A process for the preparation of methyl propionate and the methyl propionate obtained

Abstract

The application discloses a preparation method of methyl propionate and the obtained methyl propionate, and the method comprises the following steps: (1) under a hydrogen atmosphere, raw materials including methyl acetate, methanol and an aldehyde source are reacted to obtain an intermediate product; (2) under the hydrogen atmosphere, the intermediate product is subjected to hydrogenation treatment to obtain the methyl propionate. In step (1), the raw materials pass through a first catalyst bed and a second catalyst bed in series. The intermediate product is subjected to cooling treatment or heat exchange treatment before step (2), and the product is subjected to gas-liquid separation after the hydrogenation treatment in step (2), and hydrogen is separated out and recycled to step (1). The method disclosed by the application adopts the hydrogen atmosphere in the whole process, and there is no problem of separation of nitrogen and hydrogen in the later stage, energy consumption is saved, and the production cost is greatly reduced; meanwhile, the hydrogenation treatment is carried out in the gas phase, and the self-polymerization of methyl acrylate can be effectively reduced or avoided.

Description

Technical Field

[0001] This invention relates to the preparation of methyl acrylate, and more particularly to a method for preparing methyl acrylate and the methyl acrylate obtained therefrom. Background Technology

[0002] Methyl acrylate is an important fine chemical raw material with a wide range of applications. It is mainly used as an intermediate in organic synthesis and as a polymer monomer. Polymers prepared using methyl acrylate as a monomer are widely used in industries such as coatings, textiles, leather making, and adhesives.

[0003] The main methods for preparing acrylic acid and its esters include propylene oxidation, acrylonitrile hydrolysis, ketene method, propane oxidation, and methyl formate method. However, these methods suffer from drawbacks such as severe pollution, high energy consumption, and low product yield. Therefore, developing green and efficient new production processes is of great significance.

[0004] Lucite has proposed a mature α-MMA process, which uses methyl propionate and formaldehyde to produce methyl methacrylate. However, Lucite's process route is not suitable for my country's national conditions. my country has a significant overcapacity of methyl acetate. Lucite's process, which uses methyl propionate as a raw material to synthesize methyl methacrylate, cannot solve my country's overcapacity problem. Therefore, what is urgently needed in this field is a green synthesis of methyl methacrylate using methyl acetate, an industrial byproduct, as a raw material, employing a safe, environmentally friendly, and non-toxic solid alkali catalyst through a clean synthesis process.

[0005] In light of this, to address the significant overcapacity of methyl acetate in my country, a novel catalyst for the synthesis of methyl acrylate (Macrylate) has been proposed, based on the mature α-MMA process developed by Lucite. This catalyst utilizes methyl acetate and formaldehyde as raw materials. A coupled catalyst reaction-regeneration fluidized bed system has been employed, solving the problems of short catalyst lifespan, easy carbon buildup, and the need for frequent regeneration. Currently, the single-pass conversion rate of methyl acetate is generally around 15%, the selectivity of methyl acrylate can generally reach 85%, and the yield of methyl acrylate is generally around 13.5%. However, industrialization is still not feasible, primarily due to the inability to achieve a comprehensive improvement in yield, selectivity, and catalyst stability. In recent years, many domestic and international institutions have further researched the synthesis of methyl acrylate from methyl acetate and formaldehyde.

[0006] Chinese patent [CN104525176] discloses a method for preparing a Cs-based catalyst, the microspheres prepared therefrom, and a method for synthesizing methyl methacrylate using the microspheres. Specifically, a Cs salt, a salt of the auxiliary metal M, a non-essential linker, and a template agent are mixed with 20-40% silica sol to form a solution. The solution is then milled in the colloid at a speed of 5000-10000 rpm for 1-5 minutes. The microspheres are then spray-dried to obtain microspheres of 20-220 micrometers, dried at 70-120°C, and calcined at 200-600°C for 2-7 hours to obtain the condensation catalyst.

[0007] Chinese patent [CN103435483A] discloses a method for preparing methyl methacrylate (MA) from methyl acetate and formaldehyde using a fixed fluidized bed process. However, existing catalysts in fixed fluidized beds have many problems: 1. The ions are not Class A particles, increasing the risk of scale-up in fluidized bed processes and hindering industrialization; 2. The low ion packing density is unsuitable for fixed fluidized bed reaction processes; 3. Other factors, such as sphericity, activity, carbon deposition, and wear, cannot meet the requirements of fixed bed processes. The preparation of methyl methacrylate from methyl propionate and formaldehyde using the same fluidized bed process also suffers from the same problems. Therefore, in the existing technology, there is still a need for a catalyst system suitable for fluidized bed processes that can meet the requirements of fluidized beds.

