OPTIMIZED PROCESS FOR SYNTHESIZING ALKYL METHACRYLATE BY REDUCING UNWANTED BYPRODUCTS

Abstract

An improved process for synthesizing alkyl methacrylates, in particular methyl methacrylate (MMA), involves reacting acetone cyanohydrin (ACH) and sulfuric acid in a first reaction stage (amidation). The process then involves heating the first reaction mixture in a second reaction stage (conversion) such that methacrylamide (MAA) is obtained; and then esterifying methacrylamide (MAA) with alcohol and water, preferably methanol and water, in a third reaction stage such that alkyl methacrylate is formed. The sulfuric acid used has a concentration of 98.0 wt % to 100.0 wt %. A subsequent working up of the third reaction mixture involves least two distillations in which the byproducts methacrylonitrile (MeAN) and acetone are obtained as an aqueous heteroazeotrope at least in part in the top fraction. At least some of the aqueous heteroazeotrope is removed from the process and at least partially reintroduced into the third reaction stage.

Claims

1: A process for preparing alkyl methacrylate, comprising: a. reacting acetone cyanohydrin and sulfuric acid in one or more reactors I in a first reaction stage for amidation at a temperature in the range from 70 to 130 C., to obtain a first reaction mixture comprising sulfoxyisobutyramide and methacrylamide; b. converting the first reaction mixture by heating to a temperature in the range from 130 to 200 C. in one or more reactors II in a second reaction stage for conversion, to obtain a second reaction mixture comprising predominantly the methacrylamide and the sulfuric acid; c. reacting the second reaction mixture with alcohol and water, in one or more reactors III in a third reaction stage for esterification, to obtain a third reaction mixture comprising the alkyl methacrylate; and d. separating the alkyl methacrylate from the third reaction mixture obtained from the third reaction stage; wherein the sulfuric acid used in the first reaction stage has a concentration in the range from 98.0% by weight to 100.0% by weight; wherein the separation the alkyl methacrylate from the third reaction mixture comprises at least two distillation steps in which methacrylonitrile and acetone by-products are obtained at least partly in a tops fraction as a water-containing heteroazeotrope, wherein the water-containing heteroazeotrope comprising the methacrylonitrile and the acetone from at least one of the at least two distillation steps is at least partly discharged from the process, and wherein at least one stream comprising the methacrylonitrile and the acetone is at least partly recycled into the third reaction stage.

2: The process according to claim 1, wherein the at least one stream comprising the methacrylonitrile and the acetone which is at least partly recycled into the third reaction stage is a water-containing heteroazeotrope comprising the methacrylonitrile and the acetone from at least one of the at least two distillation steps.

3: The process according to claim 1, wherein at least one aqueous phase which is obtained by condensation and phase separation of the water-containing heteroazeotrope from at least one of the at least two distillation steps is recycled fully or partly into the third reaction stage, where the at least one aqueous phase is contacted with the second reaction mixture comprising predominantly the methacrylamide and the sulfuric acid.

4: The process according to claim 1, wherein at least one aqueous phase which is obtained by condensation and phase separation of the water-containing heteroazeotrope from at least one of the at least two distillation steps is discharged fully or partly from the process, optionally after an extraction step.

5: The process according to claim 1, wherein the water-containing heteroazeotrope from at least one of the at least two distillation steps is discharged fully or partly from the process, at least partly in the form of a gaseous stream, optionally after a scrubbing step.

6: The process according to claim 1, wherein the separation of the alkyl methacrylate from the third reaction mixture comprises at least one phase separation step in which the water-containing heteroazeotrope from at least one of the at least two distillation steps is separated into an aqueous phase comprising the methacrylonitrile and the acetone, and an organic phase comprising predominantly the alkyl methacrylate, wherein the aqueous phase is partly discharged from the process and/or partly recycled into the third reaction stage, and wherein the organic phase comprising predominantly the alkyl methacrylate is recycled fully or partly into the at least one of the at least two distillation steps.

7: The process according to claim 1, wherein the third reaction mixture obtained in the third reaction stage is evaporated continuously, wherein a resultant vapour stream is fed to a first distillation step K1 in which a tops fraction comprising the alkyl methacrylate, water and the alcohol, and a bottoms fraction comprising higher-boiling components are obtained, and wherein the bottoms fraction is recycled fully or partly into the third reaction stage.

8: The process according to claim 1, wherein the separation of the alkyl methacrylate from the third reaction mixture comprises (i) first distilling the third reaction mixture obtained in the third reaction stage in a first distillation step K1, to obtain a first water-containing heteroazeotrope comprising the methacrylonitrile and the acetone as a tops fraction; (ii) separating the first water-containing heteroazeotrope as condensate in a phase separation step in a phase separator I, into an aqueous phase WP-1 and an organic phase OP-1 comprising a predominant portion of the alkyl methacrylate; (iii) guiding the organic phase OP-1 into a second distillation step K2, wherein a further tops fraction obtained is a second water-containing heteroazeotrope comprising the methacrylonitrile and the acetone; (iv) separating at least a portion of the second water-containing heteroazeotrope in a phase separation step in a phase separator II, into an aqueous phase WP-2 comprising the methacrylonitrile and the acetone, and an organic phase OP-2, wherein the organic phase OP-2 is recycled fully or partly into the second distillation step K2, and wherein the aqueous phase WP-2 comprising the methacrylonitrile and the acetone is partly recycled into the third reaction stage and partly discharged from the process, optionally after an extraction step.

9: The process according to claim 8, wherein the aqueous phase WP-1 is recycled fully or partly into the third reaction stage and the organic phase OP-1 comprising the predominant portion of the alkyl methacrylate is subjected to an extraction using water as extractant, wherein an aqueous phase of the extraction is recycled into the third reaction stage and a organic phase of the extraction is guided into the second distillation step K2.

10: The process according to claim 8, wherein a portion of the aqueous phase WP-2 comprising the methacrylonitrile and the acetone is subjected to an extraction to obtain an aqueous phase WP-3 and an organic phase OP-3, wherein the aqueous phase WP-3 is discharged fully or partly from the process, and wherein the organic phase OP-3 is recycled fully or partly into the third reaction stage.

11: The process according to claim 8, wherein a portion of the aqueous phase WP-2 comprising the methacrylonitrile and acetone is subjected to an extraction to obtain an aqueous phase WP-3 and an organic phase OP-3, wherein the aqueous phase WP-3 is subjected to a further distillation step K4, wherein a tops fraction comprising the methacrylonitrile is obtained in distillation step K4 and discharged from the process, wherein a bottoms fraction comprising water is obtained in distillation step K4 and recycled fully or partly into the extraction, and wherein the organic phase OP-3 is recycled fully or partly into the third reaction stage.

12: The process according to claim 1, wherein the separation of the alkyl methacrylate from the third reaction mixture comprises (i) first distilling the third reaction mixture obtained in the third reaction stage in a first distillation step K1, to obtain a first water-containing heteroazeotrope comprising the methacrylonitrile and the acetone as a tops fraction; (ii) guiding the first water-containing heteroazeotrope as a vapour stream into a further distillation step K4, in which a further water-containing heteroazeotrope comprising the methacrylonitrile and the acetone is obtained as a tops fraction, and a bottoms fraction comprising the alkyl methacrylate is obtained, (iii) discharging the tops fraction from distillation step K4, optionally after a scrubbing step, fully or partly from the process in the form of a gaseous stream, (iv) separating the bottoms fraction from distillation step K4 in a phase separation step in phase separator II, into an aqueous phase WP-2 comprising the methacrylonitrile and acetone, and an organic phase OP-2, wherein the aqueous phase WP-2 comprising the methacrylonitrile and the acetone is recycled fully or partly into the third reaction stage, and (v) guiding the organic phase WP-2 fully or partly into a second distillation step K2, in which the tops fraction obtained is a second water-containing heteroazeotrope comprising the methacrylonitrile and the acetone, which is condensed fully or partly and guided into the phase separation step in the phase separator II according to (iv).

13: The process according to claim 1, wherein the second reaction mixture contains not more than 3% by weight of methacrylic acid, not more than 1.5% by weight of alpha-hydroxyisobutyramide and not more than 0.3% by weight of the methacrylonitrile, based in each case on the overall amount of the second reaction mixture.

14: The process according to claim 1, wherein the second reaction mixture contains 30% to 40% by weight of the methacrylamide, based on the overall amount of the second reaction mixture.

15: The process according to claim 1, wherein the conversion of the acetone cyanohydrin and the sulfuric acid in the first reaction stage is effected in at least two separate reaction zones or in at least two separate reactors.

16: The process according to claim 1, wherein, in the first reaction stage, the conversion of the acetone cyanohydrin (ACH) and the sulfuric acid is effected in at least two separate reactors, wherein the sulfuric acid and the acetone cyanohydrin are used in a first reactor in a molar ratio of the sulfuric acid to the ACH in the range from 1.6 to 3.0, and wherein the sulfuric acid and the acetone cyanohydrin are used in a second reactor in a molar ratio of the sulfuric acid to the ACH in the range from 1.2 to 2.0.

17: The process according to claim 1, wherein the reaction mixture of acetone cyanohydrin (ACH) and sulfuric acid in the first reaction mixture includes a total amount of water in the range from 0.1 mol % to 20 mol %, based on an entirety of the ACH supplied to the first reaction stage.

