JP3979256B2 - Process for producing aromatic carbonates - Google Patents

Description

[0001]
BACKGROUND OF THE INVENTION
The present invention relates to a method for producing aromatic carbonates. Specifically, the present invention relates to a method for efficiently and continuously producing an alkylaryl carbonate and / or diaryl carbonate from an aromatic hydroxy compound and a dialkyl carbonate by a transesterification reaction.
[0002]
[Prior art]
Until now, it has been well known to produce an alkylaryl carbonate by producing a transesterification reaction between a dialkyl carbonate and an aromatic hydroxy compound, or to produce a diaryl carbonate from an alkylaryl carbonate.
[0003]
These typical reactions are represented by the following formulas (1) to (3).
[0004]
[Chemical 1]
R ¹ —OCOO—R ² + ArOH → R ¹ —OCOO—Ar + R ² OH (1)
R ¹ —OCOO—Ar + ArOH → Ar—OCOO—Ar + R ¹ OH (2)
2R ¹ —OCOO—Ar → Ar—OCOO—Ar + R ¹ —OCOO—R ¹ (3)
(In the formula, R ¹ and R ² represent an aliphatic hydrocarbon group or an alicyclic hydrocarbon group, and may be the same or different. Ar represents an aromatic hydrocarbon group.)
Such a transesterification reaction is an equilibrium reaction and tends to proceed in a direction in which a highly nucleophilic substituent is replaced with a weakly nucleophilic substituent. When dialkyl carbonate having a lower aliphatic hydrocarbon group is used as a raw material and phenol is used as an aromatic hydroxy compound, the reactions of formulas (1) and (2) are particularly contrary to the principle. It is very large and biased toward the original system, and the reaction rate is generally slow. When a general transesterification catalyst such as an alkali metal hydroxide is used, a reaction represented by the following formula (4) accompanied by decarboxylation becomes dominant, and the reaction yield is significantly reduced.
[0005]
[Chemical formula 2]
R ¹ —OCOO—R ² + ArOH → R ¹ —O—Ar + R ² OH + CO ₂ (4)
In order to advance the reaction for producing aromatic carbonates efficiently, search for a highly active catalyst has been conventionally made, and various catalysts have been proposed. No. 48733) or an organic titanium compound (for example, JP-A-57-183745) can be used to advance the above reaction.
[0006]
In order to produce aromatic carbonates more efficiently, it is necessary to quickly remove the product and shift the equilibrium to the production system as much as possible. As an attempt to efficiently remove the by-produced aliphatic alcohol, a method of distilling with an azeotropic agent (for example, JP-A-54-48732 and JP-A-61-291545), and adsorption removal with a molecular sieve. Methods (for example, Japanese Patent Laid-Open No. 58-185536), methods using an osmotic vaporization method or a vapor vaporization method (for example, Japanese Patent Laid-Open No. 5-1225021) have been proposed. There are drawbacks such as being unreasonable and the process becoming complicated, and it is not suitable as an industrial method.
[0007]
In the case of an equilibrium reaction such as the above formula (1), a raw material conversion rate higher than the equilibrium composition cannot be obtained in a fully mixed reactor. Therefore, it is effective to use a series of multistage reactors to remove products from each stage and to react while increasing the conversion stepwise. For the same reason, it is known that a reactive distillation that can continuously extract a product is effective. For example, in Japanese Examined Patent Publication No. 7-91236, a continuous multi-stage distillation column is brought into contact with the aromatic hydroxy compound in liquid form from the top of the tower and dialkyl carbonate in gaseous form from the bottom of the tower in countercurrent, and by-produced alcohol and dialkyl While extracting low-boiling components containing carbonate from the top of the tower, high-boiling components containing aromatic carbonate are extracted from the bottom of the tower.
[0008]
In principle, this method allows the reaction to proceed with a relatively simple process, but the reaction is slow and is a liquid phase reaction, so that a continuous multistage distillation column must have sufficient reaction time. Is difficult. In order to secure the reaction time, a method of adding an additional reactor has been devised (for example, JP-A-4-224547 and JP-A-4-230242), but as a result, the equipment is complicated and expensive. It becomes a thing.
