25 to 35.

The absorbent may be a liquid material sufficient to absorb propylene oxide, such as C ₁ -C ₅ Is an alcohol of (a) a (c). Preferably, the absorbent is methanol. Preferably, the propylene oxide absorbent stream containing absorbent and propylene oxide is fed to the first distillation column for distillation to further enhance propylene oxide recovery. The weight ratio of the absorbent to the gas phase purge stream may be from 0.8 to 3:1, preferably 1-2.5:1, more preferably 1.2-2:1. the temperature within the absorber may be 20-40 ℃. The pressure in the absorber can be 0.01-0.1MPa, preferably 0.02-0.05MPa, the pressure being gauge pressure.

The propylene oxide stream containing propylene oxide and methanol obtained from the bottom of the second distillation column may be separated by extractive distillation to obtain a propylene oxide product and a methanol liquid phase stream containing methanol. Specifically, a propylene oxide stream containing propylene oxide and methanol may be contacted with an extractant in an extractive distillation zone of an extractive distillation column to carry out extractive distillation, thereby obtaining a propylene oxide product, and a liquid phase stream containing methanol and extractant is recovered from the bottom of the extractive distillation column.

The extractant may be a polar extractant such as: one or more of water, propylene glycol and tert-butanol. Preferably, the extractant is water. The extractant may be added to the extractive distillation zone from a location above the propylene oxide stream. Preferably, the theoretical plate number of the extractive distillation column is T, the feeding position of the propylene oxide stream is preferably (0.6 to 0.9) ×t, more preferably (0.6 to 0.85) ×t, still more preferably (0.7 to 0.8) ×t, and the feeding position of the extractant is preferably (0.15 to 0.45) ×t, more preferably (0.2 to 0.4) ×t, still more preferably (0.25 to 0.35) ×t. The theoretical plate number T of the extractive distillation column is preferably 35 to 90, more preferably 45 to 85, still more preferably 55 to 75. In the present invention, the theoretical plate number is the theoretical plate number counted from top to bottom.

The pressure at the top of the extractive distillation column may be 0.01 to 0.5MPa, preferably 0.02 to 0.3MPa, more preferably 0.03 to 0.1MPa; the overhead temperature may be from 30 to 45 ℃, preferably from 35 to 45 ℃; the bottom temperature may be 70-110 ℃, preferably 75-95 ℃; the reflux ratio may be from 1 to 10, preferably from 1.5 to 8, more preferably from 2 to 6, the overhead pressure being the gauge pressure.

From the viewpoint of further reducing the content of impurities (particularly aldehydes) in the propylene oxide product obtained by extractive distillation, the extractive distillation is preferably carried out in the presence of at least one amino group-containing compound, preferably hydrazine. At least part of the amino-containing compound may be added to the extractive distillation zone of the extractive distillation column at a position not lower than the extractant, at least part of the amino-containing compound may be added to the extractive distillation zone of the extractive distillation column at the same position as the propylene oxide stream, or a combination of both, i.e., part of the amino-containing compound may be added to the extractive distillation zone of the extractive distillation column at a position not lower than the extractant, and the remaining part of the amino-containing compound may be added to the extractive distillation zone of the extractive distillation column at the same position as the propylene oxide stream. When at least a part of the amino-containing compound is added to the extractive distillation zone of the extractive distillation column at a position not lower than the extractant, at least a part of the amino-containing compound may be added to the extractive distillation zone of the extractive distillation column at the same position as the extractant, or at least a part of the amino-containing compound may be added to the extractive distillation zone of the extractive distillation column at a position higher than the extractant. The molar ratio of the amino-containing compound to the aldehyde in the propylene oxide stream may be from 1 to 4:1, preferably 1.2-2.5:1. in a preferred embodiment, the amino-containing compound is hydrazine, the amino-containing compound is added to the extractive distillation zone of the extractive distillation column at the same location as the extractant, preferably the amino-containing compound is added to the extractant, and the extractant containing the amino-containing compound is added to the extractive distillation zone of the extractive distillation column. In this preferred embodiment, the amino group-containing compound may be contained in an amount of 0.3 to 0.8% by weight based on the total amount of the extractant.

The propylene oxide stream is preferably pretreated upstream of the extractive distillation column, based on the flow direction of the propylene oxide stream, to reduce the amount of esters (e.g., methyl formate) in the propylene oxide stream by one or both of the following methods:

in a first mode, contacting the propylene oxide stream with at least one basic substance, preferably a basic ion exchange resin and/or an aqueous solution comprising a water-soluble basic compound;

and in a second mode, the propylene oxide stream is rectified in a light component removal rectifying tower, the rectifying condition is sufficient to remove at least part of ester and/or aldehyde in the liquid phase stream, the tower top pressure of the light component removal rectifying tower is preferably 0.01-0.5MPa, the tower top temperature is preferably 30-45 ℃, the reflux ratio is preferably 100-300, the tower top pressure is gauge pressure, and the theoretical plate number of the light component removal rectifying tower is preferably 30-70.

In the case of the basic ion exchange resin, the propylene oxide stream may be pretreated by passing the propylene oxide stream through a bed of ion exchange resin packed with the basic ion exchange resin. The basic ion exchange resin may be a strong basic ion exchange resin and/or a weak basic ion exchange resin, and the basic ion exchange resin may be one or a combination of two or more of a styrene ion exchange resin, a phenolic ion exchange resin and an acrylic ion exchange resin. The temperature in the ion exchange resin bed may be in the range 35 to 80 ℃, preferably 45 to 75 ℃, more preferably 50 to 70 ℃.

In the first mode, when an aqueous solution containing a water-soluble basic compound, which may be ammonia (NH), is used, an aqueous solution containing a water-soluble basic compound may be added to the propylene oxide stream ₃ ) Containing amino groups (-NH) ₂ ) One or more of water-soluble substances (such as hydrazine), alkali metal hydroxides (such as sodium hydroxide and/or potassium hydroxide), alkali metal carbonates (such as sodium carbonate and/or potassium carbonate), alkali metal bicarbonates (such as sodium bicarbonate and/or potassium bicarbonate), and alkaline earth metal hydroxides (such as magnesium hydroxide). When an aqueous solution containing a water-soluble basic compound is used, the waterThe molar ratio of soluble basic compound to ester in the propylene oxide stream may be from 1 to 4:1, preferably 1.2-2.5:1. when an aqueous solution containing a water-soluble basic compound is employed, the propylene oxide stream may be mixed with the aqueous solution containing a water-soluble basic compound in a line mixer.

The liquid phase stream containing methanol and extractant obtained from the bottom of the extractive distillation column may be directly fed to a solvent recovery cycle for processing to recover solvent methanol, for example: directly enters the hydrogenation step and/or the methanol refining step.

The inventors of the present invention found during the course of the study that a liquid phase stream containing methanol and an extractant obtained from the bottom of an extractive distillation column contains intermediate impurities having a boiling point higher than propylene oxide and lower than methanol, such as: dimethoxyethane and acetone, the liquid phase stream containing methanol and extractant is directly fed to a solvent recovery cycle for treatment, and the recovered methanol typically also contains some amount of intermediate impurities. According to the epoxidation process of the present invention, in a preferred embodiment, at least part of the liquid phase stream comprising methanol and extractant is subjected to a third separation in a third distillation column, a distillate comprising intermediate impurities is obtained from the top of the third distillation column, a purified liquid phase stream is obtained from the bottom of the third distillation column, the intermediate impurity content of the purified liquid phase stream is lower than the intermediate impurity content of the liquid phase stream, and the purified liquid phase stream is fed to a solvent recovery cycle for treatment. According to this preferred embodiment, the content of intermediate impurities in the recovered methanol can be reduced, and the product selectivity of the epoxidation reaction can be further improved. According to this preferred embodiment, the third distillation column is operated under conditions effective to remove at least part of the intermediate impurities in the liquid phase stream comprising methanol and extractant, preferably such that the intermediate impurities in the recovered methanol obtained in the solvent recovery cycle step are present in an amount of less than 0.4 wt.%, preferably not more than 0.2 wt.%, more preferably not more than 0.1 wt.%, even more preferably not more than 0.05 wt.%, particularly preferably not more than 0.04 wt.%, such as not more than 0.03 wt.%, based on the total amount of recovered methanol. The top pressure of the third distillation column is preferably 0.01 to 0.5MPa, more preferably 0.02 to 0.3MPa, still more preferably 0.03 to 0.1MPa, and the top pressure is gauge pressure. The top temperature of the third distillation column is preferably 50-75 ℃. The reflux ratio of the third distillation column is preferably 50 to 300, more preferably 60 to 250, still more preferably 80 to 200, still more preferably 100 to 150. In this preferred embodiment, the theoretical plate number of the third distillation column is preferably 30 to 70, more preferably 40 to 60, still more preferably 45 to 55. The ratio of the theoretical plate number corresponding to the feed position of the liquid phase stream containing methanol and extractant to the theoretical plate number of the third distillation column is preferably 0.3 to 0.7, more preferably 0.35 to 0.6, still more preferably 0.4 to 0.55.

According to the epoxidation process of the present invention, the methanol stream is treated in a solvent recovery recycle process to yield recovered methanol. The methanol stream is a stream containing methanol and water separated in an epoxidation reaction product separation process, and may be, for example, a heavy stream obtained from the first separation and a liquid stream obtained from the bottom of an extractive distillation column. Preferably, the methanol stream is a heavy stream from the first separation, a purified liquid stream from the third separation, and an extractive distillation column bottoms liquid stream without the third separation. When all liquid-phase material flows obtained at the bottom of the extraction rectifying tower are sent into a third distilling tower to carry out third separation, the methanol material flows are heavy material flows obtained by the first separation and purified liquid-phase material flows obtained by the third separation; and when part of the liquid-phase material flow obtained at the bottom of the extractive distillation tower is sent to the third distillation tower for third separation, the methanol material flow is a heavy material flow obtained by the first separation, a purified liquid-phase material flow obtained by the third separation and a liquid-phase material flow obtained at the bottom of the residual part of the extractive distillation tower.

The solvent recovery and recycling process comprises a hydrogenation step, a methanol refining step, an ethanol removal step and a methanol recycling step. The whole methanol stream may be fed to a hydrogenation step for hydrotreating and then treated in a methanol refining step and an ethanol removal step in sequence to obtain recovered methanol. In a preferred embodiment, the heavy stream from the first separation is fed to a hydrogenation step for hydrotreating to obtain a liquid phase hydrotreated stream comprising methanol, which liquid phase hydrotreated stream and the remaining part of the methanol stream are separated in the methanol refining step. According to this embodiment, the amount of hydrotreatment can be effectively reduced, and the recovered methanol meeting the requirements can be obtained, thereby improving the economical efficiency of the epoxidation process of the present invention.

