Method for separating (METH)acrylic acid

Abstract

The invention relates to a method for separating (meth)acrylic acid from a gas mixture which is present as a raw product of a catalytic gas phase oxidation of at least one C3- / C4 precursor of (meth)acrylic acid, wherein the gas mixture is depleted of high-boiling components, the gas mixture depleted of high-boiling components is condensed in a separating column having installations which have a separating effect, and a liquid mixture containing the (meth)acrylic acid is guided out of the separating column, the liquid mixture being partially evaporated and the non-evaporated liquid mixture being fed, at least partially directly to an esterification system, without further purification, in order to convert the (meth)acrylic acid into (meth)acrylic acid alkyl ester.

Description

[0001] 250427W001

[0002] 1

[0003] Methods for separating (meth)acrylic acid

[0004] The present invention relates to a method and a system for separating (meth)acrylic acid.

[0005] The term "(Meth)acryl-" is used in this document as a shortened form of "acryl- or methacryl-".

[0006] (Meth)acrylic acid is generally known. One route for the large-scale production of (meth)acrylic acid is the heterogeneously catalyzed partial gas-phase oxidation of suitable C3 / C4 precursors (e.g., propylene, acrolein, isobutene, or methacrolein) with molecular oxygen (see, e.g., WO 2010 / 012586 A, US 5,198,578 A, EP 1 710 227 A, EP 1 015410 A, EP 1 484303 A, EP 1 484308 A, EP 1 484309 A, US 2004 / 0242826 A, WO 2006 / 136336 A, DE 10028582 A, and WO 2007 / 074044 A). However, this process does not yield pure (meth)acrylic acid, but a gas mixture which, in addition to (meth)acrylic acid, contains acetic acid as a by-product, the separation of which from (meth)acrylic acid, especially by rectification, is complex (see, e.g., DE 19814449 A and DE 19814421 A).

[0007] To separate (meth)acrylic acid from the product gas mixture of the heterogeneously catalyzed partial gas-phase oxidation of a C3 / C4 precursor compound, a combination of different separation methods is typically employed to achieve a level of purity suitable for the subsequent application of the (meth)acrylic acid in the most economical way possible. The specific combination used depends on the type and quantity of components other than (meth)acrylic acid contained in the product gas mixture, as well as on the subsequent application. The purity requirements for (meth)acrylic acid used in esterification and subsequent polymerization of the ester are generally lower than for the direct use of this (meth)acrylic acid in a polymerization reaction, e.g., for use as a superabsorbent.

[0008] A feature common to essentially all possible combinations of separation processes for separating (meth)acrylic acid from the product gas mixture of a heterogeneously catalyzed partial gas phase oxidation of a C3 / C4 precursor compound is that, optionally after direct and / or indirect cooling of the aforementioned product gas mixture, the (meth)acrylic acid contained in the product gas mixture is transferred to the condensed, in particular liquid, phase in a basic separation step.

[0009] The existing processes for the production of acrylic acid are presented in detail in PEP Review 2016-10 “Acrylic Acid Process Summary”, I HS Chemical. All four processes presented there have in common that the condensed acrylic acid is drawn off from the bottom of the absorber, and that this condensed acrylic acid also contains all the high-boiling components formed in the reaction. Therefore, the condensed, acrylic acid-containing mixture must undergo complex rectification in all four processes, even if only the low purity requirements of the acrylic acid for subsequent esterification need to be met. The mandatory rectification

[0010] 2

[0011] This leads to low plant availability and high energy consumption in rectification due to the frequently necessary flushing shutdowns (WO 2017 / 025391 A).

[0012] WO 1999 / 014181 A discloses a process for obtaining high-purity acrylic acid by suspension crystallization and subsequent separation of the crystals from the mother liquor. Since the purification of the acrylic acid-containing liquid mixture is carried out crystallizatively at low temperatures, high plant availability is achieved. However, the high investment costs and electricity consumption are disadvantages. The process only enables the production of high-purity acrylic acid; the purity requirements for using acrylic acid in the production of esters are exceeded, resulting in increased costs.

[0013] The object of the present invention was to provide an improved process for the separation of (meth)acrylic acid. In particular, the process was intended to have low investment requirements, high availability, and low raw material and energy consumption, as well as high plant availability and suitable product quality of the crude (meth)acrylic acid.

[0014] The problem is solved by a process for separating (meth)acrylic acid from a gas mixture that is present as a crude product of a catalytic gas-phase oxidation of at least one C3 / C4 precursor of (meth)acrylic acid, wherein

[0015] a) the gas mixture is depleted of heavy elements,

[0016] b) the gas mixture depleted of heavy elements is condensed in a separation column with separation-effective internals, and

[0017] c) a liquid mixture containing (meth)acrylic acid is removed from the separation column, characterized in that the liquid mixture is partially evaporated and the non-evaporated liquid mixture is at least partially fed directly to an esterification plant for the conversion of the (meth)acrylic acid to (meth)acrylic acid alkyl esters without further purification.

[0018] According to the invention, feeding the liquid mixture "directly into the esterification plant without further purification" means that no further thermal separation process is used to purify the (meth)acrylic acid between the separation column and the esterification plant. The term "thermal separation process" is intended to express that heat must be added to or removed from the system to achieve the desired separation effect (see DE-A 102008041573 and DE-A 1020088040799). Examples of thermal separation processes are desorptive, distillative, stripping, rectifying, azeotropic distillative, azeotropic rectifying, and crystallizing processes.

[0019] Crystallizative separation processes can be carried out in the manner described in documents DE-A 102008041573, DE-A 102008040799 and WO 2007 / 074044 A as well as DE-A 102007029053.250427W001

[0020] 3

[0021] Non-crystallizing separation processes are generally thermal separation processes in which gaseous (ascending) and liquid (descending) streams, or two liquid streams, are passed countercurrently through separation columns containing separation-effective internals. Due to the gradients between the streams, heat and mass exchange occurs, ultimately resulting in the desired separation effect in the column. Examples of such non-crystallizing thermal separation processes include rectification, azeotropic rectification, desorption, stripping, distillation, and azeotropic distillation. Such thermal separation processes (e.g.,All thermal separation processes described in WO 2011 / 000808 A2, DE-A 10 10336386, DE-A 19924532, DE-A 19924533, and DE-A 102007004960 are preferably carried out according to the invention in devices that comply with the recommendations of US 6441228 B2 and US 6966973 B2.

