Method for producing (meth) acrylic acid

Abstract

The invention relates to the production of (meth) acrylic acid in a process based on the use of two distillation columns (dehydration column and refining column) in the absence of an external organic solvent. The invention more particularly relates to performing thermal cracking after the bottom of the evaporation, hydrolysis refining column, followed by recycling the gas phase of the cracker at the bottom of the dehydration column and upgrading the cracked residue by hydrothermal gasification. The method according to the invention makes it possible to improve the energy balance of the method, while improving the material balance.

Description

Technical Field

[0001] The present invention relates to the production of (meth)acrylic acid in a process based on the use of two distillation columns (a dehydration column and a refining column) in the absence of an external organic solvent. The invention more particularly relates to carrying out thermal cracking after evaporation, hydrolysis of the refining column bottom, subsequent recycling of the cracker gas phase at the dehydration column bottom and upgrading of the cracked residue by hydrothermal gasification. The process according to the invention makes it possible to improve the energy balance of the process while improving the material balance. Background Art

[0002] The process for synthesizing acrylic acid used on a large industrial scale employs the catalytic oxidation of propylene in the presence of oxygen.

[0003] The reaction is usually carried out in the gas phase and is usually carried out in two steps: a first step in which propene is oxidized substantially quantitatively to a mixture rich in acrolein and a second step in which acrolein is selectively oxidized to acrylic acid.

[0004] The gas mixture obtained from the second step consists, besides acrylic acid, of unconverted compounds resulting from the reactants involved or impurities generated during at least one reaction step, namely:

[0005] - light compounds which are non-condensable under the temperature and pressure conditions normally used, i.e. essentially: propene, propane, nitrogen, unconverted oxygen, carbon monoxide and dioxide formed in small amounts by final oxidation;

[0006] - condensable light compounds, ie essentially: water, unconverted acrolein, light aldehydes such as formaldehyde, glyoxal and acetaldehyde, formic acid, acetic acid or acrylic acid;

[0007] - compounds having a boiling point slightly higher than that of acrylic acid: furfural, benzaldehyde, maleic acid and maleic anhydride, benzoic acid, 2-butenoic acid, phenol, protoanemonin;

[0008] - Finally, heavy compounds derived from the addition of compounds with nucleophilic nature on the double bonds of unsaturated carbonyl monomers, through the Michael reaction.

[0009] The complexity of the gas mixture obtained in this process necessitates a set of operations to recover the acrylic acid contained in the gas effluent and convert it into acrylic acid grades compatible with its final use, for example the synthesis of acrylates or the production of acrylic acid and / or acrylate polymers.

[0010] Document EP 2 066 613, based on the so-called "solvent-free" technology, describes a process for recovering acrylic acid (AA) without using external water or azeotropic solvents. This process uses only two distillation columns to purify the cooled gaseous reaction mixture: a dehydration column and a finishing column (or purification column) fed with a portion of the bottom stream from the dehydration column.

[0011] According to the process, the cooled gaseous reaction stream is subjected to dehydration in a first column. The gaseous stream distilled at the top of the column is sent to a condenser, where light compounds are partially condensed and returned as liquid reflux to the dehydration column to absorb acrylic acid, and the uncondensed gaseous effluent is at least partially returned to the reaction and the rest is removed.

[0012] The bottom stream from the dehydration column is sent to a second column called the refining column. During the purification / refining step, a stream rich in heavy compounds is removed at the bottom and a distillate containing water and light by-products is collected at the top, condensed and then recycled at the bottom of the first dehydration column, forming a recycle loop.

[0013] By lateral withdrawal from the refining column, a purified acrylic acid stream is recovered in liquid or vapor form. The acrylic acid obtained generally has a purity of more than 98.5% by mass and contains less than 0.5% by mass of water and less than 0.4% by mass of acetic acid.

