Continuous production method for N-butyl (meth)acrylate using a catalyst recycle system

Abstract

The present invention relates to a continuous process for the preparation of n-butyl (meth)acrylate by reaction of (meth)acrylic acid with n-butanol in the presence of an acidic catalyst and a polymerization inhibitor, comprising, according to a first embodiment, a step of carrying out esterification in a reactor (A) equipped with a column (B), in which the components (meth)acrylic acid and n-butanol are used in a molar ratio ranging from 1.0:1.0 to 1.0:2.0, preferably ranging from 1.0:1.1 to 1.0:1.5, and the esterification is carried out at a temperature ranging from 80 to 150°C, preferably ranging from 100 to 130°C, and at a pressure ranging from 0.2 to 5.0 bar, preferably ranging from 0.4 to 1.5 bar. - condensing the vapor stream in a condenser (C) to form an n-butyl(meth)acrylate-rich organic phase and an aqueous phase; - continuously separating the organic phase from the aqueous phase by means of a phase separator (D); - feeding the obtained reaction product (6) to a rectification column (E); - separating the azeotropes in the rectification column (E): a) water and n-butyl(meth)acrylate, b) n-butanol and n-butyl(meth)acrylate, c) n-butanol and water, d) n-butanol, n-butyl(meth)acrylate and water, wherein the rectification column (E) has a bottom temperature in the range of from 80 to 150°C and a top temperature in the range of from 70 to 130°C and a pressure in the range of from 0.2 to 5 bar, preferably from 0.4 to 1.5 bar. the rectification column (E) is operated at an absolute pressure in the range of 0.5 bar, the rectification column (E) comprising the steps of: discharging an azeotrope-enriched gas stream at the top of the rectification column (E); condensing the gas stream in a condenser (F) to form an n-butyl(meth)acrylate-enriched organic phase and an aqueous phase; continuously separating the organic phase from the aqueous phase by means of a phase separator (G); continuously discharging at least a portion of the organic phase from the phase separator (G), this discharged portion of the n-butyl(meth)acrylate-enriched organic phase representing the raw product stream (15); and discharging a high-boiling bottom product (23) from the bottom of the rectification column (E), wherein the mass flow ratio between the high-boiling bottom product (23) and the (meth)acrylic acid fed as reaction to the reactor (A) is 0.5.- feeding a high-boiling side stream (7) of the discharged high-boiling bottom product (23) to a mixer (H), in which the mass flow ratio between the high-boiling side stream (7) and the high-boiling bottom product (23) is in the range of 0.01 to 0.50, preferably in the range of 0.05 to 0.08; - feeding the mixture (10) obtained from the mixer (H) to a downstream extraction phase separator (I); - feeding the mixture (10) in the extraction phase separator (I) and a step of continuously separating the organic raffinate (11) and the catalyst-containing aqueous extract (12), the aqueous extract (12) being at least partially returned to the reactor (A) and / or the rectification column (E), wherein external water (19) is fed to the mixer (H), and the mass flow ratio of the external water (19) to the high-boiling side stream (7) of the discharged high-boiling bottom product (23) is in the range of 0.08 to 0.50, preferably in the range of 0.10 to 0.30.

Description

[Technical Field]

[0001] The present invention relates to a process for continuously preparing n-butyl (meth)acrylate by reacting (meth)acrylic acid with n-butanol in the presence of an acidic catalyst and a polymerization inhibitor. [Background technology]

[0002] n-Butyl (meth)acrylate is particularly advantageously used as a crosslinker or adhesion promoter in dispersions which are preferably used in the production of contact lenses or as adhesives, paints, e.g. inks, including printing inks, textile, leather or paper auxiliaries, and curable coatings.

[0003] The esterification of alkanols with acids generally proceeds as a typical equilibrium reaction, which is catalyzed by a strong acid and results in the removal of water of esterification as a typical condensation reaction. Typically, removal of water of esterification from the reaction mixture shifts the esterification equilibrium in the direction of the forward reaction, resulting in an increased conversion of the desired product.

[0004] Acidic catalysts commonly used in esterification reactions include inorganic acids such as sulfuric acid, organic acids such as alkanesulfonic acids, and ion exchange resins.

[0005] The water of esterification can be removed by distillation, typically as a component of an azeotrope that also contains the desired ester. Generally, the esterification reaction is carried out so that the water of esterification is continuously removed from the reaction mixture, while the majority of the desired ester formed remains in the reaction mixture.

[0006] An example of this type of esterification is one in which the water of esterification is distilled off by adding an organic solvent as an azeotropic entrainer. It is also possible for an alkanol used in excess to function as such an azeotropic entrainer.

[0007] In one variant for the preparation of n-butyl(meth)acrylate, the water of esterification is separated by distillation as a constituent of the n-butyl(meth)acrylate / n-butanol / water heteroazeotrope, and n-butanol and / or n-butyl(meth)acrylate are at least partially recycled to the esterification as organic phase.

[0008] The problem with this method is that the water of esterification formed in the esterification must be at least partially discharged from the process. Furthermore, part of the discharged catalyst is usually incinerated, which results in the release of SO 2 , which is undesirable when using, for example, sulfuric or sulfonic acids. x causing emissions.

[0009] A method for continuously preparing alkyl esters of (meth)acrylic acid by recycling an acid catalyst to the reaction zone of a reactor is disclosed in EP 0 795 536 (BASF AG). This document describes the recycling of the acid catalyst only for the alkyl ester 2-ethylhexyl acrylate (2-EHA). There is no disclosure of the applicability and transferability of the present teachings to other physical process conditions in the preparation of n-butyl acrylate, as well as the esterification reaction carried out therein with downstream specific purification to obtain the alkyl ester n-butyl acrylate.

[0010] A method for continuously preparing alkyl (meth)acrylates, particularly n-butyl acrylate, is disclosed in WO 2012 / 026661 (LG Chem). In this method, an organic acid catalyst is recycled to the reaction zone of a reactor. However, recycling is limited in that the mass flow rate of the bottom outlet from the rectification column is already biphasic. Therefore, the mass flow rate of the bottom outlet from the rectification column already contains an organic phase and an aqueous phase. [Prior art documents] [Patent documents]

[0011] [Patent Document 1] European Patent Application Publication No. 0795 536 [Patent Document 2] International Publication No. 2012 / 026661 Summary of the Invention [Problem to be solved by the invention]

[0012] Therefore, the problem addressed was to provide a continuous process for preparing n-butyl(meth)acrylate with catalyst recycling to enable improved catalyst recovery during the preparation of n-butyl(meth)acrylate using an acidic catalyst, which also reduces catalyst emissions.

[0013] A further problem that was addressed was to carry as little water as possible through the esterification and distillative purification in order to make the process energy efficient. A related effect is that fewer secondary components are formed during the process, since the reaction mixture is exposed to lower temperatures during the process.

[0014] A further problem that was addressed was achieving greater conversion and yield of n-butyl (meth)acrylate. [Means for solving the problem]

[0015] These problems are solved according to the invention by claim 1 or claim 2. The invention further relates to preferred configurations of the method according to claims 3 to 14.

[0016] In this document, reference numbers in parentheses are used to aid in better understanding during reading, but do not have a limiting effect and merely constitute one possible example in each case of multiple implementation options.

[0017] The process according to the present invention is based on the reactants n-butanol and (meth)acrylic acid. In this document, (meth)acrylic acid is used to refer to a (meth)acrylic acid quality having preferably at least 98% by weight, more preferably at least 99.5% by weight, of (meth)acrylic acid, even more preferably 0.2% by weight or less of water, and even more preferably 0.03% by weight or less of acetic acid, propionic acid, and isobutyric acid, respectively. It is preferred to use an n-butanol quality containing at least 99.5% by weight of n-butanol, 0.05% by weight or less of n-butanal, 0.02% by weight or less of dibutyl ether, 0.1% by weight or less of other alcohols, and 0.05% by weight or less of water. The color number is preferably APHA 5 or less, and the acid number is preferably 0.03 mgKOH / g or less.