In the hydrogenation step of the solvent recovery cycle process, the catalyst having a hydrogenation catalytic action may be a catalytic substance sufficient to react impurities in methanol with hydrogen.

In one embodiment, the catalyst having hydrogenation catalysis contains at least one catalytically active component, which may be selected from the group VIII metals and group IB metals, preferably one or more of ruthenium, rhodium, palladium, platinum, silver, iridium, iron, copper, nickel and cobalt, more preferably nickel. The catalyst having hydrogenation catalytic action further comprises a carrier for supporting the catalytically active component, and the carrier may be a porous heat-resistant inorganic oxide, preferably one or more of silica, titania, zirconia and alumina, more preferably alumina. The content of the catalytically active component in elemental form may be from 2 to 70% by weight, preferably from 20 to 60% by weight, based on the total amount of the catalyst having hydrogenation catalysis. In a preferred embodiment, the catalytically active component of the catalyst having hydrogenation catalysis is nickel, the support is alumina, and the content of the catalytically active component in elemental form is from 30 to 55% by weight, more preferably from 35 to 45% by weight, based on the total amount of the catalyst having hydrogenation catalysis.

In the hydrogenation step, impurities in methanol are contacted with hydrogen and a catalyst having hydrogenation catalytic action under hydrogenation reaction conditions to convert the impurities into substances which are easier to separate from the alcohol. The hydrogenation conditions may be selected according to the kind of impurities in the alcohol. In the hydrogenation step, the temperature of the contacting is preferably 50 to 175 ℃, more preferably 60 to 145 ℃, still more preferably 70 to 125 ℃, still more preferably 80 to 115 ℃, for example, may be 85 to 110 ℃. In the hydrogenation step, the contacting is performed under a pressure of preferably 0.5 to 10MPa, more preferably 1 to 6MPa, still more preferably 2 to 5.5MPa, still more preferably 3 to 5MPa, the pressure being a gauge pressure.

In the hydrogenation step, the hydrotreating may be carried out in a conventional hydrogenation reactor, for example: one or a combination of more than two of a fixed bed reactor, a slurry bed reactor and a fluidized bed reactor.

In a preferred embodiment, the hydrotreating is carried out in a fixed bed reactor, with the catalyst having a hydrogenation catalytic effect being packed in the fixed bed reactor to form a catalyst bed through which at least a portion of the methanol stream and hydrogen containing impurities are passed, the impurities and hydrogen being contacted with the catalyst having a hydrogenation catalytic effect and subjected to a hydrogenation reaction to convert the impurities into species that are more readily separated from the methanol. When contacting at least a portion of the methanol stream and hydrogen containing impurities with a catalyst having a hydrogenation catalytic effect in a fixed bed reactor, the liquid hourly space velocity (i.e., liquid volumetric flow/catalyst volume) may be in the range of 0.5 to 30 hours ^-1 Preferably 2 to 25 hours ^-1 More preferably 3 to 20 hours ^-1 . In a preferred embodiment, the liquid hourly space velocity is from 5 to 25 hours ^-1 Preferably 8-20h ^-1 More preferably 10-15h ^-1 . According to this preferred embodiment, the efficiency of hydrotreating can be effectively improved. In the case of hydrotreating in a fixed bed reactor, the methanol stream may be passed through the catalyst bed from bottom to top, or from top to bottom, preferably from top to bottom. The hydrogen and the methanol stream may be fed co-currently or counter-currently, preferably co-currently.

The hydrogenation product stream obtained in the hydrogenation step contains methanol and hydrogen, and the hydrogenation product stream can be separated by a conventional method, thereby obtaining a gas-phase hydrogenation stream containing hydrogen and a liquid-phase hydrogenation stream containing methanol.

In a preferred embodiment, the separation process of the hydrogenation product stream comprises a first gas-liquid separation step, an absorption step and an optional second gas-liquid separation step,

in the first gas-liquid separation step, carrying out gas-liquid separation on the hydrogenation product stream to obtain a first gas-phase hydrogenation stream and a first liquid-phase hydrogenation stream;

In the absorption step, the first gas phase hydrogenation stream is contacted with a liquid absorbent to obtain a second gas phase hydrogenation stream containing hydrogen and a second liquid phase hydrogenation stream containing absorbent,

in the second gas-liquid separation step, the second gas-phase hydrogenation material flow is subjected to second gas-liquid separation to obtain a third gas-phase hydrogenation material flow containing hydrogen and a third liquid-phase hydrogenation material flow,

the second vapor hydrogenation stream, the third vapor hydrogenation stream, or the second and third vapor hydrogenation streams are the vapor hydrogenation stream. The first liquid phase hydrogenation stream, the second liquid phase hydrogenation stream, and optionally the third liquid phase hydrogenation stream are liquid phase hydrogenation streams comprising methanol.

In the first gas-liquid separation step, the hydrogenation product stream is separated into a first gas phase hydrogenation stream comprising predominantly hydrogen and a first liquid phase hydrogenation stream comprising predominantly alcohol. In one example, the hydrogenation product stream is fed to a high pressure gas-liquid separation tank for gas-liquid two-phase separation. In the first gas-liquid separation step, the temperature of separation may be 80 to 135 ℃, preferably 85 to 130 ℃. In the first gas-liquid separation step, the separation may be carried out under a pressure of 0.5 to 6MPa, preferably under a pressure of 1 to 5MPa, which is a gauge pressure.

In the absorption step, the first gas phase hydrogen stream is contacted with a liquid absorbent, and the gas in the first gas phase hydrogen stream is separated from the soluble material entrained in the first gas phase hydrogen stream. The soluble material entrained in the first gas phase hydrogenation stream is predominantly methanol and the liquid absorbent may be an absorbent capable of absorbing the soluble material entrained in the gas phase hydrogenation stream. Preferably, the liquid absorbent is water. The first gas phase vapor stream may be contacted with the liquid absorbent at a temperature of from 20 to 60 c, preferably from 30 to 55 c, more preferably from 40 to 50 c. The absorption can be carried out in customary absorption apparatus. In a preferred example, a packed column is used as the absorption column. In this preferred example, the liquid absorbent may be fed from an upper portion of the absorber, the first gas phase hydrogenation stream may be fed from a location below the liquid absorbent, and the first gas phase hydrogenation stream and the liquid absorbent are countercurrently contacted in the absorber to effect separation.

The second gaseous hydrogenation stream separated in the absorption step may be output and optionally mixed with fresh hydrogen as said gaseous hydrogenation stream. In a preferred embodiment, at least a portion of the second vapor phase hydrogenation stream is fed to a second vapor-liquid step for further vapor-liquid separation to further separate out condensable materials (e.g. methanol and liquid absorbent) present in the second vapor phase hydrogenation stream. When a second gas-liquid separation step is included, fresh hydrogen may be fed to the second gas-liquid separation step along with at least a portion of the second vapor phase hydrogenation stream for separation. The second gas-liquid separation may be carried out at a temperature of 20-60 ℃, preferably at a temperature of 30-55 ℃, more preferably at a temperature of 40-50 ℃. The second gas-liquid separation may be carried out at a pressure of 0.5 to 6MPa, preferably at a pressure of 1 to 5MPa, said pressure being the gauge pressure.

In the hydrogenation step, the method for separating the hydrogenation product stream may further include a pressurizing step of pressurizing the gas-phase hydrogenation stream (i.e., the second gas-phase hydrogenation stream when the second gas-liquid separation step is not included, the second gas-phase hydrogenation stream and the third gas-phase hydrogenation stream which do not enter the second gas-liquid separation step when part of the second gas-phase hydrogenation stream is sent to the second gas-liquid step, and the third gas-phase hydrogenation stream when all of the second gas-phase hydrogenation stream is sent to the second gas-liquid separation step), so as to increase the pressure thereof, thereby meeting the requirement of the hydrogenation reaction. The degree of pressurization in the pressurization step can be selected according to the conditions of the hydrogenation reaction so as to meet the requirements.

In the hydrogenation step, part of the second gas-phase hydrogenation material flow and/or part of the third gas-phase hydrogenation material flow can be discharged out of the system, so that the accumulation of various impurities in the hydrogenation reaction system is reduced.

In the hydrogenation step, the vapor phase hydrogenation stream separated from the hydrogenation product stream may be recycled directly to the hydrogenation reaction as recycle hydrogen. The inventors of the present invention have found during research that treating a vapor phase hydrogenation stream separated from a hydrogenation product stream to reduce the carbon monoxide content, recycling the treated stream having a reduced carbon monoxide content to the hydrogenation reaction can further reduce the impurity content of the propylene oxide product, probably due to: when the methanol stream is hydrotreated, a trace amount of carbon monoxide is generated, the carbon monoxide is a poison of a catalyst with hydrogenation catalysis, and because hydrogen of hydrogenation reaction is recycled, the carbon monoxide is accumulated in the recycled hydrogen, so that the catalytic performance of the catalyst with hydrogenation catalysis is obviously reduced along with the extension of the reaction time, the impurity content in the methanol stream is difficult to keep at a lower level, and finally the impurity content in propylene oxide products is increased. Methods of reducing carbon monoxide may include, but are not limited to: membrane separation, selective adsorption and reactive removal.

In a preferred embodiment, the gas phase hydrogenation stream is contacted with a methanation catalyst under methanation reaction conditions to yield the treat stream. According to this preferred embodiment, the carbon monoxide in the gas phase hydrogenation stream undergoes methanation with hydrogen to form methane, thereby reducing the carbon monoxide content in the gas phase hydrogenation stream. The gas phase hydrogenation stream is contacted with the methanation catalyst to a degree such that the carbon monoxide content in the treat stream is preferably 5ppm or less, more preferably 3ppm or less, still more preferably 1ppm or less, for example: less than 0.5ppm and even less than 0.1 ppm. According to this preferred embodiment the carbon monoxide content of the gas phase hydrogenation stream is reduced without the additional introduction of other substances.

The methanation catalyst contains at least one catalytically active component, which may be selected from the group consisting of group VIII metals and group IB metals, preferably one or more of ruthenium, rhodium, palladium, platinum, silver, iridium, iron, copper, nickel and cobalt, more preferably nickel. The methanation catalyst further contains a carrier for supporting the catalytically active component, and the carrier may be a porous heat-resistant inorganic oxide, preferably one or more of silica, titania, zirconia and alumina, preferably alumina. The content of the catalytically active component in elemental form may be from 2 to 70% by weight, preferably from 20 to 60% by weight, more preferably from 30 to 50% by weight, based on the total amount of methanation catalyst.

