CN115504884A

CN115504884A - Preparation method of methyl propionate and obtained methyl propionate

Info

Publication number: CN115504884A
Application number: CN202110698739.4A
Authority: CN
Inventors: 王海之; 刘仲能; 余强; 刘晓曦
Original assignee: China Petroleum and Chemical Corp; Sinopec Shanghai Research Institute of Petrochemical Technology
Current assignee: China Petroleum and Chemical Corp; Sinopec Shanghai Research Institute of Petrochemical Technology
Priority date: 2021-06-23
Filing date: 2021-06-23
Publication date: 2022-12-23

Abstract

The invention discloses a preparation method of methyl propionate and the obtained methyl propionate, wherein the method comprises the following steps of (1) reacting raw materials including methyl acetate, methanol and an aldehyde source in a hydrogen atmosphere to obtain an intermediate product; (2) And carrying out hydrotreatment on the intermediate product under a hydrogen atmosphere to obtain methyl propionate. In the step (1), the raw materials sequentially pass through a first catalyst bed layer and a second catalyst bed layer which are connected in series. Cooling or heat exchange treatment is carried out on the intermediate product before the step (2), gas-liquid separation is carried out on the product after the hydrotreatment in the step (2), and the separated hydrogen is treated and recycled to the step (1). The method adopts hydrogen atmosphere in the whole process, and the problem of separation of nitrogen and hydrogen does not exist in the later stage, so that the energy consumption is saved, and the production cost is greatly reduced; meanwhile, the hydrogenation treatment is carried out in a gas phase, so that the self-polymerization of the methyl acrylate can be effectively reduced or avoided.

Description

Preparation method of methyl propionate and obtained methyl propionate

Technical Field

The invention relates to preparation of methyl acrylate, in particular to a method for preparing methyl acrylate and the obtained methyl acrylate.

Background

Methyl acrylate is an important fine chemical raw material with wide application, is mainly used for organic synthesis intermediates and high molecular monomers, and polymers prepared by taking methyl acrylate as monomers are widely used in the industries of coatings, textiles, leathers, adhesives and the like.

The acrylic acid and its ester are produced mainly by the propylene oxidation method, the acrylonitrile hydrolysis method, the vinyl ketone method, the propane oxidation method, the methyl formate method, and the like. However, the methods have the defects of serious pollution, high energy consumption, low product yield and the like. Therefore, the development of a new green and efficient production process has very important significance.

The well established alpha-MMA process is proposed by Lucite, inc., which uses methyl methacrylate prepared from methyl propionate and formaldehyde. However, the process route of the Lucite company is not suitable for the national conditions of China. The national condition of China is that the production capacity of methyl acetate is greatly surplus. However, the process route of Lu Cai Te company is to synthesize the raw material of methyl methacrylate from methyl propionate, and the consumed raw material is methyl propionate, so that the problem of surplus methyl acetate production in China cannot be solved. Therefore, what is needed in the art is to realize green synthesis of methyl acrylate by using an industrial byproduct, namely methyl acetate, as a raw material, using a safe, environment-friendly and nontoxic solid base catalyst and using a clean synthesis process.

In view of this, in order to solve the problem of the large surplus of the productivity of methyl acetate in China, the synthesis of methyl acrylate by using methyl acetate and formaldehyde as raw materials and adopting a novel catalyst is proposed on the basis of the mature alpha-MMA technology proposed by Lucite (Lucite) company at present. The catalyst reaction-regeneration fluidized bed coupling system is adopted, and the problems of short one-way service life, easy carbon deposition and frequent regeneration of the catalyst are solved. At present, the conversion per pass of methyl acetate is generally 15%, the selectivity of methyl acrylate can generally reach 85%, and the yield of methyl acrylate is generally 13.5%. But at present, industrialization still cannot be realized, and the key point is that the comprehensive improvement of yield, selectivity and catalyst stability cannot be achieved. In recent years, many units at home and abroad further research on the synthesis of methyl acrylate from methyl acetate and formaldehyde.

Chinese patent No. (CN 104525176) reports a preparation method of a Cs-based catalyst, microspheres prepared thereby, and a method for synthesizing methyl (meth) acrylate using the microspheres. Specifically, cs salt, auxiliary metal M salt, an optional linking agent and a template agent are mixed with 20-40% silica sol to prepare a solution, then the solution is ground in the colloid at the speed of 5000-10000 rpm for 1-5 minutes, microspheres of 20-220 micrometers are obtained by spraying through a spray dryer, and the condensation catalyst is obtained by drying at 70-120 ℃ and roasting at 200-600 ℃ for 2-7 hours.