[0022] According to the invention, a high-boiling component or compound is understood to have a higher boiling point than (meth)acrylic acid. According to the invention, a low-boiling component or compound is understood to have a lower boiling point than (meth)acrylic acid.

[0023] Separation columns and separation-effective internals as such are known to those skilled in the art and can be found in textbooks on process engineering, in particular on thermal separation (for example, M. Baerns et al., Technische Chemie, 2006, WILEY-VCH, Weinheim).

[0024] In the context of the present invention, a separation column is preferably understood to be an apparatus for the thermal separation of mixtures. Examples include rectification columns, absorption columns, and condensation columns, with condensation columns being preferred. The column body is typically designed as a cylindrical tube insulated to prevent heat loss. For the evaporation of the mixture to be separated, or a portion thereof, the lower end of the column can include an evaporator, or an evaporator can be implemented as an external apparatus adjacent to the column. Almost any conceivable heat source for evaporation can be used, such as hot water, electricity, microwaves, or waste heat from other processes or other process steps of the present process.For the condensation of the mixture to be separated, or a portion thereof, the column can contain a condenser at the top, or the condenser can be implemented as an external unit next to the column. Various operating media can be used for cooling in the condenser, such as water, refrigerants, air, or even media from other processes that need to be heated.

[0025] or other parts of the present process. To improve heat and mass transfer in the column, the column body can contain internals such as screen, bubble-cap, or valve trays, packing materials such as Raschig rings, or structured packings. Those skilled in the art know that (horizontal) separation-effective internals in such columns are also generally referred to as trays. In principle, columns with rotating inserts, so-called rotary columns, can also be used, which spray the reflux liquid in droplets. 250427W001

[0026] 4

[0027] In the context of the present invention, dual-flow trays and cross-flow trays are preferably used as separating internals. Dual-flow trays as such are known to those skilled in the art. In the context of the present invention, these are horizontal internals that are installed in the column at specific intervals and have openings through which vapor and liquids can pass in counterflow. These openings can be, for example, holes or slots, the opening ratio being adjustable via the number of openings. Typically, dual-flow trays do not have a drain pipe connecting them to the next tray. Naturally, each dual-flow tray can be flush with the walls of the rectification column, but can also be connected to them via webs.

[0028] Crossflow trays are also known to those skilled in the art. In the context of the present invention, these are horizontal internals installed at specific intervals in the column, in which liquid is fed onto the tray at one side, flows over the tray, and comes into contact with the rising gas. The liquid then flows to a discharge chute, is collected there, and fed to the tray below. Crossflow trays can, for example, contain sieve holes, movable valves, bells, or tunnels.

[0029] Preferably, dual-flow trays and / or cross-flow trays are used as separation-effective internals in the separation column according to the invention. Furthermore, the separation column can include additional inlets and / or outlets, for example at the top, to, for example, withdraw lighter components and / or return components.

[0030] The operating pressure prevailing in the T separation column 100 is preferably 0 to 5 bar, more preferably 0 to 3 bar and particularly preferably 0 to 1.6 bar.

[0031] The liquid mixture removed from the separation column in step c) is commonly referred to as crude acrylic acid. Crude acrylic acid streams, obtained during the purification of acrylic acid, are well known to those skilled in the art. It is also known to those skilled in the art that such a crude acrylic acid stream typically does not represent a pure product, but rather contains acrylic acid as its largest component, as well as (small) amounts of other components, such as water, lower aldehydes (e.g., furfurals, acrolein, benzaldehyde), lower carboxylic acids (e.g., acetic acid, propionic acid), and diacrylic acid.

[0032] Furthermore, the process of heterogeneously catalyzed partial gas-phase oxidation for the production of acrylic acid can be carried out as described in the prior art, as is known per se.

[0033] be performed.

[0034] If the calcium precursor compound is, for example, propylene and / or acrolein, then…

[0035] The heterogeneously catalyzed partial gas-phase oxidation can be carried out, for example, as described in WO 2005 / 042459 A, WO 2005 / 047224 A and WO 2005 / 047226 A. 250427W001

[0036] 5

[0037] If the Ca precursor compound is, for example, propane, the heterogeneously catalyzed partial gas-phase oxidation for the production of acrylic acid can be carried out, for example, as described in the documents EP-A 608838, DE-A 19835247, DE-A 10245585 and DE-A 102 46 119.

[0038] If the calcium precursor compound is, for example, glycerol, the heterogeneously catalyzed partial gas-phase oxidation for the production of acrylic acid can be carried out, for example, as described in WO 2007 / 090991.

[0039] A, WO 2006 / 114506 A, WO 2005 / 073160 A, WO 2006 / 114506 A, WO 2006 / 092272 A or WO 2005 / 073160 A are described.

[0040] The esterification of (meth)acrylic acid as such is known to those skilled in the art and can be found in various forms in the prior art. For example, the acid-catalyzed esterification of acrylic acid with n-butanol to n-butyl acrylate is described in EP 4015498 A1 or EP 1 129061 A1.

[0041] The continuous production of n-butyl acrylate is carried out by reacting acrylic acid with n-butanol in a solvent-free phase at elevated temperature and with the addition of acid, particularly sulfuric acid, as an acidic esterification catalyst. The acrylic acid, n-butanol, and esterification catalyst are fed into a reaction zone. During the residence time in the reaction zone, the water formed, as a component of an n-butanol-containing mixture, is separated from the reaction mixture in a first rectification unit placed on top of the reaction zone. The resulting distillate is separated into an organic phase containing n-butanol and an aqueous phase containing water, and the organic phase is recycled to the first rectification unit. The aqueous phase is optionally removed entirely or partially (e.g., a proportion of 50 to 100 wt.%, particularly 80 to 100 wt.%, most especially 95 to 100 wt.%, and in particular, entirely, i.e., 100 wt.%).-%), is fed to a stripping unit. The reaction mixture containing n-butyl acrylate, which is drawn off from the reaction zone, is subjected to pre-purification. In a first pre-purification stage, the majority of the esterification catalyst is extracted by water washing. In a second pre-purification stage, the acidic components are neutralized and extracted with an aqueous alkali solution by reactive extraction. Optionally, in a third pre-purification stage, residual salts and aqueous foreign phase components can be extracted with water from the organic reaction residue mixture remaining after the second pre-purification stage. The remaining organic first reaction residue mixture is fed into a separation zone comprising a further rectification unit, and the n-butyl acrylate formed is separated in this zone by feeding the remaining first reaction residue mixture to a second rectification unit.In this process, the remaining first reaction mixture is rectified into a low-boiling product containing n-butyl acrylate and components with a higher boiling point than n-butyl acrylate, and into a second reaction mixture comprising n-butyl acrylate and components with a higher boiling point than n-butyl acrylate. The low-boiling product is preferably fed into a further rectification unit, in which the low-boiling components are distilled off and the bottoms product is recycled (see EP 4015498 A1). The second reaction mixture is fed into a third rectification unit, in which the n-butyl acrylate is separated from the components with a higher boiling point than n-butyl acrylate. 250427W001.