The contact temperature of the gas phase hydrogenation stream with the methanation catalyst may be in the range of 70-250 ℃. In a preferred embodiment, the gas phase hydrogenation stream is contacted with the methanation catalyst at a temperature in the range of from 100 to 190 ℃, such as: 100. 105, 110, 115, 120, 125, 130, 135, 140, 145, 150, 155, 160, 165, 170, 175, 180, 185, or 190 ℃. According to the preferred embodiment, the carbon monoxide content in the gas phase hydrogenation stream can be reduced more effectively, and the single pass service life of the catalyst with hydrogenation catalysis can be further prolonged. According to this preferred embodiment, the contact temperature of the gas phase hydrogenation stream with the methanation catalyst is more preferably in the range of from 110 to 180 ℃, still more preferably in the range of from 130 to 180 ℃, still more preferably in the range of from 135 to 160 ℃. The pressure of contact of the gas phase hydrogenation stream with the methanation catalyst may be in the range of 0.5 to 10MPa, preferably 1 to 8MPa, more preferably 2 to 6MPa, still more preferably 3 to 5MPa, the pressure being in the gauge pressure.

The methanation reaction may be carried out in a common reactor. In a preferred embodiment, the gas phase hydrogenation stream and the methanation catalyst are carried out in a fixed bed reactor. When methanation is carried out in a fixed-bed reactor, the gas hourly space velocity (gas phase standard volume/catalyst volume) of the feed may be 500 to 10000h ^-1 Preferably 2000-8000h ^-1 More preferably 4000-6000h ^-1 . In the case of methanation in a fixed bed reactor, the gas phase hydrogenation stream may be fed in a direction through the catalyst bed from top to bottom, or from bottom to top, preferably from top to bottom.

In the hydrogenation step, at least the treatment stream is used as circulating hydrogen for the hydrogenation step, so that the recycling of hydrogen is realized. Fresh hydrogen may be replenished as necessary in the hydrogenation step. Fresh hydrogen may be mixed with the process stream for use in the hydrogenation step, for example: fresh hydrogen may be fed to the second or third vapor hydrogenation stream.

The liquid phase hydrogenation material flow obtained in the hydrogenation step can be directly sent to the methanol refining step for separation. When the pH of the liquid phase hydrogenation stream obtained in the hydrogenation step is higher than 7, the pH of the liquid phase hydrogenation stream is preferably adjusted to 7 or less, more preferably 3 to 7, and even more preferably 4 to 6.5 before the liquid phase hydrogenation stream is separated. The pH of the liquid phase hydrogenation stream may be adjusted by a variety of methods. In one embodiment, a pH adjustor can be added to the liquid phase hydrogenation stream. The pH adjustor may be an acidic substance, preferably one or two or more of an inorganic acid, an organic acid, and a salt of a strong acid and a weak base, more preferably one or two or more of hydrochloric acid, sulfuric acid, nitric acid, citric acid, oxalic acid, ammonium chloride, ammonium sulfate, and ammonium nitrate, and still more preferably sulfuric acid and/or citric acid. In another embodiment, the pH of the liquid phase hydrogenation stream may be adjusted by passing the liquid phase hydrogenation stream through a bed of acidic ion exchange resin, preferably a bed of strongly acidic ion exchange resin.

In the methanol refining step of the solvent recovery cycle process, the liquid-phase hydrogenation material flow is separated to obtain a methanol refined material flow, or the liquid-phase hydrogenation material flow is separated from the non-hydrotreated methanol material flow to obtain the methanol refined material flow.

In the methanol refining step, the liquid phase hydrogenation stream, or the liquid phase hydrogenation stream and the non-hydrotreated methanol stream, may be separated by a process sufficient to separate methanol and water, for example: the separation of methanol and water is realized by rectification.

In one embodiment (hereinafter referred to simply as "the first embodiment"), the separation is performed in a methanol rectifying column including a first rectifying column and a second rectifying column,

subjecting the liquid phase hydrogenation stream, or the liquid phase hydrogenation stream and the remaining portion of the methanol stream (i.e., the non-hydrotreated methanol stream) to a first rectification in a first rectification column at a first rectification pressure to obtain a first distillate comprising methanol from the top of the column and a first bottom stream comprising methanol and water from the bottom of the column;

subjecting the first bottoms stream to a second rectification in a second rectification column at a second rectification pressure to obtain a second distillate comprising methanol from the top of the column and a recovered water stream from the bottom of the column, the methanol refinery stream comprising either or a mixture of the first distillate and the second distillate;

The second rectification pressure is higher than the first rectification pressure, and at least part of the top gas phase of the second rectification tower is used as at least part of a heat source of a reboiler of the first rectification tower.

According to the first embodiment, not only can high-purity methanol be recovered, but also separation energy consumption can be effectively reduced.

In the first embodiment, the number of the first rectifying column and the second rectifying column is not limited to one, and one or more rectifying columns may be provided according to the throughput, and when two or more rectifying columns are used, the two or more rectifying columns may be connected in parallel, connected in series, or a combination of parallel and series. Preferably, the number of the first rectifying tower and the second rectifying tower is one.

In the first embodiment, the first rectification pressure is preferably 0.01 to 0.5MPa, more preferably 0.1 to 0.4MPa, still more preferably 0.2 to 0.3MPa, and the second rectification pressure is preferably 0.5 to 1.2MPa, more preferably 0.6 to 1MPa, still more preferably 0.7 to 0.9MPa, in terms of gauge pressure. In the present invention, the term "rectification pressure" refers to the overhead operating pressure of the rectification column.

In a first embodiment, the top temperature of the first rectification column may be in the range of 70 to 120 ℃, preferably 80 to 110 ℃, more preferably 90 to 105 ℃. The reflux ratio of the first rectification column may be from 0.5 to 2, preferably from 0.6 to 1.5, more preferably from 0.8 to 1.2. In the first embodiment, the top temperature of the second rectification column may be 100 to 150 ℃, preferably 110 to 140 ℃, more preferably 120 to 135 ℃. The reflux ratio of the second rectification column may be from 0.5 to 3, preferably from 1.5 to 2.5.

In the first embodiment, the theoretical plate numbers of the first rectifying column and the second rectifying column may each be 30 to 50, preferably 35 to 45.

In another embodiment (hereinafter referred to simply as "second embodiment"), in the methanol refining step, the separation is performed in a methanol rectifying column including a third rectifying column and a fourth rectifying column,

subjecting the liquid phase hydrogenation stream, or the liquid phase hydrogenation stream and the remaining portion of the methanol stream (i.e., the non-hydrotreated methanol stream) to a third rectification in a third rectification column at a third rectification pressure to obtain a third distillate comprising methanol from the top of the column and a third bottom stream comprising methanol and water from the bottom of the column;

subjecting the third bottoms stream to a fourth rectification in a fourth rectification column at a fourth rectification pressure to obtain a fourth distillate comprising methanol from the top of the column and a recovered water stream from the bottom of the column, the methanol refinery stream comprising either or a mixture of the third distillate and the fourth distillate;

the third rectification pressure is higher than the fourth rectification pressure, and at least part of the top gas phase of the third rectification tower is used as at least part of a heat source of a reboiler of the fourth rectification tower.

According to the second embodiment, not only can high-purity methanol be recovered, but also separation energy consumption can be effectively reduced.

In the second embodiment, the number of the third rectifying column and the fourth rectifying column is not limited to one, and one or more rectifying columns may be provided according to the throughput, and when two or more rectifying columns are used, the two or more rectifying columns may be connected in parallel, connected in series, or a combination of parallel and series. Preferably, the number of the third rectifying tower and the fourth rectifying tower is one.

In a second embodiment, the pressure of the fourth rectifying column is lower than the pressure of the third rectifying column, and the third rectifying pressure may be 1 to 2MPa, preferably 1.2 to 1.8MPa, and the fourth rectifying pressure may be 0.01 to 0.5MPa, preferably 0.1 to 0.4MPa, more preferably 0.2 to 0.3MPa, in terms of gauge pressure.

In a second embodiment, the top temperature of the third rectification column may be 140-180 ℃, preferably 142-170 ℃, more preferably 145-160 ℃. The reflux ratio of the third rectification column may be from 0.6 to 1.8, preferably from 1 to 1.5. In the second embodiment, the top temperature of the fourth rectifying column may be 70 to 120 ℃, preferably 80 to 115 ℃, more preferably 90 to 110 ℃, still more preferably 95 to 105 ℃. The reflux ratio of the fourth rectifying column may be 0.5 to 2, preferably 0.8 to 1.5.

In the second embodiment, the theoretical plate number of the third rectifying column and the fourth rectifying column may each be 30 to 50, preferably 35 to 45.

According to the method provided by the invention, the solvent recovery and recycling process comprises an ethanol removal step, and the ethanol content in the methanol refined stream is reduced by arranging the ethanol removal step, so that the product selectivity of the epoxidation reaction can be effectively improved. The entire methanol refinery stream may be subjected to the ethanol removal step or a portion of the methanol refinery stream may be subjected to the ethanol removal step. According to the process of the present invention, the content of ethanol in the recovered methanol is preferably less than 4 wt.%, such as less than 3 wt.% or less than 2 wt.%, based on the total amount of recovered methanol. The content of ethanol in the recovered methanol is more preferably 1% by weight or less, still more preferably 0.5% by weight or less, still more preferably 0.4% by weight or less, still more preferably 0.2% by weight or less, and particularly preferably 0.1% by weight or less, based on the total amount of the recovered methanol, from the viewpoint of further improving the selectivity of the product of the epoxidation reaction. In the invention, the content of ethanol in the recovered methanol is determined by a gas chromatography method. According to the process of the present invention, at least 10 wt.% of the methanol refinery stream may be fed to the ethanol rectification column for rectification, preferably at least 50 wt.% of the methanol refinery stream is fed to the ethanol rectification column for rectification, more preferably at least 80 wt.% of the methanol refinery stream is fed to the ethanol rectification column for rectification, based on the total amount of the methanol refinery stream. From the viewpoint of further reducing the ethanol content in the recovered methanol, it is preferable to feed the whole methanol refined stream to an ethanol rectifying column for rectification.