[0018] Suitable polymerization inhibitors acting as stabilizers are, for example, N-oxides (nitroxyl or N-oxyl radicals, i.e. compounds having at least one NO group), such as 4-hydroxy-2,2,6,6-tetramethylpiperidine-N-oxyl (HO-TEMPO), 4-oxo-2,2,6,6-tetramethylpiperidine-N-oxyl, 4-acetoxy-2,2,6,6-tetramethylpiperidine-N-oxyl, 2,2,6,6-tetramethylpiperidine-N-oxyl, bis(1- 4,4',4"-tris(2,2,6,6-tetramethylpiperidine-N-oxyl) sebacate, 4,4',4"-tris(2,2,6,6-tetramethylpiperidine-N-oxyl) phosphite or 3-oxo-2,2,5,5-tetramethylpyrrolidine-N-oxyl; monohydric or polyhydric phenols optionally having one or more alkyl groups, for example, alkylphenols, for example, o-, m- or p-cresol (methylphenol), 2-tert-butylphenol, 4-tert-butylphenol, phenol, 2,4-di-tert-butylphenol, 2-methyl-4-tert-butylphenol, 2-tert-butyl-4-methylphenol, 2,6-tert-butyl-4-methylphenol, 4-tert-butyl-2,6-dimethylphenol or 6-tert-butyl-2,4-dimethylphenol; quinones such as hydroquinone, hydroquinone monomethyl ether, 2-methylhydroquinone or 2,5-di-tert-butylhydroquinone; hydroxyphenols such as catechol (1,2-dihydroxybenzene) or benzoquinone; aminophenols such as p-aminophenol; nitrosophenols such as p-nitrosophenol; alkoxyphenols such as 2-methoxyphenol (guaiacol, catechol monomethyl ether), 2-ethoxyphenol, 2-isopropoxyphenol, 4-methoxyphenol (hydroquinone monomethyl ether), mono- or di-tert-butyl-4-methoxyphenol;Tocopherols, such as α-tocopherol and 2,3-dihydro-2,2-dimethyl-7-hydroxybenzofuran (2,2-dimethyl-7-hydroxycoumaran), aromatic amines, such as N,N-diphenylamine or N-nitrosodiphenylamine; phenylenediamines such as N,N'-dialkyl-p-phenylenediamines (the alkyl groups may be the same or different, may each independently consist of 1 to 4 carbon atoms, and may be linear or branched), for example, hydroxylamines such as N,N'-dimethyl-p-phenylenediamine or N,N'-diethyl-p-phenylenediamine, N,N-diethylhydroxylamine; imines such as methylethylimine or methylene violet; and N-methyl-4-toluenesulfonyl ether. sulfonamides such as toluenesulfonamide or N-tert-butyl-4-toluenesulfonamide, oximes such as aldoximes, ketoximes such as diethylketoxime, methylethylketoxime or salicylaldoxime, phosphorus compounds such as triphenylphosphine, triphenylphosphite or triethylphosphite, hypophosphorous acid or alkyl esters of phosphorous acid; sulfur compounds such as diphenylsulfide or phenothiazine; metal salts such as copper or manganese salts, cerium salts, nickel salts and chromium salts, for example, chlorides, sulfates, salicylates, tosylates, acrylates or acetates, for example, copper acetate, copper(II) chloride, copper salicylate, cerium(III) acetate or cerium(III) ethylhexanoate, or mixtures thereof;

[0019] The polymerization inhibitor or polymerization inhibitor mixture used is preferably at least one compound from the group comprising hydroquinone, hydroquinone monomethyl ether, phenothiazine, 4-hydroxy-2,2,6,6-tetramethylpiperidine-N-oxyl, 4-oxo-2,2,6,6-tetramethylpiperidine-N-oxyl, bis(1-oxyl-2,2,6,6-tetramethylpiperidin-4-yl) sebacate, 2-tert-butylphenol, 4-tert-butylphenol, 2,4-di-tert-butylphenol, 2-tert-butyl-4-methylphenol, 6-tert-butyl-2,4-dimethylphenol, 2,6-di-tert-butyl-4-methylphenol, 2-methyl-4-tert-butylphenol, hypophosphorous acid, copper(II) acetate, copper(I) chloride, copper(II) chloride, copper(II) salicylate and cerium(III) acetate.

[0020] It is particularly preferred to use phenothiazine (PTZ) and / or hydroquinone monomethyl ether (MEHQ) as polymerization inhibitors.

[0021] Very particular preference is given to using PTZ as polymerization inhibitor in the esterification and / or distillation.

[0022] The polymerization inhibitor is preferably dissolved in one or more liquid organic compounds, which are preferably n-butanol and / or n-butyl (meth)acrylate.

[0023] Useful esterification catalysts include standard mineral acids and sulfonic acids, preferably sulfuric acid, phosphoric acid, alkylsulfonic acids (e.g., methanesulfonic acid, trifluoromethanesulfonic acid) and arylsulfonic acids (e.g., benzenesulfonic acid, p-toluenesulfonic acid or dodecylbenzenesulfonic acid) or mixtures thereof.

[0024] Particularly preferred are sulfuric acid, methanesulfonic acid, p-toluenesulfonic acid, m-toluenesulfonic acid, o-toluenesulfonic acid or mixtures thereof.

[0025] Very particular preference is given to using p-toluenesulfonic acid as esterification catalyst.

[0026] The acid catalyst is particularly a homogeneous catalyst. The acid catalyst used is preferably sulfuric acid, phosphoric acid, alkylsulfonic acids (e.g., methanesulfonic acid, trifluoromethanesulfonic acid) and arylsulfonic acids (e.g., benzene-, p-toluene- or dodecylbenzenesulfonic acid) or mixtures thereof, more preferably p-toluenesulfonic acid.

[0027] According to the present invention, a method for continuously preparing n-butyl (meth)acrylate by reacting (meth)acrylic acid with n-butanol in the presence of an acidic catalyst and a polymerization inhibitor comprises: carrying out an esterification in a reactor (A) having a column (B) above it, in which the (meth)acrylic acid and n-butanol components are used in a molar ratio ranging from 1.0:1.0 to 1.0:2.0, preferably from 1.0:1.1 to 1.0:1.5, and the esterification is carried out at a temperature ranging from 80 to 150°C, preferably from 100 to 130°C, and at an absolute pressure ranging from 0.2 to 5.0 bar, preferably from 0.4 to 1.5 bar, so that a reaction effluent (6) and a vapor stream are obtained at the top of the column (B); Discharging the vapor stream at the top of the column (B); condensing the vapor stream in a condenser (C) to form an organic phase and an aqueous phase; - continuously separating the organic phase from the aqueous phase by means of a phase separator (D); feeding the reaction effluent (6) obtained into a rectification column (E); The following azeotropes in the rectification column (E): a) water and n-butyl (meth)acrylate; b) n-butanol and n-butyl (meth)acrylate; c) n-butanol and water; d) removing n-butanol, n-butyl (meth)acrylate and water, The rectification column (E) is operated at a bottom temperature in the range of from 80 to 150°C, a top temperature in the range of from 70 to 130°C and an absolute pressure in the range of from 0.2 to 5 bar; Discharging the azeotrope-enriched gas stream at the top of the rectification column (E), condensing the gas stream in a condenser (F) to form an n-butyl (meth)acrylate-rich organic phase and an aqueous phase; - continuously separating the organic phase from the aqueous phase by means of a phase separator (G); continuously removing at least a portion of the organic phase from the phase separator (G), this removed portion of the n-butyl (meth)acrylate-rich organic phase constituting the crude product stream (15); a step of discharging a high-boiling bottom discharge (23) from the bottom of the rectification column (E), wherein the mass flow ratio between the high-boiling bottom discharge (23) and the (meth)acrylic acid (1) fed as a reactant to the reactor (A) is in the range of 0.5 to 5; feeding a high-boiling substream (7) of the discharged high-boiling bottom effluent (23) to a mixer (H), the mass flow ratio of the high-boiling substream (7) to the high-boiling bottom effluent (23) being in the range of 0.01 to 0.50, preferably 0.05 to 0.08; feeding the mixture (10) obtained from the mixer (H) into a downstream extraction phase separator (I); continuous separation of the mixture (10) in the extractive phase separator (I) to obtain an organic raffinate (11) and a catalyst-containing aqueous extract (12), the aqueous extract (12) being at least partially recycled to the reactor (A) and / or the rectification column (E), External water (19) is fed to the mixer (H), and the mass flow ratio between the external water (19) and the high-boiling side stream (7) of the discharged high-boiling bottoms effluent (23) is in the range of 0.08 to 0.50, preferably 0.10 to 0.30.