Chinese patent No. (CN 103435483A) discloses a process for preparing MA from methyl acetate and formaldehyde using a fixed fluidized bed process. However, the existing catalyst has many problems in the fixed fluidized bed: 1. the ions are not the A-type particles, so that the process amplification danger of the fluidized bed is increased, and the industrialization is not facilitated; 2. the ion bulk density is small and is not suitable for the reaction process of a fixed fluidized bed; 3. other factors such as sphericity, activity, amount of carbon deposition, attrition, etc. do not suffice for fixed bed processes. The problem of the use of methyl propionate and formaldehyde in the same fluid bed process for the preparation of methyl methacrylate also exists with the problem of the use of a leishi catalyst system suitable for use in a fluid bed process and meeting the requirements of the fluid bed.

In conclusion, the research on the synthesis of methyl acrylate by using methyl acetate and formaldehyde as raw materials is not industrialized, and is only in the research stage at present, but the cost advantage is very obvious as a new process route, and the process is mainly developed by large-scale transnational chemical companies by investing manpower and material resources.

Therefore, in view of the above, there is a need in the art for a synthetic route of methyl acrylate and a corresponding solid base catalyst, which can improve yield and selectivity, and simultaneously maintain catalytic activity for a long time, thereby realizing industrialization and solving the problem of large surplus of methyl acetate production.

Disclosure of Invention

In order to overcome the problems in the prior art, the invention provides a preparation method of methyl propionate and the obtained methyl propionate, wherein the method comprises a condensation reaction and a hydrogenation reaction under a hydrogen atmosphere, and the condensation reaction is carried out by adopting a two-stage condensation process. The method can improve the yield and the selectivity, and simultaneously can keep the catalytic activity for a long time, thereby realizing industrialization and solving the problem of greatly surplus methyl acetate productivity.

One of the purposes of the invention is to provide a preparation method of methyl propionate, which comprises the following steps:

(1) Reacting raw materials including methyl acetate, methanol and an aldehyde source in a hydrogen atmosphere to obtain an intermediate product;

(2) And carrying out hydrotreatment on the intermediate product under a hydrogen atmosphere to obtain methyl propionate.

The whole process of the method is carried out in the hydrogen atmosphere, and the method is different from the technical scheme that condensation is carried out under nitrogen and then hydrogenation is carried out under hydrogen in the prior art, and the method has the advantages that the problem of separation of nitrogen and hydrogen does not exist in the later stage, so that the energy consumption is saved, and the production cost is greatly reduced.

In a preferred embodiment, the aldehyde source is selected from at least one of trioxymethylene, paraformaldehyde, methylal, and anhydrous formaldehyde; and/or the methyl acetate is selected from crude methyl acetate and/or refined methyl acetate.

Wherein, crude methyl acetate and refined methyl acetate can be obtained commercially.

In a preferred embodiment, in step (1), the raw material passes through a first catalyst bed layer and a second catalyst bed layer which are connected in series in sequence, and an intermediate product is obtained through reaction.

In a preferred embodiment, the molar ratio of methyl acetate to the aldehyde source in the feed to the first catalytic bed is (1-10: 1, preferably (2-5: 1, for example 2.

In a preferred embodiment, the weight ratio of methanol to methyl acetate in the feed to the first catalyst bed is (0-0.5): 1, preferably (0.2-0.5): 1.

In a preferred embodiment, the aldehyde source is replenished at the top of the second catalyst bed.

Wherein the supplemented aldehyde source is supplemented in a liquid state, preferably, when the supplemented aldehyde source of the second catalyst bed is trioxymethylene, solid trioxymethylene is heated until the trioxymethylene is in a liquid phase state (preferably, heated at 60-80 ℃), and then introduced into the second catalyst bed through a pump.

In a preferred embodiment, the reaction temperature of the first catalyst bed is 250 to 400 ℃; and/or the reaction pressure is 0.1-1 MPa; and/or the liquid phase volume flow rate is 0.01-1 mL/min; and/or the hydrogen flow rate is 20-150mL/min.

In a further preferred embodiment, the reaction temperature of the first catalyst bed is 300 to 360 ℃ (e.g., 300 ℃, 330 ℃, 340 ℃, 350 ℃, 360 ℃); and/or the reaction pressure is 0.1 to 0.5MPa (for example, 0.1MPa, 0.2MPa, 0.3MPa, 0.4MPa, 0.5 MPa); and/or the liquid phase volume flow rate is 0.1-0.5 mL/min (e.g., 0.1mL/min, 0.2mL/min, 0.3mL/min, 0.4mL/min, 0.5 mL/min); and/or the hydrogen flow rate is 50-150mL/min.

Wherein, the liquid phase volume flow rate on the first catalyst bed layer refers to the flow rate of liquid raw materials including methyl acetate and aldehyde source.

