CN114829347A

CN114829347A - Method for producing 1, 4-butanediol, gamma-butyrolactone and tetrahydrofuran in the gas phase while avoiding polymer deposits

Info

Publication number: CN114829347A
Application number: CN202080085262.6A
Authority: CN
Inventors: R·萍克斯; D·罗德瓦尔德; S·S·林克; T·海德曼; J·维格尼
Original assignee: BASF SE
Current assignee: BASF SE
Priority date: 2019-12-10
Filing date: 2020-11-30
Publication date: 2022-07-29
Also published as: KR20220110820A; EP4073052A1; WO2021115813A1

Abstract

The invention relates to an improved process for the production of 1, 4-butanediol, tetrahydrofuran and gamma-butyrolactone, in which an acid maleate-containing stream is hydrogenated in the presence of hydrogen at specific, different temperature levels in a catalyst layer in a single reactor comprising at least two different adjacent 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 subjected to a post-distillation treatment in order to prevent polymer deposition and catalyst ageing and damage in the reactor.

Description

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

The invention relates to an improved process for the production of 1, 4-butanediol, tetrahydrofuran and gamma-butyrolactone, in which an acid maleate-containing stream is hydrogenated in the presence of hydrogen at specific, different temperature levels in a catalyst layer in a single reactor comprising at least two different adjacent catalyst layers, wherein the first layer is a noble metal-containing catalyst layer followed by a second Zn-free Cu catalyst layer, and the post-distillation treatment is subsequently carried out in such a way that polymer deposition and ageing and damage of the catalyst in the reactor are prevented.

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), gamma-butyrolactone (GBL) and the corresponding alcohols of the esters is known and has been used on an industrial scale. This process is described, for example, in WO-A91/01960. However, a disadvantage of the process described in WO-a91/01960 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, the latter leading to a significant increase in the pressure drop, which in turn makes it necessary to interrupt the hydrogenation and replace the catalyst.

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 an alumina support, to give succinic diesters, followed by hydrogenation over a copper catalyst to give mainly GBL and THF. The GBL is hydrogenated in another second reactor to provide BDO. The hydrogenated feed stream comprises maleic acid diester and the recycle stream comprises BDO and succinic acid diester. 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 oligomeric esters and polyester deposits formed by the reaction of BDO with ester may form over time, at least in the evaporator, when BDO is recycled to the feed. Furthermore, yield losses must be expected, since BDO-maleate or-succinate are high-boiling compounds and have to be discharged during the evaporation.

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 in part succinic diesters. This can be carried out in separate reactors or in the same reactor, with subsequent further hydrogenation after complete evaporation to provide THF and optionally BDO and GBL. This may have evaporated a part of the ester mixture, especially if the hydrogenation of the double bonds is carried out in the same reactor as the reactor in which the further hydrogenation is subsequently carried out. This method, which looks perfect at first sight, has several drawbacks. The use of an additional reactor leads to an increase in cost. Furthermore, the evaporation, even if only partially carried out in the presence of the catalyst, generally causes catalyst attrition, since evaporation occurs preferentially at the highest temperatures. In this case, it therefore occurs in the catalyst pores, with the exothermic hydrogenation of the double bonds leading to the release of heat. Furthermore, evaporation leads to a large increase in volume, which can crack the catalyst shaped body, causing mechanical destruction of the catalyst over time. A further disadvantage of the process described in WO-A2005/058855 is the use of copper-containing catalysts which hydrogenate not only the double bonds of the diester but also further hydrogenate the product to form BDO as described in WO-A2005/058855. However, this has the following disadvantages: 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 in terms of conversion, yield and selectivity and requires replacement.

WO-a 01/44148 describes a process in which an ester which may still contain trace amounts of acid, such as the maleic acid diester, is hydrogenated over 3 different catalyst zones to provide BDO, THF and GBL. Each of these zones is capable of hydrogenating an ester. The first hydrogenation zone may use a Cu catalyst, but Pd catalysts are also mentioned. A disadvantage of this process using catalysts in which the respective one of these catalysts can hydrogenate the esters is, in turn, that the diols formed can form high-boiling compounds with the diesters, in particular in the presence of acids, which can then lead to deposits. Furthermore, the product streams are recovered after each hydrogenation and then introduced into the subsequent hydrogenation zone, which is inconvenient and causes additional costs.

