Continuous production method for n-butyl (meth)acrylate including catalyst recycling process

Abstract

The present invention relates to a method for continuously producing n-butyl (meth)acrylate by reacting (meth)acrylic acid with n-butanol in the presence of an acid catalyst and a polymerization inhibitor. In a first embodiment, the method comprises the following steps: carrying out an esterification in a reactor (A) comprising a column (B) installed thereon, in which the components (meth)acrylic acid and n-butanol are used in a molar ratio ranging from 1.0:1.0 to 10:2.0, 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 a pressure ranging from 0.2 to 5.0 bar, preferably from 0.4 to 1.5 bar. the reaction is carried out at an absolute pressure in the range of 100 bar, whereby the resulting reaction product (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), thereby forming an organic phase and an aqueous phase enriched in n-butyl (meth)acrylate; * continuously separating the organic phase from the aqueous phase by means of a phase separator (D); * feeding the resulting reaction product (6) to a rectification column (E); * separating the azeotropes in the rectification column (E) consisting of a) water and n-butyl (meth)acrylate, b) n-butanol and n-butyl (meth)acrylate, c) n-butanol and water, and d) n-butanol, n-butyl (meth)acrylate and water, wherein the rectification column (E) has a sump temperature in the range of 80 to 150°C, a top temperature in the range of 70 to 130°C, and a pressure of 0.2 to 50°C. bar range, preferably 0.4 to 1.* discharging a gas stream enriched in the azeotrope at the top of the rectification column (E); * condensing the gas stream in a condenser (F), thereby forming an organic phase enriched in n-butyl (meth)acrylate 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), said discharged portion of the organic phase enriched in n-butyl (meth)acrylate constituting the raw product stream (15); * discharging a high-boiling sump discharge (23) from the sump of the rectification column (E), wherein the high-boiling sump discharge (23) the mass flow ratio of the high-boiling side stream (7) to the high-boiling sump effluent (23) is in the range of 0.01 to 0.5, preferably 0.05 to 0.08; * feeding the mixture (10) obtained from the mixer (H) to an extractive phase separator (I) arranged downstream thereof; and * continuously separating the mixture in the extractive phase separator (I), thereby obtaining an organic raffinate (11) and an aqueous catalyst-containing extract (12). The aqueous extract (12) is at least partially recycled to the reactor (A) and / or the rectification column (E), wherein the aqueous phase sidestream (18) from the phase separator (G), the aqueous phase sidestream (26) from the phase separator (D), and / or the aqueous phase sidestream (5) from the phase separator (D) are fed to the phase separator (G), and then the aqueous phase sidestream (18) from the phase separator (G) is fed to the mixer (H). The mass flow ratio of the aqueous phase sidestream (18) to the high-boiling sidestream (7) of the discharged high-boiling sump effluent (23) is in the range of 0.08 to 0.50, and the mass flow ratio of the aqueous phase sidestream (5) to the high-boiling sidestream (7) of the discharged high-boiling sump effluent (23) is in the range of 0.08 to 0.50.

Description

[Technical Field]

[0001] The present invention relates to 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. [Background technology]

[0002] n-Butyl (meth)acrylate is particularly advantageously used in the manufacture of contact lenses or as a crosslinker or adhesion promoter in dispersions used preferably as adhesives, paints, inks (including e.g. printing inks), in textile, leather or paper auxiliaries and in 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 forward direction, increasing the conversion of the desired product.

[0004] Examples of acid catalysts that are 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 such 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 to add an organic solvent as an azeotropic entrainer and remove the water of esterification by distillation. An alkanol used in excess can also 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 a heterogeneous azeotrope of n-butyl(meth)acrylate / n-butanol / water, and n-butanol and / or n-butyl(meth)acrylate are at least partially recycled as organic phase to the esterification.

[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 typically incinerated, resulting in undesirable SO 2 emissions when using, for example, sulfuric or sulfonic acids. x Generate emissions.

[0009] A method for continuously preparing alkyl esters of (meth)acrylic acid while recycling an acid catalyst to the reaction zone of a reactor is disclosed in EP 0795536(A1) (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 whether this teaching is applicable or transferable to other physical process conditions in the preparation of n-butyl acrylate and to esterification reactions producing the alkyl ester n-butyl acrylate with specific downstream purification.

[0010] A method for continuously preparing alkyl (meth)acrylates, particularly n-butyl acrylate, is disclosed in WO 2012 / 026661 (A1) (LG Chem). In this process, an organic acid catalyst is recycled to the reaction zone of the reactor. However, this recycling is limited in that the mass flow at the bottom outlet from the fractionator is already biphasic. Therefore, the mass flow at the bottom outlet from the fractionator already contains an organic phase and an aqueous phase. [Prior art documents] [Patent documents]

[0011] [Patent Document 1] European Patent No. 0795536 [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, which would allow for improved catalyst recovery during the preparation of n-butyl(meth)acrylate using an acid catalyst, and would also reduce the amount of catalyst discharged.

