JP2009024000A - METHOD FOR PRODUCING alpha,beta-UNSATURATED CARBOXYLIC ACID - Google Patents

Abstract

<P>PROBLEM TO BE SOLVED: To provide a method for producing an α,β-unsaturated carboxylic acid from an olefin or an α,β-unsaturated aldehyde in a liquid phase, which is operable at lower reaction pressure than ever, is capable of reducing olefin volatilization when the olefin is used as a raw material, and is capable of markedly improving oxygen conversion. <P>SOLUTION: The method for producing the α,β-unsaturated carboxylic acid comprises a step of oxidizing the corresponding olefin or α,β-unsaturated aldehyde by feeding an oxygen-containing gas into the liquid phase in a reaction vessel, and a step of feeding an inert gas lower in oxygen concentration than the above oxygen-containing gas into the vapor phase in the reaction vessel, and concurrently discharging, out of the reaction vessel, unreacted oxygen as exhaust gas moved from the liquid phase into the vapor phase. <P>COPYRIGHT: (C)2009,JPO&INPIT

Description

  The present invention relates to a method for producing an α, β-unsaturated carboxylic acid. Specifically, the present invention relates to a method for producing an α, β-unsaturated carboxylic acid by oxidizing an olefin or an α, β-unsaturated aldehyde in a liquid phase.

  Non-Patent Document 1 describes industrial production of acetaldehyde. As the liquid phase oxidation method, a so-called Hoechst-Wacker method using direct oxidation of ethylene is the main production method, which includes a one-step method using oxygen as an oxidant and a two-step method using air. . One-stage process is characterized by simultaneous ethylene oxidation and catalyst reoxidation in one reactor.

These methods are considered to be applicable to the production of α, β-unsaturated carboxylic acids from olefins or α, β-unsaturated aldehydes. Patent Document 1 discloses a method for producing an α, β-unsaturated carboxylic acid from an olefin by air oxidation using a noble metal catalyst.
International Publication No. 2006/85540 Pamphlet Unit Process Series 1 Oxidation Chemical Industries (1963), pages 220-242

  The method of Patent Document 1 has been previously studied by the present inventors, and uses air as an oxidizing agent when producing an α, β-unsaturated carboxylic acid from an olefin using a noble metal catalyst. It was. In this case, in order to obtain a sufficient reaction rate, there is a problem that the reaction must be performed under high pressure. In air oxidation, the partial pressure of oxygen can be increased by increasing the total pressure. However, if the pressure is increased, the reactor and its maintenance costs increase, which is not economical. On the other hand, even if the total pressure was simply lowered, the reaction rate was lowered, and a catalyst was required accordingly.

  In addition, isobutylene supplied as a raw material has an extremely low boiling point compared to other substances related to this reaction. Therefore, when the reaction is performed by supplying air to the liquid phase, isobutylene is volatilized by a large amount of nitrogen gas accompanying oxygen. However, most of them were not used for the actual reaction, and they were out of the system. As a result, (1) equipment cost for isobutylene recovery increases. (2) When the amount of volatilization is compared with a constant supply amount of isobutylene, if the volatility is large, isobutylene in the reaction solution results. There was a problem that the concentration was lowered and the reaction rate was lowered, and (3) the concentration of the target substance could not be increased in the reaction vessel, and a large amount of energy was consumed for separation and purification.

  Further, air oxidation has a problem that the oxygen conversion rate is low.

  As described above, the air oxidation method as disclosed in Patent Document 1 has many problems to be solved and cannot be said to be industrially satisfactory.

  These problems can be solved in principle by using a gas having a higher oxygen concentration than air such as oxygen as the oxidant. However, when a gas having a high oxygen concentration is used, there is a risk that a combustible mixture may be formed in the gas phase portion, and a technology for industrially using a gas having a high oxygen concentration has not yet been developed.

  In the production of α, β-unsaturated carboxylic acid from olefin or α, β-unsaturated aldehyde in the liquid phase, the present invention can be operated at a reaction pressure lower than that in the prior art. It is an object of the present invention to provide a production method capable of reducing volatilization of oxygen and greatly improving the oxygen conversion rate.

  The present inventors diligently studied the oxidation of isobutylene using high-concentration oxygen as an oxidant for these problems. In this reaction, the oxygen concentration in the gas phase is restricted to avoid explosion, but it has been found that it can be realized by diluting with an inert gas in an amount corresponding to the amount of unreacted oxygen. Furthermore, through continuous evaluation studies, it was found that oxidation with high-concentration oxygen can suppress a considerable amount of isobutylene volatilization, which can reduce the reaction pressure. It has also been found that the productivity of the target substance is strongly related to the oxygen partial pressure in the bubbles.

  However, since this reaction is under high temperature and high pressure, it was impossible to accurately measure the oxygen partial pressure in the bubbles. As a result of diligent investigation, the inventors have found a method for well expressing the oxygen partial pressure in the bubbles existing in the liquid phase in consideration of the dilution of the combustible gas by the inert gas and the amount of entrained gas from the liquid phase. From the above examination results, even when high-concentration oxygen is used as the oxidant, reacting while diluting unreacted oxygen provides favorable product productivity and selectivity of the oxidation reaction, and volatilization of isobutylene. The amount was kept low and the oxygen conversion rate was found to be greatly improved, and the present invention was completed.

