DE10225928A1 - Two-stage process for the production of butanediol in two reactors - Google Patents

Abstract

The present invention relates to a process for the preparation of optionally alkyl-substituted 1,4-butanediol by two-stage catalytic hydrogenation in the gas phase of C¶4¶-dicarboxylic acids and / or their derivatives with the following steps: DOLLAR A a) introduction of a gas stream of a C¶ 4¶-dicarboxylic acid or a derivative thereof at 200 to 300 ° C. and 2 to 60 bar in a first reactor and catalytic gas phase hydrogenation to give a product which may contain alkyl-substituted gamma-butyrolactone; DOLLAR A b) introducing the product stream thus obtained into a second reactor at a temperature of 150 ° C. to 240 ° C. and a pressure of 15 to 100 bar and catalytic gas phase hydrogenation to give alkyl-substituted 1,4-butanediol, if appropriate; DOLLAR A c) separation of the desired product from intermediates, by-products and, if appropriate, unreacted starting material; DOLLAR A d) optionally recycling unreacted intermediates to one or both hydrogenation stages, DOLLAR A, a catalyst being used in each of the two hydrogenation stages, the catalyst comprising 95% by weight, preferably 5 to 95% by weight, in particular 10 to 80% by weight % CuO and> = 5% by weight, preferably 5 to 95% by weight, in particular 20 to 90% by weight, of an oxidic support, the pressure in the second reactor is higher than in the first reactor, and that in the first reactor withdrawn product mixture is introduced into the second reactor without further purification.

Description

The present invention relates to a process for the preparation of optionally alkyl substituted butanediol by catalytic hydrogenation in the gas phase of Substances that are selected from the group consisting of derivatives of maleic acid and succinic acid as well as these acids themselves present invention understood anhydrides which, like the acids, one or can have several alkyl substituents.

The known hydrogenation of MSA leads via the intermediate stage Succinic anhydride (BSA) first to γ-butyrolactone (GBL). Then the further hydrogenation to tetrahydrofuran (THF), n-butanol (BuOH) and / or n-butane. GBL is in equilibrium with butanediol (BDO) large part can be moved to the side of the butanediol. Butanediol can, however just as easy as GBL through so-called overhydrogenation to butanol and butane react; Cyclization of the butanediol leads to THF. These products cannot be converted back into BDO or GBL. If BDO is desired as a product, the Education especially of the THF can be avoided.

The gas phase hydrogenation of purified maleic anhydride (MSA) to butyrolactone (GBL) and the conversion of purified GBL to BDO have been two for many years known reactions. To carry out these two catalytic reactions are in numerous catalyst systems are described in the literature. Depending on the composition of the Catalysts and the chosen reaction parameters are used with such catalysts different product distributions achieved. Process for the direct production of Butanediol based on MSA are also already known.

If GBL and BDO are to be produced that have alkyl substituents, it offers itself, the corresponding starting materials mentioned above to use alkyl substituted species.

Those in the hydrogenation of MA to one of the products mentioned above The catalysts used, especially in older processes, often contain chromium. This is reflected in the patent literature, which contains a large number of Patents and patent applications will find the chromium-containing catalysts for the Disclose hydrogenation reaction, the hydrogenation being in most cases as MSA Educt is limited.

The following documents describe the use of chromium-containing catalysts in the hydrogenation of MSA.

EP-A 0 322 140 describes a continuous process for the production of Tetrahydrofuran (THF) and for the simultaneous production of THF and GBL by Gas phase hydrogenation of MSA and BSA disclosed. The claimed catalyst contains Copper, zinc and aluminum and a further element of group IIA, IIIA, VA, VIII, IIIB to VIIB, the lanthanum and actinium series, or Ag and Au. At 40 bar can with these catalyst systems based on pure MSA THF yields of 90-95% can be obtained at a pressure of about 20 bar mixtures of GBL and THF receive.

A similar catalyst is used in US 4,965,378 and US 5,072,009, but which can additionally contain Si, Ge, Sn and Pb. By applying such Catalysts generate large amounts of THF (95% to 31.4%) that do not result in butyrolactone or butanediol can be implemented.

EP-A 0 404 408 discloses a catalyst for MSA hydrogenation, the catalytically active material essentially corresponds to the material of US 5,072,009. It is used fixed on a support as a coated catalyst. In the Embodiments are used exclusively chromium-containing catalysts. High GBL yields can be achieved at a pressure of 2 bar, a higher one Related to pressure, the THF yield increases, the GBL yield decreases.

No. 5,149,836 describes a multi-stage gas phase process for the production of GBL and THF with variable product selectivities, wherein in a first stage one Mixture of pure MSA and hydrogen is passed over a catalyst, which Contains copper, zinc and aluminum. This raw reaction discharge then becomes Production of the THF over a chromium-containing catalyst.

WO 99/38856 describes a catalyst consisting only of copper and chromium disclosed, with the straight-through starting from pure MSA GBL selectivities from 92 to 96 mol% can be obtained.

EP-A 638 565 discloses a catalyst containing copper, chromium and silicon, the composition according to one example corresponding to approximately 78% CuO, 20% Cr ₂ O ₃ and 2% SiO ₂ . With pure MSA and nitrogen-hydrogen mixtures, a GBL yield of 98% could be obtained.

The following documents disclose the use of chrome-free catalysts the hydrogenation of MSA.

GB-A 1 168 220 discloses a gas phase process for the production of GBL, in the MSA or BSA hydrogenated to GBL on a binary copper-zinc catalyst become. In all of the exemplary embodiments, work is carried out at atmospheric pressure, GBL yields of 94 mol% could be achieved on the basis of pure MSA.

In DE-OS 24 04 493 is also a process for the preparation of GBL catalytic hydrogenation of mixtures of MSA, BSA, maleic acid (MS), Succinic acid (BS) and water on a metallic catalyst are revealed, alongside Copper-chromite catalysts also come copper-zinc and copper-zinc-aluminum Precipitation catalysts for use.

WO 91/16132 discloses the hydrogenation of MA to GBL, using a catalyst containing CuO, ZnO and Al ₂ O ₃ , which is reduced at 150 ° C. to 350 ° C. and activated at 400 ° C. Activation is intended to increase the service life of the catalyst system.

A CuO-ZnO-containing catalyst is disclosed in US 6,297,389. With him can after activation, pure MSA can be converted to GBL, in a straight pass can be obtained from pure MSA GBL yields of 92 to 96%.

WO 95/22539 discloses a process for the production of GBL by catalytic Hydrogenation of MSA and / or BSA on a catalyst made of copper, zinc and Zirconium exists. Based on pure MSA, GBL yields of up to 99% reached.

