EP1109773A1

EP1109773A1 - Method for producing acrolein and/or acrylic acid from propane

Info

Publication number: EP1109773A1
Application number: EP99941556A
Authority: EP
Inventors: Otto Machhammer; Andreas Tenten; Harald Jachow; Susanne Haupt; Heiko Arnold; Signe Unverricht
Original assignee: BASF SE
Current assignee: BASF SE
Priority date: 1998-08-19
Filing date: 1999-08-10
Publication date: 2001-06-27
Also published as: US6388129B1; DE19837518A1; JP2002523389A; CN1312789A; WO2000010960A1; AU5512599A; BR9914611A

Abstract

A method for producing acrolein and/or acrylic acid from propane, whereby the propane undergoes partial homogeneous and/or heterogeneous catalysed oxydehydrogenation with molecular oxygen in a first reaction step to form propane. The product gas mixture formed in the first reaction step is subsequently used to produce acrolein and/or acrylic acid by means of gas-phase catalytic propane oxidation. The method is characterised in that the molecular oxygen which is required in the first reaction step and is different from the recycled oxygen gas is added to the initial reaction gas mixture of the first reaction step as modified air containing nitrogen with the proviso that the nitrogen content of the modified air is lower than and the oxygen content of the modified air is greater than the corresponding air contents. The invention is also characterised in that at least one part of the molecular nitrogen contained in the product gas mixture is separated from the product gas mixture before the product gas mixture containing propane formed in the first reaction step is further processed to produce acrolein and/or acrylic acid in at least one additional reaction step.

Description

Process for the production of acrolein and / or acrylic acid from propane

description

The present invention relates to a process for the preparation of acrolein and / or acrylic acid from propane, in which the propane is subjected to a partial homogeneous and / or heterogeneously catalyzed oxydehydrogenation with molecular oxygen to propene in a first reaction stage and then the propene formed in the first reaction stage containing product gas mixture used in at least one further reaction stage for the production of acrolein and / or acrylic acid by gas phase catalytic propene oxidation.

Acrolein and acrylic acid are important intermediates that are used, for example, in the manufacture of active ingredients and polymers.

The process currently mainly used on an industrial scale for the production of acrolein and / or acrylic acid is the gas-phase catalytic oxidation of propene (e.g. EP-A 575 897), the propene being predominantly produced as a by-product of ethylene production by steam cracking naphtha.

Since the other fields of application of propene, e.g. the production of polypropylene, expanding ever further, it would be advantageous to have a large-scale, competitive process for the production of acrolein and / or acrylic acid, the raw material base of which is not propene, but e.g. propane is a naturally occurring component of natural gas.

It is known from US Pat. No. 3,798,283 that propane can be homogeneously oxidized to propene in the presence of molecular oxygen at elevated temperature. Both pure oxygen and mixtures of oxygen and inert gas can be considered as an oxygen source.

From DE-A 20 58 054 and DE-A 1 95 30 454 it is known that the oxydehydrogenation of propane to propene can also be carried out under heterogeneous catalysis. US Pat. No. 3,161,670, EP-A 117 446 and DE-A 33 13 573 relate to processes for the preparation of acrolein and / or acrylic acid, in which propane is first dehydrogenated to propene by heterogeneously catalyzed dehydrogenation with the exclusion of oxygen. 5

The product mixture containing propene is then subjected to heterogeneously catalyzed gas-phase oxidation. A disadvantage of this procedure, however, is that the catalyst required for the non-oxidative dehydrogenation of the propane is deactivated relatively quickly by carbon deposits and therefore frequently has to be regenerated. Another disadvantage of this procedure is the hydrogen formation associated with the nonoxidative propane dehydrogenation.

15 DE-A 33 13 573 mentions the basic possibility of coupling oxidative dehydrogenation of propane to propene with subsequent heterogeneously catalyzed propene oxidation. However, it does not contain any further information on the implementation of such a procedure.

20th

EP-A 293 224, US-A 5 198 578 and US-A 5 183 936 teach that an increased N ₂ content in the diluent gas is disadvantageous in the catalytic gas phase oxidation of propene to acrolein and / or acrylic acid. EP-A 293 224 also suggests that

25 Combine oxidative dehydrogenation of propane to propene and the catalytic gas phase oxidation of propene for the purpose of producing acrolein and / or acrylic acid.

In Catalysis Today 13, 1992, pp. 673 to 678, Moro-oka 30 et al combine in laboratory tests a homogeneous oxidative dehydrogenation of propane to propene with a subsequent heterogeneously catalyzed oxidation of the dehydrogenation product mixture to acrolein and / or acrylic acid. Moro-oka et al in Applied Catalysis, 70 (2), 1991, pp. 175 to 35 187 recommend the corresponding process combination. The recommendations of EP-A 293 224, US-A 5 198 578 and US-A 5 183 Correspondingly, in all cases, Moro-oka et al use either pure molecular oxygen or air depleted in nitrogen as the oxygen source for the oxide dehydrogenation stage. For the latter case, Moro-oka does not provide for any separation of nitrogen introduced into the process in the further course of his process.

CN-A 110 5352 likewise discloses a homogeneous oxidative dehydrogenation of propane to propene with a subsequent heterogeneously catalyzed oxidation of the dehydrogenation product mixture to acrolein and / or acrylic acid. Since CN-A 110 5352 is a large-scale procedure to be carried out in CN-A 110 5352, following the recommendation of EP-A 293 224, US-A 5 198 578 and US-A 5 183 936, only pure molecular oxygen is used as the oxygen source.

WO 97/36849 relates to the combination of a catalytic oxidative dehydrogenation of propane to propene with a subsequent heterogeneously catalyzed oxidation of the dehydrogenation product mixture to acrolein and / or acrylic acid in an industrial version.

WO 97/36849 does not rule out the use of nitrogen-containing oxygen (e.g. air) as a source of the molecular oxygen required in the oxydehydrogenation, but advises against it. In addition, WO 97/36849 only provides an outlet (purge) of circulating gas and no component separation from the circulating gas in the case of a continuous procedure with recycle gas to suppress undesired leveling of disadvantageous reaction gas mixture components.

For reasons of economy, essentially only air can be used as the starting material for the molecular oxygen source for large-scale gas phase oxidations.

Against this background, the above-mentioned procedures are disadvantageous insofar as, due to the similarity of 0 ₂ and N _2, starting from air, the sole measure of a preliminary nitrogen / oxygen separation for the production of pure oxygen or an air depleted in nitrogen to limit the nitrogen content subsequent oxidation of the propene contained in the dehydration product mixture is very energy-intensive.