[0009]
As a method for obtaining the same effect as reactive distillation, a method in which a reaction is carried out while continuously countercurrently contacting gas and liquid in two or more stirring tanks connected in series (for example, JP-A-6-234707, patent document) 1) and a method of performing a reaction in a bubble column reactor or a cascade of at least two bubble columns (for example, JP-A-6-298700). Although these methods have an advantage that the reaction time can be freely designed, there is a disadvantage that a sufficient conversion rate cannot be obtained if the number of reaction stages is small, and that the equipment cost increases if the number of stages is large.
[0010]
In general, a reactive distillation type reaction facility performs a multistage reaction, and performs reaction and gas-liquid separation while continuously countercurrently contacting the vapor generated in the lower or subsequent stage and the reaction solution coming from the upper or previous stage. There is an advantage that the energy given in can be efficiently transmitted to the upper stage or the previous stage. On the other hand, low-boiling products (in this case, aliphatic alcohols) are concentrated toward the upper stage or the front stage, and high-boiling products (in this case, aromatic carbonates) are concentrated toward the lower stage or the rear stage. In addition, since the raw material ratio changes continuously, it is difficult to control the reaction, and the design and operation of the reactor are difficult.
[0011]
As long as the reaction is carried out while removing the produced aliphatic alcohol, there is no need for a reactive distillation type facility. However, since the dialkyl carbonate of the low boiling point raw material is distilled off together with the generated aliphatic alcohol, the energy efficiency is extremely deteriorated unless the low boiling point raw material is recycled while effectively recovering the energy of the vapor.
[0012]
In order to reduce equipment costs while conducting multistage reactions, the liquid phase part of the reactor is separated by a partition wall, and a light boiling fraction containing aliphatic alcohol by-produced with the gas phase as a continuous phase is gas phased from the top of the reactor. There has been proposed a method in which the reaction is carried out while being continuously extracted (see, for example, JP-A-8-188558 and Patent Document 2). In this method, although the temperature and composition of the liquid phase can be changed in each reaction zone and can be a multistage reaction, since the gas phase in each reaction zone is continuous, appropriate separation according to the vapor composition is possible. There is a disadvantage that the recovery process and energy recovery cannot be performed.
[0013]
[Patent Document 1]
JP-A-6-234707 [Patent Document 2]
JP-A-8-188558 [0014]
[Problems to be solved by the invention]
In such a conventional method for producing alkylaryl carbonate and / or diaryl carbonate by transesterification, it is difficult to proceed the reaction with high selectivity and high efficiency, and complicated processes and large equipment costs are required. There was a problem that there was.
The present invention is a method for producing an alkylaryl carbonate and / or diaryl carbonate from a dialkyl carbonate and an aromatic hydroxy compound in the presence of a catalyst, without the above-mentioned drawbacks, and efficiently and continuously produced with high selectivity. The challenge is to provide a way to do this.
[0015]
[Means for Solving the Problems]
As a result of intensive investigations to solve the above problems, the present inventor is an equilibrium reaction in the liquid phase, and the progress of the reaction can be promoted by efficiently extracting the aliphatic alcohol of the low boiling point product. By using a reactor with a specific structure, the heat of the reaction vapor can be recovered by indirect heat exchange, thereby avoiding the disadvantages of reactive distillation, such as delayed reaction due to concentration of low-boiling byproducts. However, the present inventors have found that energy efficiency can be improved and reached the present invention.
[0016]
That is, the gist of the present invention is to supply an alkylaryl carbonate and / or diaryl carbonate by a transesterification reaction between a dialkyl carbonate and an aromatic hydroxy compound in the presence of a catalyst in the reactor. A reactor having a structure in which a first heat recovery section configured to indirectly exchange heat between a raw material liquid to be generated and a vapor generated from the reactor is directly connected to a gas phase section of the reactor is used. In the method for producing aromatic carbonates, the transesterification reaction is performed in the reactor while heating the raw material liquid supplied in the first heat recovery section with the condensation heat of the vapor.
[0017]
DETAILED DESCRIPTION OF THE INVENTION
Hereinafter, embodiments of the present invention will be described in detail.
The dialkyl carbonate which is a reaction raw material of the present invention is represented by the following formula (5).
[0018]
[Chemical 3]
^{R l -OCOO-R 2 (5} )
(In the formula, R ¹ and R ² represent an alkyl group having 1 to 10 carbon atoms, and R ¹ and R ² may be the same as or different from each other.)
