MXPA96006159A

MXPA96006159A - Procedure to produce but acrylate

Info

Publication number: MXPA96006159A
Application number: MXPA/A/1996/006159A
Authority: MX
Inventors: Bauer William Jr; Jesaja Venter Jeremia; Tseng Chapman Josefina; Giuseppe Luciano Mirabelli Mario
Original assignee: Rohm And Haas Company
Priority date: 1995-12-15
Filing date: 1996-12-05
Publication date: 1998-01-01

Abstract

An improved process for producing n-butyl acrylate in high yields and high purity, substantially free of acrylic acid, incorporates one or more of the following new process component, in an esterification process, catalyzed with acid, to produce the acrylate of the n-butyl acrylate. -butyl: 1. A hydrolytic recovery component, in which the heavy end adducts, produced during the acid-catalyzed esterification, are hydrolyzed, recovered and recycled as valuable reagents from a hydrolytic recovery unit ("HRU"). 2. A thermal decomposition reactor component, preferably used with the HRU unit, in which additional valuable reagents are recovered and recycled after treatment in the thermal decomposition reactor, and 3. A new distillation component, in which the current The n-butyl acrylate crude is efficiently distilled in an aqueous mode, through an acrylic acid separation column, by supplying the n-butyl acrylate substantially free of acrylic acid and in high yield. The first two components are also applicable to acid-catalyzed processes, which produce the C1-C4 alkyl acrylates. A continuous process, which produces n-butyl acrylate, which incorporates all the new components of the process, is also known

Description

PROCEDURE TO PRODUCE BUTYL ACRYLATE The present invention relates to an improved process for producing butyl acrylate. More specifically, the invention relates to a new method of distillation and recovery and recycling of normal butanol ("BuOH"), acrylic acid ("AA") and normal butyl acrylate ("BA") from a or more process streams, in an esterification process, catalyzed with acid, for the BA. The invention includes two new process components, one related to the hydrolytic recovery of valuable reagents from their higher boiling adducts, and a second component, related to the improved distillation of a crude product that supplies BA, substantially free From the a. The hydrolytic recovery component of the invention is useful in processes for producing selected acrylic esters, in addition to BA. More specifically, the invention relates to a continuous, highly efficient process for producing BA with high purity and high yield.

The direct esterification of AA with an alcohol is an equilibrium process. The equilibrium constant determines the regime and net extent of the conversion of AA and alcohol; for continued high conversion rates, the mixture should not approach equilibrium. Conventionally, a excess of alcohol on the AA is used and the esterification water is removed distillatively as its azeotrope with alcohol and ester, to maintain a high conversion rate of AA. The azeotrope is removed by means of a distillation column, mounted directly on an esterification reactor. In the case of methyl or ethyl esters, the esterification water, the excess alcohol and the product ester are removed from the head of the distillation column and are substantially free of AA. The extraction of the water removes the alcohol, which is concentrated distillatively for recycling to the reactor. The washed ester is azeotropically dehydrated and finally distilled to deliver the pure ester product. However, in the production of butyl acrylate, the separation of acrylic acid from the water of reaction, excess of alcohol and product ester is more difficult, and the distillation of the esterification reactor, in a continuous process, typically contains 1-3 % From aa. This AA is typically extracted in an aqueous caustic solution. Although it is possible to recover some of this AA from the resulting solution of aqueous salt by acidification with a strong acid, followed by extraction in an organic solvent, for example butyl acrylate or butyl acrylate / butanol mixture, no it can avoid a significant loss to a large aqueous waste stream. Butyl acrylate and excess butanol are immediately dehydrated azeotropically, where the excess butanol is separated from the ester of the product as a buta-nol azeotrope / butyl acrylate for recycling to the esterification reaction. A final distillation supplies the pure butyl acrylate. In all cases, a small bleed stream is removed from the esterification reactor and a small stream of bottoms is taken from the distillation of the final product to remove the high boiling byproducts and the inhibitor residues from the process. These streams are separated to recover the values of free AA, alcohol and alkyl acrylate, but little or none of the values present within the high-boiling by-products are recovered. Thus, conventional procedures for producing C ± -c esters suffer from yield losses to high boiling byproducts, and the C4 process also suffers from direct losses of AA, due to the difficulty of separating AA from butanol, water and ester. In the technique of recovering and recycling reagents from their major boiling adducts, formed during the process (called "heavy ends", in the production of BA, which include, for example, butyl β-butoxypropionate and sulfuric acid esters) , it only has limited success. For example, in the production of ethyl acrylate from ethylene and AA, U.S. Patent No. 4,968,834 ('834) describes a process for recovering EA from a "spent black acid" stream, containing sulfuric acid residues and other adducts bled from the bottom from a distillation column. The '834 process uses an alcohol solvent to facilitate the superior distillative recovery of ethyl acrylate, and treats the black acid residues with an aqueous alkanol mixture. No material is returned directly to the reactor that produces the EA or to the distillation column, which generates the black acid stream. The '834 process thus supplies the partial recovery of ethanol, EA and AA, but only by an aqueous treatment, which is isolated from the ethylene and AA process reactor. Other processes employ distillation units (often referred to as "bleeding separators") to partially recover the free AA, BA, and BuOH from reaction bleeds, but to the extent that the heavy ends, recovered in that operation, remain chemically the form of boiling greater (heavy end) and not transforming to the forms of the desired values of AA, BA and BuOH.

Distillation is commonly used in the production of BA. For example, the patent of E. U. A., No. 4,012,439 ('439) describes a continuous process for the BA, in which an esterification mixture of the reactor is distilled through a column of separation of the AA, to give a superior mixture of BA, butanol and water and, from the bottom of the column, a stream of concentrated AA, which is returned to the reactor. While the top AA mixture is separated, the '439 process recycles a very high proportion (> 97%) of the distillate from the aqueous phase to the top of the AA separation column. This high proportion of aqueous recycle (ie, having an aqueous reflux ratio of about 32: 1), disadvantageously requires a large column and large energy expenditures in returning large volumes of water to the process. Thus, in the acid-catalyzed production of the alkyl esters of acrylic acid ("alkyl acrylates"), particularly BA, the problems of significant energy use and recovery of the reactants remain. There are needs for a process which recovers the reagents of its adducts, heavy end, high boiling point, formed during the production of acrylic esters, for example the BA, which will recycle the recovered reagents and the ester to the esterification reactor or to another site in the process for reuse. In addition, it is necessary to include methods that make the use of water more efficient, both to facilitate the distillative separation of the acrylic ester of AA and the recovery efficient and recycled AA without reacting, particularly if these stages are achieved with the reduced use of energy. Compliance with one or more of these needs will provide increases in the process and / or efficiencies in the use of the material. Additionally, if such improved processes lead to a reduced by-product of dibutyl ether (BE), compared to known processes, this will result in even greater process efficiency. We have discovered a high performance process for producing alkyl acrylates, which use BA as a preferred example, which achieves these desired ends. Our new process provides the recovery of "values", that is, reagents and the alkyl acrylate product, from the heavy ends produced in the process. Our new process includes the use of at least one of the following process components: 1) recovery of values from the hydrolysis reactor unit ("HRU"), fed with a source of heavy ends, such as from an esterification reactor; 2) recovery of additional values from a thermal decomposition reactor, preferably used in conjunction with the hydrolysis reactor; and 3) specific to a continuous process of BA, distillation by the use of a separation column of acrylic acid in a new and efficient way and the supply of recovery.

BA, which is substantially exempt from AA. Our new process advantageously provides very low levels of DBE in the BA product, because the esterification reactor is operated under moderate conditions of temperature and pressure, and at relatively low levels of the acid catalyst.

Thus, in the broader use of the hydrolytic recovery component of the invention, a method is provided for recovering the AA, a C 1 -C 4 alkyl acrylate and a C 1 -C 4 alkanol from the heavy ends, produced during the production of the acrylate of C1-C4 alkyl, which comprises the steps of: a) feeding a heavy and heavy-weight total charge stream, comprising these heavy ends, water, residual acid catalyst and, optionally, a strong acid, selected from an acid mineral or sulphonic acid, to a hydrolysis reactor, maintained at 90 to 1402C, at a pressure of 50 to 1000 mm Hg, and a residence time of 0.5 to 20 hours, based on the total, aqueous charge stream and organic; b) distilling an upper stream containing the AA, the C1-C4 alkyl acrylate and the C1-C4 alkanol, and the water, from the hydrolysis reactor, while maintaining the concentration of the hydrolysis reactor liquid from the 5 to the 40% by weight of water and at least 1% by weight of acid, this acid comprises the residual acid catalyst and the strong acid, optional; c) condensing the upper current; d) separating from the condensed top stream an organic phase, comprising the C1-C4 alkyl acrylate, the C1-C4 alkanol, and the AA, and an aqueous phase comprising primarily water, the AA and the C1-C4 alkanol; e) remove the separated organic phase; f) recycling the separated aqueous phase to the hydrolysis reactor; and g) removing from the hydrolysis reactor from 20 to 70% by weight, based on the total, aqueous and heavy end loading stream, of a purge stream from the hydrolysis reactor.

Specifically, for the production of BA, a method is provided to recover the AA, butyl acrylate (BA) and n-butanol (BuOH), from the heavy ends produced during the acid-catalyzed esterification of AA with BuOH, this method comprises the steps of: a) feeding a total, aqueous and extreme charge current heavy, comprising the AA, BA, BuOH, water, heavy ends, residual acid catalyst and, optionally, a strong acid, selected from a mineral acid or sulphonic acid, to a hydrolysis reactor, maintained at 90 to 1400C, at a pressure of 50 to 1000 mm Hg, and a residence time of 0.5 to 20.0 hours, based on the total, aqueous and heavy ends load current; b) distilling an upper stream, containing the AA, BA, BuOH and water, from the hydrolysis reactor, while maintaining a concentration of the hydrolysis reactor liquid from 5 to 40% by weight of water and at least 1% by weight of acid, this acid comprises the residual acid catalyst and the optional strong acid; c) condensing the upper current; d) separating from the condensed top stream an organic phase comprising BA, BuOH and AA, and an aqueous phase comprising primarily water and AA and BuOH; e) remove the separated organic phase; f) recycling the separated aqueous phase to the hydrolysis reactor; Y g) removing from the hydrolysis reactor from 20 to 70% by weight, based on the total, aqueous and heavy-end charge stream, of a purge stream from the hydrolysis reactor.

Another embodiment of the invention provides a method for continuously recovering the AA, n-butyl acrylate (BA) and n-butanol (BuOH), from the heavy ends, produced during acid catalyzed esterification of AA with BuOH, which it includes the stages of: a) continuously withdrawing a purge stream from the reactor from an esterification reactor, which contains an esterification reaction mixture, comprising the AA, BA, BuOH, water, heavy ends, and residual acid catalyst, while concurrently distilling the AA, BA, BuOH and water from the esterification reaction mixture; b) feeding a total, aqueous and organic charge stream, comprising the purge stream of the reactor, water, optionally a strong acid, selected from a mineral acid or sulfonic acid and, optionally, additional heavy ends, to a reactor hydrolysis, maintained at 90 to 1402C, at a pressure of 50 to 1000 mm Hg, and a residence time of 0.5 to 20 hours, based on the total, aqueous and organic loading stream. c) distill a higher stream, which contains AA, BA, BuOH and water from the hydrolysis reactor, while maintaining a concentration of the hydrolysis reactor liquid from 5 to 40% by weight of water and at least 1% by weight of acid, this acid comprises the residual acid catalyst and the optional strong acid; d) condensing the upper current; e) separating from the condensed top stream an organic phase, comprising BA, BuOH and AA, and an aqueous phase comprising primarily water, and AA and BuOH; f) remove the separated organic phase; g) recycling the separated aqueous phase to the hydrolysis reactor, and h) removing from the hydrolysis reactor from 20 to 70% by weight, based on the total, aqueous and organic loading stream, of a bleed stream from the hydrolysis reactor.

The additional recovery of valuable reagents from the heavy ends is achieved, using a tandem thermal decomposition reactor, with the hydrolytic recovery methods, described above. This process is carried out with any of the above described hydrolytic recovery methods, including the steps of: a) feeding up to 100% of the bleed stream from the hydrolysis reactor to a thermal decomposition reactor, maintained at 90 to 1400C , a pressure of 20 to 200 mm Hg and a residence time of 0.5 to 20 hours, based on the purge current of the fed reactor; b) distill from the thermal decomposition reactor a charging current, comprising the AA, the alkyl acrylate -1-C4, the alkanol ^ -04 and the water, while maintaining the concentration of the liquid of the thermal decomposition reactor in the minus 7.5% by weight of acid; c) condensing the upper stream of the thermal decomposition reactor; Y d) recovering from the charge current of the thermal decomposition reactor, the AA, C1-C4 alkyl acrylate, C1-C4 alkanol and water.