[0008] In summary, the research on the synthesis of methyl acrylate from methyl acetate and formaldehyde has not yet been industrialized and is currently only in the research stage. However, as a new process route, its cost advantage is very obvious, and major multinational chemical companies are investing human and material resources in the research and development of this process.

[0009] Therefore, in summary, there is an urgent need in the field for a synthetic route for methyl acrylate that improves yield and selectivity while maintaining catalytic activity for a longer period of time, thereby enabling industrialization and solving the problem of excessive production capacity of methyl acetate, as well as a corresponding solid base catalyst. Summary of the Invention

[0010] To overcome the problems existing in the prior art, this invention provides a method for preparing methyl propionate and the resulting methyl propionate. The method includes a condensation reaction and a hydrogenation reaction under a hydrogen atmosphere, wherein the condensation reaction is carried out using a two-stage condensation process. The method of this invention can improve yield and selectivity, while maintaining catalytic activity for a longer period, thereby enabling industrialization and solving the problem of significant overcapacity in methyl acetate production.

[0011] One objective of this invention is to provide a method for preparing methyl propionate, comprising:

[0012] (1) Under a hydrogen atmosphere, the raw materials, including methyl acetate, methanol and aldehyde source, react to obtain an intermediate product.

[0013] (2) The intermediate product is hydrogenated under a hydrogen atmosphere to obtain methyl propionate.

[0014] In this invention, the entire process is carried out in a hydrogen atmosphere, unlike the existing technology that first condenses under nitrogen and then adds hydrogen under hydrogen. The method described in this invention does not have the problem of nitrogen and hydrogen separation in the later stage, saving energy and greatly reducing production costs.

[0015] In a preferred embodiment, the aldehyde source is selected from at least one of trioxymethylene, paraoxymethylene, methyl acetal and anhydrous formaldehyde; and / or, the methyl acetate is selected from crude methyl acetate and / or refined methyl acetate.

[0016] Both crude methyl acetate and refined methyl acetate are available for purchase.

[0017] In a preferred embodiment, in step (1), the raw material passes through a first catalyst bed and a second catalyst bed in series in sequence, and reacts to obtain an intermediate product.

[0018] In a preferred embodiment, the molar ratio of methyl acetate to aldehyde source in the feed to the first catalytic bed is (1-10):1, preferably (2-5):1, for example 2:1, 3:1, 4:1 or 5:1.

[0019] In a preferred embodiment, the weight ratio of methanol to methyl acetate in the first catalyst bed feed is (0-0.5):1, preferably (0.2-0.5):1.

[0020] In a preferred embodiment, an aldehyde source is supplemented on top of the second catalyst bed.

[0021] The aldehyde source is supplemented in a liquid state. Preferably, when the aldehyde source supplemented to the second catalyst bed is trioxymethylene, the solid trioxymethylene is heated until it is in a liquid state (preferably heated at 60-80°C) and then introduced into the second catalyst bed by a pump.

[0022] In a preferred embodiment, the reaction temperature of the first catalyst bed is 250–400°C; and / or, the reaction pressure is 0.1–1 MPa; and / or, the liquid phase volumetric flow rate is 0.01–1 mL / min; and / or, the hydrogen flow rate is 20–150 mL / min.

[0023] In a further preferred embodiment, the reaction temperature of the first catalyst bed is 300–360°C (e.g., 300°C, 330°C, 340°C, 350°C, 360°C); and / or, the reaction pressure is 0.1–0.5 MPa (e.g., 0.1 MPa, 0.2 MPa, 0.3 MPa, 0.4 MPa, 0.5 MPa); and / or, the liquid phase volumetric flow rate is 0.1–0.5 mL / min (e.g., 0.1 mL / min, 0.2 mL / min, 0.3 mL / min, 0.4 mL / min, 0.5 mL / min); and / or, the hydrogen flow rate is 50–150 mL / min.

[0024] The liquid phase volume flow rate in the first catalyst bed refers to the flow rate of liquid raw materials including methyl acetate and aldehyde source.

[0025] In a preferred embodiment, the reaction temperature of the second catalyst bed is 250–400°C; and / or, the reaction pressure is 0.1–1 MPa; and / or, the hydrogen flow rate is 50–300 mL / min; and / or, the liquid volumetric flow rate of the aldehyde source is 0–1 mL / min.

[0026] 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 300 mL / min.

[0027] In a further preferred embodiment, the reaction temperature of the second catalyst bed is 300–360°C (e.g., 300°C, 330°C, 340°C, 350°C, 360°C); and / or, the reaction pressure is 0.1–0.5 MPa (e.g., 0.1 MPa, 0.2 MPa, 0.3 MPa, 0.4 MPa, 0.5 MPa); and / or, the hydrogen flow rate is 150–300 mL / min; and / or, the liquid volumetric flow rate for replenishing the aldehyde source is 0.02–0.3 mL / min.