  That is, the present invention provides a method for producing an α, β-unsaturated carboxylic acid by oxidizing an olefin or an α, β-unsaturated aldehyde in a liquid phase in the presence of a catalyst in a reaction vessel. A step of oxidizing the olefin or the α, β-unsaturated aldehyde by supplying an oxygen-containing gas to the liquid phase part of the gas phase, and an inert gas having a lower oxygen concentration than the oxygen-containing gas in the gas phase part in the reaction vessel And a step of discharging unreacted oxygen moved from the liquid phase portion to the gas phase portion as an exhaust gas outside the reaction vessel while providing a gas-containing gas. It is about the method.

  According to the present invention, when α, β-unsaturated carboxylic acid is produced from olefin or α, β-unsaturated aldehyde in the liquid phase, it can be operated at a lower reaction pressure than before, and olefin is used as a raw material. In addition, the volatilization of olefins can be reduced. Also, the oxygen conversion rate can be greatly improved. This makes it possible to provide an industrial process with low reaction equipment and high energy efficiency.

  Hereinafter, the present invention will be described in detail.

  In the present invention, an olefin or an α, β-unsaturated aldehyde is oxidized in the liquid phase in the presence of a catalyst to produce an α, β-unsaturated carboxylic acid (this reaction is referred to as a liquid phase oxidation reaction).

  Examples of the olefin used as a raw material in the present invention include propylene, isobutylene, 2-butene and the like. Examples of the α, β-unsaturated aldehyde used as a raw material in the present invention include acrolein, methacrolein, crotonaldehyde (β-methylacrolein), cinnamaldehyde (β-phenylacrolein) and the like. Suitable for liquid phase oxidation using propylene or isobutylene as olefin, and acrolein or methacrolein as α, β-unsaturated aldehyde as raw materials. The olefin or α, β-unsaturated aldehyde used as a raw material may contain some saturated hydrocarbons and / or lower saturated aldehydes as impurities.

  The α, β-unsaturated carboxylic acid produced is an α, β-unsaturated carboxylic acid having the same carbon skeleton as the olefin when the raw material is an olefin, and when the raw material is an α, β-unsaturated aldehyde, An α, β-unsaturated carboxylic acid in which an aldehyde group of an α, β-unsaturated aldehyde is changed to a carboxyl group. Specifically, acrylic acid is obtained when the raw material is propylene or acrolein, and methacrylic acid is obtained when the raw material is isobutylene or methacrolein.

  The reaction solvent used for the liquid phase oxidation reaction is not particularly limited, and water, alcohols, ketones, organic acids, organic acid esters, hydrocarbons, and the like can be used. Examples of alcohols include tertiary butanol and cyclohexanol. Examples of ketones include acetone, methyl ethyl ketone, and methyl isobutyl ketone. Examples of the organic acids include acetic acid, propionic acid, n-butyric acid, iso-butyric acid, n-valeric acid, iso-valeric acid and the like. Examples of the organic acid esters include ethyl acetate and methyl propionate. Examples of hydrocarbons include hexane, cyclohexane, and toluene. Among these, organic acids having 2 to 6 carbon atoms, ketones having 3 to 6 carbon atoms, and tertiary butanol are preferable. A solvent may be individual or 2 or more types of mixed solvents may be sufficient as it. Further, when at least one compound selected from the group consisting of alcohols, ketones, organic acids and organic acid esters is used as a solvent, it is preferable to use a mixed solvent of this compound and water. Although content of the water in a mixed solvent is not specifically limited, 2 mass% or more is preferable with respect to the mass of a mixed solvent, and 5 mass% or more is more preferable. Moreover, 70 mass% or less is preferable, and 50 mass% or less is more preferable. The mixed solvent is desirably uniform, but may be used in a uniform state.

  In the present invention, the amount of the olefin or α, β-unsaturated aldehyde in the reaction solution for performing the liquid phase oxidation reaction is preferably 0.1% by mass or more based on the reaction solvent present in the reaction vessel, The mass% or more is more preferable. Moreover, 50 mass% or less is preferable, and 30 mass% or less is more preferable.

In order to prevent polymerization of raw materials and products due to high temperature during the liquid phase oxidation reaction, it is preferable to use a polymerization inhibitor. Examples of the polymerization inhibitor that can be used in this case include phenolic compounds such as hydroquinone and p-methoxyphenol; N, N′-diisopropyl-p-phenylenediamine, N, N′-di-2-naphthyl-p- Amine compounds such as phenylenediamine and N-phenyl-N ′-(1,3-dimethylbutyl) -p-phenylenediamine; 4-hydroxy-2,2,6,6-tetramethylpiperidine-N-oxyl, or N-oxyl compounds such as 4- [H- (OCH ₂ CH ₂ ) _n —O] -2,2,6,6-tetramethylpiperidine-N-oxyl (where n = 1 to 18); It is done. The amount of the polymerization inhibitor used can be set to an amount necessary for preventing polymerization of raw materials and products in the liquid phase oxidation reaction.

  As the catalyst used in the present invention, any catalyst having catalytic activity in the liquid phase oxidation reaction may be used. Moreover, although either a heterogeneous catalyst or a homogeneous catalyst can be used, a heterogeneous catalyst is preferred, and a noble metal-containing catalyst is more preferred. For example, it is preferable to use a noble metal-containing catalyst prepared by the following method, but a commercially available product may be used.