[0014] The temperature and pressure operating conditions of the purification column are not critical in the process and can be determined according to distillation methods known in the prior art. However, preferably, the purification column is operated at a pressure below atmospheric pressure to avoid polymerization of unsaturated products present and to minimize the formation of heavy by-products. These compounds are heavy products that reduce the recovery efficiency by consuming acrylic acid monomer. In the case of an AA production unit, it is essentially:

[0015] - derivatives of acrylic acid added to the double bond of another acrylic acid molecule: 3-acryloxypropionic acid, also called "acrylic acid dimer" or "AA dimer";

[0016] - derivatives of acrylic acid added to the double bonds of AA dimer molecules to form "AA trimers", as well as other oligomers formed by successive additions of acrylic acid to the double bonds of the aforementioned AA oligomers;

[0017] - Derivatives in which carboxylic acids formed as by-products of acrylic acid or water are added to the double bonds of AA or the above-mentioned oligomers.

[0018] Like free-radical polymerization, this covalent reaction to form Michael derivatives is strongly promoted by temperature. Therefore, the installation of a column with a large number of distillation plates in order to meet the quality requirements for acrylic acid leads to disadvantages in terms of product losses, which can only be compensated by additional high-temperature cleavage treatment of the Michael derivatives to regenerate acrylic acid monomers or by optional recycling of this stream from the bottom of the refining column to the ester unit.

[0019] In the case of heavies originating from AA production units, it is difficult to recover upgradeable monomers from the Michael-derived heavies. In fact, during the thermal cracking process of the regenerated acrylic acid (which is distilled and upgraded), a residue remains, the viscosity of which increases greatly when a high cracking efficiency is sought, until it can no longer be extracted from the cracking reactor.

[0020] As described in document FR2727964, the main factor limiting the efficiency of the regeneration of the compounds derived from the Michael reaction contained in the heavy stream from the AA unit, when the fraction rich in acrylic monomers has evaporated, is the increase in viscosity of the heavy residue obtained at the bottom of the cracker. The evaporation of light compounds during the cracking leads to a concentration of heavy products in the residue stream and an increase in the viscosity of this stream. However, the residue should remain sufficiently fluid after cooling to be transported and then treated to destroy it. The viscosity of the residue obtained at the end of the cracking increases with the residence time of the mixture to be treated at high temperature and with the amount of light monomers recovered by distillation. In order to obtain a residue viscosity compatible with normal transfer conditions and to reduce fouling phenomena, these two parameters must be limited, which has the effect of reducing the cracking yield.

[0021] Application FR2206330 achieves this improvement in the reproducibility in a continuous process, without a significant increase in the dynamic viscosity approaching 1 Pa·s, by carrying out the hydrolysis of heavy by-products before carrying out the thermal cracking, in a water:heavy acrylic acid mass ratio ranging from 0.1 to 1.3.

[0022] In the case of the presence of a production of light esters (methyl acrylate (MA) or ethyl acrylate (EA)) close to an AA production unit, the co-cracking of the corresponding heavies can improve the situation by making the cracking residue more mobile. The proposed solution makes it possible to recover the maximum amount of AA by the cracking operation while managing the viscosity of the residue formed. Thus, in document EP 717 031, it has been shown that the efficiency of recovering these valuable products that can be upgraded can be improved if the cracking is carried out with a mixture of heavies from an AA production unit and an acrylic esters (EA) production unit, compared to the separate cracking of the heavies from these units. The effect of adding the heavies originating from the ester unit (EAHP) to the heavies originating from the AA unit (HAA) is to reduce the viscosity of the final residue. The cracking reaction is carried out with a mixture of AA heavies / ester heavies in a ratio of 9 / 1 to 1 / 9 at a temperature of 180° C. to 220° C. at atmospheric pressure and a residence time of 0.5 to 3 hours.

[0023] Document FR3110571 proposes combining a partial condensation with a cleavage reactor, making it possible to increase the cleavage yield without significantly affecting the viscosity of the residue.

[0024] In the case of separate treatment of the heavies of acrylic acid (HAA), it is also envisaged to add a solvent to the residue.

[0025] Document EP 3255030 teaches that during the cracking of the residue the maleic anhydride present in the residue is converted into maleic esters which are less susceptible to polymerization.

[0026] Document US 6414183 teaches diluting the discharged residue with a solvent such as acetic acid, water and methanol. This dissolution is carried out at the bottom of the distillation column or evaporator in a ratio of 0.1 to 5 times relative to the bottom product before treatment by combustion.