Specific examples include dimethyl carbonate, diethyl carbonate, dipropyl carbonate, dibutyl carbonate, and methyl ethyl carbonate. Of these, dimethyl carbonate and diethyl carbonate are particularly preferably used.
[0019]
The aromatic hydroxy compound which is the other reaction raw material of the present invention is represented by the following formula (6).
[0020]
[Formula 4]
ArOH (6)
(In the formula, Ar represents an aromatic group having 6 to 20 carbon atoms.)
Specifically, phenol, o-, m-, or p-cresol, o-, m-, or p-ethylphenol, o-, m-, or p-propylphenol, o-, m-, or Examples include p-methoxyphenol, 2,6-dimethylphenol, 2,4-dimethylphenol, 3,4-dimethylphenol, o-, m-, or p-chlorophenol, 1-naphthol, 2-naphthol and the like. Of these, phenol is particularly preferred.
[0021]
The alkylaryl carbonate which is one of the products of the method of the present invention is represented by the following formula (7).
[0022]
[Chemical formula 5]
R ³ —OCOO—Ar (7)
(In the formula, R ³ represents the same as R ¹ or R ² in Formula (5). Ar represents the same as in Formula (6).)
Specific examples include alkyl phenyl carbonates such as methyl phenyl carbonate, ethyl phenyl carbonate, propyl phenyl carbonate, butyl phenyl carbonate, and hexyl phenyl carbonate, and methyl toluyl carbonate, ethyl toluyl carbonate, methyl xylyl carbonate, or ethyl xyl carbonate. Examples include silyl carbonate, methyl chlorophenol, and ethyl chlorophenol. Of these, methyl phenyl carbonate and ethyl phenyl carbonate are particularly preferred.
[0023]
The diaryl carbonate which is one of the products of the method of the present invention is represented by the following formula (8).
[0024]
[Chemical 6]
Ar-OCOO-Ar (8)
(In the formula, Ar represents the same as in formula (6).)
Specific examples include diphenyl carbonate, ditoluyl carbonate, dixylyl carbonate, dinaphthyl carbonate, bis (chlorophenyl) carbonate and the like. Of these, diphenyl carbonate is particularly preferred.
[0025]
As the catalyst used in the present invention, any catalyst can be used as long as it promotes a transesterification reaction between a dialkyl carbonate or an alkyl aryl carbonate and an aromatic hydroxy compound and a disproportionation reaction of the alkyl aryl carbonate. For example, the following can be mentioned.
(A) Bu ₂ SnO, Ph ₂ SnO, (C ₈ H ₁₇ ) ₂ SnO, Bu ₂ Sn (OPh) ₂ , Bu ₂ Sn (OCH ₃ ) ₂ , Bu ₂ Sn (OEt) ₂ , Bu ₂ Sn (OPh) ) Tin compounds such as O (OPh) SnBu ₂ ,
(B) lead compounds such as PbO, Pb (OPh) ₂ , Pb (OCOCH ₃ ) ₂ ,
(C) Lewis acid compounds such as AlX ₃ , TiX ₃ , TiX ₄ , ZnX ₂ , FeX 3, SnX ₄ , VX _{5 and the} like (where X represents a halogen, an acetoxyl group, an alkoxyl group, an aryloxy group), and the like Examples include AlCl ₃ , Al (OPh) ₃ , TiCl ₄ , Ti (OPh) ₄ , Ti (OEt) ₄ , Ti (OPr) ₄ , Ti (OBu) ₄ ,
(D) Zirconium compounds such as Zr (acac) ₄ and ZrO ₂ (where acac represents an acetylacetone ligand),
(E) Copper compounds such as CuCl, CuCl ₂ , CuBr, CuBr ₂ , CuI, CuI ₂ , and Cu (OAc) ₂ (where Ac represents an acetyl group).
[0026]
Of these, tin compounds or titanium compounds are particularly preferred.
The method of the present invention is characterized by using a reactor equipped with a heat recovery part by indirect heat exchange directly connected to the gas phase part (usually the upper part) of the reactor.