WO 99/35136 describes a two-stage process for the production of GBL and THF discloses, wherein MSA hydrogenates in a first stage with a copper-containing catalyst and this reaction discharge via an acidic silicon-aluminum-containing catalyst is directed.

In WO 97/24346 a copper oxide-aluminum oxide catalyst is described with which succeeds in hydrogenating MSA in yields of 92 mol% to GBL.

The conversion from GBL to BDO is also a well-known reaction. The documents mentioned below disclose this reaction using chromium-containing Catalysts.

DE 12 77 233 describes a process for the preparation of mixtures of different types Alcohols disclosed by hydrogenating lactones with hydrogen. As catalysts are copper chromite mixed with barium on an inactive aluminum oxide support used.

GB-A 1 230 276 describes a process for producing BDO from GBL on a Copper oxide-chromium oxide catalyst at a temperature between 180 ° C and 230 ° C described.

According to DE-A 22 31 986, copper-chromite catalysts containing potassium, Sodium, rubidium, aluminum, titanium, iron, cobalt or nickel are doped, the Service life of the catalysts can be extended.

According to DE-A 25 01 499, a mixture of dioxane, GBL, Water and carboxylic acids used. The reaction described takes place at high pressure (170 bar) in the liquid phase, preferably using the solvent dioxane instead and copper-chromium oxide catalysts are also used.

Copper chromite catalysts are doped with Pd according to J01121-228-A in order to to achieve higher sales.

Further copper chromite catalysts are described by Dasunin, Maeva, Z. Org. Chim. 1 (1965), No. 6, 5996-1000; YES 5366/69; YES 7240770; J4 9024-906; J49087-610; being in the exemplary embodiments of pure GBL in the liquid phase is converted to BDO.

The gas phase hydrogenation of pure GBL to butanediol is described in US 4,652,685 on copper Chromite catalysts described. Here, at a pressure of 41 bar at one Sales of 60-68% and a BDO selectivity of 92-97% can be achieved.

In US 5,406,004 and US 5,395,990 processes for the preparation of mixtures from alcohols and diols by hydrogenation of pure GBL on copper-containing Catalysts disclosed. Here, a hydrogenation zone filled with copper catalyst is added Hydrogen feed and hydrogen at a temperature of 150 to 350 ° C and pressures between 10.3 bar and 138 bar and a product consisting of alcohols and Diols isolated. In the embodiments, a number of copper, zinc and Chromium-containing catalysts described.

Finally, the documents listed below disclose the use of chrome-free Copper catalysts for the hydrogenation of GBL to BDO.

A catalyst consisting of CuO and ZnO is described in WO 82/03854. Herewith can be in the gas phase at a pressure of 28.5 bar and a temperature of 217 ° C BDO selectivity of 98.4% can be achieved. However, the turnover in pure GBL is unsatisfactorily low.

Copper impregnation catalysts doped with palladium and potassium are in the documents US 4,797,382; US 4,885,411 and EP-A 0 318 129. You are implementing suitable from GBL to butanediol.

The use of a mixture of GBL and water as a feed stream in connection with a copper oxide-zinc oxide catalyst is described in US 5,030,773 A. In here discloses that the activity of such catalysts increases if the pure GBL stream 1 to 6% of water are added and this mixture is hydrogenated in the gas phase. If pure GBL is used for this reaction, extra water must be mixed which must then be separated again. If GBL is used would result from the hydrogenation of MSA, 17% water would be in the feed. As a result, at least 11% of water would have to be separated off before the hydrogenation to BDO become.

JP-A 0 634 567 describes a copper-iron-aluminum-containing catalyst, which is suitable for the hydrogenation of pure GBL to BDO at high pressure (250 bar).

A process for the preparation of BDO starting from maleic acid esters is in WO 99/35113 listed. Here, hydrogenation is carried out in three successive stages. Starting from maleic acid ester, a catalyst containing noble metal is used Succinic acid ester produced, which is converted to GBL and THF in a second stage becomes. GBL is separated and in a third stage at higher pressure to BDO implemented.

WO 99/35114 describes a process for the preparation of BDO Liquid phase hydrogenation of GBL, succinic acid ester or mixtures of the two Pressures between 60 bar and 100 bar and temperatures between 180 ° C and 250 ° C described. A copper oxide-zinc oxide catalyst is used as the catalyst.

Another gas phase variant of the hydrogenation of GBL to BDO is described in WO 99/52845 disclosed on a copper oxide-zinc oxide catalyst. In addition to the usual Reaction feed, carbon monoxide is added to the hydrogen at the same time To produce methanol.

EP-A 0 382 050 describes the hydrogenation of pure GBL on a cobalt oxide, Copper oxide, manganese oxide and molybdenum oxide-containing catalyst worked.

The direct production of BDO based on MSA is also known. The following The documents listed describe this reaction using chromium-containing Catalysts.

DE 28 45 905 describes a continuous process for the preparation of butanediol described starting from maleic anhydride. Here, MSA is solved in monovalent aliphatic alcohols with hydrogen at pressures of 250 bar and 350 bar Copper chromite catalysts implemented.

A process for the simultaneous production of BDO and THF starting from MSA Catalysts containing copper, chromium and manganese are described in EP-A 0 373 947 disclosed. There are mixtures of MSA and GBL, mixtures of MSA and 1,4-dioxane and pure MSA. Mixtures of THF and BDO are obtained in all cases. The high yields of tetrahydrofuran are disadvantageous in this process.

In documents CN-A 1 113 831-A, CN-A 1 116 615-A, CN-A 1 138 018-A and CN- A 1 047 328 discloses chromium-containing catalysts. In CN 1 137 944 A a copper, chromium, manganese, barium, and titanium-containing catalyst is used.

According to the disclosure of CN-A 1 182 639 a copper, chromium, zinc and Titanium-containing catalyst for the hydrogenation of mixtures of GBL and MSA can be used.

CN-A 1 182 732 describes a process for the preparation of BDO by Gas phase hydrogenation of MSA on copper and chromium-containing catalysts at 200 to 250 ° C and a pressure of 30 to 70 bar. Here, MSA is in one with the hydrogenation suitable solvent.

Finally, the documents listed below disclose direct hydrogenation from MSA to BDO using chrome-free catalysts.

DE-A 24 55 617 describes a three-stage process for the production of BDO described. In a first stage, solutions of MSA in GBL hydrogenated to BSA in GBL. In a second stage at high pressure (80-200 bar) and higher temperature, this solution of BSA in GBL becomes too liquid GBL hydrogenated, then water, succinic anhydride and succinic acid from the GBL separated and the pure GBL partially recycled and in a third process stage a copper-zinc oxide catalyst in the liquid phase at high pressure to give butanediol implemented.