[0027] Application WO 2021 / 224044 describes a method for decomposing Michael adducts of acrylic acid by dilution in a solvent having a boiling point of at least 170°C at 1013 hPa and a solubility in water of at least 20 g per 100 g of water at 25°C, the solvent being selected from alcohols such as ethylene glycol, propylene glycol, diethylene glycol, triethylene glycol and 2-ethoxyethanol, formamides such as N,N-dimethylacetamide, N-methylacetamide and N,N-dimethylformamide, sulfoxides such as dimethyl sulfoxide, and sulfones such as cyclopentane.

[0028] However, these solutions have several disadvantages, such as the generation of waste to be burned, the provision of additional equipment for carrying out the mixing, or the energy consumption of evaporation if water is assumed to have been chosen as solvent. In addition, most of these solvents generate nitrogen- or sulfur-containing derivatives when burned.

[0029] Oxidation (incineration) of organic matter into carbon dioxide and water is commonly used to treat organic residues and to produce heating steam. In conventional processes, rapid oxidation of organic fuels is used to generate heat, which is then transferred to a fluid such as water in a heat exchanger. Due to the losses that necessarily occur in the discharge tower of a conventional boiler, 10-15% heat losses are expected. In addition to possible blockages due to solids supplied to the boiler, hot spots due to salt deposits on the boiler tubes or ash deposits on the tube faces exposed to the flame or hot gases reduce good heat transfer and therefore reduce the heat transfer efficiency, or even cause very expensive time losses due to rupture of the tube walls.

[0030] On the one hand, there is still a need for a solvent-free purification line for the heavies from the bottom of the refining column, which is independent of the operation of the (meth)acrylate unit. On the other hand, it is desirable to be able to upgrade the final residue from the (meth)acrylic acid production to exportable methane gas instead of converting it into CO2 by combustion. Summary of the invention

[0031] The present invention makes it possible to meet the above-mentioned needs. More specifically, the present invention provides an improved process for the production of (meth)acrylic acid, which makes it possible to better upgrade the end product, which is usually sent to incineration, by regenerating the starting materials through cleavage and by converting the residue into a combustible gas. The process according to the invention makes it possible to improve the energy balance of the process while improving the material balance.

[0032] This result is obtained by carrying out evaporation, hydrolysis of the bottom of the refining column followed by thermal cracking, then recycling the gas phase of the cracker in the bottom of the dehydration column and upgrading the cracked residue by hydrothermal gasification.

[0033] The present invention relates to a method for producing industrial-grade (meth)acrylic acid from a gaseous reaction mixture containing (meth)acrylic acid obtained by gas-phase oxidation of a (meth)acrylic acid precursor in the absence of an organic solvent, comprising the following steps:

[0034] a) a step of dehydrating the gaseous reaction mixture in a first distillation column, referred to as a dehydration column, producing an overhead stream and a bottoms stream, at least a portion of the overhead stream being condensed and returned to the dehydration column in the form of reflux, and at least a portion of the bottoms stream being returned to the lower part of the dehydration column to form a recycle loop;

[0035] b) a step of distilling at least a portion of said bottom stream coming from the dehydration column in a second distillation column, called a refining column, so that a bottom stream containing heavy compounds, a top stream containing light compounds and a side draw stream of technical grade (meth)acrylic acid can be separated, at least a portion of said top stream being returned to the dehydration column;

[0036] c) a step of concentrating the bottom stream from the refining column in an evaporator to obtain a bottom stream concentrated in Michael adduct and a top stream comprising (meth)acrylic acid, the top stream being returned to the refining column;

[0037] d) a step of hydrolyzing said bottom stream coming from the evaporator in a hydrolyzer in the presence of water, resulting in obtaining a stream of hydrolyzate;

[0038] e) a step of thermally cracking the stream of hydrolysate in a cracker, resulting in obtaining an overhead stream and a bottoms residue, the overhead stream being recycled to the dehydration column; and

[0039] f) carrying out a step of hydrothermal treatment of said residue in a hydrothermal gasification device in the presence of water, resulting in obtaining methane, hydrogen and CO2 type gases at the top and a solid residue and water at the bottom.