[0042] 6

[0043] It was surprisingly found that the inventive process with partial evaporation of the liquid mixture or crude acrylic acid stream discharged from the separation column has the advantage over the prior art that energy and raw material efficiency is improved with excellent plant availability and, on the other hand, the quality of the acrylic acid obtained is so high that the alcohol input in the subsequent esterification is acceptable and the smooth operation of the esterification is ensured.

[0044] A crude acrylic acid stream of sufficient purity for esterification can only be achieved if a removal of high-boiling components is carried out beforehand in the separation column. Without this step, operation of the esterification process would not be possible due to polymer deposits in the heat exchanger (heat exchanger fouling). Furthermore, as described in EP 0 169254 A1, problems with inhibition and discoloration would occur when using n-butyl acrylate for the production of polymer dispersions.

[0045] The method and the apparatus according to the invention are shown schematically in Figure 1 in a preferred embodiment. For the sake of clarity, non-essential details that are well known to those skilled in the art have been omitted.

[0046] 100 separation column

[0047] 101 Gas mixture in separation column

[0048] 102 Column floor

[0049] 103 Reflux fluid flow

[0050] 104 Crude acrylic acid stream (liquid mixture) from separation column

[0051] 105 Partial Evaporation

[0052] 106 Higher purity acrylic acid stream

[0053] 107 evaporators

[0054] 108 Capacitor

[0055] 109 Esterification plant

[0056] An acrylic acid-containing gas mixture 101, e.g., from a reactor for the heterogeneously catalyzed partial gas-phase oxidation of propene, is introduced into the separation column 100. In the separation column 100, which has separation-effective internals 102, e.g., column trays, the acrylic acid is condensed or purified. The quality of the crude acrylic acid stream 104 (liquid mixture) is ensured by regulating the reflux liquid flow 103. The crude acrylic acid stream 104, which still contains impurities and is discharged from the separation column, is fed to the partial evaporation stage 105.

[0057] Higher purity acrylic acid 106 is obtained by supplying sufficient energy to the evaporator 107. The unevaporated liquid or the unevaporated portion of the liquid mixture 106 is fed directly, without further purification, to an esterification plant 109 for the conversion of the acrylic acid to alkyl acrylates. 250427W001

[0058] 7

[0059] The energy input to evaporator 107 depends on the objective of the partial evaporation. If only acrolein and similar low-boiling substances are to be removed from the crude acrylic acid to ensure smooth operation of the esterification process, 0.015 kWh / kg crude acrylic acid (liquid mixture) is sufficient. However, if a significant proportion of acetic acid is also to be removed from the crude acrylic acid to ensure an improved alcohol input in the esterification process, 0.02 to 0.08 kWh / kg crude acrylic acid should be used.

[0060] Preferably, the evaporated portion of the crude acrylic acid stream 104 is condensed in the condenser 108 and returned to the separation column 100 above the extraction point of the crude acrylic acid stream 104.

[0061] The concentration of high-boiling substances in the condensate is reduced. In applications where the concentration of high-boiling substances (e.g., benzaldehyde or furfural) is critical, such as in the rectification of high-purity acrylic acid for the production of superabsorbents by polymerization of the acrylic acid, the use of the condensate from partial evaporation as a feedstock instead of recycling is also conceivable.

[0062] Below the point where the crude acrylic acid stream 104 is withdrawn from the separation column 100, the removal of the heavy-boiling components takes place, and between the point where the crude acrylic acid stream 104 is withdrawn and the point where the reflux liquid stream 103 is added, the acrylic acid condensation takes place.

[0063] A high-purity acrylic acid stream (106) for esterification is obtained with minimal equipment and without separating elements. This ensures low plant costs and high plant availability.

[0064] The reflux rate 103 in the separation column 100 can preferably be adjusted via the amount of reflux at the top of the separation column 100 or at a position far above the point where the crude acrylic acid stream 104 is drawn off. The reflux into the separation column is preferably obtained by condensing the low-boiling components of the reaction gas and consists essentially of water. Alternatively, water or other low-boiling components can also be used as reflux. The condensation of the low-boiling components is preferably carried out by cooling following the acrylic acid condensation, either in a separation column section located above or in a separate separation column.

[0065] Preferably, the liquid mixture 104 is discharged via a side outlet of the separation column 100. Alternatively, the heavy-boiling component separation can also be carried out in a separate separation column prior to the acrylic acid condensation. In this case, the liquid mixture 104 could be drawn off from the bottom of the acrylic acid condensation column.

[0066] Preferably, the separation column 100 is operated such that the temperature profile in the separation column exhibits an inverse sigmoidal curve, particularly when the liquid mixture is partially or completely fed to the esterification plant. This ensures a higher quality of the crude acrylic acid 104. It is also 250427W001

[0067] 8

[0068] It is possible to operate the separation column with more reflux than would be necessary for an inverse sigmoidal temperature profile. The temperature profile of the acrylic acid condensation begins to drop relatively quickly, and then decreases only very slowly. This preferentially leads to lower acrylic acid losses in the exhaust gas. Sufficient product quality of the acrylic acid stream 106 for direct esterification must be ensured by an adequate energy supply to the evaporator 107.

