EP1841719A1

EP1841719A1 - Method for producing butadiene from n-butane

Info

Publication number: EP1841719A1
Application number: EP06707724A
Authority: EP
Inventors: Catharina Klanner; Götz-Peter SCHINDLER; Sven Crone; Frieder Borgmeier; Mark Duda; Falk Simon
Original assignee: BASF SE
Current assignee: BASF SE
Priority date: 2005-01-17
Filing date: 2006-01-16
Publication date: 2007-10-10
Also published as: US8088962B2; TW200630333A; KR20070095335A; DE102005002127A1; CN101119949A; MY151171A; CN101119949B; EA012614B1; BRPI0606630A2; EA200701506A1; US20080183024A1; WO2006075025A1; KR101270890B1

Abstract

The invention relates to a method for producing butadiene from n-butane comprising the following steps: A) provision of a feed gas stream (a) containing n-butane; B) introduction of the feed gas stream (a) containing n-butane into at least one first dehydrogenation zone and the non-oxidative catalytic dehydrogenation of n-butane to obtain a product gas stream (b) containing n-butane, 1-butene, 2-butene, butadiene, hydrogen, low-boiling minor constituents, optionally carbon oxides and optionally water vapour; C) introduction of the product gas stream (b) of the non-oxidative catalytic dehydrogenation and a gas containing oxygen into at least a second dehydrogenation zone and the oxidative dehydrogenation of n-butane, 1-butene and 2-butene to obtain a product gas stream (c) containing n-butane, 2-butene, butadiene, low-boiling minor constituents, carbon oxides and water vapour, said stream having a higher butadiene content than the product gas (b); D) isolation of the low-boiling minor constituents and the water vapour to obtain a C<SUB>4</SUB> product gas stream (d), which essentially consists of n-butane, 2-butene and butadiene; E) separation of the C<SUB>4</SUB> product gas stream (d) into a stream (e1), which essentially consists of n-butane and 2-butene and a useful product stream (e2), which essentially consists of butadiene, by extractive distillation; F) return of the stream (e1) to the first dehydrogenation zone.

Description

Process for the preparation of butadiene from n-butane

description

The invention relates to a process for the preparation of butadiene from n-butane.

Butadiene is an important basic chemical and is used for example for the production of synthetic rubbers (butadiene homopolymers, styrene-butadiene rubber or nitrile rubber) or for the production of thermoplastic terpolymers (acrylonitrile-butadiene-styrene copolymers). Butadiene is further converted to sulfolane, chloroprene and 1,4-hexamethylenediamine (over 1,4-dichlorobutene and adiponitrile). By dimerization of butadiene, vinylcyclohexene can also be produced, which can be dehydrogenated to styrene.

Butadiene can be prepared by thermal cracking (steam cracking) of saturated hydrocarbons, usually starting from naphtha as the raw material. Steam cracking of naphtha produces a hydrocarbon mixture of methane, ethane, ethene, acetylene, propane, propene, propyne, allenes, butenes, butadiene, butynes, methylalls, C ₅ and higher hydrocarbons.

A disadvantage of the production of butadiene in the cracking process is that inevitably larger amounts of unwanted by-products occur.

The object of the invention is to provide a process for the preparation of butadiene from n-butane, incurred in the smallest possible extent coupling products.

The object is achieved by a process for the preparation of butadiene from n-butane with the steps

A) providing a n-butane containing feed gas stream a;

B) feeding the n-butane-containing feed gas stream a into at least one first dehydrogenation zone and non-oxidative catalytic dehydrogenation of n-butane, wherein a product gas stream b comprising n-butane, 1-butene, 2-butene, butadiene, hydrogen, low-boiling secondary constituents, possibly

Carbon oxides and optionally water vapor is obtained; C) feeding the product gas stream b of the non-oxidative catalytic dehydrogenation and an oxygen-containing gas into at least one second dehydrogenation zone and oxidative dehydrogenation of n-butane, 1-butene and 2-butene, wherein a product gas stream c containing n-butane, 2-butene, Butadiene, low-boiling minor components, carbon oxides and water vapor is obtained, which has a higher content of butadiene than the product gas stream b;

D) separation of the low-boiling secondary constituents and of water vapor, where a C ₄ product gas stream d consisting essentially of n-butane, 2

Butene and butadiene is obtained;

E) separation of the C ₄ product gas stream d into a stream e 1 consisting essentially of n-butane and 2-butene and a product stream e 2 consisting essentially of butadiene by extractive distillation;

F) recycling the stream e1 into the first dehydrogenation zone.

The inventive method is characterized by a particularly effective use of raw materials. Thus, losses of the raw material n-butane are minimized by recycling unreacted n-butane into the dehydrogenation. The coupling of non-oxidative catalytic dehydrogenation and oxidative dehydrogenation achieves a high butadiene yield.

In a first process part A, a feed gas stream a containing n-butane is provided. Usually, n-butane-rich gas mixtures such as liquefied petroleum gas (LPG) are assumed to be the raw material. LPG contains essentially saturated C ₂ -C ₅ hydrocarbons. It also contains methane and traces of C ₆ ⁺ hydrocarbons. The composition of LPG can vary widely. Advantageously, the LPG used contains at least 10% by weight of butane.

Alternatively, a refined C ₄ stream used for crackers or refineries.

The n-butane-containing dehydrogenation feed gas stream is usually crude butane. Preferably, it will be crude butane, which fulfills the specifications listed below (analogous to DE 102 45 585, quoting in volume fractions): Content of n-butane> 90 vol .-%, usually> 93 vol .-%, usually> 95 vol .-%; Content of isobutane <1% by volume, usually <0.5% by volume, generally <0.3% by volume; Content of 1-butene <1% by volume, usually <0.5% by volume, generally <0.3% by volume; Content of cis-butene <1 vol .-%, usually <0.5 vol .-%, usually <0.3 vol .-%; Content of trans-butene <1% by volume, usually <0.5% by volume, generally <0.3% by volume; Content of isobutene <1% by volume, usually <0.5% by volume, generally <0.3% by volume;

Total content of C ₄ hydrocarbons other than butanes and butenes <1

Vol .-%, usually <0.5 vol .-%, usually <0.3 vol .-%;

Content of methane <10 ppm, usually <5 ppm, usually 0 to 1 ppm;

Content of propane <100 ppm, usually <90 ppm, usually 50 to 80 ppm;

Total content of C ₅ hydrocarbons <5% by volume;

Total content of C ₆ -hydrocarbons <5% by volume; Total content of C ₇ -C ₈ hydrocarbons <20 ppm;

Total content of oxygen-containing organic compounds <30 ppm; Content of ionogenic Cl <0.1 mg / kg;

Total content of Cl-containing compounds and expressed as Cl <0.1 mg / kg; Total content of compounds containing F and expressed as F <0.1 mg / kg; Total content of compounds containing S and expressed as S <0.1 mg / kg;

Ag <1 μg / kg, AI <4 μg / kg, As <1 μg / kg, Au <1 μg / kg, Ba <10 μg / kg, Be <1 μg / kg, Bi <1 μg / kg, Ca <4 μg / kg, Cd <1 μg / kg, Co <1 μg / kg, Cr <1 μg / kg, Cu <0.2 mg / kg, Fe <30 μg / kg, Ga <1 μg / kg, Ge <2 μg / kg, Hg <1 μg / l, In <1 μg / kg, Ir <1 μg / kg, K <3 μg / kg, Li <1 μg / kg, Mg <1 μg / kg, Mn <2 μg / kg, Mo <1 μg / kg, Na <4 μg / kg, Nb <1 μg / kg, Ni <3 μg / kg, Pb <10 μg / kg, Pd <1 μg / kg, Pt < 1 μg / kg, Rh <1 μg / kg, Sb <1 μg / kg, Sn <1 μg / kg, Sr <1 μg / kg, Ta <1 μg / kg, Ti <1 μg / kg, T1 <1 μg / kg, V <1 μg / kg, Zn <0.1 mg / kg, Zr <1 μg / kg.

In a variant of the method according to the invention, the provision of the n-butane-containing dehydrogenation feed gas stream comprises the steps

(A1) providing a liquefied petroleum gas (LPG) stream,

(A2) Separation of propane and optionally methane, ethane and C ₅ ⁺ - hydrocarbons (mainly pentanes, besides hexanes, heptanes, benzene, toluene) from the LPG stream, wherein a stream containing butanes (n-butane and isobutane) is obtained . (A3) Separation of isobutane from the butane-containing stream, wherein the n-butane-containing feed gas stream is obtained, and optionally isomerization of the separated isobutane to a n-butane / isobutane mixture and recycling the n-butane / isobutane mixture in the isobutane -

Separation.

The separation of propane and optionally methane, ethane and C ₅ ⁺ hydrocarbons takes place, for example, in one or more conventional rectification columns. For example, in a first column low boilers (methane, ethane, propane) can be separated overhead and in a second column high boilers (C ₅ ⁺ hydrocarbons) at the bottom of the column. A stream containing butanes (n-butane and isobutane) is obtained, from which isobutane is separated off, for example in a customary rectification column. The remaining stream containing n-butane is used as feed gas stream for the subsequent butane dehydrogenation.

The separated isobutane stream is preferably subjected to isomerization. For this purpose, the isobutane-containing stream is fed into an isomerization reactor. The isomerization of isobutane to n-butane can be carried out as described in GB-A 2,018,815. An n-butane / isobutane mixture is obtained, which is fed into the n-butane / isobutane separation column.

The separated isobutane stream can also be supplied to a further use, for example, be used for the production of methacrylic acid, polyisobutene or methyl tert-butyl ether.

In a process part B, the n-butane-containing feed gas stream is fed to a dehydrogenation zone and subjected to non-oxidative catalytic dehydrogenation. In this case, n-butane is partially dehydrogenated in a dehydrogenation reactor on a dehydrogenating catalyst partially to 1-butene and 2-butene, wherein butadiene is formed. In addition, hydrogen and small amounts of methane, ethane, ethene, propane and propene fall on. Depending on the mode of dehydrogenation, carbon oxides (CO, CO ₂ ), water and nitrogen may also be present in the product gas mixture of the non-oxidative catalytic n-butane dehydrogenation. In addition, unreacted n-butane is present in the product gas mixture. The non-oxidative catalytic n-butane dehydrogenation can be carried out with or without oxygen-containing gas as a co-feed.

One feature of non-oxidative driving versus oxidative driving is the presence of hydrogen in the effluent gas. In oxidative dehydrogenation, no free hydrogen is formed in substantial quantities.

