EP1858831A2

EP1858831A2 - Method for producing propene from propane

Info

Publication number: EP1858831A2
Application number: EP06708626A
Authority: EP
Inventors: Sven Crone; Otto Machhammer; Götz-Peter SCHINDLER; Frieder Borgmeier
Original assignee: BASF SE
Current assignee: BASF SE
Priority date: 2005-03-08
Filing date: 2006-03-03
Publication date: 2007-11-28
Also published as: DE102005010586A1; NO20074272L; MX2007010515A; KR20070110080A; WO2006094938A2; CN101137605A; JP2008532975A; CA2599255A1; WO2006094938A3; BRPI0608337A2; US20080269536A1; AU2006222065A1; TW200640847A

Abstract

The invention concerns a method for producing propene from propane, said method consisting in: A) preparing an input gas stream (a) containing propane; B) introducing the input gas stream (a), containing propane, optionally water vapour and optionally an oxygen-containing gas stream, into a dehydrogenation zone then dehydrogenating propane to form propene, a product gas stream (b), containing propane, propene, methane, ethane, ethene, nitrogen, carbon monoxide, carbon dioxide, water vapour, optionally hydrogen and optionally oxygen, being obtained; C) cooling and optionally compressing the product gas stream (b) then separating water vapour by condensing a product gas stream (c) depleted in water vapour (c) being obtained; D) separating the non-condensable gas constituents or with low boiling point by contacting the product gas stream (c) with an inert absorbent then by desorbing the gases dissolved in said inert absorbent, a C3 hydrocarbon stream (d1) and a residual gas stream (d2) containing methane, ethane, ethene, nitrogen, carbon monoxide, carbon dioxide, optionally hydrogen and optionally oxygen, being obtained; E) cooling and optionally compressing the C3 hydrocarbon stream (d1), a gaseous or liquid C3 hydrocarbon stream (e1) being obtained; F) optionally introducing the C3 hydrocarbon stream (e1) into a first distillation zone then separating same by distillation into a propane and propene stream (f1) and a stream (f2) containing ethane and ethene; G) introducing the stream (e1) or (f1) into a (second) distillation zone then separating by distillation into a propene stream (g1) and a propane stream (g2), the stream (g2) being recycled at least partly to the dehydrogenation zone.

Description

Process for the preparation of propene from propane

description

The invention relates to a process for the preparation of propene from propane.

Propene is obtained industrially by dehydrogenation of propane.

In the process known as UOP Oleflex process for the dehydrogenation of propane to propene propane-containing feed gas stream is preheated to 600-700 ⁰ C and dehydrogenated in a moving bed dehydrogenation reactor on a catalyst containing platinum on alumina, wherein a predominantly propane Propene and hydrogen containing product gas stream is obtained. In addition, low-boiling hydrocarbons formed by cracking (methane, ethane, ethene) and small amounts of high boilers (C ₄ ⁺ hydrocarbons) are contained in the product gas stream. The product gas mixture is cooled and compressed in several stages. The C ₂ and C ₃ hydrocarbons and the high boilers of hydrogen and methane formed during the dehydrogenation are then separated off by condensation in a so-called "cold box." The liquid hydrocarbon condensate is then separated by distillation, the first column being filled with hydrogen peroxide C ₂ - hydrocarbons and remaining methane separated and separated in a second distillation column, the C ₃ hydrocarbon stream in a propene fraction with a high degree of purity and a propane fraction, which also contains the C ₄ ⁺ -hydrocarbons .. A disadvantage of this method is the loss of C ₃ hydrocarbons by condensing out in the "cold box". Because of the large amounts of hydrogen formed in the dehydrogenation and due to the phase equilibrium, larger amounts of C ₃ hydrocarbons are discharged with the hydrogen / methane exhaust gas flow, if not condensed out at very low temperatures. Must be so carried out at temperatures from -20 to -60 ⁰ C to the loss of C ₃ hydrocarbons to limit, which are discharged with the hydrogen / methane gas stream.

The object of the invention is to provide an improved process for the dehydrogenation of propane to propene. The problem is solved by a process for producing propene from propane with the steps

A) a propane-containing feed gas stream a is provided;

B) the propane-containing feed gas stream a, optionally an oxygen-containing gas stream and optionally water vapor are fed into a dehydrogenation zone and propane is subjected to dehydrogenation to propene, wherein a product gas stream b containing propane, propene, methane, ethane, ethene, carbon monoxide, carbon dioxide, water vapor, optionally hydrogen and optionally oxygen is obtained;

C) the product gas stream b is cooled, optionally compressed and it is separated by condensation to separate water vapor, wherein a water vapor depleted product gas stream c is obtained;

D) non-condensable or low boiling gas components are separated by contacting the product gas stream c with an inert absorbent and then desorbing the gases dissolved in the inert absorbent, wherein a C ₃ hydrocarbon stream d1 and a

Exhaust stream d2 containing methane, ethane, ethene, nitrogen, carbon monoxide, carbon dioxide, optionally hydrogen, optionally oxygen and optionally propane and propene are obtained;

E) the C ₃ -hydrocarbon stream d1 is cooled and optionally compressed to obtain a gaseous or liquid C ₃ -hydrocarbon stream e1;

