MXPA00004663A - Process for preparing gamma-butyrolactone, butane-1,4-diol and tetrahydrofuran - Google Patents

Abstract

A process for the production of at least one C4 compound selected from 1,4-butanediol,&ggr;-butyrolactone and tetrahydrofuran, in which a solution of maleic anhydride in a high boiling ester is esterified with a C1 to C4 alcohol to form the corresponding di-(C1 to C4 alkyl)maleate, which is hydrogenated to form the at least one C4 compound. The high boiling ester has a boiling point at atmospheric pressure which is about 30°C higher than that of the di-(C1 to C4 alkyl)maleate and is selected from di-(C1 to C4 alkyl)esters of alkyl dicarboxylic acids containing up to 13 carbon atoms, mono- and di-(C10 to C18 alkyl)esters of maleic acid, fumaric acid, succinic acid, and mixtures thereof, (C1 to C4 alkyl)esters of naphthalenemonocarboxylic acids, tri-(C1 to C4 alkyl)esters of tricarboxylic acids, and di-(C1 to C4 alkyl)esters of isophthalic acid.

Description

PROCEDURE FOR THE PREPARATION OF RANGE- BUTIROLACTONE. BUTANI-1.4-D1OL AND TETRAHIDROFURANO The present invention relates to the production of butane-1,4-diol, β-butyrolactone and tetrahydrofuran. Butane-1,4-diol, together with varying amounts of β-butyrolactone and tetrahydrofuran, can be produced by the hydrogenolysis of maleic acid diesters, fumaric acid and mixtures thereof. A. The main use of butane-1,4-diol is as a feedstock for the plastics industry, particularly for the production of polybutylene terephthalate. It is also used as an intermediate for the production of? -butyrolactone and of important solvent, tetrahydrofuran. The diesters of fumarate and maleate used as feedstock for the production of butane-1,4-diol by means of said hydrogenolysis route are conveniently prepared from maleic anhydride, which itself is produced by a phase oxidation. of a charge of hydrocarbons, such as benzene, mixed C4 olefins, or n-butane, in the presence of a partial oxidation catalyst. In the partial oxidation of benzene a support catalyst is typically used, vanadium pentoxide promoted with MoO3 and possibly other promoters. The reaction temperature is between 400 ° C to about 455 ° C and the present reaction is between about a bar at about 3 bar, while about 4 times the theoretical amount of air is used for the purpose of maintaining outside the explosive limits. The contact time is around 0.1 second. When the feedstock is a mixed C4 olefin charge, ie a charge of butene mixture, then the partial oxidation catalyst can be supported in vanadium pentoxide or alumina. Typical reaction conditions include the use of temperatures ranging from about 425 ° C to about 485 ° C and a pressure ranging from about 1.70 bar to about 2.05 bar. The ratio of air volumes to butenes can be around 75: 1 in order to stay below the explosive limits. Alternatively, it is possible, according to the most modern practice, to design the plant in such a way that a satisfactorily safe operation can be achieved, in spite of the fact that the mixture of air supply and butenes is within the flammable limits. In the case of using n-butane as feedstock, the catalyst is typically a vanadium toxicant and the reaction conditions include the use of a temperature ranging from about 350 ° C to about 450 ° C and a pressure that It goes from around 1 bar to around 3 bar. The volumetric relationship between air and n-butane can be from about 20: 1, even though this may be within the limits of inflammation. A reactor design for said partial oxidation reactions comprises vertical tubes surrounded by a jacket through which a molten salt is circulated for the purpose of controlling the reaction temperature. In each case a hot vapor reaction mixture is recovered at the outlet end of the reactor which comprises maleic anhydride vapor, water vapor, carbon oxides, oxygen, nitrogen and other inert gases, in addition to organic impurities such as formic acid , acetic acid, acrylic acid and the unconverted hydrocarbon feed charge. One way to recover maleic anhydride from said reaction mixture is to cool it to about 150 ° C using a steam production stream and then further cool it to about 60 ° C with water in order to condense part of the maleic anhydride, typically around from 30% to about 60% of the maleic anhydride present. The rest of the stream is then separated with water. The separation with water or with an aqueous solution or sludge is described as for example, in the patent of E.U.A. No. 2,638,481. Said separation results in a production of a solution of maleic acid which is then dehydrated by distillation with silene, for example, in order to remove the water and reform the anhydride. A disadvantage of said method is that an unacceptable proportion of the product remains in the vapor phase. In addition, part of the maleic acid is inevitably isomerized in fumaric acid. The by-product of fumaric acid represents a loss of titratable maleic anhydride and is difficult to recover from the disposition of the process since it tends to form crystalline masses that give rise to process problems. Due to this problem of isomerization a variety of other anhydrous separation liquids have been proposed. For example, dibutyl phthalate has been proposed as a separation liquid in the English patents Nos. 727,828.; 763,339 and 768,551. The use of dibutyl phthalate containing up to 10% by weight of italic anhydride is also suggested in the US patent. No. 4,118,403. The patent of E.U.A. No. 3,818,680 discloses the use of a normally liquid intramolecular carboxylic acid anhydride, such as a substituted subclinical C? 2.? 5 alkenyl branched chain anhydride, for the absorption of maleic anhydride from a reaction mixture leaving the oxidation reactor. partial. Triquesyl phosphate has been proposed for this purpose in French Patent No. 1, 025, 014. The use of dimethyl terephthalate for this task was suggested in Japanese Patent 32-8408 and Dibutyl Maleate in Japanese Patent No. 35-7460. A high molecular weight wax was described as a separation solvent in the U.S. patent. No. 3,040,059, while the patent of E.U.A. No. 2,893,927 proposes separation with diphenyl pentachloride. The use of an aromatic hydrocarbon solvent having a molecular weight between 150 and 400 and a boiling point above 240 ° C at a temperature above the dew point of water in the vaporous reaction mixture, for example dibensilbenzene, it is described in French Patent No. 2,285,386. The absorption of maleic anhydride from a vaporous partial oxidation reaction mixture into dimethyl bensophenone followed by distillation is described in US Pat. No. 3,850,758. The polymethylbensophenones, at least a portion of which contain at least 3 methyl groups, can be used as liquid absorbent for maleic anhydride according to the patent of E.U.A. No. 4,071, 540. Dialkyl phthalate esters having C4 to C8 alkyl groups and a total of 10 to 14 carbon atoms in both alkyl groups are proposed for the absorption of maleic anhydride from a reaction mixture in the US patent. No. 3,891, 680. An ester of cycloaliphatic acid, for example dibutyl hexahydrophthalate, are suggested as absorption solvents for maleic anhydride in the patent ZA-A-80/1247. It has also been proposed to effect a direct condensation of maleic anhydride from a reaction mixture leaving the partial oxidation reactor. However, this procedure is inefficient because an unacceptable proportion of the maleic anhydride remains in the vapor phase. The maleic anhydride product recovered following the condensation or by means of separation or absorption and distillation is then reacted with suitable C-C to C alkanol, such as methanol or ethanol, to leave a yield of the corresponding di- (C? A C alkyl) ) maleate. This di- (Ci to C4 alkyl) maleate may contain a minor amount of the corresponding di- (C? A C4 alkyl) fumarate, in addition to portions of the corresponding mono- (C? A C4 alkyl) maleate and / or fumarate. This is then subjected to hydrogenolysis to leave as a yield a mixture of butane-1,4-diol, together with variable amounts of β-butyrolactone and tetrahydrofuran, depending on the hydrogenolysis conditions that are selected, and of the alkanol Ci to C4 which can be used. be recycled to produce additional di- (CT to C4 alkyl) maleate. Processes and plants for the production of dialkyl maleates from maleic anhydride are described, for example, in U.S. Patent No. 4,795,824 and in WO-A90 (08127). This last mentioned document discloses a column reactor having a plurality of esterification trays, each of which has a predetermined liquid charge and contains a charge of solid esterification catalyst, such as an ion exchange resin containing pendant groups of sulfonic acid. A liquid phase containing, for example, a carboxylic acid component flows down the column from an esterification tray to the lower next tray against a vapor rising fluid stream of the lower boiling component of the esterification reagents, typically the alkanol Ci to C4. And then the esterification is removed from the top of the column reactor in a vapor stream, while the ester product is recovered from the reactor sump. As the liquid flows down the trays this liquid progressively encounters drier reaction conditions and the esterification reaction is further carried out towards 100 × 100 ester formation. This column reactor can be followed by a finishing reactor operating under liquid phase reaction conditions, with the ester-containing stream coming from the bottom of the column reactor mixed with additional C1-C4 alkanol before its admission to the reactor. of polishing. When used for the production of di- ^ a C4 alkyl) maleate, the column reactor can be preceded by a non-catalytic monoesterification reactor, in which the maleic anhydride is reacted with the alkanol Ci to C4 in the absence of a catalyst added to form the mono (C1 to C4 alkyl) maleate. The hydrogenation of di-alkyl maleates to produce butane-1,4-diol is further described in U.S. Patent Nos. 4,584,419 and 4,751, 334 and in WO-A-88/00937, which were incorporated herein by reference. reference mode. It would be desirable to simplify the production of butane-1,4-diol, β-butyrolactone, and tetrahydrofuran from maleic anhydride via the hydrogenolysis pathway of di- (C? To C4 alkyl) maleate. In particular, it would be desirable to reduce the capital cost of construction of said plant and also reduce its operating costs so as to produce butane-1,4-diol, β-butyrolactone, and tetrahydrofuran, in a more easily accessible form. It is therefore an object of the present invention to simplify the production of butane-1,4-diol, β-butyrolactone, and tetrahydrofuran from maleic anhydride via the hydrogenolysis pathway of di- (C? A C4 alkyl) maleate. An additional object is to reduce the capital cost of construction of the aforementioned plant by significantly reducing the numbers of the distillation columns and the quantity of other equipment required. In addition, it seeks to reduce the operating costs of the butane-1,4-diol production plant in order to make butane-1,4-diol, β-butyrolactone and tetrahydrofuran in the most accessible form.