The object of the present invention was therefore a process for the preparation of acrolein and / or acrylic acid from propane, in which the propane is subjected to a homogeneous and / or a heterogeneously catalyzed oxydehydrogenation with molecular oxygen to propene in a first reaction stage and then in the first reaction stage formed, containing propene, product gas mixture used in at least one further reaction stage for the production of acrolein and / or acrylic acid by gas phase catalytic propene oxidation, in which the restriction of the nitrogen content in the propene oxidation stage in a less energy-consuming manner than done in the prior art. Accordingly, a process for the preparation of acrolein and / or acrylic acid from propane was obtained, in which the propane is subjected to a partial homogeneous and / or heterogeneously catalyzed oxydehydrogenation with molecular oxygen to propene in a first reaction stage and then the one formed in the first reaction stage is subjected to Product gas mixture containing propene used in at least one further reaction stage for the production of acrolein and / or acrylic acid by gas-phase-catalyzed propene oxidation, which is characterized in that the reaction gas starting mixture of the first reaction stage requires the molecular oxygen different from cycle gas oxygen than that required in the first reaction stage Modified air containing nitrogen is added, provided that the nitrogen (expressed in mol%) of the modified air is smaller and the oxygen content of the modified air is larger than the corresponding contents of air (i.e. h of air depleted in nitrogen), and that, before the further use of product gas mixture formed in the first reaction stage and containing propene for the production of acrolein and / or acrylic acid in the at least one further reaction stage, at least part of the molecular nitrogen contained in the product gas mixture from the product gas mixture separates.

The principal advantage of the procedure according to the invention compared to the closest prior art processes is the result of in-depth research that the chemical bonding of the molecular oxygen which takes place in the course of the oxidative dehydrogenation of the propane using a nitrogen-containing oxygen source implicitly forms part of the Separation of atmospheric nitrogen and atmospheric oxygen to be carried out separating work. The difference between the resulting oxygen-containing polar compounds (e.g. H ₂ 0) and N ₂ is much more pronounced than the difference between N and 0 ₂ , which is why an at least partial subsequent N ₂ separation from the oxidative dehydrogenation product mixture is significantly less is energy-consuming, as an anticipatory N ₂ separation from air.

If the first reaction stage is designed as a homogeneous oxide dehydrogenation in the process according to the invention, this can in principle be carried out, for example, as described in US Pat. Nos. 3,798,283, CN 1,105,352, Applied Catalysis, 70 (2), 1991, p 175 to 187, Catalysis Today 13, 1992, pp. 673 to 678 and the older appendix DE-A 1 96 22 331 can be seen from this from that according to the invention modified air is to be used as the oxygen source (apart from oxygen cycle gas).

The temperature of the homogeneous oxide dehydrogenation is expediently chosen to be in the range from 300 to 700 ° C., preferably in the range from 400 to 600 ° C., particularly preferably in the range from 400 to 500 ° C. The working pressure can be 0.5 to 100 bar or 1 to 50 bar. Often it will be 1 to 20 bar or 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 tube furnace or a tube bundle reactor can be used, e.g. a counterflow tube furnace with flue gas as the heat carrier, or a tube bundle reactor with molten salt as the heat carrier.

The propane to oxygen ratio in the reaction gas starting mixture to be used can be 0.5: 1 to 40: 1. It is advantageous according to the invention if the molar ratio of propane to molecular oxygen in the reaction gas starting mixture is <6: 1 or <5: 1. As a rule, the aforementioned ratio will be> 1: 1 or> 2: 1. The nitrogen content in the reaction gas starting mixture is generally a consequence of the aforementioned requirement, since the reaction gas starting mixture normally does not comprise essentially any other gases besides propane and modified air. Of course, the reaction gas starting mixture can also comprise further, essentially inert, constituents such as H0, C0 ₂ , CO, noble gases and / or propene. Propene as a component of the reaction gas starting mixture is given, for example, if the C ₃ fraction from the refinery or the C ₃ fraction from the light hydrocarbons of the oil field, which have a propene content of up to 10% by weight, is used as the starting propane can. It can also be the case if the process according to the invention is carried out continuously and, after the last reaction step, unreacted propane and / or propene is returned to the oxydehydrogenation. Components returned to the oxydehydrogenation are generally referred to in this document as circulating gas. As a result of recycle gas recycle, the nitrogen content in the reaction gas starting mixture can be up to 60 mol% or up to 50 mol%. Recycle gas recirculation can also lead to the reaction gas starting mixture having up to 5 mol% of gases such as CO, CO ₂ , ethene and H ₂ O in a continuous procedure. 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. This is a consequence of radical mechanism of homogeneous oxidative propane dehydrogenation, reaction chamber surfaces generally act as radical scavengers. Aluminum oxide, quartz glass, borosilicate, stainless steel and aluminum are particularly favorable surface materials.

If the first reaction stage is designed as a heterogeneously catalyzed oxydehydrogenation in the process according to the invention, this can in principle be e.g. as in US-A 4, 788, 371, CN-A 10 733 893, 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-A 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-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 Corberän and S. Vic Bellόn (Editors), New Developments in Selective Oxidation II, 1994, Elsevier Science BV, pp. 305 - 313, 3rd World Congress on Oxidation

Catalysis R.K. Grasselli, S.T. Oyama, A.M. Gaffney and J.E. Lyons (Editors), 1997, Elsevier Science B.V. , P. 375 ff., It is disregarded that modified air must be used as the oxygen source (apart from oxygen cycle gas) according to the invention. In particular, all of the oxide hydrogenation catalysts mentioned in the abovementioned documents can be used. The statements made for the aforementioned documents also apply 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 .; Trifirό, 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, RHH; Seshan, K.; Ross, JRH 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, JM; Thomas, G. Catal. Lett. 1991, 10, 181;

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

g) Ind. Eng. Chem. Res. 1996, 35, 2137-2143 and

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

Oxydehydrogenation catalysts which are particularly suitable according to the invention are the multimetal oxide compositions or catalysts A of DE-A 1 97 53 817, the multimetal oxide compositions or catalysts A mentioned as preferred in the abovementioned publication being very particularly favorable. That , In particular, multimetal oxide compositions of the general formula I ^ MOi-M ^ Ox (I) come as active compositions

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 valency and frequency of the elements in I other than oxygen,

into consideration.

In principle, active compositions I suitable according to the invention can be prepared in a simple manner by generating an intimate, preferably finely divided, dry mixture composed of suitable elementary constituents according to their stoichiometry and calcining them at temperatures of 450 to 1000 ° C. The calcination can take place both under inert gas and under an oxidative atmosphere such as air (mixture of inert gas and oxygen) and also under a reducing atmosphere (eg mixture of inert gas, oxygen and NH ₃ , CO and / or H). The calcination time can range from a few minutes to a few hours and usually decreases with temperature. As sources for the elementary tare constituents of the multimetal oxide active compositions I are those compounds which are already oxides and / or those compounds which can be converted into oxides by heating, at least in the presence of oxygen.

In addition to the oxides, such starting compounds are, above all, halides, nitrates, formates, oxalates, citrates, acetates, carbonates, amine complex salts, ammonium salts and / or hydroxides (compounds such as NH ₄ OH, (NH ₄ ) ₂ CO ₃ , NH ₄ N0 ₃ , NH ₄ CH0 ₂ , CH ₃ COOH, NH ₄ CH ₃ C0 ₂ and / or ammonium oxalate, which can decompose and / or decompose to form completely gaseous compounds at the latest during later calcination, can also be incorporated into the intimate dry mixture ). The intimate mixing of the starting compounds for the production of multimetal oxide masses I can take place in dry or in wet form. If it is carried out in dry form, the starting compounds are expediently used as finely divided powders and, after mixing and optionally compacting, are subjected to the calcination. However, the intimate mixing is preferably carried out in wet form. Usually, the starting compounds are mixed together in the form of an aqueous solution and / or suspension. In the dry process described, particularly intimate dry mixtures are obtained if only sources of the elementary constituents present in dissolved form are used. Water is preferably used as the solvent. The aqueous mass obtained is then dried, the drying process preferably being carried out by spray drying the aqueous mixture at exit temperatures of 100 to 150 ° C. Particularly suitable starting compounds of Mo, V, W and Nb are their oxo compounds (molybdates, vanadates, tungstates and niobates) or the acids derived from them. This applies in particular to the corresponding ammonium compounds (ammonium molybdate, ammonium vanadate, ammonium tungstate).