[0069] The temperature profile or its course is preferably inversely sigmoidal and can be essentially determined by the function

[0070] f ( n ) = - — vT — dn + a

[0071]

[0072] V 1 1 + e (6*(nc)) «

[0073] where n = soil number, c = soil with the greatest temperature change, and a, b, d, g = adjustment constants,

[0074] The curve approximates the inverse sigmoidal shape. This means that initially the temperature drops slowly (the curve is flat) and then gradually decreases until a sharp or exponential drop occurs (the curve is very steep). Subsequently, the rate of decrease slows again, and the temperature continues to drop slowly (the curve is flat). In summary, the inverse sigmoidal curve is flat in its upper part, steep in its middle, and flat again in its lower part.

[0075] The temperature profile in the distillation column or in a distillation column section cannot be adjusted by a conventional ratio control of the column reflux, e.g., in relation to the (meth)acrylic acid flow rate in the reaction gas. It is influenced, among other things, by the reaction gas temperature, the temperature of gas mixture 101 at the inlet to the distillation column, the acrylic acid concentration in the reaction gas, and the reflux temperature and composition. The temperature profile is preferably adjusted by measuring the temperature in the distillation column, e.g., using a temperature sensor. The temperature measurement is preferably taken, or the temperature sensor is preferably located, where the inverse sigmoidal shape of the temperature profile exhibits the steepest or exponential drop (steep curve), i.e., where the temperature change is greatest. For this purpose, a temperature setpoint is preferably defined.If the measured temperature (actual temperature) is above the setpoint, the reflux rate in the separation column is increased; if the measured temperature (actual temperature) is below the setpoint, the reflux rate in the separation column is reduced. The inverse sigmoidal temperature profile preferably extends from the sampling plate of the liquid (meth)acrylic acid-containing mixture to the plate onto which the reflux is added.

[0076] Preferably, a temperature measurement is taken in the section of the column with the greatest temperature change, and the reflux rate is controlled as a function of the measured temperature (actual temperature). As preferably described above, the temperature profile is controlled by increasing the reflux rate when the temperature setpoint is exceeded and decreasing it when the temperature setpoint is undershot. 250427W001

[0077] 9

[0078] The reflux rate in the distillation column can preferably be adjusted by the amount of reflux at the top of the column or at a point above the steep temperature drop above the side outlet from which the liquid mixture, the crude acrylic acid, is drawn off. The reflux into the distillation column is preferably obtained by condensing the low-boiling components of the reaction gas and consists essentially of water. Alternatively, water or other low-boiling components can also be used as reflux. The condensation of the low-boiling components is preferably carried out by cooling following the (meth)acrylic acid condensation, either in a distillation column section located above the column or in a separate distillation column.

[0079] Preferably, the acrylic acid content in the gas mixture is more than 10 wt.%, more preferably more than 12 wt.%, further preferably more than 14 wt.%, and particularly preferably more than 16 wt.%. This has the advantage that a lower specific mass flow rate with respect to acrylic acid is required. This allows for smaller apparatus dimensions and reduced electricity consumption.

[0080] Preferably, the gas mixture is cooled before being fed into the separation column, more preferably before the spray cooler. The gas mixture is preferably cooled to at least 240°C, more preferably to at least 220°C, and particularly preferably to at least 200°C. This has the advantage that additional steam can be generated.

[0081] Preferably, the energy input during partial evaporation of the liquid mixture is at least 0.010 kWh / kg of liquid mixture, more preferably at least 0.015 kWh / kg of liquid mixture, more preferably at least 0.020 kWh / kg of liquid mixture, and particularly preferably at least 0.025 kWh / kg of liquid mixture. More preferably, the energy input during partial evaporation of the liquid mixture is between 0.015 and 0.10 kWh / kg of liquid mixture, more preferably between 0.025 and 0.09 kWh / kg of liquid mixture, and more preferably between 0.03 and 0.08 kWh / kg of liquid mixture. This allows acrolein and similar low-boiling substances to be removed from the crude acrylic acid to ensure smooth operation of the esterification process and to achieve a good alcohol yield by controlling the acetic acid content.

[0082] Preferably, the acrolein content of the unevaporated liquid mixture is less than 20 wt. ppm, more preferably less than 15 wt. ppm, and further preferably less than 10 wt. ppm.

[0083] Preferably, the 2-furfural content of the unevaporated liquid mixture is less than 500 ppm by weight, more preferably less than 400 ppm by weight, and more preferably less than 300 ppm by weight.

[0084] Preferably, the acetic acid content of the unevaporated liquid mixture is less than 0.7 wt.%, more preferably less than 0.6 wt.%, and more preferably less than 0.5 wt.%.

[0085] Preferably, the water content of the unevaporated liquid mixture is less than 2.5 wt.%, more preferably less than 2.0 wt.%, and further preferably less than 1.5 wt.%. 250427W001

[0086] 10

[0087] Preferably, the pressure during partial evaporation of the liquid mixture is less than 400 mbar absolute, more preferably less than 300 mbar absolute, and more preferably less than 200 mbar absolute.

[0088] Preferably, the temperature difference between the hot and cold sides of the evaporator during partial evaporation is less than 200°C, more preferably less than 150°C, and further preferably less than 100°C. This reduces or prevents fouling in the evaporator.

[0089] Preferably, the evaporated portion of the liquid mixture is at least partially recycled back into the separation column above the point where the liquid mixture is drawn off. This results in a further improved quality of the crude acrylic acid.

[0090] Preferably, the (meth)acrylic acid contained in the evaporated portion of the liquid mixture is at least partially further purified. Further purification here refers to additional purification beyond that achieved in the separation column. Recycling to the separation column is not considered further purification, but rather, for example, a crystallization purification. Preferably, the (meth)acrylic acid thus further purified can be reacted to form a product such as (meth)acrylic acid alkyl esters or superabsorbent polymers (polyacrylates).

[0091] Preferably, the separation column for removing the heavy elements has at least 3, preferably at least 5, and more preferably at least 7 theoretical stages. This ensures effective heavy element separation.

[0092] Preferably, no separating elements, such as dividers or packings, are used in the partial evaporation of the liquid mixture. Complex flushing procedures for removing polymer or fouling from dividers or packings (as described in EP 1 459794 A2) are therefore unnecessary. Consequently, plant availability is significantly improved compared to rectificatory purification methods in the prior art.