In the ethanol removal step, the methanol refined stream may be rectified in an ethanol rectification column, recovered methanol having a reduced ethanol content is obtained from the top of the ethanol rectification column, and a bottom stream containing ethanol is obtained from the bottom of the ethanol rectification column. From the standpoint of both reducing the ethanol content in the recovered methanol and reducing the operation energy consumption of the ethanol rectifying tower, the tower top pressure of the ethanol rectifying tower is preferably 0.01-0.5MPa, more preferably 0.02-0.3MPa, still more preferably 0.05-0.1MPa, and the tower top pressure is gauge pressure; the reflux ratio is preferably 1 to 5, more preferably 2 to 4.5; the tower top temperature of the ethanol rectifying tower is preferably 60-85 ℃, more preferably 70-82 ℃; the bottom temperature of the ethanol rectifying tower is preferably 90-120 ℃, more preferably 95-110 ℃. The theoretical plate number of the ethanol rectifying column is preferably 20 to 65, more preferably 30 to 60, still more preferably 40 to 55. In a preferred embodiment, the reflux ratio of the ethanol rectification column is higher than 2, preferably in the range of 2.5 to 4, more preferably in the range of 3 to 4. According to this preferred embodiment, the content of ethanol in the recovered methanol can be further reduced to obtain recovered methanol having an ethanol content of less than 0.2% by weight, for example, recovered methanol having an ethanol content of not more than 0.1% by weight.

According to the method of the present invention, at least a part of the recovered methanol is recycled to the epoxidation step and used as a solvent for the epoxidation reaction. Fresh methanol can be supplemented according to specific conditions so that the amount of methanol can meet the requirements of the epoxidation reaction.

The epoxidation reaction unit comprises at least one epoxidation reactor, preferably a fixed bed reactor, more preferably a tubular fixed bed reactor, in which the ratio of the inner diameter of the tubulars to the length of the tubulars (abbreviated as "aspect ratio") is preferably 50-500, more preferably 100-250, even more preferably 150-200. The number of epoxidation reactors may be one or more, preferably 2 to 10 reactors connected in series. The fixed bed reactor may be an upflow fixed bed reactor (i.e. the feed stream flows through the catalyst bed from bottom to top), or a downflow fixed bed reactor (i.e. the feed stream flows through the catalyst bed from top to bottom), preferably an upflow fixed bed reactor.

The epoxidation reaction product separation unit comprises a reaction product separation subunit, a propylene oxide refining subunit and an intermediate impurity removal subunit.

The reaction product separation subunit includes a first distillation column for separating an epoxidation reaction product stream into a light stream comprising propylene oxide, propylene, and a portion of methanol, and a heavy stream comprising water and a remaining portion of methanol. The theoretical plate number of the first distillation column may be 10 to 50, preferably 15 to 45, more preferably 20 to 30.

The propylene separation unit comprises a second distillation column and an absorber column for separating at least a portion of the propylene from the light stream. The input port of the material to be separated of the second distillation tower is communicated with the output port of the light material flow of the reaction product separation unit, and the light material flow is used for carrying out second distillation in the second distillation tower to obtain a gas phase scavenging material flow containing propylene and a propylene oxide material flow containing propylene oxide and methanol. The material input port to be separated of the absorption tower is communicated with the gas phase scavenging stream output port of the second distillation tower and is used for contacting the gas phase scavenging stream with an absorbent to obtain a gas phase hydrogenation stream containing propylene and a propylene oxide absorption stream containing the absorbent and propylene oxide. And an absorption material flow output port of the absorption tower is communicated with a material input port to be separated of the reaction product separation subunit, and is used for circularly conveying the absorption material flow into the reaction product separation subunit for separation. The theoretical plate number of the second distillation column may be 20 to 40, preferably 25 to 35.

The propylene oxide refining subunit includes an extractive distillation column and an optional third distillation column.

The extractive distillation tower is used for contacting propylene oxide material flow containing propylene oxide and methanol with an extracting agent in an extractive distillation zone of the extractive distillation tower for extractive distillation, so that a propylene oxide product and a liquid-phase material flow containing methanol and the extracting agent are obtained. The theoretical plate number of the extractive distillation column is preferably (0.6 to 0.9) ×t, more preferably (0.6 to 0.85) ×t, still more preferably (0.7 to 0.8) ×t, and the extractive distillation column is preferably (0.15 to 0.45) ×t, more preferably (0.2 to 0.4) ×t, still more preferably (0.25 to 0.35) ×t, and the theoretical plate number is the theoretical plate number counted from top to bottom. The theoretical plate number T of the extractive distillation column is preferably 35 to 90, more preferably 45 to 85, still more preferably 55 to 75.

Preferably, the propylene oxide refining subunit further comprises a preprocessor disposed upstream of the extractive distillation column for preprocessing a propylene oxide stream containing propylene oxide and methanol and feeding the preprocessed propylene oxide stream into the extractive distillation column. In one embodiment, the interior space of the preconditioner is filled with basic ion exchange resin, and the propylene oxide stream flows through the interior space of the preconditioner to contact the basic ion exchange resin, thereby performing the preconditioning. In another embodiment, the pre-processor has a hollow interior space and an alkaline material input port, the propylene oxide stream flows through the interior space of the pre-processor, and an alkaline material (e.g., an aqueous solution containing an alkaline material) is added to the interior space through the alkaline material input port to mix the alkaline material with the propylene oxide stream, in this embodiment, a preferred example of the pre-processor is a pipe mixer. In yet another embodiment, the pre-processor is at least one light ends rectification column for rectifying the propylene oxide stream in the light ends rectification column. The theoretical plate number of the light ends removal rectifying column is preferably 30 to 70.

The third distillation tower is used for carrying out third separation on a liquid-phase stream containing methanol and an extractant, a distillate containing intermediate impurities is obtained from the top of the third distillation tower, a purified liquid-phase stream is obtained from the bottom of the third distillation tower, the intermediate impurity content in the purified liquid-phase stream is lower than that in the liquid-phase stream, and the purified liquid-phase stream is sent to a solvent recovery circulation process for treatment. The theoretical plate number of the third distillation column may be 30 to 70, preferably 40 to 60, more preferably 45 to 55. The ratio of the theoretical plate number corresponding to the feed position of the liquid phase stream containing methanol and extractant to the theoretical plate number of the third distillation column is preferably 0.3 to 0.7, more preferably 0.35 to 0.6, still more preferably 0.4 to 0.55. The epoxidation system according to the present invention preferably includes a third distillation column from the viewpoint of further reducing the content of intermediate impurities in the recovered methanol and further improving the product selectivity of the epoxidation reaction.

The hydrogenation subunit comprises a hydrogenation reactor, a hydrogenation reaction product separation device and a gas phase hydrogenation material flow treatment device.

The hydrogenation reactor is used for contacting at least part of methanol stream with a catalyst with hydrogenation catalysis under hydrogenation reaction conditions to obtain a hydrogenation product stream with reduced impurity content. Preferably, the hydrogenation reactor is in communication with a heavies stream output port of the first distillation column for contacting the heavies stream with a catalyst having a hydrogenation catalytic effect. The hydrogenation reactor may be a conventional one capable of carrying out hydrogenation reactions, for example: one or a combination of more than two of a fixed bed reactor, a slurry bed reactor and a fluidized bed reactor. In a preferred embodiment, the hydrogenation reactor is a fixed bed reactor.

The to-be-separated material flow input unit of the hydrogenation reaction product separation device is communicated with a hydrogenation product material flow output pipeline of the hydrogenation reactor and is used for separating the hydrogenation product material flow to obtain a gas-phase hydrogenation material flow containing hydrogen and a liquid-phase hydrogenation material flow containing alcohol.

In a preferred embodiment, the hydrogenation reaction product separation device comprises a first gas-liquid separator, an absorption column and optionally a second gas-liquid separator. And a to-be-separated material flow input pipeline of the first gas-liquid separator is communicated with a hydrogenation product material flow output pipeline of the hydrogenation reactor and is used for carrying out gas-liquid separation on the hydrogenation product material flow to obtain a first gas-phase hydrogenation material flow and a first liquid-phase hydrogenation material flow. And the to-be-absorbed material flow input pipeline of the absorption tower is communicated with the gas phase hydrogenation material flow output pipeline of the first gas-liquid separator and is used for contacting the first gas phase hydrogenation material flow with a liquid absorbent to obtain a second gas phase hydrogenation material flow containing hydrogen and a second liquid phase hydrogenation material flow containing absorbent. And a to-be-separated material flow input pipeline of the second gas-liquid separator is communicated with a gas-phase hydrogenation material flow output pipeline of the absorption tower and is used for carrying out second gas-liquid separation on the second gas-phase hydrogenation material flow to obtain a third gas-phase hydrogenation material flow containing hydrogen and a third liquid-phase hydrogenation material flow. The first gas-liquid separator and the second gas-liquid separator can realize gas-liquid two-phase separation by adjusting pressure, temperature or a combination of the two.

The hydrogenation reaction product separation device preferably comprises a pressurizing device, wherein a material flow input pipeline to be pressurized of the pressurizing device is communicated with a gas-phase hydrogenation material flow output pipeline of the absorption tower and/or a gas-phase hydrogenation material flow output pipeline of the second gas-liquid separator, and is used for improving the pressure of the second gas-phase hydrogenation material flow and/or the third gas-phase hydrogenation material flow. The supercharging device may be pressurised using conventional components. For example: the pressure boosting means may comprise at least one compressor to compress the second and/or third vapor phase hydrogenation streams to increase their pressure.

The to-be-treated material flow input pipeline of the gas-phase hydrogenation material flow treatment device is communicated with the gas-phase hydrogenation material flow output pipeline of the hydrogenation reaction product separation device; the gas-phase hydrogenation material flow treatment device comprises a processor for reducing the content of carbon monoxide, and is used for treating the gas-phase hydrogenation material flow output by the hydrogenation reaction product separation device to obtain a treated material flow with reduced content of carbon monoxide, and sending at least the treated material flow into a hydrogenation reactor.

The processor for reducing the carbon monoxide content may be a membrane separator, an adsorption tower, or a reactor capable of chemically reacting carbon monoxide.

In a preferred embodiment, the processor for reducing the carbon monoxide content is a methanation reactor, in which. The methanation reactor may be a common reactor capable of performing methanation reaction. In a preferred embodiment, the methanation reactor is a fixed bed reactor. The fixed bed reactor can be a kettle type fixed bed reactor or a tubular fixed bed reactor, such as a tubular fixed bed reactor.

In one embodiment, the methanol rectification subunit comprises a first rectification column and a second rectification column, wherein the first rectification column is used for carrying out first rectification on a liquid phase hydrogenation stream or the liquid phase hydrogenation stream and a methanol stream which is not subjected to hydrogenation treatment at a first rectification pressure, a first distillate containing alcohol is obtained from the top of the first rectification column, and a first bottom stream containing alcohol and water is obtained from the bottom of the first rectification column;

the second rectifying tower is used for carrying out second rectification on the first tower bottom stream under a second rectifying pressure, a second distillate containing alcohol is obtained from the tower top of the second rectifying tower, and a recovered water stream is obtained from the tower bottom;

the rectification pressure of the second rectification tower is higher than that of the first rectification tower, and the output port of the gas-phase hydrogenation material flow at the top of the second rectification tower is communicated with the heat source input port of the reboiler of the first rectification tower.

