CN1275974A - Continuous process for the production of carboxylic acid esters of alkylene glycol monoalkyl ethers - Google Patents

Abstract

The present invention provides an improved method for the preparation of carboxylic acid esters of alkylene glycol monoalkyl ethers by the acid catalyzed esterification of the monoalkyl ether with a carboxylic acid. In a preferred embodiment of the invention, the carboxylic acid and alcohol are reacted in a reactor/column and the resulting ester product is distilled into an overhead decanter/extractor as a single phase. A small amount of solvent, preferably a hydrocarbon is aded to the mixture causing the resulting distillate to separate into two phases, one phase containing the desired product, the other containing primarily water. The process described is applicable to both batch and containuous operation and is not constrained by the difficulty of separating closely boiling azeotropes and results in substantially higher production rates than achieved by current processes.

Description

Production method of alkylene glycol monoalkyl ether carboxylate

field of invention

The present invention relates to a process for the production of carboxylic acid esters, in particular alkoxyalkyl esters.

Background of the invention

The present invention relates to a process for the production of carboxylic acid esters, in particular to a process for the continuous production of alkoxyalkyl esters of alkylene glycol monoalkyl ethers. In addition to unreacted reactants, water evolved from the reaction can cause handling and purification problems.

In a process for the continuous production of esters by reacting alcohols with carboxylic acids, removal of water produced during the reaction can increase conversion. Typically, the reaction process is carried out in a reactor comprising a mixture of alcohol, carboxylic acid, ester, water and acid catalyst. The reactor is heated to obtain an equilibrium mixture and the product is distilled in a fractionation column. Alcohol and carboxylic acid are added to the reactor as the product is distilled. For simple esters such as ethyl acetate and butyl acetate, water is removed as an azeotrope along with the ester and unconverted alcohol. The distillate can separate into two liquid phases. The upper layer, also known as the "oil phase", contains mainly esters with small amounts of alcohol and some water. The lower layer is called the "aqueous phase" and contains mainly water with some esters and alcohols. The aqueous phase is transferred to a distillation column, the water is withdrawn as waste from the bottom of the column, and the distillate is recycled. The oil phase is distilled in a purification column to produce a base discharge product of pure ester and a distillate recycled to the reaction product. The process has been optimized in recent years to allow the production of these simple esters in high yields.

When this method was attempted to esterify alkylene glycol monoalkyl ethers such as 1-methoxy-2-propanol, it was found to be completely or hardly feasible. The distillate from the reaction column is not easily separated into two phases, which makes it difficult to remove the water of reaction by the above-mentioned method. The reason that phase separation cannot occur is that these alcohols and esters are more stable in water than simple esters. It has also been found that the reaction column/distillation column can be operated in such a way that separation into two closely boiling azeotropes (one rich in alkylene glycol monoalkyl ether and the other rich in corresponding ester). While this can be achieved by operating the distillation column at a high reflux ratio, operating the distillation column in this manner will greatly reduce its capacity. This process produces a distillate which can be separated into oil and water phases, but the degree of separation is difficult. Furthermore, the reaction column/distillation column must be operated at very low speeds, making the overall process uneconomical.

Due to the above-mentioned solubility problems, alkylene glycol monoalkyl ether esters are usually produced by the method described in EP0119833B1. A compound such as toluene is added to the reactor and water and toluene are distilled off as an azeotrope by a distillation process. This brings the reaction towards completion. The azeotrope is separated into two phases; water is removed as the aqueous phase and the oil phase is recycled to the reactor. In this process, only water acts as the azeotrope, leaving the ester, unreacted alcohol or carboxylic acid and catalyst in the reactor. This method requires removal of the catalyst from the product prior to purification by neutralization or other means. Another drawback is that these processes typically operate in a batch rather than continuous manner, resulting in low feedstock efficiency and catalyst loss. In addition, esters prepared in this way also suffer from acidity and stability problems.