In a preferred embodiment, the reaction temperature of the second catalyst bed layer is 250-400 ℃; and/or the reaction pressure is 0.1-1 MPa; and/or the hydrogen flow is 50-300mL/min; and/or the volume flow rate of the liquid phase for supplementing the aldehyde source is 0-1mL/min.

For example, the hydrogen flow rate of the second catalyst bed is 50, 60, 70, 80, 90, 100, 120, 130, 140, 150, 160, 170, 180, 190, 200, 210, 220, 230, 240, 250, 260, 270, 280, 290, or 300mL/min.

In a further preferred embodiment, the reaction temperature of the second catalyst bed is 300 to 360 ℃ (e.g., 300 ℃, 330 ℃, 340 ℃, 350 ℃, 360 ℃); and/or the reaction pressure is 0.1 to 0.5MPa (for example, 0.1MPa, 0.2MPa, 0.3MPa, 0.4MPa, 0.5 MPa); and/or the hydrogen flow is 150-300mL/min; and/or the volume flow rate of the liquid phase of the supplemented aldehyde source is 0.02-0.3 mL/min.

Wherein the hydrogen flow in the second catalyst bed is controlled to be higher than that in the first catalyst bed.

In a preferred embodiment, a solid base catalyst loading section and an optional acid catalyst loading section are respectively and independently included in the first catalyst bed and the second catalyst bed, i.e. the solid base catalyst is optionally doped with a section of acid catalyst.

In the prior art, the aldehyde source concentration on the upper part of a bed layer is high and coked, the catalyst is inactivated, and the bed layer is easy to block, so that high-concentration methyl acetate and low-concentration aldehyde source in the lower bed layer are directly caused, and the high-concentration methyl acetate can be condensed to generate acetone, so that the separation energy consumption is increased.

In a preferred technical scheme, an acid catalyst filling section is doped in a solid alkali catalyst bed layer, and the acid catalyst can promote the decomposition of an aldehyde source, increase the concentration of the aldehyde source in a lower bed layer and provide enough formaldehyde, so that the generation of acetone is inhibited.

In a preferred embodiment, the solid base catalyst comprises a carrier I and an active component I and an optional auxiliary agent which are loaded on the carrier I.

In a further preferred embodiment, the support I is selected from at least one of silica, alumina and SBA-15 molecular sieves, preferably having a specific surface area of 50 to 500m ² The pore diameter is 6-30nm, and the porosity is 0.6-1mL/g.

In a further preferred embodiment, the active component I is at least one of cesium, potassium and rubidium (preferably cesium), preferably in a loading of 1 to 20wt%, preferably 2 to 10wt%, such as 2wt%, 3wt%, 4wt%, 5wt%, 6wt%, 7wt%, 8wt%, 9wt%, 10wt%.

The source of the active component I is at least one of carbonic acid compound, formic acid compound and nitric acid compound containing active component elements.

In a still further preferred embodiment, the auxiliary agent is selected from at least one of zirconium, bismuth and lanthanum compounds, preferably in a loading of 0 to 5wt%, preferably 0.3 to 3wt%, such as 0.5wt%, 1wt%, 2wt%, 3wt%.

Wherein the sources of zirconium, bismuth and nickel are zirconium-containing compounds, bismuth-containing compounds and lanthanum-containing compounds, such as zirconyl nitrate, bismuth chloride and lanthanum nitrate, respectively.

When the solid base catalyst is prepared, if equal-volume impregnation is adopted, as the components in the impregnation liquid are basically and completely loaded on the carrier I, the content of the active component I and the auxiliary agent in the product can be calculated according to the use amount of the raw materials by a theoretical calculation method.

In a preferred embodiment, the acidic catalyst is selected from alumina and/or molecular sieves.

In a further preferred embodiment, the acidic catalyst is selected from the group consisting of θ -Al ₂ O ₃ At least one of SAPO-34 molecular sieve and SAPO-35 molecular sieve.

In a preferred embodiment, when the first catalyst bed and the second catalyst bed comprise an acid catalyst loading section, the acid catalyst loading section is formed by loading 0.5-10% of bed volume of the acid catalyst at 1/3-3/4 position of the first catalyst bed, and the solid base catalyst loading section is formed by loading the rest positions of the first catalyst bed.

For example, 0.5%, 1%, 2%, 3%, 4%, 5%, 6%, 7%, 8%, 9%, or 10% of the bed volume of the acidic catalyst is loaded at 1/3, 1/2, 5/8, or 3/4 positions of the first catalyst bed.

In a further preferred embodiment, when the first catalyst bed and the second catalyst bed comprise an acid catalyst loading section, 1/2-5/8 position of the first catalyst bed is loaded with 1-4% bed volume of the acid catalyst to form the acid catalyst loading section, and the rest positions are loaded with the solid base catalyst to form the solid base catalyst loading section.