In another embodiment, the methanol rectification subunit comprises a third rectification column and a fourth rectification column,

the third rectifying tower is used for carrying out third rectification on the liquid-phase hydrogenation material flow or the liquid-phase hydrogenation material flow and the methanol material flow which is not subjected to hydrogenation treatment under third rectifying pressure, a third distillate containing alcohol is obtained from the top of the third rectifying tower, and a third bottom material flow containing alcohol and water is obtained from the bottom of the third rectifying tower;

the fourth rectifying tower is used for carrying out fourth rectification on the third tower bottom stream at a fourth rectifying pressure, a fourth distillate containing alcohol is obtained from the tower top of the fourth rectifying tower, and a recovered water stream is obtained from the tower bottom;

and the rectification pressure of the fourth rectification tower is higher than that of the third rectification tower, and the output port of the gas-phase hydrogenation material flow at the top of the third rectification tower is communicated with the heat source input port of the reboiler of the third rectification tower.

The present invention will be described in detail with reference to examples, but the scope of the present invention is not limited thereto.

In the following examples and comparative examples, unless otherwise specified, the pressures were gauge pressures, and the theoretical plate numbers were theoretical plate numbers from top to bottom; the composition of the various streams was determined by gas chromatography.

In the following examples and comparative examples, the single pass service life of the catalyst having hydrogenation catalytic action was evaluated by the following methods: the composition of the hydrogenation feedstock and the composition of the liquid phase hydrogenation stream separated from the hydrogenation reaction product stream are determined, the acetaldehyde conversion is calculated using the following method,

acetaldehyde conversion (%) = [1- (acetaldehyde content in liquid phase hydrogenation stream separated from hydrogenation reaction product stream/acetaldehyde content in hydrogenation feedstock) ] ×100%

When the catalyst having a hydrogenation catalytic action reaches a single-pass service life, which is considered to be measured in terms of the acetaldehyde conversion measured when the hydrogenation reaction is stably carried out for 1 hour, when the acetaldehyde conversion is reduced to 50%, the single-pass service life is measured in units of months, and if the time for which the acetaldehyde conversion is reduced to 50% is the first 15 days of one month, the month is excluded, whereas the month is included (for example, the catalyst having a hydrogenation catalytic action is reduced to 50% when the catalyst having a hydrogenation catalytic action is used for about 3 months and 10 days, the single-pass service life of the catalyst having a hydrogenation catalytic action is measured as 3 months, and for example, the catalyst having a hydrogenation catalytic action is reduced to 50% when the catalyst having a hydrogenation catalytic action is used for about 3 months and 20 days, and the single-pass service life of the catalyst having a hydrogenation catalytic action is measured as 4 months).

Examples 1-6 illustrate the invention.

Example 1

(1) Epoxidation reaction

Methanol, propylene and hydrogen peroxide (hydrogen peroxide content 50 wt%) were prepared according to propylene: hydrogen peroxide: the molar ratio of methanol is 2:1:8 mixing, feeding the mixture from the bottom into a tubular fixed bed reactor (inner diameter of tube) filled with an epoxidation catalyst (model HPO-1, which is a molded hollow titanium silicalite catalyst, manufactured by Kagaku catalyst Kagaku Co., ltd.)4 m), and the catalyst is contacted with an epoxidation catalyst to carry out epoxidation, wherein the temperature in a fixed bed reactor is controlled to be 45 ℃, and the volume space velocity of hydrogen peroxide solution is controlled to be 0.2h when the hydrogen peroxide solution is calculated ^-1 An epoxidation reaction product stream is obtained from the top of the epoxidation reactor.

(2) Epoxidation reaction product stream separation

The epoxidation reaction product stream is fed to a first distillation column for rectification, a light stream comprising propylene oxide, propylene and methanol is obtained from the top of the first distillation column and a heavy stream comprising methanol and water is obtained from the bottom of the first distillation column. The theoretical plate number of the first distillation column was 35, the column top pressure was 0.1MPaG, the column top temperature was 69 ℃, no reflux was generated, and the column top was fed.

The light stream is fed into a second distillation column for rectification, a gas phase purge stream containing propylene is obtained from the top of the second distillation column, and a propylene oxide stream containing propylene oxide and methanol is obtained from the bottom of the second distillation column. The theoretical plate number of the second distillation column was 25, the column top pressure was 0.1MPaG, the column top temperature was 36 ℃, no reflux, and the column top was fed.

The gas phase scavenging stream containing propylene enters an absorption tower to contact with methanol serving as an absorbent, a gas phase stream containing propylene is obtained from the top of the absorption tower, a propylene oxide absorption stream containing the absorbent and propylene oxide is obtained from the bottom of the absorption tower, and the propylene oxide absorption stream is sent to a first distillation tower to be separated. Wherein the feed flow rate of methanol as absorbent was 65kg/h and the feed flow rate of the gas phase purge stream containing propylene was 36.9kg/h. The temperature in the absorption column was 25 ℃, the pressure in the absorption column was 0.04MPaG, the theoretical plate number of the absorption column was 25, methanol as an absorbent was fed from the top of the absorption column, and a vapor purge stream containing propylene was fed from the bottom of the absorption column.

The propylene oxide stream containing propylene oxide and methanol was obtained from the bottom of the second distillation column and passed through a pre-treater packed with a basic ion exchange resin, which was a weak basic ion exchange resin in a LEWATIT1073 acrylic acid gel type available from the company Tianjin duplex technology limited. Liquid hourly space velocity of propylene oxide stream of 1h ^-1 The temperature in the preconditioner was 60 ℃.

The propylene oxide material flow output by the preprocessor enters an extraction rectifying tower to contact with an extractant for extraction and rectification, a propylene oxide product is obtained from the top of the extraction rectifying tower, a liquid-phase material flow containing methanol and the extractant is obtained from the bottom of the extraction rectifying tower, wherein the extractant is hydrazine water (the content of hydrazine is 0.5 weight percent), the theoretical plate number of the extraction rectifying tower is 70, the theoretical plate number corresponding to the feeding position of the extractant is 20, the theoretical plate number corresponding to the feeding position of the propylene oxide material flow is 50, the top pressure of the extraction rectifying tower is 0.04MPaG, the top temperature is 44 ℃, the bottom temperature is 80 ℃, and the reflux ratio is 3.

And (3) feeding a liquid-phase stream containing methanol and an extractant obtained from the bottom of the extractive distillation tower into a third distillation tower for rectification, obtaining a distillate containing middle impurities from the top of the third distillation tower, and obtaining a purified liquid-phase stream from the bottom of the third distillation tower. The theoretical plate number of the third distillation column was 45, the theoretical plate number corresponding to the feed position of the liquid phase stream containing methanol and extractant was 20, the column top pressure was 0.04MPaG, the column top temperature was 70℃and the reflux ratio was 100.

(3) Solvent recovery cycle

The heavy stream obtained at the bottom of the first distillation column enters a fixed bed hydrogenation reactor filled with a catalyst with hydrogenation catalysis (the catalyst with hydrogenation catalysis is EH-11 hydrogenation catalyst produced by Hunan Changling petrochemical technology development Co., ltd.) from the bottom, wherein the catalytic active component is nickel, the content of the catalytic active component is 40 weight percent based on elements, and the carrier is alumina) for hydrotreating, and a hydrogenation product stream is obtained from the top of the fixed bed hydrogenation reactor. Wherein the temperature of the hydrotreatment is 110 ℃, the pressure in the hydrogenation reactor is 4MPaG, and the liquid hourly space velocity is 12h ^-1 。

And (3) feeding the outlet material of the fixed bed hydrogenation reactor into a gas-liquid separation tank for gas-liquid separation to obtain a first gas phase hydrogenation material flow and a liquid phase hydrogenation material flow containing methanol, wherein the temperature in the gas-liquid separation tank is 115 ℃, and the pressure is 4MPaG. The first gas phase hydrogenation material flow enters the lower part of the hydrogenation tail gas absorption tower, water is added into the upper part of the absorption tower, methanol in the first gas phase hydrogenation material flow is absorbed by the water after reversely contacting, and a small part of gas phase which cannot be absorbed is discharged out of the system, and the temperature in the absorption tower is 45 ℃. Most of the gas phase enters a new hydrogen separation tank, and is mixed with fresh hydrogen in the new hydrogen separation tank, wherein the temperature in the new hydrogen separation tank is 45 ℃, and the pressure is 1.4MPag.

The mixed gas from the new hydrogen separation tank is boosted by a compressor, enters a methanation reactor with a catalyst bed layer from top to bottom and is filled with a methanation catalyst (the brand name of the methanation catalyst is BC-H-10 low-temperature methanation catalyst produced by the national institute of chemical industry, beijing, petrifaction, china), the catalytic active component of the methanation catalyst is nickel, and the content of the catalytic active component is about 30 weight percent based on the total amount of the catalystThe methanation reaction is performed in the middle stage, and the methanation reactor is a fixed bed reactor. The methanation reactor has a temperature of 135 ℃, a pressure of 4.1MPaG and a space velocity of 5000h ^-1 . The outlet gas phase hydrogenation stream from the methanation reactor is recycled to the fixed bed hydrogenation reactor.

The liquid phase output by the hydrotreatment product liquid separation tank, the hydrotreatment tail gas absorption tower and the new hydrogen liquid separation tank are mixed together as a liquid phase hydrogenation stream and sulfuric acid (with the concentration of 70 weight percent) in a mixer, and the pH value of the liquid phase hydrogenation stream is regulated to be 5.

During the reaction, the carbon monoxide content of the outlet stream of the methanation reactor was continuously monitored, wherein the carbon monoxide content in the outlet stream of the methanation reactor was kept below 0.1ppmw (ppm by weight). The single pass service life of the catalyst with hydrogenation catalysis was evaluated, and it was determined that the single pass service life of the catalyst with hydrogenation catalysis was 6 months.

The liquid phase hydrogenation material flow with the pH value adjusted and the purified liquid phase hydrogenation material flow obtained from the bottom of the third distillation tower are sent into a first distillation tower to be rectified, and methanol is extracted from the top of the first distillation tower. The theoretical plate number of the first rectifying tower is 40, the rectifying pressure is 0.25MPaG, the tower top temperature is 100 ℃, and the reflux ratio is 1.