The non-oxidative catalytic n-butane dehydrogenation can in principle be carried out in all reactor types and procedures known from the prior art. A comparatively comprehensive description of dehydrogenation processes suitable according to the invention also contains "Catalytica® ^® Studies Division, Oxidative Dehydrogenation and Alternative Dehydrogenation Processes" (Study Number 4192 OD, 1993, 430 Ferguson Drive, Mountain View, California, 94043-5272, USA).

A suitable reactor form is the fixed bed tube or tube bundle reactor. These include the catalyst (dehydrogenation catalyst and, when working with oxygen as a co-feed, optionally special oxidation catalyst) as a fixed bed in a reaction tube or in a bundle of reaction tubes. The reaction tubes are usually heated indirectly by burning a gas, for example a hydrocarbon such as methane, in the space surrounding the reaction tubes. It is advantageous to apply this indirect form of heating only to the first about 20 to 30% of the length of the fixed bed and to heat the remaining bed length by the released in the context of indirect heating radiant heat to the required reaction temperature. Typical reaction tube internal diameters are about 10 to 15 cm. A typical Dehydrierrohrbündelreaktor comprises about 300 to 1000 reaction tubes. The temperature inside the reaction tube usually ranges from 300 to 700 ° C, preferably from 400 to 700 ° C. The working pressure is usually between 0.5 and 8 bar, often between 1 and 2 bar when using a low water vapor dilution (analogous to the Linde process for propane dehydrogenation), but also between 3 and 8 bar when using a high steam dilution (analogous to so-called "steam active reforming process" (STAR process) for the dehydrogenation of propane or butane by Phillips Petroleum Co., see US 4,902,849, US 4,996,387 and US 5,389,342.) Typical catalyst loadings (GHSV) are from 500 to 2000 h ^-1 on used hydrocarbon. The catalyst geometry can be, for example, spherical or cylindrical (hollow or full). Industrially, several (eg three) such tube bundle reactors can be operated in parallel in the first dehydrogenation zone. According to the invention, one or two of these reactors may optionally be in dehydrogenation operation, while in a second (third) reactor the catalyst feed is regenerated without having to suspend operation in the second dehydrogenation zone. Such a procedure is useful, for example, in the BASF Linde propane dehydrogenation process described in the literature.

Such a procedure is also applicable to the so-called "steam active reforming (STAR) process" developed by Phillips Petroleum Co. (see, for example, US-A 4,902,849, US-A 4,996,387 and US-A 5,389 342). The dehydrogenation catalyst used in the STAR process is preferably promoters containing platinum on zinc (magnesium) spinel as a support (see, for example, US Pat. No. 5,073,662). In contrast to the BASF Linde propane dehydrogenation process, the dehydrogenated propane is diluted with steam in the STAR process. Typical is a molar ratio of water vapor to propane in the range of 4 to 6. The reactor outlet pressure is often 3 to 8 bar and the reaction temperature is conveniently 480 to 620 ⁰ C. Typical catalyst loading with propane are 200 to 4000 h ^"1 ( GHSV).

The heterogeneously catalyzed, non-oxidative n-butane dehydrogenation can also be carried out in the moving bed in the process according to the invention. For example, the catalyst moving bed can be housed in a radial flow reactor. In this, the catalyst moves slowly from top to bottom, while the reaction gas mixture flows radially. This procedure is used for example in the so-called UOP Oleflex dehydrogenation. Since the reactors are operated quasi adiabatically in this process, it is expedient to operate several reactors connected in series as a cascade (typically up to four). Within the cascade, a recycle stream containing n-butane may be supplied from the 2nd dehydrogenation zone. As a result, it is possible to avoid too great temperature differences of the reaction gas mixture at the reactor inlet and at the reactor outlet and nevertheless achieve good overall sales.

When the catalyst bed has left the moving bed reactor, it is fed to the regeneration and then reused. As the dehydrogenation catalyst, for example, a spherical dehydrogenation catalyst consisting essentially of platinum on a spherical alumina carrier can be used for this process. In the UOP variant becomes the hydrogen is added to dehydrogenating propane to avoid premature catalyst aging. The working pressure is typically 2 to 5 bar. The molar hydrogen to propane ratio is suitably 0.1 to 1. The reaction temperatures are preferably 550 to 650 ⁰ C and the residence time of the catalyst in a reactor is about 2 to 10 h.

In the described fixed-bed process, the catalyst geometry may also be spherical, but also cylindrical (hollow or full) or otherwise geometrically designed.

The non-oxidative catalytic n-butane dehydrogenation can also be carried out as described in Chem. Eng. Be. 1992 b, 47 (9-11) 2313, heterogeneously catalyzed in a fluidized bed. Conveniently, two fluidized beds are operated side by side, one of which is usually in the state of regeneration. The working pressure is typically 1 to 2 bar, the dehydrogenation temperature usually 550 to 600 ⁰ C. The heat required for the dehydrogenation is thereby introduced into the reaction system by the dehydrogenation catalyst is preheated to the reaction temperature. By adding an oxygen-containing co-feed can be dispensed with the preheater, and the heat required directly in the reactor system by combustion of hydrogen and / or hydrocarbons in the presence of oxygen are generated. Optionally, a hydrogen-containing co-feed may additionally be admixed.

The heterogeneously catalyzed n-butane dehydrogenation in a fluidized bed can also be carried out as described in Proceedings De Witt, Petrochem. Review, Houston, Texas, 1992. A, N1 described for propane. Alternatively to the procedures described above, the heterogeneously catalyzed n-butane dehydrogenation can also be realized analogously to the process developed by ABB Lummus Crest (see Proceedings De Witt, Petrochem., Review, Houston, Texas, 1992, P1).

The non-oxidative catalytic n-butane dehydrogenation can be carried out with or without oxygen-containing gas as a co-feed in a tray reactor. This contains one or more consecutive catalyst beds. The number of catalyst beds may be 1 to 20, advantageously 1 to 6, preferably 1 to 4 and in particular 1 to 3. The catalyst beds are preferably flowed through radially or axially from the reaction gas. In general, such a tray reactor is operated with a fixed catalyst bed. In the simplest case, the Fixed catalyst beds in a shaft furnace reactor arranged axially or in the annular gaps of concentrically arranged cylindrical gratings. A shaft furnace reactor corresponds to a horde. The implementation of the dehydrogenation in a single shaft furnace reactor corresponds to a preferred embodiment, wherein it is possible to work with an oxygen-containing co-feed. In a further preferred embodiment, the dehydrogenation is carried out in a tray reactor with 3 catalyst beds. In a procedure without oxygen-containing gas as a co-feed, the reaction gas mixture in the tray reactor is subjected to intermediate heating on its way from one catalyst bed to the next catalyst bed, for example by passing over heated with hot gases heat exchanger surfaces or by passing through heated with hot fuel gases pipes.

In a preferred embodiment of the process according to the invention, the non-oxidative catalytic n-butane dehydrogenation is carried out autothermally. For this purpose, oxygen is added to the reaction gas mixture of the n-butane dehydrogenation in at least one reaction zone and the hydrogen and / or hydrocarbon contained in the reaction gas mixture at least partially burned, whereby at least a portion of the required Dehydrierwärme in the at least one reaction zone is generated directly in the reaction gas mixture ,

In general, the amount of the oxygen-containing gas added to the reaction gas mixture is selected such that the amount of heat required for the dehydrogenation of the n-butane is produced by the combustion of hydrogen present in the reaction gas mixture and optionally of hydrocarbons present in the reaction gas mixture and / or of carbon present in the form of coke is produced. In general, the total amount of oxygen fed, based on the total amount of butane, is 0.001 to 0.5 mol / mol, preferably 0.005 to 0.2 mol / mol, particularly preferably 0.05 to 0.2 mol / mol. Oxygen can be used either as pure oxygen or as an oxygen-containing gas mixed with inert gases, for example in the form of air. Preferred oxygen-containing gas is air or oxygen-enriched air having an oxygen content of up to 50% by volume. The inert gases and the resulting combustion gases generally additionally dilute and thus promote the heterogeneously catalyzed dehydrogenation.

The hydrogen burned to generate heat is the hydrogen formed in the catalytic n-butane dehydrogenation and, if appropriate, the hydrogen gas additionally added to the reaction gas mixture as hydrogen. Preferably should so much hydrogen be present that the molar ratio H ₂ IO ₂ in the reaction gas mixture immediately after the supply of oxygen is 1 to 10, preferably 2 to 5 mol / mol. This applies to multi-stage reactors for each intermediate feed of oxygen-containing and possibly hydrogen-containing gas.

The hydrogen combustion takes place catalytically. The dehydrogenation catalyst used generally also catalyzes the combustion of the hydrocarbons and of hydrogen with oxygen, so that in principle no special oxidation catalyst different from this one is required. In one embodiment, the reaction is carried out in the presence of one or more oxidation catalysts which selectively catalyze the combustion of hydrogen to oxygen in the presence of hydrocarbons. The combustion of these hydrocarbons with oxygen to CO, CO ₂ and water is therefore only to a minor extent. Preferably, the dehydrogenation catalyst and the oxidation catalyst are present in different reaction zones.

In multi-stage reaction, the oxidation catalyst may be present in only one, in several or in all reaction zones.

Preferably, the catalyst which selectively catalyzes the oxidation of hydrogen is disposed at the sites where higher oxygen partial pressures prevail than at other locations of the reactor, particularly near the oxygen-containing gas feed point. The feeding of oxygen-containing gas and / or hydrogen-containing gas can take place at one or more points of the reactor.

In one embodiment of the process according to the invention, an intermediate feed of oxygen-containing gas and of hydrogen-containing gas takes place before each tray of a tray reactor. In a further embodiment of the method according to the invention, the feed of oxygen-containing gas and optionally of hydrogen-containing gas takes place before each horde except the first Horde. In one embodiment, behind each feed point is a layer of a specific oxidation catalyst, followed by a layer of the dehydrogenation catalyst. In another embodiment, no special oxidation catalyst is present. The dehydrogenation temperature is generally 400 to 1100 ⁰ C, the pressure in the last catalyst bed of the tray reactor generally 0.2 to 5 bar, preferably 1 to 3 bar. The load (GHSV) is generally 500 up to 2000 h ^{~ 1} , in high-load mode also up to 100 000 h ^{~ 1} , preferably 4000 to 16 000 h ^"1 .

A preferred catalyst which selectively catalyzes the combustion of hydrogen contains oxides and / or phosphates selected from the group consisting of the oxides and / or phosphates of germanium, tin, lead, arsenic, antimony or bismuth. Another preferred catalyst which catalyzes the combustion of hydrogen contains a noble metal of VIII. And / or I. Nebengruppe.