[0028] Specifically, the hydrogen flow rate in the second catalyst bed is controlled to be higher than that in the first catalyst bed.

[0029] In a preferred embodiment, the first catalyst bed and the second catalyst bed each independently include a solid base catalyst loading section and an optional acidic catalyst loading section, that is, an acidic catalyst section is optionally doped into the solid base catalyst.

[0030] In existing technologies, high aldehyde source concentrations in the upper part of the bed lead to coking, catalyst deactivation, and easy bed blockage, directly resulting in high concentrations of methyl acetate and low concentrations of aldehyde source in the lower bed. The high concentration of methyl acetate will condense to form acetone, increasing separation energy consumption.

[0031] In a preferred embodiment, an acidic catalyst loading section is incorporated into the solid alkali catalyst bed. The acidic catalyst promotes the decomposition of the aldehyde source, increases the concentration of the aldehyde source in the lower bed, and provides sufficient formaldehyde, thereby inhibiting the formation of acetone.

[0032] In a preferred embodiment, the solid base catalyst comprises a support I and an active component I supported on the support I, along with optional auxiliaries.

[0033] In a further preferred embodiment, the carrier I is selected from at least one of silica, alumina, and SBA-15 molecular sieve, preferably having a specific surface area of ​​50-500 m². 2 / g, pore size 6-30nm, porosity 0.6-1mL / g.

[0034] In a further preferred embodiment, the active component I is at least one of cesium, potassium and rubidium (preferably cesium), and preferably, its loading is 1 to 20 wt%, more preferably 2 to 10 wt%, for example 2 wt%, 3 wt%, 4 wt%, 5 wt%, 6 wt%, 7 wt%, 8 wt%, 9 wt%, 10 wt%.

[0035] The active component I is derived from at least one of a carbonate compound, a formic acid compound, and a nitrate compound containing an active component element.

[0036] In a further preferred embodiment, the additive is selected from at least one of zirconium, bismuth and lanthanum compounds, preferably with a loading of 0 to 5 wt%, more preferably 0.3 to 3 wt%, for example 0.5 wt%, 1 wt%, 2 wt%, 3 wt%.

[0037] Zirconium, bismuth, and nickel are derived from zirconium-containing compounds, bismuth-containing compounds, and lanthanum-containing compounds, respectively, such as zirconium oxynitrate, bismuth chloride, and lanthanum nitrate.

[0038] When preparing the solid base catalyst, if equal-volume impregnation is used, since almost all the components in the impregnation solution are loaded onto the support I, the content of active component I and auxiliary agent in the product can be calculated based on the amount of raw materials used by theoretical calculation method.

[0039] In a preferred embodiment, the acidic catalyst is selected from alumina and / or molecular sieves.

[0040] In a further preferred embodiment, the acidic catalyst is selected from at least one of θ-Al2O3, SAPO-34 molecular sieve, and SAPO-35 molecular sieve.

[0041] In a preferred embodiment, when the first catalyst bed and the second catalyst bed include an acidic catalyst loading section, 0.5-10% of the bed volume of the acidic catalyst is loaded at 1 / 3-3 / 4 of the position of the first catalyst bed to form the acidic catalyst loading section, and the remaining positions are filled with the solid base catalyst to form the solid base catalyst loading section.

[0042] For example, the acidic catalyst is loaded at 1 / 3, 1 / 2, 5 / 8, or 3 / 4 of the first catalyst bed volume in amounts of 0.5%, 1%, 2%, 3%, 4%, 5%, 6%, 7%, 8%, 9%, or 10% of the bed volume.

[0043] In a further preferred embodiment, when the first catalyst bed and the second catalyst bed include an acidic catalyst loading section, 1-4% of the bed volume of the acidic catalyst is loaded at 1 / 2-5 / 8 of the position of the first catalyst bed to form the acidic catalyst loading section, and the remaining positions are filled with the solid base catalyst to form the solid base catalyst loading section.

[0044] The positions from top to bottom of the bed are numbered 0 to 1. Specifically, in the first catalyst bed, the aldehyde source in the upper part of the bed is relatively abundant and does not require an acidic catalyst. Therefore, an acidic catalyst is only loaded in the middle and lower sections of the bed to decompose the undepolymerized aldehyde source and solve the problem of insufficient aldehyde source supply in the middle and lower parts of the bed.

[0045] In a preferred embodiment, when the first catalyst bed and the second catalyst bed include an acidic catalyst loading section, 0.5-10% of the bed volume of the acidic catalyst is loaded into the second catalyst bed at 1 / 3-3 / 4 position to form the acidic catalyst loading section, and the remaining positions are filled with the solid base catalyst to form the solid base catalyst loading section.