Preferably, the alkyl acrylate is BA.

More preferably, the thermal decomposition reactor, just described, is used in tandem with the hydrolytic reactor, in a continuous process catalyzed with acid, to produce the BA.

In a second component of the invention, related to the continuous production of the BA, a method for the continuous recovery of the n-butyl acrylate (BA), substantially free of AA, from an esterification reaction mixture, comprising The stages of: a) continuously feeding an esterification reactor, AA and BuOH, in a molar ratio of 1 to 1.1 to 1 to 1.7, and an acid catalyst; b) reacting the AA and the BuOH to supply the BA in a conversion of at least 60% in AA, and supplying the esterification reaction mixture comprising the AA, BA, BuOH, water, heavy ends and acid catalyst; c) distilling from the esterification reactor a vaporized mixture comprising AA, BA, BuOH and water; d) condensing the vaporized mixture to supply a first condensate comprising an organic phase and an aqueous phase; e) returning from 0 to 30 percent of the organic phase to a stripping separator over the esterification reactor; Y f) feeding 70 to 100 percent of the organic phase and 50 to 100 percent of the aqueous phase, to a separation column of acrylic acid; g) distilling from the acrylic acid separation column, at a pressure of 35 to 800 mm Hg, in an aqueous mode and at an aqueous reflux ratio of 8.5: 1 to 17: 1, a superior mixture comprising an azeotroped mixture - butanol of butanol, butyl acrylate and water; h) removing from the distillation column a bottom stream rich in acrylic acid; i) recycling the acrylic acid-rich bottom stream from the acrylic acid separation column to the esterification reactor; j) condensing the upper mixture to supply a second condensate; k) separating this second condensate into an organic phase, rich in butyl acrylate, and a separate aqueous phase; and 1) removing the organic phase rich in butyl acrylate, substantially free of AA. The recovery of BA, substantially free of AA, can also be carried out by feeding the vaporized reactor mixture directly to the AA separation column, bypassing immediately preceding steps d), e) and f). When the vaporized mixture is fed directly to the column, the ratio of aqueous reflux is narrowed to 13: 1 to 17: 1; all other stages are identical, except that, of course, there is no "first condensate". In a brief description of the drawings, the process incorporating both components of the invention is shown schematically in Figure 1; Figure 2 is a graph of the amount of the residual AA in an organic distillate, obtained by distilling through an acid separation column versus the rate of the aqueous reflux flow, as obtained according to the conditions described below. Figure 1 shows the equipment and flow lines, which include the esterification reactor 1, the bleeding line 3 to a unit 5 of the hydrolysis reactor (HRU); and associated streams and lines, particularly line 8 returning the organic phase to reactor 1 and line 7 returning the aqueous phase to the HRU. The thermal decomposition reactor 10 also has associated lines, for example for emptying and distillation, and supplies the return of the condensed upper stream from the separator 31 to the reactor 1 via the line 12. In a brief description of the drawing, relative to the distillation component, Figure 1 includes the line 2 feeding a vaporized mixture of the esterification reactor (in this embodiment) to a condenser 62 and the condensate to a phase separator 1 and the associated lines from the phase separator to the separation column 15 of the acrylic acid, by means of one or more lines 43, 53 and to feed, optionally, some of the aqueous phase to the HRU 5 via the line 42 and, optionally, a portion of the organic phase , to a drag separator on the reactor 1 via the line 41, when the separator 14 is used. The line 54 supplies the optional feed of the BuOH to the separation column of the AA. The lines from the acrylic acid separation column include line 17 returning the AA-rich bottom stream to reactor 1 and line 16 conveying the distilled top mixture through condenser 63 to phase separator 18 and its associated lines, the line 21 returning a controlled portion of the aqueous phase 20 to the top of the separation column 15 of the acrylic acid, the line 22 moving forward a controlled portion of the separated aqueous phase and the line 23 which carries all the organic phase 19 rich in BA, Line 52 supplies the return of the BA / BuOH to the reactor 1, during the subsequent conventional process and the isolation of the final product of BA. Figures 1 and 2 are described in more detail below. Detailed description of the invention The first component of the invention, the hydrolytic recovery component, which recovers values from the heavy ends, takes advantage of the known ability of a strong acid, for example, a mineral acid, such as sulfuric acid, to catalyze the reactions individual employees; direct esterification, ester hydrolysis, dehydration and retro-Michael reactions. Thus, the catalytic processes in which the hydrolysis of the esters and the heavy ends take place, in a reactor of hydrolysis and, in an extended mode, dehydration and retro-Michael additions carried out in a thermal decomposition reactor, are efficient new methods for the recovery, for example, of the values of BA, BuOH and AA, from of the heavy end components, formed during the above reaction in, for this example, a BA esterification reactor. The heavy ends are exemplified in detail for BA production using a sulfuric acid catalyst in the reactor and in the HRU; of these examples, a person skilled in the art will recognize the analogous "heavy end" counterparts of the production of any of the C1-C4 alkyl acrylates. The C1-C4 alkyl groups may be methyl, ethyl, propyl and isopropyl, and the isomers of butyl, preferably n-butyl. The heavy ends are adducts higher in boiling point than the reactants and, as exemplified herein, the butyl acrylate product; they include, for example, acryloxypropionic acid ("AOPA") and its butyl ester derivative, beta-hydroxy-propionic acid and its butyl ester derivative, beta-butoxy-propionic acid and its butyl ester derivative, and other non-polymeric adducts of the reactants. In addition, impurities of maleic acid and benzoic acid are present in the acrylic acid and the sulfuric acid catalyst, such as the monobutyl ester of maleic acid, butyl benzoate and monobute sulphate. linden. Likewise, the simultaneous removal of BA, BuOH and AA, by means of distillation currents of the reactors, both hydrolysis and thermal decomposition, in a continuous process, allows the recovery reactions to proceed beyond the equilibrium restrictions, present in a batch process and thus improves the yields of the process. Another advantage of the hydrolytic recovery component of the invention is that of one or more additional heavy end streams, which can be worked in the stream of the recovery process, thus providing recovery of additional values.

The following are examples of heavy end materials, present in the total, aqueous and organic charge streams (ie the heavy ends alone or a mixture thereof, reactants and product), which are hydrolysed in the hydrolysis reactor to supply valuable recoveries of the AA and the alkyl acrylate and alkanols described. The alkyl esters of the β-alkyloxy propionates are common heavy end materials. In the beta position of the alkyl esters, there may also be a hydroxyl group instead of an alkyloxy group. The β-acryloxy acid derivatives of the C 1 -C 4 alkyl esters can be present in the heavy ends: for example, butyl (β-acryloxy) -propionate is commonly present in heavy end materials, together with their corresponding acid, in the production of BA. Also present are the C1-C4 esters of the sulfuric acid catalyst, these esters are hydrolyzed to the sulfuric acid and the corresponding C1-C4 alkanol. The reactions that take place in the HRU can be generalized by the following equations 1 and 2: (2) Here, R1 is an alkyl group, -04, as defined above; R 2 is an alkyl group, -04 or H. Additionally, saturated and unsaturated esters, such as the C 1 -C 4 alkyl ester of benzoic acid and the C 1 -C 4 alkyl ester of maleic acid, as well as the C 1 -C 4 alkyl sulfate -C4, can be similarly hydrolyzed to release an equivalent alkaline lens ^ -04. Likewise, the simultaneous removal of BA, BuOH and AA by means of the distillation current of the HRU, allows the recovery reactions to proceed beyond the equilibrium limitations and improve the yields of the general process. However, the related carboxylic acid of several heavy materials can not be recovered in the HRU and, therefore, an additional recovery scheme for these materials is necessary and is carried out in a thermal decomposition reactor. The reactions that take place in the thermal decomposition reactor can be generalized by the following equations 3 and 4, where R2 is as defined above.

) The conversion of the C1-C4 alkyl ester of beta-hydroxy-propionic acid, beta-alkoxy-propionic acid and beta-acryloxy-propionic acid to a cognate acid in the HRU (via of ester hydrolysis) is very beneficial, since it is well known that these materials undergo dehydration and the addition of retro-Michael in the acid form. Therefore, under relatively dry conditions of the thermal decomposition reactor, the compounds, such as beta-n-butoxy-propionic acid and beta-hydroxy-propidnic acid, can undergo dehydration, resulting in the recovery of acrylic acid and butanol. The well-known acrylic acid dimer (AOPA) undergoes thermal decomposition to supply 2 moles of acrylic acid. Here, again, the continuous removal of the products allows the reactions to proceed beyond the limitations of the equilibrium and improve the overall performance of the process.

With reference to Figure 1, in the work of the preferred continuous hydrolytic recovery component of the invention when producing BA; the purge stream 3 of the esterification reactor is fed from the esterification reactor 1 to the unit 5 of the hydrolysis reactor ("HRU"). The method of hydrolytic recovery of any embodiment of the invention can be carried out in a multi-plate reactive distillation column or other step reactor, and is preferably carried out under continuous mixing conditions, such as in a reactor. Continuous flow stirred tank ("CSTR"). By "bleeding stream" "purge" means any process stream which is withdrawn in a controlled manner from one vessel to another, such as from one reactor to another or the distillation column Here, the bleed stream 3 of the esterification reactor contains the acid catalyst, water, AA, BA, BuOH, and heavy ends, polymerization inhibitors may be present Additional charges by way 4 may include water and also mineral acids, for example sulfuric acid or a sulfonic acid, such such as methan-, benzene-, or toluene-sulphonic acid.The mineral or sulphonic acid is added, as required, to meet the minimum use level specified in the HRU.One or more additional streams containing heavy ends of other Sources, in addition to the esterification reactor, can also be added in. These charges can be added by one or more feed lines, represented by 4. The additional heavy ends can be up to 80% by weight of the total, aqueous and organic cagas stream. Sulfuric acid is most preferred for use as both the acid catalyst of the reactor and the mineral acid, in all embodiments of the invention. The mixture of the HRU of the charge streams described is kept in a boiling state under the defined conditions. The residence time is from 0.5 to 20 hours, and is based on the total, aqueous and organic load current ("total" means the sum of water streams and heavy ends and / or reactor purge) fed to the HRU. Preferred residence times are from 0.5 to 5 hours and more preferred from 0.5 to 3 hours. An upper stream is distilled from the mixture of the HRU at 6 and condensed, 60, into a phase separator 30. The condensed charge stream is separated into the organic phase, rich in BA, BuOH and AA, and in an aqueous phase containing primarily (ie> 50%) water and some of the BuOH and AA. The separated aqueous phase is returned to the hydrolysis reactor via 7 and the separated organic phase, as stream 8, is returned in this mode to the esterification reactor 1, thus recovering the valuable BA, BuOH and AA for the subsequent reaction and product recovery. The separated organic phase can also be fed to the separator 14 for recovery via line 43 and distillation through 15. An undistilled residue, from 20 to 70% by weight of the total, aqueous and organic charge stream , is purged as the purge stream 9 from the hydrolysis reactor from this hydrolysis reactor for further handling (eg as a waste stream through line 51 or, preferably, as a charge to the thermal decomposition reactor via line 9. ) A preferred embodiment provides additional recovery of AA, BA and BuOH. As shown in Figure 1, The purge stream 9 of the hydrolysis reactor is fed to a thermal decomposition reactor 10 and treated as described now. The thermal decomposition reactor may be similar in construction to that of the HRU and is preferably a CSTR. The liquid of the thermal decomposition reactor is maintained at least 7.5% by weight of the mineral acid, preferably the sulfuric acid, and also contains a mixture of the acrylic acid, BuOH, BA, some heavy ends and residual polymerization inhibitors. The additional acid, mineral or eulphonic, can be added to the liquid of the thermal decomposition reactor (feed line not shown). The thermal decomposition reactor mixture is kept in a boiling state under the thermal decomposition conditions, previously described, while an upper stream is distilled from the thermal decomposition reactor by means of line 11 and condensed via 61 to the separator 31. The condensate contains a stream of organic distillate containing AA, BA and BuOH, and also some water; all the condensed upper stream is returned as the stream 12 to the esterification reactor, thus providing an additional recovery of the valuable AA, BA and BuOH. The stream 13 of the waste from the thermal decomposition reactor is drained for further handling, generally as waste. The most preferred residence times of the thermal decomposition reactor are the same as described for the HRU, that is, from 0.5 to 5 hours, and 0.5 to 3 hours, respectively.