The multimetal oxide materials I can be used for the process according to the invention both in powder form and in the form of certain catalyst geometries, it being possible for the shaping to take place before or after the final calcination. For example, solid catalysts can be produced from the powder form of the active composition or its uncalcined precursor composition by compression to the desired catalyst geometry (for example by tableting, extrusion or extrusion), with auxiliaries such as graphite or stearic acid optionally being used as lubricants and / or molding aids and reinforcing agents such as microfibers Glass, asbestos, silicon carbide or ca lium titanate can be added. Suitable full catalyst geometries are, for example, full cylinders or hollow cylinders 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, for example, 7 mm x 7 mm x 4 mm or 5 mm x 3 mm x 2 mm or 5 mm x 2 mm x 2 mm (each length x outer diameter x inner diameter). Of course, the full catalyst can also have a spherical geometry, the spherical diameter being 2 to 10 mm.

Of course, the pulverulent active composition or its pulverulent, not yet calcined, precursor composition can also be shaped by application to preformed inert catalyst supports. The coating of the support bodies for the production of the coated catalysts is usually carried out in a suitable rotatable container, such as that e.g. is known from DE-A 2909671 or from EP-A 293859. For coating the carrier bodies, the powder mass to be applied can expediently be moistened and, after application, e.g. by means of hot air. The layer thickness of the powder mass applied to the carrier body is expediently selected in the range from 50 to 500 μm, preferably in the range from 150 to 250 μm.

Conventional porous or non-porous aluminum oxides, silicon dioxide, thorium dioxide, zirconium dioxide, silicon carbide or silicates such as magnesium or aluminum silicate can be used as carrier materials. The carrier bodies can have a regular or irregular shape, with regularly shaped carrier bodies with a clearly formed surface roughness, e.g. Balls or hollow cylinders are preferred. It is suitable to use essentially non-porous, rough-surface, spherical supports made of steatite, the diameter of which is 1 to 8 mm, preferably 4 to 5 mm.

The reaction temperature of the heterogeneously catalyzed oxydehydrogenation of the propane will suitably be in the range from 300 to 600 ° C., frequently in the range from 350 to 500 ° C. 0.5 to 10 bar or 1 to 10 bar or 1 to 5 bar is recommended as the working pressure. Working pressures above 1 bar, e.g. at 1.5 to

10 bar have proven to be advantageous. As a rule, the heterogeneously catalyzed oxydehydrogenation of the propane takes place on a fixed catalyst bed. The latter is expediently heaped up in the tubes of a tube bundle reactor, as described, for example, in EP-A 700 893 and in EP-A 700 714 and the literature cited in these documents. The average residence time of the reaction gas mixture in the catalyst bed is normally from 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 can, according to the invention, be from 0.5: 1 to 40: 1. It is advantageous according to the invention if the molar ratio of propane to molecular oxygen in the reaction gas starting mixture is <6: 1 or <5: 1. As a rule, the aforementioned ratio will be> 1: 1 or 2: 1. The nitrogen content in the reaction gas starting mixture is generally a consequence of the aforementioned requirement, since the reaction gas starting mixture normally does not comprise essentially any other gases besides propane and modified air. Of course, the reaction gas starting mixture can also comprise further, essentially inert, constituents such as H ₂ O, CO ₂ , CO, noble gases and / or propene. Propene as a component of the reaction gas starting mixture is given, for example, if the C ₃ fraction coming from the refinery or the C ₃ fraction from the light hydrocarbons of the oil field, which have a propene content of up to 10% by weight, is used as the starting propane can. It can also be the case if the process according to the invention is carried out continuously and, after the last reaction step, unreacted propane and / or propene is recycled into the heterogeneously catalyzed oxydehydrogenation. As a result of recycle gas recycle, the nitrogen content in the reaction gas starting mixture of the heterogeneously catalyzed propane oxydehydrogenation in the process according to the invention can be up to 60 mol% or up to 50 mol% when carried out continuously. Recycle gas recirculation can also result in the reaction gas starting mixture having up to 5 mol% of gases such as CO, CO ₂ , ethane, methane, ethene and / or H ₂ 0 in a continuous procedure. The modified air to be used as an oxygen source according to the invention can contain, for example,> 0.05% by volume to <78% by volume or> 0.1% by volume to <75% by volume of N ₂ . That is, the nitrogen content of the modified air to be used as an oxygen source according to the invention can be> 1% by volume to <70% by volume or> 5% by volume to <60% by volume or> 10% by volume to <50% by volume .-% or> 15 vol.% to <40 vol.% or> 20 vol.% to <30 vol.%.

Accordingly, the oxygen content of the modified air to be used according to the invention as an oxygen source can be> 20.95

% By volume to 99.95% by volume or> 25% by volume to <99.9% by volume or> 30% by volume to <99% by volume or> 40% by volume to < 95 vol.% Or> 50 vol.% To <90 vol.% Or> 60 vol.% To <85 vol.% Or> 70 vol.% To <80 vol.%. In addition to oxygen and nitrogen, the modified air to be used according to the invention can also contain the other constituents, such as noble gases, carbon dioxide, water vapor, etc., which are usually present in small amounts. Of course, the aforementioned components may also have been partially or completely removed as part of the modification.

In the simplest way, the modified air to be used according to the invention as an oxygen source can e.g. by fractional distillation of air, preferably under pressure. Of course, the methods of EP-A 848 981 and EP-A 848639 can also be used.

Advantageously, both the 0 ₂ conversion of the homogeneous and the catalytic oxidative propane dehydrogenation in the process according to the invention (in a single pass) should be> 70 mol%. Ie, the aforementioned 0 conversion can> 75 mol. % or> 80 mol% or> 85 mol% or> 90 mol% or> 95 mol% or> 97 mol% or> 99 mol%.

It goes without saying that homogeneous and catalytic propane oxide hydrogenation can also be used in combination according to the invention.

The possible components of the product gas mixture of the propane oxide dehydrogenation may include, for example, the following components: CO ₂ , CO, HO, ^N ₂ ° ₂ 'propene, propane, ethane, ethene, methane, acrolein, acrylic acid, ethylene oxide, butane, acetic acid, formaldehyde, formic acid , Propylene oxide and butene.

The separation of at least part of the nitrogen contained therein required according to the invention prior to further use of the aforementioned propane-containing propane oxyhydrogen product gas mixture for the production of acrolein and / or acrylic acid from the same can, e.g. done in a simple manner by distillation.

It is expedient to use fractional distillation; preferably a fractional pressure distillation at low temperatures. The pressure to be used can be, for example, 10 to 100 bar. Packing columns, tray columns or packing columns can be used as rectification columns. Tray columns with dual-flow trays, bubble trays or valve trays are suitable. The reflux ratio can be, for example, 1 to 10. Other possibilities for nitrogen separation are, for example, pressure swing absorption, pressure washing and pressure extraction. Based on the total amount of nitrogen contained in the propane oxide hydrogenation product gas mixture, the amount of nitrogen to be separated off according to the invention can be 5%, or 10%, or 20%, or 30%, or 40%, or 50%, or 60%, or 70%, or 80%, or 90% or 95 to 5 100%.