[0040] The present invention makes it possible to overcome the disadvantages of the prior art. More specifically, the present invention provides a process for obtaining high-purity technical-grade (meth)acrylic acid having as a specification a purity of (meth)acrylic acid of more than 98.5%, in combination with a process for upgrading the Michael adducts as reactants recycled in the process, thereby increasing the productivity of the process and improving the energy balance by upgrading the residues to be removed. BRIEF DESCRIPTION OF THE DRAWINGS

[0041] Reference Figure 1 Other features and advantages of the present invention will become more apparent by reading the following detailed description.

[0042] Figure 1 : An overall diagram of an acrylic acid purification process having a combination of an evaporator at the bottom of a refining tower and a thermal cracker after a hydrolyzer, and a hydrothermal gasification device. DETAILED DESCRIPTION

[0043] The present invention provides an improved process for obtaining high-purity technical grade (meth)acrylic acid.

[0044] According to various embodiments, the method comprises the following features, in combination where appropriate.

[0045] According to one embodiment, the process according to the present invention is a process for producing high-purity technical grade acrylic acid.

[0046] According to one embodiment, the process according to the invention is a process for producing high-purity technical-grade methacrylic acid.

[0047] The present invention is described below by taking a method for producing acrylic acid as an example. The method according to the present invention may also include other preliminary, intermediate or subsequent steps, provided that they do not adversely affect the obtaining of purified acrylic acid.

[0048] According to one embodiment of the present invention, the acrylic acid precursor is acrolein.

[0049] According to one embodiment of the invention, acrolein is obtained by oxidation of propene or by oxidative dehydrogenation of propane.

[0050] According to one embodiment of the invention, the gaseous reaction mixture comprising acrylic acid obtained by oxidation of an acrylic acid precursor in the gas phase comprises carbon of renewable origin.

[0051] According to one embodiment of the invention, the acrylic acid precursor is derived from glycerol, 3-hydroxypropionic acid or 2-hydroxypropionic acid (lactic acid).

[0052] According to a preferred embodiment of the present invention, the gaseous reaction mixture comprises acrylic acid derived from propylene obtained according to the two-step oxidation process.

[0053] According to one embodiment, the finishing column is a conventional distillation column.

[0054] According to one embodiment, the refining column is a dividing wall column.

[0055] According to one embodiment, the refining column is operated under a reduced pressure of 5 to 60 kPa.

[0056] According to one embodiment, the aldehyde reducing chemical may be injected into the feed to the finishing column.

[0057] According to one embodiment, the evaporator placed at the bottom of the refining column is a film evaporator operating under a reduced pressure of 0.5 to 100 kPa.

[0058] According to one embodiment, the product from the top of the evaporator is recycled into the lateral withdrawal line of the finishing column.

[0059] According to one embodiment, the product from the top of the evaporator is recycled to the bottom of the finishing column, below the lateral withdrawal line.

[0060] According to one embodiment, the pressure in the hydrolyser varies between 0.1 and 2 MPa, preferably between 0.5 and 1.5 MPa.

[0061] According to one embodiment, the water / adduct mass ratio in the hydrolyser varies between 0.1 and 1.3, inclusive.

[0062] According to one embodiment, the temperature in the hydrolyser varies between 80°C and 200°C, preferably between 150°C and 200°C.

[0063] According to one embodiment, the thermal cracking reaction occurs in the absence of a catalyst.

[0064] According to one embodiment, the pyrolysis temperature is comprised between 140°C and 260°C, preferably between 160°C and 210°C.

[0065] According to one embodiment, the acrylic acid adduct is subjected to thermal cleavage.

[0066] According to one embodiment, the mixture of acrylic acid adduct and ester is subjected to thermal cleavage.

[0067] According to one embodiment, the residence time of the reaction mixture in the cleavage reactor is comprised between 0.5 h and 10 h, preferably between 4 h and 10 h.

[0068] According to one embodiment, the thermal cracking reaction is carried out at atmospheric pressure or under light pressure (maximum 0.2 MPa).

[0069] According to one embodiment, the product at the top of the cracker is recycled to the boiler of the dehydration column.