  The noble metal-containing catalyst can be prepared by dissolving a noble metal compound in a solvent and reducing it using a reducing agent. The target noble metal-containing catalyst is precipitated by this reduction. A supported catalyst in which a noble metal is supported on a carrier can be used, but a non-supported catalyst may also be used. The reduction can be performed in the gas phase, but is preferably performed in the liquid phase. Hereinafter, a liquid phase reduction method for reducing a noble metal compound in a liquid phase will be described.

  The noble metal contained in the noble metal-containing catalyst refers to palladium, platinum, rhodium, ruthenium, iridium, gold, silver, rhenium, and osmium. Of these, palladium, platinum, rhodium, ruthenium, iridium and gold are preferable, and palladium is particularly preferable. The noble metal compound used for the production of the noble metal-containing catalyst is not particularly limited. For example, noble metal chlorides, oxides, acetates, nitrates, sulfates, tetraammine complexes, acetylacetonate complexes, etc. are preferable. More preferred are oxides, oxides, acetates, nitrates and sulfates, and particularly preferred are noble metal chlorides, acetates and nitrates.

  First, the above noble metal compound is dissolved in a solvent to obtain a noble metal compound solution. As this solvent, one or two or more solvents selected from the group consisting of water, alcohols, ketones, organic acids and hydrocarbons can be used. The concentration of the noble metal compound in the noble metal compound solution is preferably 0.1% by mass or more, more preferably 0.2% by mass or more, and particularly preferably 0.5% by mass or more. Moreover, 20 mass% or less is preferable, 10 mass% or less is more preferable, and 7 mass% or less is especially preferable.

  Next, a reducing agent is added to the noble metal compound solution to reduce the noble metal compound. The reducing agent to be used is not particularly limited, and examples thereof include hydrazine, formaldehyde, sodium borohydride, hydrogen, formic acid, formic acid salts, ethylene, propylene, and isobutylene.

  When the reducing agent is a gas, it is preferably carried out in a pressurizing apparatus such as an autoclave in order to increase the solubility in the noble metal compound solution. In that case, it is preferable to pressurize the inside of a pressurization apparatus with a reducing agent. The pressure is usually 0.2 to 1.1 MPa (gauge pressure).

  When the reducing agent is a liquid, there is no limitation on the apparatus for reducing the noble metal, and the reduction can be performed by adding the reducing agent to the noble metal compound solution. Although the usage-amount of a reducing agent at this time is not specifically limited, It is about 1-100 mol normally with respect to 1 mol of noble metal compounds.

  The temperature and reduction time of the system at the time of reduction vary depending on the reduction method, the noble metal compound used, the solvent, the reducing agent, etc., but in general, in the case of the liquid phase reduction method, the reduction temperature is 0 to 100 ° C., and the reduction time is 0.5-24 hours are preferred.

  When preparing a supported catalyst, it may be reduced in the same manner as when the support is not used, except that the support is dispersed in the noble metal compound solution. Examples of the carrier include activated carbon, carbon black, silica, alumina, magnesia, calcia, titania and zirconia. Among them, activated carbon, silica and alumina are preferably used. The loading ratio of the noble metal is preferably 0.1% by mass or more, more preferably 1% by mass or more, further preferably 2% by mass or more, and particularly preferably 4% by mass or more with respect to the mass of the carrier before loading. Moreover, 40 mass% or less is preferable, 30 mass% or less is more preferable, 20 mass% or less is further more preferable, and 15 mass% or less is especially preferable.

  The precipitate (noble metal-containing catalyst) deposited by reduction is usually filtered by a method such as filtration or centrifugation. The separated noble metal-containing catalyst is appropriately dried. The drying method is not particularly limited, and various methods can be used. The physical properties of the prepared noble metal-containing catalyst can be confirmed by BET surface area measurement, XRD measurement, CO pulse adsorption method, TEM measurement, and the like.

  The noble metal-containing catalyst is preferably used in a state of being suspended in a reaction solution for performing a liquid phase oxidation reaction. The amount of the noble metal-containing catalyst in the reaction solution is preferably 0.01% by mass or more as the catalyst existing in the reaction vessel with respect to 100% by mass of the solution existing in the reaction vessel in which the liquid phase oxidation reaction is performed. More preferably 2% by mass or more. Moreover, 50 mass% or less is preferable, and 30 mass% or less is more preferable.

  In the present invention, an oxygen-containing gas is supplied to the liquid phase part of the reaction vessel, and the olefin or α, β-unsaturated aldehyde is oxidized by molecular oxygen in the oxygen-containing gas. At this time, while supplying an inert gas-containing gas having a lower oxygen concentration than the oxygen-containing gas to the gas phase portion of the reaction vessel, unreacted oxygen that has moved from the liquid phase portion to the gas phase portion is exhausted to the outside of the reaction vessel. Discharge.

  The oxygen-containing gas preferably has an oxygen concentration exceeding 21% by volume, and more preferably has an oxygen concentration of 50% by volume or more and 100% by volume or less. The oxygen-containing gas can be adjusted to a predetermined oxygen concentration by containing a gas such as nitrogen, carbon dioxide, water vapor, or the like, as required, with respect to the high-concentration oxygen gas obtained from the oxygen production facility.

