WO2005058492A1

WO2005058492A1 - Improved catalyst for hydrogenating maleic anhydride to form y-butyrolactone and tetrahydrofuran

Info

Publication number: WO2005058492A1
Application number: PCT/EP2004/013810
Authority: WO
Inventors: Markus Rösch; Rolf Pinkos; Michael Hesse; Stephan Schlitter; Olga Schubert; Gunther Windecker; Alexander Weck
Original assignee: Basf Aktiengesellschaft
Priority date: 2003-12-09
Filing date: 2004-12-04
Publication date: 2005-06-30
Also published as: DE10357716A1

Abstract

The invention relates to a chrome-free catalyst for hydrogenating C4-dicarboxylic acid and/or derivatives thereof, preferably maleic anhydride in the gaseous phase, provided in the form of catalyst shaped bodies, which have a mixture consisting of 5 to 95 % by weight of Cu oxide and 5 to 95 % by weight of an oxide having acid centers. The invention is characterized in that the volume of the individual shaped body is less than 20 mm3. The invention also relates to a method for producing mixtures of optionally substituted dicarboxylic acids and/or derivatives thereof in the presence of a catalyst of this type, during which the ratio of products THF and GBL to one another is set only by varying the temperatures inside the reactor.

Description

Improved catalyst for the hydrogenation of maleic anhydride to γ-butyrolactone and tetrahydrofuran

description

The present invention relates to an improved chromium-free catalyst for the preparation of defined BSA-free mixtures of γ-butyloctolactone (GBL) and tetrahydrofuran (THF) by catalytic hydrogenation of C4-dicarboxylic acids and / or their derivatives such as maleic anhydride (MSA) in the gas phase. Derivatives in this application are esters and anhydrides which, like the acids, may have one or more alkyl substituents. The process uses a catalyst having a uniform structure which comprises mixtures of copper oxide and optionally at least one further oxide from groups 1 to 14 of the Periodic Table of the Elements in the active composition, and an oxide having acidic centers, preferably aluminum oxide the active composition and the oxide with acidic centers (also referred to below as acidic oxide) are pressed or extruded into max.20 mm ³ large moldings.

The production of γ-butyrolactone (GBL) and tetrahydrofuran (THF) by gas-phase hydrogenation of maleic anhydride (MSA) has been a well-known reaction for many years. To carry out this catalytic reaction, numerous catalyst systems are described in the literature, the majority of which contain chromium. Depending on the composition of the catalyst and the reaction parameters chosen, different product distributions are achieved with such catalysts.

It has been shown that in the hydrogenation of MSA to THF first succinic anhydride (BSA) and then γ-butyrolactone (GBL) are formed as intermediates, which can be converted by further hydrogenation to 1, 4-butanediol (BDO). All of the abovementioned intermediates can themselves be used as starting materials for THF production. If GBL and THF are to be prepared which have alkyl substituents, it is advisable to use of the abovementioned educts also the corresponding alkyl-substituted species.

A variety of chromium-containing catalysts for the above-described hydrogenation reaction is known, wherein the hydrogenation is limited in most cases to MSA as starting material. For example, US 3,065,243 discloses a process in which copper chromite serves as the catalyst. According to the description and examples, considerable amounts of BSA still have to be circulated in this reaction procedure. As is known, this often involves procedural problems due to the crystallization of the BSA or also succinic acid resulting therefrom with consequent blockage of pipelines. The catalysts used according to US 5,072,009 correspond to the general formula CuιZn Al _c M _d O _x , in which M is at least one element which is selected from the group consisting of the groups IIA and IIIA, VA, VIII, Ag, Au, den Groups HIEB to VIIB as well as lanthanides and actinides of the Periodic Table of the Elements; b is a number between 0.001 and 500, c is a number between 0.001 and 500 and d is a number from 0 to <200 and x corresponds to the number of oxygen atoms that are necessary according to the valence criteria. In the hydrogenation of MSA with the catalysts according to the invention, according to the examples in which exclusively chromium-containing catalysts are used, THF is produced in yields of more than 90%. The high yields are observed only in straight pass. In addition, the catalysts are problematic due to the use of the highly toxic chromium oxide in their disposal.

Chromium-containing catalysts of the type Cu / Mn / Ba / Cr and Cu / Zn / Mg / Cr for the MSA gas phase hydrogenation are known from WO 97/43234. Here, BDO is mainly obtained in addition to small amounts of GBL and THF. The hydrogenation is carried out at about 150 ° C to 300 ° C and a pressure of 5 bar to 100 bar in the gas phase.

In the patent US 5,149,836 THF yields up to 83% are described in the examples. For these high THF selectivities, a two-stage catalyst concept is necessary. The catalyst in the second stage is again based on the above-mentioned reasons disadvantageous Cu-Zn-Cr oxides.