Wherein, the positions from the top to the bottom of the bed layer are 0-1 in sequence. Specifically, in the first catalyst bed, the aldehyde source at the upper part of the bed is sufficient and does not need an acid catalyst, so that the acid catalyst is filled only in the middle and lower sections of the bed to decompose the undeployed aldehyde source, and the problem of insufficient supply of the aldehyde source at the middle and lower parts of the bed is solved.

In a preferred embodiment, when the first catalyst bed and the second catalyst bed comprise an acid catalyst loading section, the acid catalyst loading section is formed by loading 0.5-10% of bed volume of the acid catalyst at 1/3-3/4 position of the second catalyst bed, and the solid base catalyst loading section is formed by loading the rest positions of the second catalyst bed.

For example, the first catalyst bed may be charged with 0.5%, 1%, 2%, 3%, 4%, 5%, 6%, 7%, 8%, 9%, or 10% of the bed volume of the acidic catalyst at 1/3, 1/2, 5/8, or 3/4 positions.

In a further preferred embodiment, when the first catalyst bed and the second catalyst bed comprise an acid catalyst loading section, 1/2-5/8 position of the second catalyst bed is respectively loaded with 1-4% of bed volume of the acid catalyst to form the acid catalyst loading section, and the rest positions are loaded with the solid base catalyst to form the solid base catalyst loading section.

Wherein, the positions from the top to the bottom of the bed layer are 0-1 in sequence. Specifically, a small amount of aldehyde source is introduced into the inlet of the second reactor, so that sufficient formaldehyde is arranged at the upper part of the reaction bed layer of the second reactor, but the aldehyde source is continuously reduced along with the reaction, the decomposed monomer formaldehyde is also continuously reduced, and the problem of insufficient formaldehyde exists in the lower bed layer, so that the acid catalyst is filled at the middle lower part of the second catalyst bed layer.

In the invention, the two beds connected in series are respectively arranged in the two reactors connected in series.

In a preferred embodiment, the hydrotreatment in step (2) is carried out in the presence of a hydrogenation catalyst comprising a support II and an active component II, wherein the BET specific surface area of the support II is from 20 to 200m ² The pore size is 2-30nm; preferably, the carrier II is at least one selected from alumina, silica, SAPO-34 and activated carbon, preferably silica and/or activated carbonAnd (4) carbon.

In a further preferred embodiment, the active component II is selected from at least one of Pd, ni and Cu, preferably 100wt% based on the hydrogenation catalyst, with an active component II loading of 0.1 to 10wt%, preferably 0.1 to 1wt%.

In a still further preferred embodiment, in step (2), in the hydrotreatment, the temperature is between 120 and 250 ℃, preferably between 120 and 150 ℃; and/or, the pressure is 0.1-1MPa, preferably 0.1-0.6MPa; and/or the hydrogen flow rate is 0 to 300mL/min, preferably 0 is excluded, preferably 0 to 200mL/min, preferably 0 is excluded.

In a preferred embodiment, the intermediate product is subjected to a temperature reduction treatment or a heat exchange treatment before step (2).

In a further preferred embodiment, the temperature of the intermediate product after the temperature reduction treatment or the heat exchange treatment is 80 to 250 ℃, preferably 100 to 150 ℃. For example, 90 ℃, 100 ℃,110 ℃, 120 ℃, 130 ℃, 140 ℃, 150 ℃.

Wherein the intermediate product obtained in step (1) has a relatively high temperature (above 300 ℃, for example 350 ℃), but the hydrogenation cannot be carried out at such a high temperature, otherwise the hydrogenation of the raw ester may occur. In the prior art, hydrogenation is generally carried out after the boiling point of MA is reduced below (i.e. MA liquid phase hydrogenation), but the intermediate product is methyl acrylate which is easy to self-polymerize at room temperature and block a reactor.

In view of the above problems, the present invention reduces the temperature or heat exchange of the intermediate product to 80 to 250 ℃, preferably 100 to 150 ℃, so that the methyl acrylate is in a gas phase in the temperature range, and the gas phase state increases the intermolecular distance, thereby effectively reducing or avoiding self-polymerization. On the other hand, compared with the method of reducing the temperature to the state that the MA is in a liquid phase, the method reduces the temperature to the state that the MA is in a gas phase, can obviously reduce the energy consumption, and is beneficial to industrial production and application.

In a preferred embodiment, the product is condensed and absorbed after the hydrotreatment in step (2), then subjected to gas-liquid separation, and the separated hydrogen is subjected to adsorption treatment with a basic compound and then recycled to step (1).