For the hydrogenation of maleic diesters to provide BDO and THF, CN-A101891592 describes two stages each containing a catalyst, wherein the first stage uses a group VIII catalyst and the second stage uses a Cu catalyst. The hydrogenation feed may also comprise 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 Zn-containing Cu catalysts, since this always causes hydrogenation in addition to hydrogenation of the double bond to provide BDO. The disadvantage of this method is that it also does not use an acid maleate-containing 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 basic and preferentially forms the corresponding Zn salt with the monoester of the acid maleate-containing stream. The Zn salts subsequently stick to the catalyst and coking occurs there as a result of the high temperatures. This in turn leads to the catalysts used having only a short service life.

It was therefore an object of the present invention to provide a process for the production of BDO, GBL and THF which makes it possible to obtain the corresponding products in an economical manner and in good yields and purity 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.

This object is achieved by a process for the production of 1, 4-butanediol, gamma-butyrolactone and tetrahydrofuran, which comprises the following steps:

a) hydrogenating an acid maleate-containing stream in the gas phase with hydrogen in a single reactor, wherein the acid maleate-containing stream is hydrogenated on a fixed catalyst bed comprising at least two different and directly adjacent catalyst layers, the catalyst layer which passes through firstly in the direction of travel of the hydrogenation being a noble metal-containing catalyst layer comprising noble metals selected from Pd and Pt and the catalyst layer which passes through subsequently being 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 ℃ relative to the inlet temperature after entering the fixed catalyst bed and subsequent further hydrogenation in the last catalyst layer of the fixed bed,

b) the resulting mixture is separated by distillation in one or more columns into three mixtures, wherein each mixture comprises one of the three products selected from the group consisting of 1, 4-butanediol, gamma-butyrolactone and tetrahydrofuran in excess.

The process of the invention is preferred when the acid-containing maleate 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.

When the noble metal in the noble metal-containing catalyst layer has been applied to a catalyst layer selected from the group consisting of carbon (e.g., activated carbon or burnt nutshell), argillaceous earth (argillaceous earth), zeolite, TiO, and mixtures thereof ₂ 、ZrO ₂ 、Al ₂ O ₃ 、SiO ₂ And mixtures of these components, the process of the invention is preferred.

The method of the present invention is preferable when the noble metal-containing catalyst layer contains palladium as the noble metal in the range of 0.001 to 15% by weight based on the total weight of the catalyst.

The process of 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 rise of at least 33% has been achieved along the fixed catalyst bed after passage through the noble metal-containing catalyst layer.

The process of the invention is preferred when the hydrogenation is carried out at a temperature in the range of 150 ℃ and 250 ℃ and a pressure of 8 to 80 bar.

The method of the present invention is preferable when the Zn-free Cu catalyst layer immediately after the noble metal-containing catalyst layer is selected from the group consisting of a reduced copper chromite catalyst and a reduced copper oxide catalyst which may further contain other metals, metalloids and/or transition metals as promoters, with Zn and its compounds excluded from the promoters.

When the Zn-free Cu catalyst comprises Al ₂ O ₃ 、SiO ₂ 、TiO ₂ 、ZrO ₂ 、Mn ₂ O ₃ 、Cr ₂ O ₃ And La ₂ O ₃ The process of the invention is preferred when other components are used.

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

The process of the present invention is preferred when the acid maleate containing stream also comprises gamma butyrolactone and succinic acid mono-or diester in a total content of up to 50 wt% based on maleic acid diester content.

The acid-containing maleate ester stream to be hydrogenated is produced by known processes, 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 conversion required, the greater the technical complexity. Therefore, there is a great need for a process for producing BDO, GBL and THF, wherein the acid-containing maleate stream resulting from production can be introduced into the further hydrogenation without additional work-up and without incurring losses in terms of yield, selectivity or catalyst lifetime.

The process of the present invention is representative of such a process. The acid-containing maleate ester stream used for the hydrogenation in step a) of the process of the present invention is understood to mean a mixture which comprises not only the maleic acid diester formed as main product in the esterification of maleic anhydride or maleic acid with the corresponding alcohol but also the secondary components formed in the esterification and unconverted starting materials. The starting materials are understood to mean alcohols and maleic anhydride or maleic acid which have not yet reacted completely in the esterification. The alcohol used for the esterification is preferably selected from methanol and ethanol, with methanol being particularly preferred. The secondary component is selected from the group consisting of maleic acid monoesters, fumaric acid diesters, fumaric acid monoesters and/or mixtures of these esters. The mixture of maleic diester and secondary components and unreacted starting materials constitutes the "acidic" maleic ester stream in step a) of the process according to the invention.