According to the present invention there is provided a process for the production of at least one C4 compound selected from butane-1,4-diol, β-butyrolactone, and tetrahydrofuran, which includes the step of hydrogenation in the vapor phase of a di- (C? a C4 alkyl) maleate in the presence of a particulate ester hydrogenation catalyst, which process comprises: (a) contacting a vaporous stream containing maleic anhydride vapor, water vapor, and carbon monoxide in an absorption zone with high-boiling ester as a solvent so as to form a solution of maleic anhydride in the high-boiling ester, said high-boiling ester having a boiling point at atmospheric pressure which is at least about 30 ° C higher than that of the di- (C? a C4 alkyl) maleate, and is selected from di- (Ci to C4 alkyl) esters of alkyl dicarboxylic acids containing up to 13 carbon atoms. carbon, mono-di- (C? 0-i8 alkyl) esters of maleic acid, fumaric acid, succinic acid and mixtures thereof, (Ci to C4 alkyl) esters of naphthalemonocarbonoxylic acids, tri- (C? to C4 alkyl) esters of tricarboxylic aromatic acids, and di- (C? a C4 alkyl) esters of isophthalic acids; (b) recovering a waste gas stream from the absorption zone; (c) reacting the maleic anhydride in the maleic anhydride solution of step a) under esterification conditions in an esterification zone with (C C4 alkanol to form the corresponding di- (C? a C alkyl) maleate; ) recovering from the esterification zone a solution of the di- (C-? a C4 alkyl) maleate in the high-boiling ester, (e) contacting the solution of the di- (C? a C4 alkyl) maleate in the high-boiling ester with a gaseous stream containing hydrogen so as to separate from it the di- (C-? a C4 alkyl) maleate and to form a vaporous stream comprising hydrogen and di- (C? a C4) alkyl) maleate, (f) contacting material of the vaporous stream of step (e) in a hydrogenation zone under ester hydrogenation conditions with a heterogeneous ester hydrogenation catalyst so as to convert the di- (C? a C4) alkyl) maleate by hydrogenation in at least one compound to C4 selected from butane-1, 4-diol, β-butyrolactone, and tetrahydrofuran; and (g) recovering from the hydrogenation zone a product stream containing said at least one C4 compound. Preferably in said process the alkanol Ci to C4 is methanol or ethanol and the di- (C? A C4 alkyl) maleate is dimethyl maleate or diethyl maleate. The use of methanol such as C4 alkanol and dimethyl maleate such as di- (C? A C4 alkyl) maleate is especially preferred. The vapor stream of step (a) of the process of the invention is preferably produced by partial oxidation of a hydrocarbon feed charge in the presence of a partial oxidation catalyst using molecular oxygen, typically in the air form. The hydrocarbon feedstock may be benzene, or a mixture of a C4 olefin stream, but more preferably is n-butane. The use of n-butane as a hydrocarbon charge is currently preferred based on the cost of the same since this is a more economic charge than benzene or butenes. Hence, in the process of the invention, the charge used for the production of the maleic anhydride vaporous stream of step (a) is more preferably n-butane and the catalyst is preferably vanadium pentoxide. Typical partial oxidation conditions in this case include the use of a temperature ranging from about 350 ° C to about 450 ° C and a pressure ranging from about 1 bar to about 3 bar, a ratio between the air and the n-butane which is between about 15: 1 to about 50: 1, for example about 20: 1 and a partial oxidation catalyst comprising vanadium pentoxide; the contact time being typically from about 0.01 second to about 0.5 seconds, for example about 0.1 second. The partial oxidation of the hydrocarbon feedstock is conveniently conducted in a reactor comprising vertical tubes surrounded by a jacket through which a molten salt is circulated for the purpose of controlling the reaction temperature. The vaporous stream from the partial oxidation reactor can then be cooled by external cooling with boiler feed water to generate steam, and possibly also by additional external cooling with cooling water to a temperature that is in the range from around from 60 ° c to around 160 ° C. In step (a) of the process of the present invention, the vaporous stream of maleic anhydride is preferably contacted with the high-boiling ester at a temperature ranging from about 60 ° C to about of 160 ° C, preferably from about 80 ° C to about 120 ° C, and a pressure ranging from about 1 bar to about 3 bar so as to form a solution comprising maleic anhydride in the high-point ester of boiling. The contacting can be carried out by bubbling the vaporous stream through the high-boiling ester body. Alternatively, the high-boiling ester can be sprayed into the vaporous stream. Countercurrent contacting devices may also be employed wherein the rising vapor stream is contacted with a downstream stream of a high boiling ester in a gas-liquid contacting device, such as a compact separation tower. or a separation tower provided with trays. In this step the high boiling ester will typically be at a lower temperature than the vaporous stream so that the latter is cooled. The resulting solution of the maleic anhydride in the high-boiling ester may be in the range from about 100 g / l to about 400 g / l. In the high-boiling ester it has a boiling point at atmospheric pressure which is at least about 30 ° C higher than, and preferably at least about 60 ° C to about 70 ° C higher than that boiling point of di- (C? a C4 alkyl) maleate. As examples of esters of alkyl dicarboxylic acids containing up to 13 carbon atoms from which a suitable high-boiling ester can be selected, di-methyl, di-ethyl, di-n- or iso-propyl can be mentioned. , di-n-, -sec-, or iso-butyl esters of suberic acid, acelaic acid, cebacic acid, undecanoic acid, dodecanodoic acid, and tridecanodoic acid. It is preferred that the alkyl portion in said ester be derived from the same alkanol as the C 1-4 alkanol used in the esterification step (c). In this way any transesterification reaction that may occur does not generate additional esters. Thus, when the alkanol used is methanol and methane, di-alkyl is di-alkyl maleate, any ester used as the high-boiling ester is also preferably a dimethyl ester, such as dimethyl sebacate. The high-boiling ester may alternatively be selected from mono- (C? 0 to C? 8) esters of one of the C4-dicarboxylic alkyl acids, ie maleic acid, fumaric acid and subclinical acid and mixtures thereof. Examples of such esters include esters and mixtures thereof derived from n-decyl alcohol, lauryl alcohol, myristyl alcohol, cetyl alcohol, stearyl alcohol, and eicosanol. In this case, some hydrolysis of the high-boiling ester may occur in the esterification zone resulting in the release of lesser propulsion of the corresponding C-10-C18 alkyl alcohol. In addition some transesterification may occur in the esterification zone resulting in the formation of a minor amount of a mono- (C -? - C alkyl) mono- (C? O-C-? 