18: The process according to claim 8, wherein a crude alkyl methacrylate product is obtained as a bottoms fraction in the second distillation step K2, wherein the crude alkyl methacrylate product is purified in a further distillation step K3, to obtain a pure alkyl methacrylate product as a tops fraction, having a methacrylonitrile content in the range from 10 to 300 ppm, based on an amount of the pure alkyl methacrylate product.

19: The process according to claim 1, wherein the process comprises a regeneration of the sulfuric acid, wherein a portion of the third reaction mixture obtained in the third reaction stage and at least one aqueous or organic waste stream comprising sulfuric acid, ammonium hydrogensulfate and sulfonated acetone derivatives that results from discharge of the water-containing heteroazeotrope comprising the methacrylonitrile and the acetone is sent to a thermal regeneration step in which the sulfuric acid is obtained, which is recycled into the first reaction stage.

20: The process according to claim 1, wherein the process comprises obtaining ammonium sulfate, wherein a portion of the third reaction mixture obtained in the third reaction stage and at least one aqueous or organic waste stream comprising sulfuric acid, ammonium hydrogensulfate and sulfonated acetone derivatives that results from discharge of the water-containing heteroazeotrope comprising the methacrylonitrile and the acetone is sent to a thermal regeneration step in which the ammonium sulfate is obtained by crystallization, which is separated off as a by-product.

Description

DESCRIPTION OF THE FIGURES

[0165] FIG. 1 shows a flow diagram of preferred embodiments of the process according to the invention. FIG. 1 shows the preferred elements of an integrated plant for continuous preparation and purification of alkyl methacrylates, especially methyl methacrylate (MMA). The integrated plant shown has various plants connected to one another, usually in a fluid-conducting manner, as elements of this integrated system. This integrated plant includes the preparation of methacrylamide or the sulfuric acid solution thereof, consisting of the process steps of amidation (A, B) and conversion (C, D), followed by an esterification (E), followed by a workup of the reaction product (F, G, H, I, J, K), followed in turn by a fine purification (L, M, N, O). Solid lines preferentially describe the flow pathways of the process according to variant A; dotted lines preferentially describe the flow pathways of the alternative process according to variant B. A combination of apparatuses and streams of matter from the two variants is likewise possible.

[0166] FIG. 2 shows a schematic flow diagram of a first preferred embodiment of the process according to the invention (variant A).

[0167] FIG. 3 shows a schematic flow diagram of a second preferred embodiment of the process according to the invention (variant B).

[0168] In FIGS. 1 to 3, the reference symbols have the following meanings:

[0169] Apparatuses [0170] (A) Stage 1 amidation reactor [0171] (B) Stage 2 amidation reactor [0172] (C) Heater [0173] (D) Gas separator/intermediate vessel [0174] (E) Esterification reactor/cascade [0175] (F) Primary column (column K1) [0176] (G) Phase separator I [0177] (H) Scrubbing column (extraction I) [0178] (I) Primary stripper column (column K4) [0179] (J) Offgas scrubbing column [0180] (K) Phase separator for crude MMA [0181] (L) Azeotrope column (column K2) [0182] (M) Condenser/vacuum system [0183] (N) Phase separator II [0184] (O) Purifying column (column K3) [0185] (P) Extraction column II (extraction of pump condensate)

[0186] Streams of Matter [0187] (1a) Acetone cyanohydrin feed to stage 1 [0188] (1b) Acetone cyanohydrin feed to stage 2 [0189] (2) Sulfuric acid feed [0190] (3) Amide mixture exiting stage 1 [0191] (4a) Offgas from stage 1 amidation reactors [0192] (4b) Offgas from stage 2 amidation reactors [0193] (5a) Optional offgas from stage 1 & 2 amidation reactors [0194] (5b) Offgas from stage 1 & stage 2 amidation reactors [0195] (6) Amide mixture exiting stage 2 [0196] (7) Converted amide mixture [0197] (8) Degassed amide mixture [0198] (9a) Gas separator offgas for esterification [0199] (9b) Optional offgas, removed from the process [0200] (10a) Alcohol feed (for MMA: methanol) [0201] (10b) Alcohol feed to offgas scrubbing column [0202] (11) Cleavage acid from esterification [0203] (12) Vapour stream from esterification [0204] (13) Liquid reflux stream from primary column [0205] (14a) Distillate stream from primary column [0206] (14b) Vapour stream from primary column [0207] (15a) Organic phase from phase separator I (OP-1) [0208] (15b) Aqueous phase from phase separator I (WP-1) [0209] (16a) Demineralized water feed to extraction [0210] (16b) Demineralized water feed [0211] (16c) Demineralized water feed to phase separator II (N) [0212] (16d) Direct steam [0213] (17a) Washed organic phase (OP-1) [0214] (17b) Aqueous phase from extraction [0215] (18) Combined aqueous phases [0216] (19a) Vapour stream from primary stripper column [0217] (19b) Bottom stream from primary stripper column [0218] (19c) Organic sidestream from primary stripper column/reflux from primary column [0219] (20a) Organic phase from phase separation I (OP-1) [0220] (20b) Aqueous phase from phase separation I (WP-1) [0221] (20c) Combined reflux stream/product water [0222] (21a) Offgas from offgas scrubbing column [0223] (21b) Bottom stream from offgas scrubbing column [0224] (22a) Vapour from azeotrope column [0225] (22b) Bottom product from azeotrope column/crude alkyl methacrylate product [0226] (23a) Condensate I to phase separator II [0227] (23b) Offgas from azeotrope column/vacuum system [0228] (23c) Circulation stream condensate [0229] (23d) Condensate II, vacuum pump condensate [0230] (23e) Offgas/inert content of the condensation/vacuum system [0231] (24a) Organic phase from phase separator II (OP-2) [0232] (24b) Aqueous phase from phase separator II (WP-2) [0233] (25a) Top product from purifying column/pure alkyl methacrylate product [0234] (25b) Bottom product from purifying column [0235] (26a) Discharge of aqueous phase from phase separator II [0236] (26b) Aqueous phase to pump condensate extraction [0237] (26c) Aqueous phase to esterification (recycle) [0238] (27) Cleavage acid [0239] (28a) Organic phase from extraction of pump condensate (OP-3) [0240] (28b) Aqueous phase from extraction of pump condensate (WP-3)/raffinate to cleavage acid [0241] (29) Methanol/methyl methacrylate mixture [0242] (30) Vacuum pump condensate

[0243] FIG. 4 describes the reaction network of the formation of methacrylic acid and/or methyl methacrylate proceeding from methane and ammonia, and acetone. Proceeding from methane (CH.sub.4) and ammonia (NH.sub.3), it is possible to prepare hydrogen cyanide via the BMA process (hydrogen cyanide from methane and ammonia) by means of catalytic dehydrogenation (CH.sub.4+NH.sub.3.fwdarw.HCN+3H.sub.2) (variant 1). Alternatively, it is possible to prepare hydrogen cyanide via the Andrussow process proceeding from methane and ammonia, with addition of oxygen (CH.sub.4+NH.sub.3+1.5 O.sub.2.fwdarw.HCN+3H.sub.2O) (variant 2). In the next step, proceeding from acetone and hydrogen cyanide, acetone cyanohydrin (ACH) is prepared with addition of a basic catalyst (e.g. diethylamine Et.sub.2NH or else alkali metal hydroxides). The hydroxyl group of acetone cyanohydrin is subsequently esterified with sulfuric acid, initially giving sulfoxyisobutyronitrile (SIBN). The nitrile group of sulfoxyisobutyronitrile (SIBN) can be hydrolysed in the next step under the action of sulfuric acid and water, giving sulfoxyisobutyramide hydrogensulfate (SIBA.Math.H.sub.2SO.sub.4). A side reaction that can proceed is the formation of methacrylonitrile (MAN) with elimination of sulfuric acid from SIBN. Sulfoxyisobutyramide hydrogensulfate (SIBA.Math.H.sub.2SO.sub.4) can additionally be partly hydrolysed to give alpha-hydroxyisobutyramide hydrogensulfate (HIBAm.Math.H.sub.2SO.sub.4). Likewise possible is the reverse reaction to give the sulfuric ester SIBA.Math.H.sub.2SO.sub.4. A by-product formed may be alpha-hydroxyisobutyric acid (HIBAc) via further hydrolysis of HIBAm.Math.H.sub.2SO.sub.4. Proceeding from SIBA.Math.H.sub.2SO.sub.4, with the elimination of sulfuric acid, methacrylamide hydrogensulfate (MAA.Math.H.sub.2SO.sub.4) is formed (conversion). The gradual reaction of HIBAm or HIBAc to give MA or MAA can likewise proceed as an elimination reaction with elimination of NH.sub.4HSO.sub.4 or water. Methacrylamide hydrogensulfate (MAA.Math.H.sub.2SO.sub.4) can subsequently be converted by hydrolysis to methacrylic acid (MA) or by esterification with methanol (MeOH) methyl methacrylate (MMA). If alpha-hydroxyisobutyric acid (HIBAc) is introduced into the esterification, it can be converted to methyl alpha-hydroxyisobutyrate (MHIB).