In a further preferred embodiment, the basic compound is selected from calcium oxide and/or calcium hydroxide, which is effective for removing carbon dioxide generated during the reaction, protecting the condensation catalyst in the condensation reactor, and recycling the gas after gas-liquid separation back to the condensation reactor.

The second object of the present invention is to provide methyl propionate obtainable by the process according to the first object of the present invention.

The endpoints of the ranges and any values disclosed in the present application are not limited to the precise range or value, and such ranges or values should be understood to encompass values close to those ranges or values. For ranges of values, between the endpoints of each of the ranges and the individual points, and between the individual points may be combined with each other to give one or more new ranges of values, and these ranges of values should be considered as specifically disclosed herein. In the following, various technical solutions can in principle be combined with each other to obtain new technical solutions, which should also be regarded as specifically disclosed herein.

Compared with the prior art, the invention has the following beneficial effects:

(1) The method adopts hydrogen atmosphere in the whole process, and does not have the problem of separation of nitrogen and hydrogen in the later period, thereby saving energy consumption and greatly reducing production cost;

(2) The composite catalyst bed layer can solve the problem of insufficient supply of aldehyde compounds in the lower bed layer, and can inhibit side reactions and reduce the content of acetone byproduct;

(3) The hydrotreatment is carried out in a gas phase, and the self-polymerization of the methyl acrylate can be effectively reduced or avoided.

Detailed Description

While the present invention will be described in detail with reference to the following examples, it should be understood that the following examples are illustrative of the present invention and are not to be construed as limiting the scope of the present invention.

It is to be noted that the various features described in the following detailed description may be combined in any suitable manner without contradiction. The invention is not described in detail in order to avoid unnecessary repetition.

In addition, any combination of the various embodiments of the present invention can be made, as long as the technical solution formed by the combination does not depart from the idea of the present invention, and the technical solution formed by the combination is part of the original disclosure of the present specification, and also falls into the protection scope of the present invention.

The raw materials used in the examples and comparative examples are disclosed in the prior art if not particularly limited, and may be, for example, directly purchased or prepared according to the preparation methods disclosed in the prior art.

The adopted solid base catalyst Cs/SiO ₂ The preparation was as follows: weighing 0.72gCs ₂ CO ₃ Dissolving the mixture in 12mL deionized water, and preparing Cs/SiO by adopting an isovolumetric method ₂ Standing the catalyst for 8 hours, drying the catalyst in a vacuum drying oven at 110 ℃ for 12 hours, and roasting the catalyst for 6 hours at 550 ℃ in air atmosphere to prepare the catalyst Cs/SiO ₂ 。

The adopted solid base catalyst Cs/Zr-SiO ₂ The preparation was as follows: weighing 0.14gZrOCl ₂ ·8H ₂ Dissolving O in 12mL deionized water, and preparing Zr-SiO by adopting an isovolumetric method ₂ Standing the catalyst for 8 hours, drying the catalyst in a vacuum drying oven at 110 ℃ for 12 hours, and roasting the catalyst for 6 hours at 550 ℃ in air atmosphere to prepare the catalyst Zr-SiO ₂ And (3) a carrier. Weighing 0.72gCs ₂ CO ₃ Dissolving in 12mL of deionized water, and adopting Zr-SiO ₂ As a carrier, preparing Cs/Zr-SiO by an isovolumetric method ₂ Standing the catalyst for 8 hours, drying the catalyst in a vacuum drying oven at 110 ℃ for 12 hours, and roasting the catalyst for 6 hours at 550 ℃ in air atmosphere to prepare the catalyst Cs/Zr-SiO ₂ 。

The hydrogenation catalyst Pd/C is obtained as follows: the Pd/C catalyst is prepared by adopting activated carbon as a carrier, palladium chloride as a precursor and 0.5wt% of loading capacity, adopting an isometric impregnation method to prepare the Pd/C catalyst, placing the catalyst at normal temperature for 8h after impregnation, drying the catalyst for 12h at 110 ℃, and roasting the catalyst for 5 h at 550 ℃ in an air atmosphere.

Hydrogenation catalyst Pd/SiO ₂ Is obtained as follows: adopting silicon dioxide as a carrier, palladium chloride as a precursor and 0.5wt% of loading capacity, and adopting an isometric impregnation method to prepare Pd/SiO ₂ And after the impregnation, placing the catalyst at room temperature for 8h, drying the catalyst at 110 ℃ for 12h, and roasting the catalyst at 550 ℃ for 5 h in an air atmosphere to prepare the Pd/SiO2 catalyst.

Sampling, adding internal standard toluene, and measuring the concentration of Methyl Propionate (MP) and Methyl Acrylate (MA) by gas chromatography.

