Process for continuous production of tert-butyl (meth) acrylate

Abstract

Process for the continuous production of tert-butyl (meth) acrylate by reacting (meth) acrylic acid in a liquid phase with gaseous isobutylene passing through the liquid phase in a reactor in the presence of an acidic catalyst at a temperature in the range of 30 DEG C to 90 DEG C and a pressure in the range of 0.1 to 20 bar absolute, wherein the two reactants are used in a molar ratio of isobutylene: (meth) acrylic acid in the range of 5 to 60, partially condensing a gas stream comprising unconverted reactants and tert-butyl (meth) acrylate leaving the reactor to obtain tert-butyl (meth) acrylate and unconverted (meth) acrylic acid as a liquid mixture, the liquid mixture is separated by subsequent distillation, and wherein tert-butyl (meth) acrylate is obtained and uncondensed isobutylene is recycled to the reactor.

Description

[0001] The present invention relates to a process for the continuous production of tert-butyl (meth)acrylate by reacting (meth)acrylic acid in the liquid phase with gaseous isobutene passing through this liquid phase in the presence of an acidic catalyst at a temperature in the range from 30 °C to 90 °C and an absolute pressure in the range from 0.1 to 20 bar in a reactor.

[0002] Tert-butyl (meth)acrylate is to be understood as meaning tert-butyl acrylate (=acrylic acid tert-butylester) produced by the reaction of acrylic acid with isobutene, or tert-butylmethacrylate (=methacrylic acid tert-butyl ester) produced by the reaction of methacrylic acid (=α-methylacrylic acid) with isobutene.

[0003] The tert-butyl esters of acrylic acid and methacrylic acid [collectively referred to as (meth)acrylic acid] are used in a wide range of applications. Tert-butyl (meth)acrylate includes, for example, important starting materials for the production of polymers which are used in particular as components of coatings, adhesives and paint resins.

[0004] It is known that tert-butyl (meth)acrylate (hereinafter referred to as TB(M)A) can be formed by the addition reaction of acrylic acid (AA) or methacrylic acid to isobutene (IB) under the influence of an acid (such as sulfuric acid) as a catalyst. This process is described in particular in WO 2016 / 156410 A1 (BASF SE), WO 2002 / 10110 A2 (BASF AG) and WO 2002 / 10109 A1 (BASF AG).

[0005] The TBA process is described in more detail below; a similar description applies to the TBMA process. The reaction is an equilibrium reaction. The reaction is carried out continuously in a vertical reactor divided into a plurality of sections and intermediately cooled, in which substantial chemical equilibrium may be obtained at the upper outlet. In this TBA process, the reaction mixture discharged at the top thus contains acrylic acid, TBA, dissolved IB and the catalyst (such as sulfuric acid). This reaction mixture is then concentrated under vacuum (such as about 60 mbar), and evaporation removal of IB and TBA and AA is preferably achieved. The TBA and AA obtained in this gas phase are condensed and fed to a distillation work-up. The uncondensed IB is recycled to the reactor via a vacuum machine. The problem is that this may be accompanied by the formation and deposition of polymers of IB and / or TBA, which impairs the operation of the vacuum machine. The liquid high-boiling fraction remaining in the evaporation concentration is also specifically recycled to the reactor.

[0006] The disadvantage of this process mode is that it heats the liquid output from the reactor for evaporation and concentration, and removes the input material (IB) from the equilibrium mixture while concentrating the catalyst, thus causing a rapid partial reverse cleavage of TBA to form IB. Due to the low pressure, the amount of recycle gas mainly containing IB to be recycled to the reactor is very large and the vacuum unit is complex. Scaling up this equipment concept to substantially exceed a specific size (e.g., more than 5 kt / a) is difficult to impossible, or at least extremely expensive, especially from the perspective of process engineering. The above similarly applies to the corresponding TBMA process.

[0007] The object of the present invention is to provide an improved process for producing TB(M)A in high yield and purity, which overcomes the above disadvantages and is thus more easily scalable.