[0244] The abbreviations in FIG. 4 have the following meanings: [0245] ACH acetone cyanohydrin; [0246] SIBN alpha-sulfoxyisobutyronitrile; [0247] SIBA alpha-sulfoxyisobutyramide; [0248] SIBA.Math.H.sub.2SO.sub.4 alpha-sulfoxyisobutyramide hydrogensulfate; [0249] MAN methacrylonitrile; [0250] HIBA alpha-hydroxyisobutyramide; [0251] HIBAm.Math.H.sub.2SO.sub.4 alpha-hydroxyisobutyramide hydrogensulfate; [0252] MAA methacrylamide; [0253] MAA.Math.H.sub.2SO.sub.4 methacrylamide hydrogensulfate; [0254] MA methacrylic acid; [0255] MMA methyl methacrylate; [0256] HIBAc alpha-hydroxyisobutyric acid; [0257] MHIB methyl alpha-hydroxyisobutyrate

[0258] Embodiment of the Process According to FIG. 2 (Variant A)

[0259] One possible embodiment of the process (variant A) relating to the preparation of alkyl methacrylate, especially MMA, according to the flow diagram in FIG. 2 is described hereinafter:

[0260] In the amidation reactors (A) and (B), which take the form of a loop reactor, ACH and sulfuric acid are converted to a sulfuric acid solution comprising SIBA, HIBAm and MAA (each predominantly in the form of the hydrogensulfates). Depending on the reaction conditions, especially in reactors (A) and (B), MAN may be formed as a by-product from ACH with release of water. The loop reactors (A) and (B) each comprise the following elements: circulation pump, static mixer, heat exchanger and gas separator.

[0261] The amidation reactor (A) of stage 1 has an ACH feed (1a) and a sulfuric acid feed (2). The ACH feed (1a) opens into the circuit of the loop reactor (A) on the pressure side of the circulation pump, but upstream of the static mixer. The sulfuric acid feed (2) opens into the circuit of the loop reactor (A) upstream of the ACH feed (1a) and on the suction side of the circulation pump, which can preferably improve the pumpability of the gas-containing reaction mixture.

[0262] The reaction mixture in loop reactor (A) is pumped in circulation within the temperature range of 70-130 C. and at a circulation ratio (ratio of circulation volume flow rate to feed volume flow rate) in the range from 5 to 110, and the temperature can be adjusted by means of secondary water-cooled shell-and-tube heat exchangers. More particularly, the heat of reaction of the strongly exothermic reaction between acetone cyanohydrin and sulfuric acid is removed. The static dwell time in the reactor circuit of the amidation reactor (A) is in the range from 5 to 35 minutes. The amidation reactor (A) is operated at standard pressure. The blended and temperature-controlled reaction mixture is then introduced into a gas separator. The selective separation of gaseous secondary components (such as carbon monoxide and other inerts/low boilers) from the amide circulation stream and the discharge of the offgas stream (4a) are effected here.

[0263] A substream (3) of the reaction mixture pumped in circulation is fed to the second loop reactor (B) by means of a discharge pump, by gravimetric means or with supply pressure from the reactor circulation pump itself, and heated up by an additional heat exchanger if required. For further conversion of the reaction mixture (3), the amidation reactor (B) is supplied with fresh acetone cyanohydrin via the ACH feed (1b). Loop reactor (B) is configured in a comparable manner to loop reactor (A) in terms of temperature, pressure, dwell time and flow pathway.

[0264] Gaseous by-products are removed from reactor (B) in the form of the offgas stream (4b).

[0265] The resultant offgas streams (4a, 4b) are combined by means of interconnection to give (5) and sent to the downstream gas separator/intermediate vessel (D) for the purpose of utilization. Alternatively, the reaction offgases (4a, 4b) are removed from the process as a combined offgas stream (5a).

[0266] The resultant liquid reaction mixture (6) is subjected to a conversion (C) for maximum conversion to MAA. The conversion is typically composed of one or more heat exchangers, with controlled heating and subsequent dwell time of the entering reaction mixture (6) maximizing the concentration of MAA in the product stream exiting from the amidation, in the converted amide mixture (7).

[0267] The converted amide mixture (7) is sent gravimetrically, for example, to the gas separator/intermediate vessel (D). The resultant offgas is separated here from the viscous and hot converted amide mixture (7). The offgas released comprises mainly carbon monoxide that forms through breakdown reactions, and additionally ultrafine droplets of methacrylamide-containing reaction mixture. The reactant-containing overall offgas (9a) from the gas separator/intermediate vessel (D) is therefore passed onward into the esterification (E). The degassed amide mixture (8) is subsequently pumped or fed gravimetrically to the esterification (E).

[0268] The offgas stream (5b) from the amidation stages can be connected on the gas side to the gas separation vessel (D), and the overall offgas (9a) from process steps (A, B, C, D) may be connected to the vapour space of the esterification (E). Alternatively, the offgas from (D) can at least partly be removed from the process as offgas stream (9b).

[0269] In the esterification (E), the reactants required for conversion of methacrylamide to the corresponding ester are fed in directly or indirectly in the form of the corresponding alcohol (10a, 10b) and of demineralized water (16a, 16b, 16c). The degassed amide mixture (8) is fed to the reaction (E) here through introduction tubes or immersed tubes, in a pumped or gravimetric manner. A direct alcohol feed (10a) (e.g. methanol for the preparation of MMA) is usually effected by means of immersed introduction tubes or static mixers in the feed to the esterification (E).

[0270] In addition, various circulation streams from the thermal workup (F, G, H, L, M, N, O, P) are connected to the esterification reactor (E) as shown in FIG. 2.

[0271] The esterification is typically conducted in one or more esterification reactors (E) that are mixed by means of a stirrer or pump and are gravimetrically connected to one another. A further form of mixing is convection, which is caused by the supply of evaporable reactants. The esterification reactors are often equipped with heat exchangers in order to assure the input of heat for the esterification reactors. For example, the heat input is achieved by jacket heating, forced circulation evaporator or direct feeding of steam.

[0272] The reaction mixture (crude ester) formed in the esterification (E) is guided out of the esterification reactor (E) by distillation as a continuous vapour stream (12). The vapour stream (12) may also be combined here from multiple reactors (E). The acid mixture (11) remaining in the esterification reactors, after intensive distillative removal of residual product, is discharged from the esterification.

[0273] The vapour stream from the esterification (12) is subjected to a counter current distillation in the primary column (F). The vapour stream may be condensed at the top of the column (F) as reflux from the primary column (F) and partly returned. The offgas (30) obtained beyond the condensation, which is generated by the supply of stream (9a) inter alia, can be removed from the process and sent to incineration, for example.

[0274] The bottom product (13) comprising MAA is returned continuously to the esterification reactor (E). It is possible to distribute the bottom product (13) between multiple esterification tanks of the esterification (E).

[0275] The vapour stream (14a) at the top of the column (F) contains the predominant proportion of the alkyl methacrylate, and also water, alcohol, acetone and MAN. Methacrylic acid forms a low-boiling azeotrope with water and is likewise present in the vapour stream (12).

[0276] The aqueous and condensed vapour stream (14a) at the top of the column (F) is subjected to a phase separation (G) in the phase separator I, in which an organic phase (15a) comprising alkyl methacrylate, methanol, acetone and MAN, and an aqueous phase (15b) are obtained.

[0277] The organic phase (15a) is subjected to a liquid/liquid extraction (H), especially in order to return a large portion of the methanol present to the esterification (E). For this purpose, the organic phase (15a) is extracted in a stirred extraction column (H) with demineralized water (16a) in countercurrent. The resultant aqueous phase (17b) is combined with the aqueous phase (15b) from the phase separator (G) in stream (18) and returned to the esterification (E). The organic phase (17a) which is present in extraction step (H) and comprises the predominant portion of alkyl methacrylate and significant proportions of low and high boilers is sent to further thermal workup (L, M, N, O).

[0278] The organic phase (17a) from extraction step (H) is subjected to an azeotropic distillation (L) under reduced pressure in a further step. The azeotrope column is implemented in the form of a stripping column, wherein the organic feed (17a) is guided preheated to the top of the column (L), which is heated indirectly with low-pressure steam by an evaporator. At the top of the column (L), a heteroazeotropic mixture (22a) comprising MMA, water, methanol, acetone, MAN and further low boilers is obtained. The bottom product (22b) separated off is purified alkyl methacrylate (crude alkyl methacrylate). The vapour stream (22a) leaves the column (L) in vaporous form and is condensed stepwise under reduced pressure in the downstream condenser (M).

[0279] The main condensation in (M) proceeds on the suction side of the vacuum unit, forming a liquid condensate (23a) which is subjected to a phase separation in the phase separator II (N). On the pressure side of the condensation unit (M), a further liquid stream (23d, vacuum pump condensate) is generated, which serves as extractant in the extraction step (P). The inert gas-containing offgas (23e) formed in the condensation on the pressure side is removed from the process.

[0280] The liquid condensate (23a) from (M) is guided into the phase separation (N) with addition of demineralized water (16c) and separated into an organic phase (24a) and an aqueous phase (24b). The organic phase (24a) contains a certain proportion of alkyl methacrylate and is guided back into the distillation step (L) via the top of the column (L).