Of course, it is also possible not to separate the nitrogen itself from the propane oxide hydrogenation product gas mixture, but together with other components of the propane oxide hydrogenation product gas mixture which may not be desired in the subsequent reaction stage. For example, one can lay the dividing line in a fractional separation of the nitrogen from the propane oxydehydrogenation product gas mixture by distillation so that, for example, at the top of the rectification column essentially all those components are separated whose boiling point is lower than the boiling point of propene. These components will primarily be the carbon oxides CO and C0 ₂ as well as unreacted oxygen and ethylene and methane. Of course, only a part of the above-mentioned components can be separated off together with the nitrogen.

Depending on the catalyst used for the subsequent catalytic gas phase oxidation of propene to acrolein and / or acrylic acid, it may also be expedient, in addition to the above-mentioned constituents, also partially or completely to remove the water contained in the propane-oxydehydrogenation product gas mixture from the propane / propene contained in the propane-oxydehydrogenation product gas mixture cut off. 0 This can easily be e.g. by propelling the propane and propene from the bottom liquid remaining in the fractional distillation separation of the constituents of the propane oxyhydrogen product gas mixture which have a lower boiling point than propene and the propene is evaporated, with the water precipitating. 5

In addition, the propane and unreacted propane-containing propane oxides and the mixture of propylene oxides, which have been depleted in nitrogen and possibly the aforementioned secondary components, can be used directly for the production of acrolein and / or acrylic acid by catalytic propene oxidation, as used in the prior art, for example in WO 97 / 36849 or CN-A 11 05 352. Of course, this gas phase catalytic propene oxidation can also be carried out as described in EP-A 117 146, US-A 5 198 578 or US-A 5 183 936 5. But you can also in analogy to DE-A 33 13 573, CA-A 12 17 502, US-A 3 161 670 or US-A 4 532 365 can be carried out.

The gas-phase catalytic oxidative conversion of the propene contained in the propane-oxydehydrogenation product gas mixture (hereinafter referred to as “residual product gas mixture”), which has been depleted in nitrogen, to acrolein and / or acrylic acid can take place, for example, in one or two subsequent oxidation stages. The propane accompanying the propene and any N ₂ , noble gas, CO, C0 ₂ , H0 and other lower organic compounds such as other lower hydrocarbons, which may be present, function essentially as an inert diluent gas. In the event that acrylic acid is the desired target product, two gas-phase catalytic oxidation stages will generally follow, although single-stage gas-phase catalytic oxidations of propene to acrylic acid are also known in the prior art. If acrolein is the desired target product, usually only one gas phase catalytic oxidation stage will follow.

The catalytic gas phase oxidation of the propene contained in the residual product gas mixture to an amount of acrolein which is predominant compared to acrylic acid is preferably described as described in EP-A 731 082, DE-A 44 31 957, DE-A 29 09 597 or EP-A 575 897 carried out.

In other words, the gas phase oxidation is expediently carried out in a multi-contact tube fixed bed reactor.

As a rule, in the case of a propene: oxygen: essentially indifferent gases, the volume (Nl) ratio is 1: (1.0 to 3.0): (5 to 25), preferably 1: (1.7 to 2.3 ): (10 to 15), worked.

To achieve the abovementioned ratios, it may be necessary to add molecular oxygen to the residual product gas mixture containing the propene before it is introduced into the propene oxidation stage. This can be in the form of air, in the form of nitrogen-depleted air or in the form of pure oxygen. Of course, additional dilution gases (eg H0), which are known to be indifferent, can be added at this point as desired. The reaction temperature is expediently chosen to be 300 ° C. to 450 ° C., preferably 320 ° C. to 390 ° C. The reaction pressure is usually 0.5 to 5 bar, preferably 1 to 3 bar. The total space load is often 1500 to 2500 Nl / l / h. Suitable catalysts for this oxidation stage are, for example, those of DE-A 29 09 592, especially those from Example 1 of this document. Alternatively, however, the multimetal oxide catalysts II or II 'of DE-A 1 97 53 817 can also be used. This particularly applies to those in these

Written exemplary embodiments listed. Especially when they are designed as hollow cylinder full catalysts as described in EP-A 575 897. Of course, the multimetal oxide catalyst ACF-2 from Nippon Shokubai containing Bi, Mo and Fe can also be used in the propene oxidation stage.

In the above-mentioned propene oxidation stage, no pure acrolein is obtained, but a mixture from whose secondary components the acrolein can be separated in a manner known per se. The acrolein so separated can be used as an intermediate for the synthesis of various end products. Of course, it can also be used for gas phase catalytic oxidation for the production of acrylic acid. When the acrolein is used for the production of acrylic acid in a further gas phase catalytic oxidation stage, the reaction gases of the propene oxidation stage containing the acrolein are generally transferred to this further oxidation stage without removal of secondary components. If necessary, they undergo intermediate cooling beforehand.

This further oxidation stage is also expediently implemented in a multi-contact tube fixed-bed reactor, as it is e.g. is described in DE-A 44 31 949, DE-A 44 42 346, DE-A 1 97 36 105 or EP-A 731 082.

As a rule, with an acrolein: oxygen: water vapor: other essentially indifferent gases, the volume (Nl) ratio of 1: (1 to 3): (> 0 to 20): (3 to 30), preferably 1: (1 to 3): (0.5 to 10): (7 to 18) worked. To achieve the above-mentioned ratios, it may be necessary to additionally add molecular oxygen to the product gas mixture from the propene oxidation stage containing acrolein before it is introduced into the acrolein oxidation stage. This can be in the form of air, in the form of nitrogen-depleted air or in the form of pure oxygen. Of course, additional dilution gases, known essentially as indifferent, can be added at this point as desired. The reaction temperature is expediently chosen to be from 200 ° C. to 300 ° C., preferably from 220 to 290 ° C. The reaction pressure is usually 0.5 to 5 bar, preferably 1 to 3 bar. The total space load is preferably 1000 to 2500 Nl / l / h. Suitable catalysts for this oxidation stage are, for example, those of the general formula I or I 'from DE-A 44 42 346. Alternatively, however, the multimetal oxide catalysts of DE-A 1 97 36 105, in particular the exemplary embodiments mentioned in this document, can also be used become. Of course, the multimetal oxide catalyst ACS-4 from Nippon Shokubai comprising Bi, Mo and Fe can also be used in the acrolein oxidation stage.

Of course, the gas mixture leaving the acrolein oxidation stage does not consist of pure acrylic acid, but of a gas mixture containing the latter, from which acrylic acid can be separated in a manner known per se.

The various known variants of acrylic acid separation are e.g. summarized in DE-A 1 96 00 955. Correspondingly, the acrolein could also be separated from the reaction gas mixture leaving the propene oxidation stage. A common feature of the separation processes is that the desired product is separated from the reaction gas mixture of the acrolein oxidation stage either by absorption with a solvent (cf. also DE-A 43 08 087) or by absorption with water or by partial condensation (the resulting absorbate or Condensate is then distilled (if necessary with the addition of an azeotropic

Entrainer) and / or worked up crystallisatively and thus essentially pure acrylic acid or pure acrolein obtained).