[ example 1 ]

Respectively adding 10mL of solid base catalyst Cs/SiO into the first catalyst bed layer and the second catalyst bed layer ₂ . The molar ratio of methyl acetate to trioxymethylene in the feed to the first catalyst bed layer is 3.

The hydrogen flow in the first catalyst bed layer is 110mL/min, the liquid phase flow rate of the mixture of methyl acetate, methanol and trioxymethylene in the first catalyst bed layer is 0.2mL/min, the reaction pressure of the first catalyst bed layer is 0.6MPa, and the reaction temperature is 350 ℃;

the hydrogen flow in the second catalyst bed layer is 120mL/min, the liquid phase flow rate for replenishing trioxymethylene in the second catalyst bed layer is 0.05mL/min, the reaction pressure of the second catalyst bed layer is 0.6MPa, and the reaction temperature is 350 ℃.

Cooling gas at the outlet of the second catalyst bed layer to 120 ℃ through a heat exchanger, enabling product gas to enter a hydrogenation reactor, enabling a catalyst in the hydrogenation reactor to be Pd/C, enabling the hydrogenation temperature to be 120 ℃, the hydrogenation pressure to be 0.6MPa and the hydrogen flow to be 140mL/min, enabling the product gas to enter a gas-liquid separator after hydrogenation, enabling liquid to be condensed and absorbed, enabling the gas to be recycled to the first catalyst bed layer after calcium oxide absorption, enabling the conversion rate of Methyl Acetate (MAC) to be 16% and enabling the product concentration of Methyl Propionate (MP) to be 11%.

[ example 2 ] A method for producing a polycarbonate

Adding into the first catalyst bed layer and the second catalyst bed layer respectively10mL of solid base catalyst Cs/SiO ₂ . The molar ratio of methyl acetate to trioxymethylene in the feed to the first catalyst bed layer is 3.

the hydrogen flow in the second catalyst bed layer is 120mL/min, the liquid phase flow rate for supplementing trioxymethylene in the second catalyst bed layer is 0.05mL/min, the reaction pressure of the second catalyst bed layer is 0.6MPa, and the reaction temperature is 350 ℃.

Cooling the gas at the outlet of the second catalyst bed layer to 120 ℃ through a heat exchanger, and feeding the product gas into a hydrogenation reactor, wherein the catalyst in the hydrogenation reactor is Pd/SiO ₂ The hydrogenation temperature is 120 ℃, the hydrogenation pressure is 0.6MPa, the hydrogen flow is 140mL/min, the hydrogenation is carried out, the hydrogenation enters a gas-liquid separator, the liquid is condensed and absorbed, the gas is recycled to the first catalyst bed after being absorbed by calcium oxide, the conversion rate of Methyl Acetate (MAC) is 17.1%, the product concentration of Methyl Propionate (MP) is 11.3%, and the product concentration of MA is 1.3%.

[ example 3 ]

Respectively adding 10mL of solid base catalyst Cs/SiO into the first catalyst bed layer and the second catalyst bed layer ₂ And filling 1% bed volume of the acidic catalyst SAPO-34 molecular sieve at the 1/3 position of the first catalyst bed to form the acidic catalyst filling section, wherein the molar ratio of methyl acetate, methanol and trioxymethylene in the feed of the first catalyst bed is 3.

The hydrogen flow in the first catalyst bed layer is 100mL/min, the liquid phase flow rate of the mixture of methyl acetate and trioxymethylene in the first catalyst bed layer is 0.2mL/min, the reaction pressure of the first catalyst bed layer is 0.6MPa, and the reaction temperature is 350 ℃;

And cooling the gas at the outlet of the second catalyst bed layer to 120 ℃ through a heat exchanger, allowing the product gas to enter a hydrogenation reactor, allowing the catalyst in the hydrogenation reactor to be Pd/C, allowing the product gas to enter a gas-liquid separator after hydrogenation, allowing the liquid to be condensed and absorbed, allowing the gas to be absorbed by calcium oxide and then circulating back to the first catalyst bed layer, wherein the conversion rate of Methyl Acetate (MAC) is 18.8%, and the product concentration of Methyl Propionate (MP) is 13.9%.

[ example 4 ]

Respectively adding 10mL of solid base catalyst Cs/SiO into the first catalyst bed layer and the second catalyst bed layer ₂ And filling 1% bed volume of the acidic catalyst SAPO-34 molecular sieve at the 1/3 position of the first catalyst bed to form the acidic catalyst filling section, wherein the molar ratio of methyl acetate to trioxymethylene in the feed of the first catalyst bed is 3.

the hydrogen flow in the second catalyst bed layer is 110mL/min, the liquid phase flow rate for supplementing trioxymethylene in the second catalyst bed layer is 0.05mL/min, the reaction pressure of the second catalyst bed layer is 0.6MPa, and the reaction temperature is 350 ℃.