8 alkyl) ester of C4 alkyl dicarboxylic acid, for example, if dilauryl maleate is used as the high boiling ester and if methanol such as C1 to C4 alkanol is used, then a minor amount of lauryl methyl maleate may be formed by transesterification. However, the formation of these minor by-products is not advantageous even if the high-boiling ester used in step (a) comprises a recycled material resulting from step (e) because any free C10 to C alkanol? they can react with fresh maleic anhydride in step (a) to form a mono- or di- (C 10 -C 18 fresh alkyl maleate.) In addition any mono- (C 1 -C alkyl) mono- (C 10 -C 18 alkyl) acid ester C4 alkyl dicarboxylic can undergo a transesterification in the next situation by passing through the esterification zone to form the ded solvent or the ded di- (C? A C4 alkyl) maleate.The high boiling ester can alternatively be selected from (Ci to C4 alkyl) esters of naphthalene monocarboxylic acids, such as methyl naphthalene-2-carboxylate, of tri- (C1 to C4 alkyl) esters of aromatic tricarboxylic acids, such as trimethyl benzene-1, 2, 4-tricarboxylate or pa of di- (C? a C4 alkyl) esters of isophthalic acid, such as dimethyl isophthalate. The high-boiling ester used in step (a) conveniently comprises material resulting from the step of hydrogen removal (e). Hence, it may already contain some di- (C? -a C4 alkyl) maleate. Taking into account that appropriate conditions are adopted in step (a), the gas stream recovered in step (b) of the process of the invention can be essentially free of maleic anhydride. The esterification of the maleic anhydride with the C 1 -C 4 alkanol is carried out in step (c) in an esterification zone. This may comprise a non-catalytic reactor, in which the maleic anhydride in the solution in the high-boiling ester undergoes a reaction in the absence of added catalyst with the C1 to C4 alkanol to form the corresponding mono- (C? A) C4 alkyl) maleate. The reaction is: wherein R is a C1 to C alkyl radical. Some conversion of the mono (C1 to C alkyl) maleate to the corresponding di- (C? A C4 alkyl) maleate may also occur. The reaction in question is: where R is as defined above. Said non-catalytic reactor can be operated under monoesterification conditions typically comprising the use of a temperature ranging from about 65 ° C to about 260 ° C and a pressure ranging from about 1 to about 50 bar. This can be followed by a step of catalytic esterification. For example, the catalytic esterification step may comprise a plurality of agitated tank reactors, such as the one described in U.S. Patent No. 4,795,824. Preferably, however, the step of catalytic esterification comprises a column reactor of the type described in WO-A-90/03127. In this case the non-catalytic monosterification step may comprise a stirred tank reactor or a column reactor containing one or more trays that do not contain any esterification catalyst and which is not fed from the bottom with methanol or other steam from the bottom. C1-C4 alkanol, while the maleic anhydride solution from step (a) is fed down through the column reactor. If the catalytic esterification step comprises a column reactor of the type described in WO-A-90/03127, then the solution of maleic anhydride (or a solution comprising the corresponding mono- (Ci to C4 akyl) maleate (if a monosterification step is used) in the high-boiling ester is fed to the upper esterification tray of the column reactor, while an excess of alkanol vapor Ci to C4 is fed to the bottom of the reactor. of column the esterification trays each contain a charge of a solid esterification catalyst Each tray has a steam ingress medium to allow steam to enter the tray below and stir the liquid mixture and solid esterification catalyst in a zone of turbulence on the tray and keep the particles of the catalyst in suspension, in order to avoid the damage of "hot spots" that are formed in the tray Through the formation of pockets of deposited catalyst particles, the floor of each tray is preferably designed so as to be inclined towards the turbulence zone with an inclination that exceeds the angle of repose of the catalyst particles under the liquid. In addition each esterification tray has a descending ingress means which allows the liquid, but not the catalyst particles, to flow down from the tray to the next lower tray. Said means of entry are usually provided with a mesh or filter to prevent the particles of the catalyst from passing down through them. Typical reaction conditions in the column reactor include the use of a temperature and pressure under which Ci-C4 alkanol is distilled. Such temperature and pressure conditions will vary depending on the Ci-C4 alkanol selected, but typically includes the use of a temperature ranging from about 65 ° C to about 185 ° C and a pressure ranging from about 1 bar to about of 3 bar. A typical solid esterification catalyst is the ion exchange resin sold under the designation Amberlyst 16, from Rohm and Hass, (Great Britain) Limjted from Linnig House, 2 Mason's Avenue, Croydon CR9 NB England or that commercially available as ion exchange resin DPT1 from Kvaemer Process Technologi Limited of The Technologi Center, Princeton Drive, Thornaby, Stockton-on-Tees TS17 6PY, England. When passing up the column from one