[0093] Unlike bottom or packing fouling in a rectification column, removing deposits in the evaporator (evaporator fouling) is straightforward. During evaporator cleaning in the partial evaporation stage, crude (meth)acrylic acid can preferably continue to be produced for esterification. For this purpose, the partial evaporation stage is preferably isolated from the rest of the plant. Even if the partial evaporation stage is shut down, the quality of the crude (meth)acrylic acid is sufficient for esterification due to the removal of the heavy components, despite increased losses of the light components during esterification. Decoupling the operation of the partial evaporation stage from the rest of the plant ensures high availability and economic efficiency of the overall process.

[0094] 11

[0095] Preferably, at least two evaporators, preferably operated in parallel, are used for partial evaporation. With this conceivable redundant design of the partial evaporator, only the evaporator affected by fouling needs to be isolated for cleaning. This technical design, with manageable additional investment costs, does not impair (meth)acrylic acid production or esterification in any way.

[0096] Preferably, during a shutdown of the partial evaporation of the liquid mixture, the process for separating (meth)acrylic acid is continued and the liquid mixture from the separation column is at least partially fed directly to the esterification plant.

[0097] Another object of the invention is a system for separating (meth)acrylic acid from a gas mixture that is present as a crude product of a catalytic gas-phase oxidation of at least one C3 / C4 precursor of (meth)acrylic acid, comprising

[0098] a separation column

[0099] an esterification plant

[0100] an evaporator,

[0101] characterized in that the separation column is connected to the evaporator via at least one liquid line and the evaporator is connected to the esterification plant via at least one further liquid line.

[0102] Preferably, the system further comprises a condenser designed to condense the evaporated portion of the liquid mixture before it is returned to the separation column.

[0103] Such a system setup according to a preferred embodiment is shown schematically in Figure 1.

[0104] Preferably, at least one temperature sensor is arranged in the separation column, configured to measure the temperature within the column. The temperature sensor is preferably positioned where the inverse sigmoidal temperature profile in the separation column is expected to exhibit a steep or exponential drop (steep curve). Preferably, the reflux rate in or to the separation column is controlled based on this measured temperature. For this purpose, a temperature setpoint is preferably defined. If the measured temperature is above the setpoint, the reflux rate in the separation column is increased; if the measured temperature is below the setpoint, the reflux rate in the separation column is reduced. The reflux rate is preferably controlled by a suitable control unit, which, for example, controls the opening degree of a valve.

[0105] The features described herein for the plant according to the invention shall also be deemed disclosed, mutatis mutandis, for the method according to the invention, and vice versa. 250427W001

[0106] 12

[0107] Example 1 (comparative example)

[0108] Figure 2 shows a schematic representation of an exemplary system according to a non-inventive embodiment. Details that are not essential and are well known to those skilled in the art have been omitted.

[0109] 200 separation column

[0110] 201 Product gas mixture (gas mixture)

[0111] 202 Crude acrylic acid stream (liquid mixture)

[0112] 203 Reflux fluid flow

[0113] 204 Esterification plant

[0114] 205 Temperature sensor

[0115] 206 column floor

[0116] 208 Spray coolers

[0117] 209 Countercurrent extraction column

[0118] 210 Rear column

[0119] 211 Desorber

[0120] The product gas mixture was obtained from a heterogeneously catalyzed gas-phase partial oxidation of polymer-grade propylene. Mass flow rate, temperature, and composition are given in Table 1.

[0121] The product gas mixture 201 is cooled to the temperature Tq specified in Table 1 in a co-current spray cooler 208, into which a mass flow of 171 t / h sump liquid is introduced.

[0122] The gas mixture 201 (101 in Fig. 1) is then fed to the separation column 200 (100 in Fig. 1). The separation column 200 has an inner diameter of 7.4 m and features separation-effective internals in the form of different trays 206. Trays 1 to 37 of the separation column 200 are dual-flow trays. Trays 38 to 75 are single-flow Thormann trays.

[0123] A liquid mixture 202 (104 in Fig. 1), the so-called crude acrylic acid, was drawn from a side outlet on the bottom 22 of the separation column 200. The composition of the drawn crude acrylic acid is shown in Table 2. The crude acrylic acid stream is fed directly to the esterification unit 204 (109 in Fig. 1) without further thermal purification.

[0124] The temperature profile of column 200 was controlled by adjusting the reflux rate 203 so that the temperature measurement on tray 46 reached the target value of 88.1°C. The reflux rate is shown in Table 3. 250427W001

[0125] 13

[0126] The floors 1 to 21, i.e. below the side outlet, serve to remove the heavy elements from the gas mixture 201.

[0127] 4193 kg / h of the sump liquid are discharged to a residue column 210. The temperature on tray 10 is 102.5°C.

[0128] Above tray 75 is a collection tray and a sour water condensation unit (trays 76 to 85). A partial flow (570 t / h) of the sour water drawn off the collection tray via a side outlet is cooled by indirect heat exchange and returned to the trays, resulting in a gas temperature of 32°C at the outlet of column 200. Another partial flow of the sour water is returned to tray 75. The flow rate is variable depending on the target temperature on tray 46.

[0129] The residual gas extracted at the top of the separation column 200 is partially recycled as dilution gas (recycled gas) into the gas-phase partial oxidation process and partially fed as stripping gas to a desorber 211 and the residue column 210. The remaining residual gas stream is fed to the exhaust gas aftertreatment system, with the flow rate adjusted to maintain an absolute pressure of 1.2 bar at the top of the separation column 200.

[0130] The difference between the condensed and recirculated acid water is fed to a countercurrent extraction column 209. The packing section of the countercurrent extraction column 209 (MontzPak B1-350) has a diameter of 1.3 m and a height of 19.5 m. The acid water is fed countercurrently below the packing section, and the extraction solvent Palatino!® A is fed countercurrently above the packing section at a mass ratio of 0.97 and a temperature of 66°C.

[0131] The acid-depleted water taken from the head of the countercurrent extraction column 209, which is depleted of acrylic acid, is thermally utilized.

[0132] The extraction solvent, enriched with acrylic acid and taken from the sump of countercurrent extraction column 209, is fed to the top tray of desorber 211. Desorber 211 contains 5 dual-flow trays in the lower section, followed by 15 single-flow cross-flow trays, and has an inner diameter of 2 m.