Cooling the gas at the outlet of the second catalyst bed layer to 120 ℃ through a heat exchanger, and feeding the product gas into a hydrogenation reactor, wherein the catalyst in the hydrogenation reactor is Pd/SiO ₂ After hydrogenation, the mixture enters a gas-liquid separator, liquid is condensed and absorbed, gas is absorbed by calcium oxide and then recycled to the first catalyst bed layer, the conversion rate of Methyl Acetate (MAC) is 19.8%, the product concentration of Methyl Propionate (MP) is 13.1%, and the product concentration of MA is 1.8%.

[ example 5 ]

Adding into the first catalyst bed layer and the second catalyst bed layer respectivelyAdding 10mL of solid base catalyst Cs/Zr-SiO ₂ . The molar ratio of methyl acetate to trioxymethylene in the feed to the first catalyst bed was 3.

Cooling gas at the outlet of the second catalyst bed layer to 120 ℃ through a heat exchanger, enabling product gas to enter a hydrogenation reactor, enabling a catalyst in the hydrogenation reactor to be Pd/C, enabling the hydrogenation temperature to be 120 ℃, the hydrogenation pressure to be 0.6MPa and the hydrogen flow to be 140mL/min, enabling the product gas to enter a gas-liquid separator after hydrogenation, enabling liquid to be condensed and absorbed, enabling the gas to be recycled to the first catalyst bed layer after calcium oxide absorption, enabling the conversion rate of Methyl Acetate (MAC) to be 17.9% and enabling the product concentration of Methyl Propionate (MP) to be 13.2%.

[ examples 6 to 9 ]

The procedure of example 1 was repeated except that the outlet gas of the second catalyst bed was cooled to a temperature of 80 c, 100 c, 150 c, 200 c, respectively, by passing through a heat exchanger, and the concentration of the Methyl Propionate (MP) product was increased as compared to the case where the temperature reduction treatment was not performed.

Comparative example 1

The procedure of example 1 was repeated except that: and cooling the gas at the outlet of the second catalyst bed layer to room temperature through a heat exchanger, and then carrying out liquid-phase hydrogenation, wherein other conditions are unchanged.

And respectively adding 10mL of solid base catalyst Cs/SiO2 into the first catalyst bed layer and the second catalyst bed layer. The molar ratio of methyl acetate to trioxymethylene in the feed to the first catalyst bed was 3.

The hydrogen flow in the first catalyst bed layer is 110mL/min, the liquid phase flow rate of the mixture of the methyl acetate and the trioxymethylene in the first catalyst bed layer is 0.2mL/min, the reaction pressure of the first catalyst bed layer is 0.6MPa, and the reaction temperature is 350 ℃;

the hydrogen flow in the second catalyst bed layer is 130mL/min, the liquid phase flow rate for supplementing trioxymethylene in the second catalyst bed layer is 0.05mL/min, the reaction pressure of the second catalyst bed layer is 0.6MPa, and the reaction temperature is 350 ℃.

Cooling gas at the outlet of the second catalyst bed layer to room temperature through a heat exchanger, carrying out gas-liquid separation, enabling a liquid phase product to enter a hydrogenation reactor, enabling a catalyst in the hydrogenation reactor to be Pd/C, carrying out hydrogenation reaction, enabling the gas to be absorbed by calcium oxide and then to be recycled to the first catalyst bed layer, enabling the conversion rate of Methyl Acetate (MAC) to be 15% and the product concentration of Methyl Propionate (MP) to be 7%, carrying out gas-liquid separation in the second condensation reactor, polymerizing the separated methyl acrylate solution, and blocking the methyl acrylate solution to enter a hydrogenation reactor pipeline.

Comparative example 2

The procedure of example 1 was repeated except that: the hydrogen flow in the first catalyst bed layer and the second catalyst bed layer is 110mL/min, and other conditions are unchanged.

It was found that when the hydrogen flow rate of the second catalyst bed was almost equal to that of the first catalyst bed, the concentration of methyl propionate in the product was significantly reduced after five days of operation because the catalyst stability was affected due to the long-term contact of the second catalyst bed with water, the content of methyl acrylate produced by condensation was reduced, and consequently the content of methyl propionate in the hydrogenation product was reduced.

Comparative example 3

The procedure of example 3 was repeated except that the amount of acidic catalyst in the bed was larger, 15% of the bed volume of acidic catalyst, and the other conditions were not changed.

It was found that after five days of operation, the methyl propionate content in the product was significantly reduced, only by half of the original, the analytical reasons being: the excessive amount of the adopted acid catalyst causes the content of formaldehyde in the condensation step to be higher, but rather causes the coking of the condensed catalyst bed layer, and finally leads to the inactivation. Thus, the reduction in the amount of methyl acrylate produced in the condensation step also affects the reduction in the amount of propionic acid in the final product after hydrogenation.