No. 4,301,077 discloses a ruthenium-containing catalyst for the hydrogenation of MA BDO used.

DE-A 37 26 510 discloses the use of a copper, cobalt and phosphorus containing catalyst for the direct hydrogenation of MA.

A pure copper oxide-zinc oxide catalyst is used with J0 2025-434-A. Loud In the exemplary embodiments, pure MSA can be implemented at a pressure of 40 bar become. However, the yield of butanediol is only 53.5 mol% Side yield is found to be 40.2 mol% GBL.

EP-A 373 946 discloses a process in which one is treated with rhenium doped copper oxide-zinc oxide catalyst MSA of the gas phase converted directly to BDO becomes.

The simultaneous production of BDO and THF is the subject of the patent applications J0 2233-627-A (a copper-zinc-aluminum catalyst is used here) J0 2233-630-A (Here a manganese, barium and silicon containing copper-chromium Catalyst used) and J0 2233-631-A (here is a copper and aluminum containing catalyst used). By using these catalysts In addition to BDO, MSA also produces large amounts of THF in addition to BDO.

A catalyst containing copper, manganese and potassium is in J0-A 2233-632 described.

EP-A 431 923 describes a two-stage process for the preparation of BDO and THF described, in a first stage GBL by liquid phase hydrogenation of MSA arises and this in a second stage by gas phase reaction on a copper Silicon-containing catalyst is converted to butanediol.

US 5,196,602 discloses a process for the preparation of butanediol Hydrogenation of MA or maleic acid with hydrogen in a two-step process. In In a first stage, MSA is hydrogenated to BSA and / or GBL, which in a second Stage is converted to BDO in the presence of an Ru-containing catalyst.

The technologies on which the above-mentioned publications are based use as a starting material for the hydrogenation reactions prepurified MA, which after its production in was generally freed of impurities by distillation. MSA is manufactured by partial oxidation of certain hydrocarbons, namely benzene, Butene mixtures and n-butane, the latter being preferably used. The In addition to the desired MA, the raw product of the oxidation mainly contains by-products such as water, carbon monoxide, carbon dioxide, unreacted Starting hydrocarbon as well as acetic and acrylic acid, these by-products are independent of the hydrocarbons used in the oxidation. mostly the by-products are separated by complex processes, for example by Distillation as mentioned above. Cleaning proves to be particularly necessary because the catalysts used in the hydrogenation processes are generally sensitive to such impurities react. The deactivation of the catalysts is already at Use of cleaned MSA is a problem, because by occupying it Polymerization products of the catalyst usually occur at relatively short intervals are often around 100 hours, must be regenerated. The tendency to Deactivation is in the presence of polymerizable impurities, such as for example, acrylic acid, increased. This fact is known to the person skilled in the art and is for example also in the patent applications EP-A 322 140, WO 91/16132 and DE-OS 24 04 493 described.

So far, there is only one document in the prior art that describes the hydrogenation of only roughly pre-cleaned MSA disclosed. From WO 97/43234 it is known Maleic anhydride from gas streams containing maleic anhydride, which from the Oxidation of hydrocarbons originate with the help of at least 30 ° C higher to absorb boiling absorbents, the maleic anhydride from these Drive out absorbents with the help of hydrogen and the Hydrogen stream containing maleic anhydride in the gas phase on a hydrogenate heterogeneous catalyst. It is mainly BDO in addition to minor Get lots of GBL and THF. The hydrogenation is at about 150 ° C to 300 ° C and a pressure of 5 bar to 100 bar in the gas phase. As catalysts promoted copper catalysts are used, as described in Journal of Catalysis 150, Pages 177 to 185 (1994). These are chromium-containing Cu / Mn / Ba / Cr and Cu / Zn / Mg / Cr catalysts. Thus, according to the Disclosure of this application Chromium-containing catalysts for the hydrogenation of MSA Qualities that have the impurities set out above are used. The stake However, chromium-containing catalysts are nowadays as far as possible due to their toxicity possibly avoided.

Newer technologies are moving away from use due to their toxicity chromium-containing catalysts. Examples of chrome-free catalyst systems can be found in the publications WO 99/35139 (Cu-Zn-Oxid), WO 95/22539 (Cu-Zn-Zr) and the US 5,122,495 (Cu-Zn-Al oxide).

In the field of hydrogenation of MSA to secondary products, especially GBL, THF and / or BDO, so there is an almost unmanageable state of the art, whereby previously only cited a selection from the entire existing state of the art has been.

In summary, it can be said that in the manufacture of BDO Hydrogenation of MSA technical problems have been solved in that satisfactory yields and selectivities to BDO- only secondary formation of THF - can be achieved. This was achieved through different measures or the Combination of different measures.

Generally, BDO was obtained through the direct hydrogenation of pure GBL - which in turn Hydrogenation of MSA and subsequent elaborate cleaning was obtained - receive. In each case, pure MSA was used as the educt, which was only minimal Contained amounts of impurities, otherwise unsatisfactory selectivity and catalyst life have been reached. There were chromium-containing catalysts used, especially in the second stage, to a high BDO selectivity and to achieve the desired service life. To avoid the use of chromium-containing Alternatively, catalysts could use catalysts containing noble metals be that with yield, selectivity and stability with chromium-containing Catalysts are comparable, but are significantly more expensive.

Furthermore, the preferred reaction was carried out in two separate ways Stages after the first hydrogenation stage, a complex cleaning of the GBL was carried out a long service life of the catalysts, especially with regard to the desired Selectivity to achieve. So far, only in the mentioned WO 97/43234 Process disclosed in which MSA, which was only roughly pre-cleaned, as a starting material for Production of BDO by hydrogenation is used. The process is one-stage carried out, thus also avoids working up after the first hydrogenation stage. however only chromium-containing catalysts are suitable for this reaction.

What is desired, however, is a process for the production of BDO from MSA in which none Pure MSA used and at least no complex cleaning of the reaction products the first stage, preferably none at all Interim cleaning is necessary. Furthermore, the process should not contain any chromium Catalysts, preferably also no catalysts, require the noble metals contain and have a high selectivity to BDO, especially little THF deliver.