[0281] The aqueous phase (24b) from (N), in a corresponding manner to the added fresh water (16c), is saturated with water-soluble components such as methanol, acetone and MAN, and is divided into two streams for avoidance of by-product enrichment. A substream (26c) is returned to the esterification reactor (E) in the form of a circulation stream. A substream (26b) is discharged from the process via (28b) after an extraction step (P).

[0282] As an alternative to stream (26b, 26c), it is likewise possible to discard stream (24b) completely in the form of stream (26a) and discharge it from the process.

[0283] Stream (26b) serves as an outlet for enriched secondary components, which, for the purpose of recovery of alkyl methacrylate, is sent to an extraction column (D) (PK extraction column). In the extraction step (P), the condensate (23d) from condenser (M) is used as extractant, with guiding of the streams (23d) and (26b) in countercurrent. In the extraction step (B), an aqueous phase (28b) and an organic phase (28a) are obtained, wherein the aqueous phase (28b) is mixed with the waste acid (11) and discharged fully from the process as cleavage acid (27), and wherein the organic phase (28a) is fed into the esterification (E) as stream of value comprising alkyl methacrylate.

[0284] Embodiment of the Process According to FIG. 3 (Variant B)

[0285] One possible embodiment of the process (variant B) relating to the preparation of alkyl methacrylate, especially MMA, according to the flow diagram in FIG. 3 is described hereinafter:

[0286] The amidation in (A) and (B), the conversion in (C), the gas separation in (D) and the esterification in (E) are effected as described in the embodiment according to FIG. 2 (variant A). In addition, various circulation streams from the thermal workup (F, I, J, K, L, M, O) are connected to the esterification reactor (E) as shown in FIG. 3.

[0287] In the embodiment according to variant B (FIG. 3), the vapour stream formed in the primary column (F) is fed uncondensed as a vapour stream (14b) to a further distillation step (1) (primary stripper column). In the primary stripper column (I), the tops fraction obtained (19a) is a low-boiling mixture comprising methanol, acetone, methacrylonitrile, methacrylic esters and water, and the bottoms fraction obtained (19b) is an azeotropically boiling mixture crude ester/water mixture. The reflux for column (I) is generated by means of a partial condenser which is adjusted such that the low-boiling mixture (19a) of methanol, acetone, alkyl methacrylate and water is discharged in the form of a vapour, while an alkyl methacrylate-rich mixture can be returned to the column.

[0288] A portion of the liquid descending within the primary stripper column (I) is drawn off in the upper region of the column (I) in the form of a liquid side stream (19c) and used as reflux for the primary column (F).

[0289] The azeotropic mixture (19b) drawn off in the bottoms from the primary stripper column (I), comprising methacrylate esters, water, small amounts of low boilers (e.g. methanol, acetone) and high boilers (e.g. hydroxyisobutyric esters), is cooled down and sent to the phase separator II (K). The bottom product (19b) is separated into an organic phase (20a) and an aqueous phase (20b) in the phase separator (K). In addition, in the phase separator (K), the reflux stream (23c) that results from the condenser (M) and the top product from the azeotrope column (L) is separated into an organic phase (20a) and an aqueous phase (20b).

[0290] The aqueous phase (20b) from the phase separation (K) comprising water, saturated with methanol, acetone and methacrylic esters, is mixed with demineralized water (16b) and fed to the esterification (E) in the form of a combined reflux stream (20c). The reflux stream (20c) may especially serve to cover the water demand of the esterification and recover reactants.

[0291] The vapour stream (19a) that leaves the primary stripper column (I) is guided into a scrubbing step (J) (offgas scrubbing column) and scrubbed with fresh alcohol (10b), e.g. methanol, as scrubbing medium, which largely frees the gas stream from alkyl methacrylate.

[0292] The offgas scrubbing column (J) is operated in countercurrent. Alkyl methacrylate-containing output air (23b) from condenser (M) is likewise fed to the offgas scrubbing column (J) in order to recover alkyl methacrylate.

[0293] The organic stream (21b) comprising methanol and alkyl methacrylate is obtained in the bottoms from the offgas scrubbing column (J), and is recycled into the esterification (E). The organic reflux stream (21b) may be distributed here between various esterification reactors.

[0294] At the top of the offgas scrubbing column (J), an offgas stream (21a) is obtained, including MAN, dimethyl ether and methyl formate, and also saturation concentrations of methanol and acetone, and containing little alkyl methacrylate. By adjusting the ratio of methanol (10b) to vapour (19a) and the top temperature in (J), it is possible to vary the composition of the offgas stream (21a). The offgas stream (19a) comprising methanol, acetone and MAN is discharged fully from the process.

[0295] The organic phase (20a) from the phase separation (K) is guided into the azeotrope column (L). The organic phase (20a) comprising the predominant proportion of alkyl methacrylate and significant proportions of low and high boilers is sent to further thermal workup (L, M, N, O).

[0296] The thermal workup (L, M, N, O) is effected essentially as described above in the embodiment according to variant A (FIG. 2). By contrast with variant A, streams (23b) and (23c) formed in the condenser (M) are returned to the process steps (J) and (K). A fully condensed and azeotropic low boiler mixture (23c) which is obtained from the tops fraction from the azeotropic column (L) via condenser (M), and which typically has two liquid phases, is guided into the phase separation (K) in the upstream step. Stream (23c) is composed of streams (23a) and (23d) from variant A. The offgas (23b) formed in the condensation/vacuum unit (M) is guided into the scrubbing column (J), wherein materials of value in particular, such as alkyl methacrylate, are recovered.

[0297] The invention is described further by the examples that follow. Example A1 (comparative example) demonstrates the operation of a process for preparing methyl methacrylate with a noninventive sulfuric acid concentration (100.3% by weight) and a low discharge of the methacrylonitrile-containing aqueous phase from a phase separator II, combined with a low overall MMA yields to obtain an MMA product having a high content of the MAN by-product.

[0298] Inventive example A2 describes the preparation of methyl methacrylate having the claimed features, with achievement of a high overall yield of MMA, characterized by reduced MAN formation in the amidation and conversion, and by a moderation of the steady-state MAN concentration in the workup section via controlled discharge, combined with the achievement of a low MAN content in the MMA target product.

[0299] Example A3 (comparative example) illustrates operation analogously to comparative example 1, but additionally without extraction of the aqueous MAN-containing stream discharged, which brings another increase in MMA losses.

[0300] Beyond that, example B1 (inventive) describes the preparation of methyl methacrylate having the claimed features, with achievement of a high overall yield of MMA, characterized by reduced MAN formation in the amidation and conversion and an exclusively distillative removal of by-products (as vapour).

EXAMPLES

Examples A1, A2 and A3 According to FIG. 2 (Variant A)

[0301] The preparation of methyl methacrylate comprising the reaction of acetone cyanohydrin with sulfuric acid in the amidation/conversion ((A), (B), (C), (D)), the subsequent esterification (E) with methanol, and distillative and extractive workup ((F), (G), (H), (L), (O), (M), (N), (P)) of the methyl methacrylate product was effected by the embodiment according to FIG. 2 as described above (variant A).

[0302] A mass balance and assessment of the discharge of methacrylonitrile (MAN) and acetone via (28b), output air streams (30) and (23e), and via the MMA product (25a) was effected (see flow diagram according to FIG. 2).

[0303] Described hereinafter are comparative examples (Examples A1 and A3) using sulfuric acid with a concentration of 100.3% by weight (formally 0.3% by weight of free SO.sub.3) and an inventive example (Example A2) using sulfuric acid with a concentration of 99.7% by weight (0.3% by weight of water).

[0304] The water content of the ACH feed streams (1a) and (1b) is calculated from the difference from the ACH content which is ascertained by means of HPLC, or via an analysis by means of gas chromatography (with thermal conductivity detector) which is quantitative and selective specifically for water.

[0305] The water content in the sulfuric acid feed (2) is calculated from the difference from the sulfuric acid content which is ascertained by measuring the density and speed of sound.

[0306] The general procedure for the process according to FIG. 2 (variant A) is described hereinafter, with differences and results shown in Tables 1 to 3.

[0307] General Process Procedure:

[0308] 1a. Reaction Stages (A), (B), (C), (D) and (E)

[0309] 5000 kg/h of acetone cyanohydrin having a composition of 98.8% by weight of acetone cyanohydrin, 0.25% by weight of acetone, 0.65% by weight by water and free sulfuric acid was divided in a mass ratio of 75/25, so as to obtain a stream (1a) of 3750 kg/h and a stream (1b) of 1250 kg/h. Feed stream (1a) was subsequently applied to the first amidation reactor (A).

[0310] The loop reactor (A) was composed of the following elements connected by pipeline: circulation pump, static mixer, heat transferrer, cooler, and a gas separator. A circulation volume flow rate of 350 m.sup.3/h was established in the reactor (A), such that effective heat transfer and effective mixing and gas separation were possible. The overall reactor circuit was operated at about 95 C. and 990 mbar(a) at slightly reduced pressure.

[0311] Feed stream (1a), continuously and at a temperature of about 20 C., was fed to the reactor circuit (A) described and mixed in.

[0312] The amount of sulfuric acid (2) needed for the optimal conversion of the reaction mixture in reactors (A) and (B) that had a concentration according to Table 1 was fed to the reactor (A) in a load-dependent manner in the specified mass ratio to the total amount of ACH (1a+1b). This achieved a sulfuric acid excess (sulfuric acid/ACH ratio of 2.6 kg/kg) in reactor (A).