  In the present invention, the oxygen-containing gas is supplied to the liquid phase part. This method includes, for example, a method of supplying the liquid phase in the form of fine bubbles using a gas disperser. Examples of the gas distributor include a perforated plate, a nozzle, and a porous plate. The superficial velocity of the oxygen-containing gas is preferably 0.2 cm / s or more, and more preferably 0.5 cm / s or more. Moreover, 30 cm / s or less is preferable and 25 cm / s or less is more preferable. The amount of molecular oxygen contained in the oxygen-containing gas is preferably 0.1 mol or more, more preferably 0.3 mol or more, more preferably 0.5 mol or more with respect to 1 mol of the olefin or α, β-unsaturated aldehyde. Is more preferable. Moreover, 30 mol or less is preferable, 25 mol or less is more preferable, and 20 mol or less is further more preferable.

  In the liquid phase oxidation reaction, unreacted oxygen that has passed through the liquid phase part in the reaction vessel and moved to the gas phase part may form a combustible mixture in the gas phase part. The exhaust gas is discharged out of the reaction vessel as exhaust gas while supplying an inert gas-containing gas having a lower oxygen concentration than the oxygen-containing gas supplied to the liquid phase. That is, after unreacted oxygen that has moved to the gas phase is diluted with an inert gas-containing gas, the gas present in the gas phase is discharged out of the reaction vessel as exhaust gas. The inert gas-containing gas preferably has an oxygen concentration of 6% by volume or less, more preferably 3% by volume or less, and more preferably 0% by volume or more. As the inert gas, nitrogen, carbon dioxide, water vapor or the like can be used, and in this range, oxygen and / or a combustible mixture is not formed within a range where the oxygen concentration is lower than that of the oxygen-containing gas used for the liquid phase. A flammable substance may be contained. Explosive limits of combustible mixtures include lower explosive lower limit concentration, upper explosive upper limit concentration focusing on combustible substance gas concentration, and upper limit oxygen concentration focusing on oxygen concentration, but the critical oxygen concentration is used for safety management . Since the critical oxygen concentration varies depending on the reaction solution composition, reaction temperature, and pressure, it is necessary to obtain a value in the system in advance and manage it below that value.

  As a method for supplying the inert gas-containing gas, from the viewpoint of avoiding an explosion, the inert gas-containing gas is supplied to the gas phase portion in advance and continuously after the start of the supply of the oxygen-containing gas to the liquid phase portion. The method of supplying to is preferable. As a method for supplying the oxygen-containing gas, a method of continuously or intermittently supplying the liquid phase portion is preferable. At this time, it is preferable that the oxygen concentration in the gas phase part is continuously monitored using an oxygen meter such as a zirconia type, a galvanic type, or a magnetic type.

  A gas obtained by diluting unreacted oxygen with the inert gas-containing gas is then discharged out of the reaction vessel as exhaust gas. Normally, low-concentration raw materials and reaction solvents exist in exhaust gas. Therefore, from the viewpoint of air pollution prevention or cost, after collecting the raw materials and reaction solvents, they are processed into an incinerator and then released into the atmosphere. The Examples of the recovery method include an absorption method and an adsorption method.

  In the present invention, as a result of investigations by the present inventors, it has been found that the rate of the oxidation reaction depends on the oxygen partial pressure in the bubbles generated in the liquid phase by the oxygen-containing gas provided to the liquid phase part (hereinafter referred to as the pressure). The unit is expressed in absolute pressure). The oxygen partial pressure in the bubbles is preferably 0.05 MPa or more, and more preferably 0.1 MPa or more. Moreover, from an economical viewpoint, 3 MPa or less is preferable and 2.5 MPa or less is more preferable. The oxygen partial pressure in the bubbles is calculated by the following formula (1).

_{p O2b = (F v p O2v} -F d p O2d) / (F v -F d -F f) ··· (1)
Here, p _O2b: oxygen partial pressure in the bubble (MPa), p _O2v: partial pressure of oxygen in the exhaust gas (MPa), p _O2d: oxygen partial pressure of the inert gas-containing gas supplied to the gas phase (MPa ), F _v : exhaust gas flow rate (mol / h), F _d : inert gas-containing gas flow rate (mol / h) used for the gas phase portion, F _f : gas flow rate accompanying the inert gas-containing gas (mol / h) h).

  Hereinafter, Formula (1) will be described in detail. FIG. 2 schematically shows the relationship between the physical quantities in a state in which the gas present in the gas phase part is discharged as exhaust gas while supplying the oxygen-containing gas and the inert gas-containing gas to the liquid phase part and the gas phase part, respectively. This is an example of the target. When the total mass balance in the gas phase is taken, the relationship of the following formula (2) is obtained. Moreover, the oxygen balance formula which paid attention to oxygen is represented by Formula (3).

Formula (1) representing the oxygen partial pressure in the bubbles is obtained by eliminating F _b from formulas (2) and (3). Note that _Fb is a physical quantity that cannot be actually measured. The value of p _O2v is usually calculated by actually measuring the oxygen concentration in the exhaust gas and multiplying this by the total pressure. The value of p _O2d is usually calculated by actually measuring the oxygen concentration in the inert gas-containing gas supplied to the gas phase and multiplying this by the total pressure. Therefore, p _O2b can be calculated based on a physical quantity that can be measured by using the equation (1) in consideration of the flow rate of the inert gas-containing gas and the gas flow accompanying the inert gas-containing gas.