[0028] Surprisingly, it has been found that supplying external water (19) to the mixer (H) means that the resulting mixture (10) always splits into two phases after the start-up phase of the plant. As a result, it is possible to maintain the catalyst under all operating conditions encountered in the plant. Furthermore, the water content in the plant, particularly in the reactor (A) or the fractionator (E), can also be kept to a minimum. This further makes the process more energy-efficient, and the level of secondary components formed during the process is reduced. Furthermore, higher conversions and higher yields of n-butyl (meth)acrylate are achieved.

[0029] An alternative process of the present invention for the continuous preparation of n-butyl (meth)acrylate by reacting (meth)acrylic acid with n-butanol in the presence of an acidic catalyst and a polymerization inhibitor includes: carrying out an esterification in a reaction zone (E1), which reaction zone (E1) is located at the bottom of the rectification column (E), in which the (meth)acrylic acid and n-butanol components are used in a molar ratio ranging from 1.0:1.0 to 1.0:2.0, preferably from 1.0:1.1 to 1.0:1.5, and in which the esterification is carried out at a temperature ranging from 80 to 150°C, preferably from 100 to 130°C, and at an absolute pressure ranging from 0.2 to 5.0 bar, preferably from 0.4 to 1.5 bar; The following azeotropes form as a result of the esterification: a) water and n-butyl (meth)acrylate; b) n-butanol and n-butyl (meth)acrylate; c) n-butanol and water; d) removing n-butanol, n-butyl (meth)acrylate and water, The removal is carried out by means of a rectification column (E) operated at a bottom temperature ranging from 80 to 150°C and a top temperature ranging from 70 to 130°C and an absolute pressure ranging from 0.2 to 5 bar; Discharging the azeotrope-enriched gas stream at the top of the rectification column (E), condensing the gas stream in a condenser (F) to form an n-butyl (meth)acrylate-rich organic phase and an aqueous phase; - continuously separating the organic phase from the aqueous phase by means of a phase separator (G); continuously removing at least a portion of the organic phase from the phase separator (G), this removed portion of the n-butyl (meth)acrylate-rich organic phase constituting the crude product stream (15); Discharging a high-boiling bottoms discharge (23) from the bottom of the rectification column (E), wherein the mass flow ratio between the high-boiling bottoms discharge (23) and the (meth)acrylic acid fed as reactant to the reaction zone (E1) is in the range of 0.05 to 0.5; feeding a high-boiling substream (7) of the discharged high-boiling bottom effluent (23) to a mixer (H), the mass flow ratio of the high-boiling substream (7) to the high-boiling bottom effluent (23) being in the range of 0.01 to 1.0, preferably 0.10 to 0.70; feeding the mixture (10) obtained from the mixer (H) into a downstream extraction phase separator (I); - continuously separating the mixture (10) in an extractive phase separator (I) to obtain an organic raffinate (11) and a catalyst-containing aqueous extract (12), the aqueous extract (12) being at least partially recycled to the rectification column (E), External water (19) is fed to the mixer (H), and the mass flow ratio between the external water (19) and the high-boiling side stream (7) of the discharged high-boiling bottoms effluent (23) is in the range of 0.08 to 0.50, preferably 0.10 to 0.30.

[0030] The same technical effect can also be observed in this alternative process of the present invention. It is recognized here that supplying external water (19) to the mixer (H) means that the resulting mixture (10) always splits into two phases after the start-up phase of the plant. As a result, the catalyst can always be maintained under all operating conditions encountered in the plant. Furthermore, the water content in the plant, particularly in the reaction zone (E1) or the rectification column (E), can also be kept to a minimum. Furthermore, the process results in a more energy-efficient process and a lower level of secondary components formed during the process. Furthermore, higher conversions and higher yields of n-butyl (meth)acrylate are achieved.

[0031] The term "rectification column" in this document should be considered as a general term for any device in which heat is supplied to generate vapor, which rises and contacts a downwardly flowing liquid phase.

[0032] Rectification columns are known in their general design and have conventional equipment, such as a bottom evaporator, a high-boiling outlet evaporator, or a low-boiling outlet condenser, with the high boiling point preferably located in the bottom region of the column and the low boiling point preferably located in the upper region. A portion of the mass stream from the high-boiling outlet is usually returned to the bottom region of the column. However, in principle, it is also possible to heat the bottom region, for example, via heating the column wall in the bottom region, and / or to integrate an evaporator in the bottom region. A vapor stream is usually withdrawn from the top of the column and fed to a condenser. This vapor stream is also usually referred to as low-boiling effluent. A portion of the vapor stream condensed in the condenser is returned to the column, and the remaining portion of the condensed vapor stream is discharged as distillate. The reflux ratio here refers to the ratio between the condensed vapor stream returned to the column and the condensed vapor stream withdrawn as distillate. A reflux ratio in the range of 10% to 200% is generally established. The column internals used in the rectification column (E) can in principle be all standard internals, such as trays, structured packings and / or random packings. Among trays, bubble cap trays, sieve trays, valve trays, Thormann trays and / or dual flow trays are preferred. Among random packings, those with rings, spirals, saddles or blades are preferred. Furthermore, the rectification column (E) can also be equipped with further standard components, for example, pressure reducers, flow regulators or sensor-based adjustments. In principle, it is also possible to connect several rectification columns to one another in series or parallel connection, which can also function collectively as a "rectification column" (E).

[0033] In this document, when using acrylic acid, a component is referred to as a low-boiling substance if its boiling point at standard pressure is lower than that of n-butyl acrylate. Similarly, a component is referred to as a high-boiling substance if its boiling point at standard pressure is equal to or greater than that of n-butyl acrylate, which has a boiling point of 147°C at standard pressure.

[0034] In this document, when using methacrylic acid, a component is referred to as a low-boiling substance if its boiling point at standard pressure is lower than that of n-butyl (meth)acrylate. Similarly, a component is referred to as a high-boiling substance if its boiling point at standard pressure is equal to or greater than that of n-butyl (meth)acrylate, which has a boiling point of 163°C at standard pressure.

[0035] In this document, the term "reactor" generally defines one reactor (A) or two or more interconnected reactors functioning as one "reactor" (A). The reactor (A) further comprises a reactor heating element for heating the reaction mixture. The reactor heating element can be, for example, an immersion heater within the reactor (A), a tubing system comprising coiled or semi-coiled tubing arranged on the outer jacket surface of the reactor and / or within the reactor (A), an electric heating system arranged on the outer jacket surface of the reactor and / or within the reactor (A), an evaporator outside the reactor (A) through which the reaction mixture flows at least partially, or a jacketed design of the reactor outer wall, in which a fluid separated from the reaction mixture, such as a liquid, gas and / or heating steam, is temperature-controlled, thereby setting a predetermined heating temperature and thereby heating the reaction mixture in the reactor (A). Two or more reactor heating elements can generally be used to heat the reaction mixture in the reactor (A). For example, the reaction mixture can be heated simultaneously or at least partially with a time delay using a jacketed design of the reactor outer wall and an evaporator outside the reactor (A). Furthermore, the reactor (A) is equipped with a column (B) at the top, which preferably separates water by distillation. Column (B) itself is a distillation column with internal structures. Such internal structures are trays, such as bubble-cap trays, perforated trays, especially dual-flow trays, random packing, structured packing, etc. In a further preferred embodiment, the reactor (A) is incorporated into a rectification column (E) so that esterification can occur at the bottom of the rectification column (E), i.e., in the reaction zone (E1).