The acid maleate-containing stream used for the hydrogenation has a residual acid content in the range from 0.001 to 2.2% by weight, preferably from 0.01 to 1.6% by weight, particularly preferably from 0.05 to 1.1% by weight, based on the proportion of maleic diester. Most of the residual acid can be attributed to the maleic acid monoester. The proportion of the maleic acid monoester is in the range from 0.001 to 2% by weight, preferably from 0.01 to 1.5% by weight, particularly preferably from 0.05 to 1% by weight, based on the proportion of the maleic acid diester.

In addition to the maleic diesters and monoesters, fumaric diesters/monoesters thereof may also be present. The content thereof is less than 0.5% by weight, preferably less than 0.2% by weight, based on the maleic diester.

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

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 further compounds. It is advantageous for the main production of BDO to also mix the acidic maleate stream with a mixture of succinate and GBL prior to hydrogenation. In the process of 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 a part of this mixture as a recycle stream to step a) of the process of the invention.

The succinic acid ester recirculated to the process of the invention in step a), which is to be understood as including both succinic acid diesters and succinic acid monoesters, and the content of GBL, based on the maleic acid diester, is in the range from 0 to 50% by weight, preferably up to 20% by weight, particularly preferably up to 10% by weight.

If GBL is not desired as a commercial product or is only not completely desired in the process of the invention, it is advantageous to recycle it. In hydrogenation processes not according to the invention, in which the fixed catalyst bed does not have a noble metal-containing catalyst layer, the recycling of the GBL leads to the formation of undesirable polymer deposits, which do not occur during recycling in the process according to the invention.

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

In the process of the present invention, the stream of acidic maleate is vaporized with steam prior to introduction into the catalyst bed. The hydrogenation requires a large amount of hydrogen. The molar ratio thereof to maleic diester at the beginning of the catalyst is from 20:1 to 600:1, preferably from 50:1 to 500:1, particularly preferably from 75:1 to 400: 1. Since a large amount of hydrogen is required to vaporize the organic components and also to absorb the heat of hydrogenation/control the temperature in the respective catalyst layers, it is advantageous to recycle hydrogen as much as possible after the reaction has occurred and the products have condensed. The hydrogen which has been chemically consumed and is discharged via the offgas is continuously replenished, preferably after condensation of the products, preferably upstream of the reactor. The proportion of the discharged hydrogen is in the range from 1 to 50 mol%, preferably from 3 to 40 mol%, particularly preferably from 5 to 20 mol%, based on the proportion of the newly supplied hydrogen.

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

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

In the process of the invention the fixed bed of the reactor consists of at least two catalyst layers. Of these two layers, the layer that passes through in the hydrogenation direction first is a noble metal-containing catalyst layer. The layer that subsequently and immediately follows the noble metal-containing catalyst layer is a Cu-containing catalyst layer that does not contain Zn.

The noble metal-containing catalyst layer can use any noble metal-containing catalyst that only allows hydrogenation of the acidic maleate ester stream to provide BDO in small amounts, if any, under normal hydrogenation conditions. Small amounts are understood here to mean amounts of BDO in the range from 0 to 5000ppm, preferably 1000ppm or less, very particularly preferably 250ppm or less, based on the entire mixture of organic compounds in the product stream leaving the layer containing the noble metal catalyst. Further catalyst layers may be present in the fixed bed of the process of the invention downstream of the Zn-free Cu-containing catalyst layer.

The noble metal-containing catalyst layer used in the process of the invention, which hydrogenates the acidic maleate stream to provide BDO in an amount of from 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.

The noble metal-containing catalyst layer particularly preferably contains Pd. The noble metal is preferably already applied to the support. The latter may be oxide in nature or made of carbon. The oxide carrier is selected from TiO ₂ 、ZrO ₂ 、Al ₂ O ₃ 、SiO ₂ Mixtures of these oxides and composites such as argillaceous soils or zeolites. Preferred noble metal-containing catalysts are selected from ZrO ₂ Pd and Al supported ₂ O ₃ Pd and TiO supported ₂ Pd on activated carbon. Particularly preferred is ZrO ₂ Pd on activated carbon, Pd or Al on activated carbon ₂ O ₃ And carrying Pd. The noble metal content is in the range from 0.001 to 15% by weight, preferably from 0.01 to 10% by weight, particularly preferably from 0.1 to 5% by weight, based on the total weight of the noble metal-containing catalyst.