esterification tray to the next higher, the upward flow of the alkanol vapor Ci-C4 carries with it water of esterification. In this way the liquid containing the di- (Ci-C4 alkyl) maleate which passes descently through the reactor of the column from an esterification tray to the lower immediate tray finds drier conditions as it continues to fall down the column. In this way the esterification reaction leads to the information of di (Ci-C4 alkyl) maleate and is thus constantly generated up to a conversion of 100% in di- (C1-C4 alkyl) maleate. Any acid by-product, such as acetic acid or acrylic acid, which is also present in the vaporous stream from the partial oxidation reactor, together with any maleic acid or fumaric acid present in the solution supplied to the esterification zone, will suffer a conversion to the corresponding Ci-C4 alkyl ester or diester, whichever is the case. The vapor phase current emerging from the uppermost esterification tray comprises Ci-C alkanol vapor and water vapor; it may also include smaller portions of its product, such as di- (C?-C4 alkyl) ether, as well as di- (C?-C 4 alkyl) maleate and C 1 -C 4 alkyl acrylate portions. An additional tray or additional trays may be provided above the uppermost esterification tray to act as a washing column shape for the purpose of returning the di- (C?-C4 alkyl) maleate to the esterification trays. The resulting vapor stream, which is now essentially free of di- (C? -C4 alkyl) maleate, exits at the top of the column. From the bottom of the column reactor, a liquid stream comprising a solution of di- (C?-C4 alkyl) maleate in the high-boiling ester is collected. This is essentially acid free. If desired this liquid can be mixed with a further C1-C4 alkanol and passed through a finishing reactor containing a bed of a solid esterification catalyst operating under liquid phase operating conditions. Such conditions typically include the use of a temperature ranging from about 65 ° C to about 135 ° C and a pressure ranging from about 1 bar to about 3 bar. A typical solid esterification catalyst is the ion exchange resin marketed under the designation Amberlyst 16, from Rohm and Hass, (Great Britain) Limited of Lennig House, 2 Mason's Avenue, Croydon CR9 NB, England; or one that is commercially available as a DPT1 ion exchange resin from Kvaemer Process Technology Limited of The Technology Center, Princeton Drive, Thornaby, Stockton-on-Tess TSL7 6PY, England. In step (e) of the process of the invention a gas stream comprising hydrogen is passed through the solution of the di- (C? -C4 alkyl) maleate. The hydrogen evolution step is preferably carried out substantially at the pressure or at a pressure slightly higher than the inlet pressure to the ester hydrogenation zone. The hydrogen evolution step is preferably carried out in a manner similar to substantially the desired inlet temperature to the hydrogenation step or at a temperature slightly below this temperature, for example from about 5 ° C to about 20 ° C. below this temperature. The temperature can then be raised to the desired inlet temperature by further mixing hot gas containing hydrogen, which has the additional benefit of diluting the ester-containing vaporous stream, and thus ensuring that it is at a temperature above its dew point, preferably at least about 5 ° C higher than its dew point. The hydrogenation step is advantageously carried out in the vapor phase, using a heterogeneous ester hydrogenation catalyst. Typical ester hydrogenation catalysts include reduced promoted copper catalysts, for example reduced copper cronite catalysts such as that sold under the designation PG81 / 1 of Davy Process Technology Limited DE 30 Eastourne Terrace, London W2 6LE. The catalyst particles preferably have a particle size that ranges from about 0.5 mm to about 5 mm. The particles can be in any convenient form for example spheres, granules, rings or forms of mounting. When using a fixed bed of catalyst the reactor can be a shell and tube reactor, which can be operated substantially in isothermal form; however, an adiabatic reactor is preferable. The use of an adiabatic reactor is advantageous since the capital of its cost is much lower than that of the shell and tube reactor and that is generally much easier to load with the chosen catalyst. Hydrogenation is carried out at an elevated temperature for example around 150 ° C to about 240 ° C and at a pressure ranging from about 5 bar to about 100 bar preferably from about 50 bar to about 70 bar. From the hydrogenation zone a mixture of hydrogenation product containing, in addition to the C1-C4 alkanol, also butane-1,4-diol, and some tetrahydrofuran and? -butyrolactone. Even if the primary product of interest is butane-1,4-diol, the presence of these minor amounts of tetrahydrofuration and β-butyrolactone is not disadvantageous since these compounds are important chemicals from a commercial point of view and therefore it is economical to recover the same in their pure form. If desired and the α-butyrolactone can be recycled to the hydrogenation zone to produce additional butane-1,4-diol. In addition, the mixture of the hydrogenolysis product will normally contain minor amounts of the corresponding di- (C? -C4 alkyl) succinate, n-butanol the corresponding dialkyl alkoxysuccinate, for example dimethyl methoxysuccinate if the di-C1-C4 alkanol is methanol and water. For further details regarding the hydrogenation of a di- (C 1 -C 4 alkyl) maleate and the subsequent purification of the resulting crude hydrogenation product mixture, reference may be made to U.S. Patent No. 4,584,419, to WO-A- 86/03189, to WO-A-88/0937, to North American patents 4,767,869; 4,945,173; 4,919,765; 5,254,758 and 5,310,954, and WO-A-91/01960. In order that the invention can be clearly understood and easily carried out, a plant for the production of butane-1,4-diol, as well as some γ-butyrolactone and tetrahydrofuran, will be described below, using a preferred process according to the present invention, this description being only by way of example with reference to the accompanying figure, which is a flow diagram of the plant.

