Process for producing 1,4-butanediol, gamma-butyrolactone and tetrahydrofuran in the gas phase while avoiding polymer deposition

Abstract

The present invention relates to an improved process for producing 1,4-butanediol, tetrahydrofuran and gamma-butyrolactone, wherein an acid-containing maleic ester stream is hydrogenated in the presence of hydrogen in a single reactor comprising at least two different adjacent catalyst layers at specific different temperature levels in the catalyst layers, wherein the first layer is a noble metal-containing catalyst layer, which is followed by a second Zn-free Cu catalyst layer, and is subsequently subjected to a distillative work-up in such a way that polymer deposition and aging and damage of the catalyst in the reactor are prevented.

Description

[0001] The present invention relates to an improved process for producing 1,4-butanediol, tetrahydrofuran and gamma-butyrolactone, wherein an acid maleate stream is hydrogenated in the presence of hydrogen in a single reactor comprising at least two different adjacent catalyst layers at specific different temperature levels in the catalyst layers, wherein the first layer is a noble metal-containing catalyst layer followed by a second Zn-free Cu catalyst layer, and is subsequently worked up by distillation in such a way that polymer deposition as well as aging and damage to the catalyst in the reactor is prevented.

[0002] The hydrogenation of maleic acid derivatives such as maleic acid esters, especially dimethyl maleate and diethyl maleate in the gas phase to provide products such as 1,4-butanediol (BDO), tetrahydrofuran (THF), γ-butyrolactone (GBL) and the corresponding alcohols of the esters is known and has been used on an industrial scale. For example, WO-A 91 / 01960 describes this process. However, a disadvantage of the process described in WO-A 91 / 01960 is that the hydrogenation of maleic acid diesters comprising maleic acid monoesters over a copper catalyst leads to the formation of deposits at the beginning of the catalyst zone over time, which leads to a significant increase in the pressure drop, which in turn makes it necessary to interrupt the hydrogenation and replace the catalyst.

[0003] WO-A 99 / 35113 describes the hydrogenation of maleic diesters in the gas phase in a first reactor zone over a noble metal catalyst, preferably palladium on alumina, to provide succinic diesters, followed by hydrogenation over a copper catalyst to provide mainly GBL and THF. In a further second reactor, GBL is hydrogenated to provide BDO. The feed stream to the hydrogenation comprises maleic diesters and the recycle stream comprises BDO and succinic diesters. No acid is mentioned in the feed. A disadvantage of this process is that the production of BDO requires the use of two different reactors and that when BDO is recycled into the feed, deposits of oligoesters and polyesters formed by the reaction of BDO with esters may form over time, at least in the evaporator. In addition, yield losses must be expected, since BDO-maleic acid esters or -succinic acid esters are high-boiling compounds and must be discharged during the evaporation process.

[0004] WO-A 2005 / 058855 describes the conversion of maleic diesters into THF and optionally BDO and GBL by hydrogenation in the liquid phase of at least an unsaturated diester substream in a pre-reactor zone to provide at least partially succinic diesters. This can be carried out in a separate reactor or in the same reactor, wherein further hydrogenation is subsequently carried out after complete evaporation to provide THF and optionally BDO and GBL. This may have evaporated part of the ester mixture, especially if double bond hydrogenation is carried out in the same reactor as the reactor in which further hydrogenation is subsequently carried out. This method, which seems perfect at first glance, has several disadvantages. The use of additional reactors leads to increased costs. In addition, evaporation—even if only partially carried out in the presence of a catalyst—usually causes catalyst wear, because evaporation occurs preferentially at the highest temperature. Therefore, in this case it occurs in the catalyst pores, where the exothermic hydrogenation of the double bonds leads to the release of heat. In addition, evaporation leads to a substantial increase in volume, which may cause the catalyst molded body to break, thereby causing mechanical damage to the catalyst over time. A further disadvantage of the process described in WO-A 2005 / 058855 is the use of a copper-containing catalyst which, as described in WO-A 2005 / 058855, not only hydrogenates the double bonds of the diester, but also further hydrogenates the product to form BDO. However, this has the disadvantage that the resulting BDO reacts with the diester to form oligomers and polymers which are difficult to evaporate and thus lead to polymer deposition on the catalyst in the gas phase, with the result that the catalyst deteriorates very rapidly with regard to conversion, yield and selectivity and requires replacement.

[0005] WO-A 01 / 44148 describes a process in which esters which may still contain traces of acid, such as maleic acid diesters, are hydrogenated over three different catalyst zones to provide BDO, THF and GBL. Each of these zones is capable of hydrogenating esters. A Cu catalyst can be used in the first hydrogenation zone, but Pd catalysts are also mentioned. A disadvantage of this process using catalysts, each of which can hydrogenate esters, is that the diols formed may form high-boiling compounds with the diesters, especially in the presence of acids, which may then lead to deposits. In addition, it is inconvenient and causes additional costs to recover the product streams after each hydrogenation and then introduce them into the subsequent hydrogenation zone.

[0006] For hydrogenating maleic diesters to provide BDO and THF, CN-A 101891592 describes two stages each comprising a catalyst, wherein the first stage uses a Group VIII catalyst and the second stage uses a Cu catalyst. The hydrogenation feed may also contain recycled intermediates. Succinic diesters and GBL are described. CN-A 101891592 also describes that polymer deposition of succinic diesters and BDO may occur when using a Zn-containing Cu catalyst, since this always causes hydrogenation to provide BDO in addition to the hydrogenation of the double bonds. The disadvantage of this process is that it also does not use an acid-containing maleate stream, but only a purified diester stream, and the Cu catalyst layer after the Group VIII catalyst always contains Zn. The Zn-containing Cu catalyst is alkaline and preferentially forms corresponding Zn salts with the monoesters of the acid-containing maleate stream. These Zn salts then stick to the catalyst and coke there due to the high temperature. This in turn results in the catalyst used having only a short service life.