And (3) delivering the bottom material flow of the first rectifying tower into a second rectifying tower for rectification, extracting methanol from the top of the second rectifying tower, delivering part of the extracted steam into a tower bottom reboiler of the first rectifying tower for heat exchange with tower bottom liquid of the first rectifying tower, and extracting water from the tower bottom of the second rectifying tower and outputting. The theoretical plate number of the second rectifying column was 35, the column top pressure was 0.7MPaG, the column top temperature was 130℃and the reflux ratio was 2.

And (3) feeding the methanol extracted from the tops of the first rectifying tower and the second rectifying tower into an ethanol rectifying tower for rectification, obtaining recovered methanol from the top of the ethanol rectifying tower, extracting a tower bottom stream containing ethanol from the tower bottom of the ethanol rectifying tower, and fully recycling the recovered methanol for epoxidation reaction. Wherein the theoretical plate number of the ethanol rectifying tower is 50, the tower top pressure is 0.05MPaG, the tower top temperature is 75 ℃, the tower bottom temperature is 95 ℃, and the reflux ratio is 3.

Steps (1) through (3) above were continuously carried out for 2000 hours, the composition of the epoxidation reaction product stream was continuously monitored and propylene oxide selectivity was calculated using the following formula,

propylene oxide selectivity (%) = [ molar amount of propylene oxide generated by reaction/(molar amount of hydrogen peroxide added-molar amount of unreacted hydrogen peroxide) ] ×100%.

Continuously monitoring the content of ethanol and intermediate impurities in the recovered methanol, continuously monitoring the composition of the propylene oxide product extracted from the top of the extraction rectifying tower, and determining the purity of the propylene oxide product.

The experimental results are listed in table 1.

TABLE 1

Comparative example 1

Propylene oxide was produced in the same manner as in example 1 except that in step (3), the ethanol rectifying column and the third rectifying column were not provided, and methanol recovered from the tops of the first rectifying column and the second rectifying column was entirely recycled for the epoxidation reaction as recovered methanol. The experimental results are listed in table 2.

TABLE 2

Example 2

Propylene oxide was produced in the same manner as in example 1 except that in step (3), 10% by weight of methanol taken out from the tops of the first rectifying column and the second rectifying column was fed into the ethanol rectifying column to be rectified, and the remaining 90% by weight of methanol was not fed into the ethanol rectifying column to be rectified, but was mixed with methanol taken out from the top of the ethanol rectifying column to be recycled as recovered methanol for the whole epoxidation reaction. The experimental results are listed in table 3.

TABLE 3 Table 3

Example 3

Propylene oxide was produced in the same manner as in example 1 except that the third distillation column was not provided in step (2), and the liquid-phase hydrogenated stream containing methanol and extractant extracted from the bottom of the extractive distillation column was directly fed into the first distillation column together with the liquid-phase hydrogenated stream having the pH adjusted for separation. The experimental results are listed in table 4.

TABLE 4 Table 4

Example 4

Propylene oxide was produced in the same manner as in example 1, except that the methanation reactor was not provided in step (3), and the mixed gas boosted by the compressor was directly circulated to the fixed bed hydrogenation reactor, resulting in a single-pass service life of the catalyst having hydrogenation catalytic action of 1 month.

Example 5

Propylene oxide was produced in the same manner as in example 1 except that in step (3), the reflux ratio of the ethanol rectifying column was adjusted to 2. The experimental results are listed in table 5.

TABLE 5

Example 6

(1) Epoxidation reaction

Methanol, propylene and hydrogen peroxide (hydrogen peroxide content 50 wt%) were prepared according to propylene: hydrogen peroxide: the molar ratio of methanol is 3:1:10, mixing, feeding the mixture from the bottom into a tubular fixed bed reactor (inner diameter of tube) filled with an epoxidation catalyst (model HPO-1, available from Kagaku catalyst Kagaku Co., ltd., which is a molded hollow titanium silicalite molecular sieve catalyst) 4 m), and the catalyst is contacted with an epoxidation catalyst to carry out epoxidation, wherein the temperature in a fixed bed reactor is controlled to be 45 ℃, and the liquid hourly space velocity calculated by hydrogen peroxide is controlled to be 0.2h ^-1 An epoxidation reaction product stream is obtained from the top of the epoxidation reactor.

(2) Epoxidation reaction product stream separation

The epoxidation reaction product stream is fed to a first distillation column for rectification, a light stream comprising propylene oxide, propylene and methanol is obtained from the top of the first distillation column and a heavy stream comprising methanol and water is obtained from the bottom of the first distillation column. The theoretical plate number of the first distillation column was 35, the column top pressure was 0.1MPaG, the column top temperature was 66 ℃, no reflux was found, and the column top was fed.

The light stream is fed into a second distillation column for rectification, a gas phase purge stream containing propylene is obtained from the top of the second distillation column, and a propylene oxide stream containing propylene oxide and methanol is obtained from the bottom of the second distillation column. The theoretical plate number of the second distillation column was 25, the column top pressure was 0.1MPaG, the column top temperature was 39 ℃, no reflux, and the column top was fed.

The propylene oxide material flow output by the preprocessor enters an extraction rectifying tower to contact with an extractant for extraction and rectification, a propylene oxide product is obtained from the top of the extraction rectifying tower, a liquid-phase material flow containing methanol and the extractant is obtained from the bottom of the extraction rectifying tower, wherein the extractant is hydrazine water (the content of hydrazine is 0.5 weight percent), the theoretical plate number of the extraction rectifying tower is 70, the theoretical plate number corresponding to the feeding position of the extractant is 18, the theoretical plate number corresponding to the feeding position of the propylene oxide material flow is 55, the top pressure of the extraction rectifying tower is 0.04MPaG, the top temperature is 44 ℃, the bottom temperature is 81 ℃, and the reflux ratio is 2.5.

And (3) feeding a liquid-phase stream containing methanol and an extractant obtained from the bottom of the extractive distillation tower into a third distillation tower for rectification, obtaining a distillate containing middle impurities from the top of the third distillation tower, and obtaining a purified liquid-phase stream from the bottom of the third distillation tower. The theoretical plate number of the third distillation column was 50, the theoretical plate number corresponding to the feed position of the liquid phase stream containing methanol and extractant was 25, the column top pressure was 0.04MPaG, the column top temperature was 56℃and the reflux ratio was 150.

(3) Solvent recovery cycle

Feeding heavy stream obtained from the bottom of the first distillation column into a fixed bed hydrogenation reactor (the catalyst with hydrogenation catalysis is EH-11 hydrogenation catalyst purchased from Hunan Changling petrochemical Co., ltd., wherein the catalytically active component is nickel, and the catalytically active component is calculated by element)The split content was 40 wt.%) and a hydrogenation product stream was obtained from the top of the fixed bed hydrogenation reactor. Wherein the temperature of the hydrotreatment is 85 ℃, the pressure in the hydrogenation reactor is 5MPaG, and the liquid hourly space velocity is 15h ^-1 。

And (3) feeding the outlet material of the fixed bed hydrogenation reactor into a gas-liquid separation tank for gas-liquid separation to obtain a first gas phase hydrogenation material flow and a liquid phase hydrogenation material flow containing methanol, wherein the temperature in the gas-liquid separation tank is 89 ℃, and the pressure is 4.95MPaG. The first gas phase hydrogenation material flow enters the lower part of the hydrogenation tail gas absorption tower, water is added into the upper part of the absorption tower, methanol in the first gas phase hydrogenation material flow is absorbed by the water after reversely contacting, and a small part of gas phase which cannot be absorbed is discharged out of the system, and the temperature in the absorption tower is 45 ℃. Most of the gas phase enters a new hydrogen separation tank and is mixed with fresh hydrogen in the new hydrogen separation tank. The temperature in the fresh hydrogen separation tank was 44℃and the pressure was 4.9MPag.

The mixed gas from the new hydrogen separation tank is boosted by a compressor, enters a methanation reactor with a methanation catalyst (the brand name is BC-H-10 low-temperature methanation catalyst produced by Beijing chemical industry institute of petrochemical industry, china) filled in a catalyst bed from top to bottom, and the methanation catalyst has a catalytic active component of nickel and a content of the catalytic active component of about 30 weight percent based on the total amount of the catalyst for methanation reaction. The methanation reactor has a temperature of 160 ℃, a pressure of 5MPaG and a space velocity of 6000h ^-1 . The outlet gas phase hydrogenation stream from the methanation reactor is recycled to the fixed bed hydrogenation reactor.

The liquid phase output by the hydrotreatment product liquid separating tank, the hydrotreatment tail gas absorption tower and the new hydrogen liquid separating tank are mixed with sulfuric acid (the concentration is 60 weight percent) in a mixer together as a liquid phase hydrogenation material flow, and the pH value of the liquid phase hydrogenation material flow is regulated to be 6.2.

During the reaction, the carbon monoxide content of the outlet stream of the methanation reactor was continuously monitored, wherein the carbon monoxide content in the outlet stream of the methanation reactor was kept below 0.1 ppmw. The single pass service life of the catalyst with hydrogenation catalysis was evaluated, and it was determined that the single pass service life of the catalyst with hydrogenation catalysis was 6 months.

The liquid phase hydrogenation material flow with the pH value adjusted and the obtained purified liquid phase material flow from the bottom of the third distillation tower are sent into a first distillation tower to be rectified, and methanol is extracted from the top of the first distillation tower. The theoretical plate number of the first rectifying column is 40, the rectifying pressure is 1.2MPaG, the tower top temperature is 148 ℃, and the reflux ratio is 1.2.

And (3) delivering the bottom material flow of the first rectifying tower into a second rectifying tower for rectification, extracting methanol from the top of the second rectifying tower, delivering part of the extracted steam into a tower bottom reboiler of the first rectifying tower for heat exchange with tower bottom liquid of the first rectifying tower, and extracting water from the tower bottom of the second rectifying tower and outputting. The theoretical plate number of the second rectifying column was 35, the column top pressure was 0.2MPaG, the column top temperature was 101℃and the reflux ratio was 1.

And (3) feeding the methanol extracted from the tops of the first rectifying tower and the second rectifying tower into an ethanol rectifying tower for rectification, obtaining recovered methanol from the top of the ethanol rectifying tower, extracting a tower bottom stream containing ethanol from the tower bottom of the ethanol rectifying tower, and fully recycling the recovered methanol for epoxidation reaction. Wherein the theoretical plate number of the ethanol rectifying tower is 55, the tower top pressure is 0.1MPaG, the tower top temperature is 82 ℃, the tower bottom temperature is 101 ℃, and the reflux ratio is 4.