Referring to the figure, the n-butane is supplied by line 1 at a pressure ranging from 1 to 3 bar and a temperature of 400 ° C to a partial oxidation plant 2 that also receives air on line 3. partial oxidation plant also receiving air on line 3. Partial oxidation plant 2 is of a conventional design and includes a partial oxidation reactor comprising tubes packed with a partial oxidation catalyst consisting of vanadium pentoxide packed inside tubes provided with a jacket or cover through which molten salt is circulated for the purpose of controlling the temperature. The partial oxidation reactor is operated with a ratio between air and n-butane of 20: 1. A stream of vaporous partial oxidation product is cooled by external cooling with boiler feed water to generate steam and then with cooling water to reduce its temperature to 138 ° C. This is collected from plant 2 on line 4. It contains 2.9% by weight of maleic anhydride, 5.8% by weight of water, 1.3% by weight of carbon dioxide, 1.0% by weight of carbon monoxide, 0.01% by weight. weight of acetic acid, 0.01% by weight of acrylic acid, 15.7% by weight of oxygen and the balance comprising essentially nitrogen and other inert gases. This is fed to the bottom of a separation tower 5, up to which it passes against a downward spray of dimethyl sebacate which is supplied at a temperature of about 68 ° C from line 6. The waste gas stream separated containing 0.03% by weight of maleic anhydride leaves the top of the separation tower 5 through the gas vent line 7 and passes to a waste gas burner. From the bottom of the separation tower 5 a liquid stream is recovered in line 8 comprising a solution of approximately 22% by weight of maleic anhydride and 0.04% by weight of acrylic acid in dimethyl sebacate. This is supplied to the top of a column reactor 9 of the type described in WO-A-90/08127. This comprises a number of esterification trays which are mounted on top of each other and each contains a charge of a solid esterification catalyst such as the Amberlyst 16 resin or the DPT1 ion exchange resin, and each has a vapor ingress. for a rising steam and a liquid lowering to allow the liquid to fall down from an esterification tray to the next lower tray. Methanol vapor is supplied to the bottom of the column reactor via line 10. The esterification water is removed in the vapor stream leaving the column reactor on line 11. Column reactor 9 is operated at a temperature which goes from around 110 ° C to around 125 ° C and at a pressure that goes from around 1 bar to around 3 bar. The residence time in the column reactor is around 3 hours. Normally the temperature in the upper tray will be somewhat higher (for example around 125 ° C than that of the lowermost tray (around 115 ° C) A solution containing about 250 g / l of dimethyl maleate in sebacate dimethyl is removed from the bottom of the column reactor 9 by line 12 and pumped to the area near the top of a release column 13 which is operated at a temperature of 170 ° C and at a pressure of 61.02 bar. 13 has a distillation tray number above the injection point of the dimethyl maleate solution within the column 13 so as to reduce the entrainment of the high-boiling dimethyl ester sebacate in the upper stream from the column. The solution of the dimethyl maleate in the dimethyl sebacate flows down into the stripping column 13 against an upward stream of hydrogen from line 14. In the after-column yield 13 against a rising stream of hydrogen from line 14. The stripped dimethyl sebacate is recycled from the bottom of the stripping column 13 by means of a line 6 towards the top of the separation tower 5. From the top from the stripping column 13 a stream of almost saturated vapor mixture, comprising dimethyl maleate in hydrogen, emerges through line 15, with a molar ratio between hydrogen and dimethyl maleate of about 320: 1. This steam mixing stream is at a temperature ranging from about 180 ° C to about 195 ° C and at a pressure of 62 bar. This is diluted with additional hot hydrogen coming from line 16 at a temperature ranging from about 180 ° C to about 195 ° C to result in a vaporous stream with a molar ratio between the hydrogen and the dimethyl maleate of around 350: 1 and is at least around 5 ° C above its dew point. This vaporous mixture passes forward through line 17 to the hydrogenation plant 18 which includes an adiabatic reactor packed with a reduced copper-based catalyst, for example, a reduced copper chromite catalyst, and operated at an inlet temperature. of 173 ° C, an inlet pressure of 61.02 bar and an outlet temperature of 190 ° C. The feeding speed range of dimethyl maleate corresponds to a liquid hourly space velocity of 0.5 h "1. The plant also includes a purification section in which the mixture of the crude hydrogenation product is distilled in various stages to leave butane-1 , Pure 4-diol in line 19. The lines for the separate recovery of β-butyrolactam and tetrahydrofuran are indicated respectively with reference numerals 20 and 21. Fresh dimethyl sebacate solvent can be added via line 22 while a purge stream of the recycled stream of solvent can be collected on line 23. The dimethyl sebacate used as solvent can be replaced for example by methyl naphthalene-2-carboxylate, trimethylbenzene-1, 2,4-tricarboxylate or dimethyl Softalate Alternatively, dimethyl sebacate can be replaced as a high-boiling ester by a di- (C? 0-C? S alkyl) maleate, fumarate, or suc or a mixture of two or more thereof, optionally in the mixture with the corresponding mono- (C-? oC? 8 alkyl) maleate, fumarate, or succinate or a mixture of two or more thereof, and / or with the corresponding free acid or mixture of acids, ie maleic acid, fumaric acid and / or succinic acid. Typically the high-boiling ester in that case predominantly comprises the diester or diester mixture, with no more than a minor amount, typically less than about 5 mol% each, of the corresponding monoester or monoester mixture and / or corresponding acid or mixture of acids. As an example of such a high boiling ester, mention may be made of dilauryl maleate, which may contain minor amounts, preferably less than about 1 mole% each and even more preferably less than about 0.25 molar% each, of 1 or more of the following, dilauryl fumarate, dilauryl succinate, monolauryl maleate, monolauryl fumarate, monolauryl succinate, maleic acid, fumaric acid and succinic acid. Further, as a result of the transesterification in the column reactor 9, the recirculation stream in line 6 can in this case contain a significant amount of methyl lauryl maleate, for example, up to about 10 mole% and / or more typically no more than about 5 mol% and frequently less than about 1 mol%, furthermore it may contain lower amounts, typically less than about 1 mol% each, and even more preferably less than about 0.25 mol% each, of lauryl alcohol, methyl lauryl fumarate, and methyl lauryl succinate. The invention will be additionally illustrated in preference to the following examples.