• a) Einführen eines Gasstroms einer C₄-Dicarbonsäure oder eines Derivates davon bei 200 bis 300°C und 2 bis 60 bar in einen ersten Reaktor und katalytische Gasphasenhydrierung zu einem hauptsächlich, gegebenenfalls alkylsubstituiertes γ- Butyrolacton enthaltenden Produkt;
• b) Einführen des so erhaltenen Produktstroms in einen zweiten Reaktor bei einer Temperatur von 150°C bis 240°C und einem Druck von 15 bis 100 bar und katalytischer Gasphasenhydrierung zu gegebenenfalls alkylsubstituiertem 1,4- Butandiol;
• c) Abtrennen des gewünschten Produkts von Zwischenprodukten, Nebenprodukten und gegebenenfalls nicht umgesetztem Edukt;
• d) gegebenenfalls Rückführen nicht umgesetzter Zwischenprodukte in eine oder beide Hydrierstufen,

wobei in beiden Hydrierstufen jeweils ein Katalysator verwendet wird, der ≤ 95 Gew.-%, vorzugsweise 5 bis 95 Gew.-%, insbesondere 10 bis 80 Gew.-% CuO und ≥ 5 Gew.-%, vorzugsweise 5 bis 95 Gew.-%, insbesondere 20 bis 90 Gew.-% eines oxidischen Trägers aufweist, im zweiten Reaktor ein höherer Druck herrscht als im ersten Reaktor, und das der ersten Hydrierstufe entnommene Produktgemisch ohne weitere Reinigung in die zweite Hydrierstufe eingeführt wird. This object is achieved by a process for the preparation of optionally alkyl-substituted 1,4-butanediol by two-stage catalytic hydrogenation in the gas phase of C ₄ dicarboxylic acids and / or their derivatives with the following steps:

• a) introducing a gas stream of a C ₄ -dicarboxylic acid or a derivative thereof at 200 to 300 ° C and 2 to 60 bar in a first reactor and catalytic gas phase hydrogenation to a mainly containing, optionally alkyl-substituted γ-butyrolactone product;
• b) introducing the product stream thus obtained into a second reactor at a temperature of 150 ° C to 240 ° C and a pressure of 15 to 100 bar and catalytic gas phase hydrogenation to optionally alkyl-substituted 1,4-butanediol;
• c) separating the desired product from intermediate products, by-products and, if appropriate, unreacted starting material;
• d) optionally recycling unreacted intermediates into one or both hydrogenation stages,

in each of the two hydrogenation stages a catalyst is used which is .- 95% by weight, preferably 5 to 95% by weight, in particular 10 to 80% by weight of CuO and 5 5% by weight, preferably 5 to 95% by weight. -%, in particular 20 to 90 wt .-% of an oxidic carrier, a higher pressure prevails in the second reactor than in the first reactor, and the product mixture removed from the first hydrogenation stage is introduced into the second hydrogenation stage without further purification.

Adherence is more certain to achieve the desired BDO selectivities Reaction parameters required in both hydrogenation stages, these parameters will be listed below.

In the process according to the invention, educts of different purity can be incorporated into the Hydrogenation reaction can be used. Of course, a starting material of high purity, especially MSA, are used in the hydrogenation reaction. An advantage of However, the method according to the invention is that the use of starting materials, especially MSA, which with the usual, in the oxidation of benzene, butenes or n-Butane resulting compounds and possibly other components contaminated is made possible. The hydrogenation process according to the invention can thus be carried out in a further Embodiment include an upstream stage that the manufacture of the hydrogenating educts by partial oxidation of a suitable hydrocarbon and the separation of the starting material to be hydrogenated from the product stream thus obtained includes. In this case, only a rough cut is preferably carried out, which is none requires unnecessarily high effort and one in the previously known Process intolerable amount of impurities remains in the educt.

This educt to be hydrogenated is in particular MSA. MSA, which originates from the partial oxidation of hydrocarbons, is preferably used. Suitable hydrocarbon streams are benzene, C ₄ olefins (e.g. n-butenes, C ₄ raffinate streams) or n-butane. N-Butane is used with particular preference since it is an inexpensive, economical starting material. Process for the partial oxidation of n-butane, for example, in Ullmann's Encyclopedia of Industrial Chemistry, 6 ^th Edition, Electronic Release, Maleic and Fumaric Acids - describes Maleic Anhydride.

The reaction discharge thus obtained is then preferably in a suitable one organic solvent or mixture added, which at atmospheric pressure boiling point at least 30 ° C higher than MSA.

This solvent (absorbent) is brought to a temperature in the range between 20 and 160 ° C, preferably between 30 and 80 ° C. The gas stream from the partial oxidation containing maleic anhydride can be brought into contact with the solvent in a variety of ways: (i) introducing the gas stream into the solvent (e.g. via gas inlet nozzles or gassing rings), (ii) spraying the solvent into the gas stream and ( iii) Countercurrent contact between the upward gas stream and the downward solvent in a tray or packing column. The gas absorption apparatus known to those skilled in the art can be used in all three variants. When choosing the solvent to be used, care must be taken to ensure that this does not react with the starting material, for example the preferably used MA. Suitable solvents are: tricresyl phosphate, dibutyl maleate, butyl maleate, high molecular waxes, aromatic hydrocarbons with a molecular weight between 150 and 400 and a boiling point above 140 ° C, such as dibenzylbenzene; Alkyl phthalates and dialkyl phthalates with C ₁ -C ₁₈ alkyl groups, for example dimethyl phthalate, diethyl phthalate, dibutyl phthalate, di-n-propyl and di-iso-propyl phthalate, undecyl phthalate, diundecyl phthalate, methyl phthalate, ethyl phthalate, butyl phthalate, n-propyl or isopropyl phthalate ; Di-C ₁ -C ₄ alkyl esters of other aromatic and aliphatic dicarboxylic acids, for example dimethyl-2,3-naphthalene-dicarboxylic acid, dimethyl-1,4-cyclohexane-dicarboxylic acid; C ₁ -C ₄ alkyl esters of other aromatic and aliphatic dicarboxylic acids, for example methyl 2,3-naphthalene dicarboxylic acid, methyl 1,4-cyclohexane dicarboxylic acid, methyl esters of long-chain fatty acids with, 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 solution resulting after treatment with the absorbent generally has an MSA content of about 5 to 400 grams per liter.

Contains the exhaust gas stream remaining after treatment with the absorbent mainly the by-products of the previous partial oxidation, such as water, Carbon monoxide, carbon dioxide, unreacted butanes, acetic and acrylic acid. The Exhaust gas flow is practically free of MSA.

The dissolved MSA is then expelled from the absorbent. this happens with hydrogen at or a maximum of 10% above the pressure of the subsequent Hydrogenation or alternatively in a vacuum with subsequent condensation of remaining MSA. A temperature profile is observed in the stripping column that changes from the boiling points of MSA on the head and the almost MSA-free absorbent at the bottom of the column at the respective column pressure and the set Dilution with carrier gas (in the first case with hydrogen) results.