Of the at least two catalyst layers used in the process of the present invention, the layer that subsequently and immediately follows the noble metal-containing catalyst layer is a Zn-free Cu-containing catalyst layer. The first noble metal-containing catalyst layer is in a significantly smaller proportion based on the total proportion of catalyst required for the hydrogenation. The proportion of the noble-metal-containing catalyst layer is generally from 0.1 to 20% by volume, preferably from 1 to 15% by volume, particularly preferably from 1.5 to 10% by volume, based on the total volume of the catalyst.

Monitoring and optionally adjusting the temperature upstream of the noble metal-containing catalyst layer is necessary to minimize the formation of by-products during hydrogenation in the noble metal-containing catalyst layer. The noble metal-containing layer hydrogenates the double bonds of the maleic acid diester to provide a succinic acid diester. This is an exothermic reaction and thus releases heat. If the temperature in the fixed bed as a whole is too high, disruptive 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 occurs and therefore leads to a shorter service life of the catalyst. The object of the process according to the invention is therefore to control the heat in the noble metal-containing catalyst layer so that the formation of by-products after passage through all catalyst layers is not increased at the end of the entire hydrogenation and the catalyst layers are protected from polymer deposits. This can be achieved when the temperature of the gas stream downstream of the noble-metal-containing layer is at least 5 ℃, preferably at least 6 ℃, particularly preferably at least 7 ℃ higher than the inlet temperature into the noble-metal-containing catalyst layer.

The hydrogenation is preferably operated such that an adiabatic temperature rise of at least 33% along the catalyst bed has been achieved after passage through the noble metal-containing catalyst layer. It is particularly preferred that an adiabatic temperature rise of at least 40% has been achieved along the catalyst bed after 10% of the catalyst length.

This means, for example, that under the stated conditions and at a total temperature rise of 21 ℃ in the entire fixed bed, the temperature has increased by 7 ℃ after the noble metal-containing catalyst layer.

The combination of the exotherm within the noble metal-containing catalyst layer and the absence of BDO formation avoids the formation of polymer deposits and significantly increases the uptime of hydrogenation, while also reducing the formation of byproducts by making it possible to lower the inlet temperature into the catalyst layer relative to a noble metal-free catalyst layer. This also reduces the outlet temperature of the reactor. It will be appreciated that this temperature reduction over the catalyst bed results overall in fewer unwanted secondary reactions and a significantly extended catalyst uptime.

The conversion of the C-C 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 maleate stream.

A Zn-free Cu catalyst layer immediately follows the noble metal-containing catalyst layer. A Zn-free Cu catalyst is understood to mean all Cu catalysts which do not comprise Zn and/or Zn-containing compounds. Catalysts which are free of Zn are understood to mean those in which no Zn component is added during the production process. Traces of Zn cannot be excluded if they are present as accompanying impurities in the feed. This means that Zn is present in the catalyst at up to a few ppm, preferably less than 1000ppm, particularly preferably less than 100ppm, very particularly preferably less than 50 ppm.

The subsequent copper catalyst layer must be Zn-free, since Zn-containing catalysts are basic and form Zn salts with monoesters of maleic and/or succinic acid, which settle as deposits on the catalyst layer and at said temperatures may cause coking of the catalyst layer and deactivation or destruction of the catalyst. This results in the catalyst used having only a short service life and the conversion and selectivity of the catalyst also decreases.

A Zn-free Cu catalyst layer is understood to mean all Zn-free Cu catalysts which are 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 be composed of one or more Zn-free Cu catalysts that may also be laminated together. Such catalyst layers are for example those described in WO/0144148. The Zn-free Cu-containing catalyst is preferably selected from reduced copper chromite catalysts and reduced copper oxide catalysts which may also comprise other metals, metalloids and/or transition metals as promoters, with Zn and its compounds excluded from the promoters. The Zn-free Cu catalyst may have been applied on various supports and may contain a wide variety of promoters in addition to Zn. Preferred supports are selected from TiO ₂ 、ZrO ₂ 、Al ₂ O ₃ 、SiO ₂ Mixtures and composites of these oxides such as argillaceous soils or zeolites and also carbon. A particularly preferred support is 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 form of an oxide in the Zn-free Cu catalyst and is therefore selected from Al ₂ O ₃ 、SiO ₂ 、Mn ₂ O ₃ 、La ₂ O ₃ And Cr ₂ O ₃ . Particularly preferred Zn-free Cu catalysts are copper chromite-reducing and copper oxide-reducing catalysts, which preferably comprise a promoter selected from Cr and Mn. Such catalysts are described, for example, in EP-A2004590 and the documents cited therein.