[0013] A further challenge was to retain as little water as possible throughout the esterification and distillation purification to make the process energy efficient. A concomitant effect is that small amounts of secondary components are formed during the process, as the reaction mixture is exposed to low temperatures in this process.

[0014] A further problem addressed was achieving higher conversions and higher yields of n-butyl (meth)acrylate. [Means for solving the problem]

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

[0016] In this document, reference numbers in parentheses are used to improve comprehension during reading. They have no limiting effect and merely constitute one possible example of 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 specification, (meth)acrylic acid is preferably used to refer to a (meth)acrylic acid quality containing at least 98% by weight, more preferably at least 99.5% by weight, of (meth)acrylic acid, preferably no more than 0.2% by weight of water, and preferably no more than 0.03% by weight 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, no more than 0.05% by weight of n-butanal, no more than 0.02% by weight of dibutyl ether, no more than 0.1% by weight of other alcohols, and no more than 0.05% by weight of water. The color number is preferably APHA 5 or less, and the acid number is preferably no more than 0.03 mgKOH / g.

[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 -oxyl-2,2,6,6-tetramethylpiperidin-4-yl) 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-butyl 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 (wherein 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), such as N,N'-dimethyl-p-phenylenediamine or N,N'-diethyl-p-phenylenediamine, hydroxylamines such as N,N-diethylhydroxylamine, imines such as methylethylimine or methylene violet, and N-methyl-4-toluenesulfonyl ether. sulfonamides such as sulfonamide or N-tert-butyl-4-toluenesulfonamide, oximes such as aldoximes, ketoximes such as diethylketoxime, methylethylketoxime or salicylaldoxime, phosphorus compounds such as triphenylphosphine, triphenylphosphite, triethylphosphite, hypophosphorous acid or alkyl esters of phosphorous acid; sulfur compounds such as diphenyl sulfide, phenothiazine, etc.; metal salts such as copper or manganese, cerium, nickel 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] As polymerization inhibitor or polymerization inhibitor mixture, preference is given to using at least one compound from the group consisting of 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] In the esterification and / or distillation, it is very particularly preferred to use PTZ as polymerization inhibitor.

[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 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 acid (e.g., methanesulfonic acid, trifluoromethanesulfonic acid), arylsulfonic acid (e.g., benzenesulfonic acid, p-toluenesulfonic acid, or dodecylbenzenesulfonic acid), or a mixture 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) equipped with a column (B) at the top, 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, and the resulting reaction output (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 using a phase separator (D); feeding the reaction output (6) obtained into a rectification column (E); In said rectification column (E) the following azeotropic mixture: 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) is operated at a column bottom temperature in the range of 80 to 150°C, a column top temperature in the range of 70 to 130°C and an absolute pressure in the range of 0.2 to 5 bar; Discharging the azeotrope-enriched gas stream at the top of the rectification column (E); condensing said gas stream in a condenser (F) to form an organic phase rich in n-butyl (meth)acrylate and an aqueous phase; continuously separating the organic phase from the aqueous phase using a phase separator (G); continuously removing at least a portion of the organic phase and the phase separator (G), the removed portion of the n-butyl (meth)acrylate-rich organic phase constituting a crude product stream (15); discharging a heavy-boiler bottom output (23) from the bottom of the rectification column (E), wherein the mass flow ratio between the heavy-boiler bottom output (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 heavy materials side stream (7) of the discharged heavy materials bottom output (23) to a mixer (H), wherein the mass flow ratio between the heavy materials side stream (7) and the heavy materials bottom output (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 said mixer (H) into a downstream extraction phase separator (I); continuously separating said mixture (10) in said extract phase separator (I) to obtain an organic raffinate (11) and a catalyst-containing aqueous extract (12), said aqueous extract (12) being at least partially recycled to said reactor (A) and / or said rectification column (E); Including, where: - a substream (18) of the aqueous phase from the phase separator (G), - a side stream (26) of the aqueous phase from the phase separator (D), and / or the aqueous phase substream (5) and the phase separator (D) are fed to the phase separator (G), from which the aqueous phase substream (18) is then The mass flow ratio between the side stream (18) of the aqueous phase fed to the mixer (H) and the side stream (7) of the heavy materials of the discharged heavy materials bottom output (23) is in the range of 0.08 to 0.50, and the mass flow ratio between the side stream (26) of the aqueous phase and the side stream (7) of the discharged heavy materials of the bottom output (23) is in the range of 0.08 to 0.50, preferably in the range of 0.1 to 0.3.

[0028] Surprisingly, it has been recognized that the introduction of the aqueous phase sidestream (18) from the phase separator (G) and / or the aqueous phase sidestream (26) from the phase separator (D) into 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 constantly recycle the catalyst under all operating conditions encountered in the plant. Furthermore, the water content in the plant, for example in the reactor (A) or the fractionator (E), can also be kept to a minimum. This further increases the energy efficiency of the process and reduces the level of secondary components formed during the process. Furthermore, higher conversions and higher yields of n-butyl (meth)acrylate are achieved.