The HRU may be a reactive distillation column of multiple plates, as long as a sufficient number of plates are incorporated to provide a specific residence time. When a reactive distillation column is used as an HRU. A separate thermal decomposition reactor unit may not be necessary to achieve acceptable value recoveries. Under most production conditions, it is preferred to use the tandem thermal decomposition reactor, with a hydrolytic reactive distillation column, similar to its use when the HRU is a CSTR. A disadvantage of the reactive distillation column on a CSTR is that the occasional accumulation of solids in the trays of the column may require an inactive time inconvenient for the cleaning of the column.

The addition of one or more additional charge streams to the purge stream of the esterification reactor or directly to the hydrolysis reactor allows the additional recovery of the AA and, for example, the BA and BuOH, through the processes that occur in the hydrolysis reactor and, when used, the thermal decomposition reactor. The liquid in the hydrolysis reactor has at least 5% water for efficient operation; preferably, the liquid HRU contains from 9 to 18% by weight, more preferably from 10 to 16% by weight of water, in order to achieve efficient hydrolysis regimes under the conditions of nominal temperature and pressure and the size of the practical equipment. The water content is maintained by a combination of the return of all the aqueous stream, condensed and separated, on line 7, to the hydrolysis reactor and adding additional water from other sources, for example by lines 4 and 42, to compensate the water losses to the organic distillate and the purge stream of the HRU. The addition of water from the distilled aqueous phase from the esterification reactor, by line 42, is a preferred source of water in a continuous BA process. In order to maintain efficient dehydration and regimes of the retro-Michael reaction in the thermal decomposition reactor, the thermal decomposition reaction mixture must have a lower aqueous content than that of the HRU mixture. Water contents typically below 5% by weight, preferably below 1% by weight, are achieved by operating the thermal decomposition reactor as a single-stage unit, i.e. by continuously distilling from the thermal decomposition reactor any water carried of the purge stream of the hydrolysis reactor and any additional water generated from the thermal decomposition reactions.

Additional acid can be added to the recovery units, as necessary, to achieve practical reaction regimes; the acid is preferably added by means of one or more charge streams. The "residual acid catalyst" is the acid catalyst which is present as an acid in the purge stream of the esterification reactor and is thus brought to the HRU. In this HRU, the concentration of the acid is preferably in the range 3.5 to 15% by weight and more preferably 5 to 8%. The concentration of acid in the thermal decomposition reactor is typically in the range of 7.5 to 20% and may be higher, for example up to 50%. The acid concentration is preferably 10 to 13% by weight, particularly for the production of BA. The amount of heavy ends in the purge stream of the esterification reactor may vary, but typically is in the range of 10 to 50% by weight, of the combined total charge stream containing the aqueous and organic phases.

The temperatures of the hydrolysis reaction ranges from 90 to 140 QC and are preferably from 105 to 125 £ > C, for efficient regimens of hydrolysis; temperatures greater than 1402C can lead to a thermally induced polymerization of the alkyl acrylates and the heavy ends bearing acryloxy, resulting in a loss of the product not desired. The residence time required for the hydrolysis reaction of the HRU is preferably 0.5 to 5 hours, more preferably 0.5 to 3 hours, shorter times are more economical. The lower temperatures and the presence of water also favor the reduced formation of DBE. Thermal decomposition reactor temperatures range from 90 to 140SC, preferably from 110 to 1252C; pressures for thermal decomposition typically range from 20 to 200 mm Hg, although higher pressures, up to 800 non Hg, can be used. The residence time for dehydration and other reactions in the thermal decomposition reactor, under these conditions, is preferably 0.5 to 3 hours. For continuous BA production, the recovery of values is maximized with two tandem CSTR reactors, one the HRU and the other a thermal decomposition reactor. In order to prevent polymerization, an effective amount of one or more polymerization inhibitors can be added at any stage in any component of the process. A process stream of the esterification reactor traditionally contains sufficient inhibitor to prevent polymerization in the HRU and the thermal decomposition reactor. If additional inhibitor is required, any of a large number of inhibitors should be added.

Known hardeners can be used, for example, hydroquinone, monome-il-ether of hydroquinone, butylated hydroxy-anisole, naphthoquinone, anthranil and derivatives thereof. The second component of the invention, the distillative component, further improves the known methods of distillation of the crude BA and supplies this BA substantially free of AA for the more efficient handling of the distillate and the aqueous reflux. Specifically, the new distillation method supplies BA in the BA rich stream that contains less than 2,000 ppm of AA to move forward to subsequent conventional insulation. The method also supplies a recycled stream of AA containing a negligible amount of BA, which specifically provides a recycled stream of AA (the bottom, AA-rich phase), which contains less than 10 ppm, preferably less than 5 ppm of BA. In the generation of the crude BA for the new distillative component of the invention, AA and BuOH are fed initially, line 70, together with the acid catalyst, to an esterification reactor, in a molar ratio of AA to BuOH in the range from 1: 1.1 to 1: 1.7, preferably from 1: 1.25 to 1: 1.45 and reacted to an AA conversion of 60 to 95%, preferably 75 to 85%, using an acid catalyst of the mineral or sulfonic acid type, previously described, or a strong acid ion exchange resin; Sulfuric acid is preferably used. The reagent ratio and BA conversion supplies a stream of crude BA, which can be processed to provide the stable "water mode" operation (discussed more fully below) of the acrylic acid separation column. The contents of the reactor are kept in a boiling state during the continuous distillation of the vaporized mixture of AA, BA, BuOH and water.

Referring to Figure 1, the mixture vaporized by line 2 from reactor 1, is condensed, 62, and is fed to a phase separator 14 (in this embodiment) to supply the condensed fineness. Alternatively, the vaporized mixture can be fed directly to the column 15 for distillation, as described above. A drag separator, or shown, can also be mounted in the reactor to reduce or eliminate the entrainment of the acid catalyst in the vaporized mixture, thus reducing the downstream corrosion potential. The phase separator 14 is particularly useful when a drag separator is employed, ensuring that an organic reflux layer returns to the drag separator, and also as an element for supplying an optional aqueous stream 42 in the HRU. The first condensate comprises an organic phase primarily (ie, more than 50%) of BA and BuOH, with some AA, and an aqueous phase primarily of water, with some BuOH and AA. All the phases can be fed to the separation column 15 of acrylic acid by one or more lines, for example 43, 53 or, optionally, up to 50% by weight of the aqueous phase can be diverted by line 42 to the unit 5 of hydrolytic recovery (when it is preferably used). The additional butanol can, optionally, be fed to the column via line 54. A superior azeotropic mixture is distilled from the acrylic acid separation column, under the conditions of pressure, temperature and aqueous reflux, previously described, and condensed by line 16 and condenser 63 in phase separator 18, providing a second condensate comprising an organic phase 19 rich in BA and an aqueous phase 20. By "rich in BA" or "rich n AA" it is understood that the BA or AA is the primary organic component (> 50% by weight) of a given phase. Concurrently, a bottom stream rich in AA, containing a negligible amount of BA, is removed from the bottom of the acrylic acid separation column and line 17 returns to the esterification reactor 1. The amount of the recycled aqueous stream 21 is adjusted to provide at least a minimum ratio of 8.5: 1 of aqueous reflux in the AA separation column 15, in order to maintain the column in "aqueous mode" operation. In this operation in aqueous mode, the AA separation column performs a surprisingly effective separation of the AA from the charging stream containing BA (ie, the first condensate stream or the vaporized mixing stream fed to the column), resulting in losses low AA in the BA distillate and consequently higher BA performance, as shown in more detail below. A small portion, from 6 to 11% by weight, of the separated aqueous phase 20, is typically fed forward as the stream 22, together with the feed in advance of the organic Lase 19 rich in BA in stream 23 for subsequent conventional isolation of the final product of BA. Conventionally, the aqueous reflux ratio is defined as the ratio of the aqueous flow returned to the aqueous flow to the next, here the ratio of the aqueous flow in 21 to that in 22. Maintenance of the specified ratio is critical to the efficient operation of the water flow. separation column of acrylic acid in the invention. The separation column of the acrylic acid may have from 20 to 50, preferably from 30 to 40, trays and is typically equipped with a bottom remelting cycle (not shown) and an upper distillation line 16 through the condenser 62 to the separator phase 18. The first condensate charge is typically fed in approximately The tenth tray in a column of 40 trays, numbered from the bottom of the column. If used optionally, the added BuOH is typically fed in the eighth to ninth trays. The column operates within the limits previously described, and preferably at a pressure of 90 to 135 mm Hg, which corresponds to a background temperature of 80 to 852C. The ratio of the aqueous reflux during the distillation of the top mixture is preferably 8.5 to 12.5 and more preferably 9.5 to 10.5. The flow rate of the bottom stream of the column at 17 is adjusted to exceed the amount of AA in the column load by 5 to 25% by weight, to ensure that all the AA remains at the bottom of the column . Stream 17 typically contains 5 to 20% water, the remainder being primarily AA and AOPA. The acrylic acid separation column operates as described, provides a substantial BA free of AA (<2,000 ppm) and a bottom stream of AA containing a negligible amount (<10 ppm) of BA.

One of the unexpected findings in the model and subsequent demonstration of the use of the separation column of acrylic acid is that the two stable states exist under the same operating conditions (ie, at the same load rate, feed composition, regime of water reflux flow and bottom flow regime). A Stable state, previously referred to as the "aqueous mode", is critical in obtaining very low levels of A in the BA-rich phase and BA in the AA-rich background current, as previously described. In the aqueous mode the acrylic acid separation column is relatively "cold", there are substantial amounts of water in the liquid in all the trays, the water is present in the bottom stream, and there is a negligible amount of BA in the stream background. However, in surprising contrast, there is a second mode, the "organic mode", the same conditions (ie, the same loading regime, charge composition, water reflux flow rate and bottom flow rate). which is inconvenient. In the organic mode, the acrylic acid separation column goes for 30-35CC more customer than in the aqueous mode, considerable amounts (> 10% by weight) of the BA are in the bottom stream, and the AA concentration in the upper blend of the BA is at least one order and magnitude greater than the maximum of 2,000 ppm of AA, achieved through the operation of the aqueous mode. Also in the inconvenient organic mode, the column is not only hotter than in the aqueous mode, but all the water is concentrated in the upper part of several trays and the bottom stream is substantially dry. Examples 1 to 6 and the model studies described below provide further details of the unexpected finding of these modes and the rational bottom operation of the acrylic acid separation column, as defined.

Finally, a more preferred continuous process, which employs all the components of the invention in combination, is provided to produce the BA substantially free of acrylic acid (AA) and the AA, BA, n-butanol (BuOH) recovered, water, ends heavy and acid catalyst, which comprises the following steps: a) feeding to an esterification reactor AA and BuOH, in a molar ratio of 1 to 1 to 1 to 1.7, and the acid catalyst; b) reacting the AA and the BuOH to supply the BA in a conversion of at least 60% in AA, and supplying the esterification reaction mixture comprising the AA, BA, BuOH, water, heavy ends and acid catalyst; c) withdrawing a purge stream from the reactor from the continuous conversion esterification reactor mixture, while concurrently distilling the AA, BA, BuOH and water from this esterification reaction mixture; d) feeding a total, aqueous and organic charge stream, comprising the reactor purge stream, water, optionally a strong acid, selected from a mineral acid or an eulphonic acid, and, optionally, adding the additional heavy ends to the hydrolysis reactor, maintained at 90 to 1402C, a pressure of 50 to 1,000 mm Hg and a residence time of 0.5 to 20 hours, based on the total, aqueous and organic loading stream; e) distill a higher stream, which contains AA, BA, BuOH and water from the hydrolysis reactor, while maintaining a concentration of the hydrolysis reactor liquid from 5 to 40% by weight, water and at least 1% by weight of acid, this acid comprises the acid catalyst and the strong acid optional; f) condensing the upper current; g) removing from the condensed upper stream an organic phase comprising the BA, BuOH and AA and an aqueous phase comprising mainly water, and AA and BuOH; h) feeding the separated organic phase to the esterification reactor; i) feeding the separated aqueous phase to the hydrolysis reactor; j) removing from the hydrolysis reactor from 20 to 70% by weight, based on the total aqueous and organic charge stream, of the purge stream of the hydrolysis reactor; k) feeding up to 100% of the purge stream from the hydrolysis reactor to a decomposition reactor l thermal, maintained at 90 to 1402C, a pressure of 20 to 200 mm Hg, and a residence time of 0.5 to 20 hours, based on the purge current of the fed reactor; 1) distilling from the thermal decomposition reactor an upper stream comprising the AA, BA, BuOH and water, while maintaining a liquid concentration of the thermal decomposition reactor of at least 7.5% by weight of the acid; m) condensing the upper stream of the thermal decomposition reactor; n) recycling the condensed upper stream of the thermal decomposition reactor, comprising AA, BA, BuOH and water to the esterification reactor; or) distilling, from the esterification reactor, concurrently with the above steps c) haeta n), a vaporized mixture that is purchased. > Cl? A, BA, BuOH and water; p) condensing the vaporized mixture to supply a first condensate comprising an organic phase and an aqueous phase; q) return from 0 to 30 percent of the organic phase to a drag separator, over the esterification reactor; Y r) feed from 70 to 100 percent of the organic phase and from 50 to loo percent of the aqueous phase to a separation column of acrylic acid; s) distill from the acrylic acid separation column, at a pressure of 35 to 800 mm Hg, in an aqueous mode and at an aqueous reflux ratio, from 8.5: 1 to 17: 1, a superior mixture comprising an azeotropic mixture of butanol, butyl acrylate and water; t) removing from the distillation column a background current rich in acrylic acid; u) recycling the acrylic acid-rich bottom stream from the acrylic acid separation column to the esterification reactor; v) condensing the upper mixture to supply a second condensate; w) separating the second condensate into a rich organic phase n butyl acrylate and a separate aqueous phase; Y x) removing the organic phase rich in butyl acrylate, substantially free of AA.