[0007] It was therefore an object of the present invention to provide a process for producing BDO, GBL and THF which makes it possible to obtain the corresponding products in an economical manner and in good yields and purities, without increasing the formation of by-products and without having to use particularly purified starting materials or having to take into account the service life of the catalyst.

[0008] The object is achieved by a method for producing 1,4-butanediol, γ-butyrolactone and tetrahydrofuran, the method comprising the following steps:

[0009] a) hydrogenating an acid-containing maleate stream with hydrogen in the gas phase in a single reactor, wherein the acid-containing maleate stream is hydrogenated on a catalyst fixed bed comprising at least two different and directly adjacent catalyst layers, the catalyst layer passed through first in the operating direction of the hydrogenation is a noble metal-containing catalyst layer comprising a noble metal selected from the group consisting of Pd and Pt and the catalyst layer passed through subsequently is a Zn-free Cu-containing catalyst layer, wherein the proportion of 1,4-butanediol after passing through the noble metal-containing catalyst layer is in the range from 0 to 0.5% by weight, based on the total amount of organic compounds in the product stream and the temperature of the gas stream after passing through the noble metal-containing catalyst layer and before entering the Zn-free Cu-containing catalyst layer is increased by at least 5° C. relative to the inlet temperature into the catalyst fixed bed and subsequently after further hydrogenation in the last catalyst layer of the fixed bed,

[0010] b) separating the resulting mixture by distillation in one or more columns into three mixtures, wherein each mixture contains one of the three products selected from 1,4-butanediol, γ-butyrolactone and tetrahydrofuran in excess.

[0011] The process according to the invention is preferred when the acid maleic ester stream has a content of maleic acid monoesters in the range from 0.001 to 2% by weight, based on the proportion of maleic acid diesters.

[0012] When the noble metal in the noble metal-containing catalyst layer has been applied to a substrate selected from carbon (e.g., activated carbon or burnt nut shells), argillaceous earth, zeolite, TiO 2 、ZrO 2 、Al 2 O 3 、SiO 2 The method of the present invention is preferred when the carrier contains a mixture of the components.

[0013] The process according to the invention is preferred when the noble metal-containing catalyst layer comprises palladium as the noble metal in the range from 0.001 to 15% by weight, based on the total weight of the catalyst.

[0014] The process according to the invention is preferred when the temperature of the gas stream is increased in the noble metal-containing catalyst layer such that an adiabatic temperature increase of at least 33% has been achieved along the catalyst fixed bed after passing through the noble metal-containing catalyst layer.

[0015] The process according to the invention is preferred when the hydrogenation is carried out at a temperature in the range from 150 to 250° C. and a pressure in the range from 8 to 80 bar.

[0016] The process of the invention is preferred when the Zn-free Cu catalyst layer following the noble metal-containing catalyst layer is selected from reduced copper chromite catalysts and reduced copper oxide catalysts which may also contain other metals, metalloids and / or transition metals as promoters, wherein Zn and its compounds are excluded from the promoters.

[0017] When the Zn-free Cu catalyst comprises a 2 O 3 、SiO 2 、TiO 2 、ZrO 2 , Mn 2 O 3 Cr 2 O 3 and La 2 O 3 The method of the present invention is preferred when other components are present.

[0018] The process according to the invention is preferred when the acid maleate stream also comprises tetrahydrofuran as a recycle stream in the range from 0.1 to 20% by weight, based on the maleic diester content.

[0019] The process according to the invention is preferred when the acid-containing maleic acid ester stream also comprises gamma-butyrolactone and succinic acid monoesters or diesters in a total content of up to 50% by weight, based on the maleic acid diester content.

[0020] The acid-containing maleate stream to be hydrogenated is produced by known methods, for example by oxidation of benzene or butane via formation of maleic anhydride or maleic acid and by esterification with an alcohol, preferably methanol or ethanol, particularly preferably methanol. This requires the removal of water to ensure that the conversion to the diester is as complete as possible. The higher the required conversion, the greater the technical complexity. Therefore, there is a great need for a process for the production of BDO, GBL and THF, in which the acid-containing maleate stream obtained by the production can be introduced into the further hydrogenation without additional post-treatment and without incurring losses in terms of yield, selectivity or catalyst service life.

[0021] The inventive method represents the method. The acid-containing maleate stream for hydrogenation in step a) of the inventive method is understood to refer to a mixture of maleic diesters formed as a main product in the esterification of maleic anhydride or maleic acid with the corresponding alcohol and secondary components formed in the esterification and unconverted raw materials. Raw materials are understood to refer to alcohols and maleic anhydride or maleic acid that have not yet fully reacted in the esterification. The alcohol used for esterification is preferably selected from methanol and ethanol, wherein methanol is particularly preferred. Secondary components are selected from maleic monoesters, fumaric diesters, fumaric monoesters and / or mixtures of these esters. The mixture of maleic diesters and secondary components and unreacted raw materials constitutes the "acidic" maleate stream in step a) of the inventive method.

[0022] The acid-containing maleic ester stream used for hydrogenation has a residual acid content in the range of 0.001-2.2% by weight, preferably 0.01-1.6% by weight, particularly preferably 0.05-1.1% by weight, based on the proportion of maleic diesters. Most of the residual acid is attributable to the maleic monoesters. Based on the proportion of maleic diesters, this proportion of maleic monoesters is in the range of 0.001-2% by weight, preferably 0.01-1.5% by weight, particularly preferably 0.05-1% by weight.