[0008] The inventors of the present invention have thus found a process for the continuous production of tert-butyl (meth)acrylate, which is carried out by reacting (meth)acrylic acid in the liquid phase with gaseous isobutene passing through the liquid phase in the presence of an acidic catalyst at a temperature in the range of 30 °C to 90 °C and an absolute pressure in the range of 0.1 to 20 bar. The process is characterized in that the two reactants are used in a molar ratio of isobutene:(meth)acrylic acid in the range of 5 to 60, a gas stream containing unreacted reactants and tert-butyl (meth)acrylate leaving the reactor is partially condensed to obtain a liquid mixture of tert-butyl (meth)acrylate and unreacted (meth)acrylic acid, the liquid mixture is separated by subsequent distillation, and tert-butyl (meth)acrylate is obtained and the uncondensed isobutene is recycled to the reactor.

[0009] Thus, when the reactant acrylic acid is used, the process provides the product tert-butyl acrylate (TBA), and when the reactant methacrylic acid is used, the process provides the product tert-butyl methacrylate (TBMA).

[0010] Thus, the process according to the invention also comprises carrying out the addition of acrylic acid (or methacrylic acid) to isobutene in a reactor, but with a significantly larger amount [based on the (meth)acrylic acid used] of gaseous isobutene passed through, which gaseous isobutene thus acts both as a reactant and as a stripping gas. The reactor output is thus achieved via a gas stream which contains isobutene, TB(M)A and acrylic acid / methacrylic acid, but does not contain the catalyst (such as in particular sulfuric acid). Thus, the reverse cracking of TBA or TBMA in the reactor output is prevented. The stripping gas is then subjected to partial condensation, where TB(M)A and acrylic acid / methacrylic acid are obtained in liquid form and are sent for further separation, for example analogous to the methods described in WO 2016 / 156410 A1 (BASF SE), WO 2002 / 10110 A2 (BASF SE) or WO 2002 / 10109 A1 (BASF SE). The uncondensed isobutene is preferably recycled mainly, in particular completely (for example to an extent of 75% to 99% by weight) to the reactor; it can also be referred to as recycle gas and / or stripping gas. Since this process is not a vacuum process as described in the prior art, the volume flow is significantly smaller and the pressure conditions are much less problematic than in the processes for the production of TB(M)A cited above and at the beginning, thus resulting in important scale-up possibilities, in particular by simplified scale-up.

[0011] The process according to the invention is carried out in a reactor, in particular a cylindrical reactor. More preferably, a reactor is used which is a stirred tank, a bubble column reactor or a loop reactor.

[0012] The isobutene is introduced into the reactor in gaseous form. The isobutene can also be used, for example, in the form of a hydrocarbon gas mixture containing isobutene. The gas mixture can in particular be a C4 gas mixture containing isobutene, isobutane, butane, 1-butene and 2-butene.

[0013] These two reactants are used in the reactor in an isobutene:(meth)acrylic acid molar ratio in the range from 5 to 60, preferably in the range from 8 to 50, more preferably in the range from 10 to 45. Thus, a characteristic feature of the process according to the invention is in particular the large molar excess of isobutene used.

[0014] The reaction of these reactants is preferably carried out in the absence of a solvent.

[0015] The acidic catalysts used are those that are at least partially soluble in the reaction mixture. Preferred catalysts are strong inorganic or organic acids, such as mineral acids, especially sulfur- or phosphorus-containing mineral acids, for example sulfuric acid, phosphoric acid and polyphosphoric acid, preferably sulfuric acid or alkylsulfonic and arylsulfonic acids, such as p-toluenesulfonic, benzenesulfonic, dodecylbenzenesulfonic and methanesulfonic acid. Very particularly preferred is sulfuric acid which also has catalytic activity in the form of the resulting monobutyl sulfate under the reaction conditions.

[0016] The amount of catalyst is preferably from 0.1% to 10% by weight, preferably from 0.5% to 3% by weight, in each case based on the weight of the two reactants [(meth)acrylic acid and isobutene].

[0017] The reaction is preferably carried out in the presence of an inhibitor which inhibits the polymerization of (meth)acrylic acid and the tert-butyl ester. Particularly suitable inhibitors are hydroquinone, hydroquinone monomethyl ether, p-benzoquinone, p-nitrosophenol, phenothiazine, 4-hydroxy-2,2,6,6-tetramethyl-1-oxyl-piperidine and methylene blue. The inhibitor is preferably used in an amount in the range from 200 to 2000 ppm, based on the weight of the two reactants [(meth)acrylic acid and isobutene].