The gas flow rate (F _f ) accompanying the inert gas-containing gas varies depending on the type of inert gas-containing gas used and the non-condensable gas by-produced, but the non-condensable gas and the inert gas (F Formulation is possible by assuming that the ratio of _d ) is the ratio of the condensable gas in the bubbles to the condensable gas (F _f ) accompanying the inert gas. For example, when pure oxygen is used as the oxygen-containing gas, nitrogen gas containing a small amount of oxygen gas is used as the inert gas-containing gas, and CO ₂ (non-condensable gas) is produced as a by-product, the distribution of the following formula (4) The ratio holds. Whether or not the gas condenses varies depending on the temperature and pressure of the target system. In the description of the present invention, the non-condensable gas means that the boiling point of a pure substance at atmospheric pressure is −50 ° C. or less. Gas.

F _f is first converted from the amount of bubbles into the amount of exhaust gas according to equation (3), then substituted into equation (2), and F _b is eliminated to eliminate equation (5). It is expressed as follows. Note that the relationship of Equation (3) holds for CO ₂ as well as O ₂ .

As for the meaning of equation (5), the first half of the right side represents the ratio of the inert gas in the total amount of non-condensable gas, and the second half of the right side (in the braces) represents the gas phase by F _b and F _f. Represents the total amount of condensable gas entrained in Each quantity on the right side is available based on actual measurement, and F _f can be calculated from these. Equation (5) is an example showing a method for calculating F _f , but F _f is obtained by determining the distribution ratio in the same manner as described above even if the type of inert gas and the gas components by-produced are changed. be able to.

  The reaction pressure (total pressure) is not particularly limited as long as it is a pressure capable of obtaining the oxygen partial pressure in the bubbles as described above, but is preferably 0.10 MPa or more, and more preferably 0.15 MPa or more. Moreover, 10 MPa or less is preferable and 8 MPa or less is more preferable.

  The temperature at which the liquid phase oxidation reaction is performed is appropriately selected depending on the reaction solvent and the raw material used. The reaction temperature is preferably 60 ° C. or higher, more preferably 70 ° C. or higher. Moreover, 200 degrees C or less is preferable and 150 degrees C or less is more preferable.

  The liquid phase oxidation reaction can be carried out by either a batch method or a continuous method, but the industrial method is preferably used industrially. When the noble metal-containing catalyst is suspended in the reaction solution and subjected to the liquid phase oxidation reaction, for example, a bubble column reactor equipped with a reaction vessel having a liquid phase part and a gas phase part, a stirred tank reactor, etc. Is preferably used. If necessary, two or more reaction vessels may be arranged in a multistage series, and the liquid phase oxidation reaction may be carried out so that the reaction solution sequentially flows through each tank. The residence time of the reaction solution in the reaction vessel can be appropriately selected depending on the amount of the noble metal-containing catalyst, the reaction temperature, the pressure, etc., but is usually 0.1 hr or more, preferably 0.2 hr or more, and 10 hr or less, preferably 8 hr or less. It is.

  When the particle size of the noble metal-containing catalyst is large and the specific gravity is large, sedimentation separation by gravity can be performed. However, when the particle size of the noble metal-containing catalyst is small and the specific gravity is small, centrifugal force is used instead of gravity. It is preferable to use filtration that separates the filter cake and the filtrate by a partition called a filter medium using centrifugal separation, gravity, pressurization, or reduced pressure. When using a series multi-stage tank reactor, place a catalyst separator at the outlet of each tank to separate the noble metal-containing catalyst and the reaction mixture, return the noble metal-containing catalyst to the original tank, It is preferable to send the solution to the next tank.

  In the method of the present invention, it is preferable to separate the olefin and the α, β-unsaturated aldehyde in the reaction solution. The olefin and the α, β-unsaturated aldehyde may be separated at the same time if possible, but it is preferable to carry out the separation of the olefin and the separation of the α, β-unsaturated aldehyde in order of the boiling point.

  Since the olefin preferably used in the present invention has a low boiling point, the separation of the olefin is preferably carried out by equilibrium flash distillation. Equilibrium flash distillation is an operation of separating the mixture into gas and liquid corresponding to temperature and pressure by flushing (injecting) the mixture into the separator through a pressure reducing valve (heating as necessary). In the present invention, the temperature of the reaction solution before flushing is preferably set within the range of reaction temperature ± 15 ° C., and the pressure is preferably set within the range of reaction pressure ± reaction pressure × 0.3. It is more preferable to flush the pressure reaction liquid as it is. The pressure in the separator is preferably in the range of 0.1 MPa to reaction pressure × 0.95. By the equilibrium flash distillation, it is separated into vapor mainly composed of olefin and other liquid. Equilibrium flash distillation is a separation operation in one stage of the separation tower, and therefore the vapor usually contains α, β-unsaturated aldehyde, a solvent component, dissolved gas, and the like. When a higher degree of separation is required, a normal distillation column such as a plate column or a packed column can be used. The vapor is preferably liquefied by compression or cooling operation or absorbed by using a solvent such as a solvent having high olefin solubility, and then recycled to the reaction vessel for use in the reaction.