EXAMPLE 1 98. 0 grams of methanol and 32.0 grams of maleic anhydride were reacted together in a rounded bottom flask. To the resulting mixture was added 130 gr of dimethyl sebacate and 26.0 grams of a DPT1 ion exchange resin. (The ion exchange resin DPT1 is a macroreticular ion exchange resin containing sulfonic acid groups and is marketed by Kvaemer Process Technology Limited of the Technology Center, Princeton Drive, Thornaby, Stockton-on-Tess TS17 6PY, England) . The mixture was heated to 110 ° C and dry methanol was fed to the flask at a rate of 6 moles hr "1. Any unconverted methanol, together with a byproduct of water and dimethyl ether, was recovered in the top after the Conversion The conversion of monoethyl maleate to dimethyl maleate was followed by periodically withdrawing samples from the reaction flask and analyzing such samples to determine the acid content The experiment was continued until the acid level fell to less than 0.5% by weight The results are shown in C10-C18 alkyl) maleate, fumarate, or succinate or a mixture of two or more thereof, Table 1 which is seen below.

TABLE 1 NOTE: DME = dimethyl ether The final sample was also analyzed using a capillary gas chromatographic technique. This allowed to determine the amount of hydrolysis of dimethyl sebacate that had been produced. The results are given in table 2 below.