Summary of the invention

The object of the present invention is to overcome the aforementioned difficulties. It has been found that the production of alkylene glycol monoalkyl ether esters such as 1-methoxy-2-propyl acetate can be carried out in a continuous process with high yields, high speeds and excellent product quality and catalyst stability. loss etc. The object of the present invention is achieved like this: adopt water as entrainer, with the azeotrope of water, carboxylic acid ester and some unreacted glycol ether alcohols, product is distilled into tower distillate and decanted in the analyzer/extractor. A small amount of inert solvent is added to the decanter/extractor to separate the distillate into oil and water phases. The oil phase mainly contains solvent, ester and a small amount of water and unreacted alcohol. The aqueous phase mainly contains water and unreacted alcohol and some esters. Unreacted carboxylic acid and catalyst remain in the reactor. The process does not have to deal with the separation difficulties of closely boiling azeotropes or high boiling ester products and operates at substantially higher productivity. Furthermore, the carboxylic acid does not distill into the overhead and contaminate the product.

Brief description of the drawings

Figure 1 illustrates a preferred embodiment of the invention and illustrates the esterification process using a reactor/reboiler, distillation column and decanter/separator.

Detailed description of the invention

The present invention provides a process for producing alkoxyalkyl esters by reacting carboxylic acids with alcohols. In a preferred embodiment of the invention, the carboxylic acid is reacted with the alcohol in a reactor/reboiler and the resulting ester product is distilled azeotropically with water as a single phase into the overhead decanter/extractor . A small amount of extraction solvent is added to the mixture so that the distillate formed separates into two phases, one comprising the desired product and the other comprising mainly water. The process of the present invention does not have to deal with the separation problems of closely boiling azeotropes or high boiling ester products, thus achieving higher productivity than conventional methods.

According to the present invention, a kind of production method of the carboxylate of alkylene glycol monoalkyl ether is provided, it comprises:

其中，n＝0-6；R₁，R₂＝H，CH₃--(CH₂)_n-，X＝Cl，Br，FX基团可在链上的任一位置；a) reacting a monocarboxylic acid or halogenated monocarboxylic acid having from 1 to about 10 carbon atoms with an alkylene glycol monoalkyl ether having the formula in the presence of an acid catalyst:

Among them, n=0-6; R ₁ , R ₂ =H, CH ₃ --(CH ₂ ) _n -, X=Cl, Br, FX group can be at any position on the chain;

b) distilling the mixture in a distillation column while using water to form an azeotrope with carboxylic acid ester and unreacted alkylene glycol monoalkyl ether;

c) directing the distillate from (b) to an overhead extractor and contacting with an effective amount of an inert solvent (also known as a phase separation agent) to form at least two liquid phases;

d) separating the two phases (aqueous and oil (product) phases) of the mixture formed;

e) distilling the oil phase to recover the (substantially pure) monocarboxylate product and inert solvent (for recycling); and

f) Distillation of the aqueous phase (to recover water for waste treatment and alcohols and esters for recycling).

The above method can be used in a continuous reaction device or a batch reaction device. Preference is given to continuous reaction plants comprising reaction columns, distillation columns and overhead decanters/extractors.

Examples of C _1-10 acids include, but are not limited to, acetic acid, formic acid, propionic acid, isobutyric acid, and n-butyric acid. Examples of glycol esters of the process product of the present invention include, but are not limited to: 1-ethoxy-2-ethyl acetate, 1-methoxy-2-propyl acetate, and 1-methoxy-2-propionate ester. Examples of useful ethers are 1-ethoxy-2-ethanol and 1-methoxy-2-propanol and the like.

