MXPA98009496A - A process to prepare (met) alkyl acrylics - Google Patents

Abstract

A process for preparing alkyl acrylates is shown. The process provides the synthesis of alkyl (meth) acrylates, the hydrolysis of impurities from the process into starting material, and the separation of the starting material and the reaction products into a reactive

Description

A Process To Prepare (Me t) Ac r i 1 a t es de Alquilo The present invention relates to a process for preparing alkyl (meth) acrylates. Specifically, the process provides the synthesis of the alkyl (meth) acrylates, the hydrolysis of the impurities from the process into starting material, and the separation of the reaction products and the starting material in a reactor. Alkyl (meth) acrylates are important monomers in commercial polymerization processes. Conventionally, alkyl (meth) acrylates, such as butyl acrylate ("AB"), are prepared commercially by means of a direct esterification process. Typically, butanol ("BuOH") and acrylic acid ("A?") Are reacted in the presence of an acid catalyst, thereby producing butyl acrylate and water. Direct esterification is usually carried out at elevated temperature and reduced pressure. During the reaction, impurities such as dibutyl ether ("DBE"), butyl-β-butoxy propionate ("PBBB"), butyl-β-hydroxy propionate ("PBBH"), butyl-acryloxypropionate ("BAOPA") are formed. ") and acryloxypropionic acid (" AAOP "). These impurities, if they are not converted back into starting material, result in poor performance. Such impurities are usually removed from the reactor, and treated to produce starting material that can be reused. As a result, these processes are less efficient and require additional capital investment costs for separate reactors. In addition, the conventional processes for preparing AB operate at reduced pressure, necessitating larger equipment. Consequently, there is a need for a more efficient and lower cost butyl acrylate process that converts the impurities from the process back into starting material, doing the above in the same reactor where the AB is produced, and that it can be operated at atmospheric pressure. The pending US patent application no. 08 / 797,380 shows a process for preparing AB, in which the process impurities are hydrolyzed in separate reactors. In addition, the separation of AA / AB is carried out in a distillation column, low pressure, separately. Although the patent application is directed to the conversion of impurities into starting material, it does not provide an AB process in which the AB • is prepared in a high reaction medium with water that allows (1) operation under atmospheric pressure, (2) the separation of AA / AB in the reactor, and (3) the recovery of the starting material of the process impurities, in one unit. The present invention shows a process for preparing alkyl (meth) acrylates, which converts the impurities from the process into starting material and furthermore makes it react in a reactor. The addition of water during the direct esterification reaction also provides a process for preparing alkyl (meth) acrylate which does not require reduced pressure, and facilitates recovery of the starting material from the impurities of the process. In addition, the separation of the reaction product, such as AB, and the starting (meth) acrylic acid, such as AA, can also be carried out in the reactor. Accordingly, the present invention provides a process that is more efficient and economical than conventional processes for preparing alkyl (meth) acrylate known in the art. One aspect of the present invention provides a process that includes: (A) loading a reactor with C1-C4 alcohol, a (meth) acrylic acid, a strong acid catalyst, and at least 5% by weight of water to form a reaction mixture; (B) reacting the reaction mixture to form the alkyl (meth) acrylate C ^ -Oj and the process impurities, wherein the process impurities formed are hydrolyzed in the reactor; and (C) separating the alkyl (meth) acrylate C ^ d and the water formed during the reaction from the reaction mixture. Another aspect of the present invention provides a process that includes: (A) loading a reactor with butanol, acrylic acid, a strong acid catalyst, and at least 5% by weight of water to form a reaction mixture; (B) reacting the reaction mixture to form the butyl acrylate and process impurities, wherein the impurities from the process are hydrolyzed in the reactor; and (C) separating the butyl acrylate and the water formed during the reaction from the reaction mixture. Another aspect of the present invention provides a process that includes: (A) loading a reactor with butanol, acrylic acid, 3.5 to 15% by weight of sulfuric acid, 6 to 18% by weight of water and at least one inhibitor for forming a reaction mixture, wherein the butanol and acrylic acid are charged to the reactor at a molar ratio of acrylic acid to butanol from 1: 1 to 1: 1.7; (B) reacting the reaction mixture to form the butyl acrylate and process impurities, wherein the impurities from the process are hydrolyzed in the reactor; and (C) separating the butyl acrylate and water formed during the reaction from the reaction mixture by azeotropic distillation. As used herein, the term "(meth) acrylic acid" includes both acrylic acid and methacrylic acid. Similarly, with the term "(meth) acrylate" includes both acrylate and methacrylate. As used herein, BuOH refers to n-butanol, i.e. 1-butanol, and the term "butanol" includes within its scope all butanol isomers, as well as mixtures thereof.