The Zn-free Cu catalyst layer was the largest catalyst layer based on the entire fixed bed. Preferably, the Zn-free Cu catalyst layer accounts for 80 to 99.9 vol% of the total volume of the fixed bed. Other catalyst layers further comprising Zn may also be installed downstream of the Zn-free Cu catalyst layer.

The Zn-free Cu catalyst layer and any other catalyst layers used after the noble-metal-containing layer are used to produce the desired product spectrum. If it is intended to produce predominantly BDO, the catalyst layer has as few acid sites as possible. If the focus is on THF, the catalysts in this Zn-free Cu catalyst layer and any other catalyst layers have acidic centers, for example acidic OH groups of the support material, which convert 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 fully desired, it can be recycled to the hydrogenation process. It has surprisingly been found that the noble metal-containing catalyst layer again ensures that no polymer formation occurs, since the recycling of GBL without the noble metal-containing catalyst layer seems to favour the formation of polymer.

The process of the invention is generally carried out at from 8 to 80 bar, preferably from 20 to 70 bar, particularly preferably from 40 to 65 bar. The hydrogenation temperature of the invention is 150-250 ℃, preferably 155-230 ℃, and particularly preferably 160-210 ℃.

After the hydrogenation the product-containing gas stream is cooled and the resulting condensate is sent for further work-up.

The gas stream containing predominantly hydrogen is recycled predominantly via the recycle gas compressor. At least an alcohol and THF are also present.

The condensed product output contains a very significant weight proportion (> 98% by weight) of THF, GBL, BDO, esterifying alcohol (methanol when methyl esters are used), water, and the undesirable products butanol, methyl butyl ether, succinic acid diester and other succinic acid esters, 4-hydroxybutyraldehyde and its acetal with esterifying alcohol and BDO.

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 3 mixtures. Mixture 1 contains methyl butyl ether and THF, mixture 2 contains succinic diester and GBL and mixture 3 contains 4-hydroxybutanal and its acetal-very notably present as BDO acetal due to the reaction in the column-and BDO. These mixtures can be separated into fuel products in further columns, with the separation of the methylbutyl ether and BDO acetal being achieved only with difficulty in the case of mixtures 1 and 3.

Examples

The process of the present invention is more specifically illustrated but not limited by reference to the following examples. The reported compositions of the feed and product are GC area% values determined by GC chromatography via FI detector. And therefore does not take into account the water content.

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

The reactor used was a metal tube of length 4m, volume 3.8 liters and with a continuously adjustable thermocouple for measuring the temperature anywhere in the reactor. The reactor tube is surrounded by two heating zones (about 200 ℃) through which constant temperature thermal oil (Marlotherm) passes. The upper heating zone surrounds 3/4 of the reactor and the lower heating zone surrounds 1/4 of the reactor. This ensures that hardly any heat loss occurs within the reactor. The liquid feed stream is brought to the desired reactor inlet temperature using hydrogen (recycle gas and fresh hydrogen) in the vaporizer. The hydrogenation direction is from top to bottom in the reactor tube. 200ml of inert packing (Raschig ring 3, 3 mm. times.3 mm-or Al with a diameter of 5 mm) are introduced in each case upstream of the actual catalyst ₂ O ₃ Balls) to make the air flow uniform. 200ml of Raschig ring No. 3 are arranged downstream of the catalyst (again Al can be used alternatively) ₂ O ₃ Balls) against which the lowest catalyst layer rests in each case.

The catalyst is activated with a mixture of nitrogen and hydrogen at a temperature of up to 170 ℃ and standard pressure before being used in the reactor. The reactor output was cooled to about 5 ℃ using a heat exchanger and the gas phase was recycled via the evaporator using a recycle gas compressor. The discharge rate of the substantially hydrogen gas was 20 normal liters/hour. The reactor system is operated under pressure control, i.e. always replenished with the same amount of hydrogen as chemically consumed and vented. The condensed portion is discharged from the gas/liquid separator under liquid level control. The liquid output was collected and analyzed.