Due to their toxicity, newer technologies are increasingly turning away from the use of catalysts containing chromium. For the gas-phase hydrogenation of MSA to GBL, WO 97/24346 (Cu-Al-oxide), WO 99/35139 (Cu-Zn-oxide), WO 95/22539 (Cu-Zn-Zr) and US 5,122,495 (US Pat. Cu-Zn-Al oxide) described chromium oxide-free catalysts. Although these catalysts provide high yields of GBL (up to 98%), THF is not formed. Although the yield of THF, which is produced as a by-product of GBL hydrogenation, can be promoted by an increase in the reaction temperature or by longer residence times in the reactor at the expense of GBL formation, however, these extreme reaction conditions also promote the formation of undesirable by-products such. As butanol, butane, ethanol or ethane. It is therefore expected that these catalysts are unsuitable for the formation of THF in high yield.

The use of chromium oxide-free catalysts based on mixed Cu-Al oxides is disclosed in JP 2-233631. The hydrogenation of MSA is to be carried out in such a way that almost exclusively THF and BDO are produced as the main products along with at most small amounts of GBL. This is then achieved by the use of the catalysts based on mixed Cu-Al oxides and by observing certain reaction conditions. According to JP 2-233 631, the MSA hydrogenation is carried out so that MSA in a solvent such as GBL, THF or Dissolved dioxane and this strongly diluted MSA solution is passed over the catalyst after their evaporation. Due to the high proportions of solvent, the process is disadvantageous because the separation of the high solvent content of the desired products is uneconomical. Example 6 describes the evaporation of MSA melt without solvent. Here, only a yield of THF of 76.3 mol% is obtained. As a by-product, 1, 4-butanediol (BDO) is formed in 20.4% yield. According to the examples, the hydrogenation is carried out at temperatures between about 210 to 230 ° C and GHSV values of about 3200 to 9600. The hydrogen / MSA ratios are comparatively unfavorably high for industrial processes, in the examples from 200 to 800.

The hydrogenation of MSA under conditions similar to those in JP 2233631 but using a different catalyst are disclosed in JP 2639463. The use of the catalyst should allow the production of BDO and THF by hydrogenation of MSA. Used here is a copper oxide / zinc oxide / alumina catalyst whose composition is not set out quantitatively in the description. The catalysts used in the examples have a composition of 20 wt .-% CuO, 43.6 wt .-% ZnO and 18.1 wt .-% Al ₂ O ₃ , 32.6 wt .-% CuO 38.1 wt % ZnO and 9.5% by weight Al ₂ O ₃ , 24.2% by weight CuO, 36.4% by weight ZnO and 17.2% by weight Al ₂ O ₃ , 26.4 Wt .-% CuO, 52.9 wt .-%, ZnO, 7.6 wt .-% Al ₂ O ₃ and 1, 4 wt .-% CaO and 22.9 wt .-% CuO, 44.8 wt % ZnO and 16.3% by weight Al ₂ O ₃ . It is generally worked in a solvent such as GBL or dioxane, with a maximum THF selectivity of 94% is achieved. When the reaction is carried out without a solvent, the THF selectivity reaches a maximum value of 83%.

EP 0404408 claims a catalyst for MSA hydrogenation to THF and GBL having a substantially inert, at least partially porous, outer surface-carrying carrier and a catalytically active, coated on the outer surface of the carrier, strong on the outer surface the oxide of the carrier comprises, wherein the catalytically active oxide material comprises a mixed oxide catalyst of the formula Cu ^ n _b Al _c M _d O _x , wherein M is at least one element selected from the group consisting of Groups IIA to VA, Group VIII, Ag, Au, groups IIIB to VIIB, the rare earth metals and the transuranic, b is a number between 0.001 and 500, c is a number between 0.001 and 500 and d is a number from 0 to 200 and x is a number of oxygen atoms necessary according to valence criteria becomes. This shell catalyst favors the formation of GBL as the preferred product arises. Another disadvantage is the large amount of BSA formed.

DE 10061555 also discloses a copper oxide and alumina-containing coated catalyst for gas-phase dehydrogenation of MSA. From DE 10061556 a method for the selective production of THF using a Cu-Al catalyst is known. A disadvantage of the method is that there are only limited possibilities of influencing the product composition. If the temperature is lowered, the content of GBL in the discharge increases, but at the same time the content of BSA in the discharge increases.

The object of the present invention is to provide a catalyst and a process by which defined mixtures of GBL and THF can be prepared by hydrogenating MSA. This process should be able to be operated continuously and with chromium-free catalysts and as large as possible

Quantities of the value products GBL and THF supply, in order to achieve a maximum economy. Furthermore, the catalyst should be able to be operated with MSA which has not been pre-cleaned in a complicated manner, for example by distillation, and nevertheless has high stability, ie does not require frequent regeneration. The method is intended to allow the operator the variability without a catalyst change, i. only by changing the operating parameters to obtain a modified product mixture in order to adapt flexibly and economically to the market requirements can. All accessible product mixtures should be as free of BSA as possible, on the one hand to prevent procedural problems (laying in pipelines), on the other hand, but also to minimize yield losses.