The term "alkyl" includes cyclic, branched-chain or straight-chain alkyl groups. As used herein, the terminology n (C? -CA) "or" (C? -C10) "means a group having 1 to 4 or 1 to 10 carbon atoms per group, as used herein, with the terms "AA-rich" or "AB-rich" are understood to be the fractions or components in which AA or AB is the largest organic component (greater than 50% by weight) of the composition Through the description and the claims, Unless otherwise indicated, references to percentages are by weight, all temperatures are in degrees centigrade, and all pressures are atmospheric, Figure 1 illustrates the equipment and flow lines used in one mode of the process. the present invention, including the direct esterification / hydrolysis reactor (1), which is a mixing reactor having a distillation column in its upper part, the line (2), which bears a vaporized distillate mixture, which includes AB , from (1) to a phase separator (3), the phase separator (3) knows the vaporized distillate in an organic phase rich in AB and an aqueous distillation phase; line (11), which carries the AB-rich organic distillate separated in (3) to a separation section; line (8), which carries the aqueous distillate separated in (3) to line (9) to be recycled in (1), and to line (10) to be carried to where it will be treated, generally to recover the material from the water waste; line (4), which carries AA-rich funds from (1) to the depletion particle (5), which is the fractionation reactor; the line (6), which carries the distillate, including the BuOH and AA recovered from (5) to be recycled in (1) through the line (22), and to the line (7) carrying the distillate from (5) ) forward to be treated, usually as waste; line 12, which carries the funds from (5) forward to be treated, usually as waste and, optionally, to line (17), which recycles the funds from (5) to (1); the line (13), which can feed the inhibitor to the reactor; the line (14), which feeds the catalyst to the reactor; line (15), which feeds the pure AA and BuOh into the reactor; an optional plug flow reactor (16); an optional line (18) to feed the AA, the BuOH and the catalyst towards (16), - an optional line (19) to carry the material from (16) to (1); the line (20), which carries the BuOH, the AB and the AA recovered in the separation section from the lines (10) and (11) back to the reactor (1); and the optional line (21) that returns the recovered material to an alternative feeding location that is located in the reactor (1). As noted above, in step (A) of the present invention, the alcohol C ^ -C ^ a (meth) acrylic acid, a strong acid catalyst and water are charged to the reactor to form a reaction mixture. Generally, the C1-C4 alcohol is a straight or branched chain alkanol having from 1 to 4 carbon atoms or a mixture thereof. Specific examples include, but are not limited to, methanol, ethanol, n-propanol, isopropanol, n-butanol, 2-butanol, tert-butanol or mixtures thereof. Furthermore, it is contemplated that the alcohol d-d may be substituted, for example, with halogen, hydroxide, alkoxide, cyano, nitro, etc. In one embodiment, the alcohol is butanol. In a preferred embodiment, the alcohol is n-butanol. Also present is (meth) acrylic acid or substituted (meth) acrylic acid with, for example, halogen, hydroxide, alkoxide, cyano, nitro, etc. In one embodiment, acrylic acid is present . or methacrylic acid or a mixture thereof. In a preferred embodiment, the unsaturated acid is acrylic acid. The (meth) acrylic acid and alcohol are present at a molar ratio of 1: 1 to 1: 1.7, preferably 1: 1.1 to 1: 1.6, more preferably 1: 1.25 to 1: 1.45. It is also contemplated that other unsaturated acids, such as crotonic acid, cinnamic acid, maleic acid, fumaric acid, etc., which may participate in a transesterification reaction with an alcohol, may be used in the process of the present invention. A strong acid catalyst is also present in the reaction mixture. Suitable examples of said acid catalyst include, but are not limited to, sulfuric acid, methane sulphonic acid, benzene sulfonic acid, p-toluene sulfonic acid, mixtures thereof, or alkyl sulfonic acids supported with polymer, such as AMBERLYST ™ 15 resin or NAFION-H ™ resin. Generally, the alkyl sulfonic acid is an alkyl sulfonic acid d-do- In one embodiment, the strong acid catalyst is a sulfur-containing acid or acid supported with sulfur-containing polymer. In a preferred embodiment, the strong acid catalyst is sulfuric acid. The concentration of the strong acid catalyst by total weight of the reaction mixture in the direct esterification / hydrolysis reactor is typically 3.5 to 15% by weight, preferably 3 to 12% by weight, and more preferably 5 to 12% by weight. 8% in weight. Water is also present in the reaction mixture provided in step (A). Generally, any type of water, such as tap water, distilled water or deionized water, that is suitable for use in a direct transesterification reaction can be used. In addition, at least part of the water provided may be recycled reaction water, which has been removed during the separation of the reaction product from the starting materials. The addition of water provides a reaction medium of water in the reactor, which allows operation under atmospheric conditions, and the hydrolysis of the reaction by-products to recover the starting materials, as well as the separation of the reaction products from of starting materials in a reactor. Also, at least one inhibitor can be charged to the reactor in step (A). Typically, from 0.001 to 1.0%, preferably 0.001 to 0.5%, and more preferably from 0.001 to 0.1% by total weight of the reaction mixture of at least one inhibitor, if used, is present during the process Direct esterification to prevent polymerization. Suitable inhibitors include hydroquinone, the mono-methyl ether of hydroquinone, 2,2,6,6-tetramethyl-1-piperidinyloxy, hydroxy butylated anisol, naphthoquinone, "anthranil, and combinations thereof. Such derivatives include, but are not limited to, piperidinyl free radical of 4-methacryloyloxy-2,2,6,6-tetramethyl and 4-hydroxy-2, 2,6,6-tetramethyl-N-hydroxy -piperidine In a preferred embodiment, the at least one inhibitor is 2,2,6,6-tetramethyl-1-piperidinyloxy In another embodiment, 2, 2, 6,6-tetramethyl-1-piperidinyloxy and another inhibitor are used such as the methyl ether of hydroquinone.