[0029] An alternative method of the present invention 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 the following steps: carrying out an esterification in a reaction zone (E1), which reaction zone (E1) is located at the bottom of the rectification column (E), the (meth)acrylic acid and n-butanol components being 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 being 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 is formed as a result of said esterification: 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 removal is carried out by the rectification column (E), the rectification column being operated at a column bottom temperature in the range of 80 to 150°C, a column top temperature in the range of 70 to 130°C, and an absolute pressure in the range of 0.2 to 5 bar; Discharging the azeotrope-enriched gas stream at the top of the rectification column (E), condensing said gas stream in a condenser (F) to form an organic phase rich in n-butyl (meth)acrylate and an aqueous phase; continuously separating the organic phase from the aqueous phase using a phase separator (G), continuously removing at least a portion of the organic phase and the phase separator (G), the removed portion of the n-butyl (meth)acrylate-rich organic phase constituting a crude product stream (15); discharging a heavy-boiler bottom output (23) from the bottom of the rectification column (E), wherein the mass flow ratio between the heavy-boiler bottom output (23) and 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 heavy-boiler side stream (7) of the discharged heavy-boiler bottom output (23) to a mixer (H), wherein the mass flow ratio between the heavy-boiler side stream (7) and the heavy-boiler bottom output (23) is in the range of 0.01 to 1.0, preferably in the range of 0.10 to 0.70; feeding the mixture (10) obtained from said mixer (H) into a downstream extraction phase separator (I), continuously separating said mixture (10) in said extract phase separator (I) to obtain an organic raffinate (11) and a catalyst-containing aqueous extract (12), said aqueous extract (12) being at least partially recycled to said rectification column (E). Here, the aqueous phase side stream (18) from the phase separator (G) is fed to the mixer (H), and the mass flow ratio between the aqueous phase side stream (18) from the phase separator (G) and the high boilers side stream (7) of the discharged high boilers bottom output (23) is in the range of 0.08 to 0.50, preferably 0.1 to 0.3.

[0030] The same technical effect can be observed with this alternative method of the present invention. It is recognized that the introduction of a side stream (18) of the aqueous phase from the phase separator (G) to the mixer (H) also 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 be recycled at all times under all operating conditions encountered in the plant. Furthermore, the water content in the plant, for example in the reaction zone (E1) or the rectification column (E), can be kept to a minimum. This process is therefore 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.

[0031] In this specification, the term "rectification column" is considered to be a general term for an apparatus in which heat is supplied to produce vapor which rises and contacts a downwardly flowing liquid phase.

[0032] The general design of a rectification column is known and includes conventional devices, such as a bottom evaporator, a high-boiling-point outlet evaporator, or a low-boiling-point outlet condenser, with the high-boiling points preferably located in the bottom region of the column and the low-boiling points preferably located in the top region. A portion of the mass flow from the high-boiling-point 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 sent to a condenser. This vapor stream is also commonly referred to as the low-boiling-point outlet. 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. The reflux ratio is generally established in the range of 10% to 200%. The column internals used in the rectification column (E) can in principle be standard internals, such as trays, structured packing, and / or random packing. Among the trays, bubble cap trays, sieve trays, valve trays, Thormann trays, and / or dual flow trays are preferred, and among the random packings, those containing rings, spirals, saddles, or blades are preferred. Furthermore, the rectification column (E) can also contain further standard components for adjustment, such as pressure reducers, flow regulators, or sensors. In principle, it is also possible to connect several rectification columns together in series or parallel, in which case they can collectively function as a "rectification column" (E).

[0033] As used herein, when acrylic acid is used, a component is referred to as a low boiler if its boiling point at standard pressure is lower than that of n-butyl acrylate. Similarly, a component is referred to as a high boiler sidestream if its boiling point at standard pressure is equal to or higher than that of n-butyl acrylate, which has a boiling point of 147° C. at standard pressure.

[0034] As used herein, when methacrylic acid is used, a component is referred to as a low boiler if its boiling point at standard pressure is lower than that of n-butyl methacrylate. Similarly, a component is referred to as a high boiler sidestream if its boiling point at standard pressure is equal to or higher than that of n-butyl methacrylate, which has a boiling point of 163° C. at standard pressure.