  Subsequent to the separation of the olefin, a separation of the α, β-unsaturated aldehyde can be carried out. This separation can be performed by a conventional distillation operation, extraction operation, membrane separation operation or the like, but it is preferable to use a continuous multistage distillation column. The continuous multi-stage distillation column is preferably a distillation column having a number of stages including a reboiler and a condenser of 3 or more, more preferably 4 or more, and even more preferably 5 or more. Examples of such distillation columns include bubble bell trays, perforated plate trays, bubble trays, jet trays, etc., in which gas and liquid are brought into contact in a cross flow; turbo grid trays in which gas and liquid are brought into contact in countercurrent , Ripple trays, dual flow trays, kittel trays, etc .; plate towers such as baffle trays, disc donut trays, etc. that are otherwise in contact with gas and liquid; and for example, Raschig rings, pole rings, interlock saddles, Dickson A generally used distillation column such as a packed column packed with an irregular packing material such as packing or McMahon packing and a regular packing material typified by a throughr packing can be used. The number of plates above indicates the number of theoretical plates for both the tower column and packed column. The column bottom temperature is preferably 150 ° C. or lower, more preferably 120 ° C. or lower, and still more preferably 90 ° C. or lower, from the viewpoint of preventing polymerization. The operating pressure is selected from atmospheric pressure, pressurized pressure, and reduced pressure within the preferable temperature range at the bottom of the column. If necessary, a part of the column top liquid can be refluxed to the distillation column, and the reflux ratio is preferably 0.05 or more, more preferably 0.1 or more, preferably 20 or less, more preferably 15 or less. In order to prevent polymerization in the tower, it is preferable to use the polymerization inhibitor. If necessary, molecular oxygen air bubbling can be used in combination. The α, β-unsaturated aldehyde obtained from the top of the column may contain a solvent. The column top liquid containing the α, β-unsaturated aldehyde is preferably recycled to the reaction vessel and used again for the reaction. In this case, α, β-unsaturated carboxylic acid may be contained in the column top liquid.

  In the method of the present invention, it is preferable that the solvent is subsequently separated from the reaction mixture after the separation of the α, β-unsaturated aldehyde. This separation can be performed by a conventional distillation operation, extraction operation, membrane separation operation, etc., in the same manner as the above-described separation of the α, β-unsaturated aldehyde. When the separation by distillation operation is described as an example, for example, separation by a continuous multistage distillation column is preferably used. As such a distillation column and its operating conditions, it is preferable to use the above-mentioned distillation column and its operating conditions. Here, the separated solvent is preferably recovered and used again for the reaction.

  In the method of the present invention, it is preferable that the α, β-unsaturated carboxylic acid is subsequently separated from the reaction mixture after separating the solvent. This separation can be carried out by a conventional distillation operation, extraction operation, membrane separation operation, etc., as in the above-described separation of the α, β-unsaturated aldehyde, but separation by distillation operation is preferably used. Separation by a continuous multistage distillation column is preferable, and as such a distillation column and its operating conditions, it is preferable to use the above-mentioned distillation column and its operating conditions.

  Thereby, crude α, β-unsaturated carboxylic acid can be obtained from the top of the column. When the target product α, β-unsaturated carboxylic acid is (meth) acrylic acid, the crude (meth) acrylic acid obtained here is esterified with methanol to produce methyl (meth) acrylate. Or purified by a crystallization operation or the like to produce high purity (meth) acrylic acid. From the bottom of the column, a high boiling point component containing α, β-unsaturated carboxylic acid anhydride or the like is obtained. The bottom liquid may contain α, β-unsaturated carboxylic acid. The content is preferably 0.1% by mass or more, more preferably 0.5% by mass or more, and preferably 10% by mass or less, more preferably 5% by mass or less, with respect to the mass of the column bottom liquid. is there.

  Next, preferred embodiments of the present invention and apparatus configurations for performing the same will be described with reference to the drawings. However, the present invention is not limited to the following contents.

  FIG. 1 is a system diagram showing the configuration of an apparatus for producing an α, β-unsaturated carboxylic acid by oxidizing an olefin. In FIG. 1, an olefin / solvent supply line 1 (the olefin and the solvent may be supplied through dedicated lines), an inert gas-containing gas supply line 2 and exhaust gas are provided above the liquid phase oxidation reactor 101. Line 4, catalyst circulation line 6, olefin circulation line 8, aldehyde circulation line 10 and solvent circulation line 12 are connected, and at the bottom of liquid phase oxidation reactor 101, oxygen-containing gas supply line 3, reaction liquid extraction line 5 is connected.

  In order to carry out the liquid phase oxidation reaction with such an apparatus, first, a solvent and a catalyst (a catalyst supply line is not shown) are charged into the liquid phase oxidation reactor 101 to prevent the formation of a squealing gas in the gas phase. An inert gas-containing gas is introduced from the line 2 into the gas phase portion of the liquid phase oxidation reactor. Next, the olefin and the solvent are supplied from the line 1 to the liquid phase oxidation reactor 101, the oxygen-containing gas is introduced from the line 3, and the inert gas-containing gas is supplied from the line 2. I do. The gas obtained by diluting the unreacted oxygen gas with the inert gas-containing gas is discharged from the line 4. An absorption tower (not shown) for recovering unreacted olefin components, solvent components, and the like can be provided in the middle of the line 4. The olefin component recovered by the absorption tower is returned to the liquid phase oxidation reactor 101. When high concentration oxygen gas is used, the load on the absorption tower can be reduced and the amount of olefin recycled can be reduced.