In one embodiment, the alcohol is butanol, the (meth) acrylic acid is acrylic acid, the acid catalyst is sulfuric acid, and the inhibitor is 2, 2,6,6-tetramethyl-1-piperidinyloxy. In another embodiment, the water is present from 6 to 18% by total weight of the reaction mixture, and the strong acid catalyst is present from 3.5 to 15% by total weight of the reaction mixture. As indicated above, the present invention utilizes a reactor wherein direct esterification of AA and BuOH, hydrolysis of reaction byproducts including PBBB, PBBH and BAOPA, and separation of AB and AA is achieved. In general, any suitable or adaptable reactor can be used for a process in which a direct esterification reaction occurs, the hydrolysis of the reaction by-products formed during the transesterification reaction, and the separation of the reaction products from the reaction materials. start in the reactor. In one embodiment, the reactor may be a mixer tank equipped with a distillation column. In a preferred embodiment, the distillation column can be placed directly on top of the reactor (as in Figure 1), and can be a fractional distillation column. In general, the distillation column contains 20 to 100 trays. It is preferred that the column contains from 20 to 70 trays. It is even more preferred that the column contain 40 to 50 trays. The column has means for feeding AA, BuOH, a strong acid catalyst, water, and at least one inhibitor. The reactor also has means to remove the funds. In step (B), the reaction mixture is reacted to form alkyl (meth) acrylate - and reaction byproducts, while the reaction by-products formed during the reaction are hydrolyzed in the same reactor. The direct esterification reaction can be carried out by feeding the AA and BuOH through the lines (15), (20) and (22) to the direct esterification / hydrolysis reactor (1), at a molar ratio of AA to BuOH ranging from 1: 1 to 1: 1.7, preferably from 1: 1.25 to 1: 1.45. The AA and BuOH can also be fed together with sulfuric acid to a plug flow reactor (16), and then to the reactor (1) of the direct esterification / hydrolysis. If used, the inhibitor is fed into the reactor using line (13). The AA, BuOH, inhibitor, strong acid catalyst and water form a reaction mixture in the direct esterification / hydrolysis reactor. The AA and BuOH are reacted at an AA conversion of 50 to 95%, preferably 60 to 95%, more preferably 70 to 95%.