This object is achieved by the use of a chromium-free catalyst for the hydrogenation of C-dicarboxylic acids and / or their derivatives, preferably maleic anhydride in the gas phase, in the form of shaped catalyst bodies, the mixture of 5 to 95 wt .-% Cu oxide and 5 to 95 wt .-% of an oxide having acidic centers, characterized in that the volume of the individual molding is less than 20 mm ³ , preferably <10 mm ³ , more preferably <6 mm ³ .

The catalyst according to the invention has a uniform structure structure. Its components are intimately mixed with each other, whereby the catalyst has no larger, differently structured components, as do, for example, coated catalysts.

By using the catalyst in the form of the inventive small moldings in a process for the hydrogenation of C ₄ dicarboxylic acids and / or their derivatives, preferably maleic anhydride in the gas phase, the quantitative ratio of optionally alkyl-substituted value products GBL and THF while retaining all other reaction parameters only can be adjusted by varying the temperature in the reactor in the ratio 0: 1 to 1: 9 and thus a product distribution according to market requirements can be obtained without catalyst change. The term C ₄ -dicarboxylic acids and their derivatives in relation to the present application maleic acid and succinic acid, optionally a or more CC ₆ alkyl substituents and the anhydrides of these optionally alkyl-substituted acids understood. An example of such an acid is citraconic acid. Preferably, the respective anhydrides of a given acid are used. In particular, the educt used is MSA.

The catalyst according to the invention is characterized in that it is able to make variable mixtures of GBL and THF accessible by gas phase hydrogenation of MSA in terms of their product ratios, which are largely free of BSA. It has a long service life and a high load capacity, so that reactivations can be omitted and reactor volume can be saved. All these points make a method practiced with this catalyst particularly economical.

The catalyst according to the invention comprises copper oxide and an acidic oxidic material as the main constituent and may optionally contain one or more further metals or their compounds, preferably oxides, from groups 1 to 14 (IA to VIIIA and IB to IVB of the ancient IUPAC nomenclature) of Periodic Table of Elements. If such a further oxide is used, preference is given to using TiO ₂ , ZrO ₂ , SiO ₂ , CaO, Na ₂ O, Mn ₂ O ₃ , BaO, and / or MgO.

The proportion of copper oxide in the total mass of the catalyst is 5 to

95 wt .-%, preferably at <70 wt .-%., Particularly preferably 10 to 65 wt .-%.

The catalyst of the invention contains an acidic oxide which must have a suitable number of acidic sites. In this case, the required proportion of the oxide having acidic centers depends on the amount of acidic centers contained therein. Suitable acid oxides having a sufficient number of acidic sites are titanium dioxide, zirconia, silica and alumina, the use of which is preferred and mixed oxides of these oxides. Particularly preferred catalyst compositions comprise <70% by weight of copper oxide, in particular 10 to 65% by weight of CuO, and> 20% by weight, preferably> 30% by weight, in particular 35 to 90% by weight, of acidic oxide, on. Low copper oxide levels are also preferred because of the cost advantage achieved thereby. Due to the acidic oxides, high yields can be achieved. The metals from groups 1 to 14 of the Periodic Table of the Elements may be present in the catalyst in an amount of up to 30% by weight.

The catalysts used may also contain one or more auxiliaries in an amount of 0 to 10 wt .-%. By auxiliaries are meant organic and inorganic substances which contribute to improved processing during catalyst preparation and / or to an increase in the mechanical strength of the catalyst bodies. Such aids are known in the art; Examples include graphite, stearic acid, silica gel, cellulose compounds, starch, polyolefins, carbohydrates (sugars), waxes, alginates, as well as alumina, alumina hydro- xid, eg in the form of boehmite; Zirconia, silica and their sols, substituted and unsubstituted siloxanes and copper powder.

The catalysts can be prepared by methods known to those skilled in the art. Preference is given to processes in which the copper oxide is finely distributed and intimately mixed with the other constituents of the active composition and the acidic oxide. Particularly preferably, the corresponding metal salts and / or hydroxides are precipitated from aqueous solution, washed, dried and calcined. Suitable metal salts are, for example, nitrates, sulfates, carbonates, chlorides, acetates or oxalates. Subsequently, this starting material will be processed by known methods to the moldings, for example extrusion, tableting or by agglomeration, optionally with the addition of auxiliaries.

It is also possible to prepare the catalysts according to the invention by applying the active material or precursor compounds of the active composition to an acidic oxide, for example by impregnation or vapor deposition. Further, catalysts of the present invention can be obtained by molding a mixture of active components and acidic oxide or their precursor compound with an oxide having acidic centers or a precursor compound thereof.

The catalysts are used as shaped bodies. Examples include strands, rib strands, other extrudate shapes, tablets, rings, balls, and chippings.

Among all moldings are particularly suitable extrudates. These are obtained by extruding the dried starting compound with an adjuvant (binder), for example boehmite or p-boehmite (AIOOH), and then calcined. The binder can be pretreated prior to extrusion. This is preferably done with acid, such as with formic acid, nitric acid or hydrochloric acid. Other adjuvants such as pore formers such as carboxymethyl cellulose, potato starch or stearic acid may additionally be added prior to extrusion.