[0018] The reactor can advantageously be equipped with internals to improve the co-mixing of the reaction mixture. Suitable internals are known to the person skilled in the art and include, for example, static mixing elements such as grids, distributor plates or sieve trays.

[0019] The reactant (meth)acrylic acid is very preferably introduced into the reactor in liquid form. This introduction can be carried out directly, for example via dip tubes, but it is preferred to use devices which allow a homogeneous distribution and co-mixing of the reactants. Such devices are known to the person skilled in the art and include, for example, distributor plates, perforated plates and pipes, nozzles etc. Gaseous isobutene is preferably introduced via an annular pipe having a plurality of outlet openings. (Meth)acrylic acid is preferably introduced via a nozzle which effects the mixing of the gas and the liquid and the co-mixing of the reactor contents. It is preferably arranged at the bottom or at the top of the reactor. Suitable nozzles are known to the person skilled in the art (spray nozzles, mixing nozzles, two-fluid nozzles etc.) and are described, for example, in Ullmann's Encyclopedia of Industrial Chemistry, volume B4, 5th edition, 1992, pages 280 ff.

[0020] It has proven advantageous to introduce the main part (in particular all) of the residue from the final distillation of fresh (meth)acrylic acid and tert-butyl esters into the reactor as a mixture. It has furthermore proven advantageous to recycle the main part (in particular all) of the reverse isobutene obtained in the distillation of tert-butyl esters back into the reactor. In the introduction of gaseous isobutene, it is particularly advantageous to introduce isobutene together with (meth)acrylic acid via the abovementioned nozzle. The nozzle causes the automatic suction of the recycled gaseous isobutene.

[0021] Part of the gas stream leaving the reactor and containing unreacted reactants and tert-butyl (meth)acrylate is condensed to obtain tert-butyl (meth)acrylate and unreacted (meth)acrylic acid as a liquid mixture.

[0022] The uncondensed isobutene is preferably recycled mainly, in particular completely, in gaseous form preferably via the abovementioned nozzle back into the reactor.

[0023] The catalyst is preferably introduced as a mixture with (meth)acrylic acid, and fresh catalyst or recycled catalyst or a mixture thereof can be used.

[0024] Since the addition of (meth)acrylic acid to isobutene is highly exothermic, it is advantageous to achieve temperature control of the reactor to regulate the reaction temperature. The temperature control of this reactor is preferably achieved via one or more internal heat exchangers or via one or more external heat exchangers or using one or more external liquid circuits.

[0025] The reaction is carried out at a temperature in the range of 30 °C to 90 °C and an absolute pressure in the range of 0.1 to 20 bar, preferably at a temperature in the range of 35 °C to 85 °C and an absolute pressure in the range of 0.2 to 15 bar, more preferably at a temperature in the range of 40 °C to 80 °C and an absolute pressure in the range of 0.5 to 13 bar.

[0026] The liquid mixture produced by partial condensation of the gas stream leaving the reactor contains a high proportion, for example ≥ 30% by weight (based on the mixture), of the desired tert-butyl ester. This liquid mixture additionally contains unreacted (meth)acrylic acid, inhibitors and other minor by-products. The mixture contains only very small amounts, in particular < 2% by weight (based on the mixture), of isobutene oligomerization products.

[0027] Figure 1 A schematic overview of a preferred configuration of the process according to the invention is shown.

[0028] It shows the interconnection of the units'reactor A', 'condenser B', 'condenser C', 'low-boiler distillation D', 'final distillation E' and'residue distillation F'.

[0029] For comparison, Figure 2The configuration of the TB(M)A method according to the prior art is shown in schematic form.

[0030] It shows the interconnection of the units'reactor A', 'condenser B', 'condenser C', 'low-boiler distillation D', 'final distillation E' and'residue distillation F'.

[0031] The following includes a preferred configuration of the method according to the invention, with particular reference by way of example to Figure 1 the description (in brackets).

[0032] Condensation (B; C):

[0033] The condensation of the vapour (4) can be carried out in a conventional manner, for example in condensers (B, C) of conventional design. It is preferred to use two condensers connected in series, in particular plate condensers or tube-bundle condensers, where the second condenser (C) operates at a lower cooling temperature. The temperature difference is typically about 30 °C to 50 °C, where the cooling temperature of the first condenser (B) is in the range of about 10 °C to 35 °C. This allows for rapid distillation and condensation and inhibits the formation of polymers.