FIG. 1 is a schematic cross-sectional view showing an example of a reaction apparatus in which such a reactor and a heat recovery unit are combined. In FIG. 1, D is a reactor, which may be a complete mixing tank, or may have a characteristic close to that of a plug flow reactor by dividing the interior with partition walls or the like. Even when the inside is divided by partition walls, the reaction vapor is taken out all at once, so that the gas phase is preferably of a type that can be freely circulated.
In the reactor, energy is required for the transesterification reaction to proceed and removal of the resulting aliphatic alcohol by evaporation. Therefore, it is preferable that the reactor is heated by installing an internal coil, an external jacket, or an external reboiler.
[0027]
Moreover, it is also preferable to stir and mix the reaction liquid in a reactor as needed. As a stirring device in this case, a mechanical stirrer using an impeller, an orifice mixer, a venturi mixer, a nozzle mixer, a jet stirrer, an airlift stirrer, an ejector stirrer, or the like can be used.
The reaction liquid is extracted from a liquid outlet line L3 installed at the bottom or side of the reactor.
A is a first heat recovery unit (heat exchange unit), in which the vapor generated in the reactor and the raw material liquid supplied to the reactor are indirectly in contact with each other across the partition wall, and the heat of the vapor Recovery is performed. The vapor generated in the reactor is cooled by the raw material liquid supplied from the line L1 and partially condensed while rising in the vapor flow path in the first heat recovery section. On the other hand, the raw material liquid is heated by the condensation heat of the vapor and then supplied to the reactor through the flow path L2.
[0028]
The structure of the first heat recovery unit A is that the raw material liquid circulates in the coiled tube, the vapor passes through the outside thereof, the vapor rises in the vertical tube, and the raw material liquid circulates outside the tube. A type or a type in which both of them exchange heat with a spiral partition wall therebetween can be used.
The heat recovery efficiency of the first heat recovery unit A depends on the temperature of the raw material liquid and the vapor, the heat transfer area, and the heat transfer coefficient, but has passed through the first heat recovery unit A without sufficiently increasing the heat recovery efficiency. The vapor may still have high energy. In that case, in addition to the first heat recovery unit, a second heat recovery unit may be installed to further recover the energy. C in FIG. 1 is such a second heat recovery unit (heat exchange unit), and usually a low-temperature fluid other than the raw material liquid is introduced from the line L5 and heated by the condensation heat of the vapor, and the heated fluid Can be removed from L6. The structure of the second heat recovery unit can be the same as that of the first heat recovery unit.
[0029]
If there is no suitable low-temperature fluid in the facility, water may be heated to steam as shown in FIG. 1 and used as a heating source for another facility. When the raw material liquid supplied to the reactor is sufficiently low in temperature, the raw material liquid is first supplied to the second heat recovery unit to perform the first heat exchange, Next, the second heat exchange can be performed by supplying the first heat recovery unit.
[0030]
In the second heat recovery unit C, the vapor that has passed through the first heat recovery unit A is cooled by the low-temperature fluid supplied from L5, and a part of the vapor is condensed. Since this condensate is lower in temperature than the vapor passing through the first heat recovery part A, if it falls as it is, it comes into contact with the vapor rising in the first heat recovery part A, lowering the temperature of the vapor and lowering the heat recovery efficiency. Therefore, a liquid collection part (B in FIG. 1) is installed at the lower part of the second heat recovery part C (the upper part of the first heat recovery part A), and the condensate is collected there, and this condensate is recovered as the first heat recovery part. Return to reactor D without contacting part A vapor.
[0031]
2 and 3 are schematic views showing examples of the structure of the liquid collection part B. FIG. 2 and 3, the collection part a constituting the liquid collection part B is installed in a slit shape with respect to the rising direction of the vapor, and the condensate falling from the second heat recovery part C is collected by the collection surface a1. The flow of the vapor which collects and rises from the 1st heat recovery part A is arrange | positioned so that it may not disturb. A hole a2 is provided in the lower part of the collection part a, and the condensate falling from the hole a2 is received by the bottle b1 of the receiving part b. The liquid collected in the receiving part b is sent to the reactor D through a pipe b2 extending downward. Alternatively, the pipe b2 may be guided to the outside, and preheated in the first heat recovery unit A together with the raw material liquid supplied from the line L1, and then sent to the reactor D.
[0032]
In the method of the present invention, it is not always necessary to use a solvent, but a solvent inert to the reaction, for example, ethers, aliphatic hydrocarbons, aromatic hydrocarbons and the like can be used.