F) the C ₃ -hydrocarbon stream e1 is optionally fed into a first distillation zone and by distillation into a stream f1 of propane and

Propene and a stream f2 containing ethane and ethene separated;

G) the stream e1 or f1 is fed into a (second) distillation zone and separated by distillation into a product stream gl from propene and a stream g2 from propane, wherein the stream g2 at least partially into the

Dehydrogenation is recycled. In a first process part A), a propane-containing feed gas stream a is provided. This generally contains at least 80% by volume of propane, preferably 90% by volume of propane. In addition, the propane-containing feed gas stream A generally still contains butanes (n-butane, iso-butane). Typical compositions of the propane-containing feed gas stream are disclosed in DE-A 102 46 119 and DE-A 102 45 585. Usually, the propane-containing feed gas stream a is obtained from liquid petroleum gas (LPG). The propane-containing feed gas stream may be subjected to a purification distillation to remove the butanes, whereby a feed gas stream a having a very high propane content (> 95% by volume) is obtained.

In a process part B), the propane-containing feed gas stream is fed to a dehydrogenation zone and subjected to a generally catalytic dehydrogenation. In this case, propane is partially dehydrogenated to propene in a dehydrogenation reactor on a dehydrogenating catalyst. In addition, hydrogen and small amounts of methane, ethane, ethene and C ₄ ⁺ hydrocarbons (n-butane, isobutane, butenes, butadiene) are obtained. In addition, carbon oxides (CO, CO ₂ ), in particular CO ₂ , water vapor and, if appropriate, small amounts of inert gases in the product gas mixture of the catalytic propane dehydrogenation generally accumulate. The dehydrogenation product gas stream generally contains water vapor which has already been added to the dehydrogenation gas mixture and / or, upon dehydrogenation in the presence of oxygen (oxidative or non-oxidative), is formed on dehydrogenation. Inert gases (nitrogen) are introduced into the dehydrogenation zone when the dehydrogenation is carried out in the presence of oxygen with the oxygen-containing gas stream fed in, provided that no pure oxygen is fed in. In addition, unreacted propane is present in the product gas mixture.

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

The dehydrogenation can be carried out as oxidative or non-oxidative dehydrogenation. The dehydration can be performed isothermally or adiabatically. The dehydrogenation can be carried out catalytically in a fixed bed, moving bed or fluidized bed reactor. The non-oxidative catalytic propane dehydrogenation can be carried out autothermally with the introduction of oxygen. However, it can also be carried out purely catalytically without supply of oxygen. During the autothermal dehydrogenation, oxygen is added to the reaction gas mixture of the propane dehydrogenation in at least one reaction zone and the hydrogen and / or hydrocarbon contained in the reaction gas mixture is at least partially combusted, thereby producing at least part of the required dehydrogenation heat in the at least one reaction zone directly in the reaction gas mixture becomes. The oxygen-containing gas used is air or oxygen-enriched air having an oxygen content of up to 70% by volume, preferably up to 50% by volume.

A feature of the non-oxidative mode of operation over an oxidative mode of operation is that free hydrogen is still present at the outlet of the dehydrogenation zone. In oxidative dehydrogenation, there is no free hydrogen at the outlet of the dehydrogenation zone.

A suitable reactor form is the fixed bed tube or tube bundle reactor. In these, the catalyst (dehydrogenation catalyst and optionally special oxidation catalyst) is a fixed bed in a reaction tube or in a bundle of reaction tubes. Typical reaction tube internal diameters are about 10 to 15 cm in the case of non-autothermal, purely catalytic dehydrogenation. A typical Dehydrierrohrbündelreaktor comprises about 300 to 1000 reaction tubes. The temperature inside the reaction tube usually moves in the range of 300 to 1200 ° C, preferably in the range of 500 to 1000 ° C. The working pressure is usually between 0.5 and 12 bar, often between 1 and 8 bar when using a low water vapor dilution, but also between 3 and 8 bar when using a high steam dilution (according to the so-called "steam active reforming process" (STAR) Process) or the Linde process) for the dehydrogenation of propane or butane by Phillips Petroleum Co. Typical catalyst loadings (GHSV) are 500 to 2000 h ^-1 , based on the hydrocarbon used The catalyst geometry can be, for example, spherical or cylindrical (hollow or full). be.

The propane dehydrogenation can also be carried out under heterogeneous catalysis in a fluidized bed, according to the Snamprogetti / Yarsintez-FBD process. 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 dehydration temperature usually 550 to 650 ° C. The heat required for the dehydrogenation can be introduced into the reaction system in that 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 is generated directly in the reactor system by combustion of hydrogen and / or hydrocarbons in the presence of oxygen. Optionally, a hydrogen-containing co-feed may additionally be admixed.

The propane dehydrogenation can be carried out in a tray reactor. If the dehydrogenation is carried out autothermally with the introduction of an oxygen-containing gas stream, it is preferably carried out in a tray reactor. This contains one or more successive 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 are arranged in a shaft furnace reactor axially or in the annular gaps of concentrically arranged cylindrical gratings. A shaft furnace reactor corresponds to a horde reactor with only one horde. The performance of dehydrogenation in a single shaft furnace reactor corresponds to one embodiment. In a further preferred embodiment, the dehydrogenation is carried out in a tray reactor with 3 catalyst beds.