The determination of the volume of the shaped bodies according to the invention is effected by calculation from the dimensions of the shaped bodies.

In the hydrogenation according to the invention, in which, in addition to MSA, other C dicarboxylic acids defined above or their derivatives can be used as starting material, the catalyst is used in reduced, activated form. Activation takes place with reducing gases, preferably hydrogen or hydrogen / inert gas mixtures, either before or after incorporation into the reactor in which the process according to the invention is carried out. If the catalyst has been incorporated into the reactor in oxidic form, the activation can take place both before the system is started up with the hydrogenation according to the invention and during the startup, ie in situ, be performed. The separate activation before starting the system is generally carried out with reducing gases, preferably hydrogen or hydrogen / inert gas mixtures at elevated temperatures, preferably between 100 and 300 ° C. In so-called in-situ activation, the activation takes place when the system starts up by contact with hydrogen at elevated temperature.

The BET surface area of the copper catalysts in the oxidic state is 10 to 400 m ² / g, preferably 15 to 200 m ² / g, in particular 20 to 150 m ² / g. The copper surface (N ₂ O decomposition) of the reduced catalyst in the installed state is> 0.2 m ² / g, preferably> 1 m ² / g, in particular> 2 m ² / g.

According to one variant of the invention, catalysts are used which have a defined porosity. As catalysts, these catalysts show a pore volume of ≥ 0.01 ml / g for pore diameters> 50 nm, preferably ≥ 0.025 ml / g for pore diameters> 100 nm and in particular ≥ 0.05 ml / g for pore diameters

> 200 nm. Furthermore, the ratio of macropores lies with a diameter

> 50 nm to the total pore volume for pores with a diameter> 4 nm at values> 10%, preferably> 20%, in particular> 30%. Often can be achieved by using these catalysts high THF yields and selectivities. The mentioned porosities were determined by mercury intrusion according to DIN 66133. The data in the pore diameter range from 4 nm to 300 nm were evaluated.

The catalysts used in the invention generally have a sufficient life. In the event that the activity and / or selectivity of the catalyst should nevertheless decrease in the course of its operating time, this can be regenerated by measures known to the person skilled in the art. This preferably includes a reductive treatment of the catalyst in a hydrogen stream at elevated temperature. Optionally, the reductive treatment may be preceded by an oxidative one. In this case, the catalyst bed is mixed with a molecular oxygen-containing gas mixture, for example air, at elevated temperature. Furthermore, it is possible to wash the catalyst with a suitable solvent, for example ethanol, THF or GBL, and then to dry in a gas stream.

In order to achieve the GBL and THF selectivities according to the invention with the catalyst according to the invention, it is furthermore necessary to observe certain reaction parameters. Furthermore, a continuous procedure is possible, which is preferred according to the invention.

Reactors in which the catalyst is arranged in a fixed bed are suitable for the process according to the invention. Particularly preferred are tube bundle reactors or plate reactors to the low heat released during the reaction dissipate. MSA is vaporized and passed through the reactor with a hydrogen-containing gas stream.

By the method according to the invention, the ratio of the optionally alkyl-substituted value products GBL and THF can be adjusted while maintaining all other reaction parameters only by varying the temperature in the reactor in a ratio of 0: 1 to 1: 9. This is achieved on the one hand by a sufficiently high inlet temperature of the starting materials in the reactor. This is at values of> 220 to 300 ° C, preferably 235 to 270 ° C. In order to obtain an acceptable or high desired product selectivity and yield, the reaction must be carried out so that there is a suitably high reaction temperature at the catalyst bed on which the actual reaction takes place. Over the length of the reactor can be adjusted at a given pressure, hydrogen / reactant ratio and load through the inlet temperature, a temperature profile, wherein several temperature maxima can be achieved in the reactor. These so-called hot-spot temperatures are at values of 240 to 310 ° C, preferably 240 to 280 ° C. The process is carried out so that the inlet temperature of the reaction gases is below this hot spot temperature. The first hot spot lies spatially in the first half of the reactor after the entry point of the reaction mixture, in particular in the presence of a tube bundle reactor. Preferably, the first hot-spot temperature is 5 to 15 ° C, in particular 10 to 15 ° C, above the inlet temperature. If a second hot spot occurs, it may be located at any point in the reactor and will be displaced spatially towards the reactor outlet by lowering the inlet temperature. Preferably, the second hot-spot temperature is 1 to 15 ° C, in particular 2 to 10 ° C, above the inlet temperature. If mainly THF-containing mixtures are to be obtained, this second hot-spot is preferably in the first 2/3 of the reactor; preferably GBL-containing mixtures are to be obtained, the second hotspot is in the last third, in particular at the reactor outlet.