The reaction uses an acid as a catalyst, such as an inorganic acid, such as concentrated sulfuric acid, hydrochloric acid, nitric acid, and the like. Lewis acids such as boron trifluoride, antimony pentafluoride and the like can also be used. Organic sulfonic acids and halogenated sulfonic acids can also be used, such as methanesulfonic acid, ethanesulfonic acid and butanesulfonic acid, trifluoromethanesulfonic acid, trichloromethanesulfonic acid, o- or p-toluenesulfonic acid, benzenesulfonic acid, etc. , and strongly acidic sulfonated aromatic ion exchange resins and perfluoroalkane sulfonic acid resins. The acid catalyst is generally used in an amount of about 0.01 to about 10 weight percent, preferably about 0.1 to about 2.0 weight percent, based on the total weight of the reaction mixture, and the amount can vary depending on the particular acid employed.

The difference between the currently used azeotropic method and the method using a phase separation agent (extractant) in the present invention is explained below. For example, when a hydrocarbon is used as an entrainer, it may be added to or present in the reactor/reboiler. The stable boiling point mixture is distilled through a distillation column to produce a distillate comprising hydrocarbons and other components, in this case primarily water. The desired product, unreacted alcohol and/or carboxylic acid and catalyst remain in the reactor/reboiler. When a hydrocarbon is used as the phase separation agent (extractant) in the process of the present invention, it can be fed to the overhead decanter/separator to separate the product into two phases. By operating in this way, no carboxylic acid is distilled off, which greatly simplifies the purification of the product.

Useful phase separating agents include those which are inert, chemically compatible with the reacting components and which separate the desired product into two phases. Solvents can be linear, branched, aromatic or cyclic hydrocarbons, esters, ethers, ketones or chlorofluoro compounds. Typically, these compounds have about 5 to 12 carbon atoms. Examples of suitable solvents include, but are not limited to: pentane, cyclopentane, hexane, cyclohexane, toluene, benzene, xylene, olefins, butyl acetate, propyl acetate, ethyl acetate, methyl t-butyl Ether, diisopropyl ether, methyl ethyl ketone, methyl propyl ketone, methyl butyl ketone, and the corresponding compounds, chlorofluoro compounds, chloroform, carbon tetrachloride, methylene chloride, and ^Freons® . Preferred phase separation agents are C ₅ -C ₁₂ hydrocarbons, especially when operating under atmospheric conditions. Hydrocarbons larger than _C12 are generally not preferred. If the hydrocarbon has too high a boiling point, the hydrocarbon will remain in the reactor and not enter the overhead receiver/decanter of the distillation column. If the hydrocarbons have too low a boiling point, it is not feasible to use such hydrocarbons under atmospheric conditions. Olefins can be used but are not preferred solvents due to their polymerization in the reactor.

The phase separation agent should be used in an effective amount to form two liquid phases over the decanter/extractor operating temperature range. Suitable amounts are from about 5 to about 70 wt%, preferably from 10 to about 50 wt%, most preferably from about 20 to about 40 wt%. Too little phase separation agent will not allow phase separation, while too much will require excessive equipment size and excessive energy consumption.

The acid catalyzed esterification reaction may be carried out in any suitable reactor having means for mixing the reactants, means for regulating the temperature of the reactor and for separating the desired ester product from unreacted components and Apparatus for water produced during the reaction. In a preferred embodiment, a distillation column, a condenser and a phase separator or decanter/extractor for removing the solvent (containing the product)-water phase, as well as for returning the solvent and Water to the device in the distillation column.

The general process of the reaction is to add glycol ether, carboxylic acid and acid catalyst to the reactor or reaction tower. The mixture is heated and maintained at the desired reaction temperature for a suitable period of time, then the distilled product mixture is transferred to an overhead phase separator. In the phase separator, the mixture is brought into contact with a phase separation solvent, whereby phase separation takes place. A product separation process is then carried out which involves separating the phases of the mixture formed (aqueous and oil (product-containing) phases) and distilling the two phases to recover the desired monocarboxylic acid ester.