[0034] The uncondensed vapour (5) is preferably mainly, preferably to an extent of 80% to 99.9% by weight, supplied to the reactor (A) (7), and the remaining part of the uncondensed vapour (6) is discharged. The uncondensed vapour (5) preferably contains ≥ 75% by weight of isobutene, ≤ 23% by weight of inert compounds and ≤ 2% by weight of the remainder.

[0035] To further reduce polymer formation, it is preferred to introduce an inhibitor soluble in the target ester [= TB(M)A] into the condenser, currently into the second condenser (C). The inhibitor used is preferably a mixture of phenothiazine (PTZ) and 4-hydroxy-2,2,6,6-tetramethyl-N-oxyl-piperidine (4-HT), which is advantageously introduced, for example, as a solution in the target ester, for example in an amount of 0.02 to 0.15 kg of 4-HT and 0.5 to 1.5 kg of PTZ in 100 l of the target ester, preferably in the top region of the condenser, in the example of the invention in the top region of the vertically arranged second condenser (C). The amount of the inhibitor solution is preferably such that the inhibitor concentration in the combined condensate (8) is, for example, in the range of 100 ppm to 500 ppm. The introduction of the inhibitor is carried out in a conventional manner, preferably by injecting the inhibitor solution. It has also proven advantageous, to reduce polymer formation, to introduce and in particular inject into the condenser a distillate in an amount of at least five times the weight part based on the weight of the vapour, currently into the first condenser (B), preferably into the top region of the vertically arranged condenser (B).

[0036] Low-boiler distillation (D)

[0037] In order to obtain the target ester from the combined condensate of condensates (B) and (C), the condensate (8) is separated into a top product (10) and a bottom product (11) in a distillation unit (D) (e.g., a conventional distillation unit including an evaporator, a column, and a condenser). The distillation temperature (bottom temperature) is typically in the range of 40 °C to 90 °C. The pressure is appropriately selected according to the product (i.e., TBA or TBMA), and the pressure is preferably the same for both target esters.

[0038] The top product (10) contains low-boiling components such as tert-butyl acetate, tert-butanol, and diisobutene. It may also contain up to 40% by weight of the target ester based on the top product. The bottom product essentially contains the target ester and (meth)acrylic acid.

[0039] The columns considered include conventional columns that include random packings or structured packings, or have bubble caps, valves, or sieve trays. However, it is preferred to use a plate column with 30 to 50 dual-flow trays. Feeding into the distillation column is typically achieved in the central region.

[0040] Condensation of the low-boiling components is achieved in a conventional manner. The condensation is preferably achieved in two condensers connected in series (e.g., shell-and-tube condensers). The cooling temperature of the second condenser is preferably about 30 °C to 50 °C lower, and the first condenser operates at a cooling temperature of preferably about 10 °C to 35 °C. The condensates are combined and partially used as column reflux. The remaining part (10) of the condensate is discharged. The uncondensed vapor (9) is preferably mainly, especially completely, recycled to the reactor (A).

[0041] The uncondensed vapor (9) preferably contains ≥85% by weight of isobutene, ≤2% by weight of inert compounds, and ≤13% by weight of the remainder.

[0042] In order to prevent the formation of polymers in the first condenser and the column, it is preferred to introduce a solution of an inhibitor in the target ester into the first condenser. The inhibitor used is preferably a mixture of PTZ and 4-HT, which is advantageously introduced as a solution in the target ester, for example, in an amount of 0.02 to 0.15 kg of 4-HT and 0.5 to 1.5 kg of PTZ in 100 l of the target ester, preferably in the top region of the vertically arranged second condenser. The amount of the inhibitor solution is such that the inhibitor concentration in the combined condensate is, for example, in the range of 100 ppm to 500 ppm.

[0043] Final distillation (E)

[0044] The target ester is obtained with a purity of preferably at least 99.5% by weight from the bottom product (11) of the low-boiling distillation (D) in a distillation unit (evaporator, column and condenser) having a conventional design. The resulting bottom product (13) usually contains at least 70% by weight of (meth)acrylic acid and is preferably recycled mainly, in particular completely, to the reactor (A). The distillation temperature is usually in the range of 50 °C to 100 °C. The pressure is selected according to the target ester to be distilled.