  The obtained reaction liquid is extracted along with the catalyst through the line 5, and the catalyst is separated in the solid-liquid separator 102. The separated catalyst is returned to the liquid phase oxidation reactor 101 from the line 6, and the reaction liquid from which the catalyst has been removed is sent from the reaction liquid line 7 to the flash column 103 after the catalyst separation. The gas that is flushed in the flash column 103 and rich in olefin components is returned to the liquid phase oxidation reactor 101 from line 8. On the other hand, the reaction liquid from which the olefin has been separated is extracted from the bottom of the flash tower 103 and sent from the reaction liquid line 9 to the multistage distillation tower 104 after the olefin separation.

  In the multistage distillation column 104, a fraction mainly containing α, β-unsaturated aldehyde is returned from the line 10 to the liquid phase oxidation reactor 101 from the top of the column. On the other hand, the reaction solution from which the α, β-unsaturated aldehyde has been separated is withdrawn from the bottom of the multistage distillation column 104 and sent from the reaction solution line 11 to the multistage distillation column 105 after the aldehyde separation.

  In the multistage distillation column 105, a fraction containing mainly a solvent is returned from the line 12 to the liquid phase oxidation reactor 101 from the top of the column. On the other hand, the reaction solution from which the solvent has been separated is withdrawn from the bottom of the multistage distillation column 105, and is sent from the reaction solution line 13 to the multistage distillation column 106 after the solvent separation. In order to remove water generated by the liquid phase oxidation reaction, a dehydration tower (not shown) can be provided in the middle of the line 13. The solvent loss is replenished from line 1.

  In the multistage distillation column 106, α, β-unsaturated carboxylic acid is extracted from the top of the column through the carboxylic acid extraction line 14. Further, a residue and a high-boiling component are extracted from the high-boiling component line 15 from the bottom of the multistage distillation column 106.

  Hereinafter, the present invention will be described in more detail by way of examples.

The conversion of isobutylene (IB), methacrolein (MAL), methacrylic acid (MAA), methacrylic anhydride (ANH) and CO ₂ selectivity, and the productivity of methacrylic acid are defined as follows.
Conversion of isobutylene (%) = (B / A) × 100
Selectivity of methacrolein (%) = (C / B) x 100
Methacrylic acid selectivity (%) = (D / B) × 100
Selectivity of methacrylic anhydride (%) = (2 × E / B) × 100
CO ₂ selectivity (%) = (1/4 × F / B) × 100
Methacrylic acid productivity (g-MAA / (g-Pd · h))
= ((D × MwM) + (E × MwA)) / PD
Here, A is the molar flow rate (mol / h) of the supplied isobutylene, B is the molar flow rate (mol / h) of the disappeared isobutylene, C is the molar flow rate (mol / h) of the produced methacrolein, and D is generated. M is the molar flow rate of methacrylic acid (mol / h), E is the molar flow rate of generated methacrylic anhydride (mol / h), F is the molar flow rate of generated CO ₂ (mol / h), MwM is the molecular weight of methacrylic acid, MwA Is the molecular weight of methacrylic anhydride and PD is the mass (g) of palladium used in the reaction.

The isobutylene volatilization rate and oxygen (O ₂ ) conversion rate are defined below.
Isobutylene volatility (%) = (G / A) × 100
Oxygen conversion (%) = (1-H / J) × 100
Here, G is the molar flow rate (mol / h) of isobutylene in the exhaust gas, H is the molar flow rate (mol / h) of oxygen in the exhaust gas, and J is the molar flow rate (mol / h) of supplied oxygen.

<Example 1>
(Liquid-phase oxidation reaction of isobutylene)
As a reaction vessel for conducting the liquid phase oxidation reaction, a jacketed stainless steel stirred tank reactor having an inner diameter of 126 mm and a capacity of 4 liters was used. The raw material is continuously supplied together with the solvent from the upper part of the reaction vessel, and the reaction liquid is continuously extracted from the system while keeping the liquid surface of the liquid phase part constant. The oxygen concentration in the exhaust gas was constantly monitored with a magnetic oxygen meter (manufactured by Yokogawa Electric).

  300 g of a palladium metal-containing silica-supported catalyst (prepared by a method in which palladium nitrate is dissolved in water and reduced with formaldehyde and supported on a silica carrier) in a reaction vessel (of which the palladium mass is 50 g), A 75% by mass tertiary butanol aqueous solution was added as a solvent so as to reach the control liquid level (the liquid level was adjusted so that the liquid volume was 3 liters). High-purity nitrogen gas (purity 99.99% by volume or more) as an inert gas-containing gas is supplied from the upper part of the reaction vessel to the gas phase at 614 g / hr to increase the pressure to 1.1 MPa (total pressure), and then pressure control The amount of exhaust gas was controlled by the apparatus so as to maintain this pressure. After raising the temperature of the liquid phase to 110 ° C. and stabilizing for about 10 minutes, 461 g / hr of liquefied isobutylene and a solvent (75% by mass tertiary butanol aqueous solution, p-methoxyphenol 200 ppm as a polymerization inhibitor) Prepared) was continuously fed to the reaction vessel at 3688 g / hr. The average residence time at this time was 1.0 hour.