[0023] In addition to the maleic acid diesters and monoesters, fumaric acid diesters / monoesters thereof may also be present in an amount of less than 0.5% by weight, preferably less than 0.2% by weight, based on the maleic acid diesters.

[0024] A portion of the acidic maleate stream can also be an incompletely reacted alcohol selected from methanol and ethanol for the esterification, preferably methanol. The proportion of alcohol used is 0.001-20% by weight, preferably 0.05-15% by weight, based on the maleic diester in the acidic maleate stream.

[0025] In an advantageous embodiment of the process according to the invention, the acidic maleate stream from step a) of the process according to the invention may also comprise other compounds. For the production of primarily BDO it is advantageous to also mix the acidic maleate stream with a mixture of succinate and GBL before the hydrogenation. In the process according to the invention, this mixture is produced after step b), since this is one of the mixtures collected after distillation of the hydrogenation output. It is advantageous to return this mixture or part of this mixture as a recycle stream to step a) of the process according to the invention.

[0026] The content of succinic esters—to be understood as including both succinic diesters and succinic monoesters—and GBL recycled to the process according to the invention in step a) is in the range from 0 to 50% by weight, preferably up to 20% by weight, particularly preferably up to 10% by weight, based on the maleic diester.

[0027] If GBL is not intended or only incompletely intended as a commercial product in the process according to the invention, it is advantageous to recycle it. In non-inventive hydrogenation processes in which the catalyst fixed bed does not have a noble metal-containing catalyst layer, the recycle of GBL leads to the formation of undesirable polymer deposits, which do not occur in the process according to the invention when recycled.

[0028] In a further advantageous embodiment of the process of the invention, the acidic maleate stream from step a) also comprises THF. The THF in the hydrogenation feed is derived from the recycle gas. Therefore, the proportion of THF in the acidic maleate stream is in the range of 0.5 to 20% by weight, preferably 1 to 15% by weight, particularly preferably 1 to 10% by weight, based on the organic stream.

[0029] In the process of the present invention, the acidic maleate stream is evaporated with steam before being introduced into the catalyst bed. The hydrogenation requires a large amount of hydrogen. The molar ratio of the maleic diester to the start of the catalyst is 20:1-600:1, preferably 50:1-500:1, and particularly preferably 75:1-400:1. Since a large amount of hydrogen is required to evaporate organic components and also to absorb hydrogenation heat / control the temperature in the corresponding catalyst layer, it is advantageous to recycle hydrogen as much as possible after the reaction occurs and the product condenses. The hydrogen that has been chemically consumed and discharged via the waste gas is continuously supplemented-preferably after the product condenses, preferably supplemented upstream of the reactor. The proportion of discharged hydrogen is based on the proportion of newly supplied hydrogen in the range of 1-50mol%, preferably 3-40mol%, and particularly preferably 5-20mol%.

[0030] The purity of the fresh hydrogen used is preferably >95% by volume, particularly preferably >99% by volume.

[0031] The hydrogenation of the acidic maleate stream is then carried out in a fixed bed in a single reactor. The hydrogenation is carried out in a tubular or axial reactor having a length in the range of 1 to 30 m, preferably 2 to 25 m.

[0032] In the process of the present invention, the fixed bed of the reactor is composed of at least two catalyst layers. Of the two layers, the layer that is first passed through in the hydrogenation direction is the noble metal catalyst layer. The layer that is subsequently passed through and immediately after the noble metal catalyst layer is the Zn-free Cu catalyst layer.

[0033] The noble metal-containing catalyst layer can use any noble metal-containing catalyst which, under normal hydrogenation conditions, allows hydrogenation of the acidic maleate stream to give BDO only in small amounts, if at all. Small amounts are to be understood here as meaning amounts of BDO in the range from 0 to ≤5000 ppm, preferably ≤1000 ppm, very particularly preferably ≤250 ppm, based on the entire mixture of organic compounds in the product stream leaving the noble metal-containing catalyst layer. Further catalyst layers can be present in the fixed bed of the process according to the invention downstream of the Zn-free Cu-containing catalyst layer.

[0034] The noble metal-containing catalyst layer used in the process of the present invention for hydrogenating the acidic maleate stream to provide BDO in an amount of 0 to ≤ 0.5 wt.-%, based on the total amount of organic compounds in the product stream leaving the noble metal-containing catalyst layer, is a reduced noble metal-containing catalyst layer and preferably comprises a noble metal selected from Pd and / or Pt.

[0035] The noble metal-containing catalyst layer particularly preferably comprises Pd. The noble metal has preferably been applied to a support. The latter can be oxidic in nature or made of carbon. The oxide support is selected from TiO 2 、ZrO 2 、Al 2 O 3 、SiO 2 , mixtures of these oxides, and composites such as clay or zeolite. Preferred noble metal-containing catalysts are selected from ZrO 2 Loaded with Pd, Al 2 O 3 Pd, TiO 2 Pd is loaded on Pd or activated carbon is loaded on Pd. ZrO is particularly preferred 2 Pd loaded, activated carbon loaded with Pd or Al 2 O 3 The noble metal content is in the range of 0.001-15 wt %, preferably 0.01-10 wt %, particularly preferably 0.1-5 wt %, based on the total weight of the noble metal-containing catalyst.

[0036] In the at least two catalyst layers used in the method of the invention, the layer that is passed subsequently and immediately after the noble metal-containing catalyst layer is a Zn-free Cu-containing catalyst layer. The first noble metal-containing catalyst layer is a significantly smaller proportion based on the total proportion of catalyst required for the hydrogenation. Based on the total volume of the catalyst, the proportion of the noble metal-containing catalyst layer is generally 0.1-20% by volume, preferably 1-15% by volume, particularly preferably 1.5-10% by volume.