[0045] The final distillation preferably uses a conventional plate column, for example a column with 30 to 50 dual-flow trays and a feed in the central column region. The pure target ester is separated at the top of the column. The condensation of the target ester is preferably carried out in two condensers arranged in series, in particular in a tube bundle condenser. The temperature of the coolant of the second condenser is preferably about 30 °C to 50 °C lower than the temperature of the first condenser, where the coolant preferably has a temperature in the range of about 10 °C to 35 °C. The combined condensate is partly used as column reflux and partly for stabilizing the top of the column and the condenser (currently the first condenser), i.e. for avoiding polymerization in the top of the column and the condenser (currently the first condenser). The other part of the target ester is obtained as a valuable product (12). To avoid polymerization in the condenser, currently especially in the second condenser, a solution of an inhibitor in the target ester is introduced. Preferably, a solution of, for example, 0.5% to 2% by weight of hydroquinone monoethyl ether (MEHQ) is used, where the amount of the inhibitor introduced is preferably selected such that the inhibitor content of the combined condensate is 10 to 20 ppm. To avoid polymerization in the first condenser, a part of the combined stabilized condensate (preferably about 5 to 10 times the weight of the discharged target ester based on the weight of the vapor) is introduced into the vapor line. The introduction of the condensate is preferably achieved by injecting counter-current to the gas flow into the vapor line above the column and / or by injecting co-current to the gas flow into the region of the condenser inlet. Polymerization in the column is first inhibited by the reflux of the condensate containing, for example, 10 ppm to 20 ppm of the inhibitor. In addition, it is preferably to introduce a solution of a further inhibitor in the target ester onto the trays in the upper column region. Preferably, a solution of PTZ and 4-HT in the target ester is used, for example 0.5 to 1.5 kg of PTZ and 0.05 to 0.15 kg of 4-HT in 100 l of the target ester, where the amount is preferably measured such that the inhibitor content in the rectifying section is 50 ppm to 500 ppm. Further preferably, a solution of MEHQ in the target ester (for example 15 to 30 g of MEHQ in 100 l of the target ester) is introduced into the column cap or into the vapor line above the column and / or into the upper condenser cap.

[0046] This is suitably and advantageously achieved by injection via a nozzle centrally mounted in the upper shroud or vapor duct or upper condenser shroud. For further stabilization, if desired, an oxygen-containing gas (such as air) can be blown into the distillation unit. These measures make it possible to prevent the formation of polymers in the condenser, vapor duct, and column.

[0047] The target ester (12) obtained has a high purity and generally has the following composition:

[0048]

[0049] The liquid reaction mixture (14) from the reactor (A) is preferably separated in the distillation (F) as a withdrawn sub-stream. The distillation is preferably carried out continuously at the same pressure as the reactor pressure. The temperature depends on the desired product (i.e., TBA or TBMA) and is generally chosen such that the target ester undergoes reverse cracking and only a small proportion of the target ester, for example less than 5% by weight of the target ester (based on the amount of bottoms), remains in the bottoms product. The resulting gaseous isobutene is preferably mainly, in particular completely, recycled to the reactor (15). The resulting bottoms product (16) essentially contains the acidic catalyst, the remaining (i.e., unreacted) (meth)acrylic acid, and high-boiling components [i.e., by-products having a boiling point higher than TB(M)A at the same pressure], in particular polymeric (meth)acrylic acid compounds.

[0050] The bottoms product (16) from the residue distillation, for example 10% to 90% by weight of the bottoms product (16), in particular the major part of the bottoms product (16), for example 75% to 90% by weight of the bottoms product (16), is fed to the reactor (A) (18), while the remaining part of the bottoms product (17) is discharged.

[0051] The distillation can be carried out in conventional equipment. However, it is preferred to use equipment that allows rapid distillation, such as a wiped-film evaporator, thin-film evaporator, or coil evaporator. Suitable wiped-film evaporators are known to those skilled in the art, see for example Ullmanns Encyclopedia of Industrial Chemistry, 5th Edition, Volume B3, pages 2-21 to 2-24 and 3-1 to 3-25, 1988.