[0035] As used herein, the term "reactor" generally defines one reactor (A) or two or more interconnected reactors functioning as one "reactor" (A). The reactor (A) further includes a reactor heating element for heating the reaction mixture. The reactor heating element may be, for example, an immersion heater within the reactor (A), a tube system including coiled or semi-coiled tubing arranged on the outer jacket surface of the reactor (A) and / or within the reactor (A), an electric heating system arranged on the outer jacket surface of the reactor (A) and / or within the reactor (A), an evaporator outside the reactor (A) (through which the reaction mixture at least partially passes), or a jacket design on the outer reactor wall. In the reactor heating element, a fluid, such as a liquid, gas, and / or heating steam, separated from the reaction mixture is under temperature control, thus setting a predetermined heating temperature, thereby heating the reaction mixture in the reactor (A). Two or more reactor heating elements may generally be used to heat the reaction mixture in the reactor (A). For example, the reaction mixture may be heated simultaneously or at least partially with a time delay using a jacket design on the reactor outer wall and an evaporator outside the reactor (A). Furthermore, the reactor (A) contains a column (B) at the top, which preferably separates water by distillation. The 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 integrated into the rectification column (E) so that the esterification can be carried out in the bottom of the rectification column (E), i.e., in the reaction zone (E1).

[0036] As used herein, the term "mass flow ratio" between real values A1 and B1 corresponds to a division ratio with A1 as the numerator and B1 as the denominator. Thus, the following formula is applicable:

number

[0037] The term "external water" is understood herein to mean water that originates outside the process and is introduced into the process through the inlet of the mixer (H). The external water is preferably demineralized water, more preferably fully demineralized water. The external water 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 contains low levels of electrolyte components, if any, and preferably contains less than 0.1% by weight of electrolyte components.

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

[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) in the range of 20 to 100°C, preferably in the range of 70 to 95°C, which has the advantage that both the aqueous and organic phases do not have to be heated later in the process, thereby saving energy.

[0040] In a preferred configuration of the process, the heavy-boiler bottom output (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 embodiment of the process, a side stream of the aqueous extract is returned to the reactor (A) or 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, which has the advantage of allowing efficient return of the catalyst without increasing the water content in the plant.

[0042] In a preferred embodiment of the process, the mass flow of the aqueous phase sidestream (18) from the phase separator (G) and / or the aqueous phase sidestream (26) from the phase separator (D) is added to the high boiler sidestream (7) of the discharged high boiler bottom output (23) in the mixer (H) in such a way that the phase ratio between the aqueous extract (12) to be obtained and the organic raffinate (11) to be obtained is in the range of 0.08 to 0.5 kg / kg, preferably 0.1 to 0.3 kg / kg, thereby ensuring a two-phase separation of the two phases, minimizing the amount of catalyst discharged from the process and minimizing the water content in the esterification and distillation.

[0043] A mass flow ratio of 0.5 kg / kg or less between the aqueous phase side stream (18) and the heavy materials side stream (7) of the discharged heavy materials bottom output (23), or between the aqueous phase side stream (26) and the heavy materials side stream (7) of the discharged heavy materials bottom output (23), is sufficient for successful acid catalyst recovery and reduces the energy requirements of the process, especially for water evaporation, compared to the use of larger amounts of water. A mass flow ratio of 0.08 kg / kg or more between the aqueous phase side stream (18) and the heavy materials side stream (7) of the discharged heavy materials bottom output (23), or between the aqueous phase side stream (26) and the heavy materials side stream (7) of the discharged heavy materials bottom output (23), is favorable 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 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, i.e., the organic raffinate (11), is fed to the cleavage reactor (J) after phase separation. This means 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 together with the extract in the cleavage reactor (G) to reactor (A) or the reaction zone (E1). This arrangement also improves phase separation in the extractive phase separator (I) due to, among other reasons, the lower viscosity of the continuous phase, i.e., the organic raffinate (11), and the higher density difference between the two phases, i.e., the organic raffinate (11) and the aqueous extract (12), compared to the extract downstream of the cleavage reactor (J).

[0045] In a preferred configuration of the process, the organic raffinate side stream (11) is fed to the cleavage reactor (J), and the mass flow ratio between the organic raffinate side stream (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, 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 embodiment of the process, the esterification is carried out at a temperature in the range of 90 to 130°C, preferably in the range of 95 to 105°C, and at a pressure in the range of 0.8 to 2.0 bar, preferably in the range of 1.0 to 1.5 bar absolute, 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 side stream (8) of the heavy boilers bottom output (23) is fed to the cleavage reactor (J) at a mass flow ratio of the heavy boilers bottom output (23) to the side stream (7) of heavy boilers in the range of 0.0 to 10.0, preferably 0.1 to 1.0, which has 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 side stream (9) of the heavy boilers bottom output (23) is fed to the reactor (A) at a mass flow ratio relative to the total heavy boilers bottom output (23) in the range of 0.1 to 0.99, preferably in the range of 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 heavy-boiler bottom output (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, resulting in less energy requirements and a lower level of secondary component formation. However, at the same time, single-phase recovery of the heavy-boiler bottom output (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 heavy-boilers bottom output (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 heavy-boilers bottom output (23) provides the advantage of a lower viscosity, which improves the subsequent phase separation in the extractive phase separator (I). The mass flow ratio between the aqueous phase side stream (18) and the heavy-boilers side stream (7) of the discharged heavy-boilers bottom output (23), or between the aqueous phase side stream (26) and the heavy-boilers side stream (7) of the discharged heavy-boilers bottom output (23), can be further reduced.