Typically, the heterogeneously catalyzed partial dehydrogenation of n-butane requires comparatively high reaction temperatures. The achievable turnover is limited by the thermodynamic equilibrium. Typical reaction temperatures are from 300 to 800 ^° C. or from 400 to 700 ^° C. High temperatures and removal of the reaction product H ₂ favor the equilibrium position in the sense of the target product.

Since the heterogeneously catalyzed dehydrogenation reaction proceeds under volume increase, the conversion can be increased by lowering the partial pressure of the products. This can be achieved in a simple manner, for example by dehydration under reduced pressure and / or by admixing essentially inert diluent gases, such as, for example, water vapor, which is normally an inert gas for the dehydrogenation reaction. A dilution with water vapor as a further advantage, as a rule, a reduced coking of the catalyst used, since the water vapor reacts with coke formed according to the principle of coal gasification. Further suitable diluents for heterogeneously catalyzed butane dehydrogenation are, for example, CO, methane, ethane, CO ₂ , nitrogen and noble gases such as He, Ne and Ar. The diluents mentioned are generally also suitable as diluents in the 2nd dehydrogenation zone.

It is favorable to carry out the non-oxidative n-butane dehydrogenation at a working pressure of 0.1 to 3 bar and to dilute the gas mixture containing n-butane to be dehydrogenated with steam. On the one hand, the water vapor acts as a heat carrier and on the other hand reduces the product partial pressure, which has a favorable effect on the equilibrium position of the n-butane dehydrogenation. Furthermore, the co-use of water vapor has an advantageous effect on the service life of the noble metal-containing dehydrogenation catalysts. Furthermore, hydrogen can be added. The molar ratio of molecular hydrogen to n-butane is usually ≤ 5. The molar ratio of water vapor to n-butane can be up to 30, more appropriate Way 0.1 to 2 and more favorably 0.5 to 1. An advantage of a procedure with lower n-butane conversion is that only a comparatively low amount of heat is consumed in a single pass of the reaction gas mixture through the first dehydrogenation and comparatively low reaction temperatures are sufficient to achieve the desired conversion.

The n-butane dehydrogenation can also be carried out (quasi) adiabatically. In this case, the reaction gas starting mixture is usually first heated to a temperature of 450 to 700 ° C (preferably from 550 to 65O ⁰ C), for example by direct firing of the reactor wall. During the adiabatic passage through the catalyst bed, the reaction gas mixture will then be cooled by about 30 ^° C. to 200 ^° C., depending on the conversion and dilution.

As the first dehydrogenation zone, a single shaft furnace reactor may be sufficient as a fixed bed reactor, which is flowed through by the reaction gas mixture axially and / or radially.

In the simplest case, this is a single closed reaction volume, for example a container whose inside diameter is 0.1 to 10 m, possibly also 0.5 to 5 m, and in which the fixed catalyst bed is placed on a carrier device (for example a grid). is applied. The catalyst-charged reaction volume, which is substantially thermally insulated in adiabatic operation, is thereby flowed through axially by the hot, n-butane-containing reaction gas. The catalyst geometry can be both spherical and annular or strand-shaped. About introduced into the catalyst bed supply lines, the originating from the 2nd dehydrogenation n-butane recycle gas can be injected. Since in this case the reaction volume can be realized by a very cost-effective apparatus, all catalyst geometries which have a particularly low pressure loss are to be preferred. These are above all catalysts, which lead to a large void volume or structured, such as monoliths or honeycomb bodies. For realizing a radial flow of the n-butane-containing reaction gas, the reactor may, for example, consist of two cylindrical grids located concentrically inside one another, and the catalyst bed may be arranged in its annular gap. In the adiabatic case, the jacket would, if necessary, again be thermally insulated. After a prolonged period of use, the aforementioned catalysts can be regenerated in a simple manner, for example, by dilute air at first at initial temperatures of from 300 to 600 ° C., often at 400 to 550 ° C., first in regeneration stages with nitrogen and / or water vapor the catalyst bed passes. The catalyst loading with regeneration gas (eg air) can be, for example, 50 to 10,000 h ^-1 and the oxygen content of the regeneration gas 0.1 to 21% by volume in subsequent further regeneration stages, air can be used as regeneration gas under otherwise identical regeneration conditions Flue gases with a residual oxygen content of 1 to 6% by volume, which are formed during combustion and still contain nitrogen, water vapor, carbon oxides and residual hydrocarbons, are also suitable It is recommended that the catalyst be regenerated with inert gas (for example N ₂ It is then generally advisable to regenerate with pure hydrogen or hydrogen diluted with inert gas (preferably water vapor) (the hydrogen content should be ≥ 1% by volume).

The heterogeneously catalyzed, non-oxidative n-butane dehydrogenation can be operated with comparatively low butane conversion (≦ 15 mol%), in which case the high selectivity and catalyst lifetime are advantageous. Preferably, the n-butane dehydrogenation should be operated at conversions> 15 mol%, more preferably> 30 mol%. The load of reaction gas may, for example, be 100 to 10,000 or 40,000 h ^-1 , for example 300 to 7000 h ^-1 or about 500 to 4000 h ^-1 .

In the simplest case, the fixed catalyst beds in a shaft furnace reactor are arranged axially or in the annular gaps of centrally arranged cylindrical gratings. However, it is also possible to arrange the annular gaps in segments one above the other and to guide the gas after radial passage in one segment in the next segment above or below it.

Conveniently, the reaction gas mixture undergoes intermediate heating on its way from one catalyst bed to the next catalyst bed, for example, by passing it over hot-surface heat exchanger surfaces (e.g., fins) or passing hot-gas-heated tubes in the tray reactor.

If the tray reactor is otherwise operated adiabatically, it is sufficient for butane conversions ≦ 30 mol%, especially when using the catalysts described in DE-A 199 37 107, in particular the exemplary embodiments, that is Reaction gas mixture preheated to a temperature of 450 to 550 ° C in the dehydrogenation reactor to lead and keep within the Horde reactor in this temperature range. For higher butane conversions, the reaction gas mixture is expediently preheated to higher temperatures into the dehydrogenation reactor (these can be up to 700 ° C.) and kept within this elevated temperature range within the tray reactor.

It is even more clever to carry out the above described intermediate heating directly (autothermal mode). For this purpose, molecular oxygen is added to the reaction gas mixture either already before flowing through the first catalyst bed (in which case molecular hydrogen should then be added to the reaction gas starting mixture) and / or between the subsequent catalyst beds. For example, a limited combustion of molecular hydrogen contained in the reaction gas mixture and formed in the course of the heterogeneously catalyzed n-butane dehydrogenation and / or added to the reaction gas mixture can be effected (as a rule catalyzed by the dehydrogenation catalysts themselves). It may be expedient to provide in the tray reactor catalyst beds, which are charged with a catalyst which selectively catalyzes the combustion of hydrogen. Examples of such catalysts are those described in US Pat. Nos. 4,788,371, 4,886,928, 5,430,209, 5,530,171, 5,527,979, and 5,563,314 described catalysts into consideration; For example, these and the dehydrogenation catalyst can be arranged in alternating catalyst beds in the tray reactor.

In general, the oxygen feed is carried out so that the oxygen content of the reaction gas mixture, based on the amount of molecular hydrogen contained therein, 0.5 to 50, preferably 10 to 25 vol .-%. Both pure molecular oxygen or inert gas, for example CO, CO ₂ , N ₂ and / or noble gases, dilute oxygen, but in particular air and oxygen-enriched air, come into consideration as oxygen source. The resulting combustion gases generally additionally dilute and thus promote the heterogeneously catalyzed n-butane dehydrogenation.

The isothermal nature of the heterogeneously catalyzed n-butane dehydrogenation can be further improved by placing in the rack reactor in the spaces between the

Catalyst beds attaches closed internals. These internals contain suitable solids or liquids that are above a certain temperature evaporate or melt while consuming heat and condense where this temperature falls below, thereby releasing heat.

One possibility is the gas mixture fed to the first dehydrogenation zone to that required for the heterogeneously catalyzed butane dehydrogenation in the first dehydrogenation zone

It also consists in adding molecular hydrogen to it and burning it by means of molecular oxygen to suitable specific combustion catalysts and by means of the thus released

Heat of combustion to cause heating to the desired reaction temperature. The resulting combustion products such as CO ₂ , H ₂ O as well as the N ₂ possibly present in the molecular oxygen required for the combustion form advantageous inert diluent gases.

In one embodiment of the invention, an intermediate feed of oxygen-containing gas is optionally carried out before each horde of the tray reactor. In a further embodiment of the method according to the invention, the feed of oxygen-containing gas takes place before each horde except the first horde. In another

Embodiment of the method according to the invention is behind everyone

Oxygen feed site, a bed of specific, suitable for H ₂ oxidation oxidation catalyst present, followed by a bed of

Dehydrogenation catalyst. If necessary, external molecular hydrogen (in pure form or diluted with inert gas) may additionally be added upstream of each horde. In a less preferred embodiment, the

Catalyst beds also contain mixtures of dehydrogenation and H ₂ oxidation catalysts.

The dehydrogenation in the tray reactor is generally from 400 to 800 ⁰ C, the pressure is generally 0.2 to 10 bar, preferably 0.5 to 4 bar and particularly preferably 1 to 3 bar. The total gas load (GHSV) is usually 500 to 10,000 h ^"1 , in high-load mode also up to 80000 h ^{" 1} , regularly at 30,000 to 40,000 h ^"1 .

In principle, all dehydrogenation catalysts known in the art are suitable for the heterogeneously catalyzed n-butane dehydrogenation. These can be roughly divided into two groups. Namely, in those which are oxidic in nature (for example chromium oxide and / or alumina) and in those which contain at least one precious metal deposited on a usually oxidic support. Among other things, all dehydrogenation catalysts which can be used in WO 01/96270, EP-A 731077, DE-A 102 11 275, DE-A 101 31 297, WO 99/46039, US-A 4 788 371, EP-AO 705 136, WO 99/29420, US-A 4 220 091, US-A 5 430 220, US-A 5 877 369, EP-AO 117 146, DE-A 199 37 196, DE-A 199 37 105 and the DE-A 199 37 107 are described.

The dehydrogenation catalysts used generally have a carrier and an active composition. The carrier is usually made of a heat-resistant oxide or mixed oxide. Preferably, the dehydrogenation catalysts contain a metal oxide selected from the group consisting of zirconium dioxide, zinc oxide, alumina, silica, titania, magnesia, lanthana, ceria and mixtures thereof as a carrier. The mixtures may be physical mixtures or chemical mixed phases such as magnesium or zinc-aluminum oxide mixed oxides. Preferred supports are zirconia and / or silica, particularly preferred are mixtures of zirconia and silica.