[0313] The resultant stirred-up mixture (3) comprising sulfoxyisobutyramide, methacrylamide and sulfuric acid was then transferred into the second amidation reactor (B), while the offgas (4a) separated off in reactor (A) was sent in the direction of conversion (D).

[0314] Reactor (B) was of analogous construction to reactor (A) and was operated under the same physical conditions and parameters. The offgas (4b) formed by side reaction was separated from the reaction mixture by means of a gas separator. The offgases from the amidation (4a) and (4b) were subsequently combined and supplied in the form of offgas stream (5b), the amount of which was about 60 m.sup.3/h, to a further gas separator/intermediate vessel (D).

[0315] The ACH stream (1b) was subsequently added to the reaction mixture in the second reactor (B). The mass ratio H.sub.2SO.sub.4/ACH established in the reaction mixture is reported in Table 1.

[0316] Over the course of the amidation reactors (A, B), a reaction mixture (6) at 95 C. comprising sulfoxyisobutyramide (SIBA), methacrylamide (MAA) and hydroxyisobutyramide (HIBAm), dissolved in the sulfuric acid reaction matrix, was obtained. This mixture was then subjected to a conversion step (C). The reaction mixture was heated therein to 155 C. within a short time and then converted thermally in a delay zone.

[0317] After the conversion (C), the methacrylamide-enriched reaction mixture (7) was supplied gravimetrically to a further gas separator or intermediate vessel (D) which was operated at a slightly reduced pressure of about 950 mbara and at a temperature of 155 C.

[0318] In the gas separator/intermediate vessel (D), the gas present in the reaction mixture was separated off and combined with offgas (5b). An overall offgas (9a) was obtained, which was supplied to the esterification (E) in gaseous form. The amount and composition of (9a), with regard to the acetone and MAN by-products, are collated in Table 1.

[0319] In addition, after the gas separation (D), a stream (8) was obtained, the mass flow rate and composition of which are reported in Table 1. The resultant amount of convertible reactants (MAA+MA) that are guided into the esterification and the amidation yield (MAA+MA) based on ACH are shown in Table 1.

[0320] The methacrylamide-containing stream (8), as well as the main components mentioned, contains acetone and MAN as important by-products in the amounts specified in Table 1. The overall mass flow rates of acetone and MAN that are fed to the esterification (E) via (8) and the gaseous feed stream (9a) are reported in Table 1.

[0321] The liquid reaction mixture (8) at more than 150 C. was fed to the esterification (E), wherein the feeding was effected into a reactor cascade consisting of three jacket-heated reactor tanks having virtually ideal mixing by free efflux by means of an immersed tube. Stream (8) was reacted in process step (E) with a total of 1540 kg/h of methanol (10a), 590 kg/h of MMA/MeOH mixture (29), 500 kg/h of direct steam (16d) and 1700 kg/h of water (feed streams (16a) and (16c)) at about 1205 C. and a slightly elevated pressure of 50-150 mbarg. The steady-state substance mixture (29) established consisted, on average, of 25% by weight of methyl methacrylate and 75% by weight of methanol, and also water, and was fed to the utilization of methanol as reactant in the esterification reaction and methyl methacrylate as product (E).

[0322] Suitable interconnection of the reactant streams (10a), (16a), (16c); (16d), (29) and of the circulation streams (13), (18), (25b), (26c), (28a) in the esterification (E) achieved a local stoichiometric excess of methanol and water based on the methacrylamide and methacrylic acid substances convertible to methyl methacrylate in each of the esterification tanks.

[0323] In a side reaction of the esterification, a portion of the amount of MAN supplied was converted to methacrylamide by hydrolysis and hence lost from the process. The higher the concentration of MAN fed in, the higher the level of hydrolysis as well. In this way, the amounts (MAN(hydrol.)) of MAN reported in Table 1 were hydrolysed in (E).

[0324] At the exit from the reactor unit (E), a post-evaporation operated with direct steam was connected, which reduces the proportion of monomers in the acid mixture to a content of <0.1% by weight of MMA, <0.1% of MA, <0.1% by weight of MAA, and also MAN and acetone, according to Table 1, while the water content of the effluxing acid mixture was kept virtually constant. There were likewise further nonvolatile organic by-products present in the waste acid (11), which were discharged in solution as TOC (total organic carbon) and to a certain degree also as polymeric solids. On account of side reactions in the esterification (E), the amounts of sulfonated acetone (Sulfo acetone) reported in Table 1 were present in the form of TOC in the process acid (11). The waste acid contained essentially NH.sub.4HSO.sub.4, H.sub.2SO.sub.4 and water. The TOC content of (11) averaged 2-3% by weight. The amount and composition of the waste acid (11) are compiled in Table 1.

[0325] The energy input into (E) for continuous evaporation of the products obtained in (E) was effected by means of 10 barg hot steam.

[0326] 1b. Prepurification (F), (G), (H)

[0327] The crude product formed in the esterification reaction and the methyl methacrylate (MMA) introduced via (29) were withdrawn continuously from the esterification cascade (E) in the form of a vapour stream (12). For this purpose, the esterification reactors were connected to a vapour conduit on the steam side, such that the vapour stream (12) was obtained as the cumulative stream from the reactors II. According to the vapour/liquid equilibrium of the reaction mixtures, the vapour stream (12) was a heteroazeotropic composition comprising MMA, water, MeOH, MA, acetone and MAN as reported in Table 2.

[0328] In addition to vapour stream (12), the offgas stream from the amidation (9a) was also fed to the bottom of the primary column (F). Vapour stream (12) was subsequently subjected to a countercurrent distillation by adding the vaporous stream (12) and the gaseous stream (9a) in the bottom region of a column (F). At the top of the column (F), full condensation was effected in condensers that were operated by means of cooling water and cold water. The biphasic distillates were combined, and a substream was guided into primary column (F) as reflux.

[0329] The offgas (30) obtained beyond the condensation, which was generated by the supply of stream (9a) inter alia, was removed and sent to an incineration. Via (30), MAN and acetone were discharged from the process in the amounts reported in Table 2.

[0330] In accordance with the reflux ratio, a liquid distillate stream (14a) and a liquid bottom stream (13) were obtained at the top of the column (F). The bottom stream (13) was recycled continuously into the esterification reaction. The heteroazeotropic distillate stream (14a), for the purpose of liquid/liquid phase separation, was fed to a phase separator (G) in which two product streams were obtained. The organic light phase (15a) comprising MMA, water, methanol, acetone, MA, MAN and high and low boilers was subjected to a liquid/liquid extraction (H).

[0331] The aqueous heavy phase (15b) was first combined with the aqueous raffinate phase (17b) to give a reflux stream (18) that included water, methanol, MMA, acetone, MA, MAN and high and low boilers. Stream (18), analogously to stream (13), was recycled continuously into the esterification reaction (E). The amounts and composition of (13), (14a), (15a), (15b) and (18) are collated in Table 2.

[0332] In the extraction step (H), stream (15a) was supplied with deionized water (16a) in order to remove further water-soluble components from the present crude MMA. For this purpose, at the top of the disc extraction column (H), deionized water was fed in continuously in countercurrent to stream (15a), such that, at the bottom of the column (H), an aqueous efflux stream (17b) and a prepurified crude MMA (17a) were obtained. The crude MMA stream (17a) contained MMA, water, methanol, acetone, MA, MAN and high and low boilers. The amounts and composition of (17a) and (17b) are collated in Table 2.

[0333] 1c. Fine Purification (L), (O), (M), (N), (P)

[0334] Stream (17a) was then subjected to a distillative purification (L). For removal of high and low boilers, stream (17a) was fed to the top region of an azeotrope column (L) operated under reduced pressure (300 mbara), which was heated indirectly with hot steam. A low boiler-enriched heteroazeotropic vapour stream (22a) was separated from a methyl methacrylate-enriched bottom stream (22b). The vapour stream (22a) contained MMA, water, MAN, acetone and further low boilers. The bottom stream (22b) contained MMA, MA, high boilers. MAN and acetone. The amounts and compositions are collated in Table 3.

[0335] At the exit from the column, on the tube side, the vapour stream (22a) was fed to a condensation/vacuum unit (M) that first subjected the vapour stream (22a) to a main condensation on the vacuum side, then compressed the residual gas in a vacuum pump and again subjected it to postcondensation on the pressure side of the compression process.

[0336] The heteroazeotropic distillate (23a) obtained after the main condensation was subjected to a phase separation (N) for further workup, while the pump distillate (23d) obtained on the pressure side was sent to the esterification reaction (E). The pump distillate (23d) contained low boilers, MMA, acetone, methanol, water and MAN. Amounts and compositions are described in Table 3.

[0337] The process offgas (23e) obtained beyond the postcondensation on the gas side was discharged continuously from the process. The process offgas (23e) contained low boilers and inert substances that are chemically reactive under given conditions, and also MAN and acetone. Amounts and compositions are described in Table 3.

[0338] For improvement of the phase separation, the distillate stream (23a) was supplied with deionized water (16c) in the phase separator (N), such that an organic phase (24a) and an aqueous phase (24b) were obtained. The light organic phase (24a) was circulated here continuously as reflux to the top of the column (L), while the aqueous phase (24b) comprising acetone and MAN was fed to the next process step. Amounts and compositions are described in Table 3.