The dividing line is drawn in all cases so that a residual gas stream essentially free of acrylic acid and / or acrolein is formed, the main components of which are carbon oxides (CO, C0 ₂ ), optionally N ₂ , noble gases, 0, water vapor, propane and unreacted Are propene.

Expediently you will be part of a continuous

Carrying out the process according to the invention recycle the propane and propene contained in this residual gas stream into the oxydehydrogenation. In the event that the at least partial nitrogen separation required according to the invention after the propoxide dehydrogenation is carried out in such a way that together with the nitrogen all gas components boiling less than propene are also separated off, it is expedient to separate the residual gas stream as such into the Attributed to propane oxide hydrogenation. By cooling the residual gas stream, it can also be used to comparatively easily condensable components such as water vapor, small amounts of acrylic acid and / or acrolein and small amounts of by-products such as, for example, before recycling Formaldehyde and acetic acid are separated (except water as so-called _S). In particular if, after the propane oxide dehydrogenation according to the invention, only a sole, partial or complete, nitrogen separation is carried out, it is advisable to recycle at least a part of the contents contained therein, other than propane and propene, before the residual gas stream is returned to the propane oxydehydrogenation To separate residual gas flow components. This can be done in the simplest way, for example by fractional distillation, in particular fractional pressure distillation.

Examples

A) Production of acrolein and / or acrylic acid starting from propane

a) 15.4 mol / h of a gas mixture (modified air) consisting of 90% by volume of O ₂ and 10% by volume of N ₂ and 79.7 mol / h of cycle gas of the composition

87.7% by volume propane, 0.4% by volume propene, 4.1% by volume 0 ₂ , 2.1% by volume N ₂ , 2.5% by volume H ₂ 0,

0.8 vol.% CO, 1.9 vol.% C0 ₂ and 0.5 vol.% Other components,

were combined to give 95.1 mol / h of reaction gas starting mixture and compressed to a pressure of 2.2 bar and heated to a temperature of 430.degree. A 3.8 m long reaction tube made of V2A steel (2.0 mm wall thickness; 2.6 cm inside diameter) was charged with the aforementioned reaction gas starting mixture, which was cooled over its entire length to a temperature of 430 ° C. using a salt bath.

In the direction of flow, the reaction tube was initially loaded with spherical (diameter = 8 mm) shaped steatite bodies over a length of 0.8 m (instead of the steatite balls, steatite rings of the geometry mentioned below 5 mm x 3 mm x 2 mm can also be used). On the remaining bed length of 3 m, the reaction tube with a bed of the multimetal oxide catalyst according to Example 1, a) / multimetal oxide mass I of DE-A 19 753 817, was pressed to form fully catalyst cylinders of geometry 5 mm × 3 mm × 2 mm (outside diameter × Height x inner diameter), filled. The inlet pressure was 1.3 bar, the outlet pressure was 1.0 bar. The product mixture (102.3 mol / h) leaving the reaction tube had the following composition:

59.9 vol.% Propane,

6.5 vol.% Propene, 4.8 vol.% 0 ₂ , 3.2 vol.% N ₂ ,

16.7 vol.% H ₂ 0, 4.3 vol.% CO

4.0 vol.% C0 ₂ , 0.2 vol.% Acrolein and 0.4 vol.% Other components.

The aforementioned product mixture was compressed to 36 bar, cooled to 70 ° C. and then fed in two phases to a rectification column operated under pressure, which had 51 trays. The product mixture was fed onto the 30th floor from below. The top pressure of the rectification column (a bell-bottom column, diameter 50 mm) was 36 bar.

The column head was cooled with the coolant Baysilone® KT3 (flow temperature -50 ° C). The bottom temperature was 92 ° C. In the bottom of the column, part of the bottom liquid removed was returned as boiling vapor.

At the top of the column, 17.4 mol / h of an offgas which had the following composition were removed:

2.7 vol.% Propane, 0.5 vol.% Propene, 28.2 vol.% 0 ₂ ,

18.8 vol.% N ₂ , 24, 9 vol.% CO

23.2% by volume of CO ₂ and 1.7% by volume of other components.

The bottom liquid continuously withdrawn was 84.9 mol / h. It was cooled to 35 ° C. and expanded to 10 bar in a storage container and brought to a temperature of 25 ° C. The gas phase of the storage container was expanded to 2.3 bar via a throttle. The aqueous liquid phase of the storage container was also continuously removed and, after evaporation, combined with the removed, relaxed gas phase. 13.6 mol / h of a mixture of 90% by volume of O ₂ and 10% by volume of N ₂ and 9.1 mol / h of propane were added to the bottom liquid which had been continuously transferred into the gas phase, which made propene gas-catalytic - Oxidation resulted in a reaction mixture with the following composition:

0.2 vol.% Acrolein, 65, 0 vol.% Propane, 6.1 vol.% Propene,

11.4 vol.% 0 ₂ ,

1.3 vol.% N ₂ , 15.9 vol.% H ₂ 0 and

0.1 vol .-% other components.

A reaction tube (V2A steel; length 3.80 m; 2.0 mm wall thickness, 2.6 cm inner diameter) was initially charged in the outflow direction over a length of 50 cm with a bed of steatite balls (diameter: 4-5 mm) . On a contact tube length of 3.00 m, a fill of the

Multimetal oxide catalyst according to Example 1, 3rd / multimetal oxide II from DE-A 19 753 817. The entire length of the reaction tube was kept at 350 ° C. with a salt bath and 107.6 mol / h of the aforementioned reaction gas starting mixture (which had a temperature of 200 ° C.) were charged.

The inlet pressure was 2.0 bar, the outlet pressure was 1.8 bar.

2.4 mol / h of a gas mixture consisting of 90% by volume of O ₂ and 10% by volume of N were mixed into the product gas mixture leaving the reaction tube of the propene oxidation stage and with the resulting reaction gas starting mixture, which was brought to a temperature of 200.degree was charged, the acrolein oxidation tube described below.

This reaction tube (V2A steel; length: 3.80 m; 2.0 mm wall thickness; 2.6 cm inner diameter) was initially at a length of 50 cm in the outflow direction with a bed of steatite balls (diameter: 4-5 mm) loaded. A bed of the multimetal oxide catalyst according to Example b, Sl of DE-A 4 442 346 followed on a contact tube length of 2.70 m. The entire length of the reaction tube was heated to 270 ° C. by means of a salt bath and charged with the reaction gas starting mixture described above. The inlet pressure was 1.8 bar and the outlet pressure was 1.7 bar. The product mixture leaving the reaction tube in an amount of 107.8 mol / h had the following composition:

0.1 vol.% Acrolein,

5.2% by volume of acrylic acid,

0.1% by volume of acetic acid,

64.9 vol.% Propane,

0.3 vol.% Propene, 3.0 vol.% 0 ₂ ,

1.5% by volume of N ₂ ,

22.7% by volume H ₂ 0,

0.6 vol.% CO,

1.4 vol.% C0 ₂ and 0.2 vol.% Of other components.