  Next, high-purity oxygen gas (purity: 99.6% by volume or more) is supplied as an oxygen-containing gas from a compressed oxygen cylinder to the liquid phase portion of the reaction vessel while maintaining the supply of nitrogen gas and maintaining the gas phase pressure. The liquid phase oxidation reaction was started by continuously feeding through a sparger made of sintered metal. In the gas phase portion, the supply amount of high-purity oxygen gas was controlled so that the oxygen concentration in a state where unreacted oxygen was diluted with nitrogen gas was maintained at about 6% by volume. At the time of reaction for 4, 5 and 6 hours after the start of the reaction, the reaction solution and exhaust gas were sampled and analyzed respectively. The raw materials and products were analyzed using gas chromatography (manufactured by Shimadzu Corporation) equipped with a FID or TCD detector.

After 4 hours from the start of the reaction, it was confirmed that the reaction solution composition and the exhaust gas composition were almost steady. Therefore, the analysis values of the reaction liquid and the exhaust gas were values obtained by arithmetically averaging the data after 4, 5 and 6 hours from the start of the reaction. Table 1 shows the results. Table 1 also shows the main conditions. The analysis results in the exhaust gas were an oxygen concentration of 6.1% by volume, an isobutylene concentration of 6.7% by volume, and a CO ₂ concentration of 8.1% by volume.

<Examples 2 and 3>
In Example 1, it implemented similarly except having changed the total pressure to 2.5 MPa and 4.9 MPa, respectively. The results are shown in Table 1.

<Comparative Example 1>
In Example 1, except that the oxygen-containing gas is changed to air (commercial compressed air cylinder; oxygen concentration 21% by volume), the total pressure is 6.3 MPa, and the inert gas-containing gas is not supplied to the gas phase part. It carried out similarly. The results are shown in Table 1. Since the oxygen partial pressure in the bubbles is almost the same as in Example 1, it seems that the selectivity and MAA productivity are equivalent. However, due to air oxidation, the total pressure must be significantly increased. Further, compared with Example 1, the oxygen conversion rate was significantly reduced.

<Comparative Example 2>
In Example 2, it implemented similarly except having changed oxygen containing gas into air (commercial compressed air cylinder; oxygen concentration 21 volume%), and not supplying an inert gas containing gas to a gaseous-phase part. The results are shown in Table 1. The total pressure was the same as in Example 2, but the isobutylene conversion rate and oxygen conversion rate were significantly reduced, resulting in an increase in isobutylene volatilization rate.

<Comparative Example 3>
In Example 3, the same procedure was performed except that the oxygen-containing gas was changed to air (commercially compressed air cylinder; oxygen concentration 21% by volume) and the inert gas-containing gas was not supplied to the gas phase part. The results are shown in Table 1. The total pressure was the same as in Example 3, but the isobutylene conversion rate and oxygen conversion rate were significantly reduced, resulting in an increase in isobutylene volatilization rate.

p _O2b and F _f are values calculated by the equations (1) and (5), respectively.

It is a systematic diagram which shows the structure of the apparatus which oxidizes an olefin and manufactures (alpha), (beta)-unsaturated carboxylic acid. The relationship between each physical quantity was schematically shown in a state where the gas present in the gas phase part was discharged as exhaust gas while supplying the oxygen-containing gas and the inert gas-containing gas to the liquid phase part and the gas phase part, respectively. It is an example.

Explanation of symbols

DESCRIPTION OF SYMBOLS 1 Olefin and solvent supply line 2 Inert gas containing gas supply line 3 Oxygen containing gas supply line 4 Exhaust gas line 5 Reaction liquid extraction line 6 Catalyst circulation line 7 Reaction liquid line after catalyst separation 8 Olefin circulation line 9 Reaction liquid after olefin separation Line 10 Aldehyde circulation line 11 Reaction liquid line after aldehyde separation 12 Solvent circulation line 13 Reaction liquid line after solvent separation 14 Carboxylic acid extraction line 15 High boiling point component line 101 Liquid phase oxidation reactor 102 Solid-liquid separator 103 Flash tower 104 Multistage distillation Tower 105 multistage distillation tower 106 multistage distillation tower

Claims (2)

In a method for producing an α, β-unsaturated carboxylic acid by oxidizing an olefin or α, β-unsaturated aldehyde in a liquid phase in the presence of a catalyst in a reaction vessel,
Providing an oxygen-containing gas to the liquid phase part in the reaction vessel to oxidize the olefin or the α, β-unsaturated aldehyde;
While supplying an inert gas-containing gas having an oxygen concentration lower than that of the oxygen-containing gas to the gas phase portion in the reaction vessel, unreacted oxygen moved from the liquid phase portion to the gas phase portion is discharged as an exhaust gas to the outside of the reaction vessel. And a step of discharging the α, β-unsaturated carboxylic acid. The α, β-unsaturated carboxylic acid according to claim 1, wherein the oxygen partial pressure in the bubbles generated in the liquid phase by the oxygen-containing gas calculated by the following formula (1) is 0.05 MPa or more and 3 MPa or less. Manufacturing method.
_{p O2b = (F v p O2v} -F d p O2d) / (F v -F d -F f) ··· (1)
Here, p _O2b: oxygen partial pressure in the bubble (MPa), p _O2v: partial pressure of oxygen in the exhaust gas (MPa), p _O2d: oxygen partial pressure of the inert gas-containing gas supplied to the gas phase (MPa ), F _v : exhaust gas flow rate (mol / h), F _d : inert gas-containing gas flow rate (mol / h) used for the gas phase portion, F _f : gas flow rate accompanying the inert gas-containing gas (mol / h) h).

Images