[0037] Monitoring and optionally regulating the temperature upstream of the noble metal-containing catalyst layer is necessary to minimize the formation of by-products during the hydrogenation process in the noble metal-containing catalyst layer. The noble metal-containing layer hydrogenates the double bonds of maleic diester to provide succinic diester. This is an exothermic reaction and therefore releases heat. If the temperature in the fixed bed is too high overall, destructive secondary components such as n-butanol are formed to a greater extent. However, if the temperature in the fixed bed is too low, polymer deposition on the catalyst will occur and the service life of the catalyst will therefore be shorter. Therefore, the purpose of the method of the present invention is to control the heat in the noble metal-containing catalyst layer so that at the end of the entire hydrogenation, after passing through all catalyst layers, by-product formation does not increase and the catalyst layer is free from polymer deposition. This can be achieved when the gas flow temperature downstream of the noble metal-containing layer is at least 5°C, preferably at least 6°C, and particularly preferably at least 7°C higher than the inlet temperature entering the noble metal-containing catalyst layer.

[0038] The hydrogenation is preferably operated such that after passing through the noble metal-containing catalyst layer an adiabatic temperature rise of at least 33% has been achieved along the catalyst bed. Particularly preferably, after 10% of the catalyst length an adiabatic temperature rise of at least 40% has been achieved along the catalyst bed.

[0039] This means, for example, that under the conditions stated and with a total temperature rise of 21° C. in the entire fixed bed, the temperature after the noble metal-containing catalyst layer has already increased by 7° C.

[0040] The combination of temperature increase within the noble metal-containing catalyst layer and the absence of BDO formation avoids the formation of polymer deposition and significantly increases the uptime of the hydrogenation, while also reducing the formation of by-products by making it possible to reduce the inlet temperature into the catalyst layer relative to a catalyst layer that does not contain noble metals. This also reduces the outlet temperature of the reactor. It should be understood that the temperature reduction on the catalyst bed overall results in fewer unwanted secondary reactions and significantly extended catalyst uptime.

[0041] The conversion of the CC double bonds to be hydrogenated in the noble metal-containing catalyst layer is at least 50%, preferably at least 80%, particularly preferably at least 95%, very particularly preferably at least 99%, based on the total amount of double bonds to be hydrogenated in the acidic maleic ester stream.

[0042] The Zn-free Cu catalyst layer follows the noble metal catalyst layer. The Zn-free Cu catalyst is understood to refer to all Cu catalysts that do not contain Zn and / or contain Zn compounds. The Zn-free catalyst is understood to refer to those in which no Zn component is added during the production method. Trace amounts of Zn cannot be excluded if they exist as associated impurities in the raw material. This means that Zn is present in the catalyst at most in a few ppm, preferably less than 1000ppm, particularly preferably less than 100ppm, and very particularly preferably less than 50ppm.

[0043] The subsequent copper catalyst layer must be free of Zn, since Zn-containing catalysts are alkaline and form Zn salts with monoesters of maleic acid and / or succinic acid, which settle as deposits on the catalyst layer and can lead to coking of the catalyst layer and deactivation or destruction of the catalyst at the temperatures mentioned. This would result in the catalyst used having only a short service life and the conversion and selectivity of the catalyst would also decrease.

[0044] The Zn-free Cu catalyst layer is understood to mean all Zn-free Cu catalysts capable of hydrogenating the acidic maleate stream in the presence of hydrogen to provide a mixture of BDO, GBL and THF. The Zn-free Cu catalyst layer may consist of one or more Zn-free Cu catalysts which may also be stacked together. Such catalyst layers are, for example, those described in WO / 0144148. The Zn-free Cu catalyst is preferably selected from reduced copper chromite catalysts and reduced copper oxide catalysts which may also contain other metals, metalloids and / or transition metals as promoters, wherein Zn and its compounds are excluded from the promoter. The Zn-free Cu catalyst may have been applied to various supports and may contain a wide variety of promoters in addition to Zn. A preferred support is selected from TiO 2 、ZrO 2 、Al 2 O 3 、SiO 2 , mixtures of these oxides and composites such as clay or zeolites and also carbon. Particularly preferred supports are TiO 2 、ZrO 2 、Al 2 O 3 and carbon. Preferred promoters are selected from Cr, Al, Mn, Si and La. The promoter is preferably present in the Zn-free Cu catalyst in oxide form and is therefore selected from Al 2 O 3 、SiO 2 , Mn 2 O 3 ,La 2 O 3 and Cr 2 O3 Particularly preferred Zn-free Cu catalysts are reduced copper chromite and reduced copper oxide catalysts, which preferably contain a promoter selected from Cr and Mn. Such catalysts are described, for example, in EP-A 2004590 and the literature cited therein.

[0045] The Zn-free Cu catalyst layer is the largest catalyst layer based on the entire fixed bed. Preferably, the Zn-free Cu catalyst layer accounts for 80-99.9% volume of the total volume of the fixed bed. Other catalyst layers that also contain Zn can also be installed downstream of the Zn-free Cu catalyst layer.

[0046] The Zn-free Cu catalyst layer and any additional catalyst layers used after the noble metal-containing layer serve to produce the desired product spectrum. If it is primarily desired to produce BDO, the catalyst layer has as few acidic centers as possible. If the focus is on THF, the catalyst in the Zn-free Cu catalyst layer and any additional catalyst layers has acidic centers, such as acidic OH groups of the support material, which convert the BDO formed as an intermediate into THF. Depending on the pressure and temperature, GBL is always also formed in equilibrium with BDO. If the GBL is not desired as a commercial product or is only not completely desired, it can be recycled into the hydrogenation process. It was surprisingly found that the noble metal-containing catalyst layer again ensures that no polymer formation occurs, because the recycling of GBL without the noble metal-containing catalyst layer seems to favor the formation of polymers.