For example, dehydrogenation catalysts comprising 0 to 99.9% by weight of zirconium dioxide, 0 to 60% by weight of alumina, silica and / or titanium dioxide and 0.1 to 10% by weight of at least one element of the first or second Main group (particularly preferably potassium and / or cesium), a third subgroup element, an element of the eighth subgroup of the Periodic Table of the Elements (particularly preferably platinum and / or palladium), lanthanum and / or tin, with the proviso that the sum % by weight gives 100% by weight.

For example, the dehydrogenation catalysts are in the form of catalyst strands (diameter typically 1 to 10 mm, preferably 1 to 5 to 5 mm, length typically 1 to 20 mm, preferably 3 to 10 mm), tablets (preferably of similar dimensions as the strands) and / or catalyst rings (outer diameter and length typically 2 to 30 mm, for example 2 to 10 mm, wall thickness expediently 1 to 10 mm, for example 1 to 5 mm or 1 to 3 mm) before.

In general, the dehydrogenation catalysts (for example those described in DE-A 199 37 107) catalyze both the dehydrogenation of n-butane and the combustion of molecular hydrogen.

The active composition of the dehydrogenation catalysts generally contain one or more elements of VIII. Subgroup, preferably platinum and / or palladium, more preferably platinum. In addition, the dehydrogenation catalysts have one or more elements of the I. and / or II. Main group, preferably potassium and / or cesium. Furthermore, the dehydrogenation catalysts may contain one or more elements of the IM. Subgroup including the lanthanides and actinides, preferably lanthanum and / or cerium. Finally, the dehydrogenation catalysts may contain one or more elements of III. and / or IV. Main group, preferably one or more elements from the group consisting of boron, gallium, silicon, germanium, tin and lead, particularly preferably tin.

In a preferred embodiment, the dehydrogenation catalyst contains at least one element of subgroup VIII, at least one element of main group I and / or II, at least one element of IM. and / or IV. main group and at least one element of IM. Subgroup including the lanthanides and actinides.

For example, according to the invention, all dehydrogenation catalysts can be used which are described in WO 99/46039, US Pat. No. 4,788,371, EP-A 705,136, WO 99/29420, US Pat. No. 5,220,091, US Pat. No. 5,430,220, US Pat. No. 5,877,369, EP 0 117 146, DE-A 199 37 106 DE-A 199 37 105 and DE-A 199 37 107 are disclosed. Particularly preferred catalysts for the above-described variants of the autothermal n-butane dehydrogenation are the catalysts according to Examples 1, 2, 3 and 4 of DE-A 199 37 107.

Characteristic of the heterogeneously catalyzed, non-oxidative dehydrogenation of n-butane is that it is endothermic. That is, the heat energy required for the adjustment of the required reaction temperature and the heat energy required for the reaction must either be supplied to the starting gas mixture in advance and / or in the course of the heterogeneously catalyzed dehydrogenation. Optionally, the required heat of reaction is removed from the reaction gas mixture itself.

Further, for the heterogeneously catalyzed, non-oxidative dehydrogenation of n-butane due to the high reaction temperatures required, it is typical to form low boiling, high molecular weight organic compounds and carbon in small amounts which deposit on and deactivate the catalyst surface. To counteract this deactivation, the n-butane-containing reaction gas mixture can be diluted with water vapor. Depositing carbon is partially or completely converted under these conditions according to the principle of coal gasification. Another way to remove deposited carbon compounds from the catalyst surface is to charge the dehydrogenation catalyst from time to time at elevated temperature with an oxygen-containing gas (conveniently in the absence of hydrocarbons) and thereby burn off the deposited carbon. However, a substantial suppression of the formation of carbon deposits on the catalyst is also possible by adding molecular hydrogen to the n-butane to be dehydrogenated before bringing it into contact with the dehydrogenation catalyst.

The dehydrogenating n-butane may also be added to a mixture of water vapor and molecular hydrogen. Addition of molecular hydrogen also minimizes the formation of allenes (1,2-butadiene, propadiene), butynes, propyne, and acetylene as by-products.

The non-oxidative n-butane dehydrogenation is preferably carried out in the presence of water vapor. The added water vapor serves as a heat carrier and supports the gasification of organic deposits on the catalysts, whereby the coking of the catalysts counteracted and the service life of the catalysts is increased. The organic deposits are converted into carbon monoxide, carbon dioxide and possibly water.

The dehydrogenation catalyst can be regenerated in a manner known per se. Thus, steam can be added to the reaction gas mixture or, from time to time, an oxygen-containing gas can be passed over the catalyst bed at elevated temperature and the deposited carbon burned off. Dilution with water vapor shifts the equilibrium to the products of dehydration. Optionally, the catalyst is reduced after regeneration with a hydrogen-containing gas.

In the non-oxidative catalytic n-butane dehydrogenation, a gas mixture is obtained which, in addition to butadiene 1-butene, 2-butene and unreacted n-butane, contains minor constituents. Common secondary constituents are hydrogen, water vapor, nitrogen, CO and CO ₂ , methane, ethane, ethene, propane and propene. The composition of the gaseous mixture leaving the first dehydrogenation zone can vary widely depending on the mode of dehydrogenation. Thus, when carrying out the preferred autothermal dehydrogenation with the introduction of oxygen, the product gas mixture has a comparatively high content of water vapor and carbon oxides. For driving without feed of Oxygen, the product gas mixture of the non-oxidative dehydrogenation a comparatively high content of hydrogen.

The product gas stream b of the non-oxidative autothermal n-butane dehydrogenation typically contains 0.1 to 15% by volume of butadiene, 1 to 20% by volume of 1-butene, 1 to 35% by volume of 2-butene (cis / trans-2-butene), 20 to 80% by volume of n-butane, 1 to 70% by volume of steam, 0 to 10% by volume of low-boiling hydrocarbons (methane, ethane, ethene, propane and propene), 0, 1 to 40% by volume of hydrogen, 0 to 70% by volume of nitrogen and 0 to 15% by volume of carbon oxides.

The product gas stream b leaving the first dehydrogenation zone is preferably separated into two substreams, wherein only one of the two substreams is subjected to the further process parts C to F and the second substream can be returned to the first dehydrogenation zone. A corresponding procedure is described in DE-A 102 11 275. However, it is also possible to subject the entire product gas stream b of the non-oxidative catalytic n-butane dehydrogenation to the further process parts C to F.

The non-oxidative catalytic dehydrogenation according to the invention is followed by an oxidative dehydrogenation (oxydehydrogenation) as process part C. Essentially, 1-butene and 2-butene are dehydrogenated to 1,3-butadiene, with 1-butene generally reacting almost completely.

This can in principle be carried out in all reactor types and operating modes known from the prior art, such as, for example, in a fluidized bed, in a tray furnace, in a fixed-bed tubular reactor or in a tube heat exchanger reactor. The latter is preferably used in the method according to the invention. To carry out the oxidative dehydrogenation, a gas mixture is required which has a molar oxygen: n-butenes ratio of at least 0.5. Preference is given to working at an oxygen: n-butenes ratio of 0.55 to 50. In order to set this value, the product gas mixture originating from the non-oxidative catalytic dehydrogenation is generally mixed with oxygen or an oxygen-containing gas, for example air. The resulting oxygen-containing gas mixture is then fed to the oxydehydrogenation.

Suitable processes are described, for example, in WO-A 2004007408, DE 10361822, DE 10361823 and DE 10361824. In the documents WO 01/96270, DE-A 10245585, DE-A 10246119, DE-A 10313210, DE-A 10313214, DE-A 10313213, DE-A 10313212, DE-A 10308824, DE-A 10313208 and DE-A A 10211275 processes for the preparation of acrolein or acrylic acid from propane or propene and suitable reactors and catalysts are described, as they can be used in an analogous manner for the production of butadiene. In DE-A 10137534 suitable reactors and catalysts are described for sub-steps of the production of maleic anhydride starting from a mixture containing n-butenes, which can also be used correspondingly in the process according to the invention. In this case, pure molecular oxygen, air, oxygen-enriched air or any other mixture of oxygen and inert gas can be added as the oxidizing agent.

Preferably, the feed gas mixture fed to the second dehydrogenation zone has the following composition, see also WO-A 04/07408, DE-A 102 45 585 and DE-A 102 46 119. Thus, the molar ratios n-butane: n Butenes: N ₂ : O ₂ : H ₂ O: Other = 0.5 to 20: 1: 0.1 to 40: 0.1 to 10: 0 to 20: 0 to 1, for example = 1 or 2 to 10: 1: 0.5 to 20: 0.5 to 5: 0.01 to 10: 0 to 1 or even = 3 to 6: 1: 1 to 10: 1 to 3: 0.1 to 2: 0 to 0.5. The load (Nl / I'h) of the oxidation catalyst with reaction gas is frequently 1500 to 2500 h ^-1 or up to 4000 h ^-1 . The load with butenes may be 50 or 80 to 200 or 300 and more Nl / lh.

The catalysts which are particularly suitable for the oxydehydrogenation are generally based on a Mo-Bi-O-containing multimetal oxide system, which as a rule also contains iron. In general, the catalyst system contains further additional components from the 1st to 15th group of the Periodic Table, such as potassium, magnesium, zirconium, chromium, nickel, cobalt, cadmium, tin, lead, germanium, lanthanum, manganese, tungsten, phosphorus, Cerium, aluminum or silicon.

Suitable catalysts for the oxydehydrogenation are disclosed in DE-A 44 31 957, the non-prepublished German patent application DE 102004025445 and DE-A 44 31 949. This applies in particular to those of the general formula I in the two abovementioned publications. Particularly advantageous catalysts for oxydehydrogenation are disclosed in DE-A 103 25 488, DE-A 103 25 487, DE-A 103 53 954, DE-A 103 44 149, DE-A 103 51 269, DE-A 103 50 812 DE-A 103 50 822.

Suitable catalysts and their preparation are described, for example, in US Pat. No. 4,423,281 (Mθi ₂ BiNi ₈ Pb ₀ , ₅ Cr ₃ Ko, ₂ O _x and Mθi ₂ Bi _b Ni 7 Al ₃ Cro, ₅ Ko, ₅ O _x ), US Pat. No. 4,336,409 (MOi ₂ BiNi ₆ Cd ₂ Cr ₃ Po ₁₅ O _x ), DE-A 26 00 128 (Mθi ₂ BiNio, ₅ Cr ₃ Po, ₅ Mg7, ₅ Ko, iOχ + SiO ₂ ) and DE-A 24 40 329 ( MOi ₂ BiCo ₄₁₅ Ni ₂₁₅ Cr ₃ Po ₁₅ K _O jO _x ).