The steps (1) to (3) described above were continuously carried out for 2000 hours, the composition of the epoxidation reaction product stream was continuously monitored and the propylene oxide selectivity was calculated, the contents of ethanol and intermediate impurities in the recovered methanol were continuously monitored, the composition of the propylene oxide product withdrawn from the top of the extractive distillation column was continuously monitored, and the purity of the propylene oxide product was determined, and the results are shown in table 6.

TABLE 6

The preferred embodiments of the present invention have been described in detail above, but the present invention is not limited thereto. Within the scope of the technical idea of the invention, a number of simple variants of the technical solution of the invention are possible, including combinations of the individual technical features in any other suitable way, which simple variants and combinations should likewise be regarded as being disclosed by the invention, all falling within the scope of protection of the invention.

Claims

1. A propylene epoxidation process comprising an epoxidation reaction step, an epoxidation reaction product separation step and a solvent recovery cycle step,

in an epoxidation reaction product separation process, separating the epoxidation reaction product stream to obtain a propylene oxide product and a methanol stream containing methanol and water, wherein the epoxidation reaction product is subjected to a first separation in a first distillation column to obtain a light stream and a heavy stream, the light stream contains propylene oxide, propylene and a part of methanol, the heavy stream contains water and the rest of methanol, the light stream is subjected to a second separation in a second distillation column to obtain a gas phase purge stream containing propylene and a propylene oxide stream containing propylene oxide and methanol, the gas phase purge stream is contacted with an absorbent to obtain a gas phase stream containing propylene and a propylene oxide absorption stream containing absorbent and propylene oxide, the propylene oxide stream containing propylene oxide and methanol is subjected to an extractive distillation in an extractive distillation zone of an extractive distillation column to obtain a propylene oxide product, a liquid phase stream containing methanol and an extractive distillation in a bottom of the extractive distillation column is recovered, a part of the liquid phase stream is subjected to a third separation in a third distillation column to obtain an intermediate liquid phase purge stream containing propylene and a rest of the liquid phase stream from the top of the third distillation column, and a liquid phase stream containing impurities in the intermediate phase stream is obtained from the bottom of the third distillation column is purified to obtain a liquid phase stream containing impurities in the intermediate phase stream containing the intermediate stream and the rest of the purified impurities;

in the ethanol removal step, at least part of the methanol refined stream is rectified in an ethanol rectifying tower, recovered methanol with reduced ethanol content is obtained from the top of the ethanol rectifying tower, and a tower bottom stream containing ethanol is obtained from the tower bottom of the ethanol rectifying tower, and the operating conditions of the ethanol rectifying tower comprise: the pressure at the top of the tower is 0.01-0.5MPa, the reflux ratio is 1-5, the temperature at the top of the tower is 60-85 ℃, the temperature at the bottom of the tower is 90-120 ℃, and the pressure at the top of the tower is gauge pressure;

In the methanol recycling step, at least part of recovered methanol, which is the recovered methanol with the reduced ethanol content or the recovered methanol with the reduced ethanol content and the remaining part of the methanol refined stream, is recycled to the epoxidation reaction process.

2. The process of claim 1, wherein the ethanol content of the recovered methanol is less than 4 wt.% based on the total amount of recovered methanol.

3. The method according to claim 1, wherein the content of ethanol in the recovered methanol is 1% by weight or less based on the total amount of the recovered methanol.

4. The method according to claim 1, wherein the content of ethanol in the recovered methanol is 0.1 wt% or less based on the total amount of the recovered methanol.

5. The method of claim 1, wherein the reflux ratio of the ethanol rectification column is 2.5-4.

6. The method of claim 1, wherein the reflux ratio of the ethanol rectification column is 3-4.

7. The method according to any one of claims 1, 5 and 6, wherein the theoretical plate number of the ethanol rectification column is 20-65.

8. The process of claim 1, wherein the hydrogenating step further comprises treating the vapor phase hydrogenation stream to produce a treated stream having a reduced carbon monoxide content, and using at least the treated stream as recycle hydrogen for the hydrogenating step.

9. The method of claim 8, wherein the method of treating the vapor phase hydrogenation stream comprises: contacting the gas phase hydrogenation stream with a methanation catalyst under methanation reaction conditions to obtain the treat stream.

10. The process according to claim 9, wherein the gas phase hydrogenation stream is contacted with the methanation catalyst to such an extent that the carbon monoxide content by weight in the treatment stream is below 5 ppm.

11. The process according to claim 9, wherein the gas phase hydrogenation stream is contacted with the methanation catalyst to such an extent that the carbon monoxide content by weight in the treatment stream is below 3 ppm.

12. The process according to claim 9, wherein the gas phase hydrogenation stream is contacted with the methanation catalyst to such an extent that the carbon monoxide content by weight in the treatment stream is below 0.5 ppm.

13. The process of any of claims 9-12, wherein the gas phase hydrogenation stream is contacted with the methanation catalyst at a temperature of from 70 ℃ to 250 ℃.

14. The process of any of claims 9-12, wherein the gas phase hydrogenation stream is contacted with the methanation catalyst at a temperature of from 100 to 190 ℃.

15. The process of any of claims 9-12, wherein the gas phase hydrogenation stream is contacted with the methanation catalyst at a temperature of 135-160 ℃.

16. The process of any of claims 9-12, wherein the gas phase hydrogenation stream is contacted with the methanation catalyst at a pressure of from 0.5 to 10MPa, the pressure being gauge pressure.

17. The process of any of claims 9-12, wherein the gas phase hydrogenation stream is contacted with the methanation catalyst at a pressure of from 1 to 8MPa, the pressure being gauge pressure.

18. The process of any of claims 9-12, wherein the gas phase hydrogenation stream is contacted with the methanation catalyst at a pressure of from 2 to 6MPa, the pressure being gauge pressure.

19. The process according to any one of claims 9-12, wherein the gas phase hydrogenation stream is contacted with the methanation catalyst in a fixed bed reactor.

20. The process of claim 19, wherein the fixed bed reactor has a gas hourly space velocity of 500-10000h ^-1 。

21. The process of claim 19, wherein the fixed bed reactor has a gas hourly space velocity of 2000-8000h ^-1 。

22. The process of claim 19, wherein the fixed bed reactor has a gas hourly space velocity of 4000-6000h ^-1 。

23. The process according to any one of claims 9-12, wherein the methanation catalyst contains at least one catalytically active component selected from group VIII metals and group IB metals.

24. The method of claim 23, wherein the catalytically active component is one or more of ruthenium, rhodium, palladium, platinum, silver, iridium, iron, copper, nickel, and cobalt.

25. The method of claim 23, wherein the catalytically active component is nickel.

26. The process according to claim 23, wherein the methanation catalyst contains a carrier for supporting the catalytically active component.

27. The method of claim 26, wherein the support is a refractory inorganic oxide.

28. The method of claim 26, wherein the support is one or more of silica, titania, zirconia, and alumina.

29. The process according to claim 23, wherein the content of the catalytically active component in elemental terms is from 2 to 70 wt.%, based on the total amount of the methanation catalyst.

30. The process according to claim 23, wherein the content of the catalytically active component in elemental terms is from 20 to 60 wt.%, based on the total amount of the methanation catalyst.

31. The process according to claim 23, wherein the content of the catalytically active component in elemental terms is from 30 to 50 wt.%, based on the total amount of the methanation catalyst.

32. The process of any one of claims 1 and 8-12, wherein in the hydrotreating step, the hydrotreating conditions comprise: the temperature is 50-175 ℃, the pressure is 0.5-10MPa, and the pressure is gauge pressure; the contacting is performed in a fixed bed reactor.

33. The method of claim 32, wherein in the hydrogenating step, the hydrotreating conditions include: the temperature is 60-145 ℃, the pressure is 1-6MPa, and the pressure is gauge pressure.

34. The method of claim 32, wherein in the hydrogenating step, the hydrotreating conditions include: the temperature is 70-125 ℃, the pressure is 2-5.5MPa, and the pressure is gauge pressure.

35. The process of claim 32, wherein the fixed bed reactor has a liquid hourly space velocity of from 0.5 to 30 hours ^-1 。

36. The method of claim 32, wherein the fixed bed reactor has a liquid hourly space velocity of from 2 to 25 hours ^-1 。

37. The method of claim 32, wherein the fixed bed reactor has a liquid hourly space velocity of 10-15h ^-1 。

38. The process according to any one of claims 1 and 8 to 12, wherein the catalyst with hydrogenation catalysis contains at least one catalytically active component selected from group VIII metals and group IB metals.

39. The method of claim 38, wherein the catalytically active component is one or more of ruthenium, rhodium, palladium, platinum, silver, iridium, iron, copper, nickel, and cobalt.

40. The method of claim 38, wherein the catalytically active component is nickel.

41. The method of claim 38, wherein the catalyst having a hydrogenation catalytic effect comprises a support for supporting the catalytically active component.

42. The method of claim 41, wherein the support is a refractory inorganic oxide.

43. The method of claim 42, wherein the support is one or more of silica, titania, zirconia, and alumina.

44. The process according to claim 38, wherein the catalytically active component is present in an amount of from 2 to 70% by weight, calculated on an elemental basis, based on the total amount of the catalyst with hydrogenation catalysis.

45. The process according to claim 38, wherein the catalytically active component is present in an amount of 20 to 60% by weight, calculated as element, based on the total amount of the catalyst having hydrogenation catalysis.

46. The process according to claim 38, wherein the catalytically active component is present in an amount of from 30 to 55% by weight, calculated as element, based on the total amount of the catalyst having hydrogenation catalysis.

47. The method according to claim 1, wherein in the methanol refining step, the separation is performed in a methanol rectifying column including a first rectifying column and a second rectifying column,

subjecting the liquid phase hydrogenation stream, or the liquid phase hydrogenation stream and the remaining part of the methanol stream, to a first rectification in a first rectification column at a first rectification pressure, obtaining a first distillate containing methanol from the top of the column, obtaining a first bottom stream containing methanol and water from the bottom of the column;

48. The method of claim 47, wherein the first rectification pressure is from 0.01 to 0.5MPa and the second rectification pressure is from 0.5 to 1.2MPa, expressed as gauge.

49. The process according to claim 47, wherein the first rectifying column has a top temperature of 70 to 120 ℃ and a reflux ratio of 0.5 to 2.

50. The process of claim 47, wherein the second rectification column has a top temperature of 100 to 150 ℃ and a reflux ratio of 0.5 to 3.