[0046] For example, the acidic catalyst is loaded at 1 / 3, 1 / 2, 5 / 8, or 3 / 4 of the first catalyst bed volume in amounts of 0.5%, 1%, 2%, 3%, 4%, 5%, 6%, 7%, 8%, 9%, or 10% of the bed volume.

[0047] In a further preferred embodiment, when the first catalyst bed and the second catalyst bed include an acidic catalyst loading section, 1-4% of the bed volume of the acidic catalyst is loaded at positions 1 / 2-5 / 8 of the second catalyst bed to form the acidic catalyst loading section, and the remaining positions are filled with the solid base catalyst to form the solid base catalyst loading section.

[0048] The positions from top to bottom of the bed are numbered 0 to 1. Specifically, because a small amount of aldehyde source is introduced at the inlet of the second reactor, there is sufficient formaldehyde in the upper part of the reaction bed of the second reactor. However, as the reaction proceeds, the aldehyde source decreases continuously, and the decomposed monomeric formaldehyde also decreases continuously, resulting in insufficient formaldehyde in the lower bed. Therefore, an acidic catalyst is packed in the middle and lower part of the second catalyst bed.

[0049] In this invention, the two beds connected in series are respectively disposed in two reactors connected in series.

[0050] In a preferred embodiment, the hydrogenation treatment in step (2) is carried out in the presence of a hydrogenation catalyst comprising a support II and an active component II, wherein the specific surface area of ​​the support II by the BET method is 20-200 m². 2 / g, with a pore size of 2-30nm; preferably, the carrier II is selected from at least one of alumina, silicon oxide, SAPO-34 and activated carbon, preferably silicon oxide and / or activated carbon.

[0051] In a further preferred embodiment, the active component II is selected from at least one of Pd, Ni and Cu, preferably, based on 100 wt% of the hydrogenation catalyst, wherein the loading of the active component II is 0.1 to 10 wt%, more preferably 0.1 to 1 wt%.

[0052] In a further preferred embodiment, in step (2), during the hydrogenation process, the temperature is 120-250°C, preferably 120-150°C; and / or, the pressure is 0.1-1 MPa, preferably 0.1-0.6 MPa; and / or, the hydrogen flow rate is 0-300 mL / min, preferably excluding 0, preferably 0-200 mL / min, preferably excluding 0.

[0053] In a preferred embodiment, the intermediate product is subjected to cooling or heat exchange treatment before step (2).

[0054] In a further preferred embodiment, the temperature of the intermediate product after cooling or heat exchange treatment is 80–250°C, preferably 100–150°C. For example, it is 90°C, 100°C, 110°C, 120°C, 130°C, 140°C, or 150°C.

[0055] The intermediate product obtained in step (1) has a high temperature (above 300°C, for example 350°C), but hydrogenation cannot be carried out at such a high temperature, otherwise hydrogenation of the raw material ester may occur. In the prior art, hydrogenation is generally carried out after the temperature is lowered to below the boiling point of MA (i.e., liquid-phase hydrogenation of MA). However, the intermediate product is methyl acrylate, which is prone to self-polymerization at room temperature, clogging the reactor.

[0056] To address the aforementioned issues, this invention cools or heats the intermediate product to 80–250°C, preferably 100–150°C. Within this temperature range, methyl acrylate (MA) is in the gaseous phase. The gaseous state increases the intermolecular distance, effectively reducing or preventing self-polymerization. Furthermore, compared to lowering the temperature to a liquid state (MA), this invention lowers the temperature to a gaseous state (MA), significantly reducing energy consumption and facilitating industrial production and application.

[0057] In a preferred embodiment, after the hydrogenation treatment in step (2), the product is condensed and absorbed, then separated by gas-liquid separation, and the separated hydrogen is adsorbed by an alkaline compound, and then recycled back to step (1).

[0058] In a further preferred embodiment, the alkaline compound is selected from calcium oxide and / or calcium hydroxide, which effectively removes the carbon dioxide generated during the reaction, protects the condensation catalyst in the condensation reactor, and the gas after gas-liquid separation is recycled back to the condensation reactor.

[0059] A second objective of this invention is to provide methyl propionate obtained using the method described in one objective of this invention.

[0060] The endpoints and any values ​​of the ranges disclosed in this invention are not limited to the precise ranges or values; these ranges or values ​​should be understood to include values ​​close to these ranges or values. For numerical ranges, the endpoint values ​​of the various ranges, the endpoint values ​​of the various ranges and individual point values, and individual point values ​​can be combined with each other to obtain one or more new numerical ranges, which 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 considered as specifically disclosed herein.