[0036] In this document, the term "mass flow ratio" between a real value A1 and a real value B1 is equivalent to the division ratio with A1 as the numerator and B1 as the denominator. Therefore, the following formula is applicable:

number

[0037] The term "external water" (19) in this document is understood to mean water that originates from outside the process and is introduced into the process by the inlet of the mixer (H). The external water (19) is preferably demineralized water, more preferably fully demineralized water. The external water (19) preferably has a pH in the range of 4.5 to 10.5, more preferably in the range of 6.5 to 10.0. Preferably, the external water (19) contains only low levels, if any, of electrolyte components, preferably less than 0.1% by weight.

[0038] The acidic esterification catalyst, i.e., the acid catalyst, used is preferably para-toluenesulfonic acid. Its content in the reaction zone (E1) or reactor (A), based on the reaction mixture present therein, is suitably 0.1% to 10.0% by weight, preferably 0.1% to 6.0% by weight. Other organic sulfonic acids, such as methanesulfonic acid, benzenesulfonic acid, dodecylbenzenesulfonic acid, and / or sulfuric acid, can also be used. Their amount is equimolar to that of para-toluenesulfonic acid. Corresponding mixtures are also possible. The content of catalytically active acid in the rectification column bottom, based on the mixture present in the rectification column (E), can advantageously be 2.5% to 50.0% by weight of para-toluenesulfonic acid or equivalent amounts of another organic sulfonic acid and / or sulfuric acid.

[0039] In a preferred configuration of the process, the mixture (10) obtained from the mixer (H) has a temperature at the outlet of the mixer (H) ranging from 20 to 100° C., preferably from 70 to 95° C. This has the advantage that both the aqueous and organic phases do not have to be heated after the process, thus saving energy.

[0040] In a preferred configuration of the process, the high-boiling bottoms effluent (23) has a water concentration of less than 10% by weight, preferably less than 5% by weight, which has the advantage that less water has to be distilled off in the rectification column (E), and therefore less energy is consumed to heat the rectification column (E), and consequently lower levels of secondary components are formed.

[0041] In a preferred configuration of the process, a side stream of the aqueous extract is returned to the reactor (A) or reaction zone (E1), the mass flow ratio of the side stream of the aqueous extract to the total mass flow of the aqueous extract (12) being in the range of 0.1 to 1.0, preferably in the range of 0.8 to 1.0, which has the advantage of allowing an efficient return of the catalyst without increasing the water content in the plant.

[0042] In a preferred configuration of the process, external water (19) is added to the high-boiling side stream (7) of the discharged high-boiling bottoms (23) in the mixer (H) in such a way that a mass flow rate of external water is achieved such that the phase ratio between the aqueous extract (12) obtained and the organic raffinate (11) obtained is in the range of 0.08 to 0.5 kg / kg, preferably 0.1 to 0.3 kg / kg, which has the advantage that a two-phase separation of the two phases is guaranteed, the smallest possible amount of catalyst is discharged from the process, and the water content in the esterification and distillation is minimized.

[0043] A mass flow ratio of 0.5 kg / kg or less between the mass flow rate of the external water (19) and the mass flow rate of the high-boiling side stream (7) of the discharged high-boiling bottoms (23) is sufficient for successful recovery of the acid catalyst and reduces the energy requirements of the process, especially for water evaporation, compared to using larger amounts of water. A mass flow ratio of at least 0.08 kg / kg between the mass flow rate of the external water (19) and the high-boiling side stream (7) of the discharged high-boiling bottoms (23) is beneficial for effective extraction and phase separation in the extraction phase separator (I).

[0044] In a preferred configuration of the present process, at least a side stream of the organic raffinate (11) is fed to the cleavage reactor (J). In particular, the cleavage reactor (J) is connected downstream of the mixer (H) and the extractive phase separator (I). This offers the advantage that the organic component after phase separation, i.e., the organic raffinate (11), is fed to the cleavage reactor (J), meaning that the bottom mixture in the cleavage reactor (J), removed via conduit 21, is anhydrous and therefore less corrosive. Furthermore, this arrangement avoids recycling by-products formed in the cleavage reactor (G) together with the extract to reactor (A) or reaction zone (E1). This arrangement also improves phase separation in the extractive phase separator (I) for reasons including the fact that the viscosity of the continuous phase, i.e., the organic raffinate (11), is lower and the density difference between the two phases, i.e., the organic raffinate (11) and the aqueous extract (12), is higher compared to the extract downstream of the cleavage reactor (J).

[0045] In a preferred configuration of the process, a side stream of organic raffinate (11) is fed to the cleavage reactor (J), the mass flow ratio of the side stream of organic raffinate (11) to the total mass flow of the organic raffinate being in the range of 0.1 to 1.0, preferably in the range of 0.95 to 1.0, which has the advantage that high boiling point substances can be cleaved and the cleavage products, such as (meth)acrylic acid and n-butanol, can be reused as reactants in the process.

[0046] In a preferred configuration of the process, the esterification is carried out at a temperature in the range of from 90 to 130° C., preferably in the range of from 95 to 105° C., and at an absolute pressure in the range of from 0.8 to 2.0 bar, preferably in the range of from 1.0 to 1.5 bar, which offers the advantage that the conversion is maximized and the formation of secondary components is kept to a minimum.

[0047] In a preferred configuration of the process, the substream (8) of high-boiling bottoms effluent (23) is fed to the cleavage reactor (J) in a mass flow ratio of high-boiling bottoms effluent (23) to high-boiling substream (7) in the range of 0.0 to 10.0, preferably 0.1 to 1.0, which provides the advantage that there is sufficient catalyst for the cleavage reaction in the cleavage reactor (J).

[0048] In a preferred configuration of the process, the substream (9) of high boiling bottoms effluent (23) is fed to the reactor (A) at a mass flow ratio to the total high boiling bottoms effluent (23) in the range of 0.1 to 0.99, preferably 0.85 to 0.95, which has the advantage that the catalyst can be returned and reused.

[0049] In a preferred configuration of the process, the high-boiling bottoms effluent (23) has a water content in the range of 0.1% to 10.0% by weight, more preferably 0.1% to less than 10.0% by weight, even more preferably 0.1% to less than 5% by weight, and especially 0.1% to 4.5% by weight. This offers the advantage that not too much water is present in the process, thus requiring less energy and forming fewer secondary components. However, at the same time, single-phase recovery of the high-boiling bottoms effluent (23) and subsequent two-phase separation of the resulting mixture (10) after adding water in the mixer (H) are guaranteed.

[0050] In a preferred configuration of the process, the high-boiling bottoms effluent (23) has an oligomer and / or polymer content in the range of 1% to 80% by weight, more preferably 10% to 65% by weight, and in particular 20% to 60% by weight. Oligomers and / or polymers are understood to mean molecules with a mass-average molar mass of more than 1000 g / mol. An oligomer and / or polymer content of 80% by weight or less in the high-boiling bottoms effluent (23) provides the advantage of a low viscosity, which results in improved subsequent phase separation in the extraction phase separator (I). The mass flow ratio of the external water (19) to the high-boiling sidestream (7) of the discharged high-boiling bottoms effluent (23) can be further reduced.

[0051] In a preferred configuration of the process, the high-boiling bottoms effluent (23) has a catalyst content in the range of 0.1 wt. % to 10.0 wt. %, which provides the advantage that not as much catalyst needs to be removed, thus allowing the process to operate more energy efficiently and allowing less catalyst to be used in the process.

[0052] In a preferred embodiment of the present method, the acidic catalyst contains p-toluenesulfonic acid in the range of 0 to 100% by weight, preferably 80 to 100% by weight, more preferably 95 to 100% by weight, which has the advantage that the catalyst has high selectivity, high reactivity, and long life in this method, and therefore the esterification proceeds very efficiently.