[0070] According to one embodiment, the product from the top of the cracker is mixed with the product from the top of the finishing column.

[0071] According to one embodiment, the bottom stream (residue) from the reactor obtained at the end of the thermal cracking operation has a dynamic viscosity of less than 1 Pa.s, preferably less than 10 Pa.s, measured at a temperature of 100°C, for example using a Brookfield "CAP 1000+" cone / plate viscometer.

[0072] According to one embodiment, a polymerization inhibitor is used in at least one step of the production process according to the invention. The polymerization inhibitor can be added at different locations, with the introduction of the reactants or at the top of the distillation column, exchanger and condenser.

[0073] As polymerization inhibitors that can be used, mention may be made of, for example, phenothiazine (PTZ), hydroquinone (HQ), hydroquinone monomethyl ether (HQME), di-tert-butyl-p-cresol (BHT), p-phenylenediamine, TEMPO (2,2,6,6-tetramethyl-1-piperidinyloxy), di-tert-butylcatechol or TEMPO derivatives such as OH-TEMPO, manganese acetate, alone or as a mixture in any ratio, which may be present in the reaction medium in an amount between 50 ppm and 5000 ppm, optionally in the presence of depleted air, but generally in an amount between 150 ppm and 1000 ppm.

[0074] In order to make the polymerization inhibitor more effective, it is appropriate to inject oxygen, air or so-called "lean" air containing 7% O2 at the bottom of the column. Preferably, the amount of oxygen injected corresponds to a content of 0.2% to 0.5% relative to the amount of organic vapor in the column.

[0075] According to one embodiment, the hydrothermal gasification device includes a first reactor, a second reactor and a liquid-gas separator.

[0076] According to one embodiment, the residue is injected as is into the gasification and the water required for the hydrothermal treatment is injected elsewhere.

[0077] According to one embodiment, the residue is mixed with the water required for the hydrothermal treatment before being introduced into the gasification.

[0078] According to this embodiment, the hydrothermal gasification is carried out at a temperature of 350-450°C and a pressure of 25 MPa.

[0079] According to this embodiment, the hydrothermal gasification comprises a gasifier making it possible to separate the salt at the bottom and the gas and liquid mixture at the top.

[0080] According to one embodiment, the hydrothermal gasification comprises a separator making it possible to separate the salts under critical conditions, a gasifier and a liquid-gas separator.

[0081] According to one embodiment, the hydrothermal gasification includes a salt separator, a gasifier containing a catalyst, and a liquid-gas separator.

[0082] According to one embodiment, the concentration of residue / water+residue in the salt separator is between 10 g / l and 400 g / l.

[0083] According to one embodiment, the water used for the hydrothermal gasification may be demineralized water, water from a wellbore or weakly mineralized water.

[0084] According to one embodiment, the water freed from organic compounds leaving the gasifier can advantageously be recycled into the feed of the separator or into the feed of the hydrolyzer.

[0085] According to one embodiment, the obtained and separated salts may be upgraded into fertilizers.

[0086] According to one embodiment, a proportion of 94% to 99% of the carbon introduced into the gasification is upgraded in gaseous form.

[0087] According to one embodiment, the gas originating from the gasification is composed of 40-70% methane, 5-20% hydrogen and 20-40% carbon dioxide.

[0088] According to one embodiment, the gas may be further fractionated to isolate methane from other compounds.

[0089] according to Figure 1 In the process shown, a gaseous reaction mixture 1 containing acrylic acid obtained by gas phase oxidation of an acrylic acid precursor is fed to a first distillation column 10. The gaseous reaction mixture containing a water / acrylic acid mass ratio of generally between 0.3 and 2, preferably between 0.3 and 1.2, may be precooled before being subjected to dehydration in the dehydration column 10.

[0090] In addition to water and acrylic acid, the reaction mixture also contains non-condensable light products such as nitrogen, oxygen, carbon monoxide and carbon dioxide, and various light or heavy by-products of different chemical nature, which may be light aldehydes such as acrolein, formaldehyde or acetaldehyde, heavy aldehydes such as furfural or benzaldehyde, light acids such as formic acid, acetic acid or propionic acid, heavy acids such as maleic acid, benzoic acid or 2-butenoic acid, and the lactone-type heavy compound protoanemonin.