[0061] Compared with the prior art, the present invention has the following beneficial effects:

[0062] (1) The method described in this invention uses a hydrogen atmosphere throughout the process, and there is no problem of nitrogen and hydrogen separation in the later stage, which saves energy and greatly reduces production costs.

[0063] (2) The composite catalyst bed can solve the problem of insufficient supply of aldehyde compounds in the lower bed, and at the same time suppress side reactions and reduce the content of by-product acetone.

[0064] (3) The hydrogenation treatment is carried out in the gas phase, which can effectively reduce or avoid the self-polymerization of methyl acrylate. Detailed Implementation

[0065] The present invention will now be described in detail with reference to specific embodiments. It should be noted that the following embodiments are only used to further illustrate the present invention and should not be construed as limiting the scope of protection of the present invention. Some non-essential improvements and adjustments made by those skilled in the art based on the content of the present invention are still within the scope of protection of the present invention.

[0066] It should also be noted that the various specific technical features described in the following embodiments can be combined in any suitable manner without contradiction. To avoid unnecessary repetition, the various possible combinations will not be described separately in this invention.

[0067] Furthermore, various embodiments of the present invention can be combined in any way, as long as they do not violate the spirit of the present invention. The resulting technical solutions are part of the original disclosure of this specification and also fall within the protection scope of the present invention.

[0068] Unless otherwise specified, the raw materials used in the examples and comparative examples are all disclosed in the prior art, such as those that can be directly purchased or prepared according to the preparation methods disclosed in the prior art.

[0069] The solid base catalyst Cs / SiO2 was prepared as follows: 0.72g of Cs2CO3 was dissolved in 12mL of deionized water, and the Cs / SiO2 catalyst was prepared by the equal volume method. After standing for 8 hours, it was dried in a vacuum drying oven at 110℃ for 12 hours, and then calcined at 550℃ for 6 hours in air atmosphere to obtain the catalyst Cs / SiO2.

[0070] The solid base catalyst Cs / Zr-SiO2 was prepared as follows: 0.14 g of ZrOCl2·8H2O was dissolved in 12 mL of deionized water, and the Zr-SiO2 catalyst was prepared by the equal volume method. The catalyst was allowed to stand for 8 hours, dried in a vacuum oven at 110 °C for 12 hours, and then calcined at 550 °C for 6 hours in air to obtain the catalyst support Zr-SiO2. Alternatively, 0.72 g of Cs2CO3 was dissolved in 12 mL of deionized water, and the Cs / Zr-SiO2 catalyst was prepared using Zr-SiO2 as the support by the equal volume method. The catalyst was allowed to stand for 8 hours, dried in a vacuum oven at 110 °C for 12 hours, and then calcined at 550 °C for 6 hours in air to obtain the catalyst Cs / Zr-SiO2.

[0071] The hydrogenation catalyst Pd / C was obtained as follows: Pd / C catalyst was prepared by using activated carbon as support and palladium chloride as precursor with a loading of 0.5 wt% and by equal volume impregnation method. After impregnation, the catalyst was placed at room temperature for 8 h, dried at 110 °C for 12 h, and then calcined at 550 °C for 5 h in air atmosphere to obtain the Pd / C catalyst.

[0072] The hydrogenation catalyst Pd / SiO2 was obtained as follows: using silica as a support and palladium chloride as a precursor with a loading of 0.5 wt%, the Pd / SiO2 catalyst was prepared by an equal-volume impregnation method. After impregnation, the catalyst was placed at room temperature for 8 hours, dried at 110℃ for 12 hours, and then calcined at 550℃ for 5 hours in air atmosphere to obtain the Pd / SiO2 catalyst.

[0073] Samples were taken, and toluene was added as an internal standard. The concentrations of methyl propionate (MP) and methyl acrylate (MA) were measured by gas chromatography.

[0074]

Example 1

[0075] 10 mL of solid base catalyst Cs / SiO2 was added to both the first and second catalyst beds. The molar ratio of methyl acetate to paraformaldehyde in the feed to the first catalyst bed was 3:1, and the amount of methanol introduced was 30% of the weight of methyl acetate.

[0076] The hydrogen flow rate in the first catalyst bed is 110 mL / min, the liquid flow rate of the mixture of methyl acetate, methanol and trioxymethylene in the first catalyst bed is 0.2 mL / min, the reaction pressure of the first catalyst bed is 0.6 MPa, and the reaction temperature is 350 °C.

[0077] The hydrogen flow rate in the second catalyst bed is 120 mL / min, the liquid flow rate of paraformaldehyde added to the second catalyst bed is 0.05 mL / min, the reaction pressure of the second catalyst bed is 0.6 MPa, and the reaction temperature is 350 °C.