During the direct esterification reaction, the reactor must have at least 5% by total weight of the water reaction mixture, for the efficient operation of the hydrolysis. Preferably, the reactor has from 6 to 18% by weight of water during the direct esterification reaction. More preferably, the reactor has from 8 to 12% by weight of water during the direct esterification reaction. The water content can be maintained by returning the condensed aqueous distillate to the reactor and separating it from the reactor through line (9). Water can also be added through any of the power lines, as necessary. The water in the reactor hydrolyzes the reaction by-products formed during the reaction. Specific examples of the hydrolysis reactions that occur include, but are not limited to, reactions where the PBBB is hydrolyzed to two BuOH and one AA, and BAOPA is hydrolyzed to one BuOH and two AA. The direct esterification reaction and the hydrolysis are carried out at a temperature of 100 ° C. at 140 ° C. , preferably 105 ° C. at 135 ° C, and more preferably at 115 ° C. at 130 ° C. The direct esterification reaction and hydrolysis are carried out at pressures of 100 mm. from Hg to 760 mm. of Hg. The atmospheric pressure is preferred. The residence time in the direct esterification / hydrolysis reactor is typically 0.5 to 5 hours, preferably 1 to 4 hours, and more preferably 2 to 3 hours. In step (C), the Cx-C4 alkyl (meth) acrylate and water formed during the reaction of the alcohol with the acid (meth) acrylic are separated from the reaction mixture by means of methods known in the art such as distillation, phase separation, etc. In a preferred embodiment, the alkyl (meth) acrylate d ~ C4 and the water formed during The reaction is separated from the reaction mixture by azeotropic distillation. In a more preferred embodiment, the d-C4 alkyl (meth) acrylate is azeotropically distilled with water (aqueous reflux) and BuOH under the conditions described above. In this way, the water added to the reaction medium, as well as the water produced by the transesterification reaction of AA and BuOH, provide an aqueous medium that increases the separation of AA and AB in the reactor. The distillate can then be carried through line (2) to a phase separator (3). In the phase separator, an organic phase, which is rich in AB and contains BuOH, and an aqueous phase, which contains water and water and separate AA. The organic phase can be carried through the line (11) to a separation section, where the pure AB is obtained. The BuOH can be recovered from the separation section, and recycled. Part of the aqueous phase is carried through line (8) to line (9) to be recycled to the reactor and maintain the appropriate amount of water in the reactor. The rest of the aqueous phase is taken through line (8) to line (10) to carry it forward and recover and treat it, usually as waste. The bottoms of the direct esterification reactor contain strong acid catalyst, AB, AA, BuOH, AAOP, APBB and APBH. A depletion primer (5) can be used to fractionate the acryloxy propionic acid ("AAOP"), the AA regulator; beta-n-butoxy propionic acid ("APBB"), and beta-hydroxy propionic acid ("APBH"). Accordingly, the funds can be taken through line (4) to the depletion splitter (5) (the fractionation reactor), where the AAOP is split into two AAs; the APBB is fractionated into a BuOH and an AA; and the APBH is fractionated into an AA. The fractionation reactor may be a continuous mix tank reactor. When the fractionation reactor is incorporated in the process, the liquid in the fractionation reactor is maintained at 5 to 25% by weight of strong acid, preferably sulfuric acid. The fractionation reactor is operated at a temperature that varies from 90 to 140 ° C. , preferably from 110 to 140 ° C. The fractionation reactor is operated at a pressure ranging from 20 to 200 mm. of Hg, although higher pressures can be used, up to 800 mm. of Hg. The residence time in the fractionation reactor is typically 0.5 to 10 hours, preferably 0.5 to 6 hours, more preferably 0.5 to 3 hours. Part of the AB, AA, BuOH and water generated in the fractionation reactor can be taken from the aerial system and recycled to the direct esterification / hydrolysis reactor through line (6) to line (22). The rest of the AB, AA, BuOH and water generated in the fractionation reactor can be taken from the air system through line (6) to the line (7) to be carried forward and treated, usually as waste. The bottoms of the fractionation reactor can be taken through the line (13) to be carried forward and be treated, generally as waste, or they can be recycled to the reactor (1) of direct esterification / hydrolysis, through the line ( 17). The abbreviations used in this application are:% = percent ° C. = degrees Celsius AB = butyl acrylate mm. = millimeters ml. = milliliters ml./min. = milliliters per minute AA = -acrylic acid BuOH = butanol AAOP = acryloxypropionic acid APBB = beta-n-butoxy propionic acid APBH = beta-hydroxy propionic acid PBBB = butyl-β-butoxypropionate PBBH = butyl-β-hydroxypropionate BAOPA = butyl-acryloxypropionate cm. = centimeters Hg = mercury gr. / hour = grams per hour The following examples illustrate the process of the present invention.