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 propane is generated by the combustion of hydrogen present in the reaction gas mixture and optionally of hydrocarbons present in the reaction gas mixture and / or of coke present in the form of coke , In general, the total amount of oxygen fed, based on the total amount of propane, is 0.001 to 0.5 mol / mol, preferably 0.001 to 0.4 mol / mol, particularly preferably 0.02 to 0.35 mol / mol. Oxygen is used as the oxygen-containing gas containing inert gases, for example, air or oxygen-enriched air.

The hydrogen burned to generate heat is the hydrogen formed during the catalytic propane dehydrogenation and, if appropriate, the hydrogen gas additionally added to the reaction gas mixture. Preferably should so much hydrogen be present that the molar ratio H ₂ IO ₂ in the reaction gas mixture immediately after the feed of the oxygen-containing gas 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 catalyzes both the combustion of the hydrocarbons and of hydrogen with oxygen, so that basically no specific oxidation catalyst different from this is required. In one embodiment, one operates in the presence of one or more oxidation catalysts that selectively catalyze the combustion of hydrogen with oxygen to water 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. The dehydrogenation catalyst and the oxidation catalyst may be present together in one or more reaction zones or separately 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 optionally 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 dehydration temperature is generally 400 to 1100 ° C, the pressure in the last Catalyst bed of the tray reactor generally 0.2 to 15 bar, preferably 1 to 10 bar, more preferably 1 to 5 bar. The load (GHSV) is generally 500 to 2000 h ^{~ 1} , in high load mode also up to 100 000 h ^{~ 1} , preferably 4000 to 16 000 h \

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.

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.

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 may comprise one or more elements of main group I and / or II, preferably potassium and / or cesium. Furthermore, the dehydrogenation catalysts may contain one or more elements of III. 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. Suitable shaped catalyst body geometries are strands, stars, rings, saddles, spheres, foams and monoliths with characteristic dimensions of 1 to 100 mm. 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 III. 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 propane dehydrogenation are the catalysts according to Examples 1, 2, 3 and 4 of DE-A 199 37 107.

The autothermal propane dehydrogenation is preferably carried out in the presence of steam. 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 and carbon dioxide. Dilution with water vapor shifts the equilibrium to the products of dehydration.

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. Optionally, the catalyst is reduced after regeneration with a hydrogen-containing gas.

The product gas stream b can be separated into two partial streams, with a partial stream being returned to the autothermal dehydrogenation, in accordance with the cycle gas method described in DE-A 102 11 275 and DE-A 100 28 582.

The propane dehydrogenation can be carried out as an oxidative dehydrogenation. The oxidative propane dehydrogenation can be carried out as a homogeneous oxidative dehydrogenation or as a heterogeneously catalyzed oxidative dehydrogenation.

If the propane dehydrogenation is designed as a homogeneous oxydehydrogenation in the context of the process according to the invention, it can be prepared in principle as described in US-A 3,798,283, CN-A 1, 105,352, Applied Catalysis, 70 (2), 1991, p to 187, Catalysis Today 13, 1992, pp. 673 to 678 and the earlier application DE-A 1 96 22 331.

The temperature of the homogeneous oxydehydrogenation is generally from 300 to 700 ⁰ C, preferably from 400 to 600 ⁰ C, particularly preferably of 400 to 500 ⁰ C. The pressure may be from 0.5 to 100 bar or from 1 to 50 bar. Often it will be at 1 to 20 bar, especially at 1 to 10 bar.

The residence time of the reaction gas mixture under oxydehydrogenation conditions is usually 0.1 or 0.5 to 20 seconds, preferably 0.1 or 0.5 to 5 seconds. a shell-and-tube reactor may be used, such as a shell-and-tube reactor with molten salt as the heat carrier, or a shaft furnace reactor with intercooling.

The propane to oxygen ratio in the starting mixture to be used can be 0.5: 1 to 40: 1. The molar ratio of propane to molecular oxygen in the starting mixture is preferably <6: 1, preferably <5: 1. In general, the abovementioned ratio will be> 1: 1, for example> 2: 1. The starting mixture may comprise further, essentially inert constituents, such as H ₂ O, CO ₂ , CO, N ₂ , noble gases and / or propene. It is favorable for a homogeneous oxidative dehydrogenation of propane to propene if the ratio of the surface of the reaction space to the volume of the reaction space is as small as possible. Particularly favorable surface materials are aluminum oxides, quartz glass, borosilicates, stainless steel and aluminum.