The invention has been described in detail with reference to specific embodiments and illustrative examples, but the description is not intended to limit the invention. Those skilled in the art will appreciate that various equivalent substitutions, modifications or improvements may be made to the embodiments and implementations of the invention without departing from the spirit and scope of the invention, and are within the scope of the invention. The scope of the invention is defined by the appended claims.

Claims

1. A preparation method of methyl propionate comprises the following steps:

2. The production method according to claim 1,

the aldehyde source is selected from at least one of trioxymethylene, paraformaldehyde, methylal and anhydrous formaldehyde; and/or the presence of a gas in the atmosphere,

in the step (1), the raw material sequentially passes through a first catalyst bed layer and a second catalyst bed layer which are connected in series.

3. The production method according to claim 2,

the molar ratio of the methyl acetate to the aldehyde source in the feed to the first catalytic bed is (1-10) to 1, preferably (2-5) to 1; and/or the presence of a gas in the atmosphere,

the weight ratio of methanol to methyl acetate in the feed to the first catalyst bed is (0-0.5): 1, preferably (0.2-0.5): 1; and/or the presence of a gas in the gas,

supplementing an aldehyde source at the top of the second catalyst bed.

4. The preparation method according to claim 2, wherein the reaction temperature of the first catalyst bed is 250 to 400 ℃; and/or the reaction pressure is 0.1-1 MPa; and/or the liquid phase volume flow rate is 0.01-1 mL/min; and/or the hydrogen flow rate is 20-150mL/min.

5. The preparation method of claim 2, wherein the reaction temperature of the second catalyst bed is 250-400 ℃; and/or the reaction pressure is 0.1-1 MPa; and/or the hydrogen flow rate is 50-300mL/min; and/or the liquid phase volume flow rate of the supplemental aldehyde source is 0-1mL/min.

6. The method of claim 2, wherein the first catalyst bed and the second catalyst bed independently comprise a solid base catalyst loading section and an optional acid catalyst loading section.

7. The preparation method according to claim 6, wherein the solid base catalyst comprises a carrier I and an active component I and an optional auxiliary agent which are loaded on the carrier I, wherein,

the carrier I is at least one selected from silicon oxide, aluminum oxide and SBA-15 molecular sieve; and/or the presence of a gas in the atmosphere,

the active component I is at least one of cesium, potassium and rubidium, and preferably, the loading amount of the active component I is 1-20 wt%; and/or the presence of a gas in the gas,

the auxiliary agent is selected from at least one of zirconium, bismuth and lanthanum compounds, and the loading amount of the auxiliary agent is preferably 0-5 wt%, and preferably 0.3-3 wt%.

8. A method according to claim 6, wherein the acidic catalyst is selected from alumina and/or molecular sieves, preferably from θ -Al ₂ O ₃ At least one of SAPO-34 molecular sieve and SAPO-35 molecular sieve.

9. The method of claim 6, wherein when the first catalyst bed and the second catalyst bed comprise an acidic catalyst loading section:

filling the acid catalyst with the bed volume of 0.5-10% at the 1/3-3/4 position of the first catalyst bed layer to form an acid catalyst filling section, and filling the solid base catalyst at the other positions to form a solid base catalyst filling section; and/or the presence of a gas in the atmosphere,

and filling the acid catalyst with the bed volume of 0.5-10% at the 1/3-3/4 position of the second catalyst bed layer to form the acid catalyst filling section, and filling the solid base catalyst at the other positions to form the solid base catalyst filling section.

10. The production method according to claim 1, wherein the hydrotreating is carried out in the presence of a hydrogenation catalyst in step (2), the hydrogenation catalyst comprising a support II and an active component II, wherein:

the carrier II is at least one selected from alumina, silica, SAPO-34 and activated carbon; and/or the presence of a gas in the atmosphere,

the active component II is selected from at least one of Pd, ni and Cu, preferably 100wt% based on the hydrogenation catalyst, wherein the loading amount of the active component II is 0.1-10 wt%.

11. The preparation method according to one of claims 1 to 10, wherein the intermediate product is subjected to a temperature reduction treatment or a heat exchange treatment before the step (2); preferably, the temperature of the intermediate product after the temperature reduction treatment or the heat exchange treatment is 80-250 ℃, and preferably 100-150 ℃.

12. The method according to claim 11, wherein the product after the hydrotreatment in the step (2) is subjected to condensation absorption and then gas-liquid separation, and the separated hydrogen is subjected to adsorption treatment with a base and then recycled to the step (1).

13. Methyl propionate obtainable by the process according to any one of claims 1 to 12.