The hot reaction gas leaving the acrolein oxidation stage was extracted from 57.4% by weight in a venturi scrubber (quench apparatus) by direct contact with quench liquid (140-150 ° C.) injected through slots in the narrowest cross section of the venturi tube. Diphenyl ether, 20.7% diphenyl and 20% by weight o-dimethyl phthalate cooled to a temperature of about 175 ° C. Subsequently, the drop-like liquid portion of the quench liquid was separated from the gas phase consisting of reaction gas and evaporated quench liquid in a downstream droplet separator (feed container with gas pipe led away at the top) and returned in a circuit I to the venturi washer. A partial stream of the recycled quench liquid was subjected to a solvent distillation, the quench liquid being distilled over and high-boiling secondary components which were burned up remaining. The distilled quench liquid was fed to the outlet of the absorption column described below.

The gas phase, which had a temperature of approximately 175 ° C., was fed into the lower part of a packed column (3 m high; double jacket made of glass; inner diameter 50 mm; three packed zones with lengths (from bottom to top) 90 cm, 90 cm and 50 cm ; the packing zones were thermostatted from bottom to top as follows: 90 ° C, 60 ° C, 20 ° C; the penultimate and the last packing zone were separated by a chimney tray; the packing units were stainless steel metal helices with a helix diameter of 5 mm and one Helix length of 5 mm; the absorbent was fed in directly above the central packing zone) and the

Claims

A flow of 2400 g / h of the absorbent, which is also composed of 57.4% by weight of diphenyl ether, 20.7% by weight of diphenyl and 20% by weight of o-dimethyl phthalate and is added at a temperature of 50 ° C. The drain of the absorption column, which in addition to acrylic acid also low-boiling by-products such as Contained acrolein and acetic acid, was indirectly heated in a heat exchanger to 100 ° C and placed on the top of a desorption column, which was also designed as a packed column with a length of 2 m (double jacket made of glass; 50 mm inner diameter;

Packing: stainless steel helices with a helix diameter of 5 mm and a helix length of 5 mm; a packing zone with a length of 1 m; thermostated to 120 ° C). In the desorption column, the components with a lower boiling point than acrylic acid, such as acrolein and acetic acid, were passed through

Stripping with 10 mol / h cycle gas (countercurrent; feed temperature 120 ° C) largely removed from the acrylic acid / absorbent mixture. The loaded circulating gas (stripping gas) leaving the desorption column was recirculated and combined with the hot reaction gas from the acrolein oxidation stage before it entered the venturi.

The non-absorbed gas mixture leaving the second packed zone in the absorption column was further cooled in the third packed zone in order to remove the easily condensable part of the secondary components contained therein, e.g. Separate water and acetic acid by condensation. This condensate is called acid water. To increase the separation effect, part of the acid water above the third packing zone of the absorption column was treated with a

Temperature of 20 ° C returned to the absorption column. The sour water was removed from the chimney floor below the uppermost packing zone. The reflux ratio was 200. The continuously removed amount of acid water was 23.0 mol / h. It contained next to

94.3% by weight of water and 3.0% by weight of acrylic acid. If necessary, this can be recovered as described in DE-A 1 96 00 955. The gas stream ultimately leaving the absorption column formed the circulating gas which, apart from the portion used for stripping in the desorption column, was recycled into the catalytic propane oxide hydrogenation and had the composition described at the beginning of the example. The amount of circulating gas returned for catalytic propane oxide hydrogenation was 79.7 mol / h. The bottom liquid of the desorption column was fed on the 8th tray from below to a tray column containing 57 dual flow trays (inside diameter: 50 mm; length 3.8 m; top pressure: 100 mbar; bottom pressure: 280 mbar: bottom temperature: 195 ° C) a pressure loss resistor was attached to the 9th floor;) and rectified in the same. 5.4 ol of a crude acrylic acid were withdrawn from the bottom from the 48th bottom per hour via side draw. The purity of the crude acrylic acid removed was> 98% by weight. At the top of the rectification column, after a partial condensation (reflux -

Ratio: 8.7) a gas stream enriched with low boilers and containing acrylic acid was drawn off and was recirculated into the absorption column above the lowest packing zone. The absorbent free of low boilers and almost free of acrylic acid was withdrawn from the bottom of the rectification column and recycled into the absorption column above the second packing zone (viewed from below). Phenothiazine was added to the reflux at the top of the rectification column as a polymerization inhibitor, in such amounts that the side draw contained 300 ppm of phenothiazine (DE-A 1 96 00 955 shows a schematic representation of the process for working up the reaction gas of the acrolein oxidation stage; in addition, the way of working up is also shown in DE-A 4308087). Based on the amount of propane used, the yield of acrylic acid was 59.3 mol%. The oxygen conversion in the oxide hydrogenation stage was 64.7 mol%, while the total oxygen conversion was 82.7 mol%.

12.0 mol / h of a gas mixture (modified air) consisting of 90% by volume of O ₂ and 10% by volume of N ₂ and 91.6 mol / h of cycle gas of the composition

88.1 vol.% Propane, 0.4 vol.% Propene,

0.2 vol.% Ethene,

3.6 vol.% 0 ₂ ,

2.0 vol.% N ₂ ,

2.5 vol.% H ₂ 0 0.8 vol.% CO,

1.9 vol.% C0 ₂ and

0.5 vol .-% other components,

were combined to 103.6 mol / h of reaction gas starting mixture, compressed to a pressure of 1.8 bar and to a

Temperature of 513 ° C warmed. With the aforementioned reaction gas starting mixture, an empty reaction tube made of V2A steel (2.0 mm wall thickness; 2.6 cm inner diameter), which was heated over its entire length to a temperature of 513 ° C (electric heating tape). The tube length was such that the mean residence time in the tube was 3 seconds. The product mixture leaving the reaction tube

(111.8 mol / h) had the following composition:

63.3 vol.% Propane,

6.8 vol.% Propene,

1.3 vol.% Ethene,

0.5 vol.% Methane,

2.0 vol. -% o ₂ ,

2.8 vol. -% N ₂ ,

15.8 vol.% H ₂ 0

3.3 vol.% CO,

3.7 vol.% C0 ₂ and

0.5 vol. - "5 other components

The aforementioned product gas mixture was compressed to 36 bar, cooled to 70 ° C. and then fed in two phases to a rectification column operated under pressure, which had 51 trays. The product mixture was fed in from the bottom on the 30th floor. The top pressure of the rectification column (a bell-bottom column, diameter 50 mm) was 36 bar. The column head was cooled with Baysilone ^® KT3

(Flow temperature: -50 ° C). The bottom temperature was 92 ° C. In the bottom of the column, part of the bottom liquid removed was returned as boiling vapor.

At the top of the column, 16.2 mol / h of an offgas which had the following composition were removed:

4.0 vol.% Propane,

0.8 vol. -% propene,

9.0 vol. -% ethene,

3.4 vol.% Methane,

13.8 vol.% 0 ₂ ,

19.3 vol.% N ₂ ,

22.7 vol.% CO,

25.2 vol.% C0 ₂ and

1.8 vol.% Other components

The bottom liquid continuously withdrawn was 95.6 mol / h. It was cooled to 35 ° C. and expanded to 10 bar in a storage container and to one

Brought temperature of 25 ° C. The gas phase of the storage container was expanded to 2.3 bar via a throttle. The aqueous liquid phase of the storage container was also continuously removed and, after evaporation, combined with the removed, relaxed gas phase. 15.0 mol / h of a gas mixture consisting of 90% by volume of O ₂ and 10% by volume of N were added to the bottom liquid thus continuously transferred into the gas phase, which resulted in a reaction gas starting mixture for the subsequent gas phase catalytic propene oxidation which had the following composition :

63.5 vol.% Propane,

6.7 vol.% Propene, 12.2 vol.% 0 ₂ ,

1.4 vol.% N ₂ , 16.0 vol.% H ₂ 0,

0.1 vol.% C0 ₂ and

0.1 vol .-% other components.