To prevent loss of solvent, rectification internals can be located above the supply of the raw MSA stream. The almost MSA-free absorbent drawn off from the sump is returned to the absorption zone. The H ₂ / MSA ratio is about 20 to 400. In the other case, the condensed MSA is pumped into an evaporator and evaporated there into the circulating gas stream.

The MSA hydrogen stream still contains by-products that are involved in the partial oxidation of n-butane, butenes or benzene with gases containing oxygen, and absorbent not separated. This is mainly acetic acid and Acrylic acid as by-products, water, maleic acid and preferably as Absorbent used dialkyl phthalates. The MSA contains acetic acid in quantities from 0.01 to 1% by weight, preferably 0.1 to 0.8% by weight and acrylic acid in amounts of 0.01 to 1 wt .-%, preferably 0.1 to 0.8 wt .-%, based on MA. In the The hydrogenation stage turns all or part of acetic acid and acrylic acid into ethanol or Hydrogenated propanol. The maleic acid content is 0.01 to 1% by weight, in particular 0.05 up to 0.3 wt .-%, based on MSA.

If dialkyl phthalates are used as absorbents, their content depends on the MSA strongly depend on the correct operation of the stripping column, in particular on the rectifying section. Phthalate contents of up to 1.0% by weight, in particular up to 0.5% by weight, should be suitable Operating mode must not be exceeded, otherwise the consumption of absorbent increases gets high.

The one preferably obtained as described above Hydrogen / maleic anhydride stream is now fed to the first hydrogenation zone and hydrogenated. The Catalyst activities and service lives are compared to the use of strong, for example by distillation, pre-cleaned MSA practically unchanged.

According to the invention, the gas stream emerging from the first reactor can have several Species are processed further.

According to a variant, it becomes gaseous to the higher pressure of the second hydrogenation stage compressed and so, with possibly recycled GBL, the second hydrogenation fed.

According to a further variant, the gas stream can be cooled to 10 to 60 ° C. there the reaction products are condensed out and passed into a separator. The uncondensed gas stream can be separated and, preferably after being fed into a Recycle gas compressors are returned to the first hydrogenation circuit. resulting By-products in the recycle gas stream can be known by those skilled in the art Measures are removed, preferably by discharging a small amount Cycle gas. The condensed reaction products are removed from the system and fed to the second hydrogenation circuit. The reaction products are there under pressure optionally recycled GBL in the gas phase and in with the second catalyst Brought in contact. The GBL which is possibly recycled can also be in gaseous form be brought to the second hydrogenation reactor.

According to a further variant, the GBL-laden gas stream is opened from the first stage compresses the pressure of the second stage and the return gas of the second stage in the Entry of the first stage expands, possibly under work.

In all reaction variants, the gas stream emerging from the second reactor cooled, preferably to 10 to 60 ° C. The reaction products condensed out and passed into a separator. The uncondensed gas stream will withdrawn from the separator and fed compressed to the cycle gas. A small Circular gas quantity is discharged. The condensed reaction products are the System preferably removed continuously and processed. As By-products are mainly found in the condensed liquid phase THF, n-butanol in addition to small amounts of propanol.

The by-products then become from the liquid hydrogenation discharge of the second stage and separated water and isolated the desired product BDO. this happens generally by fractional distillation. By-products and intermediates such as for example GBL and Di-BDO can be used in the hydrogenation of the first and / or second Stage, preferably the second stage, recycled or alternatively by distillation be worked up.

The process according to the invention can be discontinuous, semi-continuous or be carried out continuously. Continuous implementation is preferred.

An important parameter is maintaining a suitable reaction temperature in both Hydrogenation.

In the first hydrogenation stage this is preferably achieved by a sufficiently high one Temperature of the starting materials when entering the first hydrogenation reactor. This so-called The initial hydrogenation temperature is from 200 to 300 ° C., preferably 235 to 270 ° C. In order to obtain the desired selectivity and yield in the first stage, the reaction is preferably carried out so that on the catalyst bed at which the actual reaction takes place, there is a suitably high reaction temperature. This so-called hot spot temperature is determined after the educts have entered the reactor set and is preferably at values of 210 to 310 ° C, in particular 245 to 280 ° C. The process is preferably carried out so that the inlet temperature and the exit temperature of the reaction gases below this hot spot temperature lie. The hot spot temperature is advantageously in the 1st half of the Reactor, especially in the presence of a tube bundle reactor. Preferably, the Hot spot temperature 5 to 30 ° C, in particular 5 to 15 ° C, particularly preferably 5 to 10 ° C, above the inlet temperature. If the hydrogenation is below the Minimum temperatures of the input or hot spot temperature is then carried out generally increases the amount of BSA when using MSA as a starting material simultaneous reduction in the amount of GBL and BDO. Furthermore, one such temperature in the course of the hydrogenation by deactivating the catalyst Occupation with succinic acid, fumaric acid and / or BSA and a mechanical Observe damage to the catalyst. In contrast, with MSA as the starting material above of the maximum temperatures of the inlet or hot spot temperature, the drop BDO yield and selectivity generally to unsatisfactory values. It is here to observe the increased formation of THF, n-butanol and n-butane, i.e. the Products of a further hydrogenation.

In the second hydrogenation stage, the inlet temperature (initial hydrogenation temperature) is included Values from 150 ° C to 260 ° C, preferably 175 ° C to 225 ° C, especially at 180 to 200 ° C. If the hydrogenation is below the minimum temperature of the inlet temperature is carried out, the amount of BDO formed decreases. The catalyst loses Activity. Furthermore, the condensation is below the minimum temperature Input materials and damage to the copper catalyst caused by water. Becomes on the other hand with GBL as the educt above the maximum temperatures of the inlet temperature hydrogenated, the BDO yield and selectivity drop to unsatisfactory values. At these temperatures the hydrogenation equilibrium is between BDO and GBL Sides of the GBL, so that less sales is achieved, however, is the by-product formation increased by overhydrogenation to THF, n-butanol and n-butane at a higher temperature observe.

The temperature increase of the gas stream in the reactor should be 110 ° C, preferably 40 ° C not exceed, in particular not more than 20 ° C. Big ones lead here too Temperature increases often lead to overhydrogenation reactions and a (BDO + GBL) Loss of selectivity.

In the first hydrogenation stage, a pressure of 2 to 60 bar, preferably a pressure of 2 to 20 bar and particularly preferably a pressure of 5 to 15 bar. In this The pressure range in the hydrogenation of MA is initially the formation of THF from the resulting intermediate product GBL largely suppressed.