[0047] The process of the present invention is usually carried out at 8-80 bar, preferably 20-70 bar, particularly preferably 40-65 bar. The hydrogenation temperature of the present invention is 150-250°C, preferably 155-230°C, particularly preferably 160-210°C.

[0048] After the hydrogenation, the product-containing gas stream is cooled and the resulting condensate is sent for further workup.

[0049] The gas stream, which comprises predominantly hydrogen, is recycled predominantly via a cycle gas compressor. At least alcohol and THF are also present.

[0050] The condensed product output comprises very significant weight proportions (>98% by weight) of THF, GBL, BDO, the esterification alcohol (methanol when methyl esters are used), water and also the undesired products butanol, methyl butyl ether, succinic diesters and other succinic esters, 4-hydroxybutyraldehyde and its acetals with the esterification alcohol and BDO.

[0051] The mixture is separated into the desired products in a manner known per se by distillation in one or preferably more columns. This gives at least three mixtures. Mixture 1 comprises methyl butyl ether and THF, mixture 2 comprises succinic diesters and GBL and mixture 3 comprises 4-hydroxybutyraldehyde and its acetal—very predominantly present as BDO acetal due to the reaction in the column—and BDO. These mixtures can be separated into fuel products in further columns, wherein the separation of methyl butyl ether and BDO acetal in the case of mixtures 1 and 3 can only be achieved with difficulty. Example

[0052] The process of the invention is illustrated in more detail but not limited by reference to the following examples. The reported compositions of starting materials and products are GC area % values ​​determined by GC chromatography via an FI detector. The water content is therefore not taken into account.

[0053] Dimethyl maleate is produced by esterifying maleic anhydride with methanol, which is catalyzed by an acidic ion exchanger. The composition comprises 94% by area of ​​dimethyl maleate, 3% by area of ​​methanol, 0.1% by area of ​​monomethyl fumarate and 0.8% by area of ​​monomethyl maleate as well as other quantitatively minor compounds. For this purpose, methanol and THF are introduced in an amount of 1-10% by weight via the recycle gas in each case.

[0054] The reactor used is a metal tube with a length of 4 m, a volume of 3.8 liters and a continuously adjustable thermocouple for measuring the temperature anywhere in the reactor. The reactor tube is surrounded by two heating zones (about 200° C.) through which thermostatic heat transfer oil (Marlotherm) passes. The upper heating zone surrounds ¾ of the reactor and the lower heating zone surrounds ¼ of the reactor. This ensures that hardly any heat losses occur in the reactor. Hydrogen (circulating gas and fresh hydrogen) is used in the evaporator to bring the liquid feed stream to the desired reactor inlet temperature. The direction of hydrogenation is from top to bottom in the reactor tube. 200 ml of inert filler (Raschig rings No. 3, 3 mm×3 mm—or Al2O3 with a diameter of 5 mm) are introduced upstream of the actual catalyst in each case. 2 O 3 200 ml of No. 3 Raschig rings (Al 2 O rings can be used alternatively) are placed downstream of the catalyst. 2 O 3 sphere), on which the lowest catalyst layer in each case rests.

[0055] Before using it in the reactor, the catalyst is activated with a mixture of nitrogen and hydrogen at a temperature of up to 170° C. and standard pressure. The reactor output is cooled to about 5° C. using a heat exchanger and the gas phase is recycled via the evaporator using a recycle gas compressor. The discharge rate of the gas, which is essentially hydrogen, is 20 standard liters per hour. The reactor system is operated under pressure control, i.e. always replenishing the same amount of hydrogen as is chemically consumed and discharged. The condensed portion is discharged from the gas / liquid separator under liquid level control. The liquid output is collected and analyzed.

[0056] Catalyst used (per meter of reactor: about 2 liters of catalyst bed including inert material):

[0057] Catalyst 1:

[0058] ·ZrO 2 Pd-loaded (monoclinic / tetragonal phase), containing 0.5 wt% Pd (ring: 3 mm × 3 mm × 7 mm, bulk density: about 1.4 kg / m 3 )

[0059] Catalyst 2 (not of the present invention):

[0060] 5×3mm sheet, bulk density: about 1.2kg / m 3 , from BASF, containing 38 wt.% ZnO, 38 wt.% CuO, 20 wt.% Al 2 O 3 , 4 wt% C

[0061] Catalyst 3:

[0062] 1 / 8" tablets, bulk density: about 1.6, from BASF, containing 50 wt% CuO, 46 wt% Cr 2 O 3 , 4 wt% C

[0063] Catalyst 4:

[0064] 1 / 8" tablets, bulk density: 1.3, from BASF, at 30 wt% Al 2 O 3 The above comprises 58 wt% CuO and 12 wt% MnO 2

[0065] Catalyst 5:

[0066] 1 / 8" tablets, bulk density: about 1, from BASF, containing 13 wt% CuO, 84 wt% Al 2 O 3 , 3 wt% C

[0067] Inert material used: No. 3 Raschig ring (glass: 3mm×3mm)

[0068] Comparative Example 1:

[0069] The reactor was filled with 200 ml of No. 3 Raschig rings as the bottom layer, 100 ml of catalyst 5 as the bottom catalyst layer, 1.1 liters of catalyst 4, 360 ml of catalyst 2, 150 ml of catalyst 1 and another 200 ml of No. 3 Raschig rings as the top layer. The reactor was then brought to a hydrogen pressure of 57 bar over 44 hours and the feed rate of acidic maleate was brought to 396 g / h. The Pd / ZrO 2 The temperature at the beginning of the catalyst bed of the catalyst was set at 175° C. A 250 mol cycle gas / mol acid-containing maleate stream was produced. After an operating time of 44 hours, the reactor stabilized and the outlet temperature was 195° C. The temperature rise downstream of the Pd catalyst bed was 10-15° C. relative to the inlet temperature. The hydrogenation output comprised 32-34 area % methanol, 39-41 area % BDO, 6-7 area % GBL, 15-17 area % THF, 1.1-1.3 area % n-butanol, 0.08 area % dimethyl succinate and a large number of other components.