A reaction tube (V2A steel; length 3.80 m; 2.0 mm wall thickness, 2.6 cm inner diameter) was initially in the outflow direction over a length of 50 cm with a bed of Steati spheres (diameter: 4-5 mm ) loaded. A bed of the multimetal oxide catalyst according to Example 1, 3. / multimetal oxide II from DE-A 19 753 817 followed on a contact tube length of 3.00 m. The entire length of the reaction tube was heated to 350 ° C. in a salt bath and charged with 110.6 mol / h of the aforementioned reaction gas starting mixture (which had a temperature of 200 ° C.). The inlet pressure was 1.8 bar, the outlet pressure was 1.4 bar.

The product gas mixture leaving the reaction tube of the propene oxidation stage was admixed with 2.9 mol / h of a gas mixture consisting of 90% by volume of O ₂ and 10% by volume of N ₂ and with the resulting reaction gas starting mixture which was brought to a temperature of 200 ° C. the acrolein oxidation tube described below was charged. This reaction tube (V2A steel; length: 3.80 m, 2.0 mm wall thickness; 2.6 cm inner diameter) was initially at a length of 50 cm in the outflow direction with a bed of steatite balls (diameter: 4-5 mm) loaded. A bed of the multimetal oxide catalyst according to Example b, Sl of DE-A 4 442 346 followed on a contact tube length of 2.70 m. The entire length of the reaction tube was heated to 270 ° C. with a salt bath and charged with the reaction gas starting mixture described above. The inlet pressure was 1.5 bar and the outlet pressure was 1.3 bar. The product mixture leaving the reaction tube in an amount of 110.6 mol / h had the following composition:

0.1 vol.% Acrolein,

5.7 vol. Acrylic acid,

0.3% by volume of acetic acid,

62.4 vol.% Propane, 0.3 vol.% Propene, 0.2 vol.% Ethene,

3.0 vol.% 0 ₂ , 1.7 vol.% N ₂ ,

23.5 vol.% H ₂ 0, 0.7 vol.% CO, 1.6 vol.% C0 ₂ and

0.5 vol% other components.

The acrylic acid formed was separated from the hot reaction gas leaving the acrolein oxidation stage in a manner corresponding to that in Aa) and the cycle gas obtained was returned to the catalytic oxide hydrogenation. Instead of circulating gas, 11.0 mol / h of propane were used for stripping. The loaded propane (stripping gas) leaving the desorption column was recirculated and combined with the hot reaction gas of the acrolein oxidation stage before it entered the Venturi quench. The amount of acid water removed was 24.4 mol / h. It contained 94.3% by weight of water and 3.0% by weight of acrylic acid. The amount of crude acrylic acid removed in the rectification column via side draw was 6.2 mol / h.

In conclusion, it should be noted that, according to the invention, cycle gas can also be used for stripping instead of propane. In this case, the required propane would be fed directly to the oxydehydrogenation stage.

Based on the amount of propane used, the yield of acrylic acid was 56.4 mol%. The oxygen conversion in the oxide hydrogenation stage was 79.4% by volume, while the oxygen total conversion was 91.7% by volume. B) Production of 1 mol / s of propene (20 ° C., 1 bar) by oxydehydrogenation of propane. Consideration of the separation work to be carried out for nitrogen separation (to simplify the observation, complete conversion of the oxide hydrogenation and 100% selectivity of the Propene formation assumed).

a) Use of pure 0 as an oxygen source

The reaction stoichiometry of the oxydehydrogenation of propane to propene is:

1/2 0 ₂

CH ₃ CH ₂ CH ₃ CH ₃ CH = CH ₂ H ₂ 0

Accordingly, 0.5 mol / s 0 _{2 are} required to generate 1 mol / s propene by oxydehydrogenation of propane.

The starting material for the recovery of 0.5 mol / s 0 ₂ is air (78 vol.% N ₂ , 21 vol.% 0 ₂ and 1 vol.% Residual gases), the composition of which is 80 vol. N ₂ and 20% by volume 0 should be assumed. Ideal behavior should also be assumed.

To obtain the 0.5 mol / s 0, 2.5 mol / s air (= mixture of 0.5 mol / s 0 ₂ and 2.0 mol / s N) can be assumed, which have a temperature of 20 ° C and has a pressure of 1 bar.

In order to separate 0.5 mol / s of pure 0 ₂ from the aforementioned 2.5 mol / s air, the air is cooled from 1 bar to -194 ° C. while maintaining the pressure and, as a boiling liquid, is poured into the middle part of a liquid at 1 bar operated adiabatic rectification column to over

Head to be separated into pure nitrogen and as a liquid into pure oxygen.

According to "K. Sattler, Thermal Separation Process, Fundamentals,

Interpretation, apparatus, second edition, Verlag Chemie, Weinheim (1995), p. 182 "when using an ideal rectification column at the top of the column, ie having an infinite number of separation stages, a minimum reflux ratio is required (the higher the required minimum reflux ratio, the larger the separation work to be performed). While the reflux ratio at the top of the column is generally defined as the ratio of the proportion of the N quantity of gaseous gas at the top of the rectification column per unit of time which is returned to the rectification column after condensation to the proportion which is withdrawn, the minimum reflux ratio is at Column head (v _m i _n ) defines the ratio of the proportion of the amount of N ₂ produced in gaseous form at the top of the rectification column per unit of time, which is at least in order to achieve the separation task after condensation

Rectification column must be returned to the portion that is removed.

According to the above quotation after K. Sattler

α-1

With

α = relative volatility or separation factor of N and 0

(ie = ratio of the saturation vapor pressures of N and 0 ₂ at the average separation temperature along the rectification column (the separation or mixture decomposition takes place along the entire rectification column)).

The boiling temperature of N ₂ at a pressure of 1 bar = -196 ° C (= top temperature).

The boiling temperature of 0 ₂ at a pressure of 1 bar = -183 ° C (= bottom temperature).

The average boiling point along the rectification column is thus = -190 ° C. According to VDI -Wärmeatlas, 5th edition 1998 (DC6 and DC7) the saturation vapor pressure of N ₂ at -190 ° C = 2.0 bar and the saturation vapor pressure of 0 at -190 ° C = 0.48 bar.

The relevant α is thus 2.0: 0.48 = 4.15,

X ^F is the N ₂ molar fraction in the mixture to be separated (ie in air) fed to the rectification column. Ie, X ^F = 0.8 (ö 80 vol .-% N ₂ in air). X ^E is the N ₂ mole fraction in the head pull. Since pure nitrogen is to be removed at the head, X ^E = 1.

Thus v _m ι _n is 0.4.

In order to be able to remove 2 mol / s pure N at the top of the column, 0.8 mol / s N ₂ must be returned to the column.