[0052] The condensation of the vapors can be carried out in a conventional manner, for example in a condenser of conventional design. Preferably, two condensers connected in series are used, in particular plate condensers or tube bundle condensers, where the second condenser is preferably operated at a lower cooling temperature. The temperature difference is typically about 30 °C to 50 °C, where the cooling temperature of the first condenser is preferably in the range of about 10 °C to 35 °C. This allows for rapid distillation and condensation and inhibits polymer formation. To even further reduce polymer formation, an inhibitor dissolved in the target ester is introduced into the condenser (currently the second condenser). The inhibitor used is preferably a mixture of PTZ and 4-HT, which is advantageously introduced as a solution, for example in an amount of 0.02 to 0.15 kg of 4-HT and 0.5 to 1.5 kg of PTZ in 100 l of the target ester, preferably in the top region of the condenser, in the top region of the second condenser arranged vertically in the example of the present invention. The amount of the inhibitor solution is such that the inhibitor concentration in the combined condensate is, for example, in the range of 100 ppm to 500 ppm.

[0053] The introduction of the inhibitor is carried out in a conventional manner, preferably by injecting the inhibitor solution. Additionally, it has proven advantageous to introduce and in particular inject at least five times the weight parts of the distillate (crude ester) based on the weight parts of the vapors into the condenser, currently into the first condenser, preferably into the top region of the condenser arranged vertically, in order to reduce polymer formation.

[0054] In a particularly preferred embodiment, condensation can be avoided and the resulting vapors are directly recycled to the reactor (A) (15). Then, the pressure in the residue distillation (F) is equal to the pressure in the reactor.

[0055] All reported pressures refer to absolute pressure.

[0056] All reported ppm values refer to weight (ppmw).

[0057] Examples and Comparative Examples

[0058] Examples of the following processes are simulated by thermodynamic simulation. These use Aspen software (Aspen), which can be found on the website https: / / www.aspentech.com. Aspen is a widely used simulation software package that is used for modeling, simulating, and optimizing chemical processes and equipment in industry. Aspen has a widely used model database for modeling basic operations and a substance database of the physical properties of many different substances. The properties of mixtures are calculated by Aspen from the physical data of pure substances with the aid of different thermodynamic models.

[0059] Comparative Examples

[0060] According to Figure 2 a thermodynamic simulation of the entire plant was carried out by Aspen and gave the following results:

[0061] Into reactor A, isobutene with a mass flow rate of 479 kg / h is supplied via conduit 1, sulfuric acid as a catalyst with a mass flow rate of 2 kg / h is supplied via conduit 2, and acrylic acid with a mass flow rate of 605 kg / h is supplied via conduit 3.

[0062] Into reactor A, a vapor recycle stream from the condensation stage C with a mass flow rate of 343 kg / h is supplied via conduit 7, a vapor recycle stream from the low-boiler distillation D with a mass flow rate of 21 kg / h is supplied via conduit 9, a liquid recycle stream from the final distillation E with a mass flow rate of 320 kg / h is supplied via conduit 13, and a liquid recycle stream from the residue distillation F with a mass flow rate of 2256 kg / h is supplied via conduit 18.

[0063] The reaction in reactor A takes place at a temperature of 16 °C - 32 °C, an absolute pressure of 1200 mbar, and a residence time of 4 hours.

[0064] From reactor A, low boilers and non-condensable components are discharged from the process as a vapor phase via conduit 6 with a mass flow rate of 9 kg / h.

[0065] The vapor phase has the following composition:

[0066]

[0067] From reactor A, the liquid phase is transferred to the residue distillation F via conduit 19 with a mass flow rate of 4016 kg / h. The liquid phase has the following composition:

[0068]

[0069] From condensations B and C, the liquid phase is transferred to the low-boiler distillation D via conduit 8 with a mass flow rate of 1378 kg / h. The liquid phase has the following composition:

[0070]

[0071] From condensation C, the vapor phase is transferred to reactor A via conduit 7 with a mass flow rate of 343 kg / h.

[0072] The vapor phase has the following composition:

[0073]

[0074] From the low-boiler distillation D, the liquid phase is transferred from the bottom of the column to the high-purity distillation E via conduit 11 with a mass flow rate of 1351 kg / h.