[0053] In a further configuration of the process, the high boiling point bottoms effluent (23) is single phase. In this case, external water (19) is added to the mixer (H) in an amount such that the resulting mixture (10) is biphasic.

[0054] In a further configuration of the process, a side stream of the organic phase (14) from the phase separator (G) is returned to the rectification column (E) at a reflux ratio based on the organic phase ranging from 0.1 to 1.0, and a side stream of the aqueous phase (16) from the phase separator (G) at a reflux ratio based on the aqueous phase ranging from 1 to 10. This has the advantage that less organic phase has to be discharged from the process and less aqueous phase has to be removed, and both the esterification and the removal in the rectification column (E) proceed in an energy-efficient manner. Furthermore, the levels of secondary components formed are lower.

[0055] In a further configuration of the process, the acidic catalyst is present in the reaction zone (E1) of the rectification column (E) or in the resulting reaction effluent of the reactor (A) in a concentration ranging from 0.1% to 10% by weight, which has the advantage that not as much catalyst needs to be removed, and therefore the process can be run more energy-efficiently and less catalyst can be used in the process.

[0056] In a further configuration of the process, a substream of the aqueous extract (12) is fed to the reactor (A) at a mass flow ratio to the total high-boiling bottoms effluent (23) in the range of from 0.01 to 0.50, preferably from 0.01 to 0.30.

[0057] The invention will be discussed in more detail below with reference to the drawings, which should be considered as schematic representations and which do not constitute limitations of the invention, for example with respect to specific dimensions or design variations. [Brief explanation of the drawings]

[0058] [Figure 1] 1 is a first embodiment of the process of the invention for the continuous preparation of n-butyl (meth)acrylate, using a reactor A with a downstream rectification column E. In this case, the catalyst-containing aqueous extract is recycled to reactor A and / or rectification column E. [Figure 2] 2 is a second embodiment of the process according to the invention for the continuous preparation of n-butyl (meth)acrylate, in which reaction zone E1 is integrated in the bottom region of rectification column E. In this case, the catalyst-containing aqueous extract is recycled to rectification column E. [Figure 3] 3 is a third embodiment of a process for continuously preparing n-butyl (meth)acrylate as a comparative example, using a reactor A with a downstream rectification column E. In this case, there is no recycle of the catalyst-containing aqueous extract to reactor A and / or rectification column E. DETAILED DESCRIPTION OF THE INVENTION

[0059] List of reference numbers used 1. Conduit for supplying reactant streams 2 Residual vapor from condenser C 3. Conduit for recycling the organic phase from the liquid-liquid phase separator D 4. Conduit for discharging the organic phase from the liquid-liquid phase separator D 5. A conduit for supplying the aqueous phase from the liquid-liquid phase separator D to the liquid-liquid phase separator G 6. A conduit for supplying the reaction effluent obtained to the rectification column E 7. Conduit for the side stream of the high boiling point bottom discharge 23 to the mixer H or the high boiling point side stream of the high boiling point bottom discharge 23 to the mixer H 8. Conduit for the side stream of high boiling point bottoms effluent 23 to cleavage reactor J or a side stream of the high boiling point bottoms effluent 23 to the cleavage reactor J. 9 Conduit for recycling a side stream of high boiling point bottoms effluent 23 to reactor A, or a side stream of high boiling point bottoms effluent 23 to reactor A 10 Mixture obtained from mixer H 11. A conduit for feeding the organic raffinate from the extraction phase separator I to the cleavage reactor J or the organic raffinate 12 Conduit for recycling the aqueous extract from the extractive phase separator I to the reactor or rectification column E or aqueous extract 13 Residual vapor from condenser F 14 Conduit for recycling the organic phase from the liquid-liquid phase separator G to the rectification column E or the organic phase from the liquid-liquid phase separator G 15 Phase separator G or conduit for discharging a side stream of the organic phase from the crude product stream 16 Conduit for recycling the aqueous phase from the liquid-liquid phase separator G to the rectification column E or the aqueous phase from the liquid-liquid phase separator G 17 Conduit for discharging a substream of the aqueous phase from the liquid-liquid phase separator G 19 Conduit for supplying external water to mixer H or external water 21 Conduit for discharging high boiling point bottom effluent from cleavage reactor J 22 Conduit for discharging high boilers from cleavage reactor J 23 Conduit for the high-boiling bottoms discharge of rectification column E or conduit for the high-boiling bottoms discharge of rectification column E 24 Conduit for supplying polymerization inhibitor stream to condenser C 25 Conduit for feeding polymerization inhibitor stream to condenser F Reactor A B column C condenser D phase separator E. Rectification tower E1 Reactor integrated into the rectification column F Condenser G phase separator H Mixer I Extraction phase separator J Cleavage Reactor

[0060] FIG. 1 shows a schematic diagram of a process flow diagram of a chemical engineering process according to a first embodiment of the method of the present invention, in which external water 19 is added to a mixer H.

[0061] A mass flow of reactants comprising n-butanol, acrylic acid, PTZ and p-toluenesulfonic acid as an esterification catalyst is fed to reactor A through conduit 1. Column B, located above reactor A, separates the vapor mixture leaving reactor A by means of its separation internals. A downstream condenser C, which may optionally be supplemented by a postcooler, at least partially condenses the vapor stream emerging from column B. A solution of the polymerization inhibitor PTZ is added to condenser C via conduit 24. The non-condensable fraction from condenser C, which contains low-boiling impurities, is withdrawn in vapor form via conduit 2.

[0062] The condensed vapor stream flows as a condensate into liquid-liquid phase separator D, where it separates into an aqueous phase and an organic phase. The aqueous phase, primarily water, is conducted from liquid-liquid phase separator D to liquid-liquid phase separator G via conduit 5.

[0063] The resulting reaction effluent, which contains, inter alia, n-butyl acrylate, unconverted reactants, and high-boiling impurities, is withdrawn from the bottom of reactor A via conduit 6 and fed to a rectification column E having separation internals. In particular, water, n-butyl acrylate, and alcohol are separated in rectification column E. Vapors leaving the column are fed to condenser F, which may optionally be supplemented with a postcooler, and are partially condensed therein. A solution of the polymerization inhibitor PTZ is added to condenser F via conduit 25. The non-condensable fraction from condenser F, which contains low-boiling impurities, is withdrawn in vapor form via conduit 13, while the resulting condensate flows into a liquid-liquid phase separator G, where it separates into an organic phase and an aqueous phase.

[0064] The organic phase, which mainly comprises n-butyl acrylate and n-butanol, is partially returned to the rectification column E as reflux via conduit 14, the remainder being discharged via conduit 15 for further work-up. The aqueous phase, which mainly comprises water, is partially returned to the rectification column E as reflux via conduit 16, the remainder being discharged via conduit 17 for further work-up.

[0065] Mainly higher boiling impurities and spent catalyst are withdrawn from the bottom of the rectification column E via conduit 23 and returned via downstream conduits 7, 8 and 9 to the cleavage reactor J, mixer H and reactor A. In mixer H, external water 19 is mixed as the extractant and the resulting mixture is fed via conduit 10 to the extractive phase separator I.

[0066] In the extractive phase separator I, the resulting mixture is separated into an organic raffinate and a catalyst-containing aqueous extract. The aqueous extract, which mainly contains water, is fed to reactor A via conduit 12. The organic raffinate is fed at least partially to cleavage reactor J via conduit 11 for further workup. The high-boiling bottoms effluent is also fed at least partially to cleavage reactor J via conduit 8 for further workup. The bottoms effluent from cleavage reactor J is discharged from the process via conduit 21, while gaseous substances are withdrawn at the top of cleavage reactor J via conduit 22. The gaseous substances can then be condensed and finally returned to reactor A.

[0067] FIG. 2 shows a schematic process flow diagram of a chemical engineering process according to a second alternative embodiment of the method of the present invention, in which, compared to the first embodiment, external water 19 is added to mixer H and a reaction zone E1 is integrated at the bottom of rectification column E.