This method, described immediately before, can also be carried out where steps p), q) and r) are "-9 omitted and 100% of the vaporized spray is fed directly to the acrylic acid separation column of step s) and then distilled as described. When the vaporized mixture is fed directly to the column, the ratio of aqueous reflux is narrowed to 13: 1 to 17: 1; all other stages are identical, except, of course, that there is no "first condensate".

In the continuous methods, described immediately above, the acid catalyst can be selected from sulfuric acid, a sulfonic acid, preferably methan-, benzene and toluene-su Lfón 1 co, or an ion-exchange resin of fuotto acid. Sulfuric acid is preferred for its use both as the acid catalyst and as the optionally added mineral acid. A preferred range of pressure to carry out the distillation in the AA separation column is 90 to 135 mm Hg. A preferred aqueous reflux ratio is, again, from 8.5 to 12.5. The total aqueous and organic charge stream can be fed to a hydrolysis reactor, which is a multi-plate reactive distillation column, or, preferably, to a CSTR, as previously described, thus providing the low hydrolytic reaction continuous mixing conditions. The additional posed ends here can z or also comprehend the 80-s by weight of the total, aqueous and organic charge current.

Returning to Figure 1, the currents of 20 and 29 are taken forward in 23 or separately, such as 22 and 23, and the BA product is then isolated by conventional means. Thus, the process from this point forward can be completed conventionally, for example, by feeding the streams 22 and 23 to a separator, where the current is caustic-neutralized and any salt of the resulting AA is extracted by water. The organic phase free of AA is then dehydrated, according to a distillation column, removing the final traces of water. In the next column, the unreacted BuOH is recovered from the top as its azeotrope with the BA, for recycling to the esterification reactor (stream 52) and passing the bottom current, which contains the R? substantially pure and the inhibitors, to a column of final distillation of the product. In this final column, the R? Pure is distilled in the upper part in the form of a convention and a purge stream containing the inhibitors of the process is removed from the bottom for reuse. The representative purity of BA, obtained from the process, just described above, typically exceeds 99.8% ci - - 1 B? .

E j e m p l o s General data Materials: Raw acrylic acid (AA), pure n-butyl acrylate (BA), n-butanol (BuOH) and heavy end streams, were obtained from plant production streams where indicated and with quality / purity indicated. Polymerization inhibitors were used as purchased, at the indicated levels and with the hydroquinone (HQ), the methyl ether of HQ (MEHQ) and the phenothiazine (PTZ) included. The heavy end components in the Examples include the following materials: AOPA, butyl β-butoxy-propionate ("BBBP"), butyl β-hydroxypropionate ("BBHP"), butoxy AOPA ("BAOPA") , butyl maleate and DBE.

Abbreviations: These include, in addition to the already defined, the following terms: additional (ad'l); aqueous (aq.); Comparative (Comp.); Example (Ex.); Figure (Fig.); gram (g); grams per hour g / h); kilograms (kg), hour (s) (h (s)); heavy ends or heavy ends ("heavy"); weight (wt.); pressure of millimeters of mercury (mm Hg); millimoles (mmoles or mm); pounds (lbs); vaporized mixture (vap. > round bottom natraz (flask f.r.), smaller than (<); greater than (>); point (pt.); stable state (e.e). In Figure 2, the data points are abbreviated as follows: the open-box points are in the (mode) water-stable state, the triangles in the organic stable state and the data points within a circle are experimental / example operations, as they are numbered. Analysis: Standard methods were used for water determination; the monomer, BuOH, and residual impurities and heavy end levels were determined by gas / liquid chromatography (GLC) on a Varian Model 3700 chromatograph, using the flame ionization detection. Sulfuric acid determinations were obtained using a pH probe from the Orion Research ion analyzer and the alcohol tetrabutylammonium hydroxide titration substance. Unless stated otherwise, the H2SO4 concentrations given in the examples are those titrated values. The percentages are in% by weight, unless indicated otherwise. Securities Recoveries: The recovered "values" were calculated and measured as follows: representative heavy ends produced, for example, in BA production. Since all the heavy ends related to BA production are ultimately derived from AA and BuOH, the recovery data values were calculated to reflect the recovery of these reagents, although some recovery is in the form of the product BA.

For example, 100 moles of BBBP contain the equivalent of 100 moles of acrylic acid and 200 moles of BuOH. Similarly, 100 moles of BAOPA contain the equivalent of 100 moles of BuOH and 200 moles of AA. The "heavy mixture" (which is a residue not taken into account) is assumed, for the purposes of calculating the weight, is a 1: 1 molar mixture of acrylic acid and BuOH, with a molecular weight of 146 g / mol. The BA monomer contains equivalent molar amounts of AA and BuOH. The so-called "free" values are simply the same values in the "free" form (not incorporated as heavy ends). Below is a list of the representative BuOH and AA values for characterized heavy ends of the exemplified esterification reaction of BA.

Process Performance: The performance of the process was calculated as follows. The AA and BuOH present in any additional stream fed to the HRU were treated in the yield calculations as if they were fresh (i.e., raw material), these AA and BuOH were fed to the esterification reactor. The BA monomer present in the additional streams, fed to the HRU, was treated in the yield calculations as if it were the recycled BA from a downstream separation (i.e., recycle or supplementary streams); that is, no increase in performance was proven for any recycled BA. The yield in AA or BA, therefore, may exceed 100% when the values (as described below) are recovered from the heavy end streams, treated in the HRU and in the HRU / thermal decomposition reactor, as he described. So, in summary, the calculation of the yield is: % BA performance, based on AA = moles of BA (mixed vap.) - moles BA (recycled) - moles BA (additional streams) moles of AA (fresh to the reactor) + .oles AA (additional streams) and% yield of BA, based on BuOH = moles of BA (vap. mixture) - moles BA (recycled) - moles BA (additional streams) moles of BuOH (fresh to the reactor) + moles BuOH (additional streams) Equipment: In the following examples, the HRU 5 was a 4-neck round-bottomed flask, with a capacity of 1 liter, equipped with a stirrer, a water-cooled distillation head, which has an exit door leading to a separator of phase 30, fraction cutter, 250 ml. The HRU is also equipped with charge inlet ports, 3 and 4, for the purge or bleeding stream of the reactor and the heavy end current and other current additions; a dip tube made of Hastelloy material, 6 mm in external diameter, connected by a line, 9, to a thermal decomposition reactor 10 (when used) or to a purge tank on line 52. The HRU is also heated with a heating cover and is mechanically stirred. The various charging, reflow and purge streams were pumped into and out of the reactor from loading glass funnels, using dosing pumps. The thermal control of the HRU was regulated by an electronic temperature controller attached to a calibrated thermal pair. All lines of the process streams that are exposed to streams containing sulfuric acid were constructed of Hastelloy C ™ or poly-tetrafluoroethylene (PTFE) material. The thermal decomposition reactor 10 consists of a 500 ml flask, similarly configured to the HRU, with regarding the control of the temperature and the current lines of the process. The inlet 9 of the purge stream of the HRU is fed to the thermal decomposition reactor by means of a feed port and a pump. The receiver 31 is a 125 ml fraction cutter device. The separation columns of acrylic acid are described in the specific Examples. All percentages are by weight, based on the weight of the mixture in which the indicated component is contained, unless indicated otherwise. Model Experiments for the Acrylic Acid Separation Column Model studies were performed using "Aspen Plus" ™, a flowchart simulator from Aspen Technology, Inc. All data points were obtained using a column model "Aspen "which has 13 theoretical trays plus the boiler and decanter and operated at a pressure higher than 75 mm Hg. The loading tray was the fourth theoretical tray from the bottom: the bottom stream of the column had dimensions to contain 90% by weight of AA and 10% of water, during the aqueous operation. Figure 2 shows the two "stable states" (ie, the desired "aqueous" and "unwanted" organic modes, previously described) of the acrylic acid separation column for the loads shown in Table 1, which correspond to a conversion of the reactor from AA to BA of 80%, a molar ratio of AA to BuOH and 1 to 1.35. The data of Figure 2 are plotted at a concentration of AA in the organic distillate, as a function of the flow rate of the aqueous reflux in the column. The simulations indicated that the minimum aqueous reflux flow rate necessary to operate the acrylic acid separation column in the desired aqueous stable state, with the loading of Table IA, was approximately 15546 kg / hour. The location of the aqueous / organic transition was estimated recognizing that the separation of BA and AA in the AA separation column is achieved through the azeotropic distillation of BA, using water as an azeotropic agent. The following Tables IA and IB illustrate how the minimum amount of water necessary for the azeotropic reaction of the entire BA in the charge of the acrylic acid separation column is calculated. The first azeotrope that acts on the column is the ternary azeotrope of BA / butanol / water of lower boiling point, which at a pressure of 100 mm Hg boils at 46.42C and contains 36.0% BA, 26.4% BuOH and 37.6 % of water. This azeotrope exhausts the butanol in the load and takes 10429 kg / hour of the upper BA of a total of 20315 kg / hour, present in the charge. The amount of water needed to 43 satisfying this first azeotrope exceeds the amount in the load by 8198 kg / hour. Once the butanol is exhausted, the next lower boiling azeotrope that acts on the column is the BA / water binary azeotrope, which, at a pressure of 100 mm Hg, boils at 47.62C and contains 61.0% of BA and 39.0% of water. This second azeotrope takes the remaining 9885 kg / hour of the upper BA, using 6320 kg / hour of water to satisfy the azeotropic composition. The combined analysis of the two azeotropes shows that the total amount of water needed for the azeotropic reaction of the entire BA in the load exceeds the amount present in the aqueous load by 6320 kg / hr. This corresponds to the minimum amount of water that must be supplied by means of the aqueous reflux, to take all the upper BA so as to achieve the operation in the aqueous mode. An excellent concordance is shown between this estimate and the location of the aqueous / organic transition predicted by the data in Figure 2.

Tables ÍA v B Model Carsa conditions and calculations 1A: Load of the acrylic acid separation column for model conditions, calculated at 80% conversion / AA ratio: BuOH of 1: 1.35 IB: Calculation of the minimum water required for the azeotropic reaction of the entire BA in the load of Table IA, under the conditions of the model Notes: 1. Residue 1 is the load of Table 1A, reduced by the 1st azeotropic components 2. Residue 2 is Residue 1 reduced by the 2nd azeotropic components 3. Water required to achieve the azeotropic composition of azeotropic distillations respective showed deficits by the amount indicated The modeling results presented here correspond to the particular load to the separation column of the acrylic acid. However, the same analysis can be applied to any column loading, which corresponds to any particular set of reactor conditions, to estimate the minimum requirement of the aqueous reflux in the column. The ability to accurately predict the minimum water requirement for the acrylic acid separation column, based on the charge composition alone, allows the selection of an operation reflux ratio that minimizes the amount of heat and the diameter of the product. the column while ensuring a stable operation in the desired aqueous mode. In the modeling, it was possible to control in which stable state, aqueous or organic, the column operates starting from an operation at an extreme point (ie, very high reflux regimes, for the aqueous stable state or very low reflux regimes for the stable organic state), where only the objective stable state exists, and then moving along the branch, or decreasing or increasing, respectively, the flow regime of the reflux, until the objective point of operation is reached. This is achieved through the examination of the sensitivity of the process to key variables; In this study, the luxury regimen of aqueous reflux was examined do not. The two steady-state branches of Figure 2 were obtained by conducting two sensitivity studies in the program. In the first study, the rate of water reflux flow started at a very high end of 61364 kg / hr (a reflux ratio of about 40, using 1591 kg / hr as the forward regime of the aqueous load) and decreases gradually at a very low rate of 4545 kg / hr (a reflux ratio of about 3). This study generated the minor "watery branch" of Figure 2, which represents the desired aqueous mode, where the AA levels in the organic distillate are very low. (In this program, the lowest level of AA in the distillate (27 ppm) was achieved with the minimum reflux amount (approximately 15454 kg / hr, a reflux ratio of around 9, indicates operating the column in the state mode desired stable.) When the reflux flow rate becomes too low, the column becomes inoperable in an aqueous mode and, at 14090 kg / hr of reflux, a sudden and very large increase in the AA level occurred. At below 1090 kg / hr, the column operates only in the organic mode and the two modes converge to a simple solution.In a second sensitivity study, the water reflux flow rate was initiated at the low end, at 4545 kg / hr and gradually increased to about 61364 kg / hr.