[0339] In Comparative Example A1, stream (24b) was then divided in a fixed mass ratio of 80/20 as stream (26b) and stream (26c). 80% of stream (24b) was recycled directly into the esterification reaction as (26c). 20% of stream (24b) was discharged from the process in the form of stream (28b) via the intermediate step of a liquid/liquid extraction (P) for recovery of MMA. In this way, MAN and acetone were discharged from the process, and hence the enrichment thereof in the process was reduced, monitored and controlled. Amounts and compositions are described in Tables 3 and 4.

[0340] In Inventive Example A2, stream (24b) was first divided in a mass ratio of 50/50 into stream (26b) and stream (26c). Stream (26c) was recycled directly into the esterification reaction (E). Stream (26b) was partly discharged from the process via the intermediate step of a liquid/liquid extraction (P) as (28b).

[0341] In Comparative Example A3, stream (24b) was divided in a fixed mass ratio of 80/20 as stream (26a) and stream (26c). 80% of stream (24b) was recycled directly into the esterification reaction as (26c). 20% of stream (24b) was removed directly from the process as stream (26a) and not sent to any extraction step (P) for recovery of MMA. In this way, MAN and acetone were discharged from the process. Amounts and compositions are described in Table 3.

[0342] Alternatively, stream (24b) can also be removed from the process fully or partly in the form of a discharge stream (26a) prior to the division into (26b)/(26c). In that case, an even greater proportion of the MAN and acetone introduced would be removed from the process than described hereinafter.

[0343] In the extraction (P), the aqueous product stream (26b) was treated with the aid of the organic pump distillate (23d) as extractant, in order to reduce the residual content of MMA in stream (26b) prior to the discharge. For this purpose, stream (26b) was fed in at the top of the disc extraction column (P), and stream (23d) at the bottom. In (P), an aqueous raffinate (28b) at the bottom of the column (P) and an organic, methyl methacrylate-enriched extract stream (28a) were obtained. The raffinate stream (28b) contained water, methanol, acetone, methyl methacrylate, MAN and low and high boilers. Raffinate stream (28b) was then blended with the waste acid stream (11) and discharged from the process in the form of the resulting stream (27). MAN and acetone were discharged from the process via (28b). The extract stream (28a) contained methyl methacrylate, acetone, methanol, MAN and further low and high boilers, and was recycled continuously into the reaction zone of the esterification (E). Amounts and compositions are described in Table 3.

[0344] The same workup of the bottom product (22b) from the azeotrope column (L) was effected in the same way in Examples A1, A2 and A3.

[0345] The low boiler-free bottom product (22b) obtained in the azeotrope column (L), for further purification, was subjected to a reduced pressure rectification (0) that worked at 180 mbar(a) and had a rectifying section and stripping section. Stream (22b) was applied to the middle of the purifying column (O), and this was separated, in accordance with the equilibrium established, into a pure distillate phase (25a) and a high boiler-enriched bottom stream (25b).

[0346] The bottom stream (25b) that contained methyl methacrylate, methacrylic acid and high boilers was recycled continuously into the esterification reaction step (E).

[0347] The vapour stream obtained was fully condensed in (O). The offgas obtained here at about 2 m.sup.3/h was fed to process step (M). The distillate was divided in accordance with the reflux ratio required, such that the amounts of pure MMA product (25a) reported in Table 3 were obtained with >99.9% by weight purity. The pure MMA product (25a) contained acetone and MAN in the amounts reported in Table 3. MAN and acetone were removed from the process in the MMA product stream.

[0348] 1d. Process Conditions and Results

TABLE-US-00001 TABLE 1 Examples A1-A3 - data for amidation, conversion and esterification ((A), (B), (C), (D), (E)) A1* A2 A3* ACH kg/h 3750 3750 3750 (1b) ACH kg/h 1250 1250 1250 (1a) + (1b) ACH tot. kg/h 5000 5000 5000 (1a):(1b) kg/kg 75:25 75:25 75:25 (1a) + (1b) ACH conc. % by wt. 98.8 98.8 98.8 H.sub.2SO.sub.4 conc. % by wt. 100.3 99.7 100.3 Total H.sub.2SO.sub.4 kg/h 8100 8100 8100 (A) H.sub.2SO.sub.4/ACH kg/kg 2.16 2.16 2.16 (A) + (B) H.sub.2SO.sub.4/ACH kg/kg 1.62 1.62 1.62 Total m.sup.3/h 60 50 60 (9a) Total m.sup.3/h 75 65 75 (9a) Acetone g/m.sup.3 19.0 11.0 19.0 (9a) MAN g/m.sup.3 2.0 1.1 2.0 (9a) Acetone kg/h 2 1 2 (9a) MAN kg/h 0.2 0.1 0.2 (8) Total kg/h 12 914 13 015 12 914 MAA % by wt. 35.6 35.4 35.6 MA % by wt. 0.4 1.0 0.4 (8) HIBAm % by wt. 0.4 0.4 0.4 (8) Acetone % by wt. 0.25 0.2 0.25 (8) MAN ppm 500 354 500 (8) MAA + MA kg/h 4649 4734 4649 (reactants for esterification) (8) + (9a) Acetone kg/h 34 26.5 34 (8) + (9a) MAN kg/h 6.5 4.7 6.5 Amide yield MAA + MA based % 93.0 94.7 93.0 on ACH (29) Total kg/h 590 590 590 (10a) Total kg/h 1540 1570 1540 (16d) Total kg/h 500 500 500 (16a) + (16 c) Water kg/h 1700 1700 1700 MAN MAN decrease kg/h 1.5 1.0 1.5 (hydrol.) in (E) (11) Total kg/h 11 504 11 410 11 515 (11) MAN ppm 110 60 100 (11) TOC % by wt. 2.3 1.5 2.4 (11) Solid-state TOC % by wt. 0.2 0.1 0.2 (11) Sulfo acetone in ppm 470 360 465 TOC (11) Acetone kg/h 5.0 4.0 5.0 (11) MAN kg/h 1.2 0.65 1.1 *comparative example

TABLE-US-00002 TABLE 2 Examples A1-A3 - data for workup/prepurification (F), (G), (H)) A1* A2 A3* (12) Total kg/h 10810 9875 10698 (12) MMA % by wt. 67 74 63 (12) Water % by wt. 22 16 22 (12) MeOH % by wt. 6.5 5 6.7 (12) MA % by wt. 3 3 2.8 (12) Acetone % by wt. 1 0.8 1 (12) MAN % by wt. 0.15 0.09 0.1 (30) Total kg/h 83 74 85 (30) Acetone kg/h 1 0.5 1 (30) MAN kg/h 0.1 0.1 0.1 (14a) Total kg/h 10 105 9080 10 081 (15a) Total kg/h 8156 7,638 8,120 (15a) MMA % by wt. 83.9 88.8 84.1 (15a) Water % by wt. 4.1 2.9 4.1 (15a) MeOH % by wt. 5.4 3.4 5.4 (15a) MA % by wt. 1.8 1.7 1.8 (15a) Acetone % by wt. 2.1 1.1 1.9 (15a) MAN % by wt. 0.4 0.1 0.3 (15b) Total kg/h 1,949 1,447 1960 (18) Total kg/h 3,251 2,583 3261 (18) Water % by wt. 75.0 82.0 74.7 (18) MeOH % by wt. 17.0 12.5 17.8 (18) MMA % by wt. 4.0 3.4 5.0 (18) MA % by wt. 0.4 0.2 0.4 (18) Acetone % by wt. 1.5 0.8 1.4 (18) MAN ppm 470 100 330 (17a) Total kg/h 7,754 7,398 7,719 (17a) MMA % by wt. 89.5 91.2 87.7 (17a) Water % by wt. 2.9 2.3 3.0 (17a) MeOH % by wt. 1.2 1.9 3.0 (17a) MA % by wt. 1.8 1.4 1.8 (17a) Acetone % by wt. 1.9 1.0 1.7 (17a) MAN % by wt. 0.4 0.1 0.3 *comparative example