In the second hydrogenation stage, a pressure of 15 to 100 bar, preferably a pressure of 35 to 80 bar and particularly preferably a pressure of 50 to 70 bar. With increasing pressure rises at the temperatures applied in the second hydrogenation stage GBL sales to BDO. With higher pressure, a lower GBL Return rate can be selected. The pressure in the second hydrogenation stage is above the Pressure of the first hydrogenation stage.

The catalyst loading of the first hydrogenation stage is preferably in the range of 0.02 to 1, in particular 0.05 to 0.5 kg of educt / l of catalyst per hour. Will the First stage catalyst loading in the case of MA above the range mentioned increased, there is generally an increase in the proportion of BSA and succinic acid in the Watch hydrogenation discharge. The catalyst load is preferably the second Hydrogenation stage in the range from 0.02 to 1.5, in particular 0.1 to 1 kg of educt / l of catalyst. Hour. If the catalyst load is increased beyond the range mentioned, is expect incomplete sales to GBL. This can be done by a increased return rate can be compensated, but this is of course not prefers.

Furthermore, the hydrogen / educt molar ratio is a parameter that influences the product distribution and also the economy of the process according to the invention Has. From an economic point of view is a low hydrogen / educt ratio desirable. The lower limit is 5, but generally higher Hydrogen / reactant molar ratios of 20 to 600 are used. The use of the catalysts used according to the invention and compliance with the catalysts according to the invention applied temperature values allows the use of low-cost, low hydrogen / educt Ratios in the hydrogenation of the first stage, preferably at values from 20 to 200, preferably 40 to 150. The cheapest range is from 50 to 100th

In order to adjust the hydrogen / educt molar ratios used according to the invention, is usually a part, advantageously the bulk, of both hydrogen cycled in the first as well as in the second hydrogenation stage. To do this, set in generally the cycle gas compressors known to the person skilled in the art. The chemically through the Hydrogenation consumed amount of hydrogen is added. In a preferred one Embodiment, part of the circulating gas is discharged to form inert compounds, for example to remove n-butane. The hydrogen circulated can also if necessary after preheating, to evaporate the educt stream.

Together with the hydrogen cycle gas, all products are recycled not or not when cooling the gas streams emerging from the hydrogenation reactors fully condense. These are mainly THF, water and by-products such as Methane and butane. The cooling temperature is preferably 0 to 60 ° C, preferably 20 up to 45 ° C.

All reactor types come with one for heterogeneously catalyzed reactions gaseous educt and product stream suitable apparatus. Are preferred Tube reactors, shaft reactors or reactors with internal heat dissipation, for example Tube bundle reactors, it is also possible to use a fluidized bed. Especially tube bundle reactors are preferably used for the first hydrogenation stage and for the second Hydrogenation stage chimney reactors are particularly preferred. It can be used in both first and also the second hydrogenation stage several reactors in parallel or in series can be used switched. In principle, one can be placed between the catalyst beds Intermediate feed. Intercooling between or in is also possible the catalyst beds. When using fixed bed reactors, a dilution of the Catalyst possible through inert material.

An important point of the present invention is the choice of catalysts for both Steps that have copper oxide as the main catalytically active component. This is on an oxidic carrier, which may have a small number of acidic centers, appropriate. If a catalyst with too many acidic centers is used, BDO is dehydrated and THF is formed.

A suitable carrier material which has a sufficiently small number of acidic centers is a material selected from the group ZnO, Al ₂ O ₃ , SiO ₂ , TiO ₂ , ZrO ₂ , CeO ₂ , MgO, CaO, SrO, BaO and Mn ₂ O ₃ and mixtures thereof. Preferred support materials are ZnO / Al ₂ O ₃ mixtures, the delta, theta, alpha and eta modifications of Al ₂ O ₃ and mixtures which contain at least one component from the group SiO ₂ , TiO ₂ , ZrO ₂ on the one hand and from the group ZnO, MgO, CaO, SrO and BaO on the other hand. Pure ZnO, ZnO / Al ₂ O ₃ mixtures in a weight ratio of 100: 1 to 1: 2, and mixtures of SiO ₂ with MgO, CaO and / or ZnO in a weight ratio of 200: 1 to 1: 1 are particularly preferred as carrier materials.

The amount of copper oxide is values ≤ 95% by weight, preferably 5 5 to 95% by weight, in particular at 15 to 80% by weight; the carrier is used in amounts of 5% by weight, preferably 5 to 95 wt .-%, in particular 20 to 85 wt .-% used.

Chromium-containing catalysts are preferred because of the toxicity of chromium-containing catalysts Catalysts used. Of course, appropriate technical Chromium-containing catalysts known to the person skilled in the art for use in the Process according to the invention, but this does not have the desired advantages are in particular of an environmental and work-related nature.

The same catalyst can be used in both hydrogenation stages, is preferred however, the use of different catalysts.

The catalysts used according to the invention can optionally have one or more further metals or a compound thereof, preferably an oxide, from groups 1 to 14 (IA to VIIIA and IB to IVB of the old IUPAC nomenclature) of the periodic table of the elements included. If another metal is used, Pd is preferably used used in amounts of 1% by weight, preferably ≤ 0.5% by weight, in particular ≤ 0.2% by weight. However, the addition of another metal or metal oxide is not preferred.

The catalysts used can also contain an auxiliary in an amount of 0 to 10% by weight. contain. Aids are organic and inorganic substances that to improved processing during catalyst production and / or to Contribute to increasing the mechanical strength of the shaped catalyst bodies. such Aids are known to the person skilled in the art; Examples include graphite, stearic acid, Silica gel and copper powder.

The catalysts can be prepared by methods known to those skilled in the art. Methods are preferred in which the copper oxide is finely divided and intimately mixed with the other components, impregnation and Precipitation reactions.

These starting materials can be made into the shaped bodies by known methods processed, for example extruding, tableting or through Agglomeration process, optionally with the addition of auxiliary agents.

Alternatively, catalysts according to the invention can also be applied, for example the active component can be produced on a carrier, for example by Coating or vapor deposition. Furthermore, catalysts according to the invention can be used Deforming a heterogeneous mixture of active component or precursor compound thereof with a carrier component or precursor compound thereof.

In the hydrogenation according to the invention, in which, in addition to MA, other C ₄ dicarboxylic acids or their derivatives defined above can be used as starting materials, the catalyst is used in a reduced, activated form. The activation takes place with reducing gases, preferably hydrogen or hydrogen / inert gas mixtures, either before or after installation in the reactors in which the process according to the invention is carried out. If the catalyst has been installed in the reactor in oxidic form, the activation can be carried out both before starting up the plant with the hydrogenation according to the invention and during start-up, that is to say in situ. The separate activation before starting up the system is generally carried out using reducing gases, preferably hydrogen or hydrogen / inert gas mixtures, at elevated temperatures, preferably between 100 and 350 ° C. In the so-called in-situ activation, the activation takes place when the system is started up by contact with hydrogen at an elevated temperature.