If the hydrogenation is carried out below the minimum temperatures of the inlet or hot spot temperature, then in the course of the hydrogenation a deactivation of the catalyst can be observed by occupying with succinic acid, fumaric acid and / or BSA. If, on the other hand, hydrogenation with MSA as educt above the maximum temperatures of the inlet or hot spot temperature causes the GBL and THF yield and selectivity to fall to unsatisfactory levels. It is hereby observed the increased formation of n-butanol and n-butane, ie the products of a further hydrogenation.

The catalyst loading of the hydrogenation according to the invention is in the range from 0.02 to 2.0 kg of starting material / catalyst / hour. In a possible, but not preferred recycling of incomplete hydrogenation intermediate product, in the case of using MSA as starting material GBL, the catalyst loading is the sum of freshly fed educt and recycled intermediate. If the catalytic converter Tort burden increased beyond the range mentioned, is generally observed an increase in the proportion of the intermediate product in the hydrogenation. Preferably, the catalyst loading is in the range of 0.05 to 1 kg of starting material I catalyst • hour. In the case of recycling, the term "starting material" also means initially formed hydrogenation product, which is then hydrogenated further to product after recycling, for example GBL in the case of the use of MSA in the hydrogenation reaction.

The hydrogen / reactant molar ratio is likewise a parameter which has an important influence on the product distribution and also the economic viability of the process according to the invention. From an economic point of view, a low hydrogen / reactant ratio is desirable. The lower limit is a value of 5, but generally higher hydrogen / reactant mole ratios of 20 to 650 are used. The use of the catalysts according to the invention described above and the observance of the temperature values described above permit the use of favorable, low hydrogen / starting material ratios, which are preferably from 20 to 200, preferably from 40 to 150. The cheapest range is between 50 and 100.

In order to adjust the hydrogen / educt molar ratios used according to the invention, a part, advantageously the main amount, of the hydrogen is circulated. For this purpose, one generally uses the cycle gas compressor known to those skilled in the art. The amount of hydrogen chemically consumed by the hydrogenation is supplemented. In a preferred embodiment, a portion of the recycle gas is discharged to remove inert compounds, such as n-butane. The recirculated hydrogen may also be used, optionally after preheating, to evaporate the educt stream.

The pressure at which the hydrogenation according to the invention is carried out is at values of 1 to 100 bar, preferably 2 to 60 bar, in particular 5 to 15 bar.

Together with the hydrogen cycle gas, all products are circulated, which do not or not completely condense on cooling of the gas stream leaving the hydrogenation reactor. These are mainly THF, water and by-products such as methane and butane. The cooling temperature is 0 to 60 ° C, preferably 20 to 45 ° C. The THF content of the circulating gas is 0.1 to 5% by volume, in particular 1 to 3% by volume.

From the literature it is known that THF and GBL are hydrogenated with hydrogen in the presence of copper catalysts to n-butanol. The process according to the invention is characterized in that GBL and THF yields of more than 90%, in some cases even more than 95%, are achieved in spite of the high proportions of THF in the recycle gas, which are generally hydrogenated further to n-butanol become. As reactor types come all, for heterogeneously catalyzed reactions with a gaseous Edukt- and product stream suitable apparatus into consideration. Preference is given to tubular reactors, shaft reactors, plate reactors or reactors with internal heat removal, for example tube-bundle reactors. Particular preference is given to using tube bundle reactors. Several reactors can be used in parallel or in series. In principle, an intermediate feed can take place between the catalyst beds. Also possible is an intermediate cooling between or in the catalyst beds. When using fixed bed reactors, a dilution of the catalyst by inert material is possible.

The gas stream leaving the reactor is cooled to 10 to 60 ° C. The reaction products are condensed out and passed into a separator. The non-condensed gas stream is withdrawn from the separator and fed to the cycle gas compressor. A small amount of circulating gas is discharged. The condensed reaction products are continuously removed from the system and fed to the workup. As by-products are found in the condensed liquid phase mainly n-butanol in addition to small amounts of propanol.

From the hydrogenation then the azeotrope of water and optionally alkyl-substituted THF, GBL and optionally by-product, separated by fractional distillation. The water-containing THF is dehydrated in a manner known per se and worked up by distillation to specification-compliant THF.

The process according to the invention is characterized in that starting materials of different purity to be hydrogenated can be used in the hydrogenation reaction. Of course, a starting material of high purity, in particular MSA, can be used in the hydrogenation reaction. However, the catalyst used according to the invention and the other reaction conditions selected according to the invention also allow the use of educts, in particular MSA, which is contaminated with the customary compounds obtained in the oxidation of benzene, butenes or n-butane and possibly other components. Thus, in another embodiment, the hydrogenation process according to the invention may comprise an upstream stage which comprises preparing the starting material to be hydrogenated by partial oxidation of a suitable hydrocarbon and separating the starting material to be hydrogenated from the product stream thus obtained.