Particularly suitable are the multimetal oxide active compounds of the general formula I of DE-A 199 55 176, the multimetal oxide active compounds of the general formula I of US Pat

DE-A 199 48 523, the Multimetalloxidaktivmassen the general formula I of DE-A

199 48 523, the Multimetalloxidaktivmassen of the general formulas I, II and IM of

DE-A 101 01 695, the multimetal oxide active compounds of the general formulas I, II and III of DE-A 199 48 248 and the multimetal oxide active compounds of the general formulas I, II and III of DE-A 199 55 168 and those in EP-A 0 700 714 mentioned

Multimetal.

Also suitable for this oxidation step are the Mo, Bi and Fe-containing multimetal oxide catalysts described in DE-A 100 46 957, DE-A 100 63 162, DE-C 33 38 380, DE-A 199 02 562, EP-A 15 565, DE-C 23 80 765, EP-A 807 465, EP-A 27 9374, DE-A 33 00 044, EP-A 575 897, US-A 4438217, DE-A 198 55 913, WO 98/24746 DE-A 197 46 210 (of the general formula II), JP-A 91/294239, EP-A 293 224 and EP-A 700 714 are disclosed. This applies in particular to the exemplary embodiments in these documents, among which those of WO-A 04/07408, EP-A 15 565, EP-A 575 897, DE-A 197 46 210 and DE-A 198 55 913 are particularly preferred. Particularly noteworthy is the catalyst according to Example 1c of EP-A 15565 and a prepared in a corresponding way catalyst whose active material, however, the composition Mθi ₂ Ni ₆ Zn ₂ ₅ Fe ₂ biip _0, oo ₆₅ Ko, o ₆ θχ'10Si0 ₂ having. Furthermore, the catalyst according to Example No. 3 from DE-A 198 55 913 (stoichiometry: Mθi ₂ Co ₇ Fe ₃ Bio, ₆ Ko, ₈ Sii, ₆ O _x ) should be emphasized as a hollow cylinder full catalyst with the dimensions 5 mm × 3 mm × 2 mm (outer diameter × height × inner diameter) and the multimetal oxide M-unsupported catalyst according to Example 1 of DE-A 197 46 210. Further, the multimetal oxide catalysts of US Pat. No. 4,438,217 are suitable. The hollow cylinders preferably have the following dimensions: 5.5 mm × 3 mm × 3.5 mm, or 5 mm × 2 mm × 2 mm, or 5 mm × 3 mm × 2 mm, or 6 mm × 3 mm × 3 mm , or 7 mm x 3 mm x 4 mm (each outer diameter x height x inner diameter). Other suitable catalyst geometries are strands (eg 7.7 mm in length and 7 mm in diameter, or 6.4 mm in length and 5.7 mm

Diameter).

A variety of suitable for the oxydehydrogenation of n-butenes to butadiene Multimetalloxidaktivmassen can be classified under the general formula IV, Mo ₁₂ Bi _a Fe _b X ¹ _c X ² _d X ³ _Θ X ⁴ , On (IV) ₁

in which the variables have the following meaning:

X ¹ = nickel and / or cobalt,

X ² = thallium, an alkali metal and / or an alkaline earth metal,

X ³ = zinc, phosphorus, arsenic, boron, antimony, tin, cerium, lead and / or tungsten,

X ⁴ = silicon, aluminum, titanium and / or zirconium,

a = 0.5 to 5, b = 0.01 to 5, preferably 2 to 4, c = 0 to 10, preferably 3 to 10, d = 0 to 2, preferably 0.02 to 2, e = 0 to 8 , preferably 0 to 5, f = 0 to 10 and n = subsume a number determined by the valency and frequency of the elements other than oxygen in IV.

These are obtainable in a manner known per se (for example as described in DE-A 40 23 239) and are usually molded in bulk into spheres, rings or cylinders or else in the form of coated catalysts, i. used as the active composition coated, preformed, inert support body. Of course, they can also be used in powder form as catalysts.

In principle, active compounds of the general formula IV can be prepared in a simple manner by, starting from suitable sources of their elements, a very intimate, preferably finely divided, the elemental composition corresponding dry mixture and this calcined at temperatures of 350 to 650 ⁰ C. The calcination can be carried out both under inert gas and under an oxidizing atmosphere such as air (or a mixture of inert gas and oxygen) or under reducing atmosphere (eg a mixture of inert gas, NH ₃ , CO and / or H ₂ ). The calcination time may be several minutes to a few hours and usually decreases with increasing temperature. Suitable sources of the elemental constituents of the multimetal oxide active compositions IV are oxides or compounds which can be converted into oxides by heating, at least in the presence of oxygen. In addition to the oxides, suitable starting compounds are, in particular, halides, nitrates, formates, oxalates, citrates, acetates, carbonates, amine complexes, ammonium salts and / or hydroxides. Compounds such as NH ₄ OH, (NH ₄ J ₂ CO ₃ , NH ₄ NO ₃ , NH ₄ CHO ₂ , CH ₃ COOH, NH ₄ CH ₃ CO ₂ and / or ammonium oxalate, which decompose at the latest during the subsequent calcination gaseous escaping compounds and / or can be decomposed, can be incorporated into the intimate dry mixture in addition.

The intimate mixing of the starting compounds for the preparation of multimetal oxide active materials IV can be carried out in a dry or in a wet state.

If it is carried out in a dry state, the starting compounds are expediently used as finely divided powders and subjected to calcination after mixing and optionally compacting. Preferably, however, the intimate mixing takes place in the wet state. Usually, the starting compounds are mixed together in the form of an aqueous solution and / or suspension. Particularly intimate dry mixtures are then obtained in the described mixing process, if only from dissolved in

Form present sources of elemental constituents is assumed. When

Solvent is preferably used water. Subsequently, the resulting aqueous mass is dried, the drying process preferably by

Spray drying of the aqueous mixture with outlet temperatures of 100 to 150 ° C.

The multimetal oxide of the general formula IV can be used in the oxidative dehydrogenation shaped both in powder form and to certain catalyst geometries, wherein the shaping before or after the final

Calcination can be done. For example, from the powder form of the active composition or its uncalcined or partially calcined precursor composition

Compacting and shaping (e.g., by tableting, extruding, or extruding) unsupported catalysts, optionally with aids such as graphite or

Stearic acid as a lubricant and / or molding aids and reinforcing agents such

Microfibers of glass, asbestos, silicon carbide or potassium titanate can be added. Suitable Vollkatalysatorgeometrien are eg solid cylinder or hollow cylinder with an outer diameter and a length of 2 to 10 mm. In the case of the hollow cylinder, a wall thickness of 1 to 3 mm is appropriate. The full catalyst may also have spherical geometry, wherein the ball diameter may be 2 to 10 mm. A particularly favorable hollow cylinder geometry has the dimensions 5 mm × 3 mm × 2 mm (outer diameter × length × inner diameter), in particular in the case of solid catalysts.

The shaping of the pulverulent active composition or its pulverulent, not yet or only partially calcined precursor composition can also be effected by application to preformed, inert catalyst supports. The coating of the carrier bodies for the preparation of coated catalysts is usually carried out in a suitable rotatable container, e.g. in DE-A 29 09 671, EP-A 293 859 or EP-A 714 700. Conveniently, the powder to be applied is moistened to coat the carrier body and after application, e.g. by means of hot air, dried again. The layer thickness of the powder mass applied to the carrier body is suitably in the range from 10 to 1000 .mu.m, preferably from

50 to 500 microns and more preferably from 150 to 250 microns.

As carrier materials, it is possible to use customary porous or non-porous aluminum oxides, silicon dioxide, thorium dioxide, zirconium dioxide, silicon carbide or silicates, such as magnesium silicate or aluminum silicate. They are essentially inert. The carrier bodies may be regularly or irregularly shaped, with regularly shaped carrier bodies having a marked surface roughness, e.g. Spheres or hollow cylinders are preferred. Suitable is the use of substantially non-porous, surface-rough, spherical steatite supports whose diameter is 1 to 10 mm or 1 to 8 mm, preferably 4 to 5 mm. However, it is also suitable to use cylinders as support bodies whose length is 2 to 10 mm and whose outer diameter is 4 to 10 mm. In addition, in the case of rings suitable as a carrier according to the invention, the wall thickness is usually from 1 to 4 mm. Preferably used annular carrier bodies have a length of 2 to 6 mm, an outer diameter of 4 to 8 mm and a wall thickness of 1 to 2 mm. Particularly suitable according to the invention are rings of the geometry 7 mm × 3 mm × 4 mm (outer diameter × length × inner diameter) as the carrier body. The fineness of the catalytically active oxide materials to be applied to the surface of the support body is adapted to the desired shell thickness, see also EP-A 714 700.

Furthermore, multimetal oxide active compositions to be used for the oxydehydrogenation of n-butenes to butadiene are compounds of the general formula V

[Y ¹ βΥ ² bOχ-] p [Y ³ cY ⁴ _< rY ⁵ β-Y ^β rYVY ² h'O _/ ] _q (V) in which the variables have the following meaning:

Y ¹ = only bismuth or bismuth and at least one of the elements tellurium, antimony, tin and copper,

Y ² = molybdenum or molybdenum and tungsten,

Y ³ = an alkali metal, thallium and / or samarium,

Y ⁴ = an alkaline earth metal, nickel, cobalt, copper, manganese, zinc, tin, cadmium and / or mercury, Y ⁵ = iron or iron and at least one of the elements chromium and cerium,

Y ⁶ = phosphorus, arsenic, boron and / or antimony,

Y ⁷ = a rare earth metal, titanium, zirconium, niobium, tantalum, rhenium, ruthenium, rhodium, silver, gold, aluminum, gallium, indium, silicon, germanium, lead, thorium and / or uranium,

a '= 0.01 to 8, b' = 0.1 to 30, c '= 0 to 4, d' = 0 to 20, e '= 0.001 to 20, f = 0 to 6, g' = 0 to 15, h '= 8 to 16, x', y '= numbers determined by the valency and frequency of the non-oxygen elements in V, and p, q = numbers whose ratio p / q is 0.1 to 10 .

containing three-dimensionally extended areas of the chemical composition Y ¹ _a Υ ² _b 'O _X ' defined by their local composition, whose maximum diameter (longest direct line of the area of the area of two on the surface (interface) of the area Points) is 1 nm to 100 μm, often 10 nm to 500 nm or 1 μm to 50 or 25 μm.