[0078] The gas exiting the second catalyst bed is cooled to 120°C by a heat exchanger. The product gas then enters the hydrogenation reactor, where the catalyst is Pd / C, the hydrogenation temperature is 120°C, the hydrogenation pressure is 0.6 MPa, and the hydrogen flow rate is 140 mL / min. After hydrogenation, the gas enters the gas-liquid separator. The liquid is condensed and absorbed, and the gas is absorbed by calcium oxide before being recycled back to the first catalyst bed. The conversion rate of methyl acetate (MAC) is 16%, and the product concentration of methyl propionate (MP) is 11%.

[0079]

Example 2

[0080] 10 mL of solid base catalyst Cs / SiO2 was added to both the first and second catalyst beds. The molar ratio of methyl acetate to paraformaldehyde in the feed to the first catalyst bed was 3:1, and the amount of methanol introduced was 30% of the weight of methyl acetate.

[0081] The hydrogen flow rate in the first catalyst bed is 110 mL / min, the liquid flow rate of the mixture of methyl acetate, methanol and trioxymethylene in the first catalyst bed is 0.2 mL / min, the reaction pressure of the first catalyst bed is 0.6 MPa, and the reaction temperature is 350 °C.

[0082] The hydrogen flow rate in the second catalyst bed is 120 mL / min, the liquid flow rate of paraformaldehyde added to the second catalyst bed is 0.05 mL / min, the reaction pressure of the second catalyst bed is 0.6 MPa, and the reaction temperature is 350 °C.

[0083] The gas exiting the second catalyst bed is cooled to 120°C by a heat exchanger. The product gas then enters the hydrogenation reactor, where the catalyst is Pd / SiO2, the hydrogenation temperature is 120°C, the hydrogenation pressure is 0.6 MPa, and the hydrogen flow rate is 140 mL / min. After hydrogenation, the gas enters the gas-liquid separator. The liquid is condensed and absorbed, and the gas is absorbed by calcium oxide and then recycled back to the first catalyst bed. 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%.

[0084]

Example 3

[0085] 10 mL of solid base catalyst Cs / SiO2 is added to the first catalyst bed and the second catalyst bed, respectively. 1% of the bed volume of the acidic catalyst SAPO-34 molecular sieve is loaded at 1 / 3 of the position of the first catalyst bed to form the acidic catalyst loading section. The molar ratio of methyl acetate, methanol and trioxymethylene in the feed to the first catalyst bed is 3:1, and the amount of methanol introduced is 30% of the weight of methyl acetate.

[0086] The hydrogen flow rate in the first catalyst bed is 100 mL / min, the liquid flow rate of the methyl acetate and trioxymethylene mixture in the first catalyst bed is 0.2 mL / min, the reaction pressure of the first catalyst bed is 0.6 MPa, and the reaction temperature is 350 °C.

[0087] The hydrogen flow rate in the second catalyst bed is 120 mL / min, the liquid flow rate of paraformaldehyde added to the second catalyst bed is 0.05 mL / min, the reaction pressure of the second catalyst bed is 0.6 MPa, and the reaction temperature is 350 °C.

[0088] The gas exiting the second catalyst bed is cooled to 120°C by a heat exchanger. The product gas then enters the hydrogenation reactor, where the catalyst is Pd / C. After hydrogenation, the gas enters the gas-liquid separator. The liquid is absorbed by condensation, and the gas is absorbed by calcium oxide before being recycled back to the first catalyst bed. The conversion rate of methyl acetate (MAC) is 18.8%, and the product concentration of methyl propionate (MP) is 13.9%.

[0089]

Example 4

[0090] 10 mL of solid base catalyst Cs / SiO2 was added to the first catalyst bed and the second catalyst bed, respectively. 1% of the bed volume of the acidic catalyst SAPO-34 molecular sieve was loaded at the 1 / 3 position of the first catalyst bed to form the acidic catalyst loading section. The molar ratio of methyl acetate to trioxymethylene in the feed to the first catalyst bed was 3:1, and the amount of methanol introduced was 30% of the weight of methyl acetate.

[0091] The hydrogen flow rate in the first catalyst bed is 100 mL / min, the liquid flow rate of the methyl acetate and trioxymethylene mixture in the first catalyst bed is 0.2 mL / min, the reaction pressure of the first catalyst bed is 0.6 MPa, and the reaction temperature is 350 °C.

[0092] The hydrogen flow rate in the second catalyst bed is 110 mL / min, the liquid flow rate of paraformaldehyde added to the second catalyst bed is 0.05 mL / min, the reaction pressure of the second catalyst bed is 0.6 MPa, and the reaction temperature is 350 °C.