[0075] The liquid phase has the following composition:

[0076]

[0077] From the low-boiler distillation D downstream of the condenser, the low-boiler is discharged from the process as a liquid phase via conduit 10 at a mass flow rate of 19 kg / h.

[0078] The liquid phase has the following composition:

[0079]

[0080] From the low-boiler distillation D downstream of the condenser, the vapor phase is transferred to reactor A via conduit 9 at a mass flow rate of 21 kg / h.

[0081] The vapor phase has the following composition:

[0082]

[0083]

[0084] From the high-purity distillation E, the pure product tert-butyl acrylate is discharged from the process as a liquid phase via conduit 12 at a mass flow rate of 1000 kg / h.

[0085] The liquid phase has the following composition:

[0086]

[0087] From the final distillation E, the liquid phase is recycled from the bottom of the column to reactor A via conduit 13 at a mass flow rate of 320 kg / h.

[0088] The liquid phase has the following composition:

[0089]

[0090] From the residue distillation F, the liquid phase is recycled to reactor A via conduit 18 at a mass flow rate of 2256 kg / h, and the high-boiler is discharged from the process via conduit 17 at a mass flow rate of 63 kg / h.

[0091] The liquid phase has the following composition:

[0092]

[0093]

[0094] Example 1

[0095] According to Figure 1 The thermodynamic simulation of the entire plant was carried out via Aspen and gave the following results:

[0096] Into reactor A, isobutene with a mass flow rate of 455 kg / h is supplied via conduit 1, sulfuric acid as a catalyst with a mass flow rate of 7 kg / h is supplied via conduit 2, and acrylic acid with a mass flow rate of 598 kg / h is supplied via conduit 3.

[0097] Into reactor A, a vapor recycle stream from the condensation stage C with a mass flow rate of 31521 kg / h is supplied via conduit 7, a vapor recycle stream from the low-boiling distillation D with a mass flow rate of 561 kg / h is supplied via conduit 9, a liquid recycle stream from the final distillation E with a mass flow rate of 481 kg / h is supplied via conduit 13, a vapor recycle stream from the residue distillation F with a mass flow rate of 5 kg / h is supplied via conduit 15, and a liquid recycle stream from the residue distillation F with a mass flow rate of 11 kg / h is supplied via conduit 18.

[0098] The reaction in reactor A is carried out at a temperature of 50 °C, an absolute pressure of 1000 mbar, and a residence time of 2 hours.

[0099] From reactor A, the vapor phase is transferred via conduit 4 to the condensation B at a mass flow rate of 33580 kg / h.

[0100] The vapor phase has the following composition:

[0101]

[0102] From reactor A, the liquid phase is transferred via conduit 14 to the residue distillation F at a mass flow rate of 60 kg / h.

[0103] The liquid phase has the following composition:

[0104]

[0105]

[0106] From the condensations B and C, the liquid phase is transferred via conduit 8 to the low-boiling distillation D at a mass flow rate of 2054 kg / h.

[0107] The liquid phase has the following composition:

[0108]

[0109] From the condensation C, the vapor phase is transferred via conduit 7 to reactor A at a mass flow rate of 31521 kg / h, and the low-boiling components are discharged from the process via conduit 6 at a mass flow rate of 5 kg / h.

[0110] The vapor phase has the following composition:

[0111]

[0112] From the low-boiling distillation D, the liquid phase is transferred from the bottom of the column to the final distillation E via conduit 11 at a mass flow rate of 1481 kg / h.

[0113] The liquid phase has the following composition:

[0114]

[0115]

[0116] From the low-boiling distillation D downstream of the condenser, the low-boiling components are discharged from the process as a liquid phase via conduit 10 at a mass flow rate of 12 kg / h.

[0117] The liquid phase has the following composition:

[0118]

[0119] From the low-boiling distillation D downstream of the condenser, the vapor phase is transferred to reactor A via conduit 9 at a mass flow rate of 561 kg / h.

[0120] The vapor phase has the following composition:

[0121]

[0122] From the high-purity distillation E, the pure product tert-butyl acrylate is discharged from the process as a liquid phase via conduit 12 at a mass flow rate of 1000 kg / h.

[0123] The liquid phase has the following composition:

[0124]

[0125]

[0126] From the final distillation E, the liquid phase is recycled from the bottom of the column to reactor A via conduit 13 at a mass flow rate of 481 kg / h.