Particularly advantageous multimetal oxide compositions V are those in which Y ¹ is only bismuth. Among these, in turn, those of the formula VI,

[Bi _a ..Z ² _b .. O ".] P .. [Z ² ₁₂ Z ³ _c ..Z ⁴ _d ..Fe _Θ ..Z ⁵ _f .Z ⁶ _g ..Z ⁷ _h .O _y ..] _q .. (VI),

in which the variants have the following meaning:

Z ² = molybdenum or molybdenum and tungsten,

Z ³ = nickel and / or cobalt,

Z ⁴ = thallium, an alkali metal and / or an alkaline earth metal, Z ⁵ = phosphorus, arsenic, boron, antimony, tin, cerium and / or lead,

Z ⁶ = silicon, aluminum, titanium and / or zirconium,

Z ⁷ = copper, silver and / or gold,

a "= 0.1 to 1, b" = 0.2 to 2, c "= 3 to 10, d" = 0.02 to 2, e "= 0.01 to 5, preferably 0.1 to 3, f "= 0 to 5, g" = 0 to 10, h "= 0 to 1, x", y "= numbers other than the valence and frequency of oxygen

Element in VI, p ", q" = numbers whose ratio p'Vq "is 0.1 to 5, preferably 0.5 to 2,

in which Z ² _b »= (tungstenV and Z ² ₁₂ = (molybdenum) ₁₂ is very particularly preferred, those masses VI being particularly preferred.

Furthermore, it is advantageous if at least 25 mol% (preferably at least 50 mol% and particularly preferably 100 mol%) of the total fraction [Y ¹ _a 'Y ² b'O _X '] _p ([Bi _a »Z ² _b »O _x »] _p ») of the multimetal oxide materials V or of the multimetal oxide materials VI in the form of three-dimensionally extended regions of the chemical composition Y ¹ _a 'Y ² _b ' O _X ', differentiated from their local environment due to their different chemical composition [Bi _a »Z ² _b O _x -] are present whose largest diameter in the range 1 nm to 100 microns. The preparation of multimetal oxide active compositions V is described, for example, in EP-A 575 897 and in DE-A 19 855 913.

The above-recommended inert support materials come i.a. also as inert materials for dilution and / or delimitation of the corresponding fixed catalyst beds from each other, or as a protective and / or the gas mixture aufheizende infill into consideration.

The oxydehydrogenation is carried out generally at a temperature of 220 to 490 ° C and preferably from 250 to 450 ⁰ C. It chooses a reactor inlet pressure that is sufficient, existing in the system and the subsequent workup

Overcome flow resistance. This reactor inlet pressure is usually 0.005 to 1 MPa gauge, preferably 0.01 to 0.5 MPa gauge. Naturally, the gas pressure applied in the inlet area of the reactor largely drops over the entire catalyst bed.

The oxydehydrogenation of n-butenes to butadiene may be carried out with the described catalysts e.g. be carried out in a one-zone Vielkontaktrohr fixed bed reactor, as described in DE-A 44 31 957.

As the oxidizing agent, oxygen is generally used. If N _{2 is} chosen as the inert diluent gas, the use of air or oxygen-enriched air as the source of oxygen proves to be particularly advantageous.

Frequently, for an n-butene: oxygen: inert gases (including water vapor) volume (NI) ratio of 1: (1, 0 to 3.0) :( 5 to 25), preferably 1: (1, 7 to 2 , 3) :( 10 to 15) worked. The reaction pressure is usually in the range of 1 to 3 bar and the total space load is preferably 1500 to 4000 or 6000 Nl / I'h or more. The load, based on n-butenes, is typically 90 to 200 Nl / l-h or to 300 Nl / l-h or more.

Preferably, the single-zone multiple contact tube fixed bed reactor is supplied with the feed gas mixture from above. As a heat exchange agent is expediently a molten salt, preferably consisting of 60 wt .-% potassium nitrate (KNO ₃ ) and 40 wt .-% sodium nitrite (NaNO ₂ ), or from 53 wt .-% potassium nitrate (KNO ₃ ), 40 wt .-% Sodium nitrite (NaNO ₂ ) and 7 wt .-% sodium nitrate (NaNO ₃ ) used. Viewed over the reactor molten salt and reaction gas mixture can be performed both in DC and in countercurrent. The molten salt itself is preferably guided meander-shaped around the catalyst tubes.

If the contact tubes are flown from top to bottom, it is expedient to feed the catalyst tubes from the bottom to the top as follows (for the flow from bottom to top, the feed sequence is expediently reversed):

initially at a length of 40 to 80 or to 60% of the catalyst tube length either with catalyst alone, or with a mixture of catalyst and inert material, the latter, based on the mixture, constituting a weight fraction of up to 20% by weight ( Section C);

- Thereafter, to a length of 20 to 50 or 40% of the total tube length either only with catalyst, or with a mixture of catalyst and inert material, the latter, based on the mixture, a weight fraction of up to 40 wt .-% makes (Section B); and

- Finally, to a length of 10 to 20% of the total tube length with a bed of inert material (Section A), which preferably causes the lowest possible pressure loss.

Preferably, section C is undiluted.

The aforementioned feed variant is particularly useful if as catalysts such as in Example 3 of WO-A 04/07408, Example 1 of DE-A 100 46 957 or according to Example 3 of DE-A 100 46 957 and as an inert material rings of steatite with Geometry 7mm x 7mm x 4mm (outside diameter x height x inside diameter) can be used. With regard to the salt bath temperature, the statements made in DE-A 44 31 957 apply.

However, the oxydehydrogenation of n-butenes to butadiene can also be carried out with the catalysts described in a two-zone multiple contact tube fixed bed reactor, as described in DE-A 199 10 506. In both cases described above, the conversion of the n-butenes obtained with a single pass is normally at values> 85 mol%, or> 90 mol% or> 95 mol%. The salt bath temperature of multiple contact tube reactors for the oxydehydrogenation of n-butenes to butadiene is generally from 300 to 400 ⁰ C. In addition, the heat exchange means normally (preferably molten salts) in such amounts through the multiple contact tube fixed-bed reactors out that the difference between its input and their starting temperature is usually ≤ 5 ° C.

It should be noted that the oxydehydrogenation can be carried out by first passing a reaction gas mixture over the catalyst bed containing no oxygen. In this case, the oxygen required for the oxydehydrogenation is provided as lattice oxygen. In a subsequent regeneration step with an oxygen-containing gas (e.g., air, oxygen-enriched air, or deoxygenated air), the catalyst bed is regenerated to subsequently be available again for an oxygen-free reaction gas mixture. Instead of tube bundle reactors, plate heat exchange reactors with salt and / or boiling cooling, as e.g. in DE-A 19 929 487 and DE-A 19 952 964, or fluidized-bed reactors are used.

In principle, the first and the second dehydrogenation zone in the process according to the invention can also be configured as described in WO-A 04007408, DE-A 19 837 517, US Pat.

DE-A 19 910 506, DE-A 19 910 508 and DE-A 19 837 519 are described.

The external temperature control in the two dehydrogenation zones, if appropriate in multi-zone reactor systems, is customarily adapted to the specific reaction gas mixture composition and catalyst feed in a manner known per se.

An excess of molecular oxygen usually has an advantageous effect on the kinetics of the dehydrogenation. Since, in contrast to the ratios in the first dehydrogenation zone (non-oxidative dehydrogenation), the hydrogenation of the n-butenes to butadiene undergoes kinetic control, in the second dehydrogenation zone (oxydehydrogenation) also worked with a molar excess of n-butenes against molecular oxygen become. In this case, the excess n-butenes also play the role of a diluent gas.

As the source of the molecular oxygen required in the oxydehydrogenation, which is mixed, for example, with the product gas mixture b before the second dehydrogenation is charged, both pure molecular oxygen and inert gas such as CO ₂ , CO, noble gases, N ₂ and / or saturated hydrocarbons become more dilute molecular Oxygen into consideration. Conveniently, air will be used as the source of oxygen at least to meet a partial demand for molecular oxygen.

By adding cold air to hot product gas mixture b, this can be cooled to the temperature required for oxydehydrogenation.

In the long-term operation of both dehydrogenation zones, a reduction in the catalyst quality (in particular with regard to activity and selectivity) of the catalysts present in the reaction zones will normally occur. On the one hand, this can be counteracted by regenerating the respective catalyst bed from time to time (for this purpose, for example, residual gas can also be passed over the catalyst bed of the oxidation catalyst at elevated temperature in the oxidation zone), as described e.g. in DE-A 10351269, DE-A 10350812 and DE-A 10350822. Also, the reaction temperature can be increased over the operating period to compensate for a reduced catalyst activity. Surprisingly, however, it has proven to be advantageous in both dehydrogenation zones to increase the respective working pressure in the gas phase, based on an identical loading of the respective catalyst bed with reaction gas mixture in Nl / lh, during the operating life of the catalyst bed, in order to deactivate the catalyst beds counteract, as it eg is described in the non-prepublished German patent application DE 102004025445. For this purpose, a pressure regulating device (in the simplest case a throttle device, for example a throttle valve or also a swirl regulator) can be attached at a suitable point, e.g. also partially permeable apertured apertures whose holes can be successively partially or completely closed. It is in principle sufficient to introduce the pressure control device at a suitable point in the flow path of the reaction gas mixture, from where the pressure increase can propagate through backflow into the reaction zones. Typical pressure increases (which can be carried out continuously in accordance with the deactivation in progress, but also discontinuously) can amount to up to 3000 mbar and more over the course of the operating time.

By coupling the non-oxidative catalytic, preferably autothermal dehydrogenation with the oxidative dehydrogenation of the n-butenes formed, a much higher yield of butadiene, based on n-butane used, is obtained. Furthermore, the non-oxidative dehydrogenation can be operated more gently. Comparable butadiene yields would be with exclusively non-oxidative dehydrogenation only at the price of significantly lower selectivities. Only oxidative dehydrogenation only low n-butane conversions are achieved.

The product gas stream c leaving the oxidative dehydrogenation contains, in addition to butadiene and unreacted n-butane, 2-butene and water vapor. As minor constituents it generally contains carbon monoxide, carbon dioxide, oxygen, nitrogen, methane, ethane, ethene, propane and propene, optionally hydrogen and oxygen-containing hydrocarbons, so-called oxygenates. In general, it contains only small amounts of 1-butene.

Oxygenates include, for example, lower aldehydes, lower alkane carboxylic acids (e.g., acetic, formic, and propionic acids), as well as maleic anhydride, benzaldehyde, aromatic carboxylic acids, and aromatic carboxylic acid anhydrides (e.g., phthalic anhydride and benzoic acid).

In general, the product gas stream leaving the oxidative dehydrogenation comprises c 1 to 50% by volume of butadiene, 20 to 80% by volume of n-butane, 0.5 to 40% by volume of 2-butene, 0 to 20% by volume. 1-butene, 0 to 70% by volume of steam, 0 to 10% by volume of low-boiling hydrocarbons (methane, ethane, ethene, propane and propene), 0.1 to 40% by volume of hydrogen, 0 to 70% by volume. % Nitrogen, 0 to 10% by volume of carbon oxides and 0 to 10% by volume of oxygenates. Oxygenates may be, for example, furan, acetic acid, maleic anhydride, formic acid or butyraldehyde.