This study generated the superior "organic branch" in Figure 2 and represents the unwanted organic mode, where the AA levels in the organic distillate are much higher, as indicated. Moving successively along this branch (dots 4-7) to more than about 54545 kg / hr, where there is sufficient water to force the column in the aqueous mode of operation, both branches converge to a simple aqueous stable state. Within the aqueous branch, the operating region in this simulated study leads to the BA that does not have substantially AA (an objective of 2,000 ppm, preferably <1000 ppm of AA) is a small region in the aqueous bottom branch of Figure 2. The program also forecasts high levels of BA (for example 23-74% by weight) in the bottom stream, where the column operates in the organic mode. The recycling of BA to the esterification reactor. it is not convenient, because it decreases the conversion rate of AA and BuOH. In the subsequent modeling of the two stable states in the AA separation column, it was determined that by avoiding the reactor condenser and phase separator 14 and feeding a vaporized mixture directly to the column, it has the advantage of reducing the vapor requirements of the column. However, because the water in the cargo is already vaporized, it is not essentially available to form an azeotrope with BA and more water is required from reflux to compensate for this deficiency. For a vapor load to the column, the aqueous / organic transition point in Figure 2 is moved to the right by the amount of water in the charge, and the range of the ratio of the aqueous reflux to the operation in the aqueous mode it narrows from 13: 1 to 17: 1. The modeling also showed that it is detrimental to reflux any portion of the organic phase for the operation of the column, because any BA and butanol returned to the column by means of an organic reflux, will simply need to be removed again by the azeotropic distillation with additional water. In addition, the AA in the organic reflux is returned to the column at the top, leaving no trays to rectify this AA contribution to the outside of the upper vapor. These factors increase the minimum amount of water needed to operate in the aqueous mode, reduce the width of the aqueous operation window and raise the minimum AA levels that can be achieved in the distillate. The vapor-liquid balance (VLE) data indicates that butanol has the effect of decreasing the volatility of AA. According to the VLE data, modeling shows that the column charge currents that are rich in butanol give distillate streams that are low in AA. Therefore, the low conversions and high ratios of butanol to AA in the reactor which supplies effluents rich in butanol, are favorable for the separation of the BA / AA and provide wide windows for the operation of the water reflux (8.5: 1 to 17). : 1), as described. In the event that the reactor can not be operated under the above conditions, a provision for a separate fresh butanol stream, fed directly to the AA separation column, can be made to ensure a wide window of the aqueous mode, independent of the reactor conditions. Fresh butanol is better fed at or slightly below the main load. Butanol should never be fed above the main charge containing the AA (As a light component, the butanol feed above the main charge allows it to vaporize rapidly, leaving the trays between the main carya and the butanol charge with little butanol to suppress AA volatility.) Laboratory Confirmation of the Aqueous and Organic Modes The existence of two stable states in the AA separation column was experimentally confirmed in a continuous multi-day laboratory operation, from which Examples 1 to 6 and Comparative Examples 1 were taken. and 2. The material flow regimes in Table 1 and the simulations generated by Figure 2 are modeled at the plant scale; in a column and separation of the acrylic acid, in the following Examples, the flow rates were scaled down so that 250 kg / hr in the scale model of the previous plant were equivalent to 1 g / hr in this laboratory operation . The extended operation, which closely follows the points between circles in Figure 2, was initiated to demonstrate the continuous operation of the column in the aqueous mode for various reflux flow regimes, Points 1 and 2. This portion of the operation was followed by an intentional decrease of the reflux water to bring the column to operation in the organic mode, Points 3 and 4. Subsequent changes of the boiling conditions were imposed to restore the column to the aqueous operation. In Figure 2, points 1 and 5, 2 and 4, 6 and 8 represent pairs of points in correspondence, ie equal reflux flow rate points in the aqueous and organic modes, respectively. The aqueous mode, once achieved, is maintained by a reflux ratio of 8.5: 1 to 17: 1, and the distilled BA supplied having the desired level of AA, < 2,000 ppiu, and also a stream of watery, separate AA, which does not have substantially BA. The measured levels of AA in BA in points 1 and 2, were 950 ppm and 200 ppm of AA, respectively, and of BA in AA, none (<1 ppm) was measured.

Example 1 - Operation in the Aqueous Mode to a Ratio of the Aqueous Reflux of 16 (Point 1 of Figure 2) Column 15 of separation of the acrylic acid was an Oldershaw fractional distillation column, of 30 trays, with a diameter of 2.54 cm. , equipped with a glass condenser on line 16 and a stainless steel steam boiler. The column was operated at a pressure higher than 75 mm Hg. The separation column of the acrylic acid was fed, per hour, with 10.8 g of AA, 30.6 g of butanol, 82.4 g of BA and 12.0 g of water. The composition of the mixture corresponds to a reactor condensate generated in a system where reactor 1 operates at a ratio of AA to BuOH of 1: 1.35 and an 80% conversion in AA, while receiving, per hour, 0.18 g of BA recycled per gram of unreacted BuOH and 0.11 g of recycled water per gram of unreacted AA. The loading tray was 1P, to the tray from the bottom. The upper mixture distilled at a temperature of 43.5sec and condensed and separated into two phases at the receiver 18. From the organic phase 19 rich in BA, 117.3g / h, containing, by weight, 70.3% of BA, 25.0% butanol, 3.6% water and 0.1% AA. From the separated aqueous phase 20, 110.2 g / h (94.3% of the phase) was recycled to the top of the column through line 21 and 6.7 g / h (5.7%) were moved forward through the line 22, providing a relationship of aqueous reflux of 16.4. The aqueous phase contained 96.6% water, 0.2% butanol. 0.2% BA and 354 ppm AA. From the AA-rich bottom product, stream 17, 12.0 g / h was collected, which contains 89.2% AA and 10.8% water. The resulting background temperature was 60.02C. Example 1 corresponds to point 1 in Figure 2.

Example 2 - Operation of the Aqueous Mode to the Aqueous Backflow Ratio of 11 (Point 2 of Figure 2) The apparatus, load regime, load composition, load location, column pressure and general operation of the column were the same as those described in Example 1. The steam to the boiler and the return rate of the aqueous condensate were reduced in order to reduce the flow of aqueous reflux to the column. The upper product in line 16 obtained at a temperature of 42.62C, was condensed and separated into two phases in receiver 18. 117.1 g of organic phase 19, rich in BA, was collected and contained 70.4% of BA, 26.0% butanol, 3.6% water and 218 ppm AA. From the separated aqueous phase 20, 77.2 g / h (91.8%) were recycled to the top of the column - through line 21 and 6.9 g / h (8.2%) they moved forward through the line 22, providing an aqueous reflux ratio of 11.2. The aqueous phase contained 26.6% water, 3.2% butanol, 0.2% BA and 81 ppm AA. 12.0 g / h of the bottom product, by line 17, were collected and contained 90.0% AA and 10.0% water. The resulting temperature of the bottom was 60.42c. This Example corresponds to point 2 in Figure 2. Comparative Example l - Operation in the Organic Mode (Point 3 of Figure 2) The apparatus, load regime, load composition, load location, column pressure and general operation of the column, were the same as those described in Example 1 and the column was initially operated in an identical manner to Example 2. The steam to the boiler and the return rate of the aqueous condensate were reduced in order to further reduce the flow of water reflux to the column. The upper product through line 16 was obtained at a temperature of 50.12c, condensed and separated into two phases in receiver 18. 118.5 g of organic phase 19, rich in BA, was collected and contained 62.6% of BA, 25.6% butanol, 6.1 AA and 5.7% water,? E the separated aqueous phase 20, 36.0 g / h (86.2%) were recycled to the top of the column through line 21 and 5.8 g / h (13.8%) moved forward through line 22, providing an aqueous reflux ratio of 6.3. The aqueous phase contained 94.4% water, 3.1% butanol, 2.2% AA and 0.3% BA. 11.8 g / h of the bottom product, by line 17, were collected and contained 70.7% BA, 29.1% AA and 0.2% butanol. Temperature resulting from the fund was 88.32c. This Example corresponds to point 3 in Figure 2 and showed that the operation of the column in a reflux ratio below the reflux ratio regime of the present invention leads to an operation in the unwanted organic mode. The results include high levels of AA in the organic phase 19 rich in BA, high levels of BA in the bottom stream 17 and high column temperatures with reíacion to the conditions of the aqueous mode of Examples 1 and 2.

Comparative Example 2 - Confirmation ß The Two Stable States The apparatus, loading regime, charge composition, loading location, pressure, column and general operation of the column were the same as those described in Example 1 and the column was initially operated in an identical manner to Comparative Example 1. The steam to the boiler and the return rate of the aqueous condensate were increased in order to raise the aqueous reflux to the column at the same flow rate as in Example 2 (point 2) in Figure 1). The upper product through line 16 was obtained at a temperature of 43.92c, condensed and separated into two phases in receiver 18. 117.9 g / h of organic phase 19, rich in BA, was collected and contained 63.9% BA, 25.8% butanol, 5.2% water and 5.1% AA. Of the separated aqueous phase 20, 77.2% g / h (92.5%) were recycled to the top of the column via line 21 and 6.2 g / h (7.5%) moved forward through line 22, providing an aqueous reflux ratio of 12.4 . The aqueous phase contained 94.7% water, 3.1% butanol, 1.9% AA and 0.3% BA. 11.9 g / h of the bottom product, by line 17, were collected and contained 60.4% BA, 39.2% AA and 0.4% butanol. The resulting bottom temperature was 88.32c This Comparative Example corresponds to point 4 in Figure 2 and showed that even with the reflux rate of 77.2 g / h, the same as in Example 2, and an aqueous reflux ratio of 12.4, the column remains in the undesired state of operation of the organic mode and gave high levels of AA in the organic distillate and of BA in the bottom stream, and high temperatures of the 1-column in relation to the conditions of the Aqueous mode of Examples 1 and 2. By demonstrating the existence of Point 4 in Figure 2, the point of the organic mode analogous to point 2 of the aqueous mode, this Comparative Example demonstrated that there are truly stable states in the column, as forecast by the model described above. This Comparative Example also showed that the two stable state branches form a "hysteresis loop" and that once the column is operated in the unwanted organic mode, with sufficient heat input, remains in that mode of operation, even after the aqueous reflux regime has been increased to an effective level in the operation of the aqueous mode.