If, in the context of the process according to the invention, the first reaction stage is designed as a heterogeneously catalyzed oxydehydrogenation, this can in principle be carried out as described in US-A 4,788,371, CN-A 1073893, Catalysis Letters 23 (1994) 103-106, W. Zhang, Gaodeng Xuexiao Huaxue Xuebao, 14 (1993) 566, Z. Huang, Shiyou Huagong, 21 (1992) 592, WO 97/36849, DE-A 1 97 53 817, US-A 3,862,256, US-A 3,887,631, DE-A 1 95 30 454, US Pat. No. 4,341,664, J. of Catalysis 167, 560-569 (1997), J. of Catalysis 167, 550-559 (1997), Topics in Catalysis 3 (1996) 265-275, US Pat. A 5,086,032, Catalysis Letters 10 (1991) 181-192, Ind. Eng. Chem. Res. 1996, 35, 14-18, US-A 4,255,284, Applied Catalysis A: General, 100 (1993) 111-130, J. of Catalysis 148, 56-67 (1994), V. Cortes Corberan and S Vic Bellόn (Editors), New Developments in Selective Oxidation II, 1994, Elsevier Science BV, pp. 305-313, 3rd World Congress on Oxidation Catalysis RK Grasselli, ST. Oyama, AM Gaffney and JE Lyons (Editors), 1997, Elsevier Science BV, p. 375 ff. In particular, all of the abovementioned oxydehydrogenation catalysts can be used. The said for the aforementioned writings also applies to:

a) Otsuka, K .; Uragami, Y .; Komatsu, T .; Hatano, M. in Natural Gas Conversion,

Stud. Surf. Be. Catal .; Holmen A .; Jens, K.-J .; Kolboe, S., Eds .; Elsevier

Science: Amsterdam, 1991; Vol. 61, p 15;

b) Seshan, K .; Swaan, H.M .; Smits, R.H.H .; van Ommen, J.G .; Ross, J.R.H. in New Developments in Selective Oxidation; Stud. Surf. Be. Catal .; Centi, G .; Trifirö, F., Eds .; Elsevier Science: Amsterdam 1990; Vol. 55, p 505;

c) Smits, R.H.H .; Seshan, K .; Ross, J.R.H. in New Developments in Selective Oxidation by Heterogeneous Catalysis; Stud. Surf. Be. Catal .; Ruiz, P .; Delmon, B.,

eds .; Elsevier Science: Amsterdam, 1992 a; Vol. 72, p 221;

d) Smits, R.H.H .; Seshan, K .; Ross, J.R.H. Proceedings, Symposium on Catalytic Selective Oxidation, Washington DC; American Chemical Society: Washington, DC, 1992b; 1121;

e) Mazzocchia, C; Aboumrad, C; Daigne, C; Tempesti, E .; Herrmann, J. M .; Thomas, G. Catal. Lett. 1991, 10, 181;

f) Bellusi, G .; Conti, G .; Perathonar, S .; Trifirö, F. Proceedings, Symposium on Catalytic Selective Oxidation, Washington, DC; American Chemical Society: Washington, DC, 1992; p 1242;

g. Ind. Eng. Chem. Res. 1996, 35, 2137-2133 and

h) Symposium on Heterogeneous Hydrocarbon Oxidation Presented before the Division of Petroleum Chemistry, Inc. 211 th National Meeting, American Chemical Society, New Orleans, LA, March 24-29, 1996.

Particularly suitable oxydehydrogenation catalysts are the multimetal oxide materials or catalysts A of DE-A 1 97 53 817, with those mentioned as being preferred Multimetal oxide materials or catalysts A are particularly favorable. That is, as active compounds in particular multimetal oxide materials of the general formula I.

M ¹ _a M _Oi _ _b M ² _b O _x (I) ₁

With

M ¹ = Co, Ni, Mg, Zn, Mn and / or Cu,

M ² = W, V, Te, Nb, P, Cr, Fe, Sb, Ce, Sn and / or La, a = 0.5 to 1.5, b = 0 to 0.5 and x = a number which is determined by the valence and frequency of the elements other than oxygen in I,

into consideration.

Further multimetal oxide compositions suitable as oxydehydrogenation catalysts are mentioned below:

Suitable Mo-V-Te / Sb-Nb-O multimetal oxide catalysts are described in EP-A 0 318 295, EP-A 0 529 853, EP-A 0 603 838, EP-A 0 608 836, EP-A 0 608 838 EP-A 0 895 809, EP-A 0 962 253, EP-A 1 192 987, DE-A 198 35 247, DE-A 100 51 419 and DE-A 101 19 933.

Suitable Mo-V-Nb-O multimetal oxide catalysts are described inter alia in EM Thorsteinson, TP Wilson, FG Young, PH Kasei, Journal of Catalysis 52 (1978), pages 116-132 and in US 4,250,346 and EP-A 0 294 845 ,

Suitable Ni-X-O-Multimetalloxidkatalysatoren, with X = Ti, Ta, Nb, Co, Hf, W, Y, Zn, Zr, Al, are described in WO 00/48971.

In principle, suitable active compounds can be prepared in a simple manner by producing as intimate as possible, preferably finely divided, stoichiometrically composed dry mixtures of suitable sources of their components and calcining them at temperatures of 450 to 1000 ° C. The calcination can be carried out both under inert gas and under an oxidative atmosphere such as air (mixture of inert gas and oxygen) as well as under reducing atmosphere (eg mixture of inert gas, oxygen and NH ₃ , CO and / or H ₂ ) take place. Suitable sources of the components of the multimetal oxide active compounds I are oxides and / 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 complex salts, ammonium salts and / or hydroxides.