[0091] The dehydration column produces a top stream 2, at least a portion of which is condensed in a condenser 13 and returned to the dehydration column as a reflux 7 for the absorption of acrylic acid, the other portion (stream 14) comprising non-condensable light compounds being generally sent partly or completely to a purification unit or partly recycled to other steps of the process for producing acrylic acid, preferably in a step situated upstream of the reactor for producing the reaction mixture 1.

[0092] The purpose of the dehydration step is to remove most of the water present in the reaction mixture as well as non-condensable light compounds and condensable light compounds in the overhead stream. It produces an overhead stream 2, which contains most of the water and light compounds and very small amounts of acrylic acid and heavy compounds, and a bottom stream 15, which is depleted in light compounds and contains most of the acrylic acid and heavy by-products, and has a water content by mass of generally less than 10%, preferably less than 7%.

[0093] A typical mass composition of the bottom stream 15 from the dehydration column comprises essentially acrylic acid (70-90%), acetic acid (2-20%), water (2-15%) and heavy by-products.

[0094] The dehydration column generally comprises 5 to 50 theoretical plates, preferably 20 to 30 theoretical plates.

[0095] Advantageously, the dehydration tower is operated at atmospheric pressure or slightly higher pressure, up to 1.5×10 5 It operates at an absolute pressure of Pa.

[0096] Advantageously, the temperature in the upper part of the dehydration column is at least 40° C., preferably between 40° C. and 80° C. The temperature of the bottom stream from the dehydration column preferably does not exceed 120° C.

[0097] The bottom stream 15 from the dehydration column is at least partially (stream 3) sent to the top of a second distillation column 16 (called purification column or finishing column) where the top stream 8 and the bottom stream 9 are separated.

[0098] A portion 20 of the liquid stream 15 from the bottom of the dehydration tower is sent to a heat exchanger 12 (which may be a heater or a cooler) and reinjected into the dehydration tower so as to constitute a bottom recirculation loop. Preferably, a portion 11 from the bottom loop is reinjected between the feed of the gas reaction mixture and the top of the dehydration tower.

[0099] The remainder of liquid stream 15 (stream 3) is sent as feed to finishing column 16.

[0100] The finishing column 16 is generally a conventional distillation column which comprises 5 to 30 theoretical plates, preferably 8 to 20 theoretical plates. The distillation column is combined with at least one reboiler 17 at the bottom and with a condenser 19 at the top.

[0101] The temperature and pressure in column 16 are not critical and can be determined according to distillation methods known in the art. However, preferably, refining column 16 is operated at a pressure below atmospheric pressure so that it can be operated at a relatively low temperature to avoid polymerization of unsaturated products present and minimize the formation of heavy by-products.

[0102] Advantageously, the finishing column is operated at an absolute pressure ranging from 5 kPa to about 60 kPa, the temperature of the overhead stream is advantageously between 40°C and about 90°C, and the temperature of the bottom stream is between 60°C and 120°C.

[0103] The overhead gas stream 8 from the finishing column is sent to a condenser 19, and the exiting liquid stream 4 is returned to the dehydration column to mix with the stream from the bottom loop of the dehydration column. The overhead stream 8 contains water and condensable light by-products.

[0104] The lateral draw stream 5 located in the first third of the bottom of the refining column, preferably starting from the bottom above the theoretical plate 3, comprises technical-grade acrylic acid having a purity of >98.5%.

[0105] The stream 9 separated at the bottom of the refining column contains most of the heavy by-products, in particular Michael addition products, such as 3-acryloxypropionic acid, maleic anhydride / acid, benzoic acid and polymerization inhibitors. The typical mass composition of the bottom stream 9 essentially contains acrylic acid (70-90%), polymerization inhibitors (0.5-2%) and heavy by-products (5-30%).

[0106] This stream 9 can be partly recycled to the bottom of the finishing column or sent via line 6 to the falling film evaporator.

[0107] The evaporator 21 is operated under reduced pressure of 0.5 to 60 kPa and in a temperature range of 50 to 150° C. After condensation and addition of stabilizer, the gas stream 22 comprising essentially acrylic acid is returned to the column at one theoretical stage below the side draw.