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

catalyst 1:

·ZrO ₂ pd support (monoclinic/tetragonal phase) containing 0.5 wt.% Pd (ring: 3 mm. times.3 mm. times.7 mm, bulk density: about 1.4kg/m ³ )

Catalyst 2 (not according to the invention):

5 × 3mm piece, bulk density: about 1.2kg/m ³ From BASF, 38 wt.% ZnO, 38 wt.% CuO, 20 wt.% Al ₂ O ₃ 4% by weight of C

Catalyst 3:

1/8 "sheet, bulk density: about 1.6 from BASF, 50 wt.% CuO, 46 wt.% Cr ₂ O ₃ 4% by weight of C

Catalyst 4:

1/8 "sheet, bulk density: 1.3 from BASF at 30% by weight Al ₂ O ₃ Containing 58 wt.% CuO and 12 wt.% MnO ₂

Catalyst 5:

1/8 "sheet, bulk density: about 1 from BASF, 13 wt.% CuO, 84 wt.% Al ₂ O ₃ 3% by weight of C

Inert materials used: raschig ring No. 3 (glass: 3mm x 3mm)

Comparative example 1:

the reactor was filled with 200ml of Raschig ring No. 3 as the lowermost layer, 100ml of catalyst 5 as the lowermost catalyst layer, 1.1 liter of catalyst 4, 360ml of catalyst 2, 150ml of catalyst 1 and another 200ml of Raschig ring No. 3 as the uppermost layer. The reactor was then brought to a hydrogen pressure of 57 bar over 44 hours and the acid maleate feed rate to 396 g/h. In the top part, the Pd/ZrO is impacted by the raw material flow ₂ The temperature at the beginning of the catalyst bed of the catalyst was set at 175 ℃. 250mol of recycle gas per mol of acid maleate-containing stream are produced. After an operating time of 44 hours the reactor was stable and the outlet temperature was 195 ℃. The temperature rise downstream of the Pd catalyst bed was 10-15 ℃ 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 number of other components.

Comparative example 2:

similar to comparative example 1, except that the inlet temperature was decreased to 165 ℃. The reactor outlet temperature was 185 ℃. The hydrogenation output comprises 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 number of other components. It is apparent that the content of n-butanol as an impurity is decreased and the content of dimethyl succinate as an impurity is increased due to the decrease of the inlet temperature as compared with comparative example 1.

Comparative example 3:

similar to comparative example 1, except that the reactor inlet temperature was decreased from 175 ℃ to 170 ℃ after 6 days, decreased to 165 ℃ after 13 days and reset to 175 ℃ after 4 days and held for 2 days. The pressure drop over the reactor increased by more than 5% from an initial value of 32 mbar after 10 days of operation and was 20% higher than this initial value after another 10 days, with the result that the reactor was shut down after a further 14 days. Detaching the catalyst after cooling and depressurization showed adhesion in the catalyst layer, indicating that the Zn-containing catalyst 2 was damaged. At the end of the run (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 number of other components.

Comparative example 4:

150ml of catalyst 1 was installed between the respective layers of the number 3 Raschig rings above and below it, placed under a hydrogen atmosphere and heated to 175 ℃. The hydrogenation was carried out at a reaction inlet temperature of 175 ℃ then 185 ℃ and finally 165 ℃ under 57 bar and 1375g/h of recycle gas, 45g/h of fresh hydrogen and 390-400g/h of acid-containing maleate. The maximum temperature rise of the catalyst bed was 15 ℃. After the onset effect subsides, the hydrogenation output contains 93-94 area% dimethyl succinate, 3-4 area% methanol, 0.4-0.5 area% GBL, 0.4 area% THF, 0.02-0.03 area% BDO and other quantitatively minor components at 175 ℃, 93-94 area% dimethyl succinate, 3-4 area% methanol, 0.4-0.5 area% GBL, 0.4 area% THF, 0.03 area% BDO and other quantitatively minor components at 185 ℃ and 93-94 area% dimethyl succinate, 3-4 area% methanol, 0.3 area% GBL, 0.4 area% THF, 0.02 area% BDO and other quantitatively minor components at 165 ℃. It is evident that catalyst 1 hydrogenates almost only the double bonds of the acid-containing maleate.

Example 1:

similar to comparative example 1, except that instead of the Zn-containing catalyst 2 under the catalyst 1, a layer of 360ml of catalyst 3 was installed, on which a third layer of 1100ml of catalyst 4 was additionally provided, and on which a fourth layer of 100ml of catalyst 5 was additionally provided. The hydrogenation output comprised 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 number of other components. It is clear that the content of n-butanol as an impurity is reduced with respect to comparative example 1 and that the content of dimethyl succinate as an impurity is lower than in comparative example 1 and comparative example 2.