Materials: AA, AB and BuOH were obtained from a series of plant production. The inhibitors used are commercially available. Analysis: Standard methods were used for the determination of water, monomer, BuOH and residual impurities. The levels of AAOP, PBBB, PBBH and BAOPA were determined by means of gas / liquid chromatography, using the detection of flame ionization. The determination of sulfuric acid was obtained using a pH probe and alcohol tetrabutylammonium hydroxide titrant.

Example 1 Preparation of butyl acrylate A direct esterification / hydrolysis reactor was assembled using a 500 ml round bottom flask. , connected to a steam reheater, covered, with multiple tubes, Hastalloy C-276. An Oldershaw fractional distillation column, 5.08 cm. , with 45 trays, was placed directly on top of the glass reactor, and was considered part of the reactor. An aerial system, of two water-cooled condensers, of standard glass, connected in series, was connected to the fractional distillation column. A 2,000 ml glass fraction cutter. It was connected to the second capacitor. An exhaustion splitter (fractionation reactor) was mounted using a flask from 500 ml. equipped with - an electric heating blanket, temperature controllers, a stirrer, and a water-cooled distillation head that had an outlet port leading to a 125 ml fraction cutter. A peristaltic pump was provided to pump the series of funds directly from the reheater. A graduated cylinder equipped with a Teflon stopcock to easily separate the sample was used to collect the series of backgrounds. The AA, BuOH and sulfuric acid were fed to the direct esterification reactor. The direct esterification reactor was set at 128 ° C. and 760 mm. of Hg. The depletion primer was set at 130 ° C and 35 mm. of Hg with 26% by weight of sulfuric acid present. Next, the inhibitor pump was turned on to pump a solution of 0.25% 2,2,6,6-tetramethyl-1-piperidinyloxy and 0.018% hydroquinone ethyl ether into AB. Once the trays at the top of the fractional distillation column were moistened with the solution of the inhibitor, the steam was introduced into the reboiler. When the superheater reached the desired temperature and the distillate was observed in the aerial system, the following were pumped into the reactor: 3.7 g / hour of inhibitor through the positive displacement piston pump FMI, 234.4 gr./hour Total AA (including recycled) through the positive displacement piston pump FMI, 382.5 gr./hour of the total BuOH (including recycled) through the positive displacement piston pump FMI, and 4.5 gr./hour of acid sulfuric through a peristaltic pump, using appropriate tubing and compatible with strong mineral acid. A part of the water recovered from the aerial system was fed back to the reactor at a reflux feed rate of 461.8 gr./hour. The residence time in the reactor was 180 minutes. The distillate and series of bottoms were collected every hour, and analyzed with strong acid, AA and water to determine when the equilibrium state had been achieved. Once the equilibrium state was achieved, all the series of the process were carefully analyzed. Once the distillate was observed in the water-cooled distillation head of the exhauster, the feed pumps were turned on and the distillate and bottom series were collected. 66.3 gr./hour of funds from the direct esterification reactor were pumped into the depletion primer. The series of funds was collected through a spill system in which the spill point was set at a particular volume, which represented the desired residence time. Once the equilibrium state was achieved, the distillate stream was carefully analyzed. The residence time in the exhaustion splitter was 235 minutes. The distillate recovered from the aerial system was pumped back to the direct esterification reactor at a rate of 47.1 gr./hour. The process had a 108% production of AB, based on AA. The process had an AB production of 95%, based on BuOH. Excessive process production at 100% is possible because the AB can be recovered from the AAOP present in the AA feed.