TABLE-US-00003 TABLE 3 Examples A1-A3 - data for workup/fine purification (L), (O), (M), (N), (P) A1* A2 A3* (22b) Bottoms (L) kg/h // // // (22b) MMA % by wt. 96.0 96.7 95.8 (22b) MA % by wt. 2.0 1.5 2.0 (22b) High boilers % by wt. <2.0 1.8 <2.0 (22b) Acetone ppm 10 <10 9 (22b) MAN ppm 138 37 98 (23a) Condens. kg/h 7184 8238 7282 vapour (23a) MMA % by wt. 61 75 64 (23a) Water % by wt. 8.1 5.4 7.8 (23a) Acetone % by wt. 11.5 7.3 10.6 (23a) MAN % by wt. 8.4 5.6 7.4 (23a) MAN kg/h 603 461 (23d) PK kg/h // // // (23d) Low boilers % by wt. 33 >30 33 (23d) MMA % by wt. 30 40 32 (23d) MeOH % by wt. 15 11 15 (23d) Water % by wt. 1.4 1.4 1.4 (23d) Acetone % by wt. 17.0 11.0 16.2 (23d) MAN % by wt. 5.9 3.0 5.2 (23e) Offgas kg/h 28 23 27 (23e) Acetone kg/h 1.0 1.0 1.1 (23e) MAN kg/h 0.1 0.1 0.1 (24a) Total kg/h 6428 7795 6553 (24a) Water % by wt. 5.5 3.5 5.3 (24a) MeOH % by wt. 6.3 3.6 6.1 (24a) MMA % by wt. 66.7 78.6 70.0 (24a) Acetone % by wt. 10.7 6.9 9.9 (24a) MAN % by wt. 8.9 5.8 8.1 (24a) MAN kg/h 572 452 (24b) Total kg/h 1,355 1,043 1,328 (24b) Water % by wt. 60.5 73.7 61.9 (24b) MeOH % by wt. 18.6 13.8 18.5 (24b) MMA % by wt. 10.5 4.5 10.0 (24b) Acetone % by wt. 11.0 6.4 9.6 (24b) MAN % by wt. 1.9 0.8 1.7 (24b) MAN kg/h 25.7 8 22.5 (26a) Total kg/h 266 (26a) MMA kg/h 26.8 (26a) Acetone kg/h 25.5 (26a) MAN kg/h 2.75 (26b) Total kg/h 271 522 (26c) Total kg/h 1,084 522 1063 (28a) Total kg/h 72 66 (28a) Acetone % by wt. 19 14.5 (28a) MAN % by wt. 3.2 2.0 (28a) MAN kg/h 2.3 1.3 (28b) Total kg/h 252 508 (28b) Water % by wt. 63.0 76.0 (28b) MeOH % by wt. 19.6 14.0 (28b) MMA % by wt. 6.7 2.3 (28b) Acetone % by wt. 9.3 4.0 (28b) MAN % by wt. 1.1 0.5 (28b) Acetone kg/h 23 20 (28b) MAN kg/h 2.65 2.7 (25a) Total kg/h 5453 5534 5440 (25a) MMA % by wt. >99.9 >99.9 >99.9 (25a) Acetone ppm 10 <10 10 (25a) MAN ppm 176 47 118 (25a) Acetone kg/h 0.2 <0.1 0.15 (25a) MAN kg/h 0.96 0.26 0.65 *comparative example, = absent

[0349] The results in Tables 1 to 3 show that, in Inventive Example A2 using sulfuric acid with a concentration of 99.7% by weight (0.3% by weight of water), lower proportions of MAN and acetone are obtained in the amidation and conversion (see stream (8) in Table 1) and are fed into the esterification than by comparison with Examples A1 and A3 using sulfuric acid having a concentration of 100.3% by weight (0.3% by weight of free SO.sub.3).

[0350] In addition to the reduced sulfuric acid concentration, in Inventive Example A2, an elevated amount of aqueous phase from phase separator (N) is discharged (ratio of 26c/26b=50:50 vs. 80:20 in Comparative Example A1).

[0351] Moreover, the inventive procedure according to Example A2 has the effect that the pure MMA product has much lower contamination with MAN than is the case in the comparative example according to the prior art. The proportion of MAN and acetone in the MMA end product (25a) in Example A2 is distinctly reduced compared to Example A1 (see Table 3).

[0352] Table 4 below compiles the streams of matter from Table 3 relating to the enrichment of MAN in steps (L), (M) and (N), and relating to the recycling of MAN and the discharge of MAN via (M) and (P) for Examples A1 and A3 (comparative examples) and A2 (inventive).

TABLE-US-00004 TABLE 4 Overview of selected process data and overview of MAN recycling and MAN discharge in (L), (M) (N) and (P) Example A1* A2 A3* Sulfuric acid conc. (2) % by wt. 100.3 99.7 100.3 (26c)/(26b) ratio kg/kg 80/20 50/50 (26c)/(26a) ratio kg/kg 80/20 Aqueous phase to esterification (recycle) Total kg/h 1,084 522 1,063 (26c) amount Aqueous phase to esterification (recycle) Amount kg/h 20.6 4.0 16.4 (26c) of MAN Total recycle rate into the esterification Total kg/h 1,156 574 1,063 (26c) + (28a) or (26c) only amount Total recycle rate into the esterification Amount kg/h 23.0 5.3 19.8 (26c) + (28a) of MAN Raffinate to cleavage acid (discharge Total kg/h 252 508 after extraction) (28b) amount Discharge of aqueous phase after phase Total kg/h 266 separator II (26a) MAN discharge in (28b) or (26a) Amount kg/h 2.65 2.70 2.75 of MAN Quotient (MAN recycling/MAN kg/kg 8.7 2.0 7.3 discharge) MMA loss in discharge (28b) or (26a) Amount kg/h 16.9 11.6 26.8 of MMA Total MMA yield loss via discharged % 0.31 0.20 0.5 MMA Total MMA yield MMA % 90.3 91.6 90.1 based on ACH *comparative example

[0353] On the basis of the values in Table 4, it is found that, in Comparative Example A1, a greater amount of MAN as troublesome by-product in the process (esterification reactor E) is recycled (via (28a)) and less troublesome MAN can be discharged from the process (via (28b)) compared to Inventive example A2. The quotient of MAN (recycled) to MAN (discharged) is thus much lower in Inventive Example A2 and hence more advantageous than in Comparative Example A1. The loss of MMA via the discharge of aqueous phase (28b) is distinctly increased in Comparative Examples A1 and A3 compared to Inventive Example A2, with a smaller MMA loss through use of the extraction in Comparative Example A1 than in Comparative Example A3. On account of the lower enrichment of acetone and MAN in Inventive Example A2, the MMA loss is at its lowest even though the amount of aqueous phase discharged is greater than in Comparative Examples A1 and A3.

[0354] In addition, it is shown experimentally that the discharged cleavage acid (11) that has optionally been sent to a further utilization is less contaminated with sulfonated acetone (see Table 1).

Example B1 According to FIG. 3 (Variant B)

[0355] The preparation of methyl methacrylate comprising the reaction of acetone cyanohydrin with sulfuric acid in the amidation/conversion (A, B, C, D), the subsequent esterification (E) with methanol, and distillative and extractive workup (F, I, J, K, L, O, M) of the methyl methacrylate product was effected by the embodiment according to FIG. 3 as described above (variant B).

[0356] A mass balance and and assessment of the discharge of MAN and acetone via the total offgas (11), output air streams (21a) and (9b), and via the MMA product (25a) was effected (see flow diagram according to FIG. 3).

[0357] An inventive example (Example B1) using sulfuric acid having a concentration in the region of 99.7% by weight (0.3% by weight of water) is described hereinafter. The example describes the preparation of methyl methacrylate having the claimed features, with achievement of a high overall yield of MMA, characterized by reduced MAN formation in the amidation and conversion. By the exclusively distillative workup of the MMA product (without extraction) and discharge of by-products such as acetone in the offgas, it is possible to achieve a high overall MMA yield with moderate MAN content analogously to Example 2A.

[0358] The performance of the process according to the invention as per FIG. 3 (variant B) is described hereinafter.

[0359] 2a. Reaction Stages (A), (B), (C), (D) and (E)

[0360] The amidation (A) and (B), the conversion (C) and gas separation (D) were performed as described above in Example 1, with the differences specified hereinafter: [0361] The gas space of the amidation reactors (A) and (B) was operated at about 970 mbara, at slightly reduced pressure. [0362] The intermediate vessel (D) was operated at a temperature of 100 C. The mixture (7) was cooled in the region of the vessel (D) by means of a water-cooled heat exchanger through which an amide circulation stream connected to the vessel (D) flowed. [0363] Suitable interconnection of the reactant streams (10a, 10b, 16b, 16d, 29) and of the circulation streams (13, 20c, 21b, 25b) achieved a local stoichiometric excess of methanol and water based on methacrylamide and methacrylic acid in each of the six esterification tanks. [0364] An overall offgas (9b) was obtained in the gas separator (D), which contained acetone and methacrylamide and was discharged continuously from the process.

[0365] The amounts and compositions are collated in Table 5.

[0366] 2b. Prepurification (F), (I), (J)

[0367] According to the vapour/liquid equilibrium of the reaction mixtures, the vapour stream (12) left the esterification with a heteroazeotropic composition comprising MMA, water, MeOH, MA, acetone and MAN as reported in Table 6.

[0368] Vapour stream (12) was subsequently subjected to a countercurrent distillation (F), wherein the vaporous stream (12) was applied in the bottom region of a column (F). The distillation (F) was operated at a slightly elevated pressure of 100 mbar(g).

[0369] Vapour stream (12) was partially condensed at the top of the column, and an organic, liquid side stream (19c) from column (I) was applied as subcooled reflux. By this procedure, a methacrylic acid-containing bottom stream (13) was obtained, which is recycled directly into the esterification (E). The bottom stream contains MMA. MA and further high boilers according to Table 6. At the top of the column (F), a vaporous heteroazeotropic vapour stream (14b) was obtained, which was sent to further distillative separation in the middle of a rectification column (I). The heteroazeotropic stream (14b) contained MMA, MeOH, water, acetone and MAN according to Table 6.