[0093] The gas exiting the second catalyst bed is cooled to 120°C by a heat exchanger. The product gas then enters the hydrogenation reactor, where the catalyst is Pd / SiO2. After hydrogenation, the gas enters the gas-liquid separator. The liquid is absorbed by condensation, and the gas is absorbed by calcium oxide before being recycled back to the first catalyst bed. 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%.

[0094]

Example 5

[0095] 10 mL of solid base catalyst Cs / Zr-SiO2 was added to both the first and second catalyst beds. The molar ratio of methyl acetate to paraformaldehyde in the feed to the first catalyst bed was 3:1, and the amount of methanol introduced was 30% of the weight of methyl acetate.

[0096] The hydrogen flow rate in the first catalyst bed is 100 mL / min, the liquid flow rate of the methyl acetate and trioxymethylene mixture in the first catalyst bed is 0.2 mL / min, the reaction pressure of the first catalyst bed is 0.6 MPa, and the reaction temperature is 350 °C.

[0097] The hydrogen flow rate in the second catalyst bed is 110 mL / min, the liquid flow rate of paraformaldehyde added to the second catalyst bed is 0.05 mL / min, the reaction pressure of the second catalyst bed is 0.6 MPa, and the reaction temperature is 350 °C.

[0098] The gas exiting the second catalyst bed is cooled to 120°C by a heat exchanger. The product gas then enters the hydrogenation reactor, where the catalyst is Pd / C, the hydrogenation temperature is 120°C, the hydrogenation pressure is 0.6 MPa, and the hydrogen flow rate is 140 mL / min. After hydrogenation, the gas enters the gas-liquid separator. The liquid is condensed and absorbed, and the gas is absorbed by calcium oxide and then recycled back to the first catalyst bed. The conversion rate of methyl acetate (MAC) is 17.9%, and the product concentration of methyl propionate (MP) is 13.2%.

[0099]

Examples 6-9

[0100] The process of Example 1 was repeated, except that the gas exiting the second catalyst bed was cooled to 80°C, 100°C, 150°C and 200°C respectively by a heat exchanger. Compared with no cooling treatment, the concentration of methyl propionate (MP) product was increased in all cases.

[0101] Comparative Example 1

[0102] The process of Example 1 is repeated, except that the gas exiting the second catalyst bed is cooled to room temperature by a heat exchanger and then liquid-phase hydrogenation is performed, while other conditions remain unchanged.

[0103] 10 mL of solid base catalyst Cs / SiO2 was added to both the first and second catalyst beds. The molar ratio of methyl acetate to trioxymethylene in the feed to the first catalyst bed was 3:1.

[0104] The hydrogen flow rate in the first catalyst bed is 110 mL / min, the liquid flow rate of the methyl acetate and trioxymethylene mixture in the first catalyst bed is 0.2 mL / min, the reaction pressure of the first catalyst bed is 0.6 MPa, and the reaction temperature is 350 °C.

[0105] The hydrogen flow rate in the second catalyst bed is 130 mL / min, the liquid flow rate of paraformaldehyde added to the second catalyst bed is 0.05 mL / min, the reaction pressure of the second catalyst bed is 0.6 MPa, and the reaction temperature is 350 °C.

[0106] The gas exiting the second catalyst bed is cooled to room temperature by a heat exchanger, and then separated into gas and liquid phases. The liquid product enters the hydrogenation reactor, where the catalyst is Pd / C, and a hydrogenation reaction is carried out. The gas is absorbed by calcium oxide and then recycled back to the first catalyst bed. The conversion rate of methyl acetate (MAC) is 15%, and the product concentration of methyl propionate (MP) is 7%. In the second condensation reactor, after gas-liquid separation, the separated methyl acrylate solution polymerizes and blocks the pipeline leading to the hydrogenation reactor.

[0107] Comparative Example 2

[0108] The process of Example 1 was repeated, except that the hydrogen flow rate in both the first and second catalyst beds was 110 mL / min, and other conditions remained unchanged.

[0109] It was found that when the hydrogen flow rate of the second catalyst bed was approximately equal to that of the first catalyst bed, the concentration of methyl propionate in the product decreased significantly after five days of operation. This was because the long-term contact between the second catalyst bed and water affected the stability of the catalyst, resulting in a decrease in the content of methyl acrylate produced by condensation, which in turn reduced the content of methyl propionate in the hydrogenation product.

[0110] Comparative Example 3

[0111] The process of Example 3 was repeated, except that the amount of acidic catalyst in the bed was larger, which was 15% of the bed volume of acidic catalyst, while other conditions remained the same.