[0051] In a preferred configuration of the present process, the heavy boilers bottoms output (23) has a water content in the range of 0.1 wt. % to 10.0 wt. %. This has the advantage that not as much catalyst needs to be removed, thus making the process more energy efficient and allowing for the use of less catalyst 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, and more preferably 95 to 100% by weight, which provides the advantage that the catalyst has high selectivity, high reactivity, and long life in the process, allowing the esterification to proceed very efficiently.

[0053] In a further configuration of the process, the high boilers bottom output (23) is single-phase. In this case, a side stream of aqueous phase (18) from the phase separator (G) and / or a side stream of aqueous phase (26) from the phase separator (D) is added to the mixer (H) in such an amount that the resulting mixture (10) is two-phase.

[0054] In a further configuration of the process, external water is additionally added to the mixer (H), which is particularly advantageous when the mass flow of the aqueous phase (18) and the phase separator (G) and / or the mass flow of the aqueous phase (26) from the phase separator (D) to the mixer (H) is insufficient for a two-phase separation of the high boilers side stream (7) of the high boilers bottom output (23).

[0055] The organic phase sidestream (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 sidestream (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. This has the advantage that less organic phase is discharged from the process and less aqueous phase is removed, allowing the esterification and removal in the rectification column (E) to proceed in an energy-efficient manner. In addition, lower levels of secondary components are formed.

[0056] 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 output of the reactor (A) at a concentration ranging from 0.1% to 10% by weight, which has the advantage that less catalyst needs to be removed, thus making the process more energy efficient and allowing the use of less catalyst in the process.

[0057] In a further configuration of the process, the aqueous extract side stream (12) is fed to the reactor (A) at a mass flow ratio relative to the total heavy materials bottom output (23) in the range of 0.01 to 0.5, preferably in the range of 0.01 to 0.3.

[0058] 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]

[0059] [Figure 1] 1 shows a first embodiment of the process of the present invention for the continuous preparation of n-butyl (meth)acrylate, using a reactor A with a downstream fractionator E. In this case, the catalyst-containing aqueous extract is recycled to reactor A and / or fractionator E. [Figure 2] 2 shows 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 into the bottom region of rectification column E. In this case, the catalyst-containing aqueous extract is recycled to rectification column E. [Figure 3] FIG. 1 shows a third embodiment of a process for continuously preparing n-butyl (meth)acrylate as a comparative example, using reactor A with a downstream rectification column E. In this case, the catalyst-containing aqueous extract is not recycled to reactor A and / or rectification column E. DETAILED DESCRIPTION OF THE INVENTION

[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 a side stream 18 of the aqueous phase from a liquid-liquid phase separator G is added to a mixer H.

[0061] A mass stream of reactants containing 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 exiting reactor A via its separation internals. A downstream condenser C, optionally 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, containing low-boiling impurities, is withdrawn in vapor form via conduit 2. 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 containing water, is conducted from liquid-liquid phase separator D through conduit 5 to liquid-liquid phase separator G. A further substream of the aqueous phase, which in principle can also serve as an extractant, can be fed from liquid-liquid phase separator D to mixer H through conduit 26.

[0062] The resulting reaction output, 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.

[0063] 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 workup. The aqueous phase, which mainly comprises water, is partially returned to the rectification column E as reflux via conduit 16 and partially led to the mixer H as extractant via conduit 18, the remainder being discharged via conduit 17 for further workup.

[0064] 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, as extractant, the aqueous phase from liquid-liquid phase separator G via conduit 18 and / or the aqueous phase via conduit 26 is mixed with the mass flow from conduit 7 and fed to extractive phase separator I via conduit 10.

[0065] 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 heavy boiler bottoms output is also fed at least partially to cleavage reactor J via conduit 8 for further workup. The bottoms output from cleavage reactor J is discharged from the process via conduit 21, while gaseous materials are withdrawn at the top of cleavage reactor J via conduit 22. The gaseous materials can then be condensed and finally returned to reactor A.

[0066] 2 shows a schematic process flow diagram of a chemical engineering process according to a second alternative embodiment of the method of the invention, in which a side stream 18 of the aqueous phase from the liquid-liquid phase separator G is added to a mixer H. In this figure, in comparison to the first embodiment, in this embodiment a reaction zone E1 is integrated into the bottom of the rectification column E.

[0067] FIG. 3 shows a schematic diagram of the process flow diagram of the chemical engineering process as a comparative example. Compared to the first embodiment, mixer H and extractive phase separator I are not present, which means that recycling of the aqueous extract from extractive phase separator I to reactor A is not possible.

[0068] Example The following process examples are 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 (retrieved 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.