The reaction temperature in the method of the present invention is usually 50 to 300 ° C, preferably 100 to 250 ° C. As the reaction temperature is higher, the reaction rate is improved, but by-products such as alkyl aromatic ethers tend to increase, and it is not preferable to make the reaction temperature too high. The pressure in the reactor varies depending on the type of reaction raw material used, the composition in the reactor, and the like, but it can usually be performed under pressure or reduced pressure in the range of 10 to 3000 kPa. A particularly preferable range is 50 to 2000 kPa (0.5 to 20 atm).
[0033]
The catalyst is usually dissolved in the raw material solution and supplied to the reactor. The amount of the catalyst is usually 0.0001 to 10 mol%, preferably 0.001 to 5 mol%, based on the reaction raw material to be supplied. If the amount is too small, the reaction rate becomes insufficient. If the amount is too large, the amount of by-products such as alkyl aromatic ethers tends to increase. The average residence time of the liquid in the reactor is usually 0.1 to 20 hours, preferably 0.3 to 10 hours, although it depends on other reaction conditions.
[0034]
The liquid phase withdrawn from the reactor via the line L3 can be subjected to purification means such as distillation to obtain the target alkylaryl carbonate and / or diaryl carbonate.
[0035]
【Example】
Next, specific embodiments of the present invention will be described in more detail by way of examples. However, the present invention is not limited by the following examples unless it exceeds the gist.
[Example 1]
The transesterification reaction of dimethyl carbonate and phenol was performed using the reaction apparatus shown in FIG.
A raw material liquid consisting of dimethyl carbonate, phenol and a catalyst was continuously supplied from the line L1. The raw material liquid is sent to the reactor D through L2 while being heated by the vapor rising from the reactor D in the heat exchange section A. As the heat exchange part A, a spiral heat exchanger was used. The reaction solution was continuously extracted from the line L3 while keeping the liquid level in the reactor constant. The rising vapor passes through the gap of the liquid collection part B and reaches the heat exchange part C while partially condensing by heating the raw material liquid in the heat exchange part A. As the heat exchange section C, a spiral heat exchanger was used as in the heat exchange section A. In the heat exchange part C, the water supplied from L5 was heated to partially condense while generating water vapor from L6, and the non-condensed part was extracted out of the system from L4. The liquid condensed in the heat exchange part C was collected in the liquid collection part B, and returned to the reactor D without contacting the heat exchange part A with the rising vapor.
[0036]
From the line L1, 90 parts by weight of dimethyl carbonate / hr, 94 parts by weight of phenol / hr, and 2.5 parts by weight of dibutyltin oxide as a catalyst / hr were supplied. The temperature of the raw material liquid was 60 ° C. As the reactor D, an autoclave having a volume of 0.23 m ³ (0.55 mφ * 1.0 mH) was used, and the liquid level was operated at 70%. The reactor was operated at a pressure of 400 kPa, a reaction temperature of 226 ° C., and a continuous operation for 8 hours under the condition that the residence time of the reaction solution was 1 hour.
[0037]
The vapor temperature in the upper part of the heat exchange part A was 200 ° C, and the temperature of the raw material liquid in the line L2 was 180 ° C.
Hot water of 300 kPa and 100 ° C. was supplied to the heat exchange part C, and water vapor was generated by the heat of the vapor. This steam is effectively used as a heat source in an actual plant. The vapor temperature at the outlet L4 of the heat exchange part C was 150 ° C.
[0038]
80.3 parts by weight / hr of phenol remained in the reaction liquid extracted from L3, and 0.5 part by weight / hr of phenol was also contained in the vapor extracted from L4. The conversion of phenol was 14.2%.
From the amount of heat supplied to the reactor and the amount of heat recovered in the heat exchange section C, the amount of heat required for this operation was determined, and the amount of heat required to react 1 mol of phenol was calculated to be 1.35 * 10 ⁵ cal / mol. Met.