[0068] FIG. 3 shows a schematic diagram of the process flow diagram of the chemical engineering process as a comparative example, in which, compared to the first embodiment, mixer H and extractive phase separator I are not present, meaning that recycling of the aqueous extract from extractive phase separator I to reactor A is not possible.

[0069] Example The following example process is simulated by thermodynamic simulation. For this purpose, Aspen Plus® software (Aspen) is used. This software can be found on the website https: / / www.aspentech.com (accessed July 15, 2022). Aspen is a comprehensive simulation software package used for modeling, simulating, and optimizing chemical processes and plants in industry. Aspen has an extensive model database for modeling basic operations and a material database for the physical properties of many different substances. The properties of mixtures are calculated by Aspen from the physical data of the pure substances using various thermodynamic models.

[0070] Comparative Example 1 A thermodynamic simulation of the whole plant according to Figure 3 was carried out with Aspen and gives the following results:

[0071] The reactant stream is fed to reactor A via conduit 1 at a mass flow rate of 1000 kg / h, the reactant stream being a mixture having the following composition: n-butanol: 582.2 kg / h Acrylic acid: 413.9 kg / h p-Toluenesulfonic acid: 3.9 kg / h.

[0072] Column B, located above reactor A, separates the vapor mixture exiting reactor A by means of its separation internals. Downstream condenser C at least partially condenses the vapor stream formed from column B.

[0073] A solution of polymerization inhibitor is added to the condenser C via conduit 24 at a mass flow rate of 2 kg / h. The solution of polymerization inhibitor has the following composition: n-Butyl acrylate: 98% by weight Phenothiazine: 2% by weight.

[0074] The non-condensable fraction from condenser C, which contains low boiling impurities, is withdrawn in vapor form via conduit 2.

[0075] The esterification is carried out in reactor A at a temperature of 105° C., a pressure of 470 mbar and a residence time of 2 hours.

[0076] 27 kg / h of organic phase are withdrawn as organic distillate from the liquid-liquid phase separator D downstream of the condenser C via conduit 4 and 798 kg / h are returned to column B as reflux via conduit 3.

[0077] The organic phase as organic distillate has the following composition: Water: 19.6% by weight n-butanol: 77.3% by weight n-butyl acrylate: 1.5% by weight Acrylic acid: less than 0.01% by weight n-Butyl acetate: 1.2% by weight Phenothiazine: Less than 0.01% by weight Unknown: 0.4% by weight.

[0078] 145 kg / h of aqueous phase are led from the liquid-liquid phase separator D via conduit 5 to the liquid-liquid phase separator G. The aqueous phase has the following composition: Water: 94.0% by weight n-butanol: 5.9% by weight Unknown: 0.1% by weight.

[0079] At the bottom of reactor A, the reaction effluent obtained is withdrawn via conduit 6 at a mass flow rate of 1415 kg / h and fed to rectification column E with the following composition: Water: 0.2% by weight n-butanol: 10.2% by weight n-Butyl acrylate: 68.0% by weight Acrylic acid: 5.0% by weight n-Butyl acetate: less than 0.1% by weight p-Toluenesulfonic acid: 1.6% by weight Phenothiazine: 0.1% by weight Unknown: 14.8% by weight.

[0080] In a rectification column E with 13 theoretical plates, water, n-butanol and n-butyl acrylate are withdrawn from the rectification column E in vapor form at the top, condensed in a condenser F and then fed to a liquid-liquid phase separator G. In the upper region of the column, the pressure is 1059 mbar and the temperature is 95°C.

[0081] A solution of polymerization inhibitor is added to the condenser F via conduit 25 at a mass flow rate of 2 kg / h. Said solution of polymerization inhibitor has the following composition: n-Butyl acrylate: 98% by weight Phenothiazine: 2% by weight.

[0082] 803 kg / h of organic phase are withdrawn from the liquid-liquid phase separator G as organic distillate via conduit 15 for further purification. A side stream with a mass flow rate of 365 kg / h is returned to the rectification column E as reflux via conduit 14.

[0083] The organic phase has the following composition: Water: 2.7% by weight n-butanol: 16.8% by weight n-butyl acrylate: 80.3% by weight Acrylic acid: less than 0.01% by weight n-Butyl acetate: 0.1% by weight Phenothiazine: Less than 0.01% by weight Unknown: 0.1% by weight.

[0084] 83 kg / h of aqueous phase are withdrawn as aqueous distillate from the liquid-liquid phase separator G via conduit 17. A side stream of aqueous phase with a mass flow rate of 724 kg / h is returned to the rectification column E via conduit 16 as reflux.

[0085] The aqueous phase has the following composition: Water: 96.7% by weight n-butanol: 2.0% by weight n-butyl acrylate: 1.3% by weight Unknown: 0.1% by weight.

[0086] At the bottom of the rectification column E, a side stream of high-boiling bottoms 23 with a mass flow rate of 53 kg / h is led via conduit 8 to the cleavage reactor J, and a side stream of high-boiling bottoms 23 with a mass flow rate of 585 kg / h is returned to the reactor A via conduit 9. The vapor mixture formed in the cleavage reactor J is removed via conduit 22, while the bottoms mixture is removed via conduit 21.

[0087] The high boiling point bottoms effluent 23 has the following composition: Water: 3.1% by weight n-butanol: 4.3% by weight n-Butyl acrylate: 21.4% by weight Acrylic acid: 11.1% by weight p-Toluenesulfonic acid 3.4% by weight Phenothiazine: 0.2% by weight Unknown: 56.5% by weight.

[0088] The content of oligomers and / or polymers is included in the unknown. Example 1 In this Example 1, an embodiment of the present invention of a method for preparing n-butyl (meth)acrylate according to Figure 1 is simulated by Aspen software. The simulation gives the following results:

[0089] The reactant stream is fed to reactor A via conduit 1 at a mass flow rate of 1000 kg / h, the reactant stream having the following composition: n-butanol: 584.3 kg / h Acrylic acid: 414.4 kg / h p-Toluenesulfonic acid: 1.3 kg / h.

[0090] The esterification in reactor A is carried out at a temperature of 105° C., an absolute pressure of 470 mbar and a residence time of 2 hours.

[0091] Column B, located above reactor A, separates the vapor mixture exiting reactor A by means of its separation internals. Downstream condenser C at least partially condenses the vapor stream formed from column B.

[0092] A solution of polymerization inhibitor is added to the condenser C via conduit 24 at a mass flow rate of 2 kg / h. The solution of polymerization inhibitor has the following composition: n-Butyl acrylate: 98% by weight Phenothiazine: 2% by weight.

[0093] The non-condensable fraction from condenser C, which contains low boiling impurities, is withdrawn in vapor form via conduit 2.

[0094] 27 kg / h of organic phase are withdrawn as organic distillate from the liquid-liquid phase separator D downstream of the condenser C via conduit 4 and 934 kg / h are returned to column B as reflux via conduit 3.

[0095] The organic phase has the following composition: Water: 19.6% by weight n-butanol: 76.9% by weight n-butyl acrylate: 1.7% by weight Acrylic acid: less than 0.01% by weight n-Butyl acetate: 1.4% by weight Phenothiazine: Less than 0.01% by weight Unknown: 0.4% by weight.

[0096] The aqueous phase with a mass flow rate of 155 kg / h is led from the liquid-liquid phase separator D to the liquid-liquid phase separator G via conduit 5 . The aqueous phase has the following composition: Water: 94.0% by weight n-butanol: 5.9% by weight Unknown: 0.1% by weight.

[0097] At the bottom of reactor A, a liquid mixture with a mass flow rate of 1422 kg / h is withdrawn via conduit 6 and fed to rectification column E with the following composition: Water: 0.2% by weight n-butanol: 10.2% by weight n-Butyl acrylate: 68.0% by weight Acrylic acid: 5.0% by weight n-Butyl acetate: 0.0% to 0.1% by weight p-Toluenesulfonic acid 1.6% by weight Phenothiazine 0.2% by weight Unknown: 14.7% to 14.8% by weight.