The multimetal oxide compositions can be used for the process according to the invention in shaped both in powder form and to specific catalyst geometries, wherein the shaping can take place before or after the final calcination. Suitable unsupported catalyst geometries are e.g. 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. Suitable hollow cylinder geometries are e.g. 7mm x 7mm x 4mm or 5mm x 3mm x 2mm or 5mm x 2mm x 2mm (each length x outside diameter x inside diameter). Of course, the full catalyst may also have spherical geometry, wherein the ball diameter may be 2 to 10 mm.

Of course, the shaping of the powdered active composition or its powdery, not yet calcined precursor composition can also be effected by application to preformed inert catalyst supports. The layer thickness of the powder mass applied to the carrier body is expediently chosen in the range from 50 to 500 mm, preferably in the range from 150 to 250 mm. 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. The carrier bodies may be regularly or irregularly shaped, with regularly shaped carrier bodies having a marked surface roughness, e.g. Spheres, hollow cylinders or saddles, generally with dimensions in the range of 1 to 100 mm, are preferred. Suitable is the use of substantially nonporous, surface roughness, spherical steatite supports whose diameter is 1 to 8 mm, preferably 4 to 5 mm.

The reaction temperature for the heterogeneously catalyzed oxydehydrogenation of propane is generally from 300 to 600 ⁰ C, the usual manner from 350 to 500 ⁰ C. The pressure is from 0.5 to 10 bar, preferably 1 to 10 bar, for example 1 to 5 bar. pressures above 1 bar, eg 1, 5 to 10 bar, have proven to be particularly advantageous. In general, the heterogeneously catalyzed oxydehydrogenation of the propane takes place on a fixed catalyst bed. The latter is expediently poured into the tubes of a tube bundle reactor, as described, for example, in EP-A 700 893 and in EP-A 700 714 and in the literature cited in these publications. The average residence time of the reaction gas mixture in the catalyst bed is normally 0.5 to 20 seconds. The propane to oxygen ratio in the reaction gas starting mixture to be used for the heterogeneously catalyzed propane oxydehydrogenation may be 0.5: 1 to 40: 1. It is advantageous if the molar ratio of propane to molecular oxygen in the starting gas mixture is <6: 1, preferably <5: 1. In general, the aforementioned ratio will be> 1: 1, for example 2: 1. The starting gas mixture may comprise further, substantially inert constituents such as H ₂ O, CO ₂ , CO, N ₂ , noble gases and / or propene. In addition, d, C ₂ or C ₄ hydrocarbons may also be present to some extent.

In the propane dehydrogenation, a gas mixture is obtained, which generally has the following composition: 5 to 95% by volume of propane, 1 to 50% by volume of propene, 0 to 20% by volume of methane, ethane, ethene and C ₄ ⁺ Hydrocarbons, 0 to 30% by volume of carbon oxides, 0 to 70% by volume of steam and 0 to 30% by volume of hydrogen and 0 to 70% by volume of inert gases.

In the autothermal propane dehydrogenation, a gas mixture is obtained which generally has the following composition: 10 to 80% by volume of propane, 1 to 40% by volume of propene, 0 to 20% by volume of methane, ethane, ethene and C. ₄ ⁺ -hydrocarbons, 0.1 to 30% by volume of carbon oxides, 0.1 to 70% by volume of steam, 1 to 30% by volume of hydrogen and 0 to 50% by volume of inert gases (in particular nitrogen) -

The product gas stream b when leaving the dehydrogenation zone is generally under a pressure of 1 to 20 bar, preferably 1 to 10 bar, more preferably 1 to 5 bar and has a temperature in the range of 400 to 700 ° C.

In process part C), steam is first separated off from the product gas stream b, with the result that a product gas stream c depleted in water vapor is obtained. The separation of water vapor is carried out by condensation by cooling and, if appropriate, prior compressing of the product gas stream b and can be carried out in one or more cooling and optionally compression stages be performed. In general, the product gas stream b is for this purpose to a temperature in the range of 0 to 80 ° C, preferably 10 to 65 ⁰ C cooled. In addition, the product gas stream can be compressed, for example to a pressure in the range of 2 to 40 bar, preferably 5 to 20 bar, particularly preferably 10 to 20 bar.

In one embodiment of the process according to the invention, the product gas stream b is passed through a cascade of heat exchangers and initially cooled to a temperature in the range of 50 to 200 ° C and then in a quench tower with water to a temperature of 40 to 80 ° C, for example 55 ° C further cooled. Most of the water vapor condenses here, but also some of the C ₄ ⁺ hydrocarbons contained in the product gas stream b, in particular the C ₅ ⁺ hydrocarbons. Suitable heat exchangers are, for example, direct heat exchangers and countercurrent heat exchangers, such as gas-gas countercurrent heat exchangers, and air coolers.

There is obtained a water vapor depleted product gas stream c. This generally contains 0 to 10 vol .-% water vapor. For practically complete removal of water from the product gas stream c, drying using molecular sieve or membranes can be provided when using certain solvents in step D).

In a process part D) are the non-condensable or low-boiling gas components such as hydrogen, oxygen, carbon monoxide, carbon dioxide, nitrogen and the low-boiling hydrocarbons (methane, ethane, ethene) in an absorption / desorption cycle by means of a high-boiling absorbent of the C ₃ hydrocarbons separated, wherein a stream d1 is obtained, which contains the C ₃ - hydrocarbons and also small amounts of ethane and ethene, and an exhaust stream d2 is obtained, which contains the non-condensable or low-boiling gas components.