[0069] Comparative Example 1 A thermodynamic simulation of the entire plant according to Figure 3 was performed using Aspen, and the following results were obtained:

[0070] The reactant stream is fed to reactor A via conduit 1 at a mass flow rate of 1000 kg / h. The reactant stream is 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.

[0071] Column B, located above reactor A, separates the vapor mixture exiting reactor A by means of its separation internals. A downstream condenser C at least partially condenses the vapor stream formed from column B.

[0072] 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.

[0073] The non-condensable fraction from condenser C, which contains low boiling impurities, is withdrawn in vapor form via conduit 2.

[0074] 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.

[0075] 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.

[0076] 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: <0.01% by weight n-Butyl acetate: 1.2% by weight Phenothiazine: <0.01% by weight Unknown: 0.4% by weight.

[0077] 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.

[0078] At the bottom of reactor A, the reaction output obtained is withdrawn via conduit 6 at a mass flow rate of 1415 kg / h and fed to a 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 0.1% by weight Unknown: 14.7% to 14.8% by weight.

[0079] In a rectification column E with 13 theoretical plates, water, n-butanol and n-butyl acrylate are withdrawn from the 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.

[0080] 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 having the following composition: n-Butyl acrylate: 98% by weight Phenothiazine: 2% by weight.

[0081] 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.

[0082] 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: <0.01% by weight n-Butyl acetate: 0.1% by weight Phenothiazine: <0.01% by weight Unknown: 0.1% by weight.

[0083] 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.

[0084] 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: <0.05% by weight.

[0085] At the bottom of the rectification column E, a side stream with a mass flow rate of 73 kg / h is led via conduit 8 to the cleavage reactor J, and a side stream with a mass flow rate of 585 kg / h is led back to the reactor A via conduit 9. The vapor mixture formed in the cleavage reactor J is removed via conduit 22, while the bottom mixture is removed via conduit 21.

[0086] The bottom mixture in cleavage reactor J 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

[0087] The oligomeric and / or polymeric content is included as unknown.

[0088] 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] A 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. A 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, said solution of polymerization inhibitor having 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: 19.5% by weight n-butanol: 77.0% by weight n-Butyl acrylate: 1.7% by weight Acrylic acid: <0.01% by weight n-Butyl acetate: 1.4% by weight Phenothiazine: <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 .

[0097] The aqueous phase has the following composition: Water: 94.0% by weight n-butanol: 5.9% by weight Unknown: 0.1% by weight.

[0098] 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.

[0099] In the 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 top region of the rectification column E, the absolute pressure is 1059 mbar and the temperature is 95°C.

[0100] 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 having the following composition: n-Butyl acrylate: 98% by weight Phenothiazine: 2% by weight.

[0101] The organic phase with a mass flow rate of 807 kg / h is 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 367 kg / h is returned to the rectification column E as reflux via conduit 14.

[0102] 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: <0.01% by weight n-Butyl acetate: 0.1% by weight Phenothiazine: <0.01% by weight Unknown: 0.1% by weight.

[0103] The aqueous phase with a mass flow rate of 101 kg / h 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.

[0104] A further substream of the aqueous phase is conducted via conduit 18 to mixer H as extractant for catalyst extraction with a mass flow rate of 10 kg / h.

[0105] 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: <0.05% by weight.

[0106] 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.

[0107] 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

[0108] The oligomeric and / or polymeric content is included as unknown.

[0109] In mixer H, a side stream from rectification column E containing catalyst and impurities is mixed via conduit 7 with side stream 18 from liquid-liquid phase separator G containing water extractant at a phase ratio of 0.2 kg / kg, and the resulting mixture is supplied to extractive phase separator I via conduit 10. Here, the phase ratio is calculated by adding the mass flow rate of the aqueous phase from high-boiling point side stream 7 and the mass flow rate of the water extractant from side stream 18, and dividing the sum by the mass flow rate of the organic phase from side stream 7 and the mass flow rate of the organic phase from side stream 18.

[0110] 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.

[0111] The organic raffinate has the following composition: Water: 2.0% by weight n-butanol: 4.9% by weight n-Butyl acrylate: 23.8% by weight Acrylic acid: 9.1% by weight p-Toluenesulfonic acid <0.05 wt% Phenothiazine: 0.2% by weight Unknown: 59.9% by weight

[0112] The oligomeric and polymeric content is included in the unknown.

[0113] The aqueous extract is returned from the extractive phase separator I via conduit 12 to reactor A at a mass flow rate of 14 kg / h.

[0114] The aqueous extract has the following composition: Water: 71.2% by weight n-butanol: 0.7% by weight n-butyl acrylate: 0.9% by weight Acrylic acid: 9.9% by weight p-Toluenesulfonic acid 11.9% by weight Phenothiazine: 0.7% by weight Unknown: 4.7% by weight

[0115] Summarizing from the two examples, it can be stated that for the same catalyst concentration in reactor A, the following is evident: 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.