[0108] Preferably, the residue 25 also contains an acrylic acid content of >10% and <40%, in order to limit its viscosity. This stream 25 and water 34 are introduced under pressure into a reactor which allows the hydrolysis to be carried out at a temperature ranging from 100° C. to 170° C. at an autogenous pressure ranging from 1 atm to 2 MPa for a period of between 1 hour and 5 hours. The reactor may be a perfectly stirred reactor, a reactor equipped with an external recirculation loop and an exchanger, or a piston reactor. The stream 27 is then fed to a thermal cracker.

[0109] The cracker 28 comprises a liquid-gas separator and an external recirculation loop fed by a tubular exchanger heated by steam at a pressure between 1.5 and 3 MPa. The cracking time is between 1 and 10 hours at a pressure close to atmospheric pressure. After complete condensation and addition of inhibitors, the top stream 32 is mixed with the stream 4 returned to the dehydration tower.

[0110] The residue 30 and water 37 are injected into the hydrothermal gasification device 33 with a temperature range of 350°C to 450°C and a pressure of 25 MPa through two loops via a high-pressure pump. The device comprises:

[0111] - a first reactor, making it possible to separate the salts from the water and organic solution at its bottom,

[0112] - a second gasification reactor comprising a catalyst which allows the conversion of the organic products into gases to be completed, and

[0113] - a liquid-gas separator which makes it possible to recover at the bottom an aqueous phase 35 which can be recycled to the inlet of the separator or to the inlet of the hydrolyzer, and a gaseous phase 36 which is rich in methane, and which can be upgraded to produce electricity, which can bring the energy required for the gasification operation, as well as the energy required for the operation of the reaction and purification lines of the process or be exported elsewhere.

[0114] This hydrothermal gasification can be carried out in batch mode or, preferably, in continuous mode.

[0115] The following examples illustrate the invention without however limiting its scope.

[0116] Experimental Section

[0117] In the examples, unless otherwise indicated, percentages are shown by weight and the following abbreviations are used:

[0118] PTZ: Phenothiazine

[0119] AA: Acrylic acid

[0120] MA: Maleic acid

[0121] H2O: water

[0122] DiAA: Dimer of acrylic acid

[0123] AA3: Acrylic acid trimer

[0124] Heavy matter: oligomers with a mass greater than AA3

[0125] HQ: Hydroquinone

[0126] ACOH: Acetic acid

[0127] HAA: Heavy Acrylic Acid

[0128] Example

[0129] Solvent-free purification test

[0130] The characteristics of the solvent-free method are as follows:

[0131] Dehydration tower: diameter 300mm

[0132] Theoretical series: 22

[0133] Refining tower: diameter 300mm

[0134] Theoretical series: 17

[0135] Lateral removal: 14

[0136] The compositions obtained at different points in the process are shown in Table 1:

[0137] [Table 1]

[0138]

[0139] The column bottom contains about 11% DiAA, which is concentrated on the membrane evaporator.

[0140] Concentration at the bottom of the column on a membrane evaporator

[0141] The concentration at the bottom of the column was simulated in Aspen.

[0142] The operating conditions and compositions are given in Table 2 below:

[0143] [Table 2]

[0144]

[0145] Concentrating the bottom product by evaporation is effective. In fact, the AA content of the bottom decreases. In contrast, this evaporation, which produces acrylic acid with a purity of 98% at the top, is not efficient enough to mix this product with the technical grade acrylic acid obtained at the side draw and should therefore be recycled into the column.

[0146] Thermal lysis with or without prior hydrolysis

[0147] This example corresponds to examples 1 and 2 of application FR 2206330 and shows the advantages of treating the evaporator bottoms by hydrolysis followed by cleavage, in terms of the efficiency of the treatment.

[0148] It should also be noted that the top product will therefore consist of about 70% AA and 30% water, which is very close to the composition of the top of the finishing column and makes mixing these two streams very easy.

[0149] [Table 3]

[0150]

[0151] Hydrothermal gasification

[0152] The hydrothermal gasification will be illustrated by the very analogous case of the Michael adduct, butyl acrylate heavies.