[0039]
[Comparative Example 1]
A reaction distillation type reaction apparatus having a structure in which a 0.55 m diameter shelf distillation unit was installed on the top of an evaporator having the same size as the autoclave used in Example 1 was used. Eight sheave trays having a liquid depth of 3 cm on each shelf were installed on the shelf steps at a step interval of 30 cm. On the second tray of this distillation column, 94 parts by weight of phenol at 60 ° C./hr and 2.5 parts by weight of dibutyltin oxide / hr were supplied. In addition, 90 parts by weight of dimethyl carbonate / hr at 60 ° C. was evaporated by a preheater and supplied on the seventh tray in the form of vapor. When continuously operated for 8 hours at an operating pressure of 300 kPa and a reflux ratio of 0.01, the tower bottom temperature was 226 ° C., the tower top temperature was 185 ° C., and the condenser condensation temperature was 118 ° C.
In this operation, as in Example 1, the calorific value required for reacting 1 mol of phenol was 1.69 * 10 ⁵ cal / mol.
In addition, in the reactive distillation method, it is theoretically possible to further increase the energy efficiency, but in order to do so, it is necessary to increase the liquid residence time of each stage. Increasing the liquid depth to increase the liquid residence time increases the pressure loss. According to the estimation of the present inventor, in order to obtain the same efficiency as in Example 1 with a tower diameter of 0.55 m, a liquid depth of 14 cm per stage is required. The liquid depth on the sieve tray of the high-pressure distillation column is usually about 2.5 to 10 cm, and operation at a liquid depth of 14 cm is difficult.
In addition, if the tower diameter is increased to increase the liquid residence time, vapor drift tends to occur, and the construction cost also increases. According to the estimation of the present inventor, a tower diameter of 1.2 m is required to obtain the same efficiency as the embodiment at a liquid depth of 3 cm.
[0040]
【The invention's effect】
According to the method of the present invention, by recovering a part of the energy supplied to the reactor in the indirect heat exchange section directly connected to the reactor, and improving energy efficiency, without going through complicated steps, An effect is obtained that aromatic carbonates can be produced efficiently and continuously with high selectivity.
[Brief description of the drawings]
FIG. 1 is a schematic cross-sectional view showing an example of a reaction apparatus used in the method of the present invention.
FIG. 2 is a schematic view showing an example of the structure of a liquid collecting part of a reaction apparatus used in the method of the present invention.
FIG. 3 is a schematic view showing another example of the structure of the liquid collecting part of the reaction apparatus used in the method of the present invention.
[Explanation of symbols]
A: 1st heat recovery part (heat exchange part)
B: Liquid collection part C: Second heat recovery part (heat exchange part)
D: Reactor L1: Raw material liquid supply line L3: Product extraction line L5: Low temperature fluid supply line

Claims (6)

In a method for producing an alkylaryl carbonate and / or diaryl carbonate by transesterifying a dialkyl carbonate and an aromatic hydroxy compound in the presence of a catalyst in the presence of a catalyst, the raw material liquid supplied to the reactor and the reactor Using a reactor having a structure in which a first heat recovery section configured to indirectly exchange heat with the generated vapor is directly connected to a gas phase section of a reactor, the first heat recovery section A process for producing an aromatic carbonate, characterized by causing a transesterification reaction to be carried out in a reactor while heating a supplied raw material liquid with the heat of condensation of the vapor. The method for producing aromatic carbonates according to claim 1, wherein the first heat recovery section has a structure of a coil heat exchanger, a multi-tube heat exchanger, or a spiral heat exchanger. A second heat recovery unit configured to indirectly exchange heat between the fluid supplied from the outside and the vapor that has passed through the first heat recovery unit is further installed on the upper portion of the first heat recovery unit. The method for producing aromatic carbonates according to claim 1 or 2, wherein a reaction apparatus having a structure is used, and heat recovery is performed by heating the fluid with the vapor. The method for producing aromatic carbonates according to claim 3, wherein the second heat recovery unit has a structure of a coil heat exchanger, a multi-tube heat exchanger, or a spiral heat exchanger. The method for producing aromatic carbonates according to claim 3 or 4, wherein a fluid other than the raw material liquid is used as a fluid supplied from the outside to the second heat recovery unit. Using a reaction apparatus having a structure in which a liquid collecting part for collecting the condensate in the second heat recovery part is installed below the second heat recovery part, the condensate collected in the liquid collection part is The manufacturing method of aromatic carbonates in any one of Claims 3-5 which returns to a reactor, without making it contact with the vapor which passes a heat-recovery part.

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