Example 3 - Restoration of the Operation in the Aqueous Mode from the Operation of the Organic Mode The apparatus, loading regime, charge composition, loading location, column pressure and general operation of the column, were the same as those described in the Example 1. The column was initially operated at point 3 of Figure 2 in an operation identical to Comparative Example 1. To the top tray of the column was then added a stream of water at a rate of 41.2 g / h. Combined with the original water reflux of 36.0 g / h, this additional water stream supplied an effective reflux flow to the column of 77.2 g / h, the same reflux regime as in Example 2 and Comparative Example 2, ie , points 2 and 4, respectively, in Figure 2. The steam inlet to the boiler was maintained at the same level as in Comparative Example 1 (point 3 of Figure 2). Through thermal pauses .---, placed on alternative trays, a cold front was observed, primarily of liquid water, which moved down the column, starting at the upper tray and descending through a tray at a time until finally arriving at the boiler. Thus, without additional steam provided to the boiler, to handle the upper load, the additional water fed to the upper tray behaved as expected, in providing a cooling effect to all the trays. Once the cold front arrived at the boiler, indicated by an acute temperature drop of 88.32C to 57.02C, the additional stream of fresh water to the upper tray was discontinued and the steam flow rates to the boiler were increased to to raise the aqueous reflux regime from 36.0 g / h to 77.2 h / h, and the column was allowed to reach the stable state now in the aqueous mode, in the higher reflux regime. The upper mixture through line 16, obtained at a temperature of 42.62C, condensed and separated two phases in receiver 18; 117.0 g / h of the organic phase 19, rich n BA, were collected and contained 70.5% BA, 26.0% butanol, 3.5% water and 263 ppm AA. From the separated aqueous phase 20, 77.2 g / h (91.8%) were recycled to the top of the column through line 21 and 6.9 g / h (8.2%) moved forward through line 22, providing an aqueous reflux ratio of 11.2. The aqueous phase contained 96.6% water, 3.2% butanol, 0.2% BA and 75 ppm AA. 12.1 g / h of the AA-rich bottom stream 17 were collected and contained 89.5% by weight of AA and 10.5% by weight of water. The resulting background temperature was 60.22c. This result corresponds to point 2 in Figure 2 and was substantially identical to that of Example 2. Thus, Example 3 demonstrates a trimmed method for returning the column to the desired operation of the aqueous mode from a point in the branch in an organic manner. In the aqueous mode, the separation column of the acrylic acid is operated with water in all the trays and in the bottom stream 17, while in the unwanted organic mode, the concentrates in the various upper trays and in the bottom stream 17 they are devoid of water. Although in this Example, the separation column of the acrylic acid starts in the organic mode, by the treatment shown, the column becomes operable in the desired aqueous mode. Its T.-. The result is especially important in view of the findings of Comparative Example 2, which have confirmed the two stable states in this particular system to produce BA from the "hysteresis loop", as shown in Figure 2. Example 4 - Operation in Aqueous Mode at an Acid Reflux Ratio of 9.6 Using the apparatus described in Example 1, the column was operated at a pressure higher than 75 mm Hg and fed per hour, with 5.6 g (0.08 mol) of AA, 34.8 g (0.47 mol) of butanol, 96.4 g (0.75 mol) of BA and 13.1 g (0.73 g) A- -O mol) of water. This mixing composition corresponds to a vaporized mixture of the reactor, generated in an esterification of BA, in which reactor 1 operates at a ratio of AA to butanol of 1: 1.5 and a conversion of 90% in AA, 18% of recycle of BA per unit weight of the unreacted butanol, by means of the stream 52 and 7% water per unit weight of the unreacted AA by stream 17. The loading tray was the tenth tray from the bottom. The upper distillate 16, obtained at a temperature of 42.22C, was condensed and separated into two phases at the receiver 18. 135.8 g / h of the organic phase 19 rich n BA were collected; it contained 71.0% by weight of BA, 25.5% by weight of butanol, 3.5% by weight of water and 550 ppm of AA. From aqueous phase 20, 78.9 g / h (90.6%) were recycled to the top of the column through stream 21 and 8.2 g / h (9.4%) moved forward through stream 22 , thus providing a water reflux ratio of 9.6. The separated aqueous phase contained 96.7% by weight of water, 3.1% by weight of butanol, 0.2% by weight of BA and 209 ppm of AA. 6.0 g / h of the AA-rich bottom stream was collected by line 17, which contains 9.31% by weight of AA and 6.9% by weight of water.

Example 5 - Operation of Aqueous Mode in the Aqueous Reflux Ratio of 11.0 The apparatus, loading regime, charge composition, loading location and column pressure were the same as in Example 4. The top mixture in 16 was obtained at a temperature of 41.9 C, it was condensed and separated into two phases at the receiver 18. 135.9 g / h of the organic phase 19 rich in BA were collected and contained 70.9% by weight of BA, 25.4% by weight of butanol, 3.5 % by weight of water and 779 ppm of AA. From aqueous phase 20, 89.5 g / h (91.6%) was recycled back to the top of the column through stream 21 and 8.2 g / h (8.4%) moved forward through stream 22 , for an aqueous reflux ratio of 11.0. The aqueous phase contained 96.7% by weight of water, 3.1% by weight of butanol, 0.2% by weight of BA and 286 ppm of AA. 6.0 g / h of the AA-rich bottom stream 17 were collected, containing 9.28% by weight of AA and 7.2% by weight of water. This Example demonstrated an increase of AA in the organic phase rich in BA and 550 ppm at 779 ppm, under these conditions, as the amount of the aqueous reflux increases in relation to that of Example 4 (reflux ratio of 9.6).

Example 6 - Operation in the Aqueous Mode at a 9.7- Aqueous Reflux Ratio with a Column of 35 Trays A section of five trays was added to the apparatus used in Example 1, thus providing a fractional distillation column Oldershaw, of 35 trays of 2.54 cm. in diameter, equipped with a glass condenser and a stainless steel steam boiler. The loading regime, the charge composition, the loading location and the column pressure were the same as those of Example 4. The top mixture at 16 was obtained at a temperature of 42.22C and condensed and separated into two phases. at receiver 18. 135.8 g / h of organic phase 19, rich in BA, were collected and contained 71.0 & amp;; in weight of BA, 25.5% in weight of butanol, 3.5% in weight of water and 193 ppm of AA. From the separated aqueous phase 20, 78.9 g / h (90.7%) was recycled to the top of the column, through stream 21 and 8.1 g / h (9.3%) moved forward through the flow. stream 22, for an aqueous reflux ratio of 9.7. The aqueous phase contained 96.7% by weight of water, 3.1% by weight of butanol, 0.2% by weight of BA and 72 ppm of AA. 6.1 g / h of the AA-rich bottom stream 17, containing 9.25% by weight of AA and 7.5% by weight of water were collected. The resulting background temperature was 62.32c. This example shows that adding 5 trays to the rectifying section of the AA separation column further reduces the AA in the BA-rich organic phase of 550 ppm, in Example 4, to 193 ppm. Comparative Example 3 - Thermal Decomposition Reactor Process, Without the Use of an HRU This comparative example was performed in the 500 ml thermal decomposition reactor, described above, using the charging currents described and without using an HRU. Thus, 73.46 g / h of a filler, containing the composition listed in Table 2, were fed to a CFSTR maintained at 1302C, a Hg pressure of 35 mm, a residence time of 60 minutes and a catalyst concentration of 8.07% by weight of H2SO4. A total of 55.24 g / h of the single phase distillate was recovered with the composition listed in Table 3. A purge stream of 18.22 g / h was bled from the thermal decomposition reactor and discarded as a waste oil. The recovered values of AA and BuOH are summarized in Tables 10 and 11, which show that, after recovery of the free values, only 15.0% of the AA values at the heavy ends and 11.6% of the BuOH values at the heavy ends recover.

Table 2 Composition of the Load Current for Comparative Example 3 1. 47 g / h, calculated as H2SO4 Table 3 Composition of the Higher Distilled Current of Comparative Example 3 Example 7 - Evaluation »of the HRU Under the Effective Conditions A total of 73.46 g of the organic load, containing the composition listed in Table 4, was fed to the HRU, maintained at a temperature of 1082C, a pressure of 760 mm of Hq, residence time of 144 minutes. , 16% by weight of the reactor water and a catalyst concentration of 2.7% by weight of H2SO4. In addition, 48.0 g / h of the first aqueous distillate from the esterification reactor (comprising 93.0% H2O, 6.0% AA and 1.0% H2O) were also fed.

BuOH) to the HRU, to compensate the distilled water and removed with the organic distillate and reactor purge and to simulate the recycling of the aqueous distillate to the HRU. A total of 39.35 g / h of organic distillate and 38.27 g / h of aqueous distillate were collected and analyzed. The results of the analysis are summarized in Table 5. The organic phase was separated and returned to an esterification reactor, when used. A purge current of 43.24 g / h was bled from the HRU and constituted the total charge to the bleeding separator, CESTR. Table 4 Composition of the Carqa Current for the Load of the HRU, Example 7 1. 47 g calculated as H2SO4. Table 5 Compositions of the Currents of the Evaluation of the HRU of Example 7 Example 8 - Evaluation of the Thermal Decomposition Reactor Under Effective Conditions The total of 43.24 g / h of the purge stream of the HRU of Example 7 was fed to the thermal decomposition reactor, CSTR, described above and maintained at a temperature of 1302C, a pressure of 100 mm Hg and a residence time of 120 minutes. A total of 33.63 g of the side, the upper stream of the thermal decomposition reactor was obtained for its return to an esterification reactor. A total of 9.61 g / h of the waste stream from the thermal decomposition reactor was collected and discarded as waste oil. The total recoveries of AA and BuOH for the combination of both units (HRU and thermal decomposition reactor) are summarized in Tables 10 and 11. The results show that, after recovery of the free values, 68.7% of the values of AA in the heavy ends and 59.5% of the BuOH values in the heavy ends recovered. Table 6 Composition of the Upper Current of the Thermal Decomposition Reactor of Example 8 Example 9 - Evaluation of the HRU under Severe Conditions and with a Tandem Thermal Decomposition Reactor This example, under more severe operating conditions of the HRU (higher acid concentration) than in Examples 7 and 8, provided a greater recovery of AA and BuOH values. Thus, 73.46 g / h of the loading stream with the composition listed in Table 7 were fed to the HRU and were maintained at 1142C, 760 mm Hg, 5.1% by weight H2SO4 (7.5% by weight per mass balance). (MB)), 144 minutes of residence time and a water concentration of 15.5% by weight. In addition, 46.48 g of the aqueous load comprising 93% water, 6% AA and 1% butanol were fed into the HRU to simulate the recycling of the aqueous distillate plus the compensation of the loss of water to the organic distillate and the purge Of funds. A total of 41.96 g / h of the organic distillate of the HRU (composition in Table 8), 36.80 g / h of the aqueous distillate of the HRU and 28.22 g / h of the upper stream of the decomposition reactor (composition in Table 9) ) were recovered and analyzed. The upper stream of the thermal decomposition reactor was obtained by feeding the purge stream of the HRU (39.18 g / h) to the thermal decomposition reactor maintained under the following conditions: 1302C, 100 mm Hg, residence time 1120 minutes, 26.8 % H2SO4 (per MB). The recoveries qe Total AA and BuOH for this tandem combination are listed in Tables 10 and 11 and show that, after recovering the free values, 81.7% of the AA values and 65.2% of the BuOH values at the heavy ends are They recovered. Table 7 Composition of the Carqa Current of Example 9 * Calculated as H2SO4 Table 8 Compositions of the Currents of the HRU Evaluation - Example 9 Table 9 Composition of the Upper stream? E Reactor? E Thermal decomposition of Example 9 Example 10 - Continuous Process to Produce Butyl Acrylate Esterification reactor 1 was a 2-liter round bottom Pyrex flask equipped with a two-plate Oldershaw distillation column (5.0 cm in diameter) (serving as an acid catalyzed stripping separator), a condenser, thermal pair, supply ports attached to the appropriate fluid metering pumps and lines leading to a hydrolytic reactor unit (HRU, 5) and a thermal decomposition reactor 10, described more completely below. The working capacity of the reactor was 750 ml of the reaction mixture containing 2.50% by weight of sulfuric acid catalyst. The reaction temperature was 89 C and the pressure was 127 mm Hg. Reactor 1 was fed with 182.90 g / h of fresh crude AA (assay, 96% AA by weight, 2435 units / h), 182.48 g / h of fresh n-butanol (BuOH, 2466 mmoles / h) and 1.71 g / h of fresh H2SO4 (95.5% by weight of acid). The reactor was fed with a total of 655.3 g / h of a composite material of 223.85 g / h of AA (3105 mmol / h), 316.90 g / h of BuOH (4282 mmol / h), a higher organic stream, condensed and separated from the HRU, an upper condensate from the thermal decomposition reactor, a bottom stream from the separation column of AA and a mixture of BA / BuOH / H20, representing the recovery and recycle streams of the following streams: (a) the unreacted BuOH in an azeotropic downstream BuOH / BA distillation column; (b) a recovery stream of BuOH / BA from the separation of the aqueous waste streams, before sending them to a waste treatment and (c) a portion of the bottoms of the distillation column of the final product. (These flows comprise a typical and supplementary load (for example, recycled), used in a continuous process representative of a plant). The total BA thus fed to the esterification reactor of these sources was 88.07 g / h (688 mmoi / h) of which 50.85 g / h represent the recycling of the downstream separation columns. The AA and BuOH were used correspondingly in a molar ratio of 1: 1.38. The reactor was maintained at a residence time of about 60 minutes, whereby 749.8 g / h of the total material was distilled off as the upper distillate of the reactor through the Oldersha column, condensed and separated into two phases. A portion of the organic phase (160 g / h) was returned to the top of the distillation column as reflux. The remaining 563.8 g / h of the organic distillate contained 38.56 g / h of AA, 124.59 g / h of BuOH and 371.24 g / h of BA and were fed to the separation column 15 of acrylic acid.