[0070] Comparative Example 2:

[0071] Similar to Comparative Example 1, except that the inlet temperature was reduced to 165° C. The reactor outlet temperature was 185° C. The hydrogenation output contained 31-33 area % methanol, 50-52 area % BDO, 6-8 area % GBL, 5-7 area % THF, 0.8-0.9 area % n-butanol, 0.4-0.6 area % dimethyl succinate and a large amount of other components. Obviously, due to the reduction in the inlet temperature compared to Comparative Example 1, the content of impurity n-butanol decreased and the content of impurity dimethyl succinate increased.

[0072] Comparative Example 3:

[0073] Similar to comparative example 1, except that the reactor inlet temperature was reduced from 175° C. to 170° C. after 6 days, to 165° C. after a further 13 days and reset to 175° C. after a further 4 days and maintained for 2 days. After 10 days of operation the pressure drop over the reactor increased by more than 5% from the initial value of 32 mbar and after a further 10 days was 20% above the initial value, with the result that after a further 14 days the reactor was shut down. Disassembly of the catalyst after cooling and decompression showed adhesions in the catalyst layer, which indicated that the Zn-containing catalyst 2 was damaged. At the end of the test (after 24 days) the hydrogenation output comprised 31-33 area % methanol, 45-47 area % BDO, 7-8 area % GBL, 11-12 area % THF, 1.1 area % n-butanol, 0.12-0.14 area % dimethyl succinate and a large number of other components.

[0074] Comparative Example 4:

[0075] 150 ml of catalyst 1 was installed between the respective layers of 200 ml No. 3 Raschig rings above and below it, placed under a hydrogen atmosphere and heated to 175° C. The hydrogenation was carried out at 57 bar and 1375 g / h of circulating gas, 45 g / h of fresh hydrogen and 390-400 g / h of acid-containing maleate at a reaction inlet temperature of 175° C., then 185° C. and finally 165° C. The maximum temperature rise of the catalyst bed was 15° C. After the start-up effects had subsided, the hydrogenation output comprised 93-94 area % of dimethyl succinate, 3-4 area % of methanol, 0.4-0.5 area % of GBL, 0.4 area % of THF, 0.02-0.03 area % of BDO and other quantitatively minor components at 175° C., 93-94 area % of dimethyl succinate, 3-4 area % of methanol, 0.4-0.5 area % of GBL, 0.4 area % of THF, 0.03 area % of BDO and other quantitatively minor components at 185° C. and 93-94 area % of dimethyl succinate, 3-4 area % of methanol, 0.3 area % of GBL, 0.4 area % of THF, 0.02 area % of BDO and other quantitatively minor components at 165° C. It is clear that catalyst 1 hydrogenates almost exclusively the double bonds of the acid-containing maleate.

[0076] Embodiment 1:

[0077] Similar to Comparative Example 1, except that instead of the Zn-containing catalyst 2 below the catalyst 1, a layer of 360 ml of catalyst 3 was installed, on which a third layer of 1100 ml of catalyst 4 was arranged, and on which a fourth layer of 100 ml of catalyst 5 was arranged. The hydrogenation output contained 30-32 area % methanol, 55-58 area % BDO, 9 area % GBL, 2 area % THF, 0.9 area % n-butanol, 0.02 area % dimethyl succinate and a large amount of other components. It is clear that the content of impurity n-butanol is reduced relative to Comparative Example 1 and the content of impurity dimethyl succinate is lower than that in Comparative Examples 1 and 2.

[0078] Embodiment 2:

[0079] Similar to Comparative Example 2, except that instead of the Zn-containing catalyst 2 below the catalyst 1, a layer of 360 ml of catalyst 3 was installed, on which a third layer of 1100 ml of catalyst 4 was arranged, and on which a fourth layer of 100 ml of catalyst 5 was arranged. The hydrogenation output contained 31-33 area % methanol, 57-59 area % BDO, 6-7 area % GBL, 1-2 area % THF, 0.7 area % n-butanol, 0.02 area % dimethyl succinate and a large number of other components. Obviously, due to the lower inlet temperature compared with Comparative Example 1, the content of impurity n-butanol was even more significantly reduced than in Comparative Example 2 and the content of impurity dimethyl succinate was lower than in Comparative Examples 1 and 2.

[0080] Embodiment 3:

[0081] Similar to comparative example 3, except that instead of the Zn-containing catalyst 2 below the catalyst 1, a layer of 360 ml of catalyst 3 was installed, on which a third layer of 1100 ml of catalyst 4 was arranged, on which a fourth layer of 100 ml of catalyst 5 was arranged and hydrogenation was carried out over a period of 38 days, wherein the reaction temperature was reduced from 175° C. to 170° C. after 7 days, to 165° C. after another 2 days and to 160° C. after another 6 days, and then increased to 165° C. after another 5 days, to 170° C. after another 7 days and to 175° C. after another 4 days. During the adjustment of the inlet temperature between 165 and 175° C., 10 area % of GBL was added to the feed on the 5th day for 11 days. As expected, it was found that the product mixture contained 27-29 area % less methanol at an otherwise similar composition during this period. The pressure drop over the reactor did not increase significantly during the entire test. The catalysts could be disassembled without problems and no adhesion or deposition was observed in the catalyst bed. The temperature rise downstream of the catalyst bed 1 at each temperature setting was 10-15° C. relative to the inlet temperature.