The catalysts are preferably used as moldings. Examples include strands, rib strands, other extrudate forms, tablets, rings, balls and grit. The BET surface area of the copper catalysts in the oxidic state should be 10 to 300 m ² / g, preferably 15 to 175 m ² / g, in particular 20 to 150 m ² / g. The copper surface (N ₂ O decomposition) of the reduced catalyst should be> 0.2 m ² / g, preferably> 1 m ² / g, in particular> 2 m ² / g in the installed state.

According to a variant of the invention, catalysts are used which have a defined one Have porosity. As shaped bodies, these catalysts have a pore volume of 0.01 ml / g for pore diameters> 50 nm, preferably 0.025 ml / g for Pore diameter> 100 nm and in particular> 0.05 ml / g for pore diameter> 200 nm. Furthermore, the ratio of macropores with a diameter> 50 nm to Total pore volume for pores with a diameter> 4 nm at values> 10%, preferably> 20%, in particular> 30%. The porosities mentioned have been through Mercury intrusion determined according to DIN 66133. The data in the Pore diameter range from 4 nm to 300 µm evaluated.

The catalysts used according to the invention generally have one sufficient service life. In the event that the activity and / or selectivity of the If the catalyst should nevertheless fall in the course of its operating time, this can be caused by the Measures known to those skilled in the art can be regenerated. This preferably includes one reductive treatment of the catalyst in a hydrogen stream at elevated temperature. If necessary, the reductive treatment can be preceded by an oxidative treatment. Here will the catalyst bed with a gas mixture containing molecular oxygen, for example air, flows through at elevated temperature. Furthermore, the Possibility of the catalysts with a suitable solvent, for example Wash ethanol, THF, BDO or GBL, and then in a gas stream dry.

The process according to the invention is now described in more detail in the examples below explained.

example 1 a) Catalyst production

From a metal salt solution containing copper nitrate and zinc nitrate with soda at 50 ° C and a mixed basic metal carbonate precipitates around pH 6.2. The used metal salt solution contained the metals according to one Catalyst composition of 70% CuO and 30% ZnO.

The precipitate is filtered, washed, dried, calcined at 300 ° C and with 3 wt .-% Graphite pressed into tablets of 3 mm height and diameter.

b) catalyst activation

Before the start of the reaction, the catalyst is subjected to a hydrogen treatment in the hydrogenation apparatus. For this purpose, the reactor is heated to 180 ° C. and the catalyst is activated for the time specified in Table 1 with the mixture of hydrogen and nitrogen indicated in each case at atmospheric pressure. Table 1

c) hydrogenation apparatus

The pressure equipment used for the hydrogenation consists of an evaporator, a Reactor, a cooler with quench feed, a hydrogen supply, an exhaust pipe and a cycle gas blower. The pressure in the apparatus is kept constant.

The melted MSA is placed on top of the preheated (245 ° C) evaporator pumped and evaporated. A mixture also gets onto the evaporator from above from fresh hydrogen and cycle gas. Hydrogen and MSA get in from below the temperature-controlled reactor. The reactor contents consist of a mixture of glass rings and catalyst. After the hydrogenation, the resulting GBL leaves together with Water, other reaction products and hydrogen enter the reactor and is in the cooler put down by quenching. Part of the cycle gas is discharged before the The rest, mixed with fresh hydrogen, re-enters the evaporator.

The condensed liquid reaction discharge, the exhaust gas and the cycle gas analyzed by gas chromatography.

At a reactor temperature of 255 ° C, a pressure of 5 bar and a catalyst load of 0.27 kg / l _Kat h with a hydrogen: MA ratio of 85: 1, a reaction discharge of the composition: 91% GBL, 5% THF, 1% BDO, 1% BSA.

Example 2 a) Catalyst production

A solid with the composition 64% ZnO and 36% Al ₂ O ₃ (based on 100% oxide) is precipitated, filtered and filtered from an aqueous solution of zinc nitrate and aluminum nitrate with soda solution at 50 ° C and a pH of 6.8 washed. The filter cake was dried and calcined at 425 ° C for one hour.

The carrier described above is added to a nitric acid solution of copper nitrate and zinc nitrate (metal ratio corresponding to 16.6% by weight CuO and 83.4% CuO) and mixed intensively at 70 ° C. A solid is precipitated from this mixture with sodium carbonate solution at 70 ° C. and a pH of 7.4, and the suspension is left to stir at constant temperature and pH for 2 hours. The solid is filtered off, washed, dried and calcined at 430 ° C. for one hour. The catalyst powder obtained in this way was mixed with 1.5% by weight of graphite and 5% by weight of copper powder and compressed into tablets of 1.5 mm in diameter and 1.5 mm in height. The tablets were finally calcined at 330 ° C. for 1 hour, had a lateral compressive strength of 50 N and a chemical composition of 66% CuO / 24% ZnO / 5% Al ₂ O ₃ /5% Cu.

b) catalyst activation analogous to example 1b c) hydrogenation apparatus

The reactor of the hydrogenation apparatus described in Example 1c is treated with 220 ml of the Example 2a prepared catalyst and filled 130 ml glass rings. The activation was carried out as described in Example 1b.

The reaction discharge from the MA hydrogenation from Example 1 is used as starting material. At a reactor temperature of 180 ° C, a pressure of 60 bar and a catalyst load of 0.15 kg / l _cat h (hydrogen: GBL molar ratio 200: 1), a reaction discharge of the composition is obtained: 86% BDO, 7% GBL, 6% THF ,

Claims (25)