The condensed, aqueous, vaporized mixture of the reactor was separated (26.0 g / h, which contains 2.12 g / h of BuOH and 0.619 g / h of AA) and was divided as follows: 22.4 g / ha the separation column of AA by line 53 and 3.6 g / ha HRU, by way 42. The HRU 5 and the thermal decomposition reactor 10 are identical to those described in Examples 7, 8 and 9. Correspondingly, 65.5 g / h of the purge stream of the esterification reactor, which contains 4.33 g / h of BuOH, 6.19 g / h of AA, 34.23 g / h of BA and other related high boilers and inhibitors, were fed to HRU 5 by of line 3 and were maintained at 122 C, pressure of 760 mm Hg, residence time of 317 minutes, 6.26% by weight of H2O and a concentration of the acid catalyst of 7.58% by weight of H2SO4. Additionally3.6 g / h of the aqueous condensate distillate from the reactor were fed to the HRU. From this HRU, a total of 80.5 g / h of material was distilled as an upper stream, condensed and separated. All the separated aqueous phase (38.3 g / h, which contains 2.47 g / h of BuOH and 1.01 g / h of AA), was returned to the HRU as reflux, via 7. The separated organic phase (42. g / h, containing 7.07 g / h of BuOH, 3.19 g / h of AA and 29.72 g / h of BA) was returned to the esterification reactor as a recovered recycle stream, via track 8. A current of purge of 27.0 g / h was removed from the HRU by way 9, and fed to the reactor 10 of thermal decomposition, maintained at 1202C, pressure of 35 mm Hg, residence time of 815 minutes and 20.5% by weight of H2SO4 . A total of 17 g / h (containing 0.397 g / h of BuOH, 8.44 g / h of AA and 4.78 g / h of BA) of the material was distilled off and condensed in 31. This combined condensate was returned via 12 to the esterification reactor as recycled. The purge stream of the thermal decomposition reactor was discarded as a waste oil by the line 13. The separation column 15 of the acrylic acid consists of an Oldersha distillation column of 35 plates, 5.0 cm. in diameter, equipped with a stainless steel boiler, with steam jacket, and a condenser system cooled with water. Therefore, 563.8 g / h of the condensed distillate of the organic layer of the esterification reactor and 22.4 g / h of the condensed distillate of the aqueous layer of the esterification reactor (composition described above) were fed in 15, operated at a pressure of Head of 260 mm Hg, a base temperature of 822c and an aqueous reflux ratio of 9.61. A total of 907.3 g / h of the upper mixture were obtained by distillation through the column, condensed and separated in 18. A total of 400.7 g / h of the separated aqueous phase was collected, of which 363.4 g / h HE returned to the head of the column as reflux via line 21. The organic phase, rich in BA (506.60 g / h) containing the BA product is substantially free of AA (1450 ppm). A background current, rich n AA, of 42.3 g / h was removed from the column (35.35 g / h of AA) and recycled via route 17 to the esterification reactor. The recovery data are included in Tables 10 and 11. With the entire BA process operated in this manner, a quantitative BA yield in BuOH was performed and a yield of 102.7% BA in AA was realized (from a theoretical yield of 104.8%, based on AA content and AOPA of fresh loaded AA). Example 11 - Continuous Process for the Production of the BA, which includes the Recycling of Additional Currents The esterification reactor and the related process equipment, used in this example, are identical to those described in Example 10. The working capacity of the reactor was 1000 ml of the reaction mixture containing 2.25% by weight of sulfuric acid catalyst. All work units were fed as described below. Reactor 1 was fed with 183.90 q / h of fresh crude AA (assay: 96% d AA by weight, 2449 mmol / h), 207.03 g / h of fresh n-butanol (BuOH, 2798 mmol / h) and 2.05 g / h of fresh H2SO4 (95.5% by weight of acid). The HRU was fed by average of the additional currents of 30.58 g / h (424 mmol / h) of AA and 4.11 g / h (55.5 m?.? ul / h) of BuOH that carries a total fresh load of AA to the system of 207.12 g / h (2873 mmol / h) and a fresh total charge of BuOH of 211.14 g / h (2853 mmol / h) The reactor 1 was thus fed with n total of 866.9 g / h of a composite material of: 260.10 g / h of AA (3607 mmol / h), 406.6 g / h of BuOH (5495 mmol / h), an upper condensed organic layer of the HRU, an upper stream of the thermal decomposition reactor, a bottom stream of the separation column of A , and a BA / BuOH mixture representing the recovery and recycle streams of the following streams: (a) Unreacted BuOH in a downstream BuOH / BA azeotropic distillation column, (b) a recovery stream of BUOH / BA division of the aqueous waste streams before sending them to the waste treatment, and (c) a portion of the bottoms of the distillation column of the final product. (These streams include typical load and recycle streams, used in a continuous, fully integrated plant process). The total load of BA to the esterification reactor from these sources is 144.6 g / h of which 84.14 g / h represent the recycled BA from the downstream separation columns and the supplementary waste streams. AA and BuOH are therefore used in reactor 1 in a molar ratio of 1: 1.52. The reactor was maintained, at a residence time of approximately 60 minutes, whereby 995.9 g / h of the total material, such as the upper distillate of the reactor, was distilled through the Oldersha distillation column, condensed and separated into two phases. . A portion of the organic phase (216.1 g / h) was returned to the top of the distillation column as reflux. The remaining 719.8 g / h of organic distillate, containing 45.20 g / h of AA, 183.10 g / h of BuOH and 451.0 g / h of BA, were fed to the acrylic acid separation column, 15, which is described down. The aqueous distillate from the reactor (60.10 g / h, with 4.78 g / h of BuOH and 1.06 g / h of AA), was thus divided: 36.5 g / h from the separation column of AA and 23.6 g / h to the HRU. The HRU 5 and the thermal decomposition reactor 10 are identical to those described in Examples 7, 8 and 9. Therefore, 87.1 g / h of the purge stream of the esterification reactor via line 3 and 1128.7 g / h of additional streams containing 11.3 g / h of BuOH, 38.3 g / h of AA, 37.1 g / h of BA and other high-boiling components and related inhibitors via line 4, were fed to the HRU maintained at 122 C, a pressure of 760 mm Hg, a residence time of 150 minutes, 12.8% by weight of H2O, and a catalyst concentration of 9.3% by weight of H2S? 4 Additionally, 23.60 g / h of the aqueous distillate from the reactor were fed to the HRU. From this HRU, a total of 198.7 g / h of material was styled, condensed and separated. All the aqueous distillate (114.6 g / h, containing 4.35 g / h of BuOH and 6.25 g / h of AA) was returned to the HRU as reflux by line 7. The organic distillate (84.1 g / h, containing 12.2 g / h of BuOH, 9.94 g / h of AA and 56.3 g / h of BA) was returned to the esterification reactor as recycled by line 8. A purge stream of the HRU of 15.2 g / h was removed from the bottom of the HRU was fed to the thermal decomposition reactor 10, maintained at 120sc, a pressure of 35 mm Hg, a residence time of 180 minutes and 24.0% by weight of H2SO4. A total of 70.6 g / h (containing 2.00 g / h of BuOH, 34.3 g / h of AA and 9.29 g / h of BA) from the upper stream of the thermal decomposition reactor were distilled and condensed. This condensate was recovered and returned to the main esterification reactor as a recycled stream and the purge stream from the bottom of the thermal decomposition reactor was discarded as a waste oil. The separation column 15 of acrylic acid, consisting of an Oldershaw distillation column, of 35 plates, of 5.0 cm. in diameter, equipped with a stainless steel boiler, steam jacket, and a condenser system cooled with water. 719.8 g / h of the organic distillate condensate from the esterification reactor and 36.5 g / h of the aqueous distillate condensate from the reactor (already described) were charged to the acrylic acid separation column, operated at a head pressure of 260 mm. Hg, a base temperature of 82 C and an aqueous reflux ratio of 12.5. A total of 1193.8 g / h of the top mixture were obtained by distillation through the AA separation column, condensed and separated. A total of 526.4 g / h of the separated aqueous phase was collected, of which 487.9 g / h were returned to the top of the AA separation column as reflux by line 21, the remainder of the aqueous phase was moved forward with the rest of the condensate. The organic phase rich in BA (667.4 g / h) containing the BA product was also moved forward for further isolation. it was substantially free of AA, containing 1500 ppm of AA. An AA-rich bottom stream of 50.4 g / h was removed from the AA separation column (43.56 g / h of AA) and recycled via 17 to the esterification reactor. The recovery data are included in Tables 10 and 11. With the entire process to produce the BA, operated with the process described in this Example, a BA yield, based on BuOH, of 100.5% was performed; Based on the AA, the BA yield of 99.8%.

Table 10 Summary of AA Values Alived and Recovered Table 11 Summary of the Values - the BuOH Feeded and Recovered