Example 2 Preparation of butyl acrylate using the plug flow reactor.

In this example, a plug flow reactor was introduced before the direct esterification reactor. The rest of the equipment and procedures were the same as in Example 1. The direct esterification reactor was set at 115 ° C. and 760 mm. of Hg. They were pumped to the plug flow reactor: 219 g / hour of AA, 201.3 g / hour of BuOH and 2.2 g / hour of sulfuric acid. To the direct esterification reactor, 3.6 g / hour of inhibitor, 422.5 g / hour of effluent from the plug flow reactor, and 2 g / hour of sulfuric acid were pumped into the direct esterification reactor. A part of the water recovered from the aerial system was pumped back to the reactor at a reflux index of 477.7 gr./hour. The residence time in the reactor was 180 minutes. The depletion primer was driven at 130 ° C and 35 mm. of Hg with 17.5% by weight of sulfuric acid present. 96.8 gr./hour of the bottoms were pumped from the direct esterification reactor to the exhauster. The residence time in the exhauster was 195 minutes. The distillate recovered from the aerial system was pumped back to the direct esterification reactor at an index of 72.1 gr./hour. The process had an AB production of 98%, based on AA. The process had an AB production of 100%, based on BuOH.

Example 3 Preparation of butyl acrylate with recycling of the bottoms in the fractionation reactor.

In this example, the bottoms of the fractionation reactor were recycled to the direct esterification reactor. The rest of the equipment and procedures were the same as in Example 1. The direct esterification reactor was set at 130 ° C. and 760 mm. of Hg. Direct injection of 3.8 g / hour of inhibitor, 230.8 g / hour of AA, 377.4 g / hour of BuOH, and 3.2 g / hour of sulfuric acid were pumped into the direct esterification reactor. A part of the water recovered from the aerial system was pumped back to the reactor at a reflux index of 468.7 gr./hour. The residence time in the reactor was 154 minutes. The depletion primer was driven at 130 ° C. and 35 mm. of Hg with 22.2% by weight of sulfuric acid present. 65 gr./hour of the bottoms were pumped from the direct esterification reactor to the exhauster. The residence time in the exhauster was 195 minutes. The distillate recovered from the aerial system was pumped back to the direct steaming reactor at a rate of 36.4 gr./hour. The bottoms of the exhauster were recycled to the direct esterification reactor at an index of 10.5 gr./hour. The process had an AB production of 98%, based on AA. The process had an AB production of 100%, based on BuOH.