In particular, this educt to be hydrogenated is MSA. MSA is preferably used, which originates from the partial oxidation of hydrocarbons. Suitable hydrocarbon streams are benzene, C-olefins (eg, n-butenes, C ₄ raffinate streams) or n-butane. Particular preference is given to the use of n-butane, since it represents a low-cost, economical starting material. Methods for the partial oxidation of n-butane are for example, in Ullmann's Encyclopedia of Industrial Chemistry, 6 ^{, h} Edition, Electronic Release, Maleic and Furfural Acids - Maleic Anhydrides.

The resulting reaction effluent is then taken up in a suitable organic solvent or mixture which has a boiling point higher than MSA at atmospheric pressure by at least 30 ° C.

This solvent (absorbent) is brought to a temperature in the range between 20 and 160 ° C, preferably between 30 and 80 ° C. The maleic anhydride-containing gas stream from the partial oxidation can be brought into contact in many ways with the solvent: (i) introducing the gas stream into the solvent (eg via gas inlet nozzles or gassing rings), (ii) spraying the solvent into the gas stream and (iii ) Countercurrent contact between the upwardly flowing gas stream and the downwardly flowing solvent in a bottom or packed column. In all three variants, the apparatus known to those skilled in the gas absorption can be used. When choosing the solvent to be used, care must be taken that this does not react with the educt, for example the preferably used MSA. Suitable solvents are: tricresyl phosphate, dibutyl maleate, high molecular weight waxes, aromatic hydrocarbons having a molecular weight between 150 and 400 and a boiling point above 140 ° C, such as dibenzylbenzene; Dialkyl phthalates with -CC ₈ alkyl groups, for example dimethyl phthalate, diethyl phthalate, dibutyl phthalate, di-n-propyl and di-iso-propyl phthalate; Di-dC-alkyl esters of other aromatic and aliphatic dicarboxylic acids, for example dimethyl-2,3-naphthalene-dicarboxylic acid, dimethyl-1,4-cyclohexane-dicarboxylic acid, methyl esters of long-chain fatty acids having, for example, 14 to 30 carbon atoms, high-boiling ethers, for example dimethyl ether of Polyethylene glycol, for example tetraethylene glycol dimethyl ether.

The use of phthalates is preferred.

The resulting after treatment with the absorbent solution generally has an MSA content of about 5 to 400 grams per liter.

The waste gas stream remaining after treatment with the absorbent mainly contains the by-products of the preceding partial oxidation, such as water, carbon monoxide, carbon dioxide, unreacted butanes, acetic and acrylic acids. The exhaust stream is virtually free of MSA.

Subsequently, the dissolved MSA is expelled from the absorbent. This is done with hydrogen at or at most 10% above the pressure of the subsequent hydrogenation or alternatively in vacuo with subsequent condensation of remaining MSA. In the stripping column, a temperature profile is observed from the boiling points of MSA at the top and the almost MSA-free absorbent at the bottom of the column at the respective column pressure and the dilution with carrier gas (in the first case with hydrogen) results. In the case of direct stripping with hydrogen, work is carried out at a head temperature of 130 ° C. and a pressure of 5 bar.

In order to prevent losses of solvent, rectification installations may be located above the feed of the crude MSA stream. The almost MSA-free absorbent withdrawn from the bottom is returned to the absorption zone. In the case of direct stripping with hydrogen, a nearly saturated gas stream of MSA in hydrogen at a temperature of 180 ° C. and a pressure of 5 bar is withdrawn from the top of the column. The H ₂ / MSA ratio is about 20 to 650. In the other case, the condensed MSA is pumped into an evaporator and vaporized there in the recycle gas stream.

The MSA hydrogen stream still contains by-products which are formed in the partial oxidation of n-butane, butenes or benzene with oxygen-containing gases, as well as non-separated absorbent. These are, in particular, acetic acid and acrylic acid as by-products, water, maleic acid and the dialkyl phthalates which are preferably used as absorbents. The MSA contains acetic acid in amounts of 0.01 to 1 wt .-%, preferably 0, 1 to 0.8 wt .-% and acrylic acid in amounts of 0.01 to 1 wt .-%, preferably 0.1 to 0 , 8 wt .-%, based on MSA. In the hydrogenation stage, all or part of acetic acid and acrylic acid are hydrogenated to give ethanol or propanol. The maleic acid content is 0.01 to 1 wt .-%, in particular 0.05 to 0.3 wt .-%, based on MSA.

If dialkyl phthalates are used as absorbents, their content in the MSA depends strongly on the correct operation of the stripping column, in particular on the reinforcing part. Phthalate contents of up to 1, 0 wt .-%, in particular up to 0.5 wt .-% should not be exceeded with appropriate operation, otherwise the consumption of absorbent is too high.