TABLE 2 The results indicate that a small amount of dimethyl sebacate was hydrolyzed to monomethyl sebacate and sebacic acid when dry methanol was used.

EXAMPLE 2 32. 0 grams of maleic anhydride and 98.0 grams of dry methanol were reacted together and then heated at 110 ° C in 130 grams of dimethyl sebacate in the presence of 26.6 grams of a DPT1 ion exchange resin according to the procedures of Example 1 Methanol containing 30 mol% of water was then fed to the flask in the manner described in Example 1 at a range of 6 mol hr "1 until the system reached equilibrium, then the amount of water in the methanol fed to the flask was reduced to 15 mol% and the system again reached equilibrium.The experiment was continued using methanol containing 5 mol% of water and finally dry methanol was used.The results obtained are indicated in table 3 below.

TABLE 3 NOTE: acidity of the product means percentage of monomethyl maleate.

Claims (25)

NOVELTY OF THE INVENTION CLAIMS

1. A process for the production of at least one C4 compound selected from butane-1,4-diol, β-butyrolactone, and tetrahydrofuran, which includes the step of hydrogenation in the vapor phase of a di- (C? A) C4 alkyl) maleate in the presence of a particulate ester hydrogenation catalyst, which process comprises: a) contacting a vaporous stream containing maleic anhydride vapor, water vapor, and carbon oxide in an absorption zone with an ester of high boiling point as a solvent so as to form a solution of maleic anhydride in the high-boiling ester, said high-boiling ester having a boiling point at atmospheric pressure which is at least about 30 ° C higher than that of di- (C? A C alkyl) maleate, and is selected from di- (C1 to C4 alkyl) esters of alkyl dicarboxylic acids containing up to 13 carbon atoms, mono-di- (C? 0-18 alkyl) acid esters maleic acid, fumaric acid, succinic acid and mixtures thereof, (C1 to C4 alkyl) esters of naphthalemonocarbonoxylic acids, tri- (C? to C4 alkyl) esters of tricarboxylic aromatic acids, and di- (C? a C4 alkyl) esters of isophthalic acids; b) recover a waste gas stream from the absorption zone; c) reacting the maleic anhydride in the maleic anhydride solution of step (a) under esterification conditions in an esterification zone with (C1-C4 alkanol to form the corresponding di- (C? a C4 alkyl) maleate; ) recovering from the esterification zone a solution of the di- (C1 to C alkyl) maleate in the high-boiling ester; e) contacting the solution of the di- (C? a C alkyl) maleate in the high-boiling ester with a gaseous stream containing hydrogen so as to separate therefrom the di- (C? a C4 alkyl) maleate and to form a vaporous stream comprising hydrogen and di- (C1 to C alkyl) maleate; f) placing the material contact of the vaporous stream of step (e) in a hydrogenation zone under ester hydrogenation conditions with a heterogeneous ester hydrogenation catalyst in order to convert the di- (C? a C4 alkyl) maleate by hydrogenation in at least one C4 compound selected from butane-1,4-diol, β-butyrolactone, and tetrahydrofuran; and g) recovering from the hydrogenation zone a product stream containing said at least one compound C.

2. A process according to claim 1, wherein the C 1 -C 4 alkanol is methanol and the di- (C 1 -C 4 alkyl) is dimethyl maleate.

3. A process according to claim 1 or 2, wherein the vaporous stream of step (a) is produced by partial oxidation of a hydrocarbon feedstock in the presence of a partial oxidation catalyst using molecular oxygen.