Example 4 Preparation of butyl acrylate using a methane sulphonic acid catalyst.

In this example, the sulfuric acid was replaced by methane sulphonic acid as the catalyst. The rest of the equipment and procedures were the same as in Example 1. The direct esterification reactor was set at 119 ° C. and 760 mm. of Hg. Direct injection of 3.8 g / hour of inhibitor, 212.1 g / hour of AA, 383.9 g / hour of BuOH and 5.2 g / hour of methane sulphonic acid were pumped into the direct esterification reactor. A part of the water recovered from the aerial system was pumped back to the reactor at a reflux index of 470.5 gr./hour. The residence time in the reactor was 131 minutes. The depletion primer was driven at 130 ° C. and 35 mm. of Hg with 28% by weight of sulphonic methane acid present. 73 gr./hour from the bottoms of the direct esterification reactor were pumped to the exhauster. The residence time in the exhaustion splitter was 278 minutes. The distillate recovered from the aerial system was pumped back to the direct esterification reactor at a rate of 57.3 g / hour. The process had a production of 7AB of 98%, based on AA. The process had an AB production of 100%, based on BuOH.

The above examples demonstrate that the process of this invention is effective in producing alkyl (meth) acrylates in high yields by means of direct esterification, efficiently and economically. That is, the process is carried out in a high aqueous medium, where: (1) the reaction pressures may be atmospheric, and do not need to be pressures reduced; (2) AA and AB, ie start materials and reaction product, are separated in the reactor; and (3) the transesterification reaction and the reaction of the hydrolysis of the by-product occur in the same reactor.

Accordingly, the need for additional and separate process steps and equipment (for byproduct hydrolysis and / or separation of starting material / reaction product) and / or larger equipment (due to reaction pressures) is eliminated. lower) .

Claims (10)

Claims

1. A process comprising: (A) charging a reactor with a Cx-C4 alcohol, a (meth) acrylic acid, a strong acid catalyst, and at least 5% by weight of water to form a reaction mixture; (B) reacting the reaction mixture to form a (meth) d-C4 alkyl acrylate and impurities from the process, wherein the impurities from the process are hydrolyzed in the reactor; and (C) separating the C 1 -C 4 alkyl (meth) acrylate and water formed during the reaction from the reaction mixture.

2. The process according to claim 1, wherein the strong acid catalyst is sulfuric acid, aqluyl sulphonic acid or alkyl sulfonic acid supported with polymer.

3. The process according to claim 1, wherein the reactor is further charged with at least one inhibitor.

4. The process according to claim 3, wherein the inhibitor is selected from hydroquinone, the mono-methyl ether of hydroquinone, 2, 2, 6, 6-tetramethyl-1-piperidinyloxy, free radical of piperidinyl 4- methacryloyloxy-2,2,6,6-tetramethyl, 4-hydroxy-2,2,6,6-tetramethyl-N-hydroxy-piperidine, hydroxy butylated anisol, naphthoquinone, anthranil and combinations thereof.

5. The process according to claim 1, wherein the separation of water and Cx-C4 alkyl (meth) acrylate from the reaction mixture is achieved by means of azeotropic distillation, directly from the reaction mixture.

6. The process according to claim 5, wherein the azeotropic distillation is carried out in the reactor using aqueous reflux in a distillation column of 20 to 100 trays.

7. The process according to claim 6, wherein the azeotropic distillation is carried out at atmospheric pressure.

8. The process according to claim 3, wherein the d-C4 alcohol is butanol, the (meth) acrylic acid is acrylic acid, the strong acid catalyst is sulfuric acid and the inhibitor is 2, 2, 6, 6-tetramethyl -1-piperidinyloxy.

9. The process according to claim 8, wherein the butanol and the acrylic acid are charged to the reactor at a molar ratio of acrylic acid to butanol from 1: 1 to 1: 1.7.

10. The process according to claim 1, wherein the water is present from 6 to 18% by weight, and the strong acid catalyst is present from 3.5 to 15% by weight.