[0112] It was found that after five days of operation, the methyl propionate content in the product decreased significantly, to only half of its original value. The reason for this was that an excessive amount of acidic catalyst was used, leading to a higher formaldehyde content in the condensation step. This, in turn, caused coking of the catalyst bed during condensation, ultimately resulting in deactivation. Thus, the reduced methyl acrylate content produced in the condensation step also affected the reduction in propionic acid content in the final product after hydrogenation.

[0113] The present invention has been described in detail above with reference to specific embodiments and exemplary examples; however, these descriptions should not be construed as limiting the present invention. Those skilled in the art will understand that various equivalent substitutions, modifications, or improvements can be made to the technical solutions and embodiments of the present invention without departing from the spirit and scope of the invention, and all such modifications and improvements fall within the scope of the present invention. The scope of protection of the present invention is defined by the appended claims.

Claims

1. A method for preparing methyl propionate, comprising: (1) Under a hydrogen atmosphere, raw materials including methyl acetate, methanol and aldehyde source are sequentially passed through a first catalyst bed and a second catalyst bed connected in series to react. The hydrogen flow rate in the second catalyst bed is controlled to be higher than that in the first catalyst bed to obtain an intermediate product. The first catalyst bed includes a solid base catalyst loading section and an optional acid catalyst loading section. The second catalyst bed includes a solid base catalyst loading section. When the first catalyst bed includes an acid catalyst loading section: 0.5-10% of the bed volume of acid catalyst is loaded at 1 / 3-3 / 4 of the first catalyst bed to form the acid catalyst loading section, and the remaining positions are filled with solid base catalyst to form the solid base catalyst loading section. (2) The intermediate product is hydrogenated under a hydrogen atmosphere to obtain methyl propionate.

2. The preparation method according to claim 1, characterized in that, The aldehyde source is selected from at least one of trioxymethylene, paraoxymethylene, methyl acetal, and anhydrous formaldehyde.

3. The preparation method according to claim 2, characterized in that, The molar ratio of methyl acetate to aldehyde source in the feed to the first catalytic bed is (1-10):1; and / or, In the first catalyst bed feed, the weight ratio of methanol to methyl acetate is (0-0.5):1; and / or, An aldehyde source is added to the top of the second catalyst bed.

4. The preparation method according to claim 2, characterized in that, The molar ratio of methyl acetate to aldehyde source in the feed to the first catalytic bed is (2-5):1; and / or, In the first catalyst bed feed, the weight ratio of methanol to methyl acetate is (0.2-0.5):

1.

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

6. The preparation method according to claim 2, characterized in that, The reaction temperature of the second catalyst bed is 250~400℃; and / or, the reaction pressure is 0.1~1MPa; and / or, the hydrogen flow rate is 50-300mL / min; and / or, the liquid volume flow rate of the aldehyde source is 0-1mL / min.

7. The preparation method according to claim 1, characterized in that, The solid base catalyst comprises a support I and an active component I supported on the support I, along with optional auxiliary agents, wherein, The carrier I is selected from at least one of silica, alumina, and SBA-15 molecular sieve; and / or, The active component I is at least one of cesium, potassium, and rubidium; and / or, The additive is selected from at least one of zirconium, bismuth and lanthanum compounds.

8. The preparation method according to claim 7, characterized in that, The loading of active component I is 1~20 wt%; and / or, The loading of the additive is 0~5wt%.

9. The preparation method according to claim 7, characterized in that, The loading of the additive is 0.3~3wt%.

10. The preparation method according to claim 1, characterized in that, The acidic catalyst is selected from alumina and / or molecular sieves.

11. The preparation method according to claim 1, characterized in that, The acidic catalyst is selected from at least one of θ-Al2O3, SAPO-34 molecular sieve, and SAPO-35 molecular sieve.

12. The preparation method according to claim 1, characterized in that, The hydrogenation treatment in step (2) is carried out in the presence of a hydrogenation catalyst, which comprises a support II and an active component II, wherein: The carrier II is selected from at least one of alumina, silica, SAPO-34, and activated carbon; and / or, The active component II is selected from at least one of Pd, Ni and Cu.

13. The preparation method according to claim 12, characterized in that, The hydrogenation catalyst is 100 wt%, wherein the loading of active component II is 0.1~10 wt%.

14. The preparation method according to any one of claims 1 to 13, characterized in that, The intermediate product is subjected to cooling or heat exchange treatment before step (2).

15. The preparation method according to claim 14, characterized in that, The temperature of the intermediate product after cooling or heat exchange treatment is 80~250℃.

16. The preparation method according to claim 14, characterized in that, The temperature of the intermediate product after cooling or heat exchange treatment is 100~150℃.

17. The preparation method according to claim 14, characterized in that, After the hydrogenation treatment described in step (2), the product is condensed and absorbed, then separated by gas-liquid separation. The separated hydrogen is then adsorbed with alkali and recycled back to step (1).