[0153] The ABU heavy material mixture consists of:

[0154] -Butanol < 0.1%

[0155] -Butyl acrylate (5-10%)

[0156] -Butyl Hydroxypropionate (BHP): 1-3%

[0157] -Butyl butoxypropionate (BPB) 70-80%

[0158] -Acryloyloxybutyl propionate (AA / ABU) 4-6%

[0159] -Dibutyl maleate: 2-5%

[0160] - Phenothiazines: 1-3%.

[0161] 33 g / h of ABU heavies and 970 g / h of water were introduced into the separator and catalytic reactor via two different pipes, both operated at 400°C and 25 MPa. After 6 hours of testing under stable conditions, the ABU heavies were converted into a gas mixture with the following volume composition: 51% CH4; 34% CO2 and 19% H2. The energy content of this gas corresponds to 7096 kWh / ton ABU. The amount of TOC (total organic carbon) is <1 mg / l.

Claims

1. A method for producing industrial-grade (meth)acrylic acid from a gaseous reaction mixture containing (meth)acrylic acid obtained by gas-phase oxidation of a (meth)acrylic acid precursor in the absence of an organic solvent, comprising the following steps: a) a step of dehydrating the gaseous reaction mixture in a first distillation column, referred to as a dehydration column, producing an overhead stream and a bottom stream, at least a portion of which is condensed and returned to the dehydration column in the form of reflux, and at least a portion of which is returned to the lower part of the dehydration column to form a recycle loop; b) a step of distilling at least a portion of said bottom stream coming from the dehydration column in a second distillation column, called a refining column, so that a bottom stream containing heavy compounds, a top stream containing light compounds and a side draw stream of technical grade (meth)acrylic acid can be separated, at least a portion of said top stream being returned to the dehydration column; c) a step of concentrating the bottom stream from the refining column in an evaporator to obtain a bottom stream concentrated in Michael adduct and a top stream comprising (meth)acrylic acid, the top stream being returned to the refining column; d) a step of hydrolyzing said bottom stream coming from the evaporator in a hydrolyzer in the presence of water, resulting in obtaining a stream of hydrolyzate; e) a step of thermally cracking the stream of hydrolysate in a cracker, resulting in obtaining an overhead stream and a bottoms residue, the overhead stream being recycled to a dehydration column; and f) carrying out a step of hydrothermal treatment of said residue in a hydrothermal gasification device in the presence of water, resulting in obtaining methane, hydrogen and CO2 type gases at the top and a solid residue and water at the bottom.

2. The method according to claim 1, wherein the evaporator disposed at the bottom of the refining tower is a film evaporator operating under a reduced pressure of 0.5 kPa to 100 kPa.

3. The process according to any one of claims 1 and 2, wherein the product from the top of the evaporator is recycled to the bottom of the finishing column below the lateral withdrawal line.

4. The process according to any one of claims 1 to 3, wherein the temperature of the hydrolyser varies between 80°C and 200°C.

5. The process according to any one of claims 1 to 4, wherein the water / adduct mass ratio in the hydrolyser varies between 0.1 and 1.3, inclusive.

6. The process according to any one of claims 1 to 5, wherein the acrylic acid adduct is subjected to thermal cleavage.

7. The process according to any one of claims 1 to 5, wherein the mixture of acrylic acid adducts and acrylic esters is subjected to thermal cleavage.

8. The process according to any one of claims 1 to 7, wherein the product from the top of the cracker is mixed with the product from the top of the finishing column.

9. The method according to any one of claims 1 to 8, wherein the hydrothermal gasification comprises a separator for separating salts under critical conditions, a gasifier and a liquid-gas separator.

10. The process according to claim 9, wherein the concentration of residue / water+residue in the separator is between 10 g / l and 400 g / l.

11. The method according to any one of claims 1 to 10, wherein the gasification produces a gas consisting of 40% to 70% methane, 5% to 20% hydrogen and 20% to 40% carbon dioxide.

12. The process according to any one of claims 1 to 11, wherein the water free of organic compounds leaving the gasifier is recycled to the feed of the separator or to the feed of the hydrolyser.

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