In a process part D, the low-boiling minor constituents other than the C ₄ hydrocarbons (n-butane, isobutane, 1-butene, cis / trans-2-butene, isobutene, butadiene) are at least partly, but preferably in the range from Essentially completely separated from the product gas stream of n-butane dehydrogenation, wherein a C ₄ -Produktgasstrom d is obtained.

In one embodiment of the method according to the invention, first water is separated from the product gas stream c in the process part D. The separation of water can be carried out, for example, by condensation by cooling and / or compressing the product gas stream c and can be carried out in one or more cooling and / or compression stages.

The separation of the low-boiling secondary constituents from the product gas stream can be carried out by customary separation processes such as distillation, rectification, membrane process, absorption or adsorption. To separate off the hydrogen contained in the product gas stream c, the product gas mixture, optionally after cooling, for example, in an indirect heat exchanger, are passed through a usually formed as a tube membrane, which is permeable only for molecular hydrogen. The thus separated molecular hydrogen can, if necessary, at least partially used in the dehydrogenation or else be supplied to another utilization, for example, be used for generating electrical energy in fuel cells.

The carbon dioxide contained in the product gas stream c can be separated by CO ₂ gas scrubbing. The carbon dioxide gas scrubber may be preceded by a separate combustion stage in which carbon monoxide is selectively oxidized to carbon dioxide.

In a preferred embodiment of the process according to the invention, the gas stream c is compressed in at least one first compression stage and then cooled, wherein at least one condensate stream comprising water condenses out and a gas stream c 'containing n-butane, n-butenes, butadiene, hydrogen, water vapor, optionally carbon oxides and optionally inert gases remains.

The compression can be done in one or more stages. Overall, is compressed from a pressure in the range of 1.0 to 4.0 bar to a pressure in the range of 3.5 to 20 bar. After each compression stage is followed by a cooling step, in which the gas stream is cooled to a temperature in the range of 15 to 60 ⁰ C. The aqueous condensate stream may thus also comprise a plurality of streams in the case of multistage compression.

The gas stream c 'generally consists essentially of C ₄ hydrocarbons (essentially n-butane, 2-butene and butadiene), hydrogen, carbon dioxide and water vapor. In addition, the stream c 'may still contain low boilers, inert gases (nitrogen) and carbon oxides as further secondary components. The aqueous condensate stream is generally at least 80 wt .-%, preferably at least 90 wt .-% of water and also contains minor amounts of low boilers, C ₄ hydrocarbons, oxygenates and carbon dioxide. Suitable compressors are, for example, turbo, rotary piston and reciprocating compressors. The compressors can be driven, for example, with an electric motor, an expander or a gas or steam turbine. Typical compression ratios (outlet pressure: inlet pressure) per compressor stage are between 1, 5 and 3.0, depending on the design.

The cooling of the compressed gas takes place with heat exchangers, which can be designed, for example, as a tube bundle, spiral or plate heat exchanger. As coolant, cooling water or heat transfer oils are used in the heat exchangers. In addition, air cooling is preferably used using blowers.

In a preferred embodiment of the method according to the invention are the non-condensable or low-boiling gas components such as hydrogen, oxygen, carbon oxides, the low-boiling hydrocarbons (methane, ethane, ethene, propane, propene) and optionally nitrogen in an absorption / desorption cycle by means of a separated high-boiling absorbent, wherein a C ₄ product gas stream d is obtained, which consists essentially of the C ₄ - hydrocarbons. In general, the C ₄ product gas stream d is at least 80% by volume, preferably at least 90% by volume, particularly preferably at least 95% by volume, of the C ₄ hydrocarbons. Essentially, the stream d consists of n-butane, 2-butene and butadiene.

For this purpose, in an absorption stage, the product gas stream c is brought into contact with an inert absorbent after prior removal of water, and the C ₄ hydrocarbons are absorbed in the inert absorbent, whereby absorption medium laden with C ₄ hydrocarbons and an offgas containing the remaining gas constituents are obtained. In a desorption step, the C ₄ hydrocarbons are released from the absorbent again.

Inert absorbent used in the absorption stage are generally high-boiling nonpolar solvents in which the C ₄ - hydrocarbon mixture to be separated has a significantly higher solubility than the other gas constituents. The absorption can be carried out by simply passing the product gas stream c through the absorbent. But it can also be done in columns or in rotational absorbers. It can be used in cocurrent, countercurrent or cross flow. Suitable absorption columns are, for example, tray columns with bell, centrifugal and / or sieve trays, columns with structured packings, for example sheet metal packings with a specific surface area of 100 to 1000 m ² / m ³ as Mellapak ^® 250 Y, and packed columns. However, trickle and spray towers, graphite block absorbers, surface absorbers such as thick-layer and thin-layer absorbers as well as rotary columns, rags, cross-flow scrubbers and rotary scrubbers are also suitable.

Suitable absorbents are comparatively non-polar organic solvents, for example C ₅ -C ₈ -alkenoalkenes, or aromatic hydrocarbons, such as the paraffin distillation, naphtha or bulky group ethers, or mixtures of these solvents, these being a polar solvent such as 1, 2 Dimethyl phthalate may be added. Suitable absorbents are furthermore esters of benzoic acid and phthalic acid with straight-chain C 1 -C ₈ -alkanols, such as n-butyl benzoate,

Methyl benzoate, ethyl benzoate, dimethyl phthalate, diethyl phthalate, and so-called heat transfer oils, such as biphenyl and diphenyl ether, their chlorinated derivatives and triaryl alkenes. A suitable absorbent is a mixture of biphenyl and diphenyl ether, preferably in the azeotropic composition, for example, the commercially available Diphyl ^®. Frequently, this solvent mixture contains dimethyl phthalate in an amount of 0.1 to 25 wt .-%. Suitable absorbents are also pentanes, hexanes, heptanes, octanes, nonanes, decanes, undecanes, dodecanes, tridecanes, tetradecanes, pentadecanes, hexadecanes, heptadecanes and octadecanes or fractions obtained from refinery streams which contain as main components said linear alkanes. Suitable solvents are also butyrolactone, nitriles such as acetonitrile, propionitrile, methoxypropionitrile, ketones such as acetone, furfural, N-alkyl-substituted lower aliphatic acid amides such as dimethylformamide, diethylformamide, dimethylacetamide, diethylacetamide, N-formylmorpholine, N-alkyl-substituted cyclic acid amides (lactams) such as N-alkyl. Alkylpyrrolidones, especially N-methylpyrrolidone (NMP).

For desorption of the C ₄ hydrocarbons, the loaded absorbent is heated and / or expanded to a lower pressure. Alternatively, the desorption may also be by stripping or in a combination of relaxation, heating and stripping in one or more process steps. The absorbent regenerated in the desorption stage is returned to the absorption stage. In a variant of the method, the desorption step is carried out by relaxation and / or heating of the loaded desorbent.

In a preferred variant, the desorption is carried out in two stages, wherein in the first stage the desorption takes place exclusively by expansion and heating. The absorbent which, after this first stage, still has a considerable residual charge of C ₄ -hydrocarbons, is subjected to a second desorption stage, in which the remaining C ₄ -hydrocarbons are removed by combining expansion and heating and stripping. Suitable stripping gases are water vapor, air, nitrogen, carbon oxides or flue gases. Is stripped with steam, so the desorption is followed by a phase separation, in which an aqueous phase and the organic absorbent phase are obtained. The aqueous phase can be re-evaporated and reused as stripping medium. The regenerated absorbent from the second desorption stage is returned to the absorption stage. The exhaust gas of the second desorption stage can be recycled to the first or second dehydrogenation stage.

The separation D is generally not completely complete, so that in the C ₄ -PrO- duktgasstrom - depending on the nature of the separation - still small amounts or even traces of other gas components, in particular the low-boiling hydrocarbons, may be present.

The volume flow reduction also caused by the separation D relieves the subsequent process steps.

The exhaust gas containing the remaining gas constituents, which also contains oxygen, can be used, for example, for the regeneration of the catalyst bed in the first or second dehydrogenation zone.

The C ₄ -PrO duktgasstrom d consisting essentially of n-butane, 2-butene and butadiene generally contains 20 to 80% by volume of butadiene, 20 to 80% by volume of n-butane, 0 to 50% by volume. % 2-butene and 0 to 20% by volume of 1-butene.

The product gas stream c or d may still contain small amounts of oxygen. If the product gas stream c or d contains more than just slight traces of oxygen, a process stage for removing residual oxygen from the product gas stream c or d is generally carried out. The residual oxygen may have a disturbing effect insofar as it is used in downstream process steps as an initiator for Polymerization reactions can act. This danger is particularly present in the distillative removal of butadiene (step E)) and can lead to deposits of polymers (formation of so-called "popcorn") in the extractive distillation column Thereby, the oxygen removal is preferably carried out immediately after the oxidative dehydrogenation. In general, a catalytic combustion stage is carried out in which oxygen is reacted with the hydrogen contained in the gas stream c, c 'or d in the presence of a catalyst the absorption / desorption step, that is carried out only in the gas stream d, hydrogen must be added to the gas stream d.

A suitable catalyst for the oxidation of hydrogen supported on alpha alumina, 0.01 to 0.1 wt% platinum and 0.01 to 0.1 wt% tin based on the total weight of the catalyst. Platinum and tin are advantageously used in a weight ratio of 1: 4 to 1: 0.2, preferably in a ratio of 1: 2 to 1: 0.5, in particular in a ratio of approximately 1: 1. Advantageously, the catalyst contains 0.05 to 0.09 wt .-% platinum and 0.05 to 0.09 wt .-% tin based on the total weight of the catalyst. In addition to platinum and tin, it is optionally possible to use alkali metal and / or alkaline earth metal compounds in amounts of less than 2% by weight, in particular less than 0.5% by weight. Most preferably, the alumina catalyst contains only platinum and tin. The catalyst support of α-aluminum oxide advantageously has a BET surface area of 0.5 to 15 m ² / g, preferably 2 to 14 m ² / g, in particular 7 to 11 m ² / g. The carrier used is preferably a shaped body. Preferred geometries are, for example, tablets, ring tablets, spheres, cylinders, star strands or gear-shaped strands with diameters of 1 to 10 mm, preferably 2 to 6 mm. Particularly preferred are balls or cylinders, in particular cylinders.