[0098] In the rectification column E with 13 theoretical plates, water, n-butanol and n-butyl acrylate are withdrawn from the rectification column E at the top in vapor form and, after condensation in a condenser F, are fed to a liquid-liquid phase separator G. In the top region of the rectification column E, the absolute pressure is 1059 mbar and the temperature is 95°C.

[0099] A solution of polymerization inhibitor is added to the condenser F via conduit 25 at a mass flow rate of 2 kg / h. Said solution of polymerization inhibitor has the following composition: n-Butyl acrylate: 98% by weight Phenothiazine: 2% by weight.

[0100] The organic phase with a mass flow rate of 807 kg / h is withdrawn from the liquid-liquid phase separator G via conduit 15 as organic distillate for further purification. A side stream with a mass flow rate of 367 kg / h is returned to the rectification column E as reflux via conduit 14.

[0101] The organic phase has the following composition: Water: 2.6% by weight n-butanol: 16.7% by weight n-Butyl acrylate: 80.5% by weight Acrylic acid: less than 0.01% by weight n-Butyl acetate: 0.1% by weight Phenothiazine: Less than 0.01% by weight Unknown: 0.1% by weight.

[0102] A mass flow rate of 101 kg / h of the aqueous phase is withdrawn from the liquid-liquid phase separator G as aqueous distillate via conduit 17 for further purification. A side stream with a mass flow rate of 728 kg / h is returned to the rectification column E as reflux via conduit 16.

[0103] At the bottom of the rectification column E, a side stream is led via conduit 7 to mixer H with a mass flow rate of 49 kg / h, a side stream with a mass flow rate of 24 kg / h is led via conduit 8 to cleavage reactor J, and a side stream with a mass flow rate of 588 kg / h is led back to reactor A via conduit 9.

[0104] The liquid mixture has the following composition: Water: 3.1% by weight n-butanol: 4.3% by weight n-butyl acrylate: 21.7% by weight Acrylic acid: 11.1% by weight p-Toluenesulfonic acid 3.4% by weight Phenothiazine: 0.4% by weight Unknown: 56.0% by weight.

[0105] The content of oligomers and / or polymers is included in the unknown.

[0106] In mixer H, the side stream from rectification column E containing catalyst and high-boiling impurities via conduit 7 is mixed with a mass flow of external water 19 so as to establish a phase ratio of 0.2 kg / kg, the phase ratio being given by the ratio of the mass flow of external water 19 plus the mass flow of the aqueous phase of side stream 7 to the mass flow of the organic phase of side stream 7. The mixture obtained from mixer H is then fed via conduit 10 to extractive phase separator I.

[0107] 45 kg / h of organic raffinate is removed from the extractive phase separator I as an organic mixture via conduit 11 and added to the cleavage reactor J for further purification. The vapor mixture formed in the cleavage reactor J is removed via conduit 22, while the bottom mixture is removed via conduit 21.

[0108] The organic raffinate has the following composition: Water: 1.9% by weight n-butanol: 4.6% by weight n-butyl acrylate: 23.8% by weight Acrylic acid: 9.0% by weight p-Toluenesulfonic acid Less than 0.05% by weight Phenothiazine: 0.2% by weight Unknown: 60.5% by weight.

[0109] The content of oligomers and polymers is included in the unknown.

[0110] The aqueous extract is returned from the extractive phase separator I via conduit 12 to reactor A at a mass flow rate of 15 kg / h.

[0111] The aqueous extract has the following composition: Water: 71.4% by weight n-butanol: 0.6% by weight n-butyl acrylate: 0.9% by weight Acrylic acid: 10.1% by weight p-Toluenesulfonic acid 11.5% by weight Phenothiazine: 0.7% by weight Unknown: 4.8% by weight.

[0112] Summarizing from the two examples, the following is evident at the same catalyst concentration in reactor A:

[0113] The catalyst-containing reactant stream fed to reactor A via conduit 1 has a lower catalyst concentration in Inventive Example 1 than in Comparative Example 1. Thus, in Inventive Example 1, a mass flow rate of 1.3 kg / h of fresh catalyst is fed to conduit 1, whereas in Comparative Example 1, a mass flow rate of 3.9 kg / h of fresh catalyst is fed to conduit 1.

[0114] This is because in Inventive Example 1, the aqueous extract 12 is returned to reactor A at a mass flow rate of 15 kg / h and the catalyst concentration is 11.5 wt. %. In Inventive Example 1, the phase ratio between the external water 19 and the high-boiling side stream 7 in the high-boiling bottom effluent 23 resulting at the outlet of mixer H is 0.2 kg / kg, so there is phase separation, whereas in Comparative Example 1, no external water 19 is added and the water concentration in the high-boiling bottom effluent 23 is only 0.03 kg / kg. Therefore, there is no phase separation in Comparative Example 1, which means that the catalyst cannot be returned to reactor A either.

[0115] Therefore, in Example 1 of the present invention, 66.7% of the amount of catalyst is saved.

[0116] Furthermore, the influence of the placement of the cleavage reactor J during the process was investigated experimentally.

[0117] Example 2 In Example 2, as in Example 1, the cleavage reactor J was connected downstream of the mixer H and the extractive phase separator I. The high-boiling side stream 7 of the high-boiling bottom effluent 23 from the bottom of the rectification column E was fed to the mixer H. The viscosity of the organic raffinate 11 taken from the extractive phase separator I and fed to the cleavage reactor J was 0.7 mPas. External water 19 was used in the mixer H so that the mass flow ratio between the external water 19 and the high-boiling side stream 7 of the withdrawn high-boiling bottom effluent 23 was in the range of 0.13 kg / kg to 0.34 kg / kg. The separation time in the extractive phase separator I was 60 seconds, the density difference was 100 kg / m 3 It was.

[0118] Example 3 In Example 3, the cleavage reactor J was connected upstream of the mixer H, and a side stream 8 of the high-boiling bottom effluent 23 from the bottom of the rectification column E (as in Figure 3) was first fed to the cleavage reactor J. Then, the high-boiling bottom effluent 21 from the cleavage reactor J was fed to the mixer H, downstream of which was connected an extractive phase separator I. Here, the viscosity of the organic raffinate 11 taken out from the extractive phase separator I was 30 mPas. External water 19 was used in the mixer H so that the mass flow ratio between the external water 19 and the high-boiling bottom effluent 21 from the cleavage reactor J was 1 kg / kg. The separation time in the extractive phase separator I was 240-300 seconds, and the density difference was 12 kg / m 3 It was small. [Explanation of symbols]

[0119] 1 conduit 2 conduit 3 Conduit 4 Conduit 5 Conduit 6 Conduit 7 Sidestream 7 Downstream conduit 7 High boiling point substance side stream 7 High boiling point side stream 7 Conduit 8 Sidestream 8 Downstream conduit 8 Conduit 9 Conduit 10 Conduit 11 Organic Raffinate 11 Conduit 12 Aqueous extract 12 Conduit 13 Conduit 14 Conduit 14 Sidestream 15 Conduit 16 Conduit 17 Conduit 19 External Water 21 Conduit 21 High boiling bottom effluent 22 Conduit 23 High boiling bottom effluent 23 Conduit 24 Mass flow rate 24 Conduit 25 Catheter