For this purpose, in an absorption stage, the gas stream c is contacted with an inert absorbent, the C ₃ hydrocarbons and also small amounts of C ₂ hydrocarbons are absorbed in the inert absorbent and a loaded with C ₃ hydrocarbons absorbent and the others Gas components containing exhaust gas d2 are obtained. These are mainly carbon oxides, hydrogen, inert gases and C ₂ - Hydrocarbons and methane. Also certain amounts of propane and propene may still be included in the stream d2, since the separation is generally not quite complete. In a desorption step, the C ₃ hydrocarbons are released from the absorbent again.

Inert absorbent used in the absorption stage are generally high-boiling nonpolar solvents in which the separated C ₃ - hydrocarbon mixture has a significantly higher solubility than the other gas components to be separated. The absorption can be carried out by simply passing the stream c through the absorbent. But it can also be done in columns. It can be used in cocurrent, countercurrent or cross flow. Suitable absorption columns include plate columns having bubble-cap, valve, and / or sieve trays, columns with structured packings, for example fabric packings or sheet-metal packings with a specific surface area of 100 to 1000 m ² / m ³ as Mellapak ^® 250 Y, and packed columns, z. B. with balls, rings or saddles made of metal, plastic or ceramic as packing. However, there are also trickle and spray towers, graphite block absorbers, surface absorbers such as thick-film and thin-layer absorbers and bubble columns with and without internals into consideration. Preferably, the absorption column has an absorption part and a rectification part. To increase the accumulation of C ₃ hydrocarbons in the solvent in the manner of a rectification, heat can then be introduced into the bottom of the column. Alternatively, a stripping gas stream can be fed into the bottom of the column, for example from nitrogen, air, steam or propane / propene mixtures. Thus, in the rectification section of the absorption column, the laden absorbent is brought into contact with the stripping gas stream. Characterized ₂ ^~ hydrocarbons are stripped out from the laden absorbent C. A portion of the overhead can be condensed and returned to the top of the column to limit solvent losses.

Suitable absorbents are relatively nonpolar organic solvents, for example C ₄ -C ₈ -alkenes, naphtha or aromatic hydrocarbons, such as the paraffin distillation medium fractions, or bulky group ethers, or mixtures of these solvents, such as polar solvents such as 1, 2 and 3. 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 butanes, pentanes, hexanes, heptanes, octanes, nonanes, decanes, undecanes, dodecanes, tridecanes, tetradecanes, pentadecanes, hexadecanes, heptadecanes and octadecanes or fractions obtained from refinery streams containing as main components said linear alkanes. Preferred absorbents are C ₈ -C ₀ hydrocarbons, particularly preferred are Cg-hydrocarbons, in particular nonanes.

For desorption of the C ₃ hydrocarbons, the loaded absorbent is heated and / or expanded to a lower pressure. Alternatively, desorption may also be by stripping, usually with steam, or in a combination of

Relaxation, heating and stripping done in one or more steps. For example, the desorption can be carried out in two stages, wherein the second desorption stage is carried out at a lower pressure than the first desorption stage and the desorption gas of the second stage is returned to the absorption stage. The absorbent regenerated in the desorption stage is returned to the absorption stage. If necessary, a part of this

Absorbent stream, which may contain C ₄ ⁺ hydrocarbons, discharged, worked up and returned, or discarded.

In a variant of the method, the desorption step is carried out by relaxation and / or heating of the loaded absorbent. In a further process variant is additionally stripped with steam.

In a further process variant, the absorbent laden with C ₃ hydrocarbons is compressed prior to desorption to a pressure which in one of the following stages, for. B. the propane / propene separation (step G), is necessary.

The separation D) is generally not completely complete, so that in the C ₃ - hydrocarbon stream d1 - depending on the type of separation - still small amounts or even traces of other gas constituents, in particular the low-boiling hydrocarbons, may be present. To separate off the hydrogen contained in the exhaust gas stream d2, it may, if appropriate after cooling, for example in an indirect heat exchanger, be passed through a membrane, which is usually designed as a tube, which is permeable only to 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. Alternatively, the exhaust stream d2 may be burned.

In a method part E), the gas stream d1 is cooled, wherein it can additionally be compressed in one or more further compression stages. In this case, a gaseous C ₃ -hydrocarbon stream e1 or a liquid condensate stream e1 of C ₃ -hydrocarbons is obtained. The stream e1 may still contain small amounts of C ₂ - hydrocarbons. In addition, an aqueous condensate stream e2 and possibly small amounts of an exhaust gas stream e3 can be obtained. The aqueous condensate stream e2 generally accumulates when water is stripped in step D) to desorb the dissolved gases.

The compression can again be done in one or more stages. In general, a total pressure of from 1 to 29 bar, preferably from 1 to 10 bar, is compressed to a pressure in the range from 12 to 30 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 80 ⁰ C, preferably 15 to 60 ⁰ C, cooled. Subsequently, the compressed gas mixture is cooled to a temperature of -10 ° C to 60 ° C, preferably -10 ⁰ C to 30 ⁰ C. The liquid condensate flows e1 and e2 are separated in a phase separation apparatus.