4. A process according to claim 3, wherein the hydrocarbon feedstock is n-butane.

5. A process according to claim 4, wherein the partial oxidation catalyst comprises vanadium pentoxide and in which the partial oxidation conditions include the use of a temperature ranging from about 350 ° C to about 450 ° C, a pressure that goes from around 1 bar to around 3 bar, a relation between air and n-butane that goes from around 15: 1 to around 50: 1 and a contact time that goes from about 0.01 seconds to about 0.5 seconds.

6. A process according to any of claims 1 to 5, wherein in step (a) the vaporous stream of maleic anhydride is brought into contact with the high-boiling ester at a temperature that is in the range ranging from about 60 ° C to about 160 ° C and at a pressure ranging from about 1 bar to about 3 bar so as to form a solution comprising maleic anhydride in the high-boiling ester.

7. A process according to claim 6, wherein the step of contacting is carried out in a countercurrent contacting device wherein the rising vaporous current is brought into contact with a downward stream of solvent in a device of contact between gas and liquid.

8. A process according to any of the claims 1 to 7, wherein the solvent is an alkyl ester whose alkyl portion is derived from the same alkanol as the C 1 -C 4 alkanol used in the esterification step (c).

9. - A process according to any of claims 1 to 8, wherein the alkanol n ester C 1 -C 4 is methanol, the di- (C 4 -C 4 alkyl) maleate is dimethyl maleate and the high boiling solvent is also It is methyl ester.

10. A process according to claim 9, wherein the methyl ester is dimethyl sebacate.

11. A process according to claim 9, wherein the high-boiling ester is naphthalene-2-carboxylate, trimethylbenzene-1, 2,4-tricarboxylate or dimethyl isophthalate.

12. A process according to any of claims 1 to 7, wherein the high boiling ester comprises a di- (C? O-C18 alkyl) ester of maleic acid, fumaric acid, succinic acid or a mixture of two or more of them.

13. A process according to claim 12, wherein the high-boiling ester further comprises a mono- (C? 0-C-i8 alkyl) ester of maleic acid, fumaric acid, succinic acid or a mixture of two or more of them.

14. A process according to any of claims 1 to 13, wherein the high-boiling ester used in step (a) comprises recycled material resulting from the hydrogen evolution step (e).

15. A process according to any of claims 1 to 14, wherein the esterification zone comprises a non-catalytic reactor in which the maleic anhydride in the solution in the high-boiling ester is subjected to a reaction in the absence of added catalyst with the C 1 -C 4 alkanol to form the corresponding mono- (C 1 -C 4 alkyl) maleate.

16. A process according to any of claims 1 to 15, wherein the step of catalytic esterification comprises a column reactor provided with a plurality of esterification trays each of which has a charge of a solid esterification catalyst , has a means of rising vapor uptake to allow steam to enter the tray from below and to agitate the liquid mixture and solid esterification catalyst in a turbulence zone in the tray and to keep the catalyst particles in suspension , and a downstream intake means which allows the liquid, but not the catalyst particles, to flow down from the tray to the next lower tray, the column reactor being supplied below the lowermost esterification tray with a stream of Steam of (C1-C4 alkanol and to an upper esterification tray with a solution in the high-pun ester A boiling composition comprising a material selected from maleic anhydride, a mono- (C? -C alkyl) maleate wherein the C1-C4 alkyl group is derived from C1-C4 alkanol, and a mixture thereof.

17. A process according to claim 16, wherein the floor of each tray is inclined towards the zone of turbulence with an inclination that exceeds the angle of repose of the catalyst particles under a liquid.

18. A process according to any of claims 1 to 17, wherein the esterification zone comprises an autocatalytic esterification zone wherein the esterification conditions include the use of a temperature ranging from about 70 ° C to about of 250 ° C, a pressure ranging from about 1 bar to about 50 bar and where maleic anhydride is converted by reaction with C 1 -C 4 alkanol at least partially into the corresponding mono- (C 1 -C 4 alkyl) maleate.

19. A process according to any of claims 1 to 18, wherein the esterification zone includes a catalytic esterification zone wherein the esterification conditions include the use of a temperature ranging from about 65 ° C to about of 135 ° C, and of a solid esterification catalyst comprising an ion exchange resin containing pendant sulfonic acid groups.

20. A process according to any of claims 1 to 19, wherein the step of hydrogen evolution is carried out at substantially the inlet pressure in the ester hydrogenation zone.

21. A process according to any of claims 1 to 20, wherein the step of hydrogen evolution is carried out at a temperature ranging from the entry temperature in the hydrogenation zone to around of 20 ° below the entry temperature in the hydrogenation zone.

22. A process according to any of claims 1 to 20, wherein the step of hydrogenation is carried out in the vapor phase using a reduced copper catalyst promoted at a temperature ranging from about 150 ° C to around 240 ° C and a pressure that goes from around 5 bar to around 100 bar.

23. A process according to any of claims 1 to 22, wherein from the hydrogenation zone a mixture of the hydrogenation product containing, in addition to the butane-1,4-diol and C 1 -C 4 alkanol, is recovered. minor amounts of tetrahydrofuran and? -butyrolactone.

24. A process according to claim 23, wherein the hydrogenation product mixture is purified by distillation in one or more stages, including the distillation in a form of "light ends" to separate the upper parts of the components Volatile mixtures including tetrahydrofuran, C1-C4 alkanol, water, and n-butanol.

25. A process according to claim 24, wherein the bottom products from the column of "light ends" are further purified by distillation in one or more steps to give pure butane-1,4-diol.