[0133] Below the lowest tray 206, 10.7 t / h of compressed residual gas at a temperature of 162°C is fed to the desorber 211 in countercurrent flow with the enriched extraction solvent. The bottoms of the desorber 211 are heated to 162°C by heat exchange. For this purpose, a stream of 88 t / h of the bottoms is heated and returned to the desorber 211 on tray 5. The loaded stripping gas is discharged from the top of the desorber 211, mixed with 20.2 t / h of compressed residual gas at a temperature of 162°C, and fed below the first tray of the residue column 210. This column has an inner diameter of 2.4 m and 50 dual-flow trays. The bottoms liquid for the separation column 200 is fed onto tray 8.250427W001.

[0134] 14

[0135] The sump liquid is heated to 162°C by a forced-circulation evaporator. The mass flow rate through the forced-circulation evaporator is 550 t / h. A residue stream of 914 kg / h is discharged and sent for thermal treatment.

[0136] A gas mixture is extracted from the head of the residue column 210, fed to the spray cooler 208 and mixed there with reaction gas and sump contents.

[0137] A spray cooler is located at the top of residue column 210. A flow rate of 99 t / h is taken from the collecting tray located above the last dual-flow tray, indirectly cooled, and then sprayed into the gas stream, resulting in a liquid temperature of 69°C. A partial flow rate of 4000 kg / h of the liquid from the collecting tray is returned as reflux to tray 50 of residue column 210.

[0138] The production of n-butyl acrylate was carried out according to the procedure in Example 1 of EP 4015498 A1. N-butyl acrylate conforming to specifications with a purity > 99.5 wt% was produced. On one occasion, a disruption occurred during n-butyl acrylate production due to elevated acrolein levels in the wastewater.

[0139] The plant availability for the production of crude acrylic acid for esterification was 95.8%.

[0140] Example 2 (comparative example)

[0141] The procedure was the same as in Example 1, with the difference that the product gas mixture from the synthesis was cooled to 200°C while generating 4 bar of steam (Table 1). Propene oxidation was also carried out at a higher propene concentration than in Example 1 (Table 1). This resulted in a significantly lower overall mass flow rate. Mass flow rate, temperature, and composition are given in Table 1.

[0142] The product gas mixture 201 is cooled to the temperature Tq specified in Table 1 in a co-current spray cooler 208, into which a mass flow of 171 t / h sump liquid is introduced.

[0143] The gas mixture 201 (101 in Fig. 1) is then fed to the separation column 200 (100 in Fig. 1). The separation column 200 has an inner diameter of 6.5 m and features separation-effective internals in the form of different trays 206. Trays 1 to 37 of the separation column 200 are dual-flow trays. Trays 38 to 75 are single-flow Thormann trays.

[0144] A liquid mixture 202 (104 in Fig. 1), the so-called crude acrylic acid, was drawn from a side outlet on the bottom 22 of the separation column 200. The composition of the drawn crude acrylic acid is shown in Table 2. The crude acrylic acid stream 202 is fed directly to the esterification unit 204 (109 in Fig. 1) without further thermal purification.

[0145] 15

[0146] The temperature profile of the separation column 200 was controlled by adjusting the reflux flow rate 203 so that the temperature measurement on tray 46 reached the target value of 81°C. The reflux flow rate is listed in Table 3.

[0147] The trays 1 to 21, i.e. below the side outlet, serve to remove the heavy elements from the gas mixture 201.

[0148] 4193 kg / h of the sump liquid are discharged to a residue column 210. The temperature on tray 10 is 95.3°C.

[0149] Above tray 75 is a collection tray and a sour water condensation unit (trays 76 to 85). A partial flow (470 t / h) of the sour water drawn off the collection tray via a side outlet is cooled by indirect heat exchange and returned to the trays, resulting in a gas temperature of 28°C at the outlet of column 200. Another partial flow of the sour water is returned to tray 75. The flow rate is variable depending on the target temperature on tray 46.

[0150] The residual gas extracted at the top of the separation column 200 is partially recycled as dilution gas (recycled gas) into the gas-phase partial oxidation process and partially fed as stripping gas to a desorber 211 and the residue column 210. The remaining residual gas stream is fed to the exhaust gas aftertreatment system, with the flow rate adjusted to maintain an absolute pressure of 1.2 bar at the top of the separation column 200.

[0151] The difference between the condensed and recirculated acid water is fed to a countercurrent extraction column 209. The design and operation of the countercurrent extraction column 309 and the desorber 311 are identical to Example 1.

[0152] The design and operation of the residue column 310 are also identical to example 1.

[0153] A residue stream of 708 kg / h is discharged and sent for thermal treatment. Due to the cooling of the reaction gas to 200°C, the formation of acrylic acid oligomers was reduced, and thus residue losses were reduced compared to Example 1.

[0154] The process for producing n-butyl acrylate was carried out as described in Example 1 of EP 4015498 A1. N-butyl acrylate meeting specifications with a purity > 99.5 wt% was produced. However, the n-butanol input was increased by more than 11 kg n-butanol / tn-butyl acrylate. Furthermore, increased evaporator fouling occurred during operation of the n-butyl acrylate plant. On two occasions, operational disruptions occurred during n-butyl acrylate production due to elevated acrolein levels in the wastewater.

[0155] The comparison of the key figures in Table 2 shows that Example 2, with reaction gas cooling and increased propene concentration during propene oxidation, resulted in a significantly improved propene input number, increased steam production, and reduced electricity consumption compared to Example 1. The economic efficiency was therefore 250427W001

[0156] 16

[0157] The process was significantly improved, and the specific CO2 footprint was reduced by 20%. However, the considerably lower quality of the crude acrylic acid led to a significantly increased consumption of n-butanol and production problems in the production of n-butyl acrylate.

[0158] The plant availability for the production of crude acrylic acid for esterification was 96%.

[0159] Example 3

[0160] The process for separating acrylic acid is illustrated by way of example in Figure 3. Figure 3 shows a schematic representation of an exemplary plant according to a preferred embodiment. Details that are not essential and are well known to those skilled in the art have been omitted.

[0161] 300 separation column

[0162] 301 Product gas mixture (gas mixture) in separation column

[0163] 302 Crude acrylic acid stream (liquid mixture)

[0164] 303 Reflux fluid flow

[0165] 304 Esterification plant

[0166] 305 Temperature sensor

[0167] 306 Column floor

[0168] 307 Partial Evaporation

[0169] 308 Spray cooler

[0170] 309 Countercurrent extraction column

[0171] 310 Rear Column

[0172] 311 Desorber

[0173] 312 evaporators

[0174] 313 Capacitor

[0175] The features disclosed for the process shall also be deemed disclosed for the plant, mutatis mutandis, and vice versa.