However, the gas stream d1 can also only be cooled, preferably when the desorption of the dissolved gases in the process part D) takes place at high pressure.

In a process part F), the gaseous or liquid C ₃ - hydrocarbon stream e1 is fed into a first distillation zone and separated by distillation into a stream f1 containing the C ₃ hydrocarbons propane and propene and a stream f2 containing the C ₂ hydrocarbons ethane and ethene. For this, the C ₃ -hydrocarbon stream e1 is generally introduced into a C2 / C3 separation column typically having 20 to 80 theoretical plates, for example about 60 theoretical soils, fed. This is generally operated at a pressure in the range of 10 to 30 bar, for example at about 20 bar, and a reflux ratio of 2-70. The bottom temperature is generally from 40 to 100 ° C, for example, about 60 ° C, the head temperature of -20 to 10 ° C, for example, about 10 ⁰ C.

A stream f1 of propane and propene is obtained as the bottom draw stream having an ethane / ethene content of generally <5000 ppm, preferably <1000 ppm, more preferably <500 ppm. The stream f2, which is preferably obtained as a top draw stream, may still contain certain amounts of propane and propene and be recycled to the separation thereof in the absorption stage.

The process part F) can also be omitted, in particular if the stream d1 or e1 has only a low content of C ₂ hydrocarbons.

In a process part G), the C ₃ -hydrocarbon stream e1 or f1 is fed into a second distillation zone and separated by distillation into a stream gl containing propene and a stream g2 containing propane. For this purpose, the hydrocarbon stream f1 is generally fed into a C ₃ separation column ("C3 splitter") typically having from 80 to 150 theoretical plates, for example about 100 theoretical plates, which is generally at a pressure in the range from 10 to bar 30, for example at about 20 bar, and a reflux ratio of 2 -. operated 50 the bottom temperature is generally from 40 to 100 ° C, for example about 68 ⁰ C, the head temperature of 30 to 60 ⁰ C, for example ca. 60 ° C. Instead of a single C ₃ separating column, it is also possible to use two C ₃ separating columns, the first column being operated at a higher pressure, for example 25 bar, and the second column at a lower pressure, for example 18 bar (2 The head take off of the first column is liquefied in the sump heater of the second column and the bottom takeoff of the first column is fed into the second column Vapor compression possible.

In a process part H) the stream g2 and a fresh propane stream can be fed into a third distillation zone, in which a stream containing C ₄ ⁺ hydrocarbons is separated by distillation and the feed gas stream a is obtained with a very high propane content. The recycled stream g2 is vaporized before it enters the third distillation zone. In this case, a refrigerant flow be generated, which can be used for cooling elsewhere, for example, for cooling at the top of the column C2 / C3 separation column.

The invention is further illustrated by the following example.

example

The variant of the method according to the invention shown in the figure was simulated by calculation. The following process parameters were assumed.

It is assumed a capacity of the plant of 348 kt / a propene at 8000 h running time, corresponding to 43447 kg / h propene.

The fresh propane stream 1 contains about 98 wt .-% propane, about 2 wt .-% butane. The fresh propane stream 1 is mixed with the propane recycle stream 24 from the C3 splitter 37 and fed to the C3 / C4 separation column 30. In the C3 / C4 separation column 30, which has 40 theoretical stages and is operated at 10 bar operating pressure and a reflux ratio of 0.4, a high-boiler stream 4 is separated and thus a propane stream 3 with a butane content of only 0.01 Wt .-% received. The propane stream 3 is preheated to 450 ⁰ C, enters the dehydrogenation zone 31 and an autothermal dehydrogenation is subjected. For this purpose, an oxygen-containing gas 6 and water vapor 5 are fed into the dehydrogenation zone 31. The conversion of dehydrogenation is based on propane, 40%, the selectivity of the propene formation is 90%. In addition, 5% cracking products and 5% carbon oxides are formed by total combustion. The water concentration in the outlet gas of the dehydrogenation zone is about 11 wt .-%, the residual oxygen content in the outlet gas is 0 wt .-%, the outlet temperature of the product gas mixture is 600 ⁰ C. The product gas stream 7 is cooled and in the compressor 32, starting from a pressure of 2.0 bar compressed in 3 stages to a pressure of 15 bar. After the first and the second compressor stage is cooled to 55 ⁰ C in each case. This results in an aqueous condensate 9, which is discharged from the process. The compressed and cooled gas stream 8 is brought in the absorption column 33 with tetradecane 21 as an absorbent in contact. The unabsorbed gases are withdrawn as exhaust gas stream 11 via the top of the column, the absorbent laden with the C ₃ hydrocarbons is taken off via the bottom of the column and fed to the desorption column 34. In the Desorption column 34 are desorbed by depressurization to a pressure of 4 bar and stripping with 13 supplied as stream of high pressure steam, the C ₃ hydrocarbons, a stream 14 of regenerated absorbent and a stream 12 of C ₃ hydrocarbons and water vapor is obtained. The regenerated absorbent 14 is supplemented with fresh absorbent 22 and returned to the absorption column 33. At the top of the desorption column, the gas is cooled to 45 ° C, with further absorbent 14 condensed out. In addition, an aqueous phase is obtained, which is separated in a phase separator and discharged as stream 15 from the process. Subsequently, the stream 12 is compressed in two stages to a pressure of 16 bar and cooled to a temperature of 40 ⁰ C. In this case, a small waste gas stream 18, a wastewater stream 17 and a liquid C ₃ - fall to the hydrocarbon stream 16.