Claims

1. A method for continuously preparing n-butyl (meth)acrylate by reacting (meth)acrylic acid with n-butanol in the presence of an acidic catalyst and a polymerization inhibitor, comprising: carrying out an esterification in a reactor (A) having a column (B) above it, in which the (meth)acrylic acid and n-butanol components are used in a molar ratio ranging from 1.0:1.0 to 1.0:2.0, preferably from 1.0:1.1 to 1.0:1.5, and the esterification is carried out at a temperature ranging from 80 to 150°C, preferably from 100 to 130°C, and at an absolute pressure ranging from 0.2 to 5.0 bar, preferably from 0.4 to 1.5 bar, so that a reaction effluent (6) and a vapor stream are obtained at the top of said column (B); Discharging the vapor stream at the top of the column (B); condensing the vapor stream in a condenser (C) to form an organic phase and an aqueous phase; continuously separating the organic phase from the aqueous phase by a phase separator (D); feeding the resulting reaction effluent (6) to a rectification column (E); The following azeotrope in the rectification column (E): a) water and n-butyl (meth)acrylate; b) n-butanol and n-butyl (meth)acrylate; c) n-butanol and water; d) removing n-butanol, n-butyl (meth)acrylate and water, the rectification column (E) is operated at a bottom temperature in the range of from 80 to 150°C and a top temperature in the range of from 70 to 130°C and an absolute pressure in the range of from 0.2 to 5 bar; Discharging the azeotrope-enriched gas stream at the top of the rectification column (E); condensing the gas stream in a condenser (F) to form an n-butyl (meth)acrylate-rich organic phase and an aqueous phase; continuously separating the organic phase from the aqueous phase by a phase separator (G); continuously removing at least a portion of the organic phase from the phase separator (G), this removed portion of the n-butyl (meth)acrylate-rich organic phase constituting a crude product stream (15); Discharging a high-boiling bottom discharge (23) from the bottom of the rectification column (E), wherein the mass flow ratio of the high-boiling bottom discharge (23) to the (meth)acrylic acid fed as a reactant to the reactor (A) is in the range of 0.5 to 5; feeding a high boiling substream (7) of the discharged high boiling bottom effluent (23) to a mixer (H), wherein the mass flow ratio of the high boiling substream (7) to the high boiling bottom effluent (23) is in the range of 0.01 to 0.50, preferably 0.05 to 0.08; feeding the mixture (10) obtained from said mixer (H) into a downstream extraction phase separator (I); and continuously separating the mixture (10) in the extractive phase separator (I) to obtain an organic raffinate (11) and a catalyst-containing aqueous extract (12), the aqueous extract (12) being at least partially recycled to the reactor (A) and / or the rectification column (E), The method, wherein external water (19) is fed to the mixer (H), and the mass flow ratio of the external water (19) to the high-boiling side stream (7) of the discharged high-boiling bottom effluent (23) is in the range of 0.08 to 0.50, preferably 0.10 to 0.

30.

2. A method for continuously preparing n-butyl (meth)acrylate by reacting (meth)acrylic acid with n-butanol in the presence of an acidic catalyst and a polymerization inhibitor, comprising: carrying out an esterification in a reaction zone (E1), which reaction zone (E1) is located at the bottom of the rectification column (E), in which the (meth)acrylic acid and n-butanol components are used in a molar ratio ranging from 1.0:1.0 to 1.0:2.0, preferably from 1.0:1.1 to 1.0:1.5, and in which the esterification is carried out at a temperature ranging from 80 to 150°C, preferably from 100 to 130°C, and at an absolute pressure ranging from 0.2 to 5.0 bar, preferably from 0.4 to 1.5 bar; The following azeotrope forms as a result of the esterification: a) water and n-butyl (meth)acrylate; b) n-butanol and n-butyl (meth)acrylate; c) n-butanol and water; d) removing n-butanol, n-butyl (meth)acrylate and water, said removal is also carried out by means of a rectification column (E) operated at a bottom temperature ranging from 80 to 150° C. and a top temperature ranging from 70 to 130° C. and an absolute pressure ranging from 0.2 to 5 bar, preferably from 0.4 to 1.5 bar; Discharging the azeotrope-enriched gas stream at the top of the rectification column (E); condensing the gas stream in a condenser (F) to form an n-butyl (meth)acrylate-rich organic phase and an aqueous phase; continuously separating the organic phase from the aqueous phase by a phase separator (G); continuously removing at least a portion of the organic phase and the phase separator (G), this removed portion of the n-butyl (meth)acrylate-rich organic phase constituting the crude product stream (15); Discharging a high-boiling bottom discharge (23) from the bottom of the rectification column (E), wherein the mass flow ratio of the high-boiling bottom discharge (23) to the (meth)acrylic acid fed as a reactant to the reaction zone (E1) is in the range of 0.05 to 0.5; feeding a high boiling substream (7) of the discharged high boiling bottom effluent (23) to a mixer (H), wherein the mass flow ratio of the high boiling substream (7) to the high boiling bottom effluent (23) is in the range of 0.01 to 1.00, preferably 0.10 to 0.70; feeding the mixture (10) obtained from said mixer (H) into a downstream extraction phase separator (I); continuously separating the mixture (10) in the extractive phase separator (I) to obtain an organic raffinate (11) and a catalyst-containing aqueous extract (12), the aqueous extract (12) being at least partially recycled to the rectification column (E), external water (19) is fed to the mixer (H), and the mass flow ratio of the mass flow of the external water (19) to the high boiling side stream (7) of the discharged high boiling bottoms effluent (23) is in the range of 0.08 to 0.50, preferably in the range of 0.10 to 0.

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3. 3. The method according to claim 1 or 2, wherein the mixture (10) obtained from the mixer (H) has a temperature in the range of 20 to 100°C, preferably in the range of 70 to 95°C, at the outlet of the mixer (H).

4. 4. The process according to any one of claims 1 to 3, wherein a side stream of the aqueous extract is returned to the reactor (A) or to the reaction zone (E1), and the mass flow ratio between the side stream of the aqueous extract and the total mass flow of the aqueous extract (12) is in the range of 0.1 to 1.0, preferably in the range of 0.8 to 1.

0.

5. 5. The method according to claim 1, wherein the external water (19) is added to the high boiling side stream (7) of the discharged high boiling bottom effluent (23) in the mixer (H) in such a way that a mass flow rate of the external water (19) is achieved such that the phase ratio between the aqueous extract (12) obtained and the organic raffinate (11) obtained is in the range of 0.08 to 0.5 kg / kg, preferably in the range of 0.1 to 0.3 kg / kg.

6. 6. The process according to any one of claims 1 to 5, wherein a side stream of the organic raffinate (11) is fed to the cleavage reactor (J), and the mass flow ratio between the side stream of the organic raffinate (11) and the total mass flow of the organic raffinate is in the range of 0.1 to 1.0, preferably in the range of 0.95 to 1.

0.

7. 7. The process according to claim 6, wherein the sub-stream (8) of the high-boiling bottoms effluent (23) is fed to the cleavage reactor (J) at a mass flow ratio of the high-boiling bottoms effluent (23) to the high-boiling sub-stream (7) in the range of 0.0 to 10.0, preferably 0.1 to 1.

0.

8. 8. The method of any one of claims 1 to 7, wherein the high-boiling bottoms effluent (23) has a water content in the range of 0.1 wt. % to 10.0 wt. %.

9. 9. The method of any one of claims 1 to 8, wherein the high-boiling bottoms effluent (23) has a catalyst content in the range of 0.1 wt. % to 10.0 wt. %.

10. 10. The method of any one of claims 1 to 9, wherein the acidic catalyst comprises p-toluenesulfonic acid in the range of 0 wt% to 100 wt%, preferably in the range of 80 wt% to 100 wt%, more preferably in the range of 95 wt% to 100 wt%.

11. 11. The method of any one of claims 1 to 10, wherein the high-boiling bottoms effluent (23) is monophasic.

12. 12. The process according to any one of claims 1 to 11, wherein the organic phase side stream (14) from the phase separator (G) is returned to the rectification column (E) at a reflux ratio based on the organic phase in the range of 0.1 to 1.0, and the aqueous phase side stream (16) from the phase separator (G) is returned to the rectification column (E) at a reflux ratio based on the aqueous phase in the range of 1 to 10.

13. 13. The process according to any one of claims 1 to 12, wherein the acidic catalyst is present in the reaction zone (E1) of the rectification column (E) or in the resulting reaction effluent of the reactor (A) in a concentration ranging from 0.1 wt. % to 10.0 wt. %.

14. 14. The process according to claim 1 or any one of claims 3 to 13, wherein the sub-stream of aqueous extract (12) is fed to the reactor (A) at a mass flow ratio to the total high-boiling bottoms effluent (23) in the range of from 0.01 to 0.50, preferably from 0.01 to 0.30.

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