[0370] The primary stripper column (I) was heated with hot steam and operated at a slightly elevated pressure of 100 mbarg. A heteroazeotropic mixture of methyl methacrylate, water, methanol, acetone and MAN that has been depleted of low boilers in the bottom stream (19b) was separated from a low boiler-enriched top stream (19a). It was possible to control the low boiler content and hence the content of acetone and MAN among others by the energy input in the evaporator.

[0371] At the top of the column (I), the vapour stream obtained was partially condensed, the condensate was returned to column (I) as liquid reflux, and the remaining low boiler-containing vapour phase (19a) that comprised MeOH, MMA, water, acetone and MAN, and also further low boilers according to Table 6 was sent to an offgas scrubbing (J). The enrichment of low boilers in the top region of the column (I) was controlled here via the temperature in the partial condenser.

[0372] A portion of the liquid low-boiling mixture at the top of column (I) was removed continuously from the column (I) in the form of a liquid side stream draw (19c). Side stream draw (19c) contained MMA, MeOH, water, acetone and MAN according to Table 6 and, in subcooled form, served as reflux for scrubbing column (F).

[0373] The vapour stream (19a) was then fed together with offgas (23b) from the azeotrope column (L) to the bottom region of an offgas scrubbing column (J) that had a partial condenser at the top. The vapours fed in (19a) ascended within column (J) and were scrubbed by the high boiler-containing reflux formed at the top by condensation, so as to obtain an MMA-depleted offgas (21a) at the top of the column (J) and a liquid, MMA-enriched mixture (21b) at the bottom of column (J). The scrubbing process in (J) was supported by addition of fresh methanol (10b) that was applied at the top of column (J). The bottom stream (21b) from (J), comprising MeOH and MMA, was sent to the esterification (E). The offgas (21a) comprised low boilers and inerts, MeOH, acetone and MAN, and was sent to an incineration. Amounts and compositions are reported in Table 6.

[0374] The heteroazeotropic bottom stream (19b) obtained in the primary stripper (1), for separation of the liquid phases, was first cooled to 30 C. and then sent to the phase separation vessel (K). In addition to stream (19b), in the phase separator (K), an MMA-containing heteroazeotropic distillate (23c) was sent to the condensation/vacuum unit (M). The aqueous heavy phase (20b) contained water, methanol, MMA, acetone, MA and high and low boilers according to Table 6, and was sent continuously to a mixing zone. Here, the addition of deionized water (16b) afforded a diluted stream (20c) that was fed fully to the esterification (E). The organic light phase (20a) contained MMA, water, methanol, acetone, MA, MAN and high and low boilers according to Table 6, and was sent to the fine purification in the azeotrope column (L).

[0375] 2c. Fine Purification (L), (O), (M)

[0376] The organic light phase (20a), for removal of low boilers, was sent to the top region of an azeotrope column (L) operated under reduced pressure, which was indirectly heated with hot steam and separates a low boiler-enriched heteroazeotropic vapour stream (22a) from a methyl methacrylate-enriched bottom stream (22b).

[0377] The vapour stream (22a) contained MMA, water, MeOH, acetone and MAN according to Table 7. The bottom stream (22b) that was freed of low boilers in the desired manner contained MMA, MA, high boilers and the MAN and acetone by-products according to Table 7.

[0378] The bottom product (22b) obtained in the azeotrope column (L), for further purification, was subjected to a reduced pressure rectification (O) that was performed at 180 mbar(a) and was characterized by a rectifying section and stripping section. Stream (22b) was applied to the middle of the purifying column (O), which gave, in accordance with the equilibrium established, a pure MMA product (25a) and a high boiler-enriched bottom stream (25b). The resultant vapour stream was fully condensed, the resultant offgas of 4 m.sup.3/h was sent fully to process step (M), and the distillate was divided in accordance with the required reflux ratio such that, based on the esterification line (E), the amount of MMA with >99.9% purity specified in Table 7 was obtained. The proportion of acetone and MAN, and the amounts discharged, are stated in Table 7.

[0379] The resultant bottom stream (25b) containing MMA, MA and high boilers according to Table 7 was returned continuously to reaction step (E).

[0380] 2d. Process Conditions and Results

TABLE-US-00005 TABLE 5 Example B1 (inventive) - data for amidation, conversion and esterification ((A), (B), (C), (D), (E)) (1a) ACH kg/h 6,845 (1b) ACH kg/h 3,685 (1a) + (1b) ACH tot. kg/h 10,530 (1a):(1b) kg/kg 65:35 (1a) + (1b) ACH conc. % by wt. 99.0 (2) H.sub.2SO.sub.4 conc. % by wt. 99.7 (2) Total kg/h 17,060 (A) H.sub.2SO.sub.4/ACH kg/kg 2.55 (A) + (B) H.sub.2SO.sub.4/ACH kg/kg 1.62 (5b) Total m.sup.3/h 85 (9b) Total m.sup.3/h 105 (9b) Acetone g/m.sup.3 15 (9b) MAN g/m.sup.3 1 (9b) Acetone kg/h 2 (9b) MAN kg/h 0.1 (8) Total kg/h 27,475 (8) MAA % by wt. 35.7 (8) MA % by wt. 0.9 (8) HIBAm % by wt. 0.2 (8) Acetone % by wt. 0.2 (8) MAN ppm 322 (8) Acetone kg/h 53 (8) MAN kg/h 8.6 (8) MAA + MA (reactants for kg/h 10,080 esterification) Amide yield MAA + MA based on ACH % 95.7 (29) Total (MMA/MeOH) kg/h 700 (10a) Total (MeOH) kg/h 3,300 (16d) Total (direct steam) kg/h 550 (16b) Water kg/h 3450 MAN (hydrol.) MAN decrease in (E) kg/h 4.8 (11) Total kg/h 23570 (11) TOC % by wt. 3.2 (11) Solid-state TOC % by wt. 0.2 TOC/(11) Sulfo acetone ppm 265 TOC/(11) MAN ppm 120 (11) Acetone kg/h 7 (11) MAN kg/h 2.9

TABLE-US-00006 TABLE 6 Example B1 (inventive) - data for workup/prepurification (F), (I), (J)) (12) Total kg/h 32304 (12) MMA % by wt. 87.9 (12) Water % by wt. 5.3 (12) MeOH % by wt. 3.6 (12) MA % by wt. 1.3 (12) Acetone % by wt. 0.6 (12) MAN ppm 52 (14b) Total kg/h 23445 (14b) MMA % by wt. 73.9 (14b) Water % by wt. 10.3 (14b) MeOH % by wt. 13.8 (14b) Acetone % by wt. 1.5 (14b) MAN ppm 105 (19a) Total kg/h 1,684 (19a) MeOH % by wt. >60.0 (19a) MMA % by wt. 25.0 (19a) Water % by wt. 2.0 (19a) Acetone % by wt. 9.3 (19a) MAN ppm 179 (19c) Total kg/h (19c) MMA % by wt. 60 (19c) Water % by wt. 10 (19c) MeOH % by wt. 28 (19c) Acetone % by wt. 2 (19c) MAN ppm 200 (21a) Total kg/h 220 (21a) Low boilers/inerts % by wt. >60 (21a) MeOH % by wt. 20 (21a) Acetone % by wt. 18 (21a) MAN ppm 330 (21a) Acetone kg/h 37 (21a) MAN kg/h 0.1 (20b) Total kg/h 1,833 (20b) Water % by wt. 88.0 (20b) MeOH % by wt. 8.1 (20b) MMA % by wt. 2.9 (20b) MA % by wt. 0.02 (20b) Acetone % by wt. 0.03 (20b) MAN ppm 15 (20a) Total kg/h 17,998 (20a) MMA % by wt. 95.7 (20a) Water % by wt. 2.2 (20a) MeOH % by wt. 1.6 (20a) MA % by wt. 0.1 (20a) Acetone ppm 250 (20a) MAN ppm 127

TABLE-US-00007 TABLE 7 Examples B1 (inventive) - data for workup/fine purification (L), (O), (M) (22a) Total kg/h 5,814 (22a) MMA % by wt. 88.6 (22a) Water % by wt. 6.8 (22a) MeOH % by wt. 5.0 (22a) Acetone % by wt. 0.08 (22a) MAN ppm 280 (22b) Total kg/h (22b) MMA % by wt. >99.0 (22b) MA % by wt. 0.2 (22b) High boilers % by wt. 0.3 (22b) Acetone ppm <10 (22b) MAN ppm 64 (23c) Total kg/h ~5780 (23b) Total offgas kg/h (25a) Total pure MMA kg/h 11,390 (25a) MMA % by wt. >99.9 (25a) Acetone ppm <10 (25a) MAN ppm 70 (25a) Acetone kg/h <0.1 (25a) MAN kg/h 0.8 MMA yield MMA based on % 91.1 ACH (25b) Bottoms kg/h 800 (25b) MMA % by wt. 87.0 (25b) MA % by wt. 4.0 (25b) High boilers % by wt. 9.0

[0381] The results in Tables 4 to 6 show that, in Inventive Example B1 using sulfuric acid having a concentration of 99.7% by weight (0.3% by weight of water), a small proportion of MAN and acetone is obtained in the amidation and conversion and is fed into the esterification. In association with an exclusively distillative purification of the crude product and the gaseous discharge of secondary components, the effect of this procedure is that the pure MMA product has much lower contamination with MAN.