Referring to Figure 1 of the accompanying drawings, this figure illustrates a reactor/reboiler in which the reactants are brought into contact and intimately mixed using standard reaction engineering methods. If operating in a continuous mode, the feed rate of the reactants is adjusted to maintain an appropriate residence time at the reaction temperature. The mixture is fed to the bottom of a distillation column where the ester product, water and unreacted reactants are distilled. The distilled product stream comprising the desired ester product is fed to a decanter/separator where a phase separation agent is added to the mixture to separate the mixture into at least two liquid phases. The oil phase or product phase is separated from the water phase. Then, the product phase is distilled to make it more pure. The aqueous phase, which contains most of the water, some esters and some unreacted alcohol, is sent to the organics recovery column.

Typical reaction conditions for the esterification reaction of the present invention include: a temperature in the reactor/column in the range of 80 to about 160°C, a pressure of about 0.1 to 10 atmospheres, and a reactor residence time of about 0.3 to about 5 hours. Three parameters can be adjusted to optimize the reaction process, which will vary for the production of different esters. From an economic point of view, the preferred conditions are operating at approximately 1 atmosphere pressure with a reactor residence time of about 0.5 to 2 hours.

Although the process of the present invention relates to the production of alkylene glycol monoalkyl ether esters, the process can also be used to carry out esterification processes in general. Those skilled in the art will recognize that, employing suitable glycol ethers and monocarboxylic acids, the process of the present invention is broadly applicable to the production of other esters such as propylene glycol monobutyl acetate, dipropylene glycol monooctyl butyrate, formic acid Ethylene glycol monoethyl ester, etc.

The following examples are used to illustrate the present invention, but they are not to limit the protection scope of the present invention. Example Example 1

The reaction apparatus was assembled from the following: 30 trays, 2" diameter Oldershaw distillation column, reflux cooler, overhead receiver (decanter) and reboiler/reactor. Pumps were used to Send fresh feed to reboiler, and send cyclohexane to overhead receiver.Add 62.8g of 1-methoxy-2-propyl acetate, 66.0g of 1 - Methoxy-2-propyl alcohol, 132.1 g of glacial acetic acid, 67.1 g of water and 17.9 g of methanesulfonic acid catalyst. The distillation column operates at atmospheric pressure with a reflux ratio of 1.0. The composition of the fresh feed is: 44.0 The 1-methoxy-2-propanol of wt%, the glacial acetic acid of 14.0wt% and the water of 42.0wt%, this fresh material is added in the reboiler, and feed rate is 5.28g/min. Hexane is added in the distillate receiver of tower, and feed rate is 1.07g/min.The temperature of reboiler is kept on 112 ℃ during operation, and the temperature at the tray top place of distillation column is 94 ℃; Almost no cyclohexane was present in the distillation column or reactor/reboiler. Including the cyclohexane feed, the overall production rate from the overhead decanter was 6.36 g/min. Condensed distillate Separates immediately into an oily phase mainly comprising cyclohexane and 1-methoxy-2-propyl acetate, and mainly water with some 1-methoxy-2-propyl alcohol and 1-methoxy-2-propyl acetate Aqueous phase of the propyl ester. Only the aqueous phase was refluxed into the distillation column. These operating conditions were maintained for 5 hours.

The composition of the distillate before addition of cyclohexane was measured as: 33.1 wt% 1-methoxy-2-propyl acetate, 21.2 wt% 1-methoxy-2-propanol and 45.7 wt% of water. From this measurement it was determined that under these reaction conditions the mixture would not be able to separate into two phases. After cyclohexane was added, the phases were separated and the oil phase contained 44.1 wt% cyclohexane, 46.6 wt% 1-methoxy-2-propyl acetate, 7.7 wt% 1-methoxy-2- Propyl alcohol and 1.6 wt% water. Phase formation was also observed with about 10 wt% cyclohexane. Example 2

After the procedure of Example 1, the distillation column was operated at a reflux ratio of 0.68 and the composition of the fresh feed was: 47.7 wt% of 1-methoxy-2-propyl alcohol, 18.2 wt% of glacial acetic acid and 34.0 wt% This fresh material was fed to the reboiler at a feed rate of 5.85 g/min. Cyclohexane was fed to the overhead receiver at a rate of 2.01 g/min. The reboiler temperature was maintained at about 115°C during operation and the temperature at the top tray of the distillation column was about 93°C. Including the cyclohexane feed, the overall production rate from the overhead decanter was 7.78 g/min. As in the previous examples, separation into oil and water phases took place. These operating conditions were maintained for 4 hours.