Altematiwerfahren for removing residual oxygen from the product gas stream c, c 'or d include contacting the product gas stream with a mixture of metal oxides, which contains copper in the oxidation state 0 in a reduced form. In addition, such a mixture generally still contains aluminum oxides and zinc oxides, wherein the copper content is usually up to 10 wt .-%. In this way, almost complete separation of residual oxygen is possible. In addition, other methods of removing traces of oxygen can be used. Examples are the separation by means of molecular sieves using membranes, or by contacting with a NaNO ₂ solution. In a process part E, the C ₄ product gas stream d is separated by means of extractive distillation into a stream e 1 consisting essentially of n-butane and 2-butene and a product stream e 2 consisting essentially of butadiene. The stream e1 consisting essentially of n-butane and 2-butene is at least partially recycled to the first dehydrogenation stage.

Extractive distillation may be carried out as described in Petroleum and Coal Natural Gas Petrochemistry, Vol. 34 (8), pp. 343-346 or Ullmanns Enzyklopadie der Technischen Chemie, Vol. 9, 4th edition 1975, pages 1-18.

For this purpose, the C ₄ product gas stream d is brought into contact with an extraction agent, preferably an N-methylpyrrolidone (NMP) / water mixture, in an extraction zone. The extraction zone is generally carried out in the form of a wash column which contains trays, fillers or packings as internals. This generally has 30 to 70 theoretical plates, so that a sufficiently good release effect is achieved. Preferably, the wash column has a backwash zone in the column head. This backwash zone serves to recover the extractant contained in the gas phase by means of liquid hydrocarbon reflux, to which the top fraction is condensed beforehand. As a liquid hydrocarbon reflux and the process to be supplied fresh butane can be used. Typical temperatures at the top of the column are between 30 and 60 ^° C. The mass ratio of extractant to C ₄ product gas stream d in the feed of the extraction zone is generally from 10: 1 to 20: 1.

Suitable extractants are butyrolactone, nitriles such as acetonitrile, propionitrile, methoxypropionitrile, ketones such as acetone, furfural, N-alkyl-substituted lower aliphatic acid amides such as dimethylformamide, diethylformamide, dimethylacetamide, diethylacetamide, N-formylmorpholine, N-alkyl-substituted cyclic acid amides (lactams) such as N-alkylpyrrolidones , in particular N-methylpyrrolidone (NMP). In general, alkyl-substituted lower aliphatic acid amides or N-alkyl substituted cyclic acid amides are used. Particularly advantageous are dimethylformamide, acetonitrile, furfural and in particular NMP.

However, mixtures of these extractants with each other, for. As of NMP and acetonitrile, mixtures of these extractants with cosolvents and / or tert-butyl ether, for. As methyl tert-butyl ether, ethyl tert-butyl ether, propyl tert-butyl ether, n- or iso-butyl tert-butyl ether can be used. Especially suitable is NMP, preferably in aqueous solution, preferably with 0 to 20 wt .-% water, more preferably with 7 to 10 wt .-% water, in particular with 8.3 wt .-% water.

In the extractive distillation column is a gaseous, consisting essentially of n-butane and 2-butene stream e1, which is generally withdrawn via the top of the column, and obtained as Seitenabzugsstrom a mixture of extractant and butadiene. From this mixture butadiene can be subsequently obtained as a pure product. The bottom withdrawing stream is the extractant, which still contains butadiene and optionally secondary components (impurities). The bottom draw stream is recycled, optionally after carrying out further purification steps, back into the extractive distillation.

In general, the stream e1 consisting essentially of n-butane and 2-butene contains from 50 to 100% by volume of n-butane, from 0 to 50% by volume of 2-butene and from 0 to 3% by volume of further constituents, such as isobutane , Isobutene, propane, propene and C ₅ ⁺ hydrocarbons. The stream e1 is at least partially recycled to the first De hydrogenation.

For example, the extractive distillation, isolation of the pure butadiene and purification of the extractant can be carried out as follows: the side draw stream of the extractive distillation column of extractant and butadiene, which also contains impurities (acetylene, propyne, 1, 2-butadiene), is fed into a wash column, which is charged with fresh extractant. Crude butadiene, which contains, for example, 98% by weight of butadiene, is withdrawn from the column top of the wash column. The bottom draw stream is enriched with acetylene and is recycled to the extractive distillation. The crude butadiene may contain as impurities propyne and 1,2-butadiene. To remove these impurities, the crude butadiene is fed to a first purifying distillation column and a butadiene stream enriched with propyne is separated off at the top. The bottom draw stream, which is essentially propyne-free, but still contains traces of 1, 2-butadiene, is fed to a second pure distillation column, in which a substantially 1, 2-butadiene-free pure butadiene stream having a purity of, for example, at least 99, 6 wt .-% as top draw stream or side draw stream in the enrichment section of the column and a 1,2-butadiene-enriched bottom draw stream can be obtained. For the purification of the extractant, the extractant is partially or completely discharged as a bottom draw stream from the extractive distillation and regenerated as follows: The extraction solution is transferred to a desorption zone with respect to the extraction zone reduced pressure and / or elevated temperature, wherein from the extraction solution butadiene and existing acetylene traces be desorbed. The desorption zone can be designed, for example, in the form of a wash column which has 5 to 15, preferably 8 to 10 theoretical stages and a backwash zone with, for example, 4 theoretical stages. This backwashing zone is used to recover the extraction agent contained in the gas phase by means of liquid hydrocarbon reflux, to which the top fraction is condensed beforehand, or by adding water. As internals packings, trays or packing are provided. The pressure at the top of the column is, for example, 1.5 bar. The temperature in the bottom of the column is for example 130 to 150 ° C. At the bottom of the column a substantially acetylene-free extractant is obtained, which is recycled to the extractive distillation column. The top draw stream of the wash column is returned to the bottom of the extractive distillation column. The bottom draw stream of the wash column is returned to the top of the extractive distillation column. Alternatively, the desorption zone can also be operated at the same pressure as the extractive distillation.

The desired product stream e2, as obtained, for example, as the top draw stream of the second pure distillation column, can contain up to 100% by volume of butadiene.

The invention is further illustrated by the following example.

example

An n-butane containing feed gas stream (4) obtained by combining a fresh gas stream (1) and a recycle stream (15) is fed to the first autothermally operated n-butane non-oxidative catalytic dehydrogenation (BDH) stage (20) , To provide the heat necessary for the endothermic dehydrogenation, the hydrogen formed in the dehydrogenation is selectively burned, to which combustion air is supplied as stream (2). In order to counteract coking of the catalyst and to extend the service life of the catalyst also water vapor (3) is added. A dehydrogenation gas mixture (5) which, after leaving the non-oxidative n-butane dehydrogenation stage (20), is obtained. is cooled and the second, oxidative n-butane dehydrogenation (ODH) (21) is supplied. The ODH is further supplied with an air stream (6). For BDH and ODH, the conversion ratios and selectivities shown in Table 1 were assumed based on experimental results.

Table 1

The exit gas (7) of the ODH is compressed in two stages with intermediate cooling. The aqueous condensate (8) produced during the intermediate cooling is discharged from the process. From the aqueous condensate (8), an organic, C ₄ - hydrocarbons-containing phase can be separated and fed back to the first dehydrogenation. The compressed, butadiene-containing gas (9) is fed to an absorption stage (23), which is operated with tetradecane as an absorbent. In the absorption stage, an absorbent stream (11) loaded with the C ₄ hydrocarbons and an inert gas stream (10) discharged from the process are obtained. From the loaded absorbent (11) a butadiene, n-butane and 2-butene-containing stream (13) is separated in the desorption column (24), wherein the absorbent (12) is recovered. This is stripped in a stripping column (25) with air (18) to remove residues of C ₄ hydrocarbons. The gaseous top draw stream (19) of the stripping column (25) consisting of air and C ₄ hydrocarbons is fed to the ODH (21). The stream (17) of the regenerated absorbent is supplemented by fresh absorbent (17a) and returned to the absorption stage (23). The stream (13) is compressed and fed to a reactor (26) in which the residual oxygen remaining in the stream (13) is catalytically converted to water with hydrogen (13a). The substantially oxygen-free C ₄ -hydrocarbon stream (14) is used in an extractive distillation stage (27) using aqueous N-methylpyrrolidone Solution in a butadiene-containing product stream (16) and a n-butane-containing recycle stream (15), which is recycled to the BDH, separated.

The results of the simulation are shown in Table 2. The specification of the composition of the streams (1) to (19) is in mass proportions.

table

Continued table

Claims

claims

1. A process for producing butadiene from n-butane by the steps

A) providing a n-butane containing feed gas stream a;

B) feed of n-butane-containing feed gas stream a into at least one first dehydrogenation zone and non-oxidative catalytic dehydrogenation of n-butane, wherein a product gas stream b containing n-butane, 1-butene, 2-butene, butadiene, hydrogen, low-boiling

Minor constituents, optionally carbon oxides and optionally water vapor is obtained;

C) feeding the product gas stream b of the non-oxidative catalytic dehydrogenation and an oxygen-containing gas into at least one second dehydrogenation zone and oxidative dehydrogenation of n-butane, 1-butene and 2-butene, wherein a product gas stream c containing n-butane, 2-butene, Butadiene, low-boiling minor components, carbon oxides and water vapor is obtained, which has a higher content of butadiene than the product gas stream b;

D) separating the low-boiling minor constituents and water vapor to obtain a C ₄ product gas stream d consisting essentially of n-butane, 2-butene and butadiene;

F) recycling the stream e1 into the first dehydrogenation zone.

2. The method according to claim 1, characterized in that the non-catalytic dehydrogenation of n-butane is carried out autothermally with the introduction of an oxygen-containing gas.

3. The method according to claim 2, characterized in that air is fed as the oxygen-containing gas.

4. The method according to any one of claims 1 to 3, characterized in that the n-butane-containing feed stream a from liquefied petroleum gas (LPG) is obtained.

5. The method according to any one of claims 1 to 4, characterized in that in step D, the separation of water vapor by condensation by cooling and / or compressing the product gas stream c takes place in one or more cooling and / or compression stages, wherein a low-water product gas stream c 'is received.

6. The method according to claim 5, characterized in that in step D, the product gas stream c 'is brought into contact with an inert absorbent, wherein a C ₄ -hydrocarbons laden absorbent and the other gas constituents containing exhaust gas are obtained, and in a

Desorption the C ₄ hydrocarbons are released from the absorbent again.

7. The method according to any one of claims 1 to 6, characterized in that in step E, the extractive distillation is carried out with an N-methylpyrrolidone / water mixture as the extraction agent.

8. The method according to any one of claims 1 to 7, characterized in that after step C or step D in the product gas stream c or d remaining oxygen is removed.

9. The method according to claim 8, characterized in that the oxygen is removed by being catalytically reacted with hydrogen.