[0176] The product gas mixture was obtained from a heterogeneously catalyzed gas-phase partial oxidation of polymer-grade propylene. As in Example 2, but unlike Example 1, the gas mixture is cooled to 200°C while generating 4 bar of steam, and the propene oxidation is carried out at a higher propene concentration than in Example 1. This results in a significantly lower overall mass flow rate. Mass flow rate, temperature, and composition are given in Table 1. 250427W001

[0177] 17

[0178] The product gas mixture 301 is cooled to the temperature Tq specified in Table 1 in a co-current spray cooler 308, into which a mass flow of 171 t / h of sump liquid is introduced.

[0179] The gas mixture is then fed to the separation column 300 (100 in Fig. 1). The separation column 300 has an inner diameter of 6.5 m and features separation-effective internals in the form of different trays 306. Trays 1 to 37 of the separation column 300 are dual-flow trays. Trays 38 to 75 are single-flow Thormann trays.

[0180] A liquid mixture 302 (104 in Fig. 1), the so-called crude acrylic acid, was drawn from a side outlet on the bottom 22 of the separation column 300. The crude acrylic acid stream 302 of 38573 kg / h is fed to partial evaporation.

[0181] The temperature profile of column 300 was controlled by adjusting the reflux flow rate 303 so that the temperature measurement on tray 46 reached the target value of 81°C. The reflux flow rate is listed in Table 3.

[0182] The trays 1 to 21, i.e. below the side outlet, serve to remove the heavy elements from the gas mixture 301.

[0183] 4193 kg / h of the sump liquid are discharged to a residue column 310. The temperature on tray 10 is 95.3°C.

[0184] Above tray 75 is a collection tray and a sour water condensation unit (trays 76 to 85). A partial flow (470 t / h) of the sour water drawn off the collection tray via a side outlet is cooled by indirect heat exchange and returned to the trays, resulting in a gas temperature of 28°C at the outlet of column 300. Another partial flow of the sour water is returned to tray 75. The flow rate is variable depending on the target temperature on tray 46.

[0185] The residual gas extracted at the top of the separation column 300 is partially recycled as dilution gas (recycled gas) into the gas-phase partial oxidation process and partially fed as stripping gas to a desorber 311 and the residue column 310. The remaining residual gas stream is fed to the exhaust gas aftertreatment system, with the flow rate adjusted to maintain an absolute pressure of 1.2 bar at the top of the separation column 300.

[0186] The difference between the condensed and recirculated acid water is fed to a countercurrent extraction column 309. The packing section of the countercurrent extraction column 309 (MontzPak B1-350) has a diameter of 1.3 m and a height of 19.5 m. The acid water is fed countercurrently below the packing section, and the extraction solvent Palatino!® A is fed countercurrently above the packing section at a mass ratio of 0.97 and a temperature of 66°C.

[0187] The acid-depleted water taken from the top of countercurrent extraction column 309, which is depleted of acrylic acid, is thermally utilized. 250427W001

[0188] 18

[0189] The extraction solvent, enriched with acrylic acid and taken from the sump of countercurrent extraction column 309, is fed to the top tray of desorber 311. Desorber 311 contains 5 dual-flow trays in the lower section, followed by 15 single-flow cross-flow trays, and has an inner diameter of 2 m.

[0190] Below the lowest tray, 10.7 t / h of compressed residual gas at a temperature of 162°C is fed to desorber 311 in countercurrent flow with the enriched extraction solvent. The contents of desorber 311 are heated to 162°C by heat exchange. For this purpose, a flow of 88 t / h of the contents of the bottom is heated and returned to desorber 311 on tray 5.

[0191] The loaded stripping gas is discharged from the top of desorber 311, mixed with 20.2 t / h of compressed residual gas at a temperature of 162°C, and fed below the first tray of the residue column 310. This column has an inner diameter of 2.4 m and 50 dual-flow trays. The bottoms liquid for the separation column 300 is fed onto tray 8.

[0192] The sump liquid is heated to 162°C by a forced-circulation evaporator. The mass flow rate through the forced-circulation evaporator is 550 t / h. A residue stream of 716 kg / h is discharged and sent for thermal treatment. Compared to Example 1, as in Example 2, the formation of acrylic acid oligomers was reduced and residue losses were lowered due to the cooling of the reaction gas to 200°C.

[0193] A gas mixture is extracted from the head of the residue column 310, fed to the spray cooler 308 and mixed there with reaction gas and sump contents.

[0194] A spray cooler is located at the top of residue column 310. A flow rate of 99 t / h is taken from the collecting tray located above the last dual-flow tray, indirectly cooled, and then sprayed into the gas stream, resulting in a liquid temperature of 69°C. A partial flow rate of 4000 kg / h of the liquid from the collecting tray is returned as reflux to tray 50 of residue column 310.

[0195] Partial evaporation takes place in a vessel 307 without separating internals and with a diameter of 2.5 m. The capacity of the bottom evaporator 312 is 1.2 MW (0.03 kWh / kg liquid mixture) and is regulated so that 30 wt% of the crude acrylic acid stream 302 supplied to the partial evaporation is evaporated. The mass flow rate through the bottom evaporator 312 is 83 t / h. The gas phase is condensed in the condenser 313. The top pressure of the partial evaporation unit 307 was set to 100 mbar absolute by a vacuum pump. The condensate was returned to the separation column 300 via tray 50.

[0196] If the temperature difference between the hot and cold sides of the evaporator was < 100°C during partial evaporation, evaporator fouling could be significantly reduced. 250427W001

[0197] 19

[0198] As a bottoms product of the partial evaporation 307, 23144 kg / h of higher purity acrylic acid was obtained and fed directly to the esterification plant 304 (109 in Fig. 1) without further thermal purification. The composition is shown in Table 2.

[0199] The process for producing n-butyl acrylate was the same as in Example 1 of EP 4015498 A1. N-butyl acrylate conforming to specifications with a purity > 99.5 wt% was produced. The n-butanol input was not increased compared to Example 1. Unlike Example 2, there was no increased evaporator fouling. Also unlike Examples 1 and 2, there were never any elevated acrolein levels in the wastewater.