The thus obtained hydrogen / maleic anhydride stream is now fed to the hydrogenation zone and hydrogenated as described above. The catalyst activity and -Standzeit is virtually unchanged compared to the use of strong, for example by distillation, prepurified MSA. The process of the present invention allows GBL and THF yields to be at values of about 94%, preferably about 98%. It also achieves a high product selectivity. Examples

Example 1 :

Production of a Spray Powder According to the Invention

1.51 of water are placed in a heated pot equipped with stirrer and heated to 80.degree. In this Fällgefäß a metal salt solution consisting of 877g Cu (NO ₃ ) ₂ * 2.5H ₂ O and 1472g Al (NO ₃ ) ₃ ^* 9H ₂ O in 2000 ml of water and at the same time a 20 wt .-% soda solution in the course of one hour added with stirring. The sodomization is chosen so that a pH value of 6 is established in the precipitation vessel. After the complete addition of the metal salt solution, soda solution is further metered in until a pH of 8 is reached in the precipitating vessel and stirring is continued for a further 15 min at this pH. The total consumption of soda solution is 5.5 kg. The resulting suspension is filtered off and washed with water until the effluent wash water no longer contains nitrate (<25 ppm). The filter cake is slurried with water. The resulting mash is sprayed at 120 to 135 ° C.

Example 2:

Production of small moldings

a) Tablets (1, 5 × 2 mm) Shaped Volume: 3.5 mm ³ The spray powder from Example 1 is then calcined at 600 ° C. The catalyst thus prepared contains 61% by weight of CuO and 39% by weight of Al ₂ O ₃ . This is intensively mixed with 3% graphite and pressed into tablets of size 1, 5 x 2 mm. These tablets contained 59% by weight of CuO, 38% by weight of Al ₂ O ₃ and 3% by weight of carbon.

b) extrudate molding material volume: 3.5-5.3 mm ³ boehmite is etched with formic acid, mixed with spray powder from Example 1 and extruded after addition of water in an extruder to strands of length 3 mm with a diameter of 1, 5 mm. The strands are then dried and calcined at 600 ° C. The extrudates contain 50% by weight of CuO and 50% by weight of Al ₂ O ₃ . Example 3:

Carrying out the reaction with the catalyst according to the invention from Example 2a)

a) Catalyst activation

Before the start of the reaction, the catalyst is subjected to a hydrogen treatment in the pressureless hydrogenation apparatus. For this purpose, the reactor is heated to 200 ° C and the catalyst activated the time indicated in Table 1 with the particular mixture of hydrogen and nitrogen at atmospheric pressure indicated.

Table 1

The mixture is then bathed for 15 h at 250 ° C with 200 Nl / h of hydrogen.

b) Pilot plant The pressure apparatus used for the hydrogenation consists of an evaporator, a reactor, a cooler, a hydrogen supply, an exhaust gas line and a compressor. The pressure in the apparatus is kept constant.

The melted MSA is pumped from above through a plug-in tube onto the preheated (195 ° C.) evaporator and evaporated. From the bottom, a preheated mixture of fresh hydrogen and recycle gas enters the evaporator. Example hydrogen and MSA get into the tempered reactor filled with catalyst (diameter 20 mm). After hydrogenation, the resulting mixture of GBL and THF leaves the reactor with water, other reaction products and hydrogen and is precipitated in the condenser. A portion of the recycle gas is discharged before the remainder, mixed with fresh hydrogen, re-enters the evaporator.

In this pilot plant, the reactor is filled with 1800 ml of the catalyst from Example 2a (total mass 1575 g) and activated according to 3a).

The condensed liquid reaction product, the exhaust gas and the recycle gas are quantitatively analyzed by gas chromatography. The hydrogen: MSA ratio is 80: 1 The operating parameters and test results are shown in Table 2. Table 2

EXAMPLE 4 Performance of the Reaction with the Inventive Catalyst from Example 2b) a) Catalyst Activation Before the start of the reaction, the catalyst is subjected to a hydrogen treatment in the pressureless hydrogenation apparatus. For this purpose, the reactor is heated to 200 ° C and the catalyst activated the time indicated in Table 3 with the particular mixture of hydrogen and nitrogen at atmospheric pressure indicated.

Table 3

The mixture is then bathed for 15 h at 250 ° C with 200 Nl / h of hydrogen. b) Pilot plant The pressure apparatus used for the hydrogenation consists of an evaporator, a reactor, a cooler, a hydrogen supply, an exhaust gas line and a compressor. The pressure in the apparatus is kept constant.

The melted MSA is pumped from above through a plug-in tube onto the preheated (195 ° C) evaporator and evaporated. From the bottom, a preheated mixture of fresh hydrogen and recycle gas enters the evaporator. Example hydrogen and MSA get into the tempered reactor filled with catalyst (Diameter 34 mm). After hydrogenation, the resulting mixture of GBL and THF leaves the reactor with water, other reaction products and hydrogen and is precipitated in the condenser. A portion of the recycle gas is discharged before the remainder, mixed with fresh hydrogen, re-enters the evaporator.

In this pilot plant, the reactor is filled with 1200 ml of the catalyst from Example 2b and activated according to 4a). The hydrogen to MSA ratio is 80: 1. The condensed liquid reaction product, the exhaust gas and the recycle gas are quantitatively analyzed by gas chromatography.