51. The method according to claim 1, wherein in the methanol refining step, the separation is performed in a methanol rectifying column including a third rectifying column and a fourth rectifying column,

subjecting the liquid phase hydrogenation stream, or the liquid phase hydrogenation stream and the remaining part of the methanol stream, to a third rectification in a third rectification column at a third rectification pressure, obtaining a third distillate containing methanol from the top of the column, obtaining a third bottom stream containing methanol and water from the bottom of the column;

52. The method of claim 51, wherein the third rectification pressure is from 1 to 2MPa and the fourth rectification pressure is from 0.01 to 0.5MPa, expressed as gauge pressure.

53. The process of claim 51, wherein the third rectification column has a top temperature of 140-180 ℃ and a reflux ratio of 0.6-1.8; the temperature of the top of the fourth rectifying tower is 70-120 ℃, and the reflux ratio is 0.5-2.

54. The process of claim 1, wherein the recovered methanol has an intermediate impurity content of less than 0.4 wt.% based on the total amount of recovered methanol, the intermediate impurity having a boiling point higher than propylene oxide and lower than methanol.

55. The process of claim 1, wherein the recovered methanol has an intermediate impurity content of not more than 0.2 wt.% based on the total amount of recovered methanol, the intermediate impurity having a boiling point higher than propylene oxide and lower than methanol.

56. The process of claim 1, wherein the recovered methanol has an intermediate impurity content of not more than 0.05 wt.% based on the total amount of recovered methanol, the intermediate impurity having a boiling point higher than propylene oxide and lower than methanol.

57. The process of claim 1, wherein the heavies stream is hydrotreated in the hydrogenation step to provide a liquid phase hydrogenation stream comprising methanol, and the liquid phase hydrogenation stream, the purified liquid phase stream, and a remaining portion of the liquid phase stream are separated in the methanol refining step.

58. The process according to claim 1, wherein the top pressure of the third distillation column is 0.01 to 0.5MPa, the top temperature of the third distillation column is 50 to 75 ℃, the reflux ratio of the third distillation column is 50 to 300, and the top pressure is gauge pressure.

59. The process according to claim 1, wherein the theoretical plate number of the third distillation column is from 30 to 70.

60. The method of claim 1, wherein the extractant is water.

61. The process of claim 1, wherein the extractive distillation is carried out in the presence of at least one amino-containing compound.

62. The method of claim 61, wherein the amino group-containing compound is hydrazine.

63. A process as set forth in claim 61 wherein at least a portion of the amino-containing compound is added to the extractive distillation zone at a location not lower than the extractant.

64. The process of claim 61 wherein at least a portion of the amino-containing compound is added to the extractive distillation zone at the same location as the extractant.

65. A process as set forth in claim 61 wherein at least a portion of the amino-containing compound is added to the extractive distillation zone at a location above the extractant.

66. The process of claim 61 wherein at least a portion of the amino-containing compound is added to the extractive distillation zone at the same location as the propylene oxide stream.

67. A process according to claim 61 wherein the amino group-containing compound is present in an amount of from 0.3 to 0.8% by weight, based on the total amount of extractant.

68. The process of claim 1, wherein the extractant is added to the extractive distillation zone at a location above the propylene oxide stream.

69. The process of claim 1, wherein the theoretical plate number of the extractive distillation column is T, the propylene oxide stream is fed at a location of (0.6-0.9) x T, and the extractant is fed at a location of (0.15-0.45) x T.

70. The process of claim 69 wherein the theoretical plate number of the extractive distillation column is T, the propylene oxide stream is fed at a location of (0.6-0.85) x T, and the extractant is fed at a location of (0.2-0.4) x T.

71. The process according to claim 1, wherein the extractive distillation column has a theoretical plate number of from 35 to 90.

72. The process according to claim 1, wherein the extractive distillation column has a column top pressure of 0.01-0.5MPa, a column top temperature of 30-45 ℃, a column bottom temperature of 70-110 ℃, a reflux ratio of 1-10, and the column top pressure is gauge pressure.

73. The process of claim 1, wherein the propylene oxide stream is pretreated upstream of the extractive distillation column with one or both of the following based on the flow direction of the propylene oxide stream:

in one aspect, the propylene oxide stream is contacted with at least one alkaline material;

and secondly, rectifying the propylene oxide stream in a light component removal rectifying tower, wherein the pretreated propylene oxide stream is obtained from the bottom of the light component removal rectifying tower, the tower top pressure of the light component removal rectifying tower is 0.01-0.5MPa, the tower top temperature is 30-45 ℃, the reflux ratio is 100-300, the tower top pressure is gauge pressure, and the theoretical plate number of the light component removal rectifying tower is 30-70.

74. The method according to claim 73, wherein in the first mode, the basic substance is a basic ion exchange resin and/or an aqueous solution containing a basic substance.

75. The method according to claim 1, wherein in the epoxidation reaction process, methanol: propylene: the molar ratio of the hydrogen peroxide is 4-20:1.2-10:1.

76. the method of claim 75, wherein in the epoxidation reaction process, methanol: propylene: the molar ratio of the hydrogen peroxide is 6-12:1.2-5:1.

77. the process of claim 1 wherein the epoxidation reaction is carried out in a fixed bed reactor into which a feed stream comprising propylene, hydrogen peroxide and methanol enters from the bottom of the fixed bed reactor and an epoxidation reaction product stream is recovered from the top of the fixed bed reactor.

78. The method of claim 77, wherein said fixed bed reactor is a tubular fixed bed reactor.

79. The process of any one of claims 1 and 75-78, wherein in the epoxidation reaction step, the epoxidation reaction is carried out at a temperature of from 20 to 80 ℃.

80. The method of any one of claims 1 and 75-78, wherein the epoxidation catalyst is a titanium-containing molecular sieve.

81. The method of claim 80, wherein the epoxidation catalyst is a titanium silicalite.

82. The method of claim 80, wherein the epoxidation catalyst is a titanium silicalite having an MFI structure.

83. The process recited in claim 80, wherein the epoxidation catalyst is a hollow titanium silicalite molecular sieve, the crystallites of the hollow titanium silicalite molecular sieve are hollow structures having a radial length of the hollow portion of 5-300nm, and the titanium silicalite molecular sieve is at 25 ℃, P/P ₀ The benzene adsorption amount measured under the conditions of=0.10 and adsorption time of 1h is at least 70mg/g, and a hysteresis loop exists between an adsorption isotherm and a desorption isotherm of low-temperature nitrogen adsorption of the titanium silicalite molecular sieve.

84. An epoxidation system comprises an epoxidation reaction unit, an epoxidation reaction product separation unit and a solvent recovery circulation unit, wherein an epoxidation product stream output port of the epoxidation reaction unit is communicated with a material input port to be separated of the epoxidation reaction product separation unit,

the epoxidation reaction product separation unit comprises a reaction product separation subunit, a propylene oxide refining subunit and an intermediate impurity removal subunit, wherein the reaction product separation subunit is used for separating an epoxidation reaction product stream into a light stream and a heavy stream, the light stream comprises propylene oxide, propylene and part of methanol, the heavy stream comprises water and the rest part of methanol, the propylene separation subunit is used for separating at least part of propylene from the heavy stream, the propylene oxide refining subunit is used for separating at least part of propylene oxide from the heavy stream, the intermediate impurity separation subunit is used for removing intermediate impurities in residual liquid from which propylene oxide and propylene are separated, and the residual liquid from which the intermediate impurities are separated is sent to the methanol refining subunit,

The material input port to be recovered of the solvent recovery circulation unit is communicated with the recovered methanol output port of the epoxidation reaction product separation unit, the solvent recovery circulation unit comprises a hydrogenation subunit, a methanol refining subunit and a methanol circulation subunit,

the method is characterized in that the solvent recovery circulation unit further comprises an ethanol removal subunit for removing at least part of ethanol in a methanol refined stream output by the methanol refining subunit to obtain recovered methanol, and sending the recovered methanol into the epoxidation reaction unit through the methanol circulation subunit, wherein at least part of the methanol refined stream is rectified in an ethanol rectifying tower, recovered methanol with reduced ethanol content is obtained from the top of the ethanol rectifying tower, and a bottom stream containing ethanol is obtained from the bottom of the ethanol rectifying tower, and the operation conditions of the ethanol rectifying tower comprise: the pressure at the top of the tower is 0.01-0.5MPa, the reflux ratio is 1-5, the temperature at the top of the tower is 60-85 ℃, the temperature at the bottom of the tower is 90-120 ℃, and the pressure at the top of the tower is gauge pressure.

85. The system of claim 84, wherein the propylene oxide refining subunit comprises an extractive distillation column for contacting a propylene oxide stream comprising propylene oxide and methanol with an extractant in an extractive distillation zone of the extractive distillation column for extractive distillation to obtain a propylene oxide product and a liquid phase stream comprising methanol and extractant, and a third distillation column for subjecting the liquid phase stream comprising methanol, intermediate impurities and extractant to a third distillation to obtain a distillate comprising intermediate impurities from the top of the third distillation column and a purified liquid phase stream from the bottom of the third distillation column.

86. The system claimed in claim 84 and wherein said hydrogenation sub-unit comprises a hydrogenation reactor for contacting at least a portion of the methanol stream with a catalyst having a hydrogenation catalytic effect under hydrogenation conditions to produce a hydrogenated product stream having a reduced impurity content, a hydrogenation product separation device having a to-be-separated stream input port in communication with a hydrogenation product stream output line of said hydrogenation reactor for separating said hydrogenated product stream to produce a gaseous hydrogenated stream comprising hydrogen and a liquid hydrogenated stream comprising alcohol, and a gas phase hydrogenation stream processing device comprising a processor for reducing the carbon monoxide content for processing the gaseous hydrogenated stream output from the hydrogenation product separation device to produce a processed stream having a reduced carbon monoxide content and feeding at least the processed stream into the hydrogenation reactor.

87. The system of claim 86, wherein the processor for reducing the carbon monoxide content is a methanation reactor.

88. The system of claim 84, wherein the methanol rectification subunit comprises a first rectification column for subjecting the liquid phase hydrogenation stream, or the liquid phase hydrogenation stream and the non-hydrotreated methanol stream, to a first rectification at a first rectification pressure, obtaining a first distillate comprising alcohol from a top of the first rectification column, and obtaining a first bottoms stream comprising alcohol and water from a bottom of the first rectification column;

the second rectifying pressure is higher than the first rectifying pressure, and the output port of the gas phase hydrogenation material flow at the top of the second rectifying tower is communicated with the heat source input port of the reboiler of the first rectifying tower.

89. The system of claim 84, wherein the methanol rectification subunit comprises a third rectification column and a fourth rectification column,

and the fourth rectifying pressure is higher than the third rectifying pressure, and the output port of the gas phase hydrogenation material flow at the top of the third rectifying tower is communicated with the heat source input port of the reboiler of the third rectifying tower.