[0082] At the end of the run the hydrogenation output comprised 31-33 area % methanol, 55-58 area % BDO, 9 area % GBL, 2 area % THF, 1 area % n-butanol, 0.02 area % dimethyl succinate and a large number of further components.

[0083] It is clear that the process of the invention makes it possible to avoid the deposition and adhesion observed in Comparative Example 3, thereby allowing a longer catalyst uptime.

[0084] Furthermore, in order to evaluate different catalysts for the uppermost Pd and / or Pt-containing catalyst layer and various support materials for these catalysts, screening tests were carried out in a tubular reactor having a volume of 0.273 liters, an inner diameter of 20.5 mm and continuously operated at 10 bar and a temperature of 175° C., wherein the test reaction carried out was the hydrogenation of dimethyl maleate to provide dimethyl succinate and the procedure was as follows:

[0085] Dimethyl maleate (DMM for short) as a raw material is fed from a storage container into an evaporator filled with glass rings by means of a metering pump, where it is mixed with hydrogen and evaporated at a temperature of about 210°C. The hydrogen mixed with the raw material is fed via a micro-heating pipe into a reactor filled with a catalyst layer between two layers of glass rings and heated to 175°C. Downstream of the reactor, the reaction mixture is passed through a double-tube heat exchanger cooled by a low temperature thermostat via a micro-heating pipe, where the product is condensed. Subsequently, a gas / liquid separation of the product stream is carried out in a separator cooled by the same low temperature thermostat. The waste gas is passed through a cold trap cooled by cooling water to condense the residual liquid component. The unconsumed hydrogen is returned to the evaporator as a circulating gas (600l / h). The consumed hydrogen is supplemented (usually 60-70Nl / h). The liquid product stream is discharged from the separator into a collecting container by means of a leveler and its composition is analyzed by gas chromatography.

[0086] The amount of catalyst, in each case 100 ml, was placed in the reactor on a first layer of 84-86 ml No. 5 Raschig rings and subsequently covered with a further layer of 126-129 ml of the same glass rings.

[0087] Catalysts tested:

[0088] Catalyst 6:

[0089] ·ZrO 2 Pd-loaded (monoclinic / tetragonal phase), containing 0.3 wt% Pd (ring: 3 mm × 3 mm × 7 mm, bulk density: about 1.4 kg / m 3 )

[0090] Catalyst 7:

[0091] Carbon-supported Pd, containing 1 wt% Pd (pellet: 2.3-4.8 mm mesh size, bulk density: about 500 kg / m 3 )

[0092] Catalyst 8:

[0093] Carbon-supported Pd, containing 0.8 wt% Pd (3 mm extrudate, bulk density: about 450 kg / m 3 ) Catalyst 9:

[0094] Carbon-supported Pd, containing 0.5 wt% Pd (pellet: 2.3-4.8 mm mesh size, bulk density: about 450 kg / m 3 )

[0095] Catalyst 10:

[0096] ·g- / q-Al 2 O 3Pd-loaded, containing 0.5 wt% Pd (1.6 mm balls, bulk density: about 800 kg / m 3 ) Catalyst 11:

[0097] ·g- / q-Al 2 O 3 Pd-loaded, containing 0.5 wt% Pd (3.2 mm balls, bulk density: about 800 kg / m 3 ) Catalyst 12:

[0098] ·a-Al 2 O 3 Pd loaded, containing 0.3 wt% Pd (4 mm extrudate, bulk density: about 1050 kg / m 3 ) Catalyst 13:

[0099] ·a-Al 2 O 3 Pd loaded, containing 0.5 wt% Pd (4 mm extrudate, bulk density: about 1050 kg / m 3 ) Catalyst 14:

[0100] ·ZrO 2 0.5 wt% Pd and 0.5 wt% Pt (monoclinic / tetragonal phase, ring: 3 mm × 3 mm × 7 mm, bulk density: about 1.4 kg / m 3 )

[0101] Catalyst 15:

[0102] ·ZrO 2 0.4 wt% Pt (monoclinic / tetragonal phase, ring: 3mm×3mm×7mm, bulk density: about 1.4 kg / m 3 )

[0103] The table below shows the catalysts tested and the dimethyl succinate content achieved therewith in stable continuous operation ("cont. op."), reported as GC area %. Catalyst 1 described above was also tested as a comparison.

[0104] Example No. catalyst GC area % of dimethyl succinate after 2d cont.op. 4 Catalyst 1 52 5 Catalyst 6 50 6 Catalyst 7 90 7 Catalyst 8 80 8 Catalyst 9 85 9 Catalyst 10 88 10 Catalyst 11 83 11 Catalyst 12 52 12 Catalyst 13 49 13 Catalyst 14 57 14 Catalyst 15 46

[0105] Examples 4-14 show that various catalysts on various support materials achieve similarly high conversions as or higher than Catalyst 1. The reaction conditions precluded achievement of complete conversion, which is actually advantageous for providing visible differences for comparative screening. It is believed that the high levels of dimethyl succinate, greater than 80%, achieved in some cases can be attributed to the favorable distribution of the noble metal on the support.