The operating parameters and test results are shown in Table 4. Table 4

Comparative Example 1: Preparation of the Comparison Catalyst

The spray powder from Example 1 is then calcined at 600 ° C. The catalyst thus prepared contains 61% by weight of CuO and 39% by weight of Al ₂ O ₃ . This is intensively mixed with 3% graphite and pressed into tablets of size 3 × 3 mm. These tablets contained 59% by weight of CuO, 38% by weight of Al ₂ O ₃ and 3% by weight of carbon. This catalyst has a volume of 21, 2 mm ³ .

Comparative Example 2 Example 3 was repeated with the difference that 150 ml of the catalyst from Comparative Example 1 were used. The results of the catalytic test are shown in Table 6.

Table 5

Claims

1. Chromium-free catalyst for the hydrogenation of C ₄ dicarboxylic acids and / or their derivatives, preferably maleic anhydride in the gas phase, in the form of shaped catalyst bodies which are a mixture of 5 to 95% by weight of Cu oxide and 5 to 95% by weight. % of an oxide with acidic centers, characterized in that the volume of the individual shaped body is less than 20 mm ³ .

2. Catalyst according to claim 1, characterized in that the volume of the individual shaped body is <10 mm ³ , preferably <6 mm ³ .

3. Catalyst according to claim 1 to 2, characterized in that the CuO <80 wt .-%, preferably <70 wt .-%, in particular 10 to 65 wt .-% CuO and> 20 wt .-%, preferably> 30 % By weight, in particular 35 to 90% by weight, of an oxide with acidic centers are contained.

4. Catalyst according to claim 1 to 3, characterized in that the oxide with acid centers is Al ₂ O ₃ .

5. Catalyst according to one of claims 1 to 4, characterized in that one or more further metals or a compound thereof, preferably an oxide, from the group consisting of the elements of groups 1 to 14 of the periodic table of the elements are present in the catalyst , preferably a substance from the group consisting of ZrO ₂ , TiO ₂ , CaO, Na ₂ O, Mn ₂ O ₃ , BaO, SiO ₂ and MgO.

6. Catalyst according to one of claims 1 to 5, characterized in that the shaped body is in the form of a strand.

7. Catalyst according to one of claims 1 to 6, characterized in that the catalyst is activated by reduction before or after installation in the reactor and before use in the hydrogenation reaction, preferably by treatment with hydrogen or a hydrogen / inert gas mixture.

8. Catalyst according to one of claims 1 to 7, characterized in that the BET surface area of the copper catalysts in the oxidic state is 10 to 400 m ² / g, preferably 15 to 200 m ² / g, in particular 20 to 150 m ² / g. The copper surface (N ₂ O decomposition) of the reduced catalyst in the installed state is> 0.2 m ² / g, preferably> 1 m ² / g, in particular> 2 m ² / g.

9. Catalyst according to one of claims 1 to 8, characterized in that the shaped catalyst body has a pore volume of ≥ 0.01 ml / g for pore diameter knife> 50 nm, preferably ≥ 0.025 ml / g for pore diameter> 100 nm, in particular ≥ 0.05 ml / g for pore diameter> 200 nm.

10. Catalyst according to one of claims 1 to 9, characterized in that in the shaped catalyst body the ratio of macropores with a diameter> 50 nm to the total pore volume for pores with a diameter> 4 nm at values> 10%, preferably> 20% , in particular> 30%.

11. The method with a catalyst according to any one of claims 1 to 10, characterized in that maleic anhydride is used as a starting material in the reaction.

12. A process for the preparation of mixtures of optionally alkyl-substituted GBL and THF in the ratio 0: 1 to 9: 1 by hydrogenation of C ₄ -dicarboxylic acids and / or their derivatives in the presence of a catalyst according to one of claims 1 to 10, in which the ratio of the products THF and GBL to one another can only be set by varying the temperatures in the reactor.

13. The method according to any one of claims 1 to 12, characterized in that the reactor inlet temperature at values of 200 ° C to 300 ° C, preferably 235 to 270 ° C and about 5 to 15 ° C, preferably about 10 to 15 ° C, is below the first hot spot temperature.

14. The method according to claim 1 to 13, characterized in that the method is carried out continuously.

15. The method according to any one of claims 1 to 14, characterized in that the catalyst loading in the range of 0.02 to 2 kg of starting material / I catalyst • hour, in particular 0.05 to 1 kg of starting material / l catalyst hour.

16. The method according to any one of claims 1 to 15, characterized in that the molar ratio of hydrogen / starting material is from 20 to 650, preferably 20 to 200, in particular 40 to 150, most preferably 50 to 100.

17. The method according to any one of claims 1 to 16, characterized in that the first hydrogenation stage is carried out at pressures of 1 to 100 bar, preferably 2 to 60 bar, in particular 5 to 15 bar.

18. The method according to any one of claims 1 to 17, characterized in that the maleic anhydride is expelled from the absorbent in vacuo or at pressures which correspond to the pressure of the hydrogenation or are at most 10% above this pressure.