Example 2:

similar to comparative example 2, except that instead of the Zn-containing catalyst 2 under the catalyst 1, a layer of 360ml of catalyst 3 was installed, on which a third layer of 1100ml of catalyst 4 was additionally provided, and on which a fourth layer of 100ml of catalyst 5 was additionally provided. The hydrogenation output comprised 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 number of other components. It is clear that the content of n-butanol as impurity is even more significantly reduced than in comparative example 2 and the content of dimethyl succinate as impurity is lower than in comparative example 1 and comparative example 2 due to the reduced inlet temperature compared to comparative example 1.

Example 3:

in analogy to comparative example 3, except that instead of this Zn-containing catalyst 2 below catalyst 1, a layer of 360ml of catalyst 3 was installed, a third layer of 1100ml of catalyst 4 was again placed thereon, a fourth layer of 100ml of catalyst 5 was again placed thereon and the hydrogenation was carried out over a period of 38 days, wherein after 7 days the reaction temperature was reduced from 175 ℃ to 170 ℃, after a further 2 days to 165 ℃ and after a further 6 days to 160 ℃, after a further 5 days again to 165 ℃, after a further 7 days to 170 ℃ and after a further 4 days to 175 ℃. During the adjustment of the inlet temperature between 165 ℃ and 175 ℃, 10 area% GBL was added to the feed on day 5 for 11 days. As expected, the product mixture was found to contain 27-29 area% less methanol at an otherwise similar composition over this period of time. The pressure drop across the reactor did not increase significantly throughout the test. These catalysts can be disassembled without problems and no sticking or deposition in the catalyst bed was observed. The temperature rise downstream of the catalyst bed 1 at each temperature setting was 10-15 ℃ relative to the inlet temperature.

The hydrogenation output at the end of the run contained 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 number of other components.

It is evident that the process of the invention makes it possible to avoid the deposits and sticking observed in comparative example 3, thus allowing a longer catalyst uptime.

Furthermore, in order to evaluate the different catalysts for the uppermost Pd-and/or Pt-containing catalyst layer and the various support materials for these catalysts, screening tests were carried out in a tubular reactor having a volume of 0.273 liter, an internal diameter of 20.5mm and operating continuously at a temperature of 10 bar and 175 ℃, the test reaction carried out being the hydrogenation of dimethyl maleate to give dimethyl succinate and the procedure being as follows:

dimethyl maleate (DMM for short) as starting material is fed from a storage container by means of a metering pump into an evaporator filled with a glass ring, where it is mixed with hydrogen and evaporated at a temperature of approximately 210 ℃. Hydrogen mixed with the raw material was fed via a trace heating pipe into a reactor filled with a catalyst layer between two glass rings and heated to 175 ℃. Downstream of the reactor, the reaction mixture is passed via a slightly heatable line through a cryogenically and thermostatically cooled double-tube heat exchanger, in which the products are condensed. The gas/liquid separation of the product stream is then carried out in a separator cooled with the same cryostat. The off-gas is passed through a cold trap cooled with cooling water to condense residual amounts of the liquid components. Unconsumed hydrogen was returned to the evaporator as recycle gas (600 l/h). Make up the hydrogen consumed (typically 60-70 Nl/h). The liquid product stream was discharged from the separator into a collection vessel by means of a leveller and analyzed for its composition by gas chromatography.

The amount of catalyst was in each case 100ml, placed in the reactor on a first layer of 84-86ml5 Raschig rings and subsequently covered with another layer of 126-129ml of the same glass rings.

Catalysts tested:

catalyst 6:

·ZrO ₂ pd support (monoclinic/tetragonal phase) containing 0.3 wt.% Pd (ring: 3 mm. times.3 mm. times.7 mm, bulk density: about 1.4kg/m ³ )

Catalyst 7:

pd on carbon, containing 1% by weight of Pd (pellets: 2.3-4.8mm screen size, bulk density: about 500kg/m ³ )

Catalyst 8:

pd on carbon, containing 0.8% by weight of Pd (3mm extrudates, bulk density: about 450 kg/m) ³ ) Catalyst 9:

pd on carbon, containing 0.5 wt% Pd (pellet: 2.3-4.8mm screen size, bulk density: about 450kg/m ³ )

Catalyst 10:

·g-/q-Al ₂ O ₃ pd support, containing 0.5 wt.% Pd (1.6mm spheres, bulk density: about 800 kg/m) ³ ) Catalyst 11:

·g-/q-Al ₂ O ₃ pd support, containing 0.5 wt.% Pd (3.2mm spheres, bulk density: about 800 kg/m) ³ ) Catalyst 12:

·a-Al ₂ O ₃ pd support comprising 0.3 wt.% Pd (4mm extrudates, bulk density: about 1050 kg/m) ³ ) Catalyst 13:

·a-Al ₂ O ₃ pd support comprising 0.5 wt.% Pd (4mm extrudates, bulk density: about 1050 kg/m) ³ ) Catalyst 14:

·ZrO ₂ loaded with 0.5 wt% Pd and 0.5 wt% Pt (monoclinic/tetragonal phase: 3 mm. times.3 mm. times.7 mm, ring: bulk density: about 1.4kg/m ³ )

Catalyst 15:

·ZrO ₂ loaded with 0.4 wt% Pt (monoclinic/tetragonal: ring: 3 mm. times.3 mm. times.7 mm, bulk density: about 1.4kg/m ³ )

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

Example numbering	Catalyst and process for preparing same	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

Examples 4-14 show that various catalysts on various support materials achieve a conversion that is similarly high as catalyst 1 or higher than catalyst 1. The reaction conditions preclude the achievement of complete conversion, which is actually advantageous to provide a visible difference for comparative screening. It is believed that the high levels of greater than 80% dimethyl succinate achieved in some cases may be due to the favorable distribution of the noble metal on the support.

Claims

1. A process for producing 1, 4-butanediol, gamma-butyrolactone and tetrahydrofuran comprising the steps of:

a) hydrogenating an acid maleate-containing stream in the gas phase with hydrogen in a single reactor, wherein the acid maleate-containing stream is hydrogenated on a fixed catalyst bed comprising at least two different and directly adjacent catalyst layers, the catalyst layer which passes through first in the direction of travel of the hydrogenation being a noble metal-containing catalyst layer comprising noble metals selected from Pd and Pt and the catalyst layer which passes through subsequently being 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 ℃ relative to the inlet temperature after entering the fixed catalyst bed and subsequent complete hydrogenation in the last catalyst layer of the fixed bed,

b) the resulting condensed mixture is separated by distillation in one or more columns into three mixtures, each of which contains one of the three products selected from the group consisting of 1, 4-butanediol, gamma-butyrolactone and tetrahydrofuran in excess.

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

3. The method according to claim 1 or 2, wherein the noble metal in the noble metal-containing catalyst layer has been applied to a catalyst layer selected from the group consisting of carbon, argillaceous earth, zeolite, TiO ₂ 、ZrO ₂ 、Al ₂ O ₃ 、SiO ₂ And a carrier for a mixture of these components.

4. A method according to any one of claims 1 to 3, wherein the noble-metal-containing catalyst layer comprises, as the noble metal, in the range from 0.001 to 15% by weight, based on the total weight of the catalyst, of palladium.

5. The process according to any one of claims 1 to 4, wherein the temperature of the gas stream is increased in the noble metal-containing catalyst layer such that an adiabatic temperature rise of at least 33% has been achieved along the fixed catalyst bed after passage through the noble metal-containing catalyst layer.

6. Process according to any one of claims 1 to 5, wherein the hydrogenation is carried out at a temperature in the range of 150 ℃ and 250 ℃ and at a pressure of 8 to 80 bar.

7. Process according to any one of claims 1 to 6, wherein the Zn-free Cu catalyst layer immediately after the noble metal-containing catalyst layer is selected from the group consisting of reduced copper chromite catalysts and reduced copper oxide catalysts which may also comprise other metals, metalloids and/or transition metals as promoters, with Zn and its compounds being excluded from the promoters.

8. The process according to any one of claims 1 to 7, wherein the Zn-free Cu catalyst comprises a Cu selected from Al ₂ O ₃ 、SiO ₂ 、TiO ₂ 、ZrO ₂ 、Mn ₂ O ₃ 、Cr ₂ O ₃ And La ₂ O ₃ Other groups of (2)And (4) dividing.

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

10. The process according to any of claims 1 to 9, wherein the acid maleate-containing stream further comprises, based on the maleic diester content, a total content of up to 50 wt. -% of γ -butyrolactone and succinic diester.

11. The method of any of claims 1-10, wherein the catalyst-containing layer is disposed on a layer of inert material and covered by another layer of inert material.

12. The method of any of claims 1-10, wherein the catalyst-containing layer is disposed on a layer of inert material.

13. The method according to any one of claims 1 to 10, wherein the catalyst-containing layer is covered by a layer of inert material.