[0200] Table 3 shows that Example 3, like Example 2, with reaction gas cooling and increased propene concentration, led to a significantly improved propene input, increased steam production, and reduced electricity consumption compared to Example 1. Economic efficiency was therefore significantly improved, and the specific CO2 footprint was reduced by 20%. However, due to the improved acrylic acid quality resulting from partial evaporation, Example 3, unlike Example 2, did not experience increased n-butanol consumption or production problems in n-butyl acrylate production.

[0201] The plant availability for the production of crude acrylic acid for esterification was 95.8%, while the availability of partial evaporation 307, despite the purging shutdown, was only slightly lower at 94% due to the simple cleaning process (partial evaporation has no separating components). 250427EP01

[0202] 20

[0203] Table 1: Product gas

[0204] Example

[0205] 1* 2* 3 Mass flow rate kg / h 195827 155130 155134 Temperature °C 265 200 200 Nitrogen wt.% 76.1 72.2 72.3 Acrylic acid wt.% 12.2 15.5 15.4 Water wt.% 5.4 6.1 6.1 Oxygen wt.% 2.3 2.2 2.2 Carbon dioxide wt.% 2.2 2.1 2.1 Carbon monoxide wt.% 0.685 0.65 0.651 Acetic acid wt.% 0.273 0.317 0.316 Propylene wt.% 0.264 0.252 0.251 Formaldehyde wt.% 0.254 0.280 0.280 Maleic anhydride wt.% 0.098 0.124 0.124 Acrolein wt.% 0.093 0.090 0.089 Benzaldehyde wt.% 0.0101 0.013 0.013 Propionic acid wt.% 0.0040 0.0050 0.0050 2-Furfural wt.% 0.0039 0.0040 0.0041 Gas temperature Tq °C 123 102 103

[0206]

[0207] Comparison example 250427W001

[0208] 21

[0209] Table 2: Compositions of the acrylic acid stream for esterification (Example)

[0210] Component [Weight - 3

[0211] 1* 2*

[0212] %]

[0213] Acrylic acid 97.2 96.2 97.7 Water 1.88 2.68 1.19 2-Furfural 0.023 0.020 0.021 Maleic anhydride 0.01 0.002 0.002 Acetic acid 0.39 0.64 0.45 Diacrylic acid 0.33 0.32 0.47 Benzaldehyde 0.0247 0.0074 0.0086 Propionic acid 0.0327 0.0321 0.0327 Acrolein 0.0054 0.0068 0.0003

[0214]

[0215] Comparison example 250427WC01

[0216] 22

[0217] Table 3: Key figures of the examples

[0218] Example

[0219] 1* 2* 3 Acrylic acid (ACS) [kg / h] 22789 23160 32144 Propene [kg / h] 15793 15793 15793 Unit Ratio**) [wt.%] 69.3 68.2 68.2

[0220] Additional steam production (t-steam / t-ACS) - 0.4 0.3

[0221] Reduction in electricity consumption (MWh / t - 0.17 0.17 ACS)

[0222] Additional butanol consumption (kg n-butanol / t- - 11 - Butyl acrylate)

[0223] Reduction of CO2 footprint (kg-CO2 / kg-ACS) - 0.15 0.14

[0224] Return flow rate [kg / h] 30755 21182 20127

[0225]

[0226] *) Comparative example

[0227] **) Number of uses = Amount of propene used [kg] / Amount of acrylic acid obtained [kg]

Claims

250427EP01 23 Patent claims 1. Process for separating (meth)acrylic acid from a gas mixture which is present as a crude product of a catalytic gas-phase oxidation of at least one C3 / C4 precursor of (meth)acrylic acid, wherein a) the gas mixture is depleted of high boiling points, b) the gas mixture depleted of heavy elements is condensed in a separation column with separation-effective internals, and c) a liquid mixture containing (meth)acrylic acid is removed from the separation column, characterized in that the liquid mixture is partially evaporated and the non-evaporated liquid mixture is at least partially fed directly to an esterification plant for the conversion of the (meth)acrylic acid to (meth)acrylic acid alkyl esters without further purification.

2. The method according to claim 1, characterized in that the acrylic acid content in the gas mixture is more than 14 wt.%.

3. Method according to one of claims 1 or 2, characterized in that the gas mixture is cooled before being fed into the separation column.

4. Method according to one of claims 1 to 3, characterized in that the energy input during the partial evaporation of the liquid mixture is at least 0.015 kWh / kg liquid mixture.

5. Method according to any one of claims 1 to 4, characterized in that the acrolein content of the unevaporated liquid mixture is less than 10 ppm by weight.

6. Method according to any one of claims 1 to 5, characterized in that the 2-furfural content of the unevaporated liquid mixture is less than 300 ppm by weight.

7. Method according to any one of claims 1 to 6, characterized in that the acetic acid content of the unevaporated liquid mixture is less than 0.5 wt.%.

8. Method according to one of claims 1 to 7, characterized in that the pressure during partial evaporation of the liquid mixture is less than 200 mbar absolute.

9. Method according to any one of claims 1 to 8, characterized in that the temperature difference between the hot and cold sides of the evaporator during partial evaporation is less than 100°C. 250427W001 24 10. Method according to any one of claims 1 to 9, characterized in that the evaporated portion of the liquid mixture is at least partially recycled to the separation column above the extraction point of the liquid mixture and / or the (meth)acrylic acid contained in the evaporated portion of the liquid mixture is at least partially further purified.

11. Method according to one of claims 1 to 10, characterized in that the separation column for removing the heavy elements has at least 5 theoretical stages.

12. Method according to one of claims 1 to 11, characterized in that no separating components are used during the partial evaporation of the liquid mixture.

13. Method according to one of claims 1 to 12, characterized in that the separation column is operated in such a way that the temperature profile in the separation column has an inverse sigmoidal profile.

14. Method according to one of claims 1 to 13, characterized in that during a shutdown of the partial evaporation of the liquid mixture, the liquid mixture from the separation column is at least partially fed directly to the esterification plant.

15. Plant for separating (meth)acrylic acid from a gas mixture present as a crude product of a catalytic gas-phase oxidation of at least one C3 / C4 precursor of (meth)acrylic acid, comprising a separation column, an esterification plant an evaporator, characterized in that the separation column is connected to the evaporator via at least one liquid line and the evaporator is connected to the esterification plant via at least one further liquid line.

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