Claims

CLAIMS 1. A method for recovering acrylic acid, a C?-C4 alkyl acrylate and a C3_-C4 alkanol from the heavy ends produced during the production of C?-C4 alkyl acrylate, this method comprises steps of: a) feeding a total, aqueous and heavy-end charge stream, comprising these heavy ends, water, a residual acid catalyst and, optionally, a strong acid, selected from a mineral acid or sulfonic acid, a hydrolysis reactor, maintained at 90 to 1402C, at a pressure of 50 to 1000 mm Hg, and a residence time of 0.5 to 20 hours, based on the total, aqueous and organic loading stream; b) distilling a higher stream containing the acrylic acid, the C1-C4 alkyl acrylate and the C4-C4 alkanol and the aqua, from the hydrolysis reactor, while maintaining the concentration of the hydrolysis reactor liquid from 5 to 40 % by weight of water and at least 1% by weight of acid, and this acid comprises the residual acid catalyst and the strong acid, optional; c) condensing the upper current; d) separating from the condensed top stream an organic phase, comprising the C 1 -C 4 alkyl acrylate, the C 1 -C 4 alkanol, and the acrylic acid, and an aqueous phase which Primarily comprises water, acrylic acid and alkanol C1-C4; r: e) remove the separated organic phase; f) recycling the separated aqueous phase to the hydrolysis reactor; Y g) removing from the hydrolysis reactor from 20 to 70% by weight, based on the total, aqueous and heavy-end charge stream, of a purge stream from the hydrolysis reactor. 2. The method according to claim 1, further comprising the steps of: a) feed up to 100% of the purge stream of the hydrolysis reactor to a thermal decomposition reactor, maintained at 90 to 1402C, a pressure of 20 to 200 mm Hg and a residence time of 0.5 to 20 hours, with base in the purge stream of the fed reactor; b) distilling from the thermal decomposition reactor a higher stream, comprising the acrylic acid, the C1-C4 alkyl acrylate, the C1-C4 alkanol and the aqua, while maintaining the concentration of the thermal decomposition reactor liquid in the minus 7.5% by weight of acid; c) condensing the upper stream of the thermal decomposition reactor; Y d) recovering from the upper stream of the thermal decomposition reactor, acrylic acid, C 1 -C 4 alkyl acrylate, C 1 -C 4 alkanol and water. 3. The method according to claim 1, wherein the total, aqueous and heavy-end charge stream is fed to a hydrolysis reactor comprising a multi-plate reactive distillation column. 4. The method according to claim 1, wherein the total, aqueous and heavy-end charge stream is fed to a hydrolysis reactor under continuous mixing conditions. 5. The method according to claim 1, wherein the mineral acid is the sulfuric acid and the sulfonic acid is selected from the group consisting of the methanesulfonic acid, benzenesulfonic acid and toluenesulphonic acid 6. A method for recovering acrylic acid, n-butyl acrylate and n-butanol from heavy ends, produced during acid-catalyzed esterification of acrylic acid, with n-butanol, this method comprises the steps of: a) feeding a total, aqueous and heavy-end charge stream, comprising acrylic acid, n-butyl acrylate, n-butanol, water, heavy ends, catalytic residual acid and, optionally, a strong acid, selected from a mineral acid or sulfonic acid, to a hydrolysis reactor, maintained at 90 to 1402c, a pressure of 50 to 1000 mm Hg, and a residence time of 0.5 at 20.0 o'clock, based on the total, watery load and heavy ends; b) distilling a higher stream, containing the acrylic acid, n-butyl acrylate, n-butanol and water, from the hydrolysis reactor, while maintaining a concentration of the hydrolysis reactor liquid from 5 to 40% by weight of water and at least 1% by weight of acid, this acid comprises the residual acid catalyst and the optional strong acid; c) condensing the upper current; d) separating from the condensed top stream an organic phase comprising n-butyl acrylate, n-butanol and acrylic acid, and an aqueous phase comprising primarily water and acrylic acid and n-butanol; e) remove the separated organic phase; f) recycling the separated aqueous phase to the hydrolysis reactor; Y g) removing from the hydrolysis reactor from 20 to 70% by weight, based on the total, aqueous and heavy ends, of a purge stream of the hydrolysis reactor. The method according to claim 6, further comprising the steps of: a) feeding up to 100% of the purge stream of the hydrolysis reactor to a thermal decomposition reactor, maintained at 90 to 1402C, a pressure of 20 at 200 mm Hg and a residence time of 0.5 to 20 hours, based on the purge stream of the fed reactor; b) distill from the thermal decomposition reactor a charging stream, comprising acrylic acid, n-butyl acrylate, n-butanol and water, while maintaining the concentration of the thermal decomposition reactor liquid at least 7.5 % by weight of acid; c) condensing the upper stream of the thermal decomposition reactor; Y d) recovering from the upper stream of the thermal decomposition reactor, the acrylic acid, the n-butyl acrylate, the n-butanol and the water. The method according to claim 6, in which the total, aqueous and heavy end charge stream is fed to a hydrolysis reactor, comprising a multi-plate reactive distillation column. 9. The method according to claim 6, wherein the total aqueous and organic charge stream is fed to the hydrolysis reactor under continuous mixing conditions. 10. The method according to claim 6, wherein the mineral acid is selected from the sulfuric acid and the sulfonic acid is selected from the methan-, benzene or toluenesulfonic acid. 11. The method according to claim 6, wherein the residual acid catalyst and the mineral acid are each sulfuric acid. 12. A method for continuously recovering acrylic acid, n-butyl acrylate and n-butanol from heavy ends, produced during acid catalyzed esterification of acrylic acid with n-butanol, this method comprises the steps of: a) continuously withdrawing a purge stream from the reactor from an esterification reactor, which contains an esterification reaction mixture, comprising acrylic acid, n-butyl acrylate, n-butanol, heavy ends and residual acid catalyst, while concurrently the acrylic acid, n-butyl acrylate, n-butanol, and water are distilled from the esterification reaction mixture; 9? b) feeding a total, aqueous and organic charge stream, comprising the purge stream of the reactor, water, optionally a strong acid, selected from a mineral acid or sulfonic acid and, optionally, additional heavy ends, to a reactor hydrolysis, maintained at 90 to 1402C, at a pressure of 50 to 1000 mm Hg, and a residence time of? 5 to 20 hours, based on the total, aqueous and organic loading stream. c) distilling a top stream, which contains the acrylic acid, n-butyl acrylate, n-butanol, and water from the hydrolysis reactor, while maintaining a concentration of the hydrolysis reactor liquid from 5 to 40% by weight of water and at least 1% by weight of acid, this acid comprises the residual acid catalyst and the optional strong acid; d) condensing the upper current; e) separating from the condensed top stream an organic phase, comprising n-butyl acrylate, n-butanol, and acrylic acid, and an aqueous phase comprising primarily water, and acrylic acid and n-butanol; f) remove the separated organic phase; g) recycling the separated aqueous phase to the hydrolysis reactor, and h) removing from the hydrolysis reactor from 20 to 70% by weight, based on the total aqueous and organic charge stream, of a purge stream from the hydrolysis reactor. 13. The method according to claim 12, further comprising the steps of: a) feed up to 100% of the purge stream of the hydrolysis reactor to a thermal decomposition reactor, maintained at 90 to 1402C, a pressure of 20 to 200 mm Hg and a residence time of 0.5 to 20 hours, based on the purge current of the fed reactor; b) distill from the thermal decomposition reactor a charging stream, comprising acrylic acid, n-butyl acrylate, n-butanol and water, while maintaining the concentration of the thermal decomposition reactor liquid in at least 7.5 % by weight of acid; c) condensing the upper stream of the thermal decomposition reactor; Y d) recovering from the charge stream of the thermal decomposition reactor, the acrylic acid, n-butyl acrylate, n-butanol and water. 14. The method according to claim 12, wherein the total aqueous and organic load stream is fed to a hydrolysis reactor, comprising a reactive distillation column of multiple plates. 15. The method according to claim 12, wherein the total aqueous and organic charge stream is fed to the hydrolysis reactor under continuous mixing conditions. 16. The method according to claim 12, wherein the additional heavy ends comprise up to 80% by weight of the total, aqueous and organic charge stream. 17. The method according to claim 12, wherein the mineral acid is the sulfuric acid and the sulfonic acid is selected from the methan-, benzene or toluene sulfonic acid. 18. The method according to claim 12, wherein the acid catalyst and the mineral acid are each sulfuric acid. 19. A method for the continuous recovery of n-butyl acrylate, substantially free of acrylic acid, from an esterification reaction mixture, this method comprises the steps of: a) continuously feeding to an esterification reactor, acrylic acid and n-butanol, in a molar ratio of 1 to 1.1 to 1 to 1.7, and an acid catalyst; b) reacting the acrylic acid and the n-butanol to supply the n-butyl acrylate with a conversion of less than 60% based on the acrylic acid, and supplying the esterification reaction mixture comprising acrylic acid, n-butyl acrylate, n-butanol, water, heavy ends and acid catalyst; c) distilling from the esterification reactor a vaporized mixture comprising acrylic acid, n-butyl acrylate, n-butanol and water; d) condensing the vaporized mixture to supply a first condensate comprising an organic phase and an aqueous phase; e) returning from 0 to 30 percent of the organic phase to a drag separator placed on the esterification reactor; Y f) feeding 70 to 100 percent of the organic phase and 50 to 100 percent of the aqueous phase, to a separation column of acrylic acid; g) distilling from the acrylic acid separation column, at a pressure of 35 to 800 mm Hg, in an aqueous mode and at an aqueous reflux ratio of 8.5: 1 to 17: 1, a superior mixture comprising an azeotroped mixture - butanol of butanol, butyl acrylate and water; h) removing from the distillation column a bottom stream rich in acrylic acid; i) recycling the bottom stream, rich in acrylic acid, from the acrylic acid separation column to the esterification reactor; j) condensing the upper surface to supply a second condensate; k) separating this second condensate into an organic phase, rich in butyl acrylate, and a separate aqueous phase; Y

1) remove the organic phase rich in butyl acrylate, substantially free of acrylic acid. 20. The method, according to claim 19, wherein the acid catalyst is selected from the group consisting of sulfuric acid, sulfonic acid and an ion exchange resin of a strong acid. 21. The method according to claim 19, wherein the acid catalyst is sulfuric acid. 22. The method according to claim 19, wherein the distillation in the separation column of the acrylic acid is carried out at a pressure of 90 mm to 135 mm Hg. 23. The method according to claim 19, wherein the ratio of the aqueous reflux is 8.5 to 12.5. The method, according to claim 19, further comprising the steps of: a) omit steps d), e) and f); and b) feeding 100 percent of the vaporized mixture directly to the acrylic acid separation column, from step f), and distilling, wherein the ratio of the aqueous reflux is from 13: 1 to 17: 1. 25. A continuous process for producing n-butyl acrylate substantially free of acrylic acid and for recovering acrylic acid, n-butyl acrylate, n-butanol and water from a mixture of the esterification reactor, which contains the acrylic acid, n-butyl acrylate, n-butanol, water, heavy ends and the acid catalyst, comprising the steps of: a) feeding an esterification reactor with acrylic acid and n-butanol, in a molar ratio of 1 to 1.1 up to 1 to 1.7, and the acid catalyst; b) reacting the acrylic acid and the n-butanol to supply the n-butyl acrylate at a conversion of at least 60% to acrylic acid, and supplying the esterification reaction mixture comprising the acrylic acid, n-acrylate and butyl, n-butanol, water, heavy ends and acid catalyst; c) withdrawing a purge stream from the reactor from the continuous conversion sterilization reactor mixture, while acrylic acid is concurrently distilled, n-butyl acrylate, n-butanol and water, of this esterification reaction mixture; d) feeding the total load, aqueous and organic stream, comprising the purge stream of the reactor, water, optionally a strong acid, selected from a mineral acid or a sulfonic acid, and, optionally, additional heavy ends, to the reactor of hydrolysis, maintained at 90 to 1402C, a pressure of 50 to 1,000 mm Hg and a residence time of 0.5 to 20 hours, based on the total, aqueous and organic load current; e) distilling a higher stream, containing the acrylic acid, n-butyl acrylate, n-butanol and water from the hydrolysis reactor, while maintaining a concentration of the hydrolysis reactor liquid of 5 to 40% by weight, aqua and at least 1% by weight of acid, this acid comprises the acid catalyst and the optional strong acid; f) condensing the upper current; g) separating from the condensed upper stream an organic phase comprising n-butyl acrylate, n-butanol and acrylic acid and an aqueous phase comprising mainly water, and acrylic acid and n-butanol; h) feeding the separated organic phase to the esterification reactor; i) feeding the separated aqueous phase to the hydrolysis reactor; j) removing from the hydrolysis reactor from 20 to 70% by weight, based on the total aqueous and organic charge stream, of the purge stream of the hydrolysis reactor; k) feeding up to 100% of the purge stream of the hydrolysis reactor to a thermal decomposition reactor, maintained at 90 to 1402C, a pressure of 20 to 200 mm Hg, and a residence time of 0.5 to 20 hours, based on the purge current of the fed reactor; 1) distill from the thermal decomposition reactor an upper stream comprising acrylic acid, n-butyl acrylate, n-butanol and water, while maintaining a thermal decomposition reactor liquid concentration of at least 7.5 % by weight of the acid; m) condensing the upper stream of the thermal decomposition reactor; n) recycling to the esterification reactor the condensed upper stream of the thermal decomposition reactor, comprising acrylic acid, n-butyl acrylate, n-butanol and water; o) distilling from the esterification reactor, concurrently with the preceding steps c) to n), a vaporized mixture comprising acrylic acid, n-butyl acrylate, n-butanol and water; p) condensing the vaporized mixture to supply a first condensate comprising an organic phase and an aqueous phase; q) returning from 0 to 30 percent of the organic phase to a drag separator, placed on the esterification reactor; Y r) feeding from 70 to 100 percent of the organic phase and from 50 to 100 percent of the aqueous phase to a separation column of acrylic acid; s) distilling from the acrylic acid separation column, at a pressure of 35 to 800 mm Hg, in an aqueous mode and at an aqueous reflux ratio, from 8.5: 1 to 17: 1, a top mixture comprising a mixture azeotro-butanol of butanol, butyl acrylate and water; t) removing a bottom stream rich in acrylic acid from the distillation column; u) recycling the acrylic acid-rich background stream from the acrylic acid separation column to the esterification reactor; v) condensing the upper mixture to supply a second condensate; w) separating the second condensate into an organic phase rich in butyl acrylate and a separate aqueous phase; and x) removing the organic phase rich in butyl acrylate, substantially free of acrylic acid. 26. The method according to claim 25, further comprising the steps of: a) omit steps p), q) and r); Y b) feeding 100 percent of the vaporized mixture directly to the separation column of acrylic acid from step s) and distilling, in which the ratio of aqueous reflux is 13: 1 17: 1. 27. The method according to claim 25, wherein the acid catalyst is selected from the group consisting of sulfuric acid, a sulfonic acid, and a strong acid ion exchange resin. 28. The method according to claim 27, wherein the sulphonic acid is selected from the group consisting of methanesulfonic acid, benzenesulfonic acid and toluenesulfonic acid. 29. The method according to claim 25, wherein the distillation in the separation column of the acrylic acid is carried out at a pressure of 90 to 135 mm Hg. 30. The method according to claim 25, wherein the ratio of aqueous reflux is 8.5: 12.5. 31. The method according to claim 25, wherein the total aqueous and organic charge stream is fed to a hydrolysis reactor comprising a multi-plate reactive distillation column. 32. The method according to claim 25, wherein the total aqueous and organic charge stream is fed to the hydrolysis reactor under continuous mixing conditions. The method according to claim 25, wherein the additional heavy ends comprise up to 80% by weight of the total, aqueous and organic charge stream. 34. The method according to claim 25, wherein the acid catalyst and the mineral acid are each sulfuric acid. SUMMARY OF THE INVENTION An improved process for producing n-butyl acrylate in high yields and high purity, substantially free of acrylic acid, incorporates one or more of the following new process component, in an esterification process, catalyzed with acid, to produce n-butyl acrylate: 1. A hydrolytic recovery component, in which heavy-end adducts, produced during acid-catalyzed esterification, are hydrolyzed, recovered and recycled as valuable reagents from a recovery hydrolytic unit (" HRU ").

2. A thermal decomposition reactor component, preferably used with the HRU unit, in which additional valuable reagents are recovered and recycled after treatment in the thermal decomposition reactor; Y

3. a new distillation component, in which the crude stream of n-butyl acrylate is efficiently distilled in an aqueous mode, through a separation column of acrylic acid, supplying the n-butyl acrylate substantially free of acrylic acid and with high performance. The first two components are also applicable to acid-catalyzed processes, which produce the C1-C4 alkyl acrylates. A continuous process, which produces n-butyl acrylate, which incorporates all the new components of the process, is also disclosed.