Claims

1. A method for producing 1,4-butanediol, γ-butyrolactone and tetrahydrofuran, mainly producing 1,4-butanediol, comprising the following steps: a) hydrogenating an acid-containing maleate stream with hydrogen in the gas phase in a single reactor, wherein the acid-containing maleate stream is hydrogenated over a catalyst fixed bed comprising at least two different and directly adjacent catalyst layers, the catalyst layer passed through first in the operating direction of the hydrogenation is a noble metal-containing catalyst layer comprising a noble metal selected from the group consisting of Pd and Pt and the catalyst layer passed through next to the noble metal-containing catalyst layer is a Zn-free Cu-containing catalyst layer, wherein the proportion of 1,4-butanediol after passing through the noble metal-containing catalyst layer is in the range from 0 to 0.5% by weight, based on the total amount of organic compounds in the product stream and the temperature of the gas stream after passing through the noble metal-containing catalyst layer and before entering the Zn-free Cu-containing catalyst layer is increased by at least 5° C. relative to the inlet temperature after entering the catalyst fixed bed and subsequently being fully hydrogenated in the last catalyst layer of the fixed bed, b) separating the resulting condensed mixture by distillation in one or more columns into three mixtures, wherein Each mixture contains one of the three products selected from 1,4-butanediol, γ-butyrolactone and tetrahydrofuran in excess, and the γ-butyrolactone-containing mixture of γ-butyrolactone and succinic ester or a portion of the mixture is returned to step a) as a recycle stream.

2. The process according to claim 1, wherein the acid maleic acid ester stream has a maleic acid monoester content in the range from 0.001 to 2% by weight, based on the proportion of maleic acid diester.

3. The process according to claim 1, wherein the noble metal in the noble metal-containing catalyst layer has been applied to a support selected from the group consisting of carbon, clay, zeolite, TiO2, ZrO2, Al2O3, SiO2 and mixtures of these components.

4. The method according to claim 2, wherein the noble metal in the noble metal-containing catalyst layer has been applied to a carrier selected from the group consisting of carbon, clay, zeolite, TiO2, ZrO2, Al2O3, SiO2 and mixtures of these components. 5 . The method according to claim 1 , wherein the noble metal-containing catalyst layer comprises palladium as the noble metal in a range of 0.001 to 15 wt % based on the total weight of the catalyst.

6. The method according to claim 2, wherein the noble metal-containing catalyst layer contains palladium as the noble metal in a range of 0.001 to 15 wt% based on the total weight of the catalyst.

7. The method according to claim 3, wherein the noble metal-containing catalyst layer contains palladium as the noble metal in a range of 0.001 to 15 wt% based on the total weight of the catalyst.

8. The method according to claim 4, wherein the noble metal-containing catalyst layer contains palladium as the noble metal in a range of 0.001 to 15 wt% based on the total weight of the catalyst.

9. The process according to any one of claims 1 to 8, wherein the temperature of the gas stream is increased in the noble metal-containing catalyst layer such that after passing through the noble metal-containing catalyst layer an adiabatic temperature rise of at least 33% has been achieved along the catalyst fixed bed.

10. The process according to any one of claims 1 to 8, wherein the hydrogenation is carried out at a temperature in the range of 150 to 250°C and a pressure in the range of 8 to 80 bar.

11. The process according to claim 9, wherein the hydrogenation is carried out at a temperature in the range of 150-250°C and a pressure in the range of 8-80 bar.

12. The method according to any one of claims 1 to 8, wherein the Zn-free Cu catalyst layer following the noble metal-containing catalyst layer is selected from reduced copper chromite catalysts and reduced copper oxide catalysts which may also contain other metals, metalloids and / or transition metals as promoters, wherein Zn and its compounds are excluded from the promoters.

13. The method according to claim 11, wherein the Zn-free Cu catalyst layer following the noble metal-containing catalyst layer is selected from a reduced copper chromite catalyst and a reduced copper oxide catalyst which may also contain other metals, metalloids and / or transition metals as promoters, wherein Zn and its compounds are excluded from the promoters.

14. The process according to any one of claims 1 to 8, wherein the Zn-free Cu catalyst comprises further components selected from the group consisting of Al2O3, SiO2, TiO2, ZrO2, Mn2O3, Cr2O3 and La2O3.

15. The method according to claim 13, wherein the Zn-free Cu catalyst comprises other components selected from the group consisting of Al2O3, SiO2, TiO2, ZrO2, Mn2O3, Cr2O3 and La2O3.

16. The process according to any one of claims 1 to 8, wherein the acid maleate-containing stream further comprises tetrahydrofuran in the range from 0.1 to 20% by weight, based on the maleic diester content.

17. The process according to claim 15, wherein the acid maleate-containing stream further comprises tetrahydrofuran in the range from 0.1 to 20% by weight, based on the maleic diester content.

18. The process according to claim 1, wherein the acid-containing maleic acid ester stream further comprises a total content of up to 50% by weight of gamma-butyrolactone and succinic acid diesters, based on the maleic acid diester content.

19. The process according to claim 17, wherein the acid maleate-containing stream further comprises gamma-butyrolactone and succinic diesters in a total content of up to 50% by weight, based on the maleic diester content.

20. The method according to any one of claims 1 to 8, wherein the catalyst-containing layer is disposed on a layer of inert material and is covered by another layer of inert material.

21. The method of claim 19, wherein the catalyst-containing layer is disposed on a layer of an inert material and is covered by another layer of an inert material.

22. The method according to any one of claims 1 to 8, wherein the catalyst-containing layer is disposed on a layer of inert material.

23. The method of claim 19, wherein the catalyst-containing layer is disposed on a layer of inert material.

24. The method according to any one of claims 1 to 8, wherein the catalyst-containing layer is covered by a layer of inert material.

25. The method of claim 19, wherein the catalyst-containing layer is covered by a layer of an inert material.