[0116] This is because in Example 1, the aqueous extract 12 is returned to reactor A at a mass flow rate of 14 kg / h and the catalyst concentration is 11.9 wt. %. In Example 1, the phase ratio between the aqueous phase 18 and the heavy boilers side stream 7 in the heavy boilers bottom output 23 resulting at the outlet of mixer H is 0.2 kg / kg, which causes phase separation, whereas in Comparative Example 1, there is no mixer (H) and the water concentration in the heavy boilers bottom output 23 is only 0.03 kg / kg. Therefore, there is no phase separation in Comparative Example 1.

[0117] As a result, the amount of catalyst in Example 1 of the present invention is saved by 66.7%.

[0118] Furthermore, the influence of the placement of the cleavage reactor J during the process was investigated experimentally.

[0119] 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 heavy boiler side stream 7 of the heavy boiler bottom output 23 from the bottom of the rectification column E was fed to the mixer H. The organic raffinate 11 taken from the extractive phase separator I and fed to the cleavage reactor J had a viscosity of 0.7 mPas. The side stream 18 from the liquid-liquid phase separator G containing the water extractant was used in the mixer H, with the mass flow ratio between the aqueous phase side stream 18 and the heavy boiler side stream 7 of the discharged heavy boiler bottom output 23 ranging from 0.13 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.

[0120] Example 3 In Example 3, the cleavage reactor J was connected upstream of the mixer H, and a side stream 8 (similar to FIG. 3) of the heavy boilers bottom output 23 from the bottom of the rectification column E was first fed to the cleavage reactor J. Then, the heavy boilers bottom output 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 withdrawn from the extractive phase separator I was 30 mPas. A side stream 18 from the liquid-liquid phase separator G containing a water extractant was fed to the mixer H at a mass flow ratio of 1 kg / kg between the side stream of the aqueous phase 18 and the heavy boilers bottom output 21 from the cleavage reactor J. 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]

[0121] 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, or 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 obtained reaction output to the rectification column E 7. A conduit for leading a side stream of the heavy-boiler bottom output 23 to the mixer H, or a side stream of the heavy-boiler bottom output 23 leading to the mixer H. 8. A conduit for conducting a side stream of the heavy-boiler bottom output 23 to the cleavage reactor J. or a side stream of the high boiler bottoms output 23 to the cleavage reactor J 9 Conduit for recycling a side stream of the heavy materials bottom output 23 to reactor A, or a side stream of the heavy materials bottom output 23 going to reactor A 10 Mixture obtained from mixer H 11 Conduit for feeding the organic raffinate from the extraction phase separator I to the cleavage reactor J or 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 A conduit for recycling the organic phase from the liquid-liquid phase separator G to the rectification column E, or a conduit for recycling the organic phase from the liquid-liquid phase separator G to the rectification column E 15 Conduit for discharging the organic phase side stream or crude product stream from the liquid-liquid phase separator G 16 A conduit for recycling the aqueous phase from the liquid-liquid phase separator G to the rectification column E, or a conduit for recycling the aqueous phase from the liquid-liquid phase separator G to the rectification column E 17 Conduit for discharging a substream of the aqueous phase from the liquid-liquid phase separator G 18 A conduit for feeding a substream of the aqueous phase from the liquid-liquid phase separator G to the mixer H, or a conduit for feeding a substream of the aqueous phase from the liquid-liquid phase separator G to the mixer H 21 conduit for discharging the bottoms output of high boilers from cleavage reactor J 22 Conduit for discharging high boiling point materials from cleavage reactor J 23 Conduit for the bottom output of heavy materials from rectification column E or the bottom output of heavy materials from rectification column E 24 Conduit for supplying polymerization inhibitor stream to condenser C 25 Conduit for feeding polymerization inhibitor stream to condenser F 26 A conduit for supplying the aqueous phase from the liquid-liquid phase separator D to the mixer H, or a conduit for supplying the aqueous phase from the liquid-liquid phase separator D to the mixer H Reactor A B tower 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