In total, 2.8 mol / s of liquid N are fed to the column per unit of time (2.0 mol / s as a component of the liquid air and 0.8 mol / s as reflux at the top of the column), the total amount of which must be evaporated continuously.

In order to evaporate this amount of liquid, a corresponding amount of liquid must be returned as vapor into the column bottom. At the same time, 0.5 mol / s of 0 ₂ must be removed from the bottom of the column.

The oxygen withdrawn from the bottom of the column is heated to 20 ° C. at 1 bar, mixed with 1 mol / s propane (20 ° C., 1 bar), heated to 500 ° C. from 1 bar while maintaining the pressure and oxidized (catalytically and / or homogeneous).

The reaction gas mixture consisting of 1 mol / s propene and 1 mol / s water vapor is then retained while maintaining the

Pressure cooled to 20 ° C, the water vapor condenses completely due to the high boiling point difference and

1 mol / s pure propene (1 bar, 20 ° C) is obtained.

An overview of the above is shown in FIG. 1.

Use of a mixture of 80 vol.% 0 and 20 vol.% N

To obtain the mixture mentioned in the heading, 2.5 mol / s air (^ mixture of 0.5 mol / s 0 ₂ and 2.0 mol / s N ₂ ), the temperature of 20 ° C and a pressure from 1 bar. This is separated in a first rectification column.

The rectificative separation takes place as in a), but in pure nitrogen (top product) and a mixture of 80% by volume

0 and 20 vol .-% N ₂ (bottom product).

The boiling temperature of N ₂ at a pressure of 1 bar = -196 ° C (= top temperature). The boiling temperature of a mixture of 80 vol .-% 0 ₂ and 20 vol .-% N ₂ is at a pressure of 1 bar = -187 ° C (= sump temperature). T / EP99 / 05791

28

The mean boiling point along the rectification column is thus = -192 ° C.

According to VDI-Warmth Atlas, 5th edition 1998 (DC6 and DC7) the saturation vapor pressure of N ₂ at -192 ° C = 1.62 bar and the saturation vapor pressure of oxygen (0 ₂ ) = 0.37 bar. The relevant α is thus 1.62: 0.37 = 4.36. X ^F is the N ₂ molar fraction in the mixture to be separated (ie in air) fed to the rectification column. Ie, X ^F = 0.8 (= 80 vol .-% N ₂ in air).

X ^E is the N ₂ mole fraction in the head pull. Purer on the head

If nitrogen is to be removed, X ^E = 1.

This makes v _m i _n 0.37.

That is, in order to be able to remove 1.88 mol / s pure N ₂ at the top of the column, 0.70 mol / s N ₂ must be returned to the column in liquid form.

In total, 2.7 mol / s of liquid N _{2 are} supplied to the column per unit of time (2.0 mol / s as a component of the liquid air and 0.7 mol / s as return at the top of the column), of which the sum of each unit of time 1.88 mol / s N (the

Withdrawal quantity at the top) and 0.70 mol / s N (the return quantity) must be evaporated.

In order to evaporate this amount of liquid, a corresponding amount of liquid must be returned in vapor form to the bottom of the column. At the same time, a mixture of 0.5 mol / s O ₂ (80% by volume) and 0.12 mol / s N (20% by volume) must be taken continuously from the bottom of the column.

The mixture of 0.50 mol / s 0 ₂ and 0.12 mol / s N ₂ thus obtained is heated from 1 bar to 20 ° C. while maintaining the pressure and also at 1 mol at 1.0 mol / s containing propane mixed. The mixture is heated to 500 ° C. under pressure and catalytically and / or homogeneously oxydehydrogenated under these conditions (complete conversion, ideal selectivity).

The resulting reaction product mixture is cooled from 1 bar to 20 ° C. while maintaining the pressure, the water of reaction completely condensing out due to the high boiling point difference. The remaining mixture of 0.12 mol / s N ₂ and 1 mol / s propene, at 20 ° C. and 1 bar, is cooled (-171 ° C.) while maintaining pressure so that it boils liquid in the middle part of a second at 1 bar adiabatically operated rectification column (infinite number of stages) can be carried out to be separated overhead into pure nitrogen and as bottom liquid into pure propene. The boiling point of N at a pressure of 1 bar = -196 ° C (= head temperature).

The boiling temperature of propene at a pressure of 1 bar = -48 ° C (= bottom temperature).

The mean boiling point along the rectification column is thus = -122 ° C.

According to the VDI-Warmth Atlas, 5th edition 1998 (DC6 and DC7) the saturation vapor pressure of N ₂ at -122 ° C = 94.8 bar and the saturation vapor pressure of propene at -122 ° C = 0.00476 bar. This results in the relevant α in the second rectification column being 94.8: 0.00476 = 19900.

X ^F is the N mole fraction in the mixture to be separated which is fed to the rectification column. That is, X ^F = 0.12 / 1.12 = 0.107. X ^E is the N ₂ mole fraction in the head pull. Since pure nitrogen is to be removed at the head, X ^E = 1.

Thus v _{min is} calculated to be 0.00047. That is, in order to be able to remove 0.12 mol / s N ₂ at the top of the column, 0.000056 mol / s N liquid must be recycled into the column.

Overall, 0.120556 mol / s of liquid N _{2 are} fed to the column per unit of time (0.12 mol / s as part of the liquid mixture supplied and 0.000056 mol / s at the top of the column), the total amount of which must be evaporated continuously.

In order to evaporate this amount of liquid, a corresponding amount of liquid must be returned in vapor form to the bottom of the column. At the same time, 1.0 mol / s of propene are withdrawn from the column bottom and heated to 20 ° C. while maintaining pressure (1 bar).

An overview of the above is shown in FIG. 2.

Comparing the paths a), b) with each other, this shows less v _min , in case b (∑v _min = 0.37 + 0.00047 = 0.3705) in contrast to v _m i _n = 0.4 im Case a) the lower separation work to be performed when using only partially de-aired nitrogen. claims

1. A process for the preparation of acrolein and / or acrylic acid from propane, in which the propane is subjected to a partial homogeneous and / or heterogeneously catalyzed oxydehydrogenation with molecular oxygen to propene in a first reaction stage and then the propene formed in the first reaction stage containing, containing product gas mixture in at least one further reaction stage for the production of acrolein and / or acrylic acid by gas phase catalytic propene oxidation, characterized in that the reaction gas starting mixture of the first reaction stage requires the molecular oxygen as nitrogen required in the first reaction stage, which is different from cycle gas oxygen containing modified air, with the proviso that the nitrogen content of the modified air is smaller and the oxygen content of the modified air is greater than the corresponding contents of air, and that prior to reuse in the first reac tion stage formed, containing propene, for the production of acrolein and / or acrylic acid in the at least one further reaction stage, at least part of the molecular nitrogen contained in the product gas mixture is separated from the product gas mixture. . A method according to claim 1, characterized in that before further use of product gas mixture formed in the first reaction stage and containing propene for the production of acrolein and / or acrylic acid in at least one further reaction stage, the total amount of nitrogen contained in the product gas mixture is separated off. . A method according to claim 1 or 2, characterized in that the nitrogen separation is heard by fractional distillation. . A method according to claim 3, characterized in that with the separation of the nitrogen, all components of the product gas mixture which have a lower boiling point than propene are separated off.