[0127] The liquid phase has the following composition:

[0128]

[0129] From the residue distillation F, the liquid phase is recycled to reactor A via conduit 18 at a mass flow rate of 11 kg / h, and the high-boiling components are discharged from the process via conduit 17 at a mass flow rate of 44 kg / h.

[0130] The liquid phase has the following composition:

[0131]

[0132] From the residue distillation F, the vapor phase is transferred via conduit 15 to reactor A at a mass flow rate of 5 kg / h.

[0133] The vapor phase has the following composition:

[0134]

[0135]

[0136] In the example of the present invention, the low-boiler ratio output is 0.012 kg of low boilers / kg of tert-butyl acrylate and the high-boiler ratio output is 0.044 kg / kg of tert-butyl acrylate.

[0137] In the comparative example, the low-boiler ratio output is 0.019 kg of low boilers / kg of tert-butyl acrylate, and the high-boiler ratio output is 0.063 kg / kg of tert-butyl acrylate.

[0138] Compared with the comparative example, in the example of the present invention, the low-boiler output is 37% lower and the high-boiler output is 30% lower.

[0139] Due to the lower ratio outputs of low boilers and high boilers, the process of the present invention is significantly more economical than the conventional process.

[0140] The concentration of the polymerizable component tert-butyl acrylate present in the recycle gas (conduit 5) is 3.72% by weight in the comparative example and only 0.18% by weight in the example of the present invention. This significantly reduces the risk of polymerization in the recycle gas blower warmed at the operating temperature.

Claims

1. A process for the continuous production of tert-butyl (meth)acrylate, which process is carried out by reacting (meth)acrylic acid in the liquid phase with gaseous isobutene passing through this liquid phase in the presence of an acidic catalyst at a temperature in the range from 30 °C to 90 °C and an absolute pressure in the range from 0.1 to 20 bar, wherein the two reactants are used in an isobutene:(meth)acrylic acid molar ratio in the range from 5 to 60, a gas stream leaving the reactor and containing unreacted reactants and tert-butyl (meth)acrylate is partially condensed to obtain tert-butyl (meth)acrylate and unreacted (meth)acrylic acid as a liquid mixture, this liquid mixture is separated by subsequent distillation, and wherein the tert-butyl (meth)acrylate is obtained and the uncondensed isobutene is recycled to the reactor.

2. The method according to the preceding claim, wherein, The reactor is a stirred tank, a bubble column reactor or a loop reactor.

3. The method according to any one of the preceding claims, wherein, These reactants are used in an isobutene:(meth)acrylic acid molar ratio in the range from 8 to 50.

4. The method according to any one of the preceding claims 1 to 2, wherein These reactants are used in an isobutene:(meth)acrylic acid molar ratio in the range from 10 to 45.

5. The method according to any one of the preceding claims, wherein, The reaction is carried out at a temperature in the range from 35 °C to 85 °C and an absolute pressure in the range from 0.2 to 15 bar.

6. The method according to any one of the preceding claims 1 to 4, wherein The reaction is carried out at a temperature in the range from 40 °C to 80 °C and an absolute pressure in the range from 0.5 to 13 bar.

7. The method according to any one of the preceding claims, wherein, The (meth)acrylic acid obtained from the liquid mixture by distillation separation after partial condensation of the gas stream is recycled to the reactor.

8. The method according to any one of the preceding claims, wherein The temperature control of the reactor is achieved via one or more internal heat exchangers or via one or more external heat exchangers or using one or more external liquid circuits.

9. The method according to any one of the preceding claims, wherein, The catalyst is a sulfur- or phosphorus-containing mineral acid, an alkylsulfonic acid or an arylsulfonic acid.

10. The method according to the preceding claim, wherein, The catalyst is sulfuric acid.

11. The method according to any one of the preceding claims, wherein, The isobutene is used in the form of a hydrocarbon gas mixture containing isobutene.

12. The method according to the preceding claim, wherein, The gas mixture is a C4 gas mixture containing isobutene, isobutane, butane, 1-butene and 2-butene.

13. The method according to any one of the preceding claims, wherein High-boiling by-products are discharged from the liquid reaction mixture in the reactor via the bottom product of the distillation of an off-stream of the liquid reaction mixture.

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