Claims

1. 1. 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, comprising the steps of: carrying out an esterification in a reactor (A) equipped with a column (B) at the top, wherein 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 wherein 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 resulting reaction output (6) and a vapor stream being obtained at the top of the column (B); - discharging said vapor stream at the top of said column (B); condensing said vapor stream in a condenser (C) to form an organic phase and an aqueous phase; - continuously separating the organic phase from the aqueous phase using a phase separator (D); feeding the reaction output (6) obtained into a rectification column (E); In the rectification column (E), the following azeotrope is formed: 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) is operated at a bottom temperature in the range of from 80 to 150°C and at a top temperature in the range of from 70 to 130°C and at an absolute pressure in the range of from 0.2 to 5 bar; - discharging a gas stream enriched in said azeotrope at the top of said rectification column (E); condensing said gas stream in a condenser (F) to form an organic phase rich in n-butyl (meth)acrylate and an aqueous phase; - continuously separating the organic phase from the aqueous phase using a phase separator (G); continuously removing at least a portion of the organic phase from the phase separator (G), the removed portion of the n-butyl (meth)acrylate-rich organic phase constituting a crude product stream (15); withdrawing a heavy boilers bottom output (23) from the bottom of the rectification column (E), wherein the mass flow ratio between the heavy boilers bottom output (23) and the (meth)acrylic acid fed as reactant to the reactor (A) is in the range of 0.5 to 5; feeding a heavy materials side stream (7) of the discharged heavy materials bottom output (23) to a mixer (H), the mass flow ratio between the heavy materials side stream (7) and the heavy materials bottom output (23) being 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 said mixer (H) into a downstream extraction phase separator (I); - continuously separating said mixture (10) in said extract phase separator (I) to obtain an organic raffinate (11) and a catalyst-containing aqueous extract (12), said aqueous extract (12) being at least partially recycled to said reactor (A) and / or said rectification column (E); Including, where: a side stream (18) of said aqueous phase from said phase separator (G), a side stream (26) of the aqueous phase from the phase separator (D), and / or The aqueous phase substream (5) from the phase separator (D) is fed to the phase separator (G), and then the aqueous phase substream (18) from this phase separator (G) is fed to the mass flow ratio between the minor stream (18) of the aqueous phase and the minor stream (7) of the discharged heavy materials bottom output (23) is in the range of 0.08 to 0.5, and the mass flow ratio between the minor stream (26) of the aqueous phase and the minor stream (7) of the discharged heavy materials bottom output (23) is in the range of 0.08 to 0.5, method.

2. 1. 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, comprising the steps of: carrying out an esterification in a reaction zone (E1), said reaction zone (E1) being located at the bottom of the rectification column (E), said (meth)acrylic acid and n-butanol components being 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, said esterification being 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 is formed 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) n-butanol, n-butyl (meth)acrylate and water; wherein said removal is also carried out by means of said rectification column (E), which is operated at a bottom temperature in the range from 80 to 150°C, a top temperature in the range from 70 to 130°C and an absolute pressure in the range from 0.2 to 5 bar, preferably in the range from 0.4 to 1.5 bar; - discharging a gas stream enriched in said azeotrope at the top of said rectification column (E); condensing said gas stream in a condenser (F) to form an organic phase rich in n-butyl (meth)acrylate and an aqueous phase; - continuously separating the organic phase from the aqueous phase using a phase separator (G); continuously removing at least a portion of the organic phase and the phase separator (G), the removed portion of the n-butyl (meth)acrylate-rich organic phase constituting a crude product stream (15); withdrawing a heavy boilers bottom output (23) from the bottom of the rectification column (E), wherein the mass flow ratio between the heavy boilers bottom output (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 heavy boiler side stream (7) of the discharged heavy boiler bottom output (23) to a mixer (H), the mass flow ratio between the heavy boiler side stream (7) and the heavy boiler bottom output (23) being in the range of 0.01 to 1.0, preferably in the range of 0.10 to 0.70; feeding the mixture (10) obtained from said mixer (H) into a downstream extraction phase separator (I); - continuously separating said mixture in said extractive phase separator (I) to obtain an organic raffinate (11) and a catalyst-containing aqueous extract (12), said aqueous extract (12) being at least partially recycled to said rectification column (E); Including, where: the aqueous phase sub-stream (18) from the phase separator (G) is fed to the mixer (H), and the mass flow ratio between the aqueous phase sub-stream (18) from the phase separator (G) and the high boilers sub-stream (7) of the discharged high boilers bottom output (23) is in the range of 0.08 to 0.5; method.

3. 3. The process 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. the high boilers substream (7) of the discharged high boilers bottom output (23) in the mixer (H) with the aqueous phase substream (18) from the phase separator (G); and / or the mass flow rate of the sub-stream (26) of the aqueous phase from the phase separator (D) is 5. The method according to claim 1, wherein the organic raffinate (11) is added so that the phase ratio between the aqueous extract to be obtained and the organic raffinate (12) to be 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 (11) of the organic raffinate is fed to a cleavage reactor (J), and the mass flow ratio between the side stream (11) of the organic raffinate 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 side stream (8) of the heavy boilers bottom output (23) is fed to the cleavage reactor (J) at a mass flow ratio of the heavy boilers bottom output (23) to the heavy boilers side stream (7) in the range of 0.0 to 10.0, preferably in the range of 0.1 to 1.

0.

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

9. 9. The process of any one of claims 1 to 8, wherein the heavy boilers bottoms output (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 boilers bottoms output (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 obtained reaction output of the reactor (A) in a concentration ranging from 0.1 wt.% to 10.0 wt.%.

14. 14. The method according to any one of claims 1 to 13, wherein external water is further added to the mixer (H).

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