1. A process for the preparation of optionally alkyl-substituted 1,4-butanediol by two-stage catalytic hydrogenation in the gas phase of C ₄ dicarboxylic acids and / or their derivatives, with the following steps: a) introducing a gas stream of a C ₄ -dicarboxylic acid or a derivative thereof at 200 to 300 ° C and 2 to 60 bar in a first reactor and catalytic gas phase hydrogenation to a mainly, optionally alkyl-substituted γ-butyrolactone containing product; b) introducing the product stream thus obtained into a second reactor at a temperature of 150 ° C to 240 ° C and a pressure of 15 to 100 bar and catalytic gas phase hydrogenation to optionally alkyl-substituted 1,4-butanediol; c) separating the desired product from intermediate products, by-products and, if appropriate, unreacted starting material; d) optionally recycling unreacted intermediates into one or both hydrogenation stages, in each of the two hydrogenation stages a catalyst is used which is .- 95% by weight, preferably 5 to 95% by weight, in particular 10 to 80% by weight of CuO and 5 5% by weight, preferably 5 to 95% by weight. -%, in particular 20 to 90 wt .-% of an oxidic carrier, a higher pressure prevails in the second reactor than in the first reactor, and the product mixture removed from the first hydrogenation stage is introduced into the second reactor without further purification. 2. The method according to claim 1, characterized in that the inlet temperature in the first reactor at values from 235 to 270 ° C and the inlet temperature in the second reactor at values from 175 ° C to 225 ° C, especially at 180 to 200 ° C. 3. The method according to claim 1 or 2, characterized in that the hot spot Temperature in the first reactor at values from 210 to 310 ° C, preferably 245 to 280 ° C and the procedure is carried out so that the hot spot temperature above the inlet temperature and the outlet temperature of the reaction gases the hot spot temperature is 5 to 30 ° C, in particular 5 to 15 ° C, particularly preferably 5 to 10 ° C, is above the inlet temperature. 4. The method according to any one of claims 1 to 3, characterized in that in the first hydrogenation stage a pressure of 2 to 20 bar, preferably a pressure of 5 to 15 bar, and in the second hydrogenation stage a pressure of 35 to 80 bar, preferably there is a pressure of 50 to 70 bar. 5. The method according to any one of claims 1 to 4, characterized in that the Catalyst loading of the first hydrogenation stage in the range from 0.02 to 1, in particular 0.05 to 0.5 kg of educt / l of catalyst.hour, and the Catalyst loading of the second hydrogenation stage in the range from 0.02 to 1.5 in particular 0.1 to 1 kg of educt / l of catalyst per hour 6. The method according to any one of claims 1 to 5, characterized in that the Hydrogen / reactant molar ratio in both reaction stages at values> 5. preferably 20 to 600. 7. The method according to claim 6, characterized in that the hydrogen / starting material Ratio in the hydrogenation of the first stage at values from 20 to 200, preferably 40 to 150, in particular 50 to 100. 8. The method according to any one of claims 1 to 7, characterized in that the reactors used are selected from the group consisting of Tubular reactors, shaft reactors, reactors with internal heat dissipation, Shell and tube reactors and fluidized bed reactors. 9. The method according to claim 8, characterized in that in the first Hydrogenation stage a tube bundle reactor is used. 10. The method according to claim 8 or 9, characterized in that in the second Hydrogenation stage a shaft reactor is used. 11. The method according to any one of claims 1 to 10, characterized in that in the first and / or the second hydrogenation stage several reactors in parallel or can be used in series. 12. The method according to any one of claims 1 to 11, characterized in that the catalyst support material selected from the group ZnO, Al ₂ O ₃ , SiO ₂ , TiO ₂ , ZrO ₂ , CeO ₂ , MgO, CaO, SrO, BaO and Mn ₂ O ₃ and mixtures thereof, preferably from the group ZnO / Al ₂ O ₃ mixtures, delta, theta, alpha and eta modifications of Al ₂ O ₃ and mixtures which contain at least one component from the group SiO ₂ , TiO ₂ , ZrO ₂ on the one hand and from the group ZnO, MgO, CaO, SrO and BaO on the other hand. 13. The method according to any one of claims 1 to 12, characterized in that the carrier material is selected from ZnO, ZnO / Al ₂ O ₃ mixtures in a weight ratio of 100: 1 to 1: 2 and mixtures of SiO ₂ with MgO, CaO and / or ZnO in a weight ratio of 200: 1 to 1: 1. 14. The method according to any one of claims 1 to 13, characterized in that the Catalyst one or more other metals, preferably Pd, or one Connection of one or more other metals, preferably an oxide, from the Contains groups 1 to 14 of the Periodic Table of the Elements. 15. The method according to any one of claims 1 to 14, characterized in that the Catalyst is used as a shaped body, preferably in the form of strands, Ribbed strands, tablets, rings, balls or grit. 16. The method according to any one of claims 1 to 15, characterized in that the BET surface area of the copper catalysts in the oxidic state is 10 to 300 m ² / g, preferably 15 to 175 m ² / g, in particular 20 to 150 m ² / g , 17. The method according to any one of claims 1 to 16, characterized in that the copper surface 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. 18. The method according to any one of claims 1 to 17, characterized in that the im the first and second reactors used the same or different from each other are different, preferably are different from each other. 19. The method according to any one of claims 1 to 18, characterized in that the used catalyst as a shaped body a pore volume of ≥ 0.01 ml / g for Pore diameter> 50 nm, preferably ≥ 0.025 ml / g for pore diameter> 100 nm and in particular ≥ 0.05 ml / g for pore diameters> 200 nm. 20. The method according to any one of claims 1 to 19, characterized in that the Ratio of macropores with a diameter> 50 nm to Total pore volume for pores with a diameter> 4 nm at values> 10%, preferably> 20%, in particular> 30%. 21. The method according to any one of claims 1 to 20, characterized in that Maleic anhydride is used as a starting material in the reaction. 22. The method according to any one of claims 1 to 21, characterized in that maleic anhydride is used, which was prepared by oxidation of benzene, C ₄ -olefins or n-butane, the crude maleic anhydride obtained by oxidation with a solvent from the crude product mixture extracted and then expelled from this solvent with hydrogen. 23. The method according to any one of claims 1 to 22, characterized in that the absorbent is selected from the group consisting of tricresyl phosphate, dibutyl maleate, high molecular waxes, aromatic hydrocarbon with a molecular weight between 150 and 400 and a boiling point above 140 ° C, preferably dibenzene , Di-C ₁ -C ₄ -alkyl esters of aromatic and aliphatic dicarboxylic acids, preferably dimethyl-2,3-naphthalene-dicarboxylic acid and / or dimethyl-1,4-cyclohexane-dicarboxylic acid, methyl esters of long-chain fatty acids with 14 to 30 carbon atoms, high-boiling ethers, preferably dimethyl ether of polyethylene glycol, preferably of tetraethylene glycol, and alkyl phthalates and dialkyl phthalates with C ₁ -C ₁₈ alkyl groups, preferably from the group of dimethyl phthalate, diethyl phthalate, dibutyl phthalate, di-n-propyl and di-iso-propyl phthalate, undecyl phthalate, diundecyl phthalate, methyl phthalate , Ethyl phthalate, butyl phthalate, n-propyl and isopropyl p hthalat. 24. The method according to any one of claims 1 to 23, characterized in that the Maleic anhydride from the absorbent in vacuo or at pressures that correspond to the pressure of the hydrogenation or at most 10% above it Pressure, is expelled. 25. The method according to any one of claims 1 to 24, characterized in that the Process carried out batchwise, semi-continuously or continuously is, preferably continuously.