The composition of the distillate before adding cyclohexane to the overhead distillate receiver is: 32.5wt% of 1-methoxyl-2-propyl acetate, 19.3wt% of 1-methoxyl-2 - Propyl alcohol and 48.2% by weight of water. From this experiment it was determined that the mixture would not separate into two phases.

After addition of cyclohexane and phase separation, the composition of the product phase is: 50.8 wt% cyclohexane, 43.5 wt% 1-methoxy-2-propyl acetate, 5.7 wt% 1-methoxy - 2-Propanol and 0.0% by weight of water. comparative example

Numerous experiments without the addition of cyclohexane have shown that no phase separation is possible in the overhead receiver unless the reflux ratio in the column is at least 3.0, preferably greater than 5.0 when a 30 plate column is used. However, under these operating conditions, the feed to the reboiler may only be maintained at a rate of about 1.0 g/min. In practice, even at this rate, when recycling unreacted 1-methoxy-2-propyl ether is impossible, consideration must be given to operating a dehydration column.

These examples show that the desired ester product can be withdrawn at a rate five times higher than when no phase former is used. The results also showed that the distillable product obtained contained the desired ester and that it contained no carboxylic acid reactants. The results also show that neutralization of the acid catalyst is not required in practice to recover pure product.

Claims (8)

1, a kind of production method of the carboxylate of alkylene glycol monoalkyl ether, comprising: a) In a reaction tower, in the presence of an acid catalyst, a monocarboxylic acid or a halogenated monocarboxylic acid having 1 to 10 carbon atoms is reacted with an alkylene glycol monoalkyl ether having the formula:

Wherein, n=0-6; R ₁ , R ₂ =H, CH ₃ -(CH ₂ ) _n -, X=Cl, Br, F b) the mixture is distilled in a distillation tower while utilizing water of reaction to form an azeotrope with carboxylic acid ester and unreacted alkylene glycol monoalkyl ether; c) directing the distillate from (b) to an overhead extractor/separator and contacting with an effective amount of an inert solvent to form at least two phases; d) separating the two phases of the resulting mixture to form an aqueous phase and an oily phase (product phase); e) Distilling the oil phase to recover the monocarboxylate product and inert solvent. 2. The method according to claim 1, wherein the acid catalyst is selected from the group consisting of: concentrated sulfuric acid, hydrochloric acid, nitric acid, boron trifluoride, antimony pentafluoride, sulfonic acids such as methanesulfonic acid, ethanesulfonic acid and butanesulfonic acid, trifluoro Methanesulfonic acid, trichloromethanesulfonic acid, o- or p-toluenesulfonic acid, benzenesulfonic acid, and sulfonated aromatic ion exchange resins and perfluoroalkanesulfonic acid resins. 3. The method according to claim 1, wherein the ether is 1-methoxy-2-propanol, 1-ethoxy-2-ethanol. 4. The method according to claim 1, wherein the inert solvent is linear, branched, aromatic, cycloaromatic hydrocarbons, esters, ethers, ketones and chlorofluoro compounds. 5. The method according to claim 4, wherein the solvent is selected from C ₅ -C ₁₂ , pentane, cyclopentane, hexane, cyclohexane, toluene, benzene, xylene; butyl acetate, propyl acetate, acetic acid Ethyl esters, MTBE, diisopropyl ether, MEK, MPK, MBK, and corresponding branched compounds, dichloromethane, carbon tetrachloride. 6. The method according to claim 5, wherein the solvent is used in an amount of about 5-70 wt%. 7. The method according to claim 6, wherein the solvent is used in an amount of about 10-50 wt%. 8. The method according to claim 7, wherein the solvent is used in an amount of about 20-40 wt%.