Of the liquid C ₃ hydrocarbon stream 16 is in a C2 / C3 separation column 36 with 30 theoretical stages at 16 bar and a reflux ratio of 63 on

Head a C ₂ -hydrocarbon stream 20, which also contains certain amounts of C ₃ -hydrocarbons, separated. The stream 20 is recycled to the absorption column 33 where C ₃ hydrocarbons contained in the stream 20 are separated. The bottom temperature in the C2 / C3 separation column 36 is 42 ⁰ C, the head temperature is -4 ⁰ C. The residual ethane content of the bottom draw stream 19 0.01

Wt .-%. The bottom draw stream 19 is fed to a propane / propene separation column with 120 theoretical plates, which at 16 bar with a reflux ratio of

22 is operated. The bottom temperature is 46 ° C, the head temperature 38 ° C. At the

At the top, a propene stream 23 having a purity of 99.5% by weight of propene is obtained. The bottom take-off stream 24 contains about 98.5% by weight of propane and is introduced into the

Dehydrogenation zone 31 returned.

table

Continued table

Claims

claims

1. A process for the preparation of propene from propane with the steps

A) a propane-containing feed gas stream a is provided;

B) the propane-containing feed gas stream a, optionally water vapor and optionally an oxygen-containing gas stream are fed into a dehydrogenation zone and propane is subjected to dehydrogenation to propene, wherein a product gas stream b containing propane, propene, methane, ethane, ethene, nitrogen, carbon monoxide, carbon dioxide, Water vapor if necessary

Hydrogen and optionally oxygen is obtained;

D) non-condensable or low-boiling gas components are by contacting the product gas stream c with an inert

Absorbing agent and subsequent desorption of the gases dissolved in the inert absorbent separated, wherein a C ₃ - hydrocarbon stream d1 and an exhaust stream d2 containing methane, ethane, ethene, nitrogen, carbon monoxide, carbon dioxide, optionally hydrogen and optionally oxygen are obtained;

E) the C ₃ -hydrocarbon stream d1 is cooled and optionally compressed to obtain a gaseous or liquid C ₃ - hydrocarbon stream e1;

F) the C ₃ -hydrocarbon stream e1 is optionally fed into a first distillation zone and separated by distillation into a stream f1 of propane and propene and a stream f2 containing ethane and ethene;

G) the stream e1 or f1 is fed into a (second) distillation zone and by distillation into a product stream gl from propene and a stream g2 separated from propane, wherein the stream g2 is at least partially recycled to the dehydrogenation zone.

2. The method according to claim 1, characterized in that the dehydrogenation in step B) is carried out as oxidative or non-oxidative dehydrogenation.

3. The method according to claim 1, characterized in that the dehydrogenation in step B) is performed adiabatically or isothermally.

4. The method according to claim 1, characterized in that the Deyhdrierung in step B) is carried out in a fixed bed reactor, moving bed reactor or fluidized bed reactor.

5. The method according to claim 1, characterized in that in step B) an oxygen-containing gas stream is fed.

6. The method according to claim 5, characterized in that the dehydrogenation is carried out as autothermal dehydrogenation or as oxydehydrogenation.

7. The method according to any one of claims 1 to 7, characterized in that in step C) the product gas stream b is cooled to a temperature in the range of 10 to 80 ⁰ C.

8. The method according to any one of claims 1 to 7, characterized in that in a step H) the stream g2 and fresh propane are fed into a third distillation zone and are separated by distillation into the feed gas stream a and a stream containing C ₄ ⁺ hydrocarbons.

9. The method according to any one of claims 1 to 8, characterized in that the absorbent used in step D) is selected from the group consisting of C ₄ -C ₈ -alkanes, naphtha and the middle oil fraction from the paraffin distillation.

10. The method according to any one of claims 1 to 9, characterized in that in step D) the product gas stream c in an absorption column with a

Absorption part and a rectification part with the inert absorbent in Contact and the loaded absorbent in the rectification is brought into contact with a stripping gas stream.

11. The method according to any one of claims 1 to 10, characterized in that in step D) for the desorption of the gases dissolved in the absorbent with

Steam is stripped.

12. The method according to any one of claims 1 to 11, characterized in that in step E) the product gas stream d is compressed to a pressure of 5 to 25 bar.

13. The method according to any one of claims 1 to 11, characterized in that in step D) before carrying out the desorption the loaded with C ₃ hydrocarbons absorbent is compressed to a pressure of 5 to 25 bar.

14. The method according to any one of claims 1 to 13, characterized in that in step E) the product gas stream d is cooled to a temperature in the range of -10 to 60 ⁰ C.

15. The method according to any one of claims 1 to 14, characterized in that additionally obtained in step E) an aqueous condensate stream e2 and separated in a phase separation apparatus of the liquid C ₃ hydrocarbon stream.

16. The method according to any one of claims 1 to 15, characterized in that the oxygen-containing gas stream used in step B) is used air or oxygen-